SELECTIVE OXIDATIVE DEHYDROGENATION CATALYSTS AND METHOD OF OXIDATIVE DEHYDROGENATION

CROSS-REFERENCE TO RELATED .APPLICATION [0001] This application claims the benefit of priority to U.S. Provisional Application Number 62/962,091 filed on January' 16, 2020, and entitled “Selective Oxidation Catalysts and Systems”, the contents of which are hereby incorporated by reference.

TECHNICAL FIELD

[0002] The present disclosure relates generally to selective oxidation (SO) catalysis and systems. More specifically, catalysts and systems are provided for oxidative dehydrogenation (ODH).

BACKGROUND

[0003] Olefins like ethylene, propylene, and butylene, are basic building blocks for a variety of commercially valuable polymers. Since naturally occurring sources of olefins do not exist in commercial quantities, polymer producers rely on methods for converting the more abundant lower alkanes into olefins. The method of choice for today ’s commercial scale producers is steam cracking, a highly endothermic process where steam-diluted alkanes are subjected very briefly to a temperature of at least 800 °C. The fuel demand to produce the required temperatures and the need for equipment that can withstand that temperature add significantly to the overall cost. In addition, the high temperature promotes the formation of coke, which accumulates within the system, resulting in the need for cosily periodic reactor shutdown for maintenance atsd coke removal.

[0004] Selective oxidation processes, such as oxidative dehydrogenation (ODH), are an alternative to steam cracking that are exothermic and produce little or no coke. In ODH, a lower alkane, such as ethane, is mixed with oxygen in the presence of a catalyst and optionally an inert diluent, such as carbon dioxide or nitrogen or steam, which may be performed at temperatures as low' as 300 °C, to produce the corresponding a!kene. Various other oxidation products may be produced in this process, including carbon dioxide and acetic acid, among others. ODH suffers from lower conversion rates when compared to steam cracking, a fact that when combined with lower selectivity and the risk of thermal explosion due to mixing of a hydrocarbon with oxygen, may have prevented ODH from achieving widespread commercial implementation

BRIEF DESCRIPTION OF DRAWINGS

[0005] Figures 1 A and IB are block diagrams of an example selective oxidation system for the selective oxidation of light hydrocarbons.

[0006] Figure 2 is a graphic depiction of an example chemical complex.

[0007] Figure 3 is a graphic depiction of an example chemical complex.

[0008] Figure 4 is a schematic of an example experime tal reactor unit.

[0009] Figure 5 is a schematic diagram of an example apparatus and process flow for providing mixing of reagents.

[0010] Figure 6 is a schematic diagram of an example reactor in which oxidative dehydrogenation can take place

[0011 ] Figure 7 is a schematic diagram of an example reactor in which the oxidative dehydrogenation can take place

[0012] Figures 8, 9, and 10 are schematic diagrams of various configurations of three example fixed bed catalysts that may be used to scavenge oxygen from the product stream of an oxidative dehydrogenation (ODH) reactor.

[0013] Figure 11 shows a reaction profile (Dynamics) of ethylene formation as a function of time at different temperatures after a gas flow switch (air to ethane) for a catal st with Ti(½ as support.

[0014] Figure 12 shows the reaction profile (Dynamics) of CO formation as a function of time at different temperatures after a gas flow switch (air to ethane) for a catalyst with Ti 0 ₂ as a support.

[0015] Figure 13 shows the selectivity of ethylene formation as a function of time at different temperatures after a gas flow switch (air to ethane) for a catalyst with Ti0 ₂ as a support.

[0016] Figure 14 shows the reaction profile (Dynamics) of 0 ₂ removal from the model gas mixture at different temperatures by a pre -reduced catalyst w ith Ti0 ₂ as support.

[0017] Figures 15 and 16 show the reaction profiles (Dynamics) of C0 ₂ and CO formation (respectively) after feeding the model gas mixture at different temperatures by a pre-reduced catal st with Ti0 ₂ as support. [0018] Figure 17 is a schematic representation of an example gas mixer.

[0019] Figure 18 is a schematic representation of an example gas mixer, according to a pipe-in-pipe embodiment.

[0020] Figure 19 is a cross-sectional view of the pipe-in-pipe embodiment of Figure ! 8, taken through the line X-X in Figure i 8.

[0021] Figure 20 is a schematic representation of an example gas mixer, according to a standing pipe embodiment.

[0022] Figure 21 is a schematic representation of an example gas mixer.

[0023] Figure 22 is a schematic representation of an example twinned gas mixer unit.

[0024] Figure 23 is a plot depicting a long-term microreactor unit (MRU) experimental n with dimethyl disulfide (DMDS) injections.

[0025] Figure 24 is a schematic diagram of a conventional C ₂ splitter (cryogenic distillation tower)

[0026] Figure 25 Is a schematic diagram of an example C ₂ splitter integrated with an oxidative dehydrogenation unit at the overhead stream (ethylene product stream). [0027] Figure 26 is a schematic diagram of an example oxidative dehydrogenation reactor integrated with the bottom product stream (ethane) from a C _? splitter.

[0028] Figure 27 is a plot showing selectivity of ethane formation by an ODH process as a function of temperature.

[0029] Figure 28 is a schematic diagram of an example integration of an ODH unit within the C splitter.

[0030] Figure 29 is a schematic diagram of an example ODH unit that is integrated into the feed front the cracker to the C spliter.

[0031] Figure 30 is a schematic diagram of an example ODH unit that is integrated downstream of the acetylene hydrogenation unit.

[0032] Figure 31 is a schematic diagram of an example ODH reactor containing three fixed beds of catalyst.

[0033] Figure 32 is a schematic diagram of an example apparatus that can be used to mix oxygen and hydrocarbons.

[0034] Figure 33 is a schematic diagram of an example fixed bed reactor unit that can be used for ODH.

[0035] Figure 34 is a plot of the temperature at which there is a 25% conversion of ethane to ethylene against the amount of 30% H O for a catalyst.

[0036] Figure 35 is a plot of the selectivity for conversion to ethylene at the temperature at which there is a 25% conversion to ethylene against the volume of 30% H O for a cataly st.

[0037] Figure 36 is a plot of an experimental run showing product selectivity versus steam added for an ODH reaction.

[0038] Figure 37 is a schematic diagram of an example reactor system with two fixed bed ODH reactors [0039] Figures 38, 39, 40, 41, 42, and 43 show tables that include results of experimental runs of ODH reactions.

[0040] Figure 44 is a block diagram of an example reactor system with acetic acid added to the feed.

[0041 [ Figure 45 is a block diagram of an example reactor system with acetic acid added to the feed and acetic acid recycled from after a separation step.

[0042] Figure 46 is a plot showing the results from simulations with and without acetic acid added to the feed.

[0043] Figure 47 is a block diagram of an example ODH coproduction system including an ODH reactor system having two ODH reactors in series

[0044] Figure 48 is a flow chart of an example method of coproduction in an ODH system that converts ethane to ethylene.

[0045] Figure 49 is a block diagram of an example ODH reactor system.

[0046] Figure 50 is a schematic diagram of an example circulating fluidized bed (CFB) reactor.

[0047] Figures 51a and 51b are plots showing the conversion and selectivity (respectively) performance of a cataly st over time.

[0048] Figures 52a and 52b are plots showing the conversion and selectivity (respectively) performance of a cataly st over time.

[0049] Figures 53a and 53b are plots showing the conversion and selectivity (respectively) performance of a catalyst over time. [005 ⁽ 1] Figures 54a and 54b are plots showing the conversion and selectivity (respectively) performance of a catalyst over time.

[0051] Figure 55 is a plot showing time dependence of the selectivity of ethylene formation in the presence of a catal st at different How rates.

[0052] Figure 56 is a plot showing dependence of ethane conversion on the amount of ethane supplied into the reactor at different rates in the presence of a catalyst.

[0Q53] Figures 57a and 57b are plots showing the time dependence of ethane conversion and 0 residual content (57a) and selectivity of ethylene formation (57b) in the presence of a catalyst.

[0054] Figures 58a and 58b are plots showing the time dependence of ethane conversion and 0; residual content (58a) and selectivity of ethylene formation (58b) in the presence of a catalyst.

[0055] Figure 59 is a schematic diagram of an example gas mixer.

[0056] Figure 60A is a cross-sectional side view of a section of the gas mixer of Figure 59 with random packing.

[0057] Figure 6QB is a cross-sectional side view of a section of the gas mixer of Figure 59 with structured packing.

[0058] Figure 60C is a cross-sectional side view of a section of the gas mixer of Figure 59 with an impeller.

[0059] Figure 61 is a schematic diagram of an example gas mixer with supply nozzles at a top end of the gas mixer.

[0060] Figure 62A is a cross-sectional view of a postion of an example gas mixer with supply nozzles that comprise circular concentric spargers.

[0061 ] Figure 62B is a cross-sectional top view of the gas mixer of Figure 62A.

[0062] Figure 63 A is a schematic diagram of an example oxygen separation module in which hydrocarbons are directed to a permeate side of the oxygen separation module.

[0063] Figure 63B is a cross-sectional top view of the oxygen separation module of Figure 63 A.

[0064] Figure 63C is a schematic diagram of an example oxygen separation module in which hydrocarbons are directed to a retentate side of the oxygen separation module.

[0065] Figure 63D is a schematic diagram of an example oxygen separation module in which hydrocarbons can be disected to either or both of the permeate side and the retentate side of the oxygen separation module.

[0066] Figure 64 is a flow diagram of an example of a system that can be used to convert an alkane to an alkene.

[0067] Figure 65 is a flow diagram of an example of a system that can be used to separate an oxy genate from a stream.

[0068] Figure 66 is a flow diagram of an example of a system comprising a separation vessel.

[0069] Figure 67 is a flow diagram of an example of a system comprising an oxygen remover.

[0070] Figure 68 is a flow diagram of an example of a system comprising an amine tower

[0071] Figure 69 is allow diagram of an example of a system comprising a polymerization reactor.

[0072] Figure 70 is a schematic diagram of an example ODH reactor and associated downstream piping showing where fouling may develop. [0073] Figure 71 is a side view of a pipe-in-pipe arrangement.

[0074] Figure 72 is a side view of an outlet pipe with an instreani atomizer.

[0075] Figure 73A is a side view of an outlet pipe with inner surface jets.

[0076] Figure 73B is a cross-sectional view of the outlet pipe with inner surface jets of Figure 73 A along dotted line X-X.

[0077] Figure 74A is a side view' of an outlet pipe with a plurality of holes

[0078] Figure 74B is a cross-sectional view' of the outlet pipe with a plurality of holes of Figure 74A along doited line X-X

[0079] Figure 75 is a photo of a catalyst sample prior to the addition of the oxalic acid.

[0080] Figure 76 is a photo of a solubilized catalyst and a dispersed support.

[0081 [ Figure 77 is a photo of a stainless steel rod contaminated with catalyst.

[0082] Figure 78 is a photo of the stainless steel rod of Figure 77 cleaned of catalyst by treatment with oxalic acid.

[0083] Figure 79 is a block flow diagram of an example fixed bed ODH operation with an air separation unit.

[0084] Figure 80 is a block flow' diagram of an example sw ing bed ODH operation with an air separation unit.

[00851 Figure 81 is a block flow' diagram of an example fluidized bed ODH operation with an air se aration unit

[0086] Figure 82 is a block flow' diagram of an example fluidized bed ODH operation with an air separation unit and two fluidized bed catalyst regenerators.

[0087] Figure 83 is a block flow' diagram of an example swing bed ODH operation w ith an air separation unit.

[0088] Figure 84 is a flow diagram of an example of a system that can be used to convert an alkane to an a!kene.

[0089] Figure 85 is a flow diagram of an example of a system comprising a separation tower.

[0090] Figure 86 is a flow' diagram of an example of a system comprising an oxygen remover.

[0091] Figure 87 is a flow diagram of an example of a system comprising an amine tower

[0092] Figure 88 is a flow' diagram of an example of a system comprising a polymerization reactor.

[0093] Figure 89 is a flow diagram of an example of a system that can be used to convert; an alkane to an alkene.

[0094] Figures 90, 91, 92, 93, 94, 95, 96, and 97 are schematic diagrams of example ODH reactor systems, each having three ODH reactors operationally disposed in series.

[0095] Figure 98 is a block flow diagram of an example method for ODH.

[0096] Figure 99 is a block flow diagram of a portion of an example chemical complex.

[0097] Figure 100 is a block flow diagram of a portion of an example chemical complex

]0098] Figure 101 is a block flow' diagram of a portion of an example chemical complex.

[0099] Figure 102 is a block flow diagram of a portion of an example chemical complex

[0100] Figure 103 is a block flow diagram of a portion of an example chemical complex.

[0101 ] Figure 104 is a block diagram illustrating an example chemical process. [0102] Figure i 05 is a block diagram of an example reactor system including a reactor with a catalyst for the conversion of C(¾ into acetic acid or carbon monoxide (CO).

[0103] Figure 106 is a block flow diagram of an example method of processing CO (e.g., conversion of CO ) with a catalyst in a reactor system.

[0104] Figure 107 is a block flow diagram of an example reactor system.

[0105] Figure 108 is a block flow diagram of at! example of integrated sy stems.

[0106] Figure 109 is a block flow diagram of an example chemical complex

[0107] Figure 110 is a schematic diagram of an example reactor system.

[0108] Figure 111 is an X-ray powder diffraction (XRD) plot along with peak locations for various example catalysts

[0109] Figure 112 is an XRD plot along with peak locations for various example catalysts

[0110] Figure 113 is an XRD plot along with peak locations for various example catalysts.

[0111 ] Figure 114 is an XRD plot along with peak locations for various example catalysts.

[0112] Figure 115 is an XRD plot along with peak locations for an example catalyst.

[0113] Figure 116 shows scanning electron microscopy (SEM) images of an example catalyst at different scales.

[0114] Figure 117 shows SEM images of various example catalysts at different scales.

[0115] Figure 118 shows SEM images of various example catal sts at different scales.

[0116] Figure i 19 shows SEM images of various example catalysis at different scales.

[0117] Figure 120 is a plot showing pore volume versus pore width for various example catalysts.

[0118] Figure 12 i is a plot showing pore volume versus pore width for various example catalysts

[0119] Figure 122 is a plot showing pore volume versus pore width for various example catalysts.

[0120] Figure 123 is a plot showing pore area versus pore width for various example catalysts

[0121] Figure 124 is a Fourier-transform infrared spectroscopy (FTIR) plot for various example catal sts.

[01221 Figure 125 is a set of various schematics depicting experimental setups for an example catalyst and various example catalyst materials.

[0123] Figure 126 is a schematic diagram of at! experimental setup of an example nucroreactor unit (MRU)

[0124] Figure 127 is a schematic diagram of an experimental setup of an example MRU.

[0125] Figure 128 shows MRU raw data for various example catalyst materials.

[0126] Figure 129 shows MRU raw data for various example catalyst materials.

[0127] Figure 130 is an SEM image of an example catalyst

[0128] Figure 131 is an SEM image of an example catalyst material.

[0129] Figure 132 is an SEM image of an example catalyst material.

[0130] Figure 133 is an SEM image of an example catalyst material.

[0131] Figure 134 is an SEM image of an example catalyst material.

[0132] Figure 135 is an SEM image of an example catalyst material.

[0133] Figure 136 is an SEM image of beryllium oxide

[0134] Figure 137 is an SEM image of an example catalyst

[0135] Figure 138 is an SEM image of an example catalyst material. [0136] Figure 139 is an SEM image of an example catalyst material.

[0137] Figure 140 is anFTIR plot for various example catalyst materials.

[0138] Figure 141 is an XRD plot for various example catalyst materials.

[0139] Figure 142 is an XRD plot along with peak locations lor various example catalyst materials.

DETAILED DESCRIPTION

[0140] Selective oxidation (SO) is generally used in ODH reactions to form ethylene, or other alpha- olefins, from ethane. Embodiments described herein provide a catalyst system for a selective oxidation reaction. [0141] The Cattily st System

[0142] Provided herein is an oxidative dehydrogenation catalyst material that includes molybdenum (Mo), vanadium (V), oxygen (O), and at least one element selected from iron (Fe), aluminum (Alt, and beiyllium (Be). [0143] As used herein, the term “catalyst material” refers to a material that includes an active catalyst that can promote the oxidative dehydrogenation of ethane to ethylene, for example, on a support. The catalyst material can be a plurality of particles or a formed catalyst material. Non-limiting examples of formed catalyst materials include extruded catalyst materials, pressed catalyst materials, and cast catalyst materials. Non- limiting examples of pressed and cast catalyst materials includes pellets — such as tablets, ovals, and spherical particles.

[0144] As used herein, the tern “catalyst” generally refers to the active catalyst portion of a catalyst material. The catalyst is generally processed in further steps to form a catalyst material. The catalyst material may also be processed in further steps to form a final catalyst material.

[0145] In some embodiments, the catalyst tnaterial does not include niobium, tellurium, or both

[0146] Catalyst Materials Including Molybdenum, Vanadium, Oxygen, and Iron [0147] In some embodiments, the catalyst tnaterial includes molybdenum, vanadium, oxygen, and iron. The molar ratio of molybdenum to vanadium its the catalyst material can be from 1:0.1 to 1:0.5. The molar ratio of molybdenum to iron in the catalyst material can be from 1 :0.25 to 1:5.5. Further, oxygen can be present in the catalyst tnaterial at least in amount to satisfy the valency of any present metal oxides

[0148] Unless stated otherwise, the molar ratios of molybdenum, vanadium, iron, aluminum, beryllium, and optionally other elements in the catalyst materials described herein can be determined by employing inductively coupled plasma mass spectrometry' ICP-MS. For instance, the molar ratio of molybdenum, vanadium, and iron in the catalyst materials described herein can be determined by a method that includes (1) first digesting (e.g , fully dissolving) the catalyst material (e.g., 10 rug of catalyst material) in (a) a solution that includes 10 wt. % to 15 wt. % NaOCl or (b) a solution that includes 25 mol/L NaOH; (2) heating, agitating, or heating and agitating the solution including to the catalyst material to facilitate digestion of the catalyst material, and (3) diluting the solution as necessary to prepare a sample suitable for ICP-MS analysis.

[0149] In some embodiments, the molar ratio of molybdenum to vanadium in the catalyst material is from 1:0.30 to 1:0.45. For example, the molar ratio of molybdenum to vanadium can be from 1:0.30 to 1:0.35 or from 1:0.35 to 1:0.45. In some embodiments, the molar ratio of molybdenum to vanadium is from 1:0.37 to 1:0.41. [0150] In some embodiments, at least a portion of the molybdenum and vanadium in the catalyst material can be present as a mixed metal oxide having the empirical formula:

MO _l Vo25-05oO _d wherein d is a number lo satisfy the valence of the oxide.

[0151] In some embodiments, the molar ratio of molybdenum to iron in the catalyst material is from 1:3 to 1:5.5. For example, the molar ratio of molybdenum to iron can be from 1 :4.0 to 1:5.0 or from 1 :4.25 to 1 :4.75.

In some embodiments, the molar ratio of molybdenum to iron is from 1 :4 45 to 1:4.55. [0152] In some embodiments, the molar ratio of molybdenum to vanadium in the catalyst material is from 1:0.30 to 1:0.45 and the molar ratio of molybdenum to iron is from 1:4.25 to 1:4.75.

[0153] In some embodiments, the molar ratio of molybdenum to iron in the catalyst material is from 1:0.25 to 1:1.0. For example, the molar ratio of molybdenum to iron can be from 1:0.25 to 1:0.75. In some embodiments, the molar ratio of molybdenum to iron is from 1:0.35 to 1:0.65. For example, the molar ratio of molybdenum to iron can be from 1:0.35 to 1:045 or from 1:0.55 to 1:0.65. In some embodiments, the molar ratio of molybdenum to iron is about 1 :0.4. In some embodiments, the molar ratio of molybdenum lo iron is about 1:0.6.

[0154] in some embodiments, the molar ratio of molybdenum to vanadium in the catalyst material is from 1:0.30 to 1 :0 45 and the molar ratio of molybdenum to iron is from 1 :0.35 to 1:0.65.

[0155] in some embodiments, the molar ratio of molybdenum to iron in the catalyst material is from 1:1.3 to 1:2.2. For example, the molar ratio of molybdenum to iron can be from 1: 1.6 to 1:2.0. In some embodiments, the molar ratio of molybdenum to iron is from 1 : 1.80 to 1 : 1.90.

[0156] In some embodiments, the molar ratio of molybdenum to vanadium in the catal st material is from 1:0.30 to 1:0.45 and the molar ratio of molybdenum to iron is from 1:1.80 to 1:1.90.

[0157] In some embodiments, the molar ratio of molybdenum to iron in the catalyst material is from 1:2.0 to 1:2.5. For example, the molar ratio of molybdenum to iron can be from 1 :2 3 to 1:2.7 In some embodiments, the molar ratio of molybdenum to iron is from 1 : 1.80 to 1:1.90.

[9158] In some embodiments, the molar ratio of molybdenum to vanadium in the catalyst material is from 1 :0 30 to 1:0.45 and the molar ratio of molybdenum to iron is from 1:2.45 to 1 :26.

[0159] In some embodiments, the catalyst material has a 35% conversion temperature from about 300 °C to about 400 °C. For example, the catalyst material can have a 35% conversion temperature from about 300 °C to about 350 °C. In some embodiments, the catalyst material has a 35% conversion temperature from about 315 °C to about 335 °C. As used herein, conversion temperature is the temperature at which the stated percentage of ethane and oxygen are converted to final products.

[9169] As used in this disclosure, the phrase “35% conversion temperature” refers to the temperature at which 35% of ethane in a gas stream is converted to a product other than ethane. The 35% conversion temperature of an oxidative dehydrogenation catalyst can be determined by using a microreactor unit (MRU). In a microreactor unit, the 35% conversion temperature of a catalyst can be determined by passing a feed gas over a catalyst bed in a reactor tube. The MRU reactor tube has an outer diameter of about 0 5 inches and an internal diameter of about 0.4 inches and length of about 15 inches. For example, the reactor tube can be stainless-steel SWAGELOK ^® Tubing with a wall thickness of about 0.049 inches.

[0161] The feed gas can include ethane and oxygen having a molar ratio of 70:30 to 90: 10. For example, the feed gas can include ethane and oxygen having a molar ratio of 82:18. Alternatively, the feed gas can include ethane, oxygen, and nitrogen. The molar ratio of ethane to oxygen to nitrogen can be 18 : 18 :64 to 54:18:28 For example, the molar ratio of ethane to oxygen to nitrogen can be 36:18:46 or 35:17.5:47.5. The flow rate of the feed gas can be about 70 standard cubic centimeters per minute (seem) to about 80 seem. For example, the How rate of the feed gas can be about 75 seem (e.g., 74.6 seem). The catalyst bed consists of the oxidative dehydrogenation catalyst and a filler, such as sand, I :() 5 to 1 :3 volume ratio, ith the total weight for the oxidative dehydrogenation catalyst being 1 .96 to 200 g. Any remaining space in the reactor tube (e.g., below or above the catalyst bed) is packed with an additional filler, such as quartz sand. The 35% conversion temperature is determined at a weight hourly space velocity (WHSV) of 2.90 h ¹ , with the WHSV based on the weight of Mo VO* or MoVFeO _x in the sample, and a gas hourly space velocity (GHSV) of about 2,000 to 3,000 hX In cases where catalyst materials include an alumina — such as alumina hydroxide oxide (e.g., hoehmite) — the original weight percentage ofMoVO* or MoVFeO _x produced from the by -weight combinations of these materials with alumina are assumed to be unchanged from post-mixing workups involving heat treatments. Therefore, the feed gas flow can be adjusted to a WHSV target of 2.90 lr ^! based on the original weight percentage of (i) MoVO _x in a catalyst material that includes molybdenum, vanadium, oxygen, and aluminum (e.g , MoVAlO _x ) or (it) MoVFeO _x in a catalyst material that includes molybdenum, vanadium, oxygen, iron, and aluminum (e.g., MoVFeAIO _x ) at the time of mixing. Typically, the inlet pressure is in the range of about 1 pound per square inch gauge (psig) to about 2.5 psig and the outlet pressure is in the range of about 0 psig to about 0.5 psig. The gas feed exiting the catalyst bed is analyzed by gas chromatography to determine the percent of various hydrocarbons (e.g., ethane and ethylene) and, optionally other gases such as O2, CO2, and CO. Conversion of the feed gas is calculated as a mass flow rate change of ethane in the product compared to feed ethane mass flow rate using the following formula: wherein C is the percent (molar percent) of feed gas that has been converted from ethane to another product (i.e., ethane conversion) and X is the molar concentration of the corresponding compound in the gaseous effluent exiting the reactor. The ethane conversion is then ploted as a function of temperatures to acquire a linear algebraic equation. The linear equation for ethane conversion is solved to determine the temperature in which the ethane conversion is 35% (i.e. the 35% conversion temperature). Not taken into account for calculating the 35% conversion of ethane temperature or selectivity to ethylene, were reaction the products exiting the reactor in an aqueous stream such as, but not limited to, acetic acid, maleic acid, propionic acid, ethanol, and acetaldehyde.

[0162] In some embodiments, the catalyst material has a selectivity to ethylene from about 65% to 99%. For example, the catalyst material can have a selectivity to ethylene from about 75% to 95%. In some embodiments, the catalyst material has a selectivity to ethylene from about 77% to about 85%.

[0163] As used in this disclosure, the phrase “selectivity to ethylene” refers to the percentage on a molar basis of converted or reacted ethane that forms ethylene. An oxidative dehydrogenation catalyst’s selectivity to ethylene can be determined using an MRU as discussed above. An oxidative dehydrogenation catalyst’s selectivity to ethylene can be determined using to the following equation: wherein _Ethyi is the selectivity to ethylene, andX is the molar concentration of the corresponding compound in the gaseous effluent exiting the reactor. Notably, the selectivity to ethylene is determined at the 35% conversion temperature, unless otherwise indicated. As such, after the 35% conversion temperature is determined, the above equation for selectivity is solved using the corresponding values for XCX _B , and

Xco at the 35% conversion temperature.

[0164] In some embodiments, the catalyst material has a selectivity to acetic acid of less than 15 wt. % in a process for the oxidative dehydrogenation of ethane to ethylene. For example, the catalyst material can have a selectivity to acetic acid of about 1 wt. % to about 15 wt. %, about 3 wt. % to about 12 wt. %, or about 7 wt. % to about 12 wt. % in a process for the oxidative dehydrogenation of ethane to ethylene. In sotne embodiments, the catalyst material has a selectivity to acetic acid of about 7 wt. %, 8 wt. %, 9 wt %, 10 wt %, 11 wt %, 12 wt % or about 13 wt. % in a process for the oxidative dehydrogenation of ethane to ethylene 0165 In some embodiments, at least a portion of the iron present in the catalyst material can be present as iron (IP) (i.e., Fe ³⁺ ), The presence of iron (III) in the catalyst material can be detected by X-ray photoelectron spectroscopy (XPS).

[0166] In some embodiments, at least a portion of the iron in the catalyst material can be present as amorphous iron. The presence of amorphous iron in the catalyst material can be detected by XPS.

[0167] In some embodiments, at least a portion of the iron present in the catalyst material is present as an iron oxide, an iron oxide hydroxide, or a combination thereof. The presence of iron oxides and iron oxide hydroxides can be determined by X-ray powder diffraction (XRD)

[0168] When at least a portion of the iron in the catalyst material is present as an iron oxide, the iron oxide cars include an iron oxide selected from hematite (a-FWOs), maghemite (y-Fe2<¾), magnetite (FesCti), or a combination thereof in some embodiments, at least a portion of the iron in the catalyst material is present as hematite.

[0169] When at least a portion of the iron i the catalyst material is present as an iron oxide hydroxide, the iron oxide hydroxide can include an iron oxide hydroxide selected from a goethite (a-FeO(OH)), an akaganeite (P-FeO(OH)), a lepidocrocite (y-FeO(OH)), or a combination thereof. In some embodiments, at least a portion of the iron in the catalyst material is present as a goethite.

[0179] In some embodiments, at least a portion of the iron in the catalyst material is present as a goethite and at least a portion of the iron in the catalyst material is present as hematite. 0171 ] The catalyst materials provided above that include molybdenum, vanadium, oxygen, and iron can prepared by a method that includes preparing an aqueous mixture including (i) a catalyst that includes molybdenum, vanadium, and oxygen; (ii) an iron compound, and (iii) a water. The method further includes removing a substantial amount of the water from the mixture to provide a precatalyst material. Subsequently, the precatalyst material is heated to provide the catalyst material.

[0172] In some embodiments, the method further includes preparing the catalyst including molybdenum, vanadium, and oxygen.

[0173] In some embodiments, the molar ratio of molybdenum to vanadium in the provided catalyst is from 1:0.25 to 1:0.6. For example, the molar ratio of molybdenum to vanadium in the catalyst can be from 1:0.35 to 1:0.55. In some embodiments, the molar ratio of moly bdenum to vanadium in the provided catalyst is from 1:0.40 to 1:0.49

[0174] In some embodiments, the provided catalyst includes a mixed metal oxide having the empirical formula:

MOiVo _{25-0 5} oOij wherein d is a number to satisfy the valence of the oxide

[0175] In some embodiments, the iron compound in the aqueous mixture includes an iron (III) compound. In some embodiments, the iron compound in the mixture of the catalyst, the iron compound, and the water includes an iron compound selected from an iron oxide, an iron oxide hydroxide, or a combination thereof. The iron oxide be an iron oxide selected from hematite (a-FeaCb), maghemite (y-Fe C> ), magnetite (FeaCb), or a combination thereof. The iron oxide hydroxide can be an iron oxide hydroxide selected from a goethite, an akageneite, a lepidocrocite, or a combination thereof. In some embodiments, the iron compound in the mixture of the catalyst, the iron compound, and the water includes hematite. In some embodiments, the iron compound its the mixture of the catalyst, the iron compound, and the water includes goethite. in some embodiments, the iron compound in the snixtsue of the catalyst, the isosi compound, and the water includes goethite and hematite. [0176] Removing a substantial amount of water from the aqueous mixture including the catalyst, the iron compound, and the water to provide a precatalyst material can include removing from about 50 wt. % to about 99 wt % of the water. For example, about 50 wt. % to about 75 wt. % or about 75 wt. % to about 99 wt. % of the water can be removed to provide the precatalyst material. For example, enough water can be removed from the aqueous mixture such that the provided precatalyst material has a paste-iike consistency.

[0177] In sense embodiments, removing a substantial amount of the water from the mixture of the catalyst, the iron compound, and the water to provide a precata!yst material includes heating the aqueous mixture at a temperature from about 50 °C to about 100 °C. For example, the aqueous mixture can be heated at a temperature of about 80 ^C 'C

[0178] in some embodiments, heating the precatalyst material to provide the catalyst material includes heating the precatalyst material at a temperature from about 300 °C to about 500 °C. For example, heating the precatalyst material to provide the catalyst material can include heating the precatalyst material at a temperature from about 350 °C to about 450 °C. In some embodiments, heating the precatalyst material to provide the catalyst material includes heating the precatalyst material at a temperature of about 350 °C to about 375 °C, about 375 °C to about 400 °C, about 400 °C to about 425 °C, or about 425 °C to about 450 °C.

[0179] Further, heating the precatalyst material to provide the catalyst material can include heating the precatalyst material in the presence of air, an oxidizing atmosphere, an inert atmosphere, or a combination thereof. In some embodiments, heating the precatalyst material to provide the catalyst material includes heating the precatalyst material the presence of air.

[018ft] In some embodiments, after the catalyst material is prepared, the catalyst material can be treated with a gas that includes ethane, oxygen, ethylene, or a combination thereof. For example, the catalyst material can be treated with a gas that includes ethane, oxygen, ethylene, or a combination thereof in an oxidative dehydrogenation reactor.

[0181] In some embodiments, the catalyst material is treated with a gas including ethane, oxygen, ethylene, or a combination thereof at an elevated temperature. For example, the catalyst material can be treated with a gas including ethane, oxygen, eth lene, or a combination thereof at a temperature from about 100 °C to about 500 °C, about 200 °C to about 450 °C, or about 300 °C to about 400 °C. In some embodiments, the catalyst material is treated with a gas including ethane, oxygen, ethylene, or a combination thereof for about 6 hours to about 144 hours or about 18 hours to about 72 hours. In some embodiments, the catalyst material is treated with the gas including ethane, oxygen, ethylene, or combination thereof in an oxidative dehydrogenation reactor. [0182] Catalyst Materials Including Molybdenum, Vanadium, Oxygen, and Aluminum [0183] In some embodiments, the catalyst material includes molybdenum, vanadium, oxygen, and aluminum. The molar ratio of molybdenum to vanadium can be from 1:0.1 to 1:0.5. The molar ratio of molybdenum to aluminum can be from 1 : 1.5 to 1:6.5. Further, oxy gen can be present at least in an amount to satisfy the valency of any present metal oxides.

[0184] In some embodiments, the molar ratio of molybdenum to vanadium in the catalyst material is from i :0.25 to 1:0.50. In some embodiments, the molar ratio of molybdenum to vanadium in the catalyst material is from 1:0.30 to 1:0.45 For example, the molar ratio of molybdenum to vanadium can be from 1:0.30 to i :0 35 or from 1:0.35 to 1:0.45. In some embodiments, the molar ratio of molybdenum to vanadium is from 1 :0 37 to 1:0.41.

[0185] In some embodiments, the molar ratio of molybdenum to vanadium in the catalyst material is from 1:2.5 to 1:6.0.

[0186] In some embodiments, the molar ratio of molybdenum to vanadium in the catalyst material is from 1:5.5 to 1:6.5, as determined by energy -dispersive X-ray spectroscopy (EDS). In some embodiments, the molar ratio of molybdenum to vanadium in the catalyst material is 1 ^' roin 1:2.5 to 1:3.5, as determined by EDS.

[0187] In some embodiments, at least a portion of the molybdenum and vanadium in the catalyst material can be present as a mixed metal oxide having the empirical formula: wherein d is a number to satisfy the valence of the oxide

[0188] In some embodiments, the molar ratio of molybdenum to aluminum is from i :3 0 to 1 :6.5. For example, the molar ratio of molybdenum to aluminum can be from 1:3.25 to 1:5.5. In some embodiments, the molar ratio of molybdenum to vanadium is from 1:3.5 to 1 :4.1. For example, the molar ratio of molybdenum to aluminum can be from 1:3.9 to 1:4.3. In some embodiments, the molar ratio of molybdenum to vanadium is from 1:4.8 to 5.2. For example, the molar ratio of molybdenum to aluminum can be from 1:4.95 to 1:5.05.

[0189] In some embodiments, the molar ratio of molybdenum to vanadium in the catalyst material is front 1:0.30 to 1:0 45 and the molar ratio of molybdenum to aluminum is from 1:3.8 to 1:4 3. In some embodiments, the molar ratio of molybdenum to vanadium in the catalyst material is from 1 :0.30 to 1 :0.45 and the molar ratio of molybdenum to aluminum is from 1:4.8 to 5.2.

[0199] in some embodiments, the molar ratio of molybdenum to aluminum is from i : 1 5 to 1 :3 5. For example, the molar ratio of molybdenum to aluminum can be from 1:1.5 to 1 :2.5 or from 2.5 to 3.5. In some embodiments, the molar ratio of molybdenum to aluminum is from 1:2.0 to 1:2.2 or from 1:2.9 to 1:3.1.

[9191] In some embodiments, the molar ratio of molybdenum to vanadium in the catalyst material is from 1:0.30 to 1:0.45 and the molar ratio of molybdenum to aluminum is 1 ^' roin 1:1.5 to 1:2.5. In some embodiments, the molar ratio of molybdenum to vanadium in the catalyst material is from 1:0.30 to 1:0.45 and the molar ratio of molybdenum to aluminum is from 1:2.5 to 1:3.5.

[9192] In some embodiments, the catalyst material has a 35% conversion temperature from about 300 °C to about 400 °C. For example, the catalyst material can have a 35% conversion temperature from about 300 °C to about 350 °C. In some embodiments, the catalyst material has a 35% conversion temperature from about 315 °C to about 335 °C. [0193] In some embodiments, the catalyst material has a selectivity to ethylene from about 65% to 99%. For example, the catalyst material can have a selectivity to ethylene from about 75% to 95%. In some embodiments, the catalyst material has a selectivity to ethylene from about 77% to about 85%.

[0194] In some embodiments, the catalyst material has a selectivity to acetic acid of less than 15 wt. % in a process for the oxidative dehydrogenation of ethane. For example, the catalyst material can have a selectivity to acetic acid of about 1 wt. % to about 15 wt %, about 3 wt. % to about 12 wt. %, or about 7 wt. % to about 12 wt. % in a process for the oxidative dehydrogenation of ethane. In some embodiments, the catalyst material has a selectivity to acetic acid of about 7 wt. %, 8 wt. %, 9 wt. %, 10 wt. %, 11 wt. %, 12 wt. % or about 13 wt. % in a process for the oxidative dehydrogenation of ethane.

[0195] In some embodiments, at least a portion of the aluminum in the catalyst material is present as an aluminum oxide. In some embodiments, the aluminum oxide is an aluminum oxide hydroxide. The aluminum oxide hydroxide can include a gibbsite, a bayerite, a boehmite, or a combination thereof. In some embodiments, at least a portion of the aluminum in the catalyst material is present as a boehmite. In some embodiments, at least a portion of the aluminum in the catalyst material is present as gamma alumina.

[0196] The catalyst materials provided herein that include molybdenum, vanadium, oxygen, and aluminum cars prepared by a method that includes preparing an aqueous mixture including (i) a catalyst that includes molybdenum, vanadium, and oxygen; (ii) an aluminum compound, and (iii) a water. The method further includes removing a substantial amount of the waster from the mixture to provide a precataiyst material. The method also includes the heating the precataiyst material to provide the catalyst material.

[0197] In some embodiments, the method further includes preparing the cataly st including molybdenum, vanadium, and oxygen.

[0198] In some embodiments, the molar ratio of molybdenum to vanadium in the catalyst is from 1 :0.25 to 1 :0.6. For example, the molar ratio of molybdenum to vanadium in the catalyst can be from 1 :0.35 to 1:0.55. In some embodiments, the molar ratio of molybdenum to vanadium in the catalyst is from 1:0.40 to 1 :0.49. [0199] In some embodiments, the catalyst includes a mixed metal oxide having the empirical formula: wherein d is a number to satisfy the valence of the oxide.

[9200] In some embodiments, the aluminum compound in the mixture including the catalyst, the aluminum compound, and the water includes an aluminum oxide. In some embodiments, the aluminum oxide includes an aluminum oxide hydroxide. The aluminum oxide hydroxide can include a gibbsite, a bayerite, a boehmite, or a combination thereof. In some embodiments, the aluminum compound used to prepare the aqueous mixture includes a boehmite. In some embodiments, the boehmite includes a pseudoboehmite such as VERSAL™ 250. VERSAL™ 250 has a dispersibility index (%<lntu) of 20-30, a bulk density of 12-16 pounds per cubic foot (Ibs/ft ³ ), a surface area of about 320 meters squared per gram (m ² 7g), and a loss on ignition (LOI) of about 26 wt. %. The dispersibility index for VERSAL™ 250 can be determined by using 8 grants of sample on a volatile free basis and 96 milliliters (mL) of 0.22 normal (N) nitric acid solution, which is approximately 260 meq nitric acid per 100 grants (g) of alumina, mixing the acidic alumina slurry in a WARIN G ^® blender at low speed (17000 rpm) for 5 min, and then determining particle size distribution by using a SEDIGRAPH ^®

PSA — with the results reported as wt. % submicron particles. In some embodiments, the boehmite includes CATAPAL ^® B. CATAPAL ^® B is an alumina hydrate that has a loose bulk density of 670 to 750 g/L, a packed bulk density of 800 to 1100 g/L, a particle size (dso) of 60 gnu, a surface asea (BET) after activation at 550 ^C 'C for 3 hours of 250 m ² /g, a pore volume after activation at 550 °C for 3 hours of 0.5 ml/g, and a crystallite size ( 120) of about 4.5 nm.

[0201 ] Removing a substantial amount of water from the aqueous mixture including the catalyst, the aluminum compound, and the water to provide a precatalyst material can include removing from about 50 wt. % to about 99 wt. % of the water. For example, about 50 wt. % to about 75 wt % or about 75 wt. % to about 99 wt. % of the water can be removed to provide the precatalyst material. For example, enough water can be removed from the aqueous mixture such that the provided precatalyst material has a paste-like consistency. [0202] in some embodiments, removing a substantial amount of the water from the aqueous mixture of the catalyst, the aluminum compound, and the water to provide a precatalyst material includes heating the aqueous mixture at a temperature from about 50 °C to about 100 °C. For example, the aqueous mixture can be heated at a temperature of about 80 °C.

[0203] In some embodiments, heating the precatalyst material to provide the catalyst material includes healing the precatalyst material at a temperature from about 300 °C to about 500 °C. For example, heating the precatalyst material to provide the catalyst material can include heating the precatalyst material at a temperature from about 350 °C to about 450 °C. In some embodiments, heating the precatalyst material to provide the catalyst material includes heating the precatalyst material at a temperature of about 350 °C to about 375 °C, about 375 °C to about 400 °C, about 400 °C to about 425 °C, or about 425 °C to about 450 °C.

[0204] Further, heating the precatalyst material to provide the cataly st material can include heating the precatalyst material in the presence of air, an oxidizing atmosphere, an inert atmosphere, or a combination thereof. In some embodiments, heating the precatalyst material to provide the cataly st material includes heating the precatalyst material the presence of air.

[0205] In some embodiments, after the catal st material is prepared, the catalyst material can be treated with a gas that includes ethane, ox gen, ethylene, or a combination thereof. For example, the catalyst material can be treated with a gas that includes ethane, oxygen, ethylene, or a combination thereof in an oxidative dehydrogenation reactor.

[0206] In some embodiments, the catalyst material is treated with a gas including ethane, oxygen, ethylene, or a combination thereof at an elevated temperature. For example, the catalyst material can be treated with a gas including ethane, oxygen, ethylene, or a combination thereof at a temperature from about 100 ^C 'C to about 500 °C, about 200 °C to about 450 °C, or about 300 °C to about 400 °C. In some embodiments, the catalyst material is treated with a gas including ethane, oxygen, ethy lene, or a combination thereof for about 6 hours to about 144 hours or about 18 hours to about 72 hours. In some embodiments, the catalyst material is treated with the gas including ethane, oxygen, ethylene, or combination thereof in an oxidative dehydrogenation reactor. [0207] Catal st Materials Including Molybdenum, Vanadium, Oxygen, Aluminum, and Iron [0208] In some embodiments, the catalyst material includes molybdenum, vanadium, oxygen, aluminum, and iron. The molar ratio of molybdenum to vanadium can be from 1:0.1 to 1 :0.5. The molar ratio of molybdenum to iron can be from 1:0.25 to 1:5.5. The molar ratio of molybdenum to aluminum can be from 1:1.5 to 1:6.0. Further, oxygen can be present at least in an amount to satisfy the valency of any present metal oxides. [0209] In some embodiments, the molar ratio of molybdenum to vanadium in the catalyst material is from 1 :0.25 to 1:0.50. In some embodiments, the molar ratio of molybdenum to vanadium in the catalyst material is from 1:0.30 to 1 :0.45. For example, the molar ratio of molybdenum to vanadium can he from 1:0.30 to 1 :0.35 or from 1:0.35 to 1:0.45. In some embodiments, the molar ratio of molybdenum to vanadium is from 1:0.37 to 1:0.41.

[0210] In some embodiments, at least a portion of the molybdenum and vanadium in the catalyst material can be present as a mixed metal oxide having the empirical formula: wherein d is a number to satisfy the valence of the oxide.

[0211] In some embodiments, the molar ratio of molybdenum to iron is from 1 :0.25 to i : i 0 and the molar ratio of molybdenum to aluminum is from i :3,5 to 1 :5 5. For example, the molar ratio of molybdenum to iron can be from 1 :0.25 to 1:0.75 and the molar ratio of molybdenum to aluminum can be from 1:3.75 to 1:5.25. In some embodiments, the molar ratio of molybdenum to iron is from 1 :0.35 to 1 :0.65 and the molar ratio of molybdenum to aluminum is from 1:3.75 to 1:5.25. In some embodiments, the molar ratio of molybdenum to iron is from about 1 :0.35 to about 1 :0.45 and the molar ratio of molybdenum to aluminum is from 1:3.9 to 1 :4.0. In some embodiments, the molar ratio of molybdenum to iron is about 1 :0.55 to about 0:65, and the molar ratio of molybdenum lo aluminum is from 1 :4.95 to 1 :5.05

[0212] In some embodiments, the molar ratio of molybdenum to vanadium is from 1:0.30 to 1 :0.45, the molar ratio of molybdenum to iron is from 1 :0.25 to 1 :1, and the molar ratio of molybdenum to aluminum is from 1:3.5 to 1 :5 5. For example, the molar ratio of molybdenum to vanadium can be from 1 :0 30 to 1:0.45, the molar ratio of molybdenum to iron can be from 1:0.25 to 1 :0.75, and the molar ratio of molybdenum to aluminum can be from 1:3.75 to 1:5.25. In some embodiments, the molar ratio of molybdenum to vanadium is from 1:0.30 to 1:0.45, the molar ratio of molybdenum to iron is from 1:0.35 to 1:0.65, and the molar ratio of molybdenum to aluminum is from 1:3.75 to 1:5.25.

[0213] In some embodiments, the molar ratio of molybdenum to vanadium is from 1:0.30 to 1:0.35, the molar ratio of molybdenum to iron is from about 1 :0 35 to about 1 :0.45, and the molar ratio of molybdenum to aluminum is from 1:3.9 to 1:4.0. In some embodiments, the molar ratio of molybdenum to vanadium is from 1:0.35 to 1:0 45, the molar ratio of molybdenum to iron is about 1 :0.55 to about 0:65, and the molar ratio of molybdenum to aluminum is from 1:4.95 to i :5.05. For example, the molar ratio of molybdenum to vanadium can be from 1 :0 37 to i :0.41, the molar ratio of molybdenum to iron can be from about i :0.55 to about 0:65, and the molar ratio of molybdenum to aluminum can be from 1:4.95 to 1:5.05.

[0214! In some embodiments, the molar ratio of molybdenum to iron is from 1:1.3 to 1:2.2 and the molar ratio of molybdenum to aluminum is from 1:2.0 to 1:4.0. For example, the molar ratio of molybdenum to iron can be from 1 : 1.6 to 1 :2.0 and the molar ratio of molybdenum to aluminum can be from 1 :2.5 to 1 :3.5. In some embodiments, the molar ratio of molybdenum to iron is from about 1: 1.80 to about 1: 1.90 and the molar ratio of molybdenum to aluminum is from 1:2.9 to 1:3.1.

[0215] In some embodiments, the molar ratio of molybdenum to vanadium is from 1 :0.35 to 1 :0.45, the molar ratio of molybdenum to iron is from 1:1.3 to 1:2.2, and the molar ratio of molybdenum to aluminum is from 1 :2.0 to 1:4.0. For example, the molar ratio of molybdenum to vanadium can be from 1:0.35 to 1 :0.45, the molar ratio of molybdenum to iron can be from i : 1 6 to 1 :2 0, and the molar ratio of molybdenum to aluminum can be from ! :2.5 to 1:3.5. In some embodiments, the molar ratio of molybdenum to vanadium is from 1:0.35 to 1 :0.45, the molar ratio of molybdenum to iron is from about 1: 1.80 to about 1: 1.90, and the molar ratio of molybdenum to aluminum is from 1:2.9 to 1:3.1. For example, the molar ratio of molybdenum to vanadium can be from 1:0.37 to 1:0.41, the molar ratio of molybdenum to iron is from about 1:1.80 to about 1:1.90, and the molar ratio of molybdenum to aluminum is from 1:2.9 to 1:3.1.

[0216] In some embodiments, the molar ratio of molybdenum to iron is from 1 :2.2 to 1 :2.8 and the molar ratio of molybdenum to aluminum is from 1 : 1.8 to 1:2.4. For example, the molar ratio of molybdenum to iron can be from 1:2.4 to 1:2.6 and the molar ratio of molybdenum to aluminum can be from 1:1.9 to 1:2.3. in some embodiments, the molar ratio of molybdenum to iron is from about 1:2.45 to about 1:2.6 and the molar ratio of molybdenum to aluminum is from 1 :20 to 1 :2 2.

[0217] in some embodiments, the molar ratio of molybdenum to vanadium is from 1 :0.35 to 1 :0.45, the molar ratio of molybdenum to iron is from is from 1 :2.2 to 1:2.8, and the molar ratio of molybdenum to aluminum is from 1: 1.8 to 1:2.4. For example, the molar ratio of molybdenum to vanadium can be from 1:0.35 to 1 :0.45, the molar ratio of molybdenum to iron can be from 1 :2.4 to 1:2.6, and the molar ratio of molybdenum to aluminum can be from 1 : 1.9 to 1:2.3.

[0218] In some embodiments, the catalyst material has a 35% conversion temperature from abort! 300 °C to about 400 °C. For example, the catalyst material can have a 35% conversion temperature from about 300 °C to about .350 C. In some embodiments, the catalyst material has a 35% conversion temperature from about 315 °C to about 335 °C.

[0219] In some embodiments, the catalyst material lias a selectivity to ethylene from about 65% to 99% For example, the cataly st material can have a selectivity to ethylene from about 75% to 95%. In some embodiments, the catalyst material has a selectivity to ethylene from about 77% to about 85%.

[0220] In some embodiments, the catalyst material has a selectivity to acetic acid of less than 15 wt. % in a process for the oxidative dehydrogenation of ethane. For example, the catalyst material can have a selectivity to acetic acid of about 1 wt. % to about 15 wt. %, about 3 wt. % to about 12 wt. %, or about 7 wt. % to about 12 wt. % in a process for the oxidative dehydrogenation of ethane In some embodiments, the cataly st material has a selectivity to acetic acid of about 7 wt. %, 8 wt %, 9 wt. %, 10 wt. %, 11 wt. %, 12 wt. % or about 13 wt. % in a process for the oxidative dehydrogenation of ethane. Selectivity to acetic acid can be determined as described its the Example section.

[0221] In some embodiments, at least a portio of the iron present in the catalyst material can be present as iron (III) (i.e., Fe ³⁺ ). In some embodiments, at least a portion of the iron in the catalyst materia! can be present as amorphous iron. In some embodiments, at least a portion of the iron present in the catalyst material is present as an iron oxide, an iron oxide hydroxide, or a combination thereof.

[0222] When at least a portion of the iron in the catalyst material is present as an iron oxide, the iron oxide can include an iron oxide selected from hematite, maghemite, magnetite, or a combination thereof. In some embodiments, at least a portion of the iron in the catalyst material is present as hematite.

[0223] In some embodiments, at least a portion of the iron in the catalyst material is present as a hematite- like iron

[0224] When at least a portion of the iron in the catalyst material is present as an irots oxide hydroxide, the iron oxide hydroxide can include an iron oxide hydroxide selected from goethite, an akaganeite, a !epidoerocite, or a combination thereof. In some embodiments, at least a portion of the iron in the catalyst material is present as goethite.

[0225] In some embodiments, at least a portion of the iron in the catalyst material is present as goethite and at least a portion of the iron in the catalyst material is present as hematite.

[0226] In some embodiments, at least a portion of the aluminum in the catalyst material is present as an aluminum oxide. In some embodiments, the aluminum oxide is an aluminum oxide hydroxide. The aluminum oxide hydroxide can include a gibbsite, a bayerite, a boehmite, or a combination thereof In some embodiments, at least a portion of the aluminum in the catalyst material is present as a boehmite. In sotne embodiments, at least a portion of the aluminum in the catalyst material is present as gamma alumina.

[0227] In some embodiments, at least a portion of the iron in the catalyst material is present as goethite and at least a portion of the aluminum in the catalyst material is present as a boehmite. In some embodiments, at least a portion of the iron in the catalyst material is present as goethite, at least a portion of the iron in the catalyst material is present as hematite, and at least a portion of the aluminum in the catalyst material is present as a boehmite.

[0228] The catalyst materials provided herein that include molybdenum, vanadium, oxygen, iron, and aluminum cars prepared by a method that includes preparing an aqueous mixture including (i) a catalyst that includes molybdenum, vanadium, and oxygen; (ii) an aluminum compound; (iii) an iron compound; and (iv) a water. The method further includes removing a substantial amount of the water from the mixture to provide a precatalyst material. The method also includes heating the precatalyst material to provide the catalyst material [0229] In some embodiments, the method further includes preparing the catalyst including molybdenum, vanadium, and oxygen.

[0230] In some embodiments, the molar ratio of molybdenum to vanadium in the provided catalyst is from 1:0.25 to 1:0.6. For example, the molar ratio of molybdenum to vanadium in the catalyst can be from 1:30 to 1:35 orfroml:0.35 to 1:0.55. In some embodiments, the molar ratio of molybdenum to vanadium in the provided catalyst is from 1 :0.35 to 1 :0.42.In some embodiments, the molar ratio of molybdenum to vanadium in the provided catalyst is from 1:0.40 to 1:0.49.

[0231 ] in some embodiments, the provided catalyst includes a mixed metal oxide having the empirical formula: wherein d is a number to satisfy the valence of the oxide

[0232] In some embodiments, the iron compound in the aqueous mixture including the catalyst, the iron compound, the aluminum compound, and the water includes an iron (III) compound. In some embodiments, the iron compound in the aqueous mixture includes an iron oxide, an iron oxide hydroxide, or a combination thereof. The iron oxide can include hematite, maghemite, magnetite, or a combination thereof. In some embodiments, the iron compound in the aqueous mixture of the catalyst, the iron compound, the aluminum compound, and the water includes goethite. In some embodiments, the iron compound in the aqueous mixture of the catal st, the iron compound, the aluminum compound, and the wrater includes goethite and hematite

[0233] In some embodiments, the aluminum compound in the mixture including the catalyst, the iron compound, the aluminum compound, and the water includes an aluminum oxide. In some embodiments, the aluminum oxide includes an aluminum oxide hydroxide. The aluminum oxide hydroxide can include a gibbsite, a bayerite, a boehmite, or a combination thereof. In some embodiments, the aluminum compound used its the aqueous mixture includes a boehmite. In some embodiments, the boelunite includes a pseudoboehmite such as VERSAL™ 250. In some embodiments, the boehmite includes CATAPAL ^® B. In some embodiments, the boehmite includes PB 250 alumina.

[0234] In some embodiments, the iron compound in the aqueous mixture includes goethite and the aluminum compound includes a boelunite (e.g., VERSAL™ 250, CATAPAL ^® B, or both). In some embodiments, the iron compound in the aqueous mixture includes goethite, hematite, or both, and the aluminum compound includes a boehmite (e.g., VERSAL™ 250, CATAPAL ^® B, or both)

[0235] Removing a substantial amount of water from the aqueous mixture including the catalyst, the iron compound, the aluminum compound, and the water to provide a precatalyst material can include removing from about 50 wt, % to about 99 wt. % of the water. For example, about 50 wt. % to about 75 wt. % or about 75 wt.

% to about 99 wt. % of the water can be removed to provide the precatalyst material. For example, enough water can be removed from the aqueous mixture such that the provided catalyst material has a paste-like consistency. [0236] In some embodiments, removing a substantial amount of the water from the aqueous mixture of the catalyst, the iron compound, the aluminum compound, and the water to provide a precatalyst material includes heating the aqueous mixture at a temperature from about 50 °C to about 100 C. For example, the aqueous mixture can be heated at a temperature of about 80 °C

[0237] In some embodiments, heating the precatalyst material to provide the catalyst material includes heating the precatalyst material at a temperature from about 300 °C to about 500 °C. For example, heating the precatalyst material to provide the catalyst material can include heating the precatalyst material at a temperature from about 350 °C to about 450 ^C, C. In some embodiments, heating the precatalyst material to provide the catalyst material includes heating the precatalyst material at a temperature of about 350 °C to about 375 °C, about 375 °C to about 400 °C, about 400 °C to about 425 °C, or about 425 °C to about 450 °C.

[0238] Further, heating the precatalyst material to provide the catalyst material can include heating the precatalyst material in the presence of air, an oxidizing atmosphere, an inert atmosphere, or a combination thereof. In some embodiments, heating the precatalyst material to provide the catalyst material includes heating the precatalyst material the presence of air.

[0239] In some embodiments, after the catalyst material is prepared, the catalyst material can be treated with a gas that includes ethane, oxygen, ethylene, or a combination thereof. For example, the catalyst material can be treated with a gas that includes ethane, oxygen, ethylene, or a combination thereof in an oxidative dehydrogenation reactor.

[0240] In some embodiments, the catalyst material is treated with the gas including ethane, oxygen, ethylene, or a combination thereof at an elevated temperature. For example, the catalyst material can be treated with the gas at a temperature from about 100 °C to about 500 °C, about 200 °C to about 450 °C, or about 300 °C to about 400 °C. In some embodiments, the catalyst material is treated with the gas for about 6 hours about 144 hours or about 18 hours to about 72 hours. In some embodiments, the catalyst material is treated with the gas including ethane, oxygen, ethylene, or combination thereof in an oxidative dehydrogenation reactor.

[0241] Catalyst Materials Including Molybdenum, Vanadium, Oxygen, and Beryllium

[0242] In some embodiments, the catalyst material includes molybdenum, vanadium, beryllium, and oxy gen. The molar ratio of molybdenum to vanadium can be from 1 :0 25 to 1 0.65 The molar ratio of inolybdenuin to beryllium can be from 1:0.25 to 1 :85. Further, oxy gen is present at least in an amount to satisfy the valency of any present metal oxides.

[0243] In some embodiments, the molar ratio of molybdenum to vanadium is from 1 :0.35 to 1 :0.55. For example, the molar ratio of molybdenum to vanadium can be from 1:0.38 to 1:0.48.

[0244] In some embodiments, the molar ratio of molybdenum to beiyllium is from 1:0.35 to 1:0.75. For example, the molar ratio of molybdenum to beiyllium can be from 1:0.45 to 1:0.65.

[0245] The catalyst (material can have a 35% conversion temperature from about 300 °C to about 400 °C. For example, the catalyst material can have a 35% conversion temperature from about 310 °C to about 375 °C.

In some embodiments, the catalyst material has a 35% conversion temperature from about 15 °C to about 345

°C.

[0246] In some embodiments, the catalyst material lias a selectivity to ethylene from about 65% to 99% For example, the catalyst material can have a selectivity to ethylene from about 75% to 97%. In some embodiments, the catalyst material has a selectivity to ethylene from about 85% to 95%.

[0247] In some embodiments, the catalyst material has an amorphous phase of from 45 wt. % to 75 wt. %. For example, the catalyst material can have an amorphous phase of from 55 wt. % to 65 wt. %.

[0248] The weight percent of amorphous phase of a catalyst material can be determined using XRI) analysis

[0249] The Ortho-Mo VO _x phase contributed to the crystalline area and therefore needed to be quantified in order to determine the amorphous area. To compensate for the fact that different materials and backgrounds would have different effects, a sample of MoVTeNbO _x was used to calibrate some constants needed for the DOC method. Samples containing Mo VO _x phases had the ortho-Mo VO _x phase weight percentages qualitatively determined using only two elements (Mo and V) based on the MoVTeNbOx calibration.

[0250] In some embodiments, the catalyst material has an average crystallite size of greater than 50 nm. For example, the catalyst material can have an average crystallite size of greater than 100 n . In some embodiments, the catalyst material lias an average crystallite size from 75 n to 150 nm. For example, the catalyst material can have a mean particle size from 0.5 pm to 10 p .

[02511 The average crystallite size of catalyst material can be determined by using XRD analysis

[0252] In some embodiments, the catalyst material has a mean particle size from 2 pm to 8 pm. For example, the catalyst material can have a mean particle size from 3 pm to 5 pm.

[0253] The mean particle size of catalyst material can be determined can be determined SEM analysis [0254] In some embodiments, the catalyst material is characterized by having at least one or more XRD diffraction peaks (2Q degrees) chosen from 6.5 ± 0.2, 7.8 ± 0.2, 8.9 ± 0.2, 10.8 ± 0.2, 13.2 ± 0.2, 14.0 ± 0.2,

22.1 ± 0.2, 23.8 ± 0.2, 25.3 ± 0.2, 26.3 ± Q.2, 26.6 ± 0.2, 27.2 ± 0.2, 27.6 ± Q.2, 28.2 ± 0.2, 29.2 ± 0.2, 30.5 ±

0.2, 31.4 ± 0.2 wherein the XRD is obtained using CuKa radiation.

[0255] Catalyst Materials Including Molybdenum, Vanadium, Oxygen, Aluminum, and Beryllium [0256] In some embodiments, the catalyst tnaterial includes molybdenum, vanadium, beryllium, aluminum, and oxygen. The molar ratio of molybdenum to vanadium can be from 1:0.25 to 1:0.65 The molar ratio of molybdenum to beryllium can be from 1:0.25 to 1:7. The molar ratio of molybdenum to aluminum can be from 1 : 1 to 1:9. Further, oxy gen can be present at least in an amount to satisfy the valency of any present metal oxides. [0257] In sonie embodiments, the molar ratio of molybdenum to vanadium is from i :025 to 1:0.85 or from 1:0.35 to 1:0.55. For example, the molar ratio of molybdenum to vanadium can be from 1:0.38 to 1:0.48. [0258] In some embodiments, the molar ratio of molybdenum to beryllium is from 1:0.35 to 1:0.75. For example, the molar ratio of molybdenum to beryllium can be from 1:0.45 to 1:0.65.

[0259] In some embodiments, the molar ratio of molybdenum to aluminum is from 1 :2 to 1:8. For example, the molar ratio of molybdenum to aluminum cars be from 1:4 to 1:6.

[0268] In some embodiments, at least a portion of the aluminum in the catal st material is present as art aluminum oxide. In some embodiments, the aluminum oxide is an aluminum oxide hydroxide. In so e embodiments, the aluminum oxide hydroxide includes an aluminum oxide hydroxide selected from a gibbsite, a bayerite, a boehmite, or a combination thereof. In some embodiments, the aluminum oxide hydroxide includes a boeh ite. In some embodiments, the boehmite includes a pseudoboehmite such as VERSAL™ 250 In some embodiments, the boehmite includes CAT ARAL ^® B.

[0261 ] In some embodiments, at least a portion of the aluminum in the catalyst material is present as gamma alumina.

[0262] In some embodiments, the catalyst material has a 35% conversion temperature from about 300 °C to about 400 °C. For example, the catalyst material can have a 35% conversion temperature from about 310 °C to about 375 °C. In some embodiments, the catalyst material has a 35% conversion temperature from about 315 °C to about 345 °C.

[0263] In some embodiments, the catalyst material has a selectivity to ethylene front about 65% to 99%. For example, the catalyst material can have a selectivity to ethylene from about 75% to 97%. In some embodiments, the catalyst material has a selectivity to ethylene from about 85% to 95%

[0264] In some embodiments, the catalyst material lias an amorphous phase from 50 wt. % to 80 wt %. For example, the catalyst material can have an amorphous phase from 55 wt. % to 75 wt. %. In some embodiments, the catalyst material has an amorphous phase from 60 wt. % to 70 wt. %.

[0265] In some embodiments, the catalyst material has an average crystallite size of greater than 75 run. For example, the catalyst material can have an average crystallite size of greater than 125 nm. In some embodiments, lire catalyst material Iras an average crystallite size from 75 nm to 250 nnt For example, the catalyst material can have an average crystallite size from 125 run to 175 nm.

[0266] in some embodiments, the catalyst material lias a mean particle size from 0.5 mhi to 20 mhi For example, the catalyst material can have a mean particle size from 5 pm to 15 pm. In some embodiments, the catalyst material has a mean particle size from 7 pm to 11 pm.

[0267] In some embodiments, the catalyst material is characterized by having at least one or more XRD diffraction peaks (2Q degrees) chosen from 6.6 ± 0.2, 6.8 ± Q.2, 8.9 ± 0.2, 10.8 ± Q.2, 13.0 ± 0.2, 22.1 ± 0.2, 26.7 ± Q.2, 27.2 ± 0.2, and 28.2 ± 0.2, wherein the XRD is obtained using CuKa radiation.

[0268] In some embodiments, the catalyst material has a longitudinal crush strength from Q.66 N/mm to 200 N/mm, from 0.66 N/ntm to 150 N/nrni, from 0.66 N/nun to 100 N/mni, from 0 66 N/mm to 50 N/mm, 0.66 N/nrni to 6.67 N/ntm. For example, the catalyst material can have a longitudinal crush strength front 0.2N/mm to 4 N/mm.

[0269] In some embodiments, the catalyst material has a bulk density from 0. 1 g/mL to 2 g/niL. For example, the catalyst material can have a bulk density from 0 3 g/mL to 07 g/mL. [027(1] In sonie embodiments, the catalyst material further includes calcium. For example, the catalyst material can include from about 0.8 wt. % to about 30 wt. %, about 0.8 wt % to about 20 wt. %, about 0.8 wt. % to about 10 wt. %, or from about 2 wt. % to about 6 wt. % calcium. In some embodiments, the catalyst material includes from about 0.15 wt. % to about 2.8 wt. % calcium. In some embodiments, the catalyst material includes from about 5 wt. % to about 10 wt. % calcium, about 10 wt. % to about 15 wt. %, about 15 wt. % to about 20 wt. %, about 20 wt. % to about 25 wt. %, or from about 25 wt. % to about 30 wt. % calcium.

[0271] In some embodiments, the catalyst material further includes calcium carbonate. For example, the catalyst material can include from about 2 wt. % to about 75 wt. %, about 2 wt. % to about 50 wt. %, about 2 wt % to about 25 wt. %, or about 5 wt. % to about 15 wt. % calcium carbonate. In some embodiments, the catalyst material includes about 0.5 wt. % to about 7 wt. % calcium caibonate. In some embodiments, the catalyst material includes about 15 wt. % to about 25 wt. %, about 25 wt. % to about 35 wt. %, about 35 wt. % to about 45 wt. %, about 45 wt. % to about 55 wt. %, about 55 wt. % to about 65 wt. %, or about 65 wt. % to about 75 wt. % calcium carbonate.

[0272 ] Cataly sts Including Molybdenum and Vanadium

[0273] The catalyst materials disclosed herein can be prepared by a method that includes providing a catalyst that includes molybdenum, vanadium, and oxygen. The molar ratio of molybdenum to vanadium in the provided catalyst can be from 1:0.25 to 1:0.55. For example, the molar ratio of molybdenum to vanadium in the catalyst can be from 1:0.35 to 1:0.55. In some embodiments, the molar ratio of molybdenum to vanadium in the provided catalyst is from 1:0.40 to 1 :0.49.

[0274] The molar ratio of molybdenum to vanadium can be determined by ICP-MS.

[0275] In some embodiments, the provided cataly st is a mixed metal oxide having the empirical formula:

MO _L Vo25-05oO _d wherein d is a number to satisfy the valence of the oxide. For example, the provided catalyst can be a mixed metal oxide having the empirical formula:

MO _l Vo42-04sO _d wherein d is a number to satisfy the valence of the oxide.

[0276] In some embodiments, providing the cataly st can include preparing an aqueous mixture including molybdenum and vanadium; hydrothermally reacting the aqueous mixture to form a precalcined catalyst; and calcining the precalcined catalyst to provide the cataly st. The pH of the hydrothermal reaction can be from 2 5 to 3.5. For example, the pH of the hydrothermal reaction can be about 2 85

[0277] The aqueous mixture of molybdenum and vanadium can be prepared by combining an aqueous mixture that includes molybdenum and an aqueous mixture that includes vanadium. The aqueous mixture including molybdenum can prepared from at least ( fyieMoyC^ * 4H ₂ 0 and a first water. The aqueous mixture including vanadium can be prepared from at least VOSO * XH O and a second water. In some embodiments, the molar the molar ratio of (NH ) ₆ Mq ₇ q ₂ · 4H ₂ 0 to VOSO * XH ₂ 0 used to prepare the aqueous mixture including molybdenum and the aqueous mixture including vanadium is from 1:1.5 to 1:2. For example, the molar ratio of (NR _C MG _T C^ · 4H ₂ 0 to VOSO * XH ₂ 0 can be about 1: 1.75.

[0278] In some embodiments, the concentration of molybdenum in the aqueous mixture of molybdenum and vanadium is less than 6.3 ^c 10 ¹ snol/L. For example, the concentration of molybdenum in the aqueous mixture can be from 2.09 MO ¹ tnol/L to 3. 13 ^c !O ¹ mol/L. In some embodiments, the concentration of vanadium its the aqueous mixture of molybdenum and vanadium is less than 15.60 ^x 10 ² mol/L. For example, the concentration of vanadium in the aqueous mixture can be from 5.20 ^c 10 ² mol/L to 7.80 ^x lO ² mol/L.

[0279] In some embodiments, providing the catalyst further includes contacting the aqueous mixture of molybdenum and vanadium with a templating agent. For example, a templating agent such as a surfactant, a catalyst seed, or a combination thereof can be added to the aqueous mixture of molybdenum and vanadium. In some embodiments, the templating agent is added to the aqueous mixture of molybdenum before it is combined with the aqueous mixture of vanadium to provide the aqueous mixture of molybdenum and vanadium. In some embodiments, the templating agent is added to the aqueous mixture of vanadium before it is combined with the aqueous mixture of molybdenum to provide the aqueous mixture of molybdenum and vanadium

[0280] In some embodiments, the catalyst seed can be a catalyst seed including molybdenum and vanadium; a catalyst seed including molybdenum, vanadium, tellurium, and niobium, or a combination thereof. The catalyst seed including molybdenum and vanadium can be a previously synthesized catalyst including molybdenum and vanadium. For example, catalyst seed including molybdenum and vanadium can be a mixed metal oxide having the empirical formula:

MO _l Vo _42-048 0 _d wherein d is a number to satisfy the valence of the oxide. The catalyst seed including molybdenum and vanadium can also be an orthorhombic, trigonal, tetragonal, or amorphous solid that includes molybdenum and vanadium In some embodiments, the catal st seed including molybdenum, vanadium, tellurium, and niobium is an orthorhombic, trigonal, tetragonal, or amorphous solid that includes molybdenum, vanadium, tellurium, and niobium

[0281] In some embodiments, the weight ratio of the catalyst seed to the (NFuLMo-iCb * 4H ₂ 0 used to prepare the aqueous mixture is about 0.5: 100 to about 4.0: 100.

[0282] In some embodiments, the templating agent includes a surfactant. For example, the templating agent can include sodium dodecyl sulfate (SDS). In some embodiments, the surfactant molar loading is from about 0.005 to about 0.2 or about 0.015 to 0.2. As used herein, the term “surfactant molar loading” refers to the moles of surfactant per moles of molybdenum and moles of vanadium in the aqueous mixture including molybdenum and vanadium.

[0283] In some embodiments, the first and the second water of the aqueous mixture including molybdenum — prepared from at least (NI-D6M07O24 · 4H ₂ 0 — and the aqueous mixture including vanadium — prepared from at least VOSO4 · X¾0 — are selected independently from a distilled water, a deionized water, a demineralised water, a mineral water, or a combination thereof. In some embodiments, the first and second water include a distilled water.

[0284] In some embodiments, hydrothermally reacting the solution to form a precalcined catalyst includes heating the aqueous mixture of molybdenum and vanadium at a temperature of about 150 °C to about 300 °C. For example, hydrothermally reacting the solution to form a precalcined catalyst can include heating the aqueous mixture of molybdenum and vanadium at a temperature of about 200 °C to about 250 °C. In some embodiments, hydrothermally reacting the solution to form a precalcined catalyst includes heating the aqueous mixture of molybdenum and vanadium at a temperature of about 220 °C to about 230 °C. [0285] In sonie embodiments, hydrottiermally reacting the solution to form a precalcined catalyst includes heating the aqueous mixture of molybdenum and vanadium at a temperature of about 150 °C to about 300 °C, while maintaining a pressure above the saturated vapor pressure of the water at the corresponding temperature. [0286] In some embodiments, hydrothermally reacting the solution to form a precalcined catalyst includes contacting the solution with a glass liner, a steel liner, or a Teflon liner. For example, hydrothermally reacting the aqueous mixture to form a preeaicined catalyst can include contacting the aqueous mixture with a glass liner. In some embodiments, hydrothermally reacting the aqueous mixture to form a precalcined catalyst includes contacting the aqueous mixture with a Teflon liner. In some embodiments, hydrothermally reacting the aqueous mixture to form a preeaicined catalyst includes contacting the aqueous mixture with a steel liner. In some embodiments, the steel liner is a HASTELLOY ^® steel liner, an INCONEL ^® steel liner, or a stainless liner.

[0287] In some embodiments, the method of preparing the catalyst further includes treating the precalcined catalyst in air at a temperature from about from 250 °C to 300 °C. For example, the preeaicined catalyst can be treated in air at a temperature of about 280 °C. Treating the preeaicined catalyst at an elevated temperature can serve to remove any residual volatile compounds in the precalcined catalyst such as surfactant, if used. In some embodiments, the method of preparing the catalyst further includes treating the preeaicined catalyst in air at a temperature from about from 250 °C to 300 °C for about 0.5 hours to about 120 hours or for about 2 to about 76 hours

[0288] In some embodiments, the precalcined catalyst is calcined at about 300 °C to about 500 °C to provide the catalyst. For example, the preeaicined catalyst can be calcined at about 375 °C to about 425 °C. In some embodiments, the preeaicined catalyst is calcined at about 300 °C to about 500 °C for about 1 hour to about 24 hours. In some embodiments, precalcined catalyst is calcined at about 375 °C to about 425 °C for about I hour to about 4 hours.

[0289] In some embodiments, the precalcined catalyst is calcined in air, an inert atmosphere, or a combination thereof. In some embodiments, the precalcined catalyst is calcined in an inert atmosphere. For example, the precalcined catalyst can calcined in an inert atmosphere including N ₂ .

[0290] The Support

[0291 ] There are several ways the oxidative dehydrogenation catalyst may be supported.

[0292] In one embodiment, the support may have a low' surface area, preferably, less than 50 m ² /g, more preferably, less than 20 m ² /g. The support may be prepared by compression molding. At higher pressures, the interstices within the ceramic precursor being compressed collapse. Depending on the pressure exerted on the support precursor, the surface area of the support may be from about 20 to 5 m ² /g, preferably 18 to 10 nr/g. [Q293] There is a safety advantage using low ^r surface area supports in that there is a reduced probability that an intersti tial space may be filled only with oxidant providing a source of ignition.

[0294] The low surface area support could be of any conventional shape, such as, spheres, rings, saddles, etc. These types of supports would be used in more conventional reactors where a mixed stream or sequential stream of gaseous reactants pass over the supported catalyst and the ethane is converted to ethylene. There are a number of other approaches in the prior art where, for example, a mixed bed of supported catalyst and a reversible metal oxide may be passed together through a reaction zone to release oxide to the reaction and then regenerate the oxide. In some embodiments, the reversible metal oxide may contact a screen or permeable membrane having the supported catalyst on the other side together with a stream of ethane to release oxygen to the reaction.

[0295] In an alternate embodiment described below, the catalyst may be supported on a surface of a permeable membrane defining at least part of the flow path for one reactant and the other reactant flow's over the opposite surface of the ceramic to permit the oxidant and ethane to react on the ceramic surface.

[0296] It is important that the support be dried prior to use. Generally, the support may be heated at a temperature of at least 200 °C for up to 24 hours, typically, at a temperature front 500 °C to 800 °C for about 2 to 20 hours, preferably 4 to 10 hours. The resulting support wall be free of adsorbed water and should have a surface hydroxyl content from about 0.1 to 5 minoi/g of support, preferably, from 0.5 to 3 mmol/g of support. [0297] The dried support may th n be compressed into the required shape by compression molding. Depending on the particle size of the support, it may be combined with an inert binder to hold the shape of the compressed part.

[0298] The support for the catalyst may be a ceramic or ceramic precursor formed from oxides, dioxides, nitrides, carbides and phosphates selected from the group consisting of silicon dioxide, fused silicon dioxide, aluminum oxide, titanium dioxide, zirconium dioxide, thorium dioxide, lanthanum oxide, magnesium oxide, calcium oxide, barium oxide, tin oxide, cerium dioxide, zinc oxide, boron oxide, boron nitride, boron carbide, boron phosphate, zirconium phosphate, yttrium oxide, aluminum silicate, silicon nitride, silicon carbide and mixtures thereof

[0299] Preferred components for forming ceramic membranes include oxides of titanium, zirconium, aluminum, magnesium, silicon and mixtures thereof

[0300] Loadings

[0301 ] Typically, the catalyst loading on the support provides from 0.1 to 20 weight % typically from 5 to 15 weight %, preferably from 8 to 12 weight % of the catalyst and from 99.9 to 80 weight %, typically, from 85 to 95 weight %, preferably, from 88 to 92 weight % of the support.

[0302 ] The catalyst may be added to the support in any number of ways. For example the catal st could be deposited from an aqueous slurry onto one of the surfaces of the low' surface area support by impregnation, wash-coating, brushing or spraying. The catalyst could also be co -precipitated from a slurry' with the ceramic precursor (e.g., alumina) to form the low surface area supported catalyst.

[0303] The support and catalyst may be combined and then comminuted to produce a fine particulate material having a particle size ranging from 1 to 100 micron. The comminution process may be any conventional process including ball and bead mills, both rotary, stirred and vibratory, bar or tube mills, hammer mills, and grinding discs. A preferred method of comminution is a ball or bead mill.

[0304] The particulate catalyst may be used in an oxidative dehydrogenation reactor. The reactor may- have a single or multiple beds, preferably, multiple beds.

[0305] The Reaction

[0306] The oxidative dehydrogenation in the main reactor may be conducted at temperatures from 300 °C to 550 °C, typically, from 300 °C to 500 °C, preferably, from 350 °C to 450 °C, at pressures from 0.5 to 100 psig (3.447 to 68947 kPag), preferably, from 15 to 50 psig (103 4 to 344.73 kPag), and the residence time of the paraffin (e.g. ethane) in the reactor is typically from 0.002 to 30 seconds, preferably, from 1 to 10 seconds. The paraffin (e.g ethane) feed should be of purity of preferably, 95%, most preferably , 98%. Preferably, the process has a selectivity for olefin (ethylene) of greater than 95%, preferably, greater than 98%. The gas hourly space velocity (GHSV) will be from 500 to 30000 h ‘, preferably greater than 1000 h ¹ . The space-time yield of ethylene (productivity) in g/hoiir per kg of the catalyst should be not less than 900, preferably, greater than 1500, most preferably, greater than 3000, most desirably, greater than 3500 at 350 to 40Q °C. It should be noted that the productivity of the catalyst will increase with increasing temperature until the selectivity is sacrificed.

[0307] The conversion of ethane to ethylene should be not less than 80%, preferably, greater than 90%, most preferably, 95% or greater.

[0308] The oxygen feed may be pure oxygen, but this is expensive. The feed may comprise about 95 vol. % of oxygen and about 5 vol. % of argon. This stream is a bi-produet of nitrogen production and relatively in expensive. Argon, being inert, should not interfere with any downstream reactions

[0309] Other than in the operating examples or where otherwise indicated, all numbers or expressions referring to quantities of ingredients, reaction conditions, etc. used in the specification and claims are to be understood as modified in all instances by the term “about”. Accordingly, unless indicated to the contrary, the numerical parameters set forth in the following specification and attached claims are approximations that can vary depending upon the desired properties, which the present disclosure desires to obtain. At the veiy least, arid not as an attempt to limit the application of the doctrine of equivalents to the scope of the claims, each numerical parameter should at least be construed in light of the number of reported significant digits and by applying ordinary rounding techniques

[0310] Notwithstanding that the numerical ranges and parameters setting forth the broad scope of the disclosure are approximations, the numerical values set forth in the specific examples are reported as precisely as possible. Any numerical values, however, inherently contain certain errors necessarily resulting from the standard deviation found in their respective testing measurements.

|Q311] In addition, it should be understood that any numerical range recited herein is intended to include all sub-ranges subsumed therein. For example, a range of “1 to 10” is intended to include all sub-ranges between and including the recited minimum value of 1 and the recited maximum value of 10; that is, having a minimum value equal to or greater than 1 and a maximum value of equal to or less than 10 Because the disclosed numerical ranges are continuous, they include every value between the minimum and maximum values. Unless expressly indicated otherwise, the various numerical ranges specified in this application are approximations. [0312] As used herein, the term “alkane” refers to an acyclic saturated hydrocarbon In many cases, an alkane consists of hydrogen and carbon atoms arranged in a linear structure in which all of the carbon-carbon bonds are single bonds. Alkanes have tire general chemical formula C„H _2n+ 2. In some embodiments, alkane refers to one or more of ethane, propane, butane, pentane, hexane, octane, decane and dodecane. In particular embodiments, alkane refers to ethane and propane and, in some embodiments, ethane.

[0313] As used herein, the term “alkene” refers to unsaturated hydrocarbons that contain at least one carbon-carbon double bond. In many embodiments, alkene refers to alpha olefins. In some embodiments, alkene refers to one or more of ethylene, propylene, 1 -butene, butadiene, pentene, pentadiene, hexene, octene, decene and dodecene. In particular embodiments, alkene refers to ethylene and propylene and, in some embodiments, ethylene. [0314] As used herein, the terns “alpha olefin” or “ct-olefiit” refer to a family of organic compounds, which are alkenes (also known as olefins) with a chemical formula C _x H _2x , distinguished by laving a double bond at the primary or alpha (a) position. In some embodiments, alpha olefin refers to one or more of ethylene, propylene, 1 -butene, 1-pentene, 1-hexene, 1-octene, 1-decene and, 1-dodecene. In particular embodiments, alpha olefins refer to ethylene and propylene and, in some embodiments, ethylene.

[0315] As used herein, the term “essentially free of oxygen” means the amount of oxygen present, if any, remaining in a process stream after the one or more ODH reactors, and in many embodiments after the second reactor as described herein, is low enough that it will not present a flammability or explosive risk to the downstream process streams or equipment.

[031 ] As used herein, the term “fixed bed reactor” refers to one or more reactors, in series or parallel, often including a cylindrical tube filled with catalyst pellets with reactants flowing through the bed and being converted into products. The catalyst in the reactor may have multiple configurations including, but not limited to, one large bed, several horizontal beds, several parallel packed tubes, and multiple beds in their own shells. [0317] As used herein, the term “fluidized bed reactor” refers to one or more reactors, in series or parallel, often including a fluid (gas or liquid) which is passed through a solid granular catalyst, which can be shaped as tiny spheres at high enough velocities to suspend the solid and cause it io behave as though it were a fluid. [0318] As used herein, the term “.gas phase polyethylene process” refers to a process where a mixture of ethylene, optional alpha olefin comonomers and hydrogen is passed over a catalyst in a fixed or fluidized bed reactor. The ethylene and optional alpha olefins polymerize to form grains of polyethylene, suspended its the flowing gas, which can pass out of the reactor. In some embodiments, two or more of the individual reactors are placed in parallel or in series, each of which are under slightly different conditions, so that the properties of different polyethylenes from the reactors are present in the resulting polyethylene blend. In many cases the catalyst system includes, but is not limited to, chromium catalysts, Ziegler-Natta catalysts, zirconocene catalysts, and metallocene catalysts and combinations thereof.

[03191 As used herein, the term “HOPE” refers to high-density polyethylene, which generally has a density of greater or equal to 0.941 g/cm ³ . HOPE has a low degree of branching. HOPE is often produced using chromium/silica catalysis, Ziegler-Natta catalysts or metallocene catalysts.

[0320] As used herein, the term “high pressure polyethylene process” refers to converting ethylene gas into a white solid by heating it at very high pressures in the presence of minute quantities of oxygen (about < 10 ppm oxygen) at about 1000 - 3000 bar and at about 80 - 300 ^C 'C. In many cases, the high-pressure polyethylene process produces LDPE.

[0321] As used herein, the tern “LDPE” refers to low density' polyethylene, which is a polyethylene with a high degree of branching with long chains. Often, the density of a LDPE will range from 0.910 - 0.940 g/cm ³ . LDPE is created by free radical polymerization.

[0322] As used herein, the term “LLDPE” refers to linear low-density polyethylene, which is a polyethylene that cars have significant numbers of short branches resulting from copolymerization of ethylene with at least one a-olefin comonomer. In some cases, LLDPE has a density in the range of 0.915 - 0.925 g cm ³ . In many cases, the LLDPE is an ethylene hexene copolymer, ethylene octene copolymer or ethylene butene copolymer. The amount of comonomer incorporated can be from 0.5 to 12 mole %, in some cases from 1.5 to 10 mole %, and in other cases from 2 to 8 mole % relative to ethylene. [0323] As used herein, the tern “long-chain branching” refers to a situation where during a-olefin polymerization, a vinyl terminated polymer chain is incorporated into a growing polymer chain. Long branches often have a length that is longer than the average critical entanglement distance of a linear (no long chain branching) polymer chain. In many cases, long chain branching affects melt rheological behavior.

[0324] As used herein, the term “low pressure polyethylene process” refers to polymerizing ethylene using a catal st that in many cases includes aluminum at generally lower pressures than the high-pressure polyethylene process. In many cases, the low-pressure polyethylene process is carried out at about 10 to 80 bar and at about 70 to 300 °C. In many cases, the low-pressure polyethylene process provides HOPE. In particular cases, an a-olefin comonomer is included in the low-pressure polyethylene process to provide LLDPE

[0325] As used herein, the term “MDPE” refers to medium density polyethylene, which is a polyethylene with some short and/or long chain branching and a density in the range of 0.926 to 0 940 g/enr\ MDPE can be produced using chromium/ silica catalysts, Ziegler-Natta catalysts or metallocene catalysts.

[0326] As used herein, the tern “monomer” refers to small molecules containing at least one double bond that reacts in the presence of a free radical polymerization initiator to become chemically bonded to other monomers to form a pol er.

[0327] As used herein, the term “moving bed reactor” refers to reactors in which the catalytic material Slows along wish She reactants and is then separated from the exit stream and recycled

[0328] As used herein, the term “MoVOx catalyst” refers to a mixed metal oxide having She empirical formula Mo _{& 5-?} oV:iO , where d is a number So satisfy the valence of the oxide; a mixed metal oxide having the empirical formula

Mo _f , _25-725 ViO _c s, where d is a number to satisfy tise valence of the oxide, or combinations thereof

[0329] As used herein, the term, “olefinic monomer” includes, without limitation, a-olefins, and in particular embodiments ethylene, propylene, 1-butene, 1-hexene, 1-octene and combinations thereof.

[0330] As used herein, the term, “oxidative dehydrogenation” or “ODH” refers to processes that couple the endothermic dehydration of an alkane with the strongly exothermic oxidation of hydrogen as is further described herein.

[0331] As used herein, the term “polyolefin” refers to a material, which is prepared by polymerizing a monomer composition containing at least one olefinic monomer.

[0332] As used herein, the term “polyethylene” includes, without limitation, homopolymers of ethylene and copolymers of ethylene and one or more a-olefins.

[0333] As used herein, the tern “polypropylene” includes, without limitation, homopolymers of propylene, including isotactic polypropylene and syndiotactic polypropylene and copolymers of propylene and one or more a-olefins.

[0334] As used herein, the term “polymer” refers to macromolecules composed of repeating structural units connected by covalent chemical bonds and is meant to encompass, without limitation, homopolyniers, random copolymers, block copolymers, and graft copolymers.

[0335] In the disclosure, “reactive oxygen” means oxygen taken up by the oxidative dehydrogenation catalyst, which is available to be used in the oxidative dehydrogenation reaction and removed from the catalyst. [0336] In the disclosure the tern “reactive oxygen depleted” when referring to the catalyst in the pre reactor is not intended to mean absolute oxygen depletion. Rather it means that the levels of residual reactive oxygen in the catalyst is sufficiently low so that there is less than 25%, preferably less than 15% most preferably less than 10% of the maximum amount of oxygen, which has been be taken up by the catalyst. After giving up reactive oxygen the catal sts comprises metal oxides, which do not give up oxygen.

[0337] Substantially saturated with reactive oxygen means that not less than 60%, preferably more than 70%, most preferably more than 85% of the reactive oxygen has been compJexed with the oxidative dehydrogenation cataly st.

[0338] As used herein, the term “short; chain branching” refers to copolymers of ethylene with an a-olefin or with branches of less than about 40 carbon atoms. In many cases, the a-olefin or branches are present at less than 20 wt. %, in some cases less than 15 wt. % of the polyethy lene. In many cases, the presence of short chain branches interferes with the formation of the polyethylene crystal structure and is observed as a lower density compared with a linear (no short chain brandling) polyethylene of the same molecular weight.

[0339] As used herein, the term “solution polyethylene process” refers to processes that polymerize ethylene and one or more optional a-oiefins in a mixture of lower alkane hydrocarbons in the presence of one or more catalysts. In some embodiments, two or more of the individual reactors are placed in parallel or in series, each of which can be under slightly different conditions, so that the properties of different polyethylenes from the reactors are present in the resulting polyethylene blend. In many cases the catalysts include, but are not limited to, chromium catalysis, Ziegier-Natta catalysis, zireonocene catalysts, hafuocene catalysts, phosphinimine catalysts and metallocene catalysts and combinations thereof

[034ft] As used herein, the term “slurty polyethylene process” refers to single-tube loop reactors, double tube loop reactors or autoclaves (stirred-tank reactors) used to polymerize ethylene and optional a-oiefins in the presence of a catalyst system and a diluent. Non-limiting examples of diluents include isobutane, n-hexane or n-heptane. In some embodiments, two or more of the individual reactors are placed in parallel or in series, each of which can be under slightly different conditions, so that the properties of different polyethylenes from the reactors are present in the resulting polyethylene blend. In many cases the catalyst system includes, but is not limited to, chromium catalysts, Ziegier-Natta catalysts, zireonocene catalysts, hafnocene catalysts, phosphinimine catalysts and metallocene catalysts arid combinations thereof.

[0341] As used herein, the term “substantially free of acetylene” means the amount of acetylene present, if any, remaining its a process stream after the one or more ODH reactors, and its many embodiments after the second reactor as described herein, is undetectable using the analytical techniques described herein or zero vppm.

[0342] As used herein, the term “swing bed type reactor arrangement” is a gas phase reactor system where a first vessel effectively operates as a reactor and a second v essel effectively operates as a regenerator for regenerating the catalyst system. This arrangement can be used with fixed bed as well as fluidized bed gas phase polyethylene reactors.

[0343] As used herein, the term “thermoplastic” refers to a class of polymers that soften or become liquid when heated and harden when cooled. In many cases, thermoplastics are high-molecular-weight polymers that can be repeatedly heated and remolded. In some embodiments, thermoplastic resins include polyolefins and elastomers that have thermoplastic properties. [0344] As used herein, the terns “thermoplastic elastomers” and “TPE” refer to a class of copolymers or a blend of polymers (in many cases a blend of a thermoplastic and a rubber) which includes materials having both thermoplastic and elastomeric properties.

[0345] As used herein, the terms “thermoplastic olefin” or “TPO” refer to polymer/filler blends that contain some fraction of polyethylene, polypropylene, block copolymers of polypropylene, rubber, and a reinforcing filler. The fillers can include, without limitation, talc, fiberglass, caibon fiber, wollastonite, and/or metal oxy sulfate. The rubber can include, without limitation, ethylene-propylene rubber, EPDM (ethylene- propylene-diene robber), ethylene-butadiene copolymer, styrene-ethylene-butadiene-styrene block copolymers, styrene-butadiene copolymers, ethylene-vinyl acetate copolymers, ethylene-alkyl (meth)aciylate copolymers, very low density polyethylene (VLDPE) such as those available under the Ftexomer ^® resin trade name from the Dow' Chemical Co , Midland, MI, styrene-ethylene-eihylene-propylene-slyrene (SEEPS). These can also be used as the materials to be modified by the interpolymer to tailor their rheological properties.

[0346] As used herein, the term “VLDPE” refers to very low-density polyethylene, which is a polyethylene with high levels of short chain branching with a typical density in the range of Q.880 - 0.915 g/cc. In many cases, VLDPE is a substantially linear polymer. VLDPE is ty pically produced by copolymeiization of eth lene with a-olefins VLDPE is often produced using metallocene catalysts.

[0347] Unless otherwise specified, all molecular weight values are determined using gel permeation chromatography (GPC). Molecular weights are expressed as polyethylene equivalents with a relative standard deviation of 2.9% for the number average molecular weight (“Mn”) and 5.0% for the weight average molecular weight (“Mw”) Unless otherwise indicated, the molecular weight values indicated herein are weight average molecular weights (Mw).

[0348] In some embodiments disclosed herein, the degree to which carbon monoxide is produced during the ODH process can be mitigated by converting it to carbon dioxide, which can then act as an oxidizing agent. The process can be manipulated so as to control the output of carbon dioxide from the process to a desired level. Using the methods described herein, a user may choose to operate in carbon dioxide neutral conditions such that surplus carbon dioxide need not be flared or released into the atmosphere.

[0349] Figures 1A and IB are block diagrams of a selective oxidation system 100 for the selective oxidation of light hydrocarbons, in accordance wdth examples. The oxidizing agent generally used in the process is air 102, although oxygen, generally mixed with a diluent, may also be used. The air 102 is flowed into an air separation unit (ASU) 104 In the ASU 104, the oxygen 106 is separated from other gases, such as nitrogen and carbon dioxide, among others. The oxygen 106 may then be mixed with a diluent, for example, in a steam dilution system 108.

[0350] To avoid process upsets, in many embodiments, mixtures of one or more alkanes with oxy gen are employed using ratios that fall outside of the flammability envelope of the one or more alkanes and oxy gen. In some embodiments, the ratio of alkanes to oxygen may Ml outside the upper flammability envelope. In these embodiments, the percentage of oxygen in the mixture can be less than 30 wt %, in some cases less than 25 wt. %, or in other cases less than 20 wt. %, but greater than zero.

[0351] In embodiments with higher oxygen percentages, alkane percentages can be adjusted to keep the mixture outside of the flammability envelope. While a person skilled in the art would be able to determine an appropriate ratio level, in many cases the percentage of alkane is less than about 40 vol. % and greater than zero. As a non-limiting example, where the mixture of gases prior to ODH includes 20 voi. % oxygen and 40 vol. % alkane, the balance can be made up with an inert diluent. Non-limiting examples of useful inert diluents in this embodiment include, but are not limited to, one or more of steam, nitrogen, and carbon dioxide, among others. In some embodiments, the inert diluent should exist in the gaseous state at the conditions w ithin the reactor and should not increase the flammability of the hydrocarbon added to the reactor, characteristics that a skilled w'orker would understand when deciding on which inert diluent to employ. The inert diluent cats be added to either of the alkane containing gas or the oxygen containing gas prior to entering the ODH reactor or may be added directly into the ODH reactor.

[0352] Although a number of different hydrocarbons may be used, in an oxidative dehydration process, generally ethane is provided to the reactor along witis oxygen. In some embodiments, the volumetric feed ratio of oxygen to etisane (CVCzHe) provided to the one or more ODH reactors can be at least about 0.3, in some cases at least about 0.4, and in other cases at least about 0.5 and can be up to about 1, in some cases up to about 0.9, in other cases up to about 0.8, in some instances up to about 0.7 and in other instances up to about 0.6. The volumetric feed ratio of oxy gen to ethane can be any of the values or range between any of the values recited above.

[0353] In some embodiments, mixtures that fall within the flammability envelope may be employed, for example, in instances here the mixture exists in conditions that prevent propagation of an explosive event. In these non-limiting examples, the flammable mixture is created within a medium where ignition is immediately quenched. As a further non-limiting example, a user may design a reactor where oxygen and the one or more alkanes are mixed at a point where they are surrounded by a flame arresting material. Any ignition would be quenched by the surrounding material. Flame arresting materials include, but are not limited to, metallic or ceramic components, such as stainless steel walls or ceramic supports. In some embodiments, oxygen and alkanes can be mixed at a low temperature, w here an ignition event would not lead to an explosion, then introduced into the reactor before increasing the temperature. The flammable conditions do not exist until the mixture is surrounded by the flame arrestor material inside of the reactor.

[0354] The amount of steam added to the SO process in the steam dilution system 108 affects the degree to which carbon dioxide acts as an oxidizing agent. In some embodiments, steam may be added directly to the SO reactor 110, or steam may be added to the individual reactant components — the lower alkane, oxygen, or inert diluent — or combinations thereof and subsequently introduced into the SO reactor 1 10 along with one or more of the reactant components. Alternatively, steam may be added indirectly as water mixed with either the lower alkane, oxygen or inert diluent, or a combination thereof, with the resulting mixture being preheated before entering the reactor. When adding steam indirectly as water, a heater 112 is used to increase the temperature so that the water is entirely converted to steam before entering the reactor.

[0355] Increasing the amount of steam added to the SO reactor 110 increases the degree to which carbon dioxide acts as an oxidizing agent. Decreasing the amount of steam added to the SO reactor 110 decreases the degree to which carbon dioxide acts as an oxidizing agent. In some embodiments, a user monitors the carbon dioxide output and compares it to a predetermined target carbon dioxide output. If the carbon dioxide output is above the target, a user can then increase the amount of steam added to the ODH process. If the carbon dioxide output is below the target, a user can decrease the amount of steam added to the ODH process, provided steam has been added. Seting a target carbon dioxide output level is dependent on the requirements for the user. In some embodiments, increasing the steam added will have the added effect of increasing the amount of acetic acid and other by-products produced in the process. As larger amounts of acetic acid from the output of the ODH may be generated by higher levels of steam, reducing steam levels will decrease the amount generated. Conversely, higher levels of steam will increase the amount of carbon dioxide consumed.

[0356] In some embodiments, the amount of steam added to the SO reactor 110 can be up to about 5Q vol. %, in some circumstances up to about 40 vol %, m some cases up to about 35 vol. %, in other cases up to about 30 vol. %, and in some instances up to about 25 vol. % and can be zero, in some cases at least 0.5 vol. %, in other cases at least 1 vol %, in other cases at least 5 vol. %, m some instances at least 10 vol. % and in other instances at least 15 vol. % of the stream entering the SO reactor 110 The amount of steam in the stream entering the SO reactor 1 10 can be any value or range between any of the values recited above. As used herein, the SO reactor 110 may include a single reactor, or multiple reactors.

[0357] In some embodiments when using two or more SO reactors a user may choose to control carbon dioxide output in only one, or less than the whole complement of reactors. For example, a user may opt to maximize carbon dioxide output of an upstream reactor so that the higher level of carbon dioxide can be part of the inert diluent for the subsequent reactor. In that instance, maximizing carbon dioxide output upstream minimizes the amount of inert diluent that would need to be added to the stream prior to the next reactor.

[0358] There is no requirement for adding steam to an SO process, as it is one of many alternatives for the inert diluent. For processes where no steam is added, the carbon dioxide output is maximized under the conditions used with respect to ethane, oxygen, and inert diluent inputs. Decreasing the carbon dioxide output can then be a matter of adding steam to tise reaction until carbon dioxide output drops to the desired level. In embodiments where selective oxidation, or oxidative dehydrogenation, conditions do not include addition of steam, and the carbon dioxide output is higher than tire desired carbon dioxide target level, steam may be introduced into the reactor while keeping relative amounts of the main reactants and inert diluent — lower alkane, oxygen and inert diluent — added to the reactor constant, and monitoring the carbon dioxide output, increasing the amount of steam until carbon dioxide decreases to the target level.

[0359] In some embodiments, a carbon dioxide neutral process can be achieved by increasing steam added so that any carbon dioxide produced in the oxidative dehydrogenation process can then be used as an oxidizing agent such that there is no net production of carbon dioxide. Conversel , if a user desires net positive carbon dioxide output tisen the amount of steam added to the process can be reduced or eliminated to maximize carbon dioxide production As the carbon dioxide levels increase there is potential to reduce oxygen consumption, as carbon dioxide is competing as an oxidizing agent. The skilled person would understand that using steam to increase the degree to which carbon dioxide acts as an oxidizing agent can affect oxygen consumption. The implication is that a user can optimize reaction conditions with lower oxygen contributions, which may assist in keeping mixtures outside of flammability limits.

[0360] From the heater 112, the feed is introduced into the SO reactor 110. The SO reactor 110 may be any of the known reactor types applicable for an SO process, such as the ODH of alkanes. In some embodiments, the SO reactor 110 is a conventional fixed bed reactor. In a typical fixed bed reactor, reactants are introduced into the reactor at one end, and flow past an immobilized catalyst, during which products are formed. The products leave the SO reactor 110 at the opposite end from where the feed is introduced. Designing a fixed bed reactor suitable for the methods disclosed herein can follow techniques known for reactors of this type. A person skilled its the art would know which features are required with respect to shape and dimensions, inputs for reactants, outputs for products, temperature, and pressure control, and means for immobilizing the catalyst. [0361] In some embodiments, the use of inert non-catalytic heat dissipative particles can be used within one or more of the SO reactors. In various embodiments, the heat dissipative particles are present within the bed and include one or more non catalytic inert particulates having a melting point at least 30 °C, in some embodiments at least 250 C, in further embodiments at least 500 °C above the temperature upper control limit for the reaction; a particle size in the range of 0.5 to 75 nun, in sotne embodiments 0.5 to 15, in further embodiments in the range of 0.5 to 8, in further embodiments in the range of 0 5 to 5 mm; and a thermal conductivity of greater than 30 W/rnK (watts/meter Kelvin) within the reaction temperature control limits. In some embodiments, the particulates are metal alloys and compounds having a thermal conductivity of greater than 50 W/rnK (waits/meter Kelvin) within the reaction temperature control limits. Non-limiting examples of suitable metals that can be used in these embodiments include, but are not limited to, silver, copper, gold, aluminum, steel, stainless steel, molybdenum, and tungsten.

[0362] The heat dissipative particles can have a particle size of from about 1 nun to about 15 mm. In some embodiments, the particle size can be from about 1 min to about 8 rum. The heat dissipative particles can be added to the fixed bed in an amount from 5 to 95 wt. %, in some embodiments from 30 to 70 wt. %, in other embodiments from 45 to 60 wt. % based outlie entire weight of the fixed bed The particles are employed to potentially improve cooling homogeneity and reduction of hot spots in the fixed bed by transferring heat directly to the walls of the reactor. As described herein, in embodiments the SO reactor 110 may be cooled by the generation of high-pressure steam 114, for example, in a jacket around or coils within the SO reactor 110

[0363] Additional embodiments include the use of a fluidized bed reactor, where tise catalyst bed can be supported by a porous structure, or a distributor plate, located near a bottom end of the reactor and reactants flow through at a velocity sufficient to fluidize the bed (e.g. the catalyst rises and begins to swirl around in a fluidized manner). The reactants are converted to products upon contact with the fluidized catalyst and the reactants are subsequently removed from Ore upper end of the reactor. Design considerations those skilled in the art can modify and optimize include, but are not limited to, the shape of the reactor, the shape and size of the distributor plate, the input temperature, the output temperature, and reactor temperature and pressure control. [0364] Some embodiments include using a combination of both fixed bed and fluidized bed reactors, each with the same or different ODH catalyst. The multiple reactors can be arrayed in series or in parallel configuration, the design of which falls within the knowledge of the worker skilled in the art.

[0365] In some embodiments, the stream exiting the one or more SO reactors can be treated to remove or separate water and w ^r ater-soluble hydrocarbons from the stream exiting the one or more SO reactors. In some embodiments, this stream is fed to a second reactor.

[0366] In some embodiments, the stream exiting the SO reactor 110 is directed to a quench fewer 118 to be cooled and condensed. This facilitates the removal of oxygenates, such as water 120 and acetic acid 122, via a bottom outlet that feeds an acetic acid separator 124 The acetic acid separator 124 separates an acetic acid stream 122 front the water stream 120, as well as separating a gas stream that is returned to an acetic acid scrubber 126 The water stream 120 may be treated in a bio oxidation unit 128 to remove any remaining catbon compounds, such as traces of acetic acid, among others. From the bio oxidation unit 128, the purified water stream 120 may be fed to a cooling tower 130 as a makeup stream. [0367] The remaining gases from the quench tower 118 are fed to the acetic acid scrubber 126, along with separated gases from the acetic acid separator 124. The acetic acid scrubber 126 may remove traces of acetic acid, and other carbon compounds, from these gas streams by oxidation or adsorption.

[0368] A stream 132 containing unconverted lower alkane (such as ethane), corresponding alkene (such as ethylene), unreacted x ' gen, carbon dioxide, carbon monoxide, optionally acetylene and inert diluent, are allowed to exit the acetic acid scrubber 126 and are fed to an oxygen removal system 134 (Figure IB), or to a second reactor, as described with respect to Figures 2 and 3.

[0369] The oxygenates removed via the quench tower 118 and/or acetic acid scrubber 126 can include carboxylic acids (for example acetic acid), aldehydes (for example acetaldehyde) and ketones (for example acetone). The amount of oxygenate compounds remaining in the stream 132 exiting the scrubber and fed to the oxygen removal system 134 will often be zero, for example, below the detection limit for analytical test methods typically used to detect such compounds. When oxygenates can be detected they can be present at a level of up to about 1 per million by volume (ppmv), in some cases up to about 5 ppmv, in other cases less than about 10 ppmv, in some instances up to about 50 ppmv and in other instances up to about 100 ppmv and can be present up to about 2 vol. %, in some cases up to about 1 vol. %, and in other cases up to about 1,000 ppmv. The amount of oxygenates or acetic acid in the stream exiting the scrubber and fed to the oxygen removal system 134 can be any value, or range between any of the values recited above.

[0370] In the oxygen removal system 134, as described herein, a high temperature membrane may be used to remove oxygen from the stream 132 exiting the acetic acid scrubber 126 The high temperature membrane may be heated by combusting access hydrocarbons in the stream 132, by combusting fuel added to the oxygen removal system 134, or both. A stream i 01 exiting the acetic acid scrubber 126 can be recy cled to the steam dilution system 108.

[0371] From the oxygen removal system, the stream 132 may be compressed, for example, in a first compressor system 136. The first compressor system 136 may include a single compressor or a series of compressors that sequentially boost the pressure of the stream 132. The compressed stream may then be fed to an amine scnibber 138 to remove CO 140 front the compressed stream, as described in further detail herein. From the amine scrubber 138, the compressed stream may be fed to a caustic wash tower 142. The caustic wash tower 142 further reduces the concentration of CO in the compressed gas stream, sending the CO in a rich caustic stream 144 The rich caustic stream 144 may then be treated to form a lean caustic stream, which is returned to the caustic wash tower 142.

[0372] The purified gas stream from the caustic wash tower 142 may include unconverted lower alkane (such as ethane) and the corresponding alkene (such as ethylene), and excess inert diluent, such as nitrogen. The purified gas stream may be compressed in a second compressor system 146. The second compressor system 146 may include a single compressor or a chain of compressors that sequentially boost the pressure of the purified gas. The compressed purified gas may then be passed to a dryer 148 to remove excess water vapor from the amine scrubber 138 and the caustic tower 142. The diver 148 may include molecular sieves to absorb the water, or may include a senes of heat exchangers and chillers to physically condense the water, or both

[0373] The dried stream is then passed to a chiller 150. The chiller 150 may include a series of heat exchangers, such as propane chilled heat exchangers, compressed nitrogen chilled eat exchangers, and heat exchangers cooled by fluids from other portions of the process. The chiller 150 may be integrated with, or feed, a depropanizer (C3R) 152, a deethanizer (C2R) 154, or both.

[0374] Returning to Figure 1 A, the chilled gas stream is fed to a demethanizer 156. From the demethanizer 156, an off gas stream 158 is sent to waste or to downstream processes. The off gas stream 158 includes the remainder of the inert diluent as well as methane removed from the chilled gas stream. Further, the demethanizer 156 returns a portion of the Ci compounds, such as ethylene and ethane, lo the process upstream of the first compressor system 136. A Ci stream from the demethanizer 156 is fed to a C splitter 160.

[0375] The Ci splitter 160 divides the Ci stream into an ethylene product stream 162 and an ethane feed stream 164. The ethane feed stream 164 is vaporized in a heat exchanger 166 to form an ethane gas feed stream. An ethane feed 168 from another ethane source may be vaporized in a heat exchanger 170 and blended into the ethane gas feed stream.

[0376] The ethane gas feed stream is then passed through a high temperature heat exchanger 172 to be superheated. The superheated ethane gas feed stream is then fed to the steam dilution system 108 for use in the process. The core reaction process, including the separation of oxygenates, amine washing, and caustic washing me described further with respect to Figures 2 and 3, below.

[0377] Figure 2 is a graphic depiction of a chemical complex 200, according to some embodiments. In the following description, like parts are designated by like reference numbers.

[0378] In some embodiments, the chemical complex, shown in one embodiment schematically in Figure 2, includes, in cooperative arrangement, an SO reactor, such as an ODH reactor 202, a quench tower and/or acetic acid scrubber 204, a second reactor 206 (as described herein), an amine wash tower 208, a drier 210, a distillation tower 212, and an oxygen separation module 214 ODH reactor 202 includes an SO catalyst capable of catalyzing, in the presence of oxygen, which may be introduced via oxygen line 216, the oxidative dehydrogenation of alkanes introduced via alkane line 218. Although second reactor 2Q6 is shown directly after quench tower or acetic acid scrubber 204, it can be placed further downstream. In many cases, the process configuration can be more energy efficient if second reactor 206 is placed after the input stream has been compressed.

[0379] The ODH reaction may also occur m the presence of an inert diluent, such as carbon dioxide, nitrogen, or steam, that is added to ensure the mixture of oxygen and hydrocarbon are outside of flammability limits. Determination of whether a mixture is outside of the flammability limits, for the prescribed temperature and pressure, is within the knowledge of the skilled worker. An ODH reaction that occurs within ODH reactor 202 may also produce, depending on the catalyst and the prevailing conditions within ODH reactor 202, a variety of other products which may include carbon dioxide, carbon monoxide, oxygenates, and water. These products leave ODH reactor 202, along with unreacted alkane, corresponding aikene, residual oxygen, carbon monoxide, and inert diluent, if added, via ODH reactor product line 220.

[0389] ODH reactor product line 220 is directed to quench tow'er or acetic acid scrubber 204, which quenches the products from product line 220 and facilitates removal of oxygenates and w'ater via quench tower bottom outlet 222 Unconverted lower alkane, corresponding alkene, unreacted oxygen, carbon dioxide, carbon monoxide, and inert diluent added to acetic acid scrubber (quench tower) 204 exit through quench tower overhead line 224 and are directed into second reactor 206 [0381] Second reactor 206 contains the group 11 metal with optional promoter and optional support as described above, which causes unreacted oxygen to react with carbon monoxide to form carbon dioxide or, optionally, reacts acetylene to reduce or eliminate it. In second reactor 206, most or all of the unreacted oxygen and acetylene is consumed. All or a portion of the carbon dioxide in second reactor 206 can be recycled back to ODH reactor 202 via recycle lines 226 and 227 to act as an oxidizing agent as described above. The remaining unconverted lower alkane, corresponding alkene, unreacted oxygen (if present), all or part of the carbon dioxide, carbon monoxide (if present), and inert diluent are conveyed to amine wash tower 208 via line 228.

[0382] Any carbon dioxide present in line 228 is isolated by amine wash tower 208 and captured via carbon dioxide bottom outlet 230 and may be sold, or, alte natively, may be recycled back to ODH reactor 202 as described above. Constituents introduced into amine wash tower 208 via line 228, other than carbon dioxide, leave asnine wash tower 208 through amine wash tower overhead line 232 and are passed through the dryer 210 before being directed to distillation tower 212 through line 234, where C /C ₊ hydrocarbons are isolated and removed via C /C ₊ hydrocarbons bottom outlet 236. The remainder includes mainly Ci hydrocarbons, including remaining inert diluent and carbon monoxide (if any), which leave distillation tower 212 via overhead stream 238 and is directed to oxygen separation module 214.

[0383] Oxygen separation module 214 includes a sealed vessel having a retentate side 240 and a permeate side 242, separated by oxygen transport membrane 244. Overhead stream 238 may be directed into either of retentate side 240 or permeate side 242. Optionally, a flow controlling means, as discussed herein, may be included that allows for flow into both sides at varying levels. In that instance an operator may choose what portion of the flow from overhead stream 238 enters retentate side 240 and what portion enters pernseate side 242. Depending upon conditions, an operator may swatch between the two sides, to allow equivalent amounts to enter each side, or bias the amount directed to one of the two sides. Oxygen separation module 214 also includes air input 246 for the introduction of atmospheric air, or other oxygen containing gas, into the retentate side 240. Combustion of products introduced into retentate side 240, due to the introduction of oxygen, may contribute to raising the temperature of oxy gen transport membrane 244 to at least about 850 °C so that oxygen can pass from retentate side 240 to permeate side 242. Components within the atmospheric air, or other oxygen containing gas, other than oxygen, cannot pass from retentate side 240 to permeate side 242 and can only leave oxygen separation module 214 via exhaust 248.

[0384] As a result of oxygen passi ng from retentate side 240 to permeate side 242, there is separation of oxygen from atmospheric air, or other oxygen containing gas, introduced into retentate side 240 The result is production of oxygen enriched gas on permeate side 242, which is then directed via oxygen enriched bottom line 220 to ODH reactor 202, either directly or in combination with oxygen line 216 (as shown in Figure 1). When overhead stream 238 is directed into retentate side 240, the degree of purity of oxygen in oxygen-enriched botom line 220 can approach 99%. Conversely, when overhead stream 238 is directed into permeate side 242 the degree of purity of oxygen in oxygen enriched botom line 220 is lower, with an upper limit ranging from 80 vol. % 90 vol. % oxygen, the balance in the form ofcaibon dioxide, water, and remaining inert diluent, all of which do not affect the ODH reaction as contemplated by the present disclosure and can accompany the enriched oxygen into ODH reactor 202. Water arid carbon dioxide can be removed by quench tower 204 and amine wash tower 208, respectively. In some embodiments of the disclosure, some or ail of the carbon dioxide can be captured for sale as opposed to being flared where it contributes to greenhouse gas emissions. In other embodiments, when carbon dioxide is used in the ODH process, any carbon dioxide captured in the amine wash can be recycled back to ODH reactor 202.

[0385] Oxygen transport membrane 244 is temperature dependent, only allowing transport of oxygen when the temperature reaches at least about 850 °C. In some embodiments, the components in overhead stream 238 by themselves are not capable, upon combustion in the presence of oxygen, to raise the temperature of oxygen transport membrane 244 to the required level. In this embodiment, the chemical complex of the present disclosure also includes fuel enhancement line 250, upstream of oxygen separation module 214, whe re combustible fuel, as a non-limiting example methane, may be added to supplement the combustible products from overhead stream 238.

[0386] In some embodiments, a concern for ODH processes is the mixing of a hydrocarbon with oxygen. Under certain conditions, the mixture may be unstable and lead to an explosive event. Mixers may be used to mix a hydrocarbon containing gas with an oxygen containing gas in a flooded mixing vessel. By mixing in this way, pockets of unstable compositions are surrounded by a non-flammable liquid so that even if an ignition event occurred it would be quenched immediately. Provided addition of the gases to the ODH reaction is controlled so that homogeneous mixtures fall outside of the flammability envelope, for the prescribed conditions with respect to temperature and pressure, the result is a safe homogeneous mixture of hydrocarbon and oxygen. [0387] In some embodiments, there is a Hooded gas mixer 302 (Figure 3) upstream of ODH reactor 202.

In this instance, oxygen line 216 and alkane line 218 feed directly into flooded gas mixer 302. A homogeneous mixture that includes hydrocarbon and oxygen, and optionally an inert diluent, can be introduced into ODH reactor 202 from flooded gas mixer 302 via mixed line 304 (Figure 3). Oxygen enriched bottom line 227 may feed directly into the flooded gas mixer 302 or in combination with oxygen line 216 into the flooded gas mixer 302.

[0388] The temperature of the contents within product line 220 in a typical ODH process can reach about 450 °C. It can be desirable to lower the temperature of the stream before introduction into quench tower or acetic acid scrubber 204 as described above. In that instance, the present disclosure contemplates the use of a heat exchanger immediately downstream of each ODH reactor 202 and immediately upstream of acetic acid scrubber 204 Use of a heat exchanger to lower temperatures in this fashion is well known in the art.

[0389] Figure 3 is a graphic depiction of a chemical complex according to some embodiments. As indicated above, with reference to Figure 2, in the ODH process configuration depicted in Figure 3, although second reactor 206 is sShowm directly after quench tower or acetic acid scrubber 204, it can be placed further downstream. In many cases, the process configuration can be more energy' efficient if second reactor 206 is placed after the input stream has been compressed.

[0390] In some embodiments, the olefins produced using the one or more ODH reactors, or any of the processes or complexes described herein, can be used to make various olefin derivatives. Olefin derivatives include, but are not limited to polyethylene, polypropylene, ethylene oxide, propylene oxide, polyethylene oxide, polypropylene oxide, vinyl acetate, vinyl chloride, aciylic esters (e.g. methyl methacrylate), thermoplastic elastomers, thermoplastic olefins and blends and combinations thereof

[0391] In some embodiments, ethylene and optionally a-olefins are produced in the one or more ODH reactors, or any of the processes or complexes described herein, and are used to make polyethylene. The polyethylene made from the ethylene and optional a-olefrns descsibed herein can include homopolymers of ethylene, copolymers of ethylene and ct-olefins, resulting in HOPE, MDPE, LDPE, LLDPE, and VLDPE.

[0392] The polyethylene produced using the ethylene and optional a-olefins described herein can be produced using any suitable polymerization process and equipment. Suitable ethylene polymerization processes include, but are not limited to gas phase polyethylene processes, high pressure polyethylene processes, low pressure polyethylene processes, solution polyethylene processes, slurry polyethy lene processes and suitable combinations of the above arranged either in parallel or in series.

[0393] The present disclosure also contemplates use of various tools commonly used for chemical reactors, including flowmeters, compressors, valves, and setssors for measuring parameters such as temperature, pressure and flow rates. It is expected that the person of ordinary' skill in the art would include these components as deemed necessaiy for safe operation.

[0394] Figure 4 is a schematic of an experimental reactor unit (ERU) 400 as described in examples. The ERU 400 was used to produce feed gas for evaluating the catalysts. The ERU 400 consists of fixed bed tube reactor 402, which is surrounded by two-zone electric heater 404. Reactor 402 is a 316L stainless steel tube, which has an outside diameter of 0.5 inches (about 1.27 cm) and inside diameter of 0.4 inches (about 1 cm) and a length of 14.96 inches (about 38 cm). Two main feed gas lutes are attached to reactor 402. One tine 406 is dedicated for a bulk nitrogen purge gas and the other line 408 is coimected to a dual solenoid valve, which can be switched front ODH process feed gas (gas mixture of ethane/ oxygen/Nhrogen at a molar ratio of about 36/18/46) to compressed tor when regenerating catalyst bed 410.

[0395] For safety reasons the unit is programmed its a way that prevents air from mixing with the feed gas. This is accomplished through safety interlocks and a mandatory 15-minute nitrogen purge of the reactor when switching between feed gas 406 and air 412. The flow of gases is controlled by mass flow controllers. A 6-point thermocouple 414 is inserted through reactor 702, which is used to measure and control the temperature within catalyst bed 416. The catalyst is loaded in the middle zone of reactor 402 and located in between points 3 and 4 of thermocouple 414, which are the reaction temperature control points. The remaining 4 points of thermocouple 414 are used for monitoring purposes. Catalyst bed 416 consists of a one to one volume ratio of catalyst to quartz sand, a total of 3 ml. The rest of reactor 402, below and above catalyst bed 416 is packed with 100% quartz sand and the load is secured with glass wool on the top and the bottom of reactor 402. A glass tight sealed condenser 418 is located after reactor 402 at room te perature to collect water/acidic acid and the gas product can flow to either vent 420 or sampling loop/vent 422 by a three-way solenoid valve

[0396] As described herein, feeding a mixture of oxygen and hydrocarbons to an SO reactor may be problematic. The present techniques provide a process for the catalytic oxidative dehydrogenation of one or more C alkanes. The process includes two or more pre-reactors and one or more main downstream oxidative reactors. As described herein, the oxidative dehydrogenation of alkanes is performed in the presence of a mixed inetal oxide oxidative dehydrogenation catalyst system, which takes up oxygen in the catalyst. The process includes passing a feed stream that includes one or more C alkanes through one or tnore of n- 1 of the pre reactors at a temperature from 300 °C to 500 °C and a pressure from 3 447 kPag to 689.47 kPag (0.5 to 100 psig). The pre -reactors oxidatively dehydrogenate at least a portion of the feed stream until the oxidative dehydrogenation catalyst is depleted of reactive oxy gen. [0397] The feed stream from the pre-reactor(s) in which the oxidative dehydrogenation catalyst lias beets depleted of reactive oxygen is fed to a pre-reactor in which the oxidative dehydrogenation catalyst is substantially saturated with reactive oxygen.

[0398] The product stream from the n-1 pre-reactor(s), together with additional oxygen, is fed to one or more downstream reactors at a temperature from 300 °C to 500 °C and a pressure from 3.447 kPag to 689.47 kPag (0.5 to 100 psig) for the oxidative dehydrogenation of the one or more C alkanes. A product stream irons the one or more downstream reactors that includes corresponding CM alkenes, umeacted CM alkanes, unreacted oxygen arsd water [vapor] is passed through one or more pre-reactors depleted of reactive oxygen at a temperature from 50 °C to 300 °C and had a pressure from 3.447 kPag to 689.46 kPag to complex the oxygen in the product stream and increase the reactive oxygen saturation of the oxidative dehydrogenation catalyst and recovering a product stream substantially free of oxygen.

[0399] This is continued until either there is another pie-reactor or depleted of reactive oxygen than that through which the product stream is being passed, for the oxidative dehydrogenation catalyst in the pre-reactor is substantially complexed with reactive oxygen. At that point, the flow of the product stream is switched from the current pre-reactor to another pre-reactor that has a catalyst that is more reactive oxygen depleted. The formerly reactive oxygen depleted pre-reactor may include an SO catalyst that is completely saturated with reactive oxygen, and the formerly ox gen depleted pre-reactor may be brought back online. In various embodiments, the SO catalyst in any of the reactors is independently selected from any of the catalysts described above

[0409] One or more downstream reactors, termed main reactors, herein, may be supported on a surface of a permeable ceramic membrane defining at least part of the flow path for one reactant. In this configuration, the other reactant flows over the opposite surface of the ceramic to pennit the oxidant and etliane to react on the ceramic surface.

[0401] In some embodiments, the support is dried prior to use. To dry the support, the support may be heated at a temperature of at least 200 °C for up to 24 hours, typically, at a temperature from 500 °C to 800 °C for about 2 to 20 hours, preferably 4 to 10 hours. The resulting support will be free of adsorbed water and should have a surface hydroxyl content from about 0.1 to 5 mmol/g of support, preferably, from 0.5 to 3 mmol/g of support.

[0402] The dried support; may then be compressed into the required shape by compression molding. Depending on the particle size of the support, it may be combined with an inert binder to hold the shape of the compressed part.

[0493] The support for the catalyst may be a ceramic or ceramic precursor formed from oxides, dioxides, nitrides, carbides and phosphates selected from the group consisting of silicon dioxide, fused silicon dioxide, aluminum oxide, titanium dioxide, zirconium dioxide, thorium dioxide, lanthanum oxide, magnesium oxide, calcium oxide, barium oxide, tin oxide, cerium dioxide, zinc oxide, boron oxide, boron nitride, boron carbide, boron phosphate, zirconium phosphate, yttrium oxide, aluminum silicate, silicon nitride, silicon carbide and mixtures thereof. If necessaiy, the support could include a binder to help shape it.

[0404] Preferred components for forming ceramic membranes include oxides of titanium, zirconium, aluminum, magnesium, silicon and mixtures thereof. [0405] As noted above, the support in the main reactor should have a low surface area, preferably, less than 10 m ² /g, more preferably, less than 5 m ² /g, most preferably less than 3 m ² /g. The support may be prepared by compression molding. At higlier pressures, the interstices within the ceramic precursor being compressed collapse. Depending on the pressure exerted on the support precursor, the surface area of the support should be less than 15 m ² /g. The support will be porous and will have a pore volume from about 0.1 to 3.0 ml/g, typically, from 0.3 to 1.0 ml/g. The pore size of the ceramic may be small. Preferred pore size (diameter) ranges from about 3 to 10 run. The small pore diameter is helpful in the ceratnic membrane application as it helps maintain the pressure drop across the membrane so that a break in the membrane is readily detected by a sudden change its pressure. Additionally, the small pore diameter promotes a more uniform distribution of the reaction over the entire catalyzed surface of the membrane. That is, if larger pores are used, a majority of the oxygen tends to diffuse through the portion of the ceramic the oxygen containing gas initially cosnes in contact with. The remaining portion of the ceramic is largely unused.

[0406] The ceramic support may be prepared from the ceramic material using conventional techniques.

For example, the starting material may be cleaned, washed and dried (or spray dried) or produced from a sol/gel of the ceramic and where necessary' ground or milled to the appropriate particle size. The powder may be subjected to benefication, such as, acid or base washing to alter the pore size of the ceramic

[0407] The resulting powder is dried or calcined to remove associated water as noted above (water of hydration, etc.) and may be formed into a suitable substrate, preferably, tubular, by, for example, compression molding or isostatic compaction at pressures from about 5 to 200 MPag (725 to 29,000 psig), with or without a binder and sintering at temperatures to fuse the particles (e.g. at temperatures from about 0.5 to 0.75 of the melting temperature of the ceramic material)

[0408] Other techniques may be used, such as, tape casting or slip casting of slurries and the subsequent “punching of’ the required shape, such as, circular, square or annular, etc. For example, annular sections could be “stacked” to produce a “tube”.

[0409] While a tube is generally considered cylindrical, it could have any cross section shapes, such as, square, rectangular, hexagonal or stars, etc. In the case of a non-cylindrical tube, wall sections could be made by slip casting and then hermetically joining the wall sections together to form a central passage defined by an outer ceramic wall. The joints need to be hermetically sealed to prevent oxygen coming in contact with the ethane feed and forming an explosive mixture. Glass cement or a ceramic cement or slip would be used for this purpose. A hermetic seal also needs to be at the ends of the tube where it enters and exits the reactor or joins to the steel parts of the reactor.

[0410] In some embodiments, once the ceramic tube is prepared, the catalyst may he deposited on the surface of the tube, which will be in contact with the ethane.

[0 11] The ceramic membrane may have a thickness from about 0.1 to 10 cm, typically, from 1 to 8 cm, preferably, from 2 to 7 cm.

[0412] While ceramics are strong, they can be brittle. It is preferred to have a supporting structure at least on one side, preferably, the outside of the ceramic tube. Most preferably, there is a support structure on the outside and inside of the tube. The structure should be in the form of a mesh or a web having holes there through to permit the oxygen containing gas to pass through tise support and the ceramic to react at tise surface of the tube bearing the catalyst. The support may be any material suitable for use at the reactor operating temperatures. From a cost point of view, a steel mesh is likely most cost effective. Preferably, the steel is a stainless steel. The support structure should provide sufficient integrity to the tube to permit a shutdown of the reactor, if the ceramic is breached (e.g., becomes cracked, etc.).

[0413] One or more tubes are then placed inside the reactor. In one embodiment, the reactor is designed to have a plug flow of feedstock (e.g., primarily, ethane) through a passage between the reactor shell and the ceramic tube and a How of oxygen containing gas through the ceramic tube. There are a number of arrangements that come to mind. The reactor could comprise several shorter tubes placed end to end to provide a tube of appropriate length Alternatively, the design could be similar to a core shell heat exchanger with a number of parallel tubes through which the oxygen containing gas is passed with and an enclosed shell providing a passage between the external wall of the reactor and the ceramic tubes defining a flow' path for the ethane. The flow paths might be reversed (ethane on the interior and oxy gen on the exterior of the tube).

[0414] In one embodiment of the techniques, the catalyst in the main reactor is on a ceramic membrane and in the case of the pre-reactor, on a high surface particulate support as described below.

[0415] Figure 5 is a schematic diagram of an apparatus and process flow for providing mixing of reagents for an inherently safe process. In an embodiment, the catalyst in the main reactor is in the form of granular beds having a low surface area as described in association with Figure 5. Alkane feedstock, preferably, ethane, flows through a line 502 to a valve set 504 and through line 506 to pre-reactor 508, one of a pair of pre-reactors 508 and 510. In pre-reactors 508 and 510 there is a single fixed bed of catalyst, which is not shown. The bed is held in place between two porous membranes, or open metallic meshes of a small enough mesh size so that the particles ill not pass out of the bed, which are not shown. The feed passes through pre-reactor 508 and is partially oxidatively dehydrogenated and the catalyst bed is depleted of reactive oxygen. The partially dehydrogenated feed passes through exit line 512 from the pre-reactor 508 to another valve set 514. The partially dehydrogenated feed flows from valve set 514 via a feed line 516 to the top of the downstream reactor 518. The oxidant, typically air, oxygen, or a mixture of oxy gen and an inert gas such as nitrogen or argon flows through a line 520 and enters the feed line 516 near the top of reactor 518. The mixed feed of oxygen and parbally dehydrogenated feed flow through three fixed beds of catalyst 522, 524, and 526. There is a space between the catalyst beds and additional oxygen is fed via lines 528 and 530 into the space between the catalyst beds. The substantially dehydrogenated product stream containing small amounts of oxygen (typically less than 5 vol. %, preferably less than 3 vol %) is fed via line 532 to valve set 514. The dehydrogenated feed passes through valve set 514 via line 534 to pre- reactor 510, which is depleted or substantially depleted of reactive oxygen. As the product stream passes through pre -reactor 510 oxygen is extracted from it and the oxidative dehydrogenation catalyst becomes more saturated with reactive oxygen. The product stream substantially depleted of oxy gen is fed via line 536 to valve set 504. The product passes from valve set 504 to line 538 for recovery and further processing.

[0416] When pre-reactor 508 is depleted of reactive oxygen then valve sets 504 and 514 are switched so that alkane feed is fed to pre-reactor 510 and product from reactor 518 is fed to pre-reactor 508 so that it becomes more saturated (charged) with reactive oxygen

[0417] In some instances (e.g. on startup) oxygen may be fed to the pie-reactor containing the catalyst supported on a support having a high surface area topically greater than 100ni2/g, preferably greater than 150ni2/g a to “charge” it with oxygen. This is snore to balance the reaction times between various pre-reactors so that a pre-reactor dehydrogenating feedstock will have a sufficiently long operation to permit full “charging” of a pre-reactor depleted of reactive oxy gen.

[0418] It is important to minimize the potential for oxidizing product stream from line 532 and producing one or more of carbon monoxide and carbon dioxide. Such an oxidation consumes valuable feed and product stocks, introduces undesirable byproducts, and reduces the conversion and selectivity of the process. To minimize the undesirable further oxidation of the feed and product it is important that temperatures in the pre reactor when adsorbing oxygen from the product stream (e.g. chemisorption) is kept below the temperature for oxidative dehydrogenation (e.g. from 50 °C to about 300 °C, preferably less than 270 °C). In the pre-reactor during the chemisorption or oxygen scavenging process from the product stream the temperature should be below about 270 °C, preferably from 50 °C up to about 270 ^C 'C typically from 100 °C to 250 °C. In view of the temperature difference between the pre-reactors its oxidative dehydrogenation mode and chemisorption or oxygen scavenging mode it may necessary to cool the feed to the pre-reactor to be used for chemisorption or oxygen scavenging to an appropriate temperature before entering the pre-reactor in chemisorption or oxygen scavenging mode. Hence, it is preferable to have several pie-reactors to permit the pre-reactor time to cool prior to putting it into service for scavenging. Oxygen scavenging is exothermic and the reactor will heat up and depending on the catalyst system, oxygen release could be exothermic so the heat requirements may not be to significant (e g. near neutral)

[0419] Figure 6 is a schematic diagram of a main reactor 600 in which the oxidative dehydrogenation takes place on the surface of ceramic membrane tabes 602, which are functioning as catalytic membranes. The reactor comprises an inlet 604 into which a stream of ethane or an ethane containing gas 606 flows. The ethane 606 passes through the cerasnic membrane tubes 602 to a collector 608 A stream of oxygen or an oxygen containing gas, 610 is fed to the tube bundle so the oxygen is on the outside of the tabes. The ethane or ethane containing gas 606 reacts with the oxygen as it passes down the tube to form ethylene. The ethylene is collected in the collector (footer) 608 and exits the reactor at 612.

[0429] Figure 7 is a schematic diagram 700 of a main reactor 702 in which the oxidative dehydrogenation takes place on the surface of ceramic tubes or membranes. In the schematic diagram 700, the ethane, or ethane containing gas, 704 enters the main reactor 702 through an inlet or header 706. The oxygen or oxygen containing gas 708 enters a tube and shell t e plate shown at 710, which includes a series of ceramic membrane tabes 712 encased in a steel shell 714. The ceramic membrane tubes 712 extend up to the header 706. As a result, the ethane or ethane containing gas 704 flows down the interior of the ceramic membrane tabes and the oxygen flows down the annular space between the exterior of each of the ceramic membrane tubes 712 and the steel shell 714. The ethane is converted to ethylene, or other products, and exits the ceramic membrane tabes into collector (footer) 716 and exits at 718. One advantage of this design is if a ceramic membrane loses integrity, excess oxygen only enters that tube. This is easily detected by an oxygen detector (not shown) which may be at the exit of each of the ceramic membrane tubes 712 or in the collector 716. Then the reactor can be safely shut down and the damaged tube may be located.

[9421] The flows of the reactants may be co-current or counter current (e.g., ethane up the outside of the tube and oxygen down the inside of the tube) The feed to the reactor comprises two separate flows to opposite sides of a tube. In one further embodiment, one flow, preferably, to the internal surface of the tube is an oxygen containing gas, which is selected front the group consisting of oxygen, mixtures comprising from 100 to 2 i vol. % of oxygen and from 0 to 79 vol. % of one or more inert gases. Some inert gases may be selected from the group consisting of nitrogen, helium and argon and mixtures thereof. The oxygen containing gas could be air. [0422] Oxygen Scavenging

[0423] The amount of oxygen that is entrained in the product ethylene stream should be minimized for further processing. However, there will likely be some small amount of oxygen in the product stream. It is highly desirable that the oxygen be removed from the product stream prior to further processing of the product stream. Immediately downstream of the oxidative dehydrogenation reactor may be a low temperature (below about 270 °C ) pre-reactor in which the oxidative dehydrogenation catalyst has a reduced reactive oxygen content to take up residual oxygen from the product stream without oxidizing more than about 5 wt %, preferably less than I wt. % of the ethy lene produced. The low temperature oxygen scavenging reactor operates at temperatures less than or equal to 300 °C, typically from 50 ^C 'C to 300 °C, more preferably from 50 °C to 270 °C, generally from 50 °C to 270 °C, preferably form 50 °C to 250 °C, desirably form 100 °C to 250 °C.

[0424] In operation, it will be necessary to balance the oxygen feed to the main reactor depending on the conversion in pre-reactor 508 (Figure 5). There may be several “pre-reactors” also used as scavengers to accommodate the product flow' out of the main reactor. It may not be so much of an issue with the pre-reactor operating in oxidative dehydrogenation mode since any excess alkane not dehydrogenated in the pre-reactor will be converted in the mam reactor(s). The key issue is the scavenging of oxygen from the product stream.

[0425] Preferably, at the exit of the main oxidative dehydrogenation reactor is art oxygen sensor

Additionally, there should be an oxygen sensor at the exit for the dehydrogenated product from each pre-reactor to determine the oxy gen level leaving the process chairs. When the oxygen level rises at the dehydrogenated product outlet of the pre-reactor (i.e. scavenger reactor) it indicates the catalyst have substantially taken up reactive oxygen (and may be returned to use as a pre-reactor). The amount of reactive oxygen uptake by the oxygen depleted catal st in the pre-reactor operation in oxygen scavenging or chemisorption mode should be not less than about 1.5 wt. %, typically about 2 wt. % of the total oxygen in the catalyst (this will also correspond to the amount of reactive oxygen available for release from the catalyst in the pre-reactors in oxidative dehydrogenation mode)

[0426] Figures 8, 9, and 10 illustrate how a series of three fixed bed catalysts may be used to scavenge oxygen from the product stream in an oxidative dehydrogenation reactor. Each of the reactors are labeled, with the first reactor labeled ‘A’, a second reactor labeled Έ’, and a third reactor labeled ‘C’. Accordingly, the numerical labels in the Figures 8, 9, and 10 indicate the functional name of the individual reactors. Further, the results are shown in Table 1. In these figures, the valves used to switch the functionality of the individual reactors, and described with respect to the previous figures, are not shown. The main reactor is the same throughout Figures 8, 9, and 10. However, the switching of the valves causes the pre-reactor, scavenger reactor and the guard reactor to appear to “switch” places. One pre-reactor operates as such and converts part of the feed stream to ethylene. One 0x5' gen depleted pre-reactor acts as a primary oxygen scavenger or chemisorption reactor and a second pre-reactor (also the oxygen depleted pre-reactor acts as a guard or secondary oxygen scavenger or chemisorption reactor). [0427] Operation

[0428] Process Step Process streams flow sequence

[0429] Step 1 : As shown in Figure 8, in a first step or operation, ethane 802 (50) is routed to a pre-reactor 8Q4 (A in this example). The pre-reactor 804 is preferably oxygen saturated. Some of the ethane is converted to ethylene and the product together with oxygen 806 is routed to the main reactor 808, where most or all ethane is converted to ethylene, or other products. The product is cooled in a condenser 810 to a temperature from 50 °C to 270 °C Optionally, water is knocked out of the product stream m knock out drum 812 (55), or is adsorbed by one or more guard beds. The cooled product stream is routed to a primary oxygen depleted pre-reactor (B in the example of Figure 8) which acts as a lead oxygen scavenger reactor 814 As the oxygen scavenging, or chemisorption reaction, is exothermic, the product stream from the primary oxygen scavenging reactor may be cooled in a condenser 816 and routed through a water knock out drum 818 (58), prior to being fed to the secondary or guard oxy gen scavenger reactor 820 (oxygen depleted pre-reactor C in the example of Figure 8). In some embodiments, cooling is not used, since the only reason for cooling is to reduce oxidation reaction of the final product 822, producing CO or CO or both. As the secondary' or guard oxygen depleted pre-reactor 820 initially has a very low' level of reactive oxygen (typically less about 50 ppm, or less than about 25 ppm, or less than about 10 ppm of reactive oxygen in the feed stream) this may not be an issue. Further, a slightly elevated temperature (2 °C to 5 °C higher) may help to remove oxygen to very low' level without converting the product to CO and CO . Oxygen sensors, not shown, are active on inlets to the lead oxygen scavenger reactor 814 and guard oxygen scavenger reactor 820 and on the outlet of the guard oxygen scavenger reactor 820 The next operation or step is started when the oxygen content in the product stream exiting the guard oxygen scavenger reactor 820 exceeds a specified value.

[0439] Step 2 : Oranges from Step 1 : In the second operation or step, shown in Figure 9, the former pre reactor (reactor A) now becomes guard scavenger 820. The former guard scavenger (reactor C) now becomes lead oxygen scavenger reactor 814. Finally, the former lead scavenger (reactor B) becomes the pre-reactor 804. After the reactors me switched, the operation is the same as described for the previous operation or step. Once an increased oxygen content is measured, the next operation step is started

[0431 ] Step 3 : Changes from Step 2 (Figure 10): in the third operation or step, the pre-reactor (reactor B) becomes the guard oxygen scavenger reactor 820. The former guard scavenger (reactor A) becomes lead scavenger reactor 814. The former lead scavenger (reactor C) becomes the pre-reactor 804 Operation is the same as described for the Step 1 .

[0432] Once an increased oxygen content is measured, operations returns to the initial operation or step, described with respect Figure 8.

[0433] EXAMPLES

[0434] Scavenging (Post -Removal) of Residual Oxygen From the Ethylene Product Gas Mixture by the Periodical Redox Cycle

[0435] Scavenging of residual oxygen from the product (outgoing gas) mixture was realized by cyclical periodical redox operation mode In this case, a Moi-Viu-Nbo _.j -Teo i-Ox catalyst or other Oxygen Storage Material (OSM) can be used in the two-step process. Step 1 provides the reduction of the OSM layer by pure ethane at temperatures -TOO °C, and Step 2 supports absorptive removal of the residual O from the outgoing product mixture by the pre- reduced layer working as an OSM at a reduced temperature. It was shown that the Moi-Vo ₃ -Nbo ₂ -Teo i-Ox catalyst itself served as a rather effective OSM at 300-400 °C.

[0436] a) Step 1 : reduction of the M01-Vo.3-Nbo.2-Teo.1-Qx catalyst iayer by pure ethane [0437] Measurements w ^r ere done using a fresh sample [20 wt. % Mo-V-Te-Nb-O _x + 80 wt. % T O (support)] prepared by mechanical methods (grinding and compaction/extrusion). In tins testing, the sample (2.0 cm ³ ; 2.97 g, particle size 0.2-0.4 mm) was placed into a quartz reactor and heated to a specified temperature (375° and 400 °C) in an air flow, for 15 nun, then the gas flow (900 crrVh) was switched to pure ethane, and a probe of the outgoing mixture was taken for analysis after a given time. After reoxidation of the sample by air for 15 min, measurements were repeated several times with varying the time interval, and resulting response curves of products were obtained (up to 7.5 min). Figures 11, 12, and 13 demonstrate the time dependence of ethylene and Ci¾ formation rates as well as the selectivity of ethylene formation upon the catalyst reduction by pure ethane at two different temperatures. The CO formation curves are quite similar to those observed for CO (Figure 5).

[0438[ Figure 11 shows the reaction profile (Dynamics) of ethylene formation as a function of time at 375 °C and 400 °C after the gas flow switch (air — > ethane) for a Moi-Vo -Nbo 2-Teo.i-Ox catalyst with 80% of T1O2 as support.

[0439] Figure 12 shows the reaction profile (Dynamics) of CO formation as a function of time at 375 °C and 400 °C after the gas flow switch (air — ethane) for a Moi-Vo.3-Nbo 2-Teo.x-Ox catalyst with 80% of T O as a support.

[0440] Figure 13 shows the selectivity of ethy lene formation of time at 375 °C and 400 °C after the gas flow switch (air -» ethane) for a M -Vo -Nbo -Teo -Ox catalyst with 80 wt. % of T O as a support. Thus, the step of the catalyst reduction by pure ethane at 380-400 °C is accompanied by the formation of ethylene with a selectivity >92% (Figure 13).

[0441] The results obtained permit one to calculate the total amount of the “reactive” latice oxygen in the catalyst working as an OSM. Integration of the response curves (Figures 11 and 12) being produced at a constant ethane flow rate of 37.5 mmol/h may be used evaluate the overall amount of oxygen reacted during the catalyst reduction step. Taking into account that 1 g of the mixed oxide (M -Vo -Nbo -Teo -Ox contains about 333 ing of oxygen, it can be concluded that about 1.9 wt. % of the oxygen from this amount (about 6.3 iiig/g) can be removed from the active phase by reduction. Thus, the oxygen storage (absorption) capacity upon the subsequent reoxidation step cannot exceed this number.

[9442] (b) Step 2: absorption of the residual O2 from the outgoing mixture by the pre-reduced layer w'orking as an oxygen-storage material at a reduced temperature

[9443] In this testing, the sample of catalyst after reduction by pure ethane at 400 °C for 15 mitt was cooled to a given temperature (270 °C) in the ethane flow', then model product gas flow ([49.5 vol. % C- ₆ + 46.7 vol. % C2H4 + 3.8 vol. % O2 + CX -traces] ; 720 cmVii) was switched on, and the probe of the outgoing mixture was taken for analysis lifter a given time. After subsequent reduction of the sample by ethane (15 nun, 40Q °C), measurements were repeated several times with varying the time interval, and the resulting response curves of products w ere produced (up to 5 min). The same testing w ^r as repeated at 40Q °C for comparison. Figures 7 and 8 demonstrate the time dependence of the total O removal, as well as the variation of the CO and CO2 concentrations in the product flow from the scavenging reaction flow at 270 °C and 400 °C. [0444] Figure 14 shows the reaction profile (Dynamics) of O _? removal from the model gas mixture by the pro-reduced catalyst of (M0 ₁ -Vo ₃ -Nbo ₂ -Teo ₁ -Ox supported on 80% of Ti(¼ at 270 °C and 400 °C.

[0445] Figures 15 and 16 show the reaction profile (Dynamics) of CO ₂ and CO formation after feeding the model gas mixture by the pro-reduced catalyst of (M0 ₁ -Vo ₃ -Nbo ₂ -Teo ₁ -Ox supported on 80% of T1O2 at 270 °C and 400 °C.

[0446] Accordingly, both CO and CO2 are fonned by catalytic oxidation at 400 °C, even on the pre- reduced catalyst. After the temperature is reduced. For example, at 270 °C the absorptive O2 removal becomes the tnain process, with a minor contribution of both CO and CO2 formation.

[0447] This mode of operation is beneficial, because the efficiency of most adsorption / chemisorption processes is limited by mass transfer front or zone (MTZ). As a result, a significant part of the oxy gen scavenging materia] remains not saturated with oxygen, and consequentially once the scavengers is switched into the pre-reactor mode, the pre-reactor will have shorter run time compared to the fully oxygen saturated pre reactor (on start-up). Having lead and guard scavengers permits a better take up of oxygen in the lead scavenger chemisorption reactor (pre-reactor). This option also gives the benefit of having an oxygen sensor between lead and guard scavengers and to having the option to switch the operation exactly at the point in time when the lead is fully saturated or to keep it on stream slightly longer, if there is a process upset of any nature requiring longer operation without switching. Another benefit of this option is that lead scavenger has to be significantly colder than She main reactor; it is so avoid ethylene oxidation reaction from occurring. The guard scavenger should preferably be hotter than the lead scavenger, since most of the oxygen is removed and to remove trace oxygen higher temperature is beneficial. Oxidation of the product stream in the guard scavenger is not expected to occur to any significant extend, since only traces of oxygen are present. Because of the operation as described above, when pre-reactor (converter) switches to be a guard scavenger, it is still hot, which is very beneficial for the guard, when guard switched to be the lead, it is already cooled with ethylene product in the most efficient way, by direct contact of ethylene product with the surface of the catalyst.

[0448] The resulting product is then passed down stream for further separation if necessary'. The separation requirements are minimized in the present reaction as the catalyst in the main reactor has a selectivity above 95% preferably above 98% and no or a minimum amount of bi products are produced in the oxygen scavenging step. The product can be sent directly to polymerization plant or other ethylene derivatives plants, (such as ethylene glycol, acetic acid, vinyl acetate, etc.) as they can utilize ethylene of a lower purity, alternately only CO, CO2 may be separated or CO2 only, if needed. However as noted above it is preferable to operate the pre-reactor in chemisorption or oxy gen scavenging mode at a temperature to minimized further generation of carbon dioxide, carbon monoxide or both.

[0449] IND U STRIAL APPLIC AB ILITY

[0450] The present techniques seeks to improve the conversion and selectivity of oxidative dehydrogenation reactors by providing scavenger beds downstream of the reaction to remove residual oxygen from the product stream stopping conversion of desired product(s) to products such as CO and CO2.

[0451] This section relates to an apparatus and method for the safe mixing of gases More specifically, the section relates to a method for the safe mixing of a hydrocarbon, such as ethane, with an oxidant, such as oxygen, in a manner that risk of a process disruption is minimized. Additionally, a gas mixer suitable for use with the method is described. The method and the gas mixer are applicable for use in oxidative processes such as the catalytic oxidative dehydrogenation of ethane into ethylene.

[0452] In one aspect, a method of mixing a hydrocarbon-containing gas with an oxidant-containing gas is provided. The method comprises introducing a first gas and optionally a first inert diluent into a first pre-mix zone flooded with a non-flammable liquid to form a first saturated gas. The method comprises introducing a second gas and optionally a second inert diluent into a second pie-mix zone flooded with the non-flammable liquid to form a second saturated gas. The method comprises mixing the first saturated gas and the second saturated gas in a common mixing zone partially flooded with the non-flammable liquid and comprising a flooded region and a headspace, to form a homogeneous gas mixture comprising the first gas, optionally the first inert diluent, the second gas, optionally the second inert diluent, and non-flammable liquid, that is outside the flammability limit, whereby the term homogeneous is not limited to one phase mixtures, but rather includes aerosols and or suspensions and refers to even distribution of all the phases in the mixture through the process equipment space. The method comprises recovering the homogeneous gas mixture from the headspace. The first gas and the second gas are different and me selected from a hydrocarbon-containing gas and an oxidant- containing gas.

[0453] Further, in accordance with another aspect, a gas mixer for mixing a hydrocarbon-containing gas with an oxidant Containing gas is provided. The gas mixer comprises a first pre-mix zone flooded with a non flammable liquid and configured to receive a firs! gas from a first gas supply line and form a first saturated gas. The gas mixer comprises a second pre-mix zone, physically separated from the first pre-mix zone, flooded with a non-flammable liquid, and configured to receive a second gas from a second gas supply line and form a second saturated gas. The gas mixer comprises a common mixing zone partially flooded with a non-flammable liquid resulting in a headspace and a flooded region, the flooded region separately fluidly connected to the first pre-mix zone and second pre-mix zone and configured to separately receive the first saturated gas and the second saturated gas. The gas mixer comprises a homogeneous gas mixture outlet for removing from the headspace a homogeneous gas mixture comprising the first gas and the second gas.

[0454] Techniques are provided for a method for mixing a hydrocarbon-containing gas with at! oxidant- containing gas and a gas mixer apparatus useful for the method described.

[0455] Techniques are provided for applications that require the mixing of a hydrocarbon-containing gas with an oxidant-containing gas. It is well known that gaseous compositions containing a hydrocarbon and oxygen in ratios that fall within the flammability envelope are potentially hazardous. An ignition event, such as a spark, can ignite the mixture and potentially lead to an explosion. While applications that require mixing of hydrocarbons and oxygen normally do so with ratios that are safe and not susceptible to ignition there are moments during initial mixing, where heterogeneous pockets of unfavorable hydrocarbon/oxygen compositions exist and may ignite if a spark occurs.

[0456] Techniques are provided for a method for mixing a hydrocarbon-containing gas with an oxidant- containing gas that is simple and safe in that ignition events are unlikely to occur. The method comprises introducing a hydrocarbon-containing gas into a first pre-mix zone, introducing at! oxidant-containing gas into a second pre-mix zone, the first pre-mix zone and the second pre-mix zone flooded with a non-flammable liquid. The hydrocarbon-containing gas and the oxidant-containing gas form bubbles within the first pre-mix zone and second pre-mix zone, respectively, that travel through the respective pre-mix zone, becoming saturated with the non-flammable liquid before entering a common mixing zone that is partially flooded with the non-flammable liquid mid comprises a flooded region and a head space. Within the flooded region, bubbles of the hydrocarbon- containing gas saturated with non-flammable liquid mix with bubbles of the oxidant -containing gas saturated with non-flammable liquid, to form bubbles of a homogeneous gas mixture of the hydrocarbon-containing gas and the oxidant-containing gas, the homogeneous gas mixture saturated with the non-flammable liquid. The bubbles of the homogeneous gas mixture eats exit the flooded region into the headspace where the homogeneous .gas mixture can be removed for use downstream in a process that requires a mixture of a hydrocarbon with an oxidant. The term homogeneous in reference to the gas mixture is intended to mean that the components of the gas, including aerosol components if present, are generally e venly dispersed throughout.

[0457] As the term suggests, the non-flammable liquid used with the method is not be flammable in the process. Not flammable or non-flammable in this context means that the material cannot be ignited under the process conditions described herein. Suitable non-flammable liquids imder the process conditions include water, ethylene glycol, silicon oils, and carbon tetrachloride. The non-flammable liquid must also be non-reactive with either the hydrocarbon or the oxidant. In an embodiment, the non-flammable liquid is water.

[0458] While any non-flammable liquid may be used with the present techniques, it is important to consider that the homogeneous .gas mixture removed from the headspace will comprise the hydrocarbon- containing gas, oxidant-conlaining gas, and some carry-over of non-flammable liquid in vapor form, and inert diluent if used. For this reason, selection of a non-flammable liquid must consider any potential effects the carry-over of non-flammable liquid may have on downstream applications. For example, some catalysts used in ethane ODH are sensitive to water (in liquid form) atsd therefore the water is preferably removed from the homogeneous gas mixture before use in ethane ODH Alternatively, the homogeneous gas mixture is heated to a temperature where the water present is converted to steam before use in ethane ODH. In some embodiments, the non-flammable liquid is water and the temperature of the homogeneous gas mixture is raised to at least 100 °C, preferably to a temperature of at least 150 °C, but not exceeding 300 °C, preferably not exceeding 250 °C, before use in ethane ODH. It is contemplated that non-flammable liquid within the homogeneous gas mixture is captured and recycled back to one or more of the first pre-mix zone, the second pre-mix zone, and the common mixing zone

[0459] In certain embodiments, the non-flammable liquid is wnter and the homogeneous gas mixture removed from the headspace is subjected to a separation step to, at least partially, remove the non-flammable liquid from the homogeneous gas mixture.

[0460] In certain embodiments, the non-flammable liquid is water and the homogeneous gas mixture is passed through a condenser to condense and subsequently separate water present in the homogeneous gas mixture from the hydrocarbon and the oxidant, the separated water recycled back to one or more of the first pre mix zone, the second pre-mix zone, and the common mixing zone.

[0461] In working these techniques, the amount of the gases introduced into the pre-mix zones results in a homogeneous gas mixture that comprises a ratio of hydrocarbon-containing gas to oxidant-containing gas that is preferably outside of the flammability envelope. The chosen ratio will depend on the nature of the gases and the application for w hich the mixture will be used. For example, for an ethane ODH application, the ratio of ethane to oxygen chosen will depend on whether under the proposed ODH reaction conditions the ratio is above the higher explosive limit or below the lower explosive limit. In comparison, an application that requires or includes the addition of ethylene, the ratio of ethylene to oxygen added to the reactor would be different because ethylene is more reactive than ethane. Farther, the flammability limit of ethylene is higher than that of ethane. The temperature of the ODH process to be employed must also be taken into consideration as higher temperatures correspond to a much smaller window' of safe ratios of ethane to oxygen. For example, at temperatures below 300 °C and pressures below 450 kPag a molar ratio of about 80:20 ethane to oxygen for cattily tic ODH would fall above the upper explosive limit, while a ratio of about 1.5:98.5 ethane to oxygen would fall below the lower explosive limit, with each ratio safe enough in that ignition events would not lead to an explosion or flame propagation under ODH reaction conditions. Ratios falling between that, for example, 50:50, may be flammable or explosive, posing a rick of process disruptions.

[0462] In certain embodiments, the hydrocarbon-containing gas comprises ethane.

[0463] In processes such as ODH of ethane, the conversion rates ofte do not exceed 70%, leaving unconverted ethane in the product stream. The unconverted ethane can be isolated from the product stream, using distillation for example, and recycled back to be used in the ODH process. In certain embodiments, the hydrocarbon-containing gas comprises ethane recycled from an oxidative process. In certain embodiments, the hydrocarbon-containing gas comprises ethane recycled from an ODH process.

[0464] In certain embodiments, the oxidant-containing gas comprises oxygen.

[0465] In certain embodiments, the oxidant-containing gas comprises air.

[0466] Techniques are provided for the use of inert diluents for the hydrocarbon-containing gas and the oxidant-containing gas. As used herein, art inert diluent does not participate in a combustion reaction with the hydrocarbon-containing gas or the oxidant-containing gas. However, so e inert diluents, such as steam or carbon dioxide, may participate in the reactions in the present of the catalyst. One or both of the hydrocarbon- containing gas and the oxidant-containing gas may he diluted with an inert diluent prior to entering or within each of the respective pre-mix zone. Use of inert diluents when mixing an oxidant with a hydrocarbon is well known. Example diluents include carbon dioxide, steam, nitrogen, helium, and argon. The use of inert diluents has an impact on the flammability' ratio, a property of mixed gases that Mis within the knowledge of the person skilled in the art Both the hydrocarbon-containing gas and the oxidant-containing gas may be diluted with the saute or different inert diluents or mixture of different diluents

[9467] Another consideration is determining the flow rate at which each gas is added to the respective pre-mix zones. The flow rate of the gases and the corresponding pressure ithi the gas supply line would need to be higher than the pressure of the non-flammable liquid. In the absence of a pressure differential, the gases cannot enter the non-flammable liquid and consequently the pre-mix zone. Furthermore, if the pressure of the non-flammable liquid is higher than the line containing the gas to he introduced there may be, in the absence of a back-flow prevention mechanism, back flow of non-flammable liquid into the gas supply lines. This should be avoided. Back-flow prevention may include methods known in the art, including, but not limited to, use of one way valves, check valves, pressure controlling devices, and knock-out/flush vessels on the gas supply line equipped with a level indicator alarm.

[9468] When determining flow rates, the skilled worker must correlate the flow rates with the pressure and temperature used within the gas mixer The conditions within the gas mixer are chosen to reflect the amount of canyover of non-flammable liquid into the homogeneous gas mixture removed from the headspace. The flow rate of the incoming gases must be sufficient to allow entry into the non-flammable liquid at the predetermined temperature and pressure. Calculating flow rates and residence time for a gas mixer with particular dimensions would fall within the knowledge of the person skilled in the art.

[0469] In some embodiments, the gas mixer comprises a first pre-mix zone flooded with a non-flammable liquid, a second pre-mix zone physically separated from the first pre-mix zone and flooded with the non flammable liquid, and a common mixing zone partially flooded with the non-flammable liquid to form a flooded region atsd a head space, wherein the first pre-mix zone and second pre-mix zone are separately fluidly connected to the flooded region. The first pre-mix zone is configured to receive a first gas from a first gas supply line and the second pre-mix zone is configured to receive a second gas from a second gas supply line. First gas received in the first pre-mix zone atsd second gas received its the second pre-mix zotse may form bubbles within the non-flammable liquid within the respective zones. A small fraction of the non-flammable liquid may convert to a gaseous state and enter into the bubbles of the first gas and second gas within the pre mix zones. As the bubbles pass through the first pre-mix zone and the second pre-mix zone, they may become saturated with the non-flammable liquid, the gaseous non-flammable liquid diffusing within the bubbles, resulting in the formation of bubbles of a first saturated gas mixture in the first pre-mix zone and bubbles of a second saturated gas mixture in the second pre-mix zone. Upon entering the flooded region of the common mixing zone, the bubbles of the first saturated gas mixture combine with bubbles of the second saturated gas mixture to form bubbles of homogeneous gas mixture comprising first gas atsd second gas, saturated with non flammable liquid. The bubbles of the homogeneous gas mixture maty pass into the headspace for removal.

[0470] The first gas and second gas are different and are selected from a hydrocarbon-containing gas and an oxidant-containing gas. Preferably, the ratio of hydrocarbon to oxidant in the homogeneous gas mixture falls outside of the flammability limit. Use technique allows for safe mixing of a hydrocarbon with an oxidant, which is beneficial for oxidative processes that require these mixtures.

[0471] As mentioned, the non-flammable liquid used within the gas mixer must not be flammable.

Suitable non-flammable liquids include water, ethylene glycol, silicon oils, and carbon tetrachloride. The non flammable liquid should also be non-reactive with the hydrocarbon or the oxidant. In a preferred embodiment, the non-flammable liquid comprises water.

[0472] Reference to a first gas and a second gas is intended to indicate that the pre-mix zones in which the gases are introduced are not specific for either a hydrocarbon-containing gas or an oxidant-containing gas. That is, when the first gas is a hydrocarbon-containing gas the second gas is an oxidant-containing gas. Conversely, whe the first gas is an oxidant-containing gas the second gas is a hydrocarbon-containing gas

[0473] The first pre-mix zone and the second pre-mix zone are configured to receive a first gas and a second gas, respectively, which enter the non-flammable liquid as bubbles of gas at a front end of the respective zone, and to produce bubbles of gas saturated with non-flammable liquid at a back end which is fluidly connected to the flooded region of the common mixing zone. Use of the term saturation refers to the interior of the bubbles of the gas and the degree to which the gaseous form of the non-flammable liquid enters the interior of the bubbles during the passage from the front end to the back end of the respective zones. The term front end refers to the position within the first pre-mix zone and the second pre-mix zone that is at or near the most distant location relative to the flooded region. The back end refers the position in the first pre-mix zone and the second pre-mix zone that fluidly connects with the flooded region. The dimensions of the first pre-mix zone and the second pre-mix zone are configured so that under operating conditions the bubbles are close to being or are completely saturated with non-flammable liquid prior to reaching the flooded regiots of the common mixing zone. Diffusion rates of non-flammable liquid into the bubbles vary according to temperature, pressure, and the size of the bubbles. Higher temperatures and pressures size are associated with higher rates of diffusion of non flammable liquid into the bubbles in comparison to lower temperatures and pressures. In addition, smaller bubbles can reach saturation sooner than larger bubbles. It is contemplated to include mechanisms for limiting the size of the bubbles, including the use of static mixers or by placing one or more mesh screens along the path from the front end to the back end of the corresponding pre-mix zone. These mechanisms promote the separation of large bubbles into smaller bubbles and selection and configuration of a particular mechanism can limit bubble to a predetermined size. For example, a No.18 mesh provides an opening of approximately 1 mm, which would theoretically limit bubbles to around 1 mm in diameter. Choosing a size means considering the size of the pre-mix zones, and how long bubbles take to pass from the front end to the back end. Bubbles should ideally be less than 1 mm in diameter, preferably below' 0.5mm in diameter, more preferably below 0.1 mm in diameter.

[0474] The path for bubbles within the first pre-mix zone and second pre-mix zone may take any form, including a straight horizontal or vertical path, a spiral path, or a combination. The first pre-mix zone and second pre-mix zones need not be similar. In certain embodiments, the first pre-mix zone and the second pro- mix zone are vertical and fluidly connected to the bottom of a vertical common mixing zone, the first pro -mix zone and second pre-mix zone extending downwardly with respect to the common mixing zone. In certain embodiments, one or both of the first pre-mix zone and the second pre-mix zone comprise a pipe with a front end that is vertically above the flooded regiots of the common mixing zone. In certain embodiments, one or both of the first pte-mix zone and the second pre-mix zone are horizontally oriented atsd extend away from the common mixing zone.

[0475] Movement of the bubbles through the first pre-mix zone, the second pre-mix zone, and the flooded region may occur by gravity. Bubbles of gas naturally rise vertically through the higher density non-flammable liquid. Movement may also be encouraged using pressure and or other mechanisms known in the art for allowing flow of gases within a liquid medium.

[0476] For the present techniques, there is a requirement for a means of introducing gases into the pre-mix zones and non-flammable liquid into one or more of the first pre-mix zone, the second pre-mix zone, and the common tnixs ng zone. Any means known in the art may be employed including but not limited to a nozzle, spout, a valve, or a sparger. For simplicity, use of the terns nozzle within the description in relation to the introduction of gases refers to the point where contact between the gases and the non-flammable liquid first occurs, and the means for introduction includes any means knowm within the art. Use of the term nozzle in relation to the introduction of non-flammable liquid refers to the point where the non-flammable liquid enters any of the first pre-mix zone, the second pre-mix zone, and the common mixing zone. Connection of a nozzle to a gas supply line would be required for both the first and second gas. It is contemplated that the gas supply line for either gas cats include fresh gas or recycled gas obtained from a process downstream of the gas mixer. For example, unreacted ethane separated downstream of an oxidative dehydrogenation process can be directed back to the gas mixer for use via a rec cle line that connects with either the first or seeotsd gas supply line, depending upon which line is sised to introdsice the hydrocarbon-containing gas. [0477] In certain embodiments, at least one of the first pre-mix zone and the second pre-mix zone are configured to receive the first gas and second gas, respectively, via a nozzle comprising a cylindrical or round spout at the end of a gas supply line.

[0478] In certain embodiments, at least one of the first pre-mix zone and the second pre-mix zone are configured to receive the first gas and second gas, respectively, via a valve at the end of a gas supply line.

[0479] In certain embodiments, at least one of the first pre-mix zone and the second pre-mix zone are configured to receive the first gas and second gas, respectively, via a sparger.

[0480] In certain embodiments, at least one of the first pre-mix zone and the second pre-mix zone are configured to receive the first gas and second gas, respectively, through a porous sintered metal

[0481] In certain embodiments, non-flammable liquid is introduced into at least one of the first pre-snix zone, the second pre-snix zone, and the common mixing zone via a nozzle comprising a cylindrical or round spout at the end of a non-flammable liquid supply tube.

[0482] In certain embodiments, non-flammable liquid is introduced into at least one of the first pre-mix zone, the second pre-mix zone, and the common mixing zone via a valve at the end of a non-flammable liquid supply tube.

[0483] Flooding of each of the first pre-mix zone and the second pre-mix zone, and the partial Hooding of the common mixing zone can be achieved by direct introduction of non-flammable liquid into the zones via supply nozzles present within the respective zones. It is also contemplated that a single non-flammable liquid supply nozzle, present in one of the first pre-mix zone, the second pre-snix zone, and the common mixing zone may be used to flood the first pre-snix zone and the second pre-mix zone and partially flood the common mixing zone. For example, a single non-flammable liquid supply nozzle present in the first pre-mix zone may supply non-flammable liquid into the first pre-snix zone to the point that even after complete flooding of the first pre mix zone supply of non-flammable liquid is continued until non-flammable liquid from the first pre-mix zone overflows into the common mixing zone and the second pre-mix zone. A person skilled in the art would understand that non-flammable liquid supplied to one of the first pre-mix zone, the second pre-mix zone, and the common mixing zone will spread out to the other zones if the supply is continued A person skilled in the art would appreciate that the introduction of non-flammable liquid would need to be controlled so that the common mixing zone is only partially flooded, allowing for the creation of a headspace where the mixed gas may be removed.

[0484] Introduction of the first gas and the second gas into the first pre-mix zone and the second pre-mix zone, respectively, may comprise a single entry' point or multiple entry points.

[0485] Back -flow of non-flammable liquid from any of the first pre-mix zone, the second pre-mix zone, and the common mixing zone and into entry points for the first gas, the second gas, and the non-flammable liquid can be prevented using any means known in the art. For example, gas supply nozzles for introducing the gas into the pre-mix zones may comprise a check valve that prevents the non-flammable liquid from entering the gas supply line through the nozzle. In addition, the pressure difference between the line supplying the first gas or second gas is preferably higher than the pressure within the non-flammable liquid. This arrangement would also limit the flow of the non-flammable liquid out of the first pre-mix zone and the second pre-mix zone and into a gas supply line. [0486] The continual use of the gas mixer gives rise to the potential for buildup of contaminants and fouling within the gas mixer. For this reason, the non-flammable liquid should ideally be periodically replaced, either completely or by continual addition and removal of non-flammable liquid. In certain embodiments, drains within at least one of the first pre-mix zone and second pre-mix are located in close proximity to the front end of the first pre-mix zone and or second pre-mix zone. These drains can be opened to allow drainage of the gas mixer. In certain embodiments, the one or more drains can be controlled so as to limit the rate of drainage, and by coordinating addition of non-flammable liquid with removal a relatively constant level of non-flammable liquid within the gas mixer can be maintained. This configuration allows for continual refreshing of the non flammable liquid.

[0487] In certain embodiments, the gas mixer may be flushed with non-flammable liquid during periods where no gas is introduced into the mixer. Flushing with non-flammable liquid may result in the solubilization and removal of fouling. For example, water-soluble fouling may be removed by extended flushing with water.

In certain embodiments, multiple gas mixers may be employed such that during flusliing of one or more mixers, additional mixers remain online and capable of accepting gases for mixing.

[0488] Construction of the gas mixer can be accomplished with a variety of materials including stainless steel, carbon steel, and any other material chemically compatible with the hydrocarbon to be mixed. Furthermore, the lining of the first pre-mix zone, the second pre-mix zone, and the common mixing zone may be coated with a spark suppressing material such as Teflon, sapphire, or oxide-based ceramic liners or the like. [0489] The temperature, along with the pressure, play a role in determining what fraction of the non flammable liquid may enter the gaseous state, joining the hydrocarbon and oxygen gas present in bubbles that are mixing and moving toward the flooded region of the common mixing zone. The temperature and pressure can be controlled to minimize the carryover of non-flammable liquid into the gas mixture leaving the mixer from the headspace. Temperature control using a heater, either surrounding the gas mixer or within one or more of the first pre-mix zone, the second pre-mix zone and the common mixing zone, is contemplated for use with the provided techniques. Heaters for use in mixing vessels similar to that of those described here are well known. In certain embodiments, the gas mixer is temperature controlled using a heater that is external to the gas mixer. In certain embodiments, the gas mixer is temperature controlled rising a heater that is located within at least one of the first pre-mix zone, the second pre-mix zone, and the common mixing zone.

[0490] The efficiency of mixing of the first saturated gas with the second saturated gas within the flooded region of the common mixing zone is dependent upon, among other things, the residence time, and the frequency of interactions between bubbles of gas. In other words, how often do bubbles collide, break, and reform together, permitting mixing of the gas compositions from each of the bubbles, which combine to fonn the mixed gas. While mixing can occur naturally given sufficient time, it is not likely that the homogeneous gas mixture will comprise a relatively homogeneous mixture of the hydrocarbon and oxygen without internal mixing where collisions between bubbles are promoted. Without interna! mixing, the vessel would need to be of such height as to be not economically feasible. Means for promoting mixing are well known in the art and include use of a static mixers, random packing, structured packing, and impellers

[0491] In certain embodiments, the common mixing zone comprises internal mixing means.

[0492] Static mixers promote mixing by creating a multitude of tortuous pathways that increase the distance that bubbles need to travel to reach the top of the vessel and consequently static mixers act partly by increasing the residence iinre. In addition, the pathways comprise limited space that results in an increased probability that bubbles collide and ultimately mix to combine their gaseous contents.

[0493] Random and structured packing work by providing for increased residence time and probability of interaction between bubbles by creation of a plethora of winding pathways. Random packing involves filling at least a part of the common mixing zone with a packing material that comprises objects 1768 of varying shape and size (Figure 17) that create random pathways for the bubbles to follow as they rise to the top (see dashed arrow' in Figure 17). An example of commonly used random packing is glass beads of varying diameter.

[0494] Structured packing also increases residence time and probability of contact between bubbles but differs from random packing in that the structured packing has an ordered arrangement so that most of the pathways are of a similar shape and size. For example, use of corrugated metal plates 1873 (Figure 18) provides a structured, as opposed to random, array of pathways. In certain embodiments, the internal mixing means comprises a structured bed. Figures 17 and 18 are provided as simplified examples for random and structured packing and should not be seen to limit the scope of this disclosure in any way. In addition, while not shown in the Figures, random and/or structured packing need to be supported wi thin the gas mixer using means known in the art.

[0495] Techniques are provided for the use of power-driven mixers, which can promote interactions by creating flow within the vessel. Impellers include a rotating component 2074 (direction of rotation shown by solid circular arrow), driven by a motor 2075, that may force the non-flammable liquid, and associated bubbles of salurated gas, to the outside wall and away from the center of rotation. Impellers can create axial flow or radial flow' depending upon design, and can be further sub-typed as propellers, paddies, or turbines.

Furthermore, the position of the impeller may be subject to cisange through vertical movement throughout the common mixing zone. Motor driven pumping of an impeller further improves mixing.

[0496] In certain embodiments, the common mixing zone comprises internal mixing means selected from the group comprising a static mixer, a structured bed, random packing, and an impeller.

[0497] The internal mixing means, whether a static mixer, random or structured packing, or an impeller may be comprised of any material that is chemically compatible with the hydrocarbon to be mixed.

[0498] The shape and design of the gas mixer impacts the residence time of the bubbles within each of the first pre-mix zone, the second pre-mix zone, and the common mixing zone. The overall shape of the vessel is not critical, but the distance between where the first gas and the second gas enter the first pre-mix zone and tise second pre-mix zone, respectively , and enter the flooded region, and the distance from tise fluid connection between the common mixing zone and the first and second pre-mix zones and the headspace are. The point of first contact between the gases and the non-flammable liquid in the pre-mix zones should be a distance from common mixing zone that allows for a residence time that permits complete or near complete saturation prior to entering the common mixing zone. Furthermore, the size of the flooded region must be such that the residence time of the saturated bubbles of first gas and second gas is sufficient to the point where a relatively homogeneous mixture of the hydrocarbon-containing gas and the oxidant-containing gas is formed.

[9499] Another consideration for the optimum mixing of the bubbles of gas is the surface area over which the bubbles are dispersed within the common mixing zone A larger surface area of dispersion promotes better mixing. More thorough mixing occurs when a larger number of smaller bubbles are dispersed over a larger surface area. [050(1] The final consideration is the removal of the homogeneous gas mixture from headspace of the gas mixer, which can be accomplished with any variety of means for removal well known in the art.

[0501] Figure 17 shows a schematic representation of the gas mixer 1700, according to an embodiment. Reference is now' made to Figure 17, which show s an embodiment of the gas mixer 1700. The gas mixer 1700 comprises, as shown inFigurel7, first pre-mix zone 1702, the second pre-mix zone 1705, and the common mixing zone 1708, contained within a closed vessel 1720 composing atop end 1721 and a bottom end 1722.

The first pre-mix zone 1702 is flooded with a non-flammable liquid (shaded) and comprises a first front end 1703 and a firs! back end 1704 and the second pre-mix zone 1705 also flooded with the non-flammable liquid comprises second front end 1706 and a second back end 1707. The first pre-mix zone 1702 and the second pre-mix zone 1705 are physically separated by partition 1711 extending from the bottom end 1722 of the closed vessel 1720 to the common mixing zone 1708, the common mixing zone 1708 partially flooded with the non flammable liquid creating a flooded region 1709 and a headspace 1710. In the embodiment shown in Figure 17, common mixing zone 1708 also comprises random packing 1768, shown as a collection of solid gray circles. Bubbles of gas are shown as open circles. The size of the bubbles is for illustrative purposes and is not intended to limit the scope of this disclosure in any way'. The first pre-mix zone 1702 and the second pre-mix zone 1705 are fluidly connected at the first back end 1704 and second back end 1707, respectively, to the flooded region 1709 A first gas supply nozzle 1712, located in close proximity to the first front end 1703, allows introduction of the first gas into pre-mix zone 1702, and a second gas supply nozzle 1713 comprising a sparger, located in close proximity to the second front end 1706, which allows for introduction of the second gas into pre-mix zone 1705. First gas introduced into first pre-mix zone 1702 and second gas introduced into second pre-mix zone 1705 form bubbles within the non-flammable liquid and the interiors of the bubbles may become saturated with gaseous non-flammable liquid as they pass from the first front end 1703 and the second front end 1706 towards the first back end 1704 and second back end 1707, respectively. Preferably, the interior of the bubbles of first gas and second gas are near complete saturation, forming bubbles of a first saturated gas and bubbles of a second saturated gas, which may separately enter the flooded region 1709. Within the flooded region 1709 the bubbles of the first saturated gas arid bubbles of the second saturated gas mix lo form bubbles containing a mixed gas comprising the first gas and the second gas, saturated with the gaseous form of the non-flammable liquid. The bubbles of mixed gas move through the flooded region (see dashed arrow's within the random packing) and pass into the headspace 1710, where they can be removed via mixed gas line 1716.

[0502] In the embodiment shown in Figure 17, it can be seen that the length of first pre-mix zone 1702 and second pre-mix zone 1705 are determined by the length of partition 1711, which ideally extends a distance from the bottom end 1722 that creates pre-mix zones long enough to provide for saturation, or near saturation, of the bubbles w ithin the first pre-mix zone 1702 and second pre-mix zone 1705 with non-flammable liquid prior to leaching the flooded region 1709.

[0503] In the embodiment shown in Figure 17 non-flammable liquid may be added to the common mixing zone 1708, via inlet 1714 located at the top end 1721. While not shown in Figurel7, non-flammable liquid may also be added directly into the first pre-mix zone 1702 or into the second pre-mix zone 1705. The non flammable liquid may be removed through drains (not shown) preferably located in dose proximity to the bottom end L 722 in one or both of the pre-mix zone 1702 and the pre-mix zone 1705. [0504] In certain embodiments, non-flammable liquid is introduced directly into the first pre-mix zone 1702.

[0505] In certain embodiments, non-flammable liquid is introduced directly into the second pre-mix zone 1705.

[0506] Figure 18 shows a schematic representation of the gas mixer 1800, according to a pipe-in-pipe embodiment. Figure 19 shows a cross-section view of the pipe-in-pipe embodiment, taken through the line X-X II! Figure 18.

[0507] Reference is now made to Figures 18 and 1 , which illustrate another embodiment of the gas mixer 1800 In certain embodiments, the gas mixer 1800 encompasses a pipe-in-pipe design comprising an inner pipe 1825 comprising an inner pipe end 1828, an inner pipe inner surface 1829, and an inner pipe outer surface 1830. The inner pipe i 825 is nested inside an outer pipe i 826 comprising an outer pipe end 183 i, an enter pipe inner surface 1832, and an outer pipe outer surface 1833. The tern "nested” is intended to describe the insertion of inner pipe 1825 within the hollow interior of outer pipe 1826 such that inner pipe end 1828 end is positioned within the hollow interior of outer pipe 1826. The diameter of outer pipe 1826 reduces in size to the point w here outer pipe end 1831 contacts inner pipe outer surface 1830 and is sealed along the entire circumference of inner pipe 1825 The reduction in the diameter of outer pipe 1826 may be abrupt, creating an inward inflection or may comprise a rounded end (as shown Figure 18) or a gradual taper. The degree to which the outer diameter of inner pipe 1825 is smaller than the inner diameter of outer pipe 1826 is such that a space is c reated between inner pipe outer surface ! 830 and outer pipe inner surface i 832. The second pre-mix zone ! 805 in this embodiment is formed by the space between inner pipe outer surface i 830 and outer pipe inner surface i 832 and extends from the seal between inner pipe outer surface i 830 and outer pipe end 183 i to inner pipe end 1828. First pre-mix zone 1802 is formed by the hollow' interior of inner pipe 1825 and common mixing zone 1808, comprising the flooded region 1809 and headspace 1810, is formed by the hollow' interior of outer pipe 1826 past inner pipe end 1828. The homogeneous gas mixture may be removed from the headspace 1810 via mixed gas outlet line 1816.

[9598] Non-flammable liquid may be added via an inlet in at least one of the pre-mix zone 1802, the second pre -mix zone 1805, and the common mixing zone 1808.

[9509] In certain embodiments of the gas mixer 1800, the first pre -mix zone i 802 may be flooded with the non-flammable liquid by direct addition of non-flammable liquid via an inlet at any point along the inner pipe inner surface 1829. In certain embodiments of the pipe-in-pipe design, the first pre-mix zone 1802 is flooded with the non-flammable liquid through a valve 1840 at the front end 1841 of inner pipe 1825. Back-flow of non-flammable liquid out of the hollow interior and into the non-flammable liquid supply line 1842 may he prevented by the valve 1840, by higher pressure in the non-flammable liquid supply line 1842, or both.

[0519] In certain embodiments of the gas mixer 1800, the second pre-mix zone 1805 may be flooded with the non-flammable liquid by direct addition of non-flammable liquid via one or more inlets 1843 present at one or more locations along the outer pipe inner surface 1832.

[9511] In certain embodiments of the gas mixer 1800, the common mixing zone 1808 may be partially flooded with non-flammable liquid by direct addition of non-flammable liquid via one or more inlets 1844 at one or more locations along the outer pipe inner surface i 832 past inner pipe end 1828 [0512] In certain embodiments of the gas mixer i 800, the first pre-mix zone i 802 may be flooded with the non-flammable liquid by overflooding the second pre-mix zone 1805 by direct addition of non-flammable liquid via an inlet at any point along the outer pipe inner surface 1832 such that the non-flammable liquid spills over inner pipe end 1828 and into first pre-mix zone 1802 to the point where first pre-mix zone 1802 is entirely flooded. In such embodiments, common mixing zone 1808 may also be partially flooded by overflooding second pre-mix zone 1805 to the point where the level of non-flammable liquid rises past inner pipe end 1828, thereby creating the flooded region 1809 and leaving the head space 1810.

[0513] In certain embodiments of the gas mixer 1800, the second pre-mix zone 1805 may be flooded with the non-flammable liquid by overflooding the first pre-mix zone ! 802 by direct addition of non-flammable liquid via an inlet at any point along the inner pipe inner surface 1829 such that the non-flammable liquid spills over inner pipe end ! 828 atsd into second pre-mix zone i 805 to the point where second pre-mix zone ! 805 is entirely flooded. In such embodiments, common mixing zone 1808 may also be partially flooded by overflooding first pre-mix zone 1802 and overflooding second pre-mix zone 1805 to the point where the level of non-flammable liquid rises past inner pipe end 1828, thereby creating the flooded region 1809 and leaving the head space 1810.

[0514] In certain embodiments of the gas mixer 1800, the first gas is introduced into the first pre -mix zone 1802 via one or more nozzles located on inner pipe inner surface 1826

[0515] In certain embodiments of the gas mixer 1800, inner pipe 1825 comprises a section of porous material 1850, shown as gaps in the wall of inner pipe 1825, and the first gas is introduced into the first pre-mix zone 1802 via passage through the porous material 1850 from a first gas supply line 185 i The porous material may extend all the way along the circumference of inner pipe 1825. Gas supply line 1851 can lead to a manifold 1852 that encompasses the porous material 1850. The pressure of the first gas within the gas supply line 1851 and within the manifold 1852 exceeds the pressure within the hollow interior of inner pipe 1825, promoting passage of the first gas across the porous material 1850. Backflow of non-flammable liquid into the gas supply line 1851 can be prevented by inclusion of a check valve upstream of manifold 1852.

[0516] In certain embodiments of Hie gas mixer 1800, the second gas is introduced into the second pre- mix zone 1805 via one or more nozzles located in close proximity to the front end of second pre-mix zone 1805 on outer pipe inner surface 1832.

[0517] In certain embodiments of the gas mixer 1800, the second gas is introduced into the second pre mix zone 1805 via one or more spargers located in close proximity to the front end of second pre-mix zone 1805.

[0518] In certain embodiments of the gas mixer 1800, as shown in Figure 18 and in cross-section of Figure 19, the second gas is introduced into the second pre-mix zone 1805 via a ring sparger 1849 located in close proximity to the front end of second pre-mix zone 1805. Ring sparger 1849 encircles inner pipe 1825 and is connected to a second gas supply line 1860 comprising a check valve 1861 for preventing backflow of non flammable liquid into gas supply line 1860 past check valve 1861. The second gas enters second pre-mix zone 1805 through sparger holes 1854 arranged throughout the upper surface of ring sparger 1849 (see Figure 19). [0519] Figure 20 shows a schematic representation of the gas mixer 2000, according to a standing pipe embodiment. Reference is now made to Figure 20 illustrating another embodiment of the gas mixer 2000 The common mixi ng zone 2008 of gas mixer 2000 is contained within a standalone vessel 2035 having a bottom end 2036 and a top end 2037, and first pre-mix zone 2002 and second pre-mix zone 2005 comprise the hollow interior of a first pipe 2038 and a second pipe 2039, respectively, that extend from the standalone vessel 2035 in close proximity to the bottom end 2036 of the standalone vessel 2035. The fluid connection between the flooded region 2009 and the first pre-mix zone 2002 and the second pre-mix zone 2005 is created by the point of contact between the open end of first pipe 2038 and the open end of the second pipe 2039, respectively, and the bottom end 2036 of standalone vessel 2035 The homogeneous gas mixture may be removed from the headspace 2010 via mixed gas line 2016.

[0520] In certain embodiments of the gas mixer 2000, first pre-mix zone 2002 is flooded with non flammable liquid by direct introduction of non-flammable liquid via one or more inlets at one or snore locations within the hollow interior of first pipe 2038. In Figure 20, the front end of first pipe 2038 comprises a valve 2045 from which non-flammable liquid cart be introduced from non-flammable liquid supply line 2046.

[0521] In certain embodiments of the gas mixer 2000, second pre-mix zone 2005 is flooded with non flammable liquid by direct introduction of non-flammable liquid via an inlet at any point within the hollow' interior of second pipe 2039. In Figure 20, the end of second pipe 2038 comprises a valve 2047 from which non flammable liquid can be introduced from non-flammable liquid supply line 2048.

[0522] In certain embodiments of the gas mixer 2000, the common mixing zone 2008 is partially flooded with non-flammable liquid by direct introduction of non-flammable liquid via one or more inlets at one or more locations within the hollow interior of standalone vessel 2035 (not shown).

[0523] In certain embodiments of the gas mixer 2000, the common mixing zone 2008 is partially flooded by direct introduction of non-flammable liquid into the hollow' interior of standalone vessel 2035, and first pre mix zone 2002 and second pre-mix zone 2005 are flooded with non-flammable liquid by direct introduction of non-flammable liquid from the common mixing zone 2008 via the fluid connection between the flooded region 2009 and the hollow interior of first pipe 2038 and the hollow' interior of second pipe 2039, respectively.

[0S24] In certain embodiments of the gas mixer 2000, one or both of first pipe 2038 and second pipe 2039 extend vertically downward from the botom end 2036 of standalone vessel 2035. For example, in Figure 20, the first pipe 2038 extends vertically downward from standalone vessel 2035.

[0525] In certain embodiments of the gas mixer 2000, one or both of first pipe 2038 and second pipe 2039 extend horizontally from the bottom end 2036 of standalone vessel 2035. For example, in Figure 20 the second pipe 2039 extends horizontally from standalone vessel 2035. In addition, second pipe 2039 extends horizontally from standalone vessel 2035 and bends upwards so that a section of the second pre-mix zone 2005 is at a vertical position above the headspace 2010.

[Q526] In certain embodiments of the gas mixer 2000, the first gas is introduced into the first pre-mix zone 2002 via one or more nozzles located on the inner surface of first pipe 2038.

[0527] In certain embodiments of the gas mixer 2000, as shown in Figure 20, the first pipe 2038 comprises a section of porous material 2062 and the first gas is introduced into the first pre-mix zone 2002 via passage through the porous material 2062 from a first gas supply line 2063. The porous material may extend all the way along the circumference of first pipe 2038 similar to the configuration shown in Figure 18 in relation to the pipe-in-pipe design. First gas supply line 2063 can lead to a manifold 2064 that encompasses the porous material 2062. The pressure of the first gas within the gas supply line 2063 and within the manifold 2064 exceeds the pressure within the hollow interior of first pipe 2038, promoting passage of the first gas across the porous material 2062.

[0528] In certain embodiments of the gas mixer 2000, the second gas is introduced into the second pre mix zone 2005 via one or more nozzles located on the inner surface of second pipe 2039.

[0529] In certain embodiments of the gas mixer 2000, as shown in Figure 20, the second pipe 2039 comprises a section of porous material 2065 and the second gas is introduced into the second pre-mix zone 2005 via passage through the section of porous material of second pipe 20.39 and from a second gas supply line 2066. The porous material may extend all the wav along the circumference of second pipe 2039. The second gas supply line can lead to a second gas manifold 2067 that encompasses the section of porous material of second pipe 2039. The pressure of the second gas within the second gas supply line 2066 and within the second gas manifold may exceed the pressure within the hollow interior of the second pipe 2039, promoting passage of the second gas across the section of porous material of second pipe 2039.

[0539] The means for introducing first gas into the first gas pre-mix zone and the second gas into the second gas pre-mix zone need not be the same.

[0531] The embodiments of the gas mixer 2000 with respect to orientation of the pipes, the introduction of non-flammable liquid into one or more of the zones, and the introduction of the first and second gas are meant as examples only, and not intended lo limit the scope of this disclosure in any way.

[0532] Techniques are provided for emergency shutdown procedures common to oxidative reaction processes. It is well known that when undesirable conditions occur in an oxidative reaction process an emergency shutdown procedure can he initiated to limit damage to equipment, reduce likelihood of personal injury, and prevent or minimize release of chemicals into the surrounding environment. Known emergency shutdown procedures include the cessation of adding reactants while at the same time providing a flow of an inert material, such as nitrogen, to the reaction zone to displace the reactants from the reactor.

[0533] It is contemplated that for an additional safety component an inert material inlet may be included for the introduction of a flow of an inert material into the headspace. In addition, a suppression outlet leading from the headspace to any known explosion suppression system may be included in the gas mixer. When an unsafe operating condition is detected at any point in the oxidative process, flow of an inert material through the inert material inlet can be initiated while the suppression outlet can be opened. These events can be coordinated with a reduction or termination of the flow of the first and second gas into tise first and second pre-mix zones, respectively Tise result is that any mixed gases within the gas mixer are displaced to the explosion suppression system or to downstream components of the oxidative process. The flow of inert material acts as diluent and promotes movement in a single direction so that backflow of materials from the oxidation reactor into the gas mixer are prevented.

[0534] In certain embodiments, the gas mixer (for example, the gas mixer 1700, 1800, or 2000) further comprises an inert material inlet for introducing an inert material into the head space of the gas mixer and a suppression outlet for removing gaseous mixtures from the head space arid directing removed gaseous mixture to an explosion suppression system.

[0535] Oxidative dehydrogenation of paraffins to olefins is an alternative to the costly, energy intensive and environmentally unfriendly thermal cracking method currently used. In ODH, a stream of one or more alkanes are passed over a catalyst in the presence of oxygen, to produce corresponding olefins and a variety of byproducts that cats be removed in downstream processing steps. Since i ts ODH the conversion of paraffins to olefins is assisted by a catalyst the required operating temperatures are significantly lower than the temperature required for thermal cracking. In addition, for conversion of ethane to ethylene, ODH provides for higher conversion and selectivity rates. Despite these advantages, ODH is not employed commercially due to the risk of thermal runaway of the reaction and consequential loss of containment. Tins risk is due to the requirement for mixing a hydrocarbon-containing gas with oxygen or an oxidant-containing gas.

[0536] Provided here is a process for the oxidative dehydrogenation of a paraffin to a correspondmg olefin More specifically, provided here is a process for oxidative dehydrogenation of ethane into ethylene comprising mixing of ethane and oxy gen in a ratio that falls outside of the flammability envelope in a gas mixer having a first pre-mix zone, a second pre-mix zone, and a common mixing zone, to form a gas mixture of tise etisane and oxygen, and optionally passi ng the mixture of ethane and oxy ge through a heat excisanger to raise the temperature to at least 250 °C, introducing the gas mixture into an ODH reactor containing an ODH catalyst to produce ethylene, carbon monoxide, carbon dioxide, water, acetic acid, and possibly 02, in addition to other hydrocarbons, such as methane and unconverted ethane, and directing the reactor effluent through an oxygen removal unit to remove unconverted oxygen, and a quench tower to remove water and acetic acid. The residual products are then directed through an amine wash to remove carbon dioxide, optionally followed by a caustic wash and dtying, and finally through a demethanizer to remove methane, and if present, carbon monoxide and other compounds which have a boiling point lower than hydrocarbons The products may then be fed through a C spliter to separate ethylene and unconverted ethane.

[0537] The use of the term ODH reactor includes use of more than otse reactor and may include one or more of a fixed bed, fluidized bed, or microchannel reactor. Fixed bed reactors commonly described for ODH include shell-and-tube ty pe reactors, including use of multiple zones separated by partitions, each zone comprising the same or different ODH catalysts. More than one ODH catalyst may be used, either within a single reactor or with different catalysts in separate reactors. Finally, the terms quench tower, amine wash, caustic wash, and dryer are meant to include use of more than one of each, in series. By in series it is meant that use of more than one of each means additional quench towers follow a quench tower, and additional caustic washes follow a caustic wash, and additional divers follow a dtyer

[0538] By using the gas mixer and method of mixing a hydrocarbon -containing gas and an oxidant- containing gas discussed above the iniserent risks of cataly tic ODH are minimized. The mix of ethane and oxygen entering the reactor is outside the flammability envelope so that thermal runaway and subsequent explosion is not likely. Furthermore, by pie-mixing the gases a user can ensure consistent conversion due to the homogeneous nature of the ethane and oxygen mix, wherein the hydrocarbon-containing gas and tire oxidant- containing gas pass through first and second pre-mix zones, respectively, and become saturated with nonflammable liquid before passing into the mixing zone where bubbles of hydrocarbon-containing gas and bubbles of oxidant-containing gas mix within the flooded region and combine to form bubbles comprising the homogeneous gas mixture, and the bubbles containing the gas mixture rise through and eventually exil the flooded region, release the gas mixture into the head space where it can be removed for use in the ODH process [0539] This section relates generally to oxidative dehydrogenation (ODH) of lower alkanes into corresponding alkenes. More specifically, the section relates to a chemical complex for ODH that includes two upstream gas mixer units and a method for cleaning sulfur-containing deposits irons the gas mixers and feed lines to the ODH reactor.

[0540] Provided here is a chemical complex for oxidative dehydrogenation of lower alkanes, the chemical complex including in cooperative arrangement i) at least two mixers for premixing an oxygen containing gas and a lower alkane containing gas to produce a mixed feedstock stream and additionally including a cleaning loop; ii) at least one oxidative dehydrogenation reactor; wherein the at least two mixers are connected in parallel to the at least one oxidative dehydrogenation reactor so that either a first gas mixing unit or a second gas mixing unit is connected to the at least one oxidative dehydrogenation reactor during normal operations; and where an oxidative dehydrogenation catalyst contained within the at least one oxidative dehydrogenation reactor reacts with the mixed feed stock stream to produce a product stream including the corresponding alkene.

[0541 ] Also provided here is a process for removing sulfur-containing deposits during the operation of an oxidative dehydrogenation reactor complex. The process includes operating a chemical complex including in cooperative arrangement: a) at least two mixers for premixing an oxygen containing gas and a lower alkane containing gas to produce a mixed feedstock stream, and b) at least one oxidative dehydrogenation reactor. The at least two mixers are connected in parallel to the at least one oxidative deh drogenation reactor so that either a first gas mixing unit or a second gas mixing un t is connected to the at least one oxidative dehydrogenation reactor during normal operations. An oxidative dehydrogenation catalyst contained within the at least one oxidative dehydrogenation reactor reacts with the mixed feed stock stream to produce a product stream including the corresponding alkene. The process includes monitoring the pressure within the chemical complex during normal operation. The process includes switching from a first mixer for premixing the oxygen containing gas and the lower alkane containing gas to a second mixer when a pressure drop is observed. The process includes purging the first mixer of the flammable hydrocarbons and oxygen by the means of gaseous of liquid purge. The process includes introducing a cleaning solvent into the first mixer and cycling the cleaning solvent through a cleaning loop until the sulfur-containing deposits are removed. The process includes continuing to monitor the pressure within the complex during normal operation. The process includes switching back to the first mixer when a pressure drop is observed. The process includes introducing the cleaning solvent into the second mixer and c cling the cleaning solvent through a cleaning loop until the sulfur-containing deposits are removed. The process includes repeating the aforementioned steps during continued operation of the chemical complex.

[0542] Also provided here is a process for removing sulfur-containing deposits during the operation of an oxidative dehydrogenation reactor complex. The process includes operating a chemical complex including in cooperative arrangement: a) at least two mixers for premixing an oxygen containing gas and a lower alkane containing gas to produce a mixed feedstock stream, b) at least one oxidative dehydrogenation reactor, and c) a feedline connecting each of the at least two mixers to the at least one oxidative dehydrogenation reactor, where the feedlines are fitted with sprayers to introduce a cleaning solvent to internal walls of the feedline. The at least two mixers are connected by the feedline in parallel to the at least one oxidative dehydrogenation reactor so that either a first gas mixing unit or a second gas mixing unit is connected to the at least one oxidative dehydrogenation reactor during normal operations. An oxidative dehydrogenation catal st contained within the at least one oxidative dehydrogenation reactor reacts with the mixed feed stock stream to produce a product stream including the corresponding alkene. The process includes monitoring the pressure within the chemical complex during normal operation. The process includes introducing the cleaning solvent into the feedline through the sprayer to remove sulfur-containing deposits when a pressure drop is observed in the chemical complex. The process includes continuing to monitor the pressure within the chemical complex during operations and while the cleaning solvent is being introduced. The process includes stop the cleaning solvent flow once the pressure in the chemical complex returns to normal operating levels.

[0543] Techniques are provided for oxidative dehydrogenation (ODH) of lower alkanes into corresponding alkenes. In some embodiments, there is a chemical co plex useful for ODH and in another aspect there is described a process for ODH that may be performed in the chemical complex outlined in the first aspect. Lower alkanes are intended to inclnde saturated hydrocarbons with from 2 to 6 carbons, and the corresponding a!kene includes hydrocarbons with the same number of carbons, but with a single double carbon to carbon bond For ethane, ethylene is its corresponding alkene.

[0544] Figure 21 is a schematic representation of an embodiment of the gas mixer 2100. The gas mixer 2100 includes a closed mixing vessel 2110 having a top end 2109 and a bottom end 2107. The closed mixing vessel 2110 is flooded with a non-flammable liquid, the choice of which depends on the application for which the mixed gas is to be used. Non-flammable liquid may be added to the dosed mixing vessel 2110 via a nozzle or inlet 2102 located at the top end 2109, while non-flammable liquid may be removed from the outlet 2103 located at the bottom end 2107.

[05451 Construction of the mixing vessel 2110 can be accomplished with a variety of materials including stainless steel, carbon steel, and any other material chemically compatible with the hydrocarbon to be mixed. Furthermore, the lining of mixing vessel 2 i 10 may be coated with a spark suppressing material such as Teflon, sapphire, or oxide-based ceramic liners or the like.

[0546] Lower alkane containing gas may be introduced into the closed mixing vessel 2110 through the lower alkane containing gas supply nozzle 2104, while the 0x5' gen containing gas may be introduced via oxygen containing gas supply nozzle 2105. The lower alkane containing gas supply nozzle 2104 and the oxygen containing gas supply nozzle 2105 cooperate with the closed mixing vessel 2110 in a w ay so that introduction of the gases directly into the non-flammable liquid occurs at or near the bottom end 2107 of the closed mixing vessel 2110. For the purposes of this disclosure, the term “nozzle" refers simply to tire point where contact between the gases and the non-flammable liquid within the closed mixing vessel 2110 first occurs, and can include any means know'» within the art;. W hile not essential, the lower alkane containing gas supply nozzle 2104 and the oxygen containing gas supply nozzle 2105 may be orientated such that streams of the lower alkane containing gas and the oxy gen containing gas impinge upon one another immediately upon entering the mixer. The introduced gases rise and are mixed through mixing zone 2108 and are available for removal after exiting the non-flammable liquid at the top of the closed mixing vessel 2110 through the mixed gas removal line 2106. The mixed gas is optionally passed through a heat exchanger 2111 and the optionally heated mixture then passes into a reactor, for example an ODH reactor.

[0547] As the term suggests, the non-flammable liquid used to flood the closed mixing vessel 2110 is not flammable Thai is, the non-flammable liquid is not capable of igniting or burning, for example, under conditions experienced within the reactor. Examples of suitable non-flammable liquids include water, ethylene glycol, silicon oils, and casbon tetrachloride. In some embodiments, water is used as the non-flammable liquid. While any non-flammable liquid may be used with the various embodiments disclosed herein, it is important to consider that mixed gas removed from the gas mixer 2100 will include the lower alkane containing gas, oxygen containing gas, and in some instances carryover of non-flammable liquid. For tins reason, selection of a non flammable liquid also considers any potential effects the cany over may have on downstream applications. Catal sts used for oxidative reactions may be sensitive to catalytic poisoning by specific non-flammable liquids that tire carried over in a gaseous state.

[0548] The temperature, along with the pressure, play a role in determining what fraction of the non flammable liquid may enter the gaseous state, joining the hydrocarbon and oxygen gas present in bubbles that are mixing and rising to the top end of the closed mixing vessel 2110. The temperature and pressure can be controlled to minimize the carryover of non-flammable liquid into the gas mixture leaving through mixed gas removal line 2106 Temperature control using a heater, within or external the closed mixing vessel 2110 is contemplated for use with the provided techniques. In some embodiments, tise closed mixing vessel 2100 is temperature controlled using a heater that is external to the closed mixing vessel 2110. In another embodiment, the closed mixing vessel 2110 is temperature controlled using a heater that is located within the closed mixing vessel 2110.

[0549] In some instances, it may be desirable for recycling purposes, to include a secondary lower alkane containing gas supply nozzle or product supply nozzle 2115. For example, some oxidative reactions are not as efficient as others are and may include conversion rales below an acceptable level. In those cases, it may be desirable to send a product line containing product and unreacted hydrocarbon back to start the oxidative reaction process again, with the intent of maximizing conversion of the starting hydrocarbon — the hydrocarbon originally mixed in the gas mixer before passage through an oxidative process. The product stream, similar to and containing unreacted starting hydrocarbon would need to be mixed with oxidant before entering the reactor. If the product contained in the product stream is more reactive to oxygen than the starting hydrocarbon, it would be safer to introduce the product stream into the reactor at a point where the oxygen is already partially mixed and diluted. To this end, in some embodiments, the secondary lower alkane containing gas supply nozzle 2115 is at a position distant from the 0x5' gen containing gas supply nozzle 2105. The position of the secondary lower alkane containing gas supply nozzle 2115 is no! critical, provided it is in a position where the oxygen present in the closed mixing vessel 2110 has begun mixing with the lower alkane containing gas, and there is sufficient residence time for the product gas to mix thoroughly with the added oxygen and lower alkane containing gases. In some embodiments, the position of the secondary' lower alkane containing gas supply nozzle is near a point equidistant from the oxygen containing gas supply nozzle 2105 and the point where mixed gas removal line 2106 leaves the top end 2109 of the closed mixing vessel 2110. The secondary' lower alkane containing gas supply nozzle 2115 may also be used as an additional input location for the introduction of the lower alkane containing gas. In some embodiments, there is a secondary low er alkane containing gas supply nozzle 2115 for introducing a product stream from an oxidative process or additional lower alkane containing gas into the closed mixing vessel 2110 at a point distant from oxygen containing gas supply nozzle 2105.

[0550] In embodiments where there is recycling of an oxidative process such that a product line is fed back to the gas mixer 2100 tor introduction into the closed mixing vessel 2110 via the secondary lower alkane containing gas supply nozzle 2115, it is contemplated that heat fro the product line may be used in temperature control of the closed mixing vessel 2110. The heat provided from an oxidative process, for example ODH, may be used its this fashion and would therefore assist in reducing the cost associated with providing heat through an internal or external heater. In another embodiment, the closed mixing vessel 2 HO is temperature controlled using heat from a product line leaving an exothermic oxidation process.

[0551] Internal mixing

[0552] The efficiency of mixing of the gases within 2108 is dependent upon, among other things, the residence time, and the frequency of interactions between bubbles of gas. In other words, how often do bubbles collide, break, and reform together, permitting mixing of the gas compositions front each of the bubbles, which combine to form a homogeneous mixture. Means for promoting mixing are well known in the art and include use of a static mixers, random packing, structured packing, and impellers.

[0553] Static mixers promote mixing by creating a multitude of tortuous pathways that increase the distance that bubbles need to travel to reach the top of the vessel and consequently static mixers act partly by increasing the residence time in addition, the pathways include limited space that results in an increased probability that bubbles collide and ultimately mix to combine their gaseous contents. In some embodiments, the internal mixing means includes a static mixer.

[0554] Random and structured packing act similar to static mixers in that they provide for increased residence time and probability of interaction between bubbles by creation of a plethora of winding pathways. Random packing involves filling at least a part of the closed mixing vessel with a packing material that includes objects of varying shape and size that create random pathways for the bubbles to follow' as they rise to the top. An example of commonly used random packing is glass beads of varying diameter. In some embodiments, the internal mixing means includes a packed bed.

[0555] Structured packing also increases residence time and probability of contact between bubbles, but differs from random packing in that the structured packing has an ordered arrangement so that most of the pathways are of a similar shape and size. Random and structured packing are supported within the gas mixer using means know n in the art. In some embodiments, the internal mixing means includes structured packing. |0556] Techniques are provided for the use of pow¾r driven mixers, which can promote interactions by creating flow within the vessel. Impellers include a rotating component, driven by a motor that may force the non-flammable liquid, and associated bubbles of gas, to the outside wall and away from the center of rotation. Impellers can create axial flow or radial flow depending upon design, and can be further sub-typed as propellers, paddles, or turbines. Furthermore, the position of the impeller may be subject to change through vertical movement throughout the mixing zone. Motor driven pumping of an impeller further improves mixing. In some embodiments, the closed mixing vessel includes an impeller

[0557] Similar to the closed mixing vessel, the internal mixing means, -whether a static mixer, random or structured packing, or an impeller may be comprised of any material that is chemically compatible with the hydrocarbon to be mixed.

[0558] The shape and design of the closed mixing vessel impacts the residence time. The overall shape of the vessel is not critical, but the distance betw een where the gas enters and exits the mixing zone should be considered when designing the unit. The point of first contact between the gases and the water in the closed mixing vessel should be a distance from the top that allow s for a residence time that permits complete mixing before removal. In some embodiments, the entry point is near the bottom of the vessel. Where the lines containing the gas enter the vessel is not important, provided the nozzle, e.g,, the point where the gas contacts the water in the vessel, is in the position where residence time is sufficient. [0559] Another consideration for the optimum mixing of the gases is the surface area over which the gases are dispersed. A larger surface area of dispersion promotes better mixing. While injection through a single inlet is feasible, provided sufficient residence time, more thorough mixing occurs when a larger number of smaller bubbles are dispersed over a larger surface area. Having multiple lower alkane containing gas supply nozzles and multiple oxygen containing gas supply nozzles allows each of the gases to be introduced in multiple locations. Conversely, a single nozzle may include multiple exit points where gas can enter the vessel, effectively dispersing the gas over a greater surface area compared to dispersion front a nozzle with a single exit point. In some embodiments, at least one of the lower alkane containing gas supply nozzle 2104 and the oxygen containing gas supply nozzle 2105 itscludes a sparger

[0560] In some embodiments, the lower alkane containing gas supply nozzle 2104 atsd the oxygen containing gas supply nozzle 2105 are arranged as spargers in the form of concentric rings. Furthermore, the exit points for the lower alkane containing gas and the oxygen containing gas from their respective nozzles me arranged such that the streams of gas impinge on one another, initiating mixing as early as possible after introduction into the mixer. The arrangement of the gas supply nozzles is not limited to examples provided here. As another example, a series of concentric rings, with alternating oxygen and low er alkane containing gas supply nozzles, is also contemplated

[0561] Emergency Shutdown

[0562] Another embodiment relates to emergency shutdown procedures common to oxidative reaction processes. It is well known that whets undesirable conditions occur in an oxidative reaction process an emergency shutdown procedure can be initiated to limit damage to equipment, reduce likelihood of persotsal injury, and prevent or minimize release of chemicals into the surrounding environment. Known emergency shutdown procedures include the cessation of adding reactants while at the same time providing a flow of an inert material, such as nitrogen, to the reaction zone to displace the reactants from the reactor.

|0563[ In some embodiments, it is contemplated that for an additional safety component an inert material inlet, located near the top end and above the liquid level, may be included for the introduction of a flow of an inert material. In addition, a suppression outlet leading to any known explosion suppression system may be included near the top end of the gas mixer. When an unsafe operating condition is detected at any point in the oxidative process, flow' of an inert material through the inert material inlet can be initiated while the suppression outlet can be opened. These events can be coordinated with a reduction or termination of the hydrocarbon and oxidant reactants. The result is that any mixed gases within the mixer are displaced to the explosion suppression system or to downstream components of the oxidative process. The flow of inert material acts as diluent and promotes movement in a single direction so that backflow of materials from the oxidation reactor into the gas mixer are prevented,

[0564] In some embodiments, the gas mixer further includes an inert material inlet, located near the top end of the gas mixer, for introducing an inert material into the gas mixer above the level of the non-flammable liquid, and a suppression outlet for removing gaseous mixtures, located near the top end of the gas mixer and leading to an explosion suppression system

[0565] Method for mixing a lower alkane containing gas and an oxygen containing gas

[0566] Techniques are provided for applications that itsclude the mixing of a lower alkane containing gas with an oxygen containing gas. It is well known that gaseous compositions containing a hydrocarbon and oxygen in ratios that fall within the flammability envelope are potentially hazardous. An ignition event, such as a spark, can ignite the mixture and potentially lead to an explosion. While applications that require mixing of hydrocarbons and oxygen normally do so with ratios that are safe and not susceptible to ignition there are moments during initial mixing, where heterogeneous pockets of unfavorable iiydragen/oxygen compositions exist and may ignite if a spark occurs.

[0567] Techniques are provided for a method for mixing a lower alkane containing gas with art ox gen containing gas that is simple and safe in that ignition events are unlikely to occur. The method includes introducing, separately and simultaneously, a lower alkane containing gas and an oxygen containing gas directly into a closed mixing vessel having a top end and a bottom end and flooded with a non-flammable liquid, i n close proximity' (e.g , within 15% or 10% of the length of the reactor) to the bottom end, allowing the bubbles of gas to mix while surrounded by the non-flammable liquid, and removing from the top of the vessel, after mixing is complete, a homogeneous mixture of the lower alkane containing gas and the oxygen containing gas in a ratio that is outside of the flammability envelope.

[0568] In some embodiments, the amount of the gases introduced into the botom end of the closed mixing vessel 2110 will result in a final composition that includes a ratio of lower alkane containing gas to oxygen containing gas that is outside of the flammability envelope. The chosen ratio will depend on the nature of the gases and the application for which the mixture will be used. Fo r example, for an ODH application, the ratio of ethane to oxygen chosen will depend on whether under the proposed ODH reaction conditions the ratio is above the higher explosive limit or below' the lower explosive limit. In comparison, the ratio of ethylene to oxygen added to the reactor would be different because ethylene is more reactive than ethane. Further, the flammability' limit of ethylene is higher than that of ethane. The temperature of the ODH process to be employed should also be taken into consideration as higher temperatures correspond to a much smaller window of safe ratios of ethane to oxygen. For example, a molar ratio of about 80:20 ethane to oxygen for catalytic ODH would fall above the upper explosive limit, while a ratio of about 1.5:98.5 ethane to oxygen would fall below the lower explosive limit, with each ratio safe enough in that ignition events would not lead to an explosion or flame propagation under ODH reaction conditions. Ratios falling between that, 50:50, for example, may be potentially flammable/explosive posing a risk of process disruptions.

[0569] The next consideration after determining the desired final ratio of hydrocarbon to oxygen is determining the flow rate at which each gas is added to the bottom of the closed mixing vessel 2110. The flow rate of the gases and the corresponding pressure would need to be higher than the pressure of the non-flammable liquid in the closed mixing vessel 2110. In the absence of a pressure differential, the gases cannot enter the closed mixing vessel 2110 and consequently the mixing zone 2108. Furthermore, if the pressure of the non flammable liquid is higher than the line containing the gas to be introduced there may be, in the absence of a one-way valve, flow back of non-flammable liquid into the gas supply lines. This should be avoided.

[0570] When determining flow rates, the skilled worker correlates the flow rates with the pressure and temperature used within the closed mixing vessel 2110. The conditions within the closed mixing vessel 2110 are chosen to reflect the amount of carryover of non-flammable liquid into the gas mixture removed through mixed gas removal outlet 2106. In some embodiments, flow rates of the incoming gases allow entry' into the non flammable liquid at the predetermined temperature and pressure. [0571] As a further safety precaution, in embodiments where the dilution of the oxygen containing gas with non-flammable liquid prior occurs to entry into the closed mixing vessel 2110. The prior dilution of the oxygen containing gas permits the saturation of incoming oxygen molecules with molecules of the non flammable liquid that discourage ignition events igniting any hydrocarbons that interact with the oxygen during the early stages of mixing. Dilution of the oxygen containing gas with non-flammable liquid can be accomplished by directing a non-flammable liquid line into the oxygen containing gas line poor to the oxygen containing gas nozzle. Non-flammable liquid present within the closed mixing vessel 2110 that is ejected via 2103 may be suitable for this purpose, provided this non-flammable liquid passes through a filter to remove pasticuiate mater prior to introduction into the oxygen containing gas line. In some embodiments, the oxygen containing gas is diluted with non-flammable liquid prior to introduction into the closed mixing vessel 2110. [0572] The choice of gas mixer and associated design of the closed mixing vessel should consider the factors discussed above. In some embodiments, gas mixers allow for a residence time that allows complete, or near complete, mixing to create a homogeneous composition of gas where there are no potentially unsafe pockets of gas with undesirable ratios of hydrocarbon to oxygen.

[0573] The final consideration is the removal of the mixed gas from the top of the closed mixing vessel, which can be accomplished with any variety of means for removal well known in the art.

[0574] Figure 22 is a schematic representation of a twinned gas mixer unit 2200. In some embodiments, at least two gas mixer units 2100 are associated with and integrated into the complex including the ODH reactor.

In these embodiments, the mixed gas from either of the mixer units 2100 can be introduced to the ODH reactor after exiting the top of the closed mixing vessel through the respective mixed gas removal line 2106. A valve configuration 2212 allows for either switching between the two gas mixer units or allowing both gas mixer units to feed into tire reactor (after passing tlirough the optional heat exchanger 2211) at the same time.

[0575] Sulfur containing Deposits

[0S76] Techniques are provided for the ability' to remove sulfur and sulfur containing deposits that are created as a result of mixing the feed gases. A very common contaminant in ethane feeds to petrochemical plants is ¾S and in some cases elemental sulfur (refinery' paraffin / olefin sources). It is known that when ¾S is combined with oxygen at low temperatures one result may be formation of deposits including elemental sulfur or solid sulfur-rich compounds. In a reactor environment, this can lead to severe equipment fouling and potential shutdown of the equipment. Considering the tight specifications typically in place for H ₂ S concentration in feed streams, the rate of fouling is usually rather low, yet it is nonetheless very likely to occur and to build up over time. In some instances, pretreatment steps are put in place to remove any ¾S prior to exposure of the feed streams to oxygen. However, even the best of technologies may result in breakthrough of ¾S to downstream equipment. As such, methods to address the removal of those deposits are useful in, for example, ODH reactor complexes.

[0577] There are known methods to remove sulfur based fouling, and/or coke deposits from reactors. Disclosed herein however, are methods for removal of deposits in premixers as well as the feed lines that lead from the mixer unit into a reactor. These methods are not specifically intended to address deposits within a reactor. The methods disclosed use a combination of a twinned 0 _? /HC mixer tower as detailed herein above and shown in Figure 22, or any other kind of mixer unit, wherein the mixerunit includes injection ports prior to the inlei of the reactor to introduce a solvent that would have little or no impact on the ODH reactor section performance but would dissolve and/or remove the fouling deposits.

[0578] A cleaning solvent is any solvent that dissolves or loosens or dislodges or suspends in the cleaning solvent, the sulfur containing deposits and does not affect the operation of the ODH reactor. One such compound, which has been demonstrated to dissolve elemental sulfur as well as sulfur-rich organic fouling compounds, is dimethyl disulfide (DMDS). This solvent also meets the requirement of allowing the ODH catalyst and ODH process to proceed as desired when used to remove deposits while the reactor remains in operation. It is speculated that DMDS is a good material to dissolve solid sulfur fouling, as it does not act as a true solvent, but rather as a reactant. The sulfur-rich organic fouling enters an equilibrium with the DMDS solvent, which allows it to remain in the liquid phase regardless of temperature. It is stated in the literature that DMDS is capable of taking up as much as 600 weight percent (wt. %) of elemental sulfur as poly sulfides at 80 °C. Other potentially useful solvents or reactants for dissolving sulfur-rich organic fouling compounds include carbon disulfide and warm or hot toluene. In some embodiments, the toluene is warmed to temperatures below the boiling point. In some embodiments, the toluene is heated to about 80 °C.

[0579] Determining when the mixer units or feed lines have sulfur-rich organic fouling that requires cleaning is something that is known to a person of ordinary skill in the art and can be done by monitoring the pressure within the complex at various points in the system. When there is a pressure difference at two different measured points, which indicates that fouling has occurred and cleaning may be needed. In other embodiments, the amount of ¾S or total sulfur on the inlet to the mixer, on the outlet of the mixer, and inlet to the reactor, can be monitored. The measured values can be used to indicate the size the fouling in the corresponding section of the equipment

[0580] An example of a chemical complex for oxidative dehydrogenation of lower alkanes includes in cooperative arrangement: i) at least two mixers for premixing an oxy gen containing gas and a lower alkane containing gas to produce a mixed feedstock stream and additionally including a cleaning loop; and is) at least one oxidative dehydrogenation reactor. In some embodiments, the at least two mixers are connected in parallel to the at least one oxidative dehydrogenation reactor so that either a first gas mixing unit or a second gas mixing unit is operating and connected to the at least one oxidative dehydrogenation reactor during normal operations. In some embodiments, the at least two mixers are connected in parallel to the at least one oxidative dehydrogenation reactor so that both the first gas mixing unit and the second gas mixing unit are operating and connected to the at least one oxidative dehydrogenation reactor during normal operations. The oxidative dehydrogenation catalyst contained within the at least one oxidative dehydrogenation reactor reacts with the mixed feed stock stream to produce a product stream including the corresponding a!kene.

[0581 ] While it is most likely that the complex will be operated using a single mixer unit, which is alternated with the other mixing unit once fouling is detected, it is also contemplated that the complex can be operated while both mixer units are online. Both mixers can be shirt down and cleaned, but the presently disclosed apparatus and methods allows for the advantage of cleaning the mixer units while the cotnplex continues to operate. In some embodiments, when fouling is detected a single unit can be isolated and cleaned while the other continues operating.

[0582] The cleaning loop is an arrangement of inlets and outlets on the mixer unit that provide for the i) injection of cleaning solvents into the mixer unit ii) circulation of solvent in the mixer unit, iii) removal of cleatsing solvent from the mixer unit. The cleaning loop inlets may be located at any position in the mixer unit(s) that allows for cleaning. In some embodiments, the cleaning loop inlets may be at or near the lower alkane containing gas supply nozzle or at or near the oxygen containing gas supply nozzle (2214 in Figure 22). In other embodiments, the cleaning loop inlets may be at or near the mixed gas removal line. In some embodiments, the cleaning loop outlets are located at or near the mixed gas removal line (2215 in Figure 22). In other embodiments, the cleatsing loop outlets may be at or near the lower alkane containing gas supply nozzle or at or near the oxygen containing gas supply nozzle

[Of5S3] Its some embodiments, the cleaning loop further includes a pump 2213, and/or a filter, and/or a small heating vessel. Its some embodiments, the solvent is heated to about 60 °C, or for example to about 80 °C degrees during the cleaning process. The iempesatsrre should be kept below the boiling point of the solvent used for cleaning (e.g., the boiling point of DMDS is 110 °C).

[0584] In some embodiments, the complex further includes a knockout vessel, after the mixed feedstock stream outlet mid in close proximity (e.g. within about the length of the dehydrogenation reactor) to the at least one oxidative deh drogenation reactor, wherein the knockout vessel is configured to receive condensed cleaning solvent. The condensed cleaning solvent may also contain the dissolved sulfur fouling material.

[0585] In some embodiments, either in addition to or instead of the cleaning loop, the complex further includes sprayers that are fitted on to the feedlines between the mixer units and the at least one oxidative dehydrogenation reactor, which allows the solvent to be sprayed onto the internal walls of the feedline.

[0586] The feedlines are any of the pipes or feeds between the mixer unit 2 i 10, the optional heat exchange unit 221 i, and the reactor, shown but not numbered in Figures 21 and 22

[0587] The sprayer, also commonly referred to as an atomizer, can take numerous forms depending on the cleaning solvent properties, receiving fluid (e.g., mixed feedstock) properties and flow rates and the local geometry' (e.g., pipe diameter, pipe length, bends or elbow s). The sprayer can be flush to the pipe wall or inserted on a small pipe or lance to position it in an optimal way to maximize coverage of the walls by the spray. The sprayer typically will have the cleaning solvent supplied to it at a pressure significantly higher than the pressure in the feed stock piping. This pressure is used with the geometry' of the sprayer nozzle to atomize the solvent into droplets that will coat the walls of the receiving pipe. The sprayer nozzle may have multiple holes, use swirl or, in some embodiments, use a high -pressure gas to obtain the required solvent droplet size and droplet spray pattern to cover the internal walls of the feedline. SPRAYING SYSTEMS CO ® is a company that sells numerous spray nozzles designs and spray nozzle holders (also referred to as quills, lances, or injectors). [0588] In some embodiments, the process for removing sulfur-containing deposits during the operation of an oxidative dehydrogenation reactor complex includes operating a chemical complex including in cooperative arrangement: at least two mixers for premixing an oxygen containing gas and a lower alkane containing gas to produce a mixed feedstock stream and at least one oxidative dehydrogenation reactor, in which the at least two mixers are connected in parallel to the at least one oxidative dehydrogenation reactor so that either a first gas mixing unit or a second gas mixing unit is connected to the at least one oxidative dehydrogenation reactor during normal operations. An oxidative dehydrogenation catalyst contained within the at least one oxidative dehydrogenation reactor reacts with the mixed feed stock stream to produce a product stream including the corresponding alkene. The process includes monitoring the pressure within the chemical complex during normal operation. The process includes switching from a first mixer for premixing the oxygen containing gas and the lower alkane containing gas to a second mixer when a pressure drop is observed. The process includes purging the first mixer of the flammable hydrocarbons and oxygen by the means of gaseous or liquid purge. The process includes introducing cleaning solvent into the first mixer and cycling cleaning solvent through a cleaning loop until the sulfur-containing deposits are removed. The process includes continuing to monitor the pressure within the complex during normal operation. The process includes switching back to the first mixer when a pressure drop is observed. The process includes introducing cleaning solvent into the second mixer and cycling cleaning solvent through a cleaning loop until the sulfur-containing deposits are removed. The process includes repeating the aforementioned steps during continued ope ration of the chemical complex

[0589] In some embodiments, the process for removing sulfur-containing deposits during the operation of an oxidative de ydrogenation reactor complex includes operating a chemical complex including in cooperative arrangement: at least two mixers for pre mixing an oxygen containing gas and a lower alkane containing gas to produce a mixed feedstock stream and at least one oxidative dehydrogenation reactor, in which the at least two mixers are connected in parallel to the at least one oxidative dehydrogenation reactor and both a first gas mixing unit and a second gas mixing unit are connected to the at least one oxidative dehydrogenation reactor during normal operations. An oxidative dehydrogenation catalyst contained within the at least one oxidative dehydrogenation reactor reacts with the mixed feed stock stream to produce a product stream including the corresponding alkene. The process includes monitoring the pressure within the chemical complex during normal operation. The process includes, when a pressure drop is observed, isolating at least one of the at least two mixers. The process includes purging the isolated mixer of the flammable hydrocarbons and oxygen by the means of gaseous of liquid purge. The process includes introducing cleaning solvent into the isolated mixer from the previous step and cycling cleaning solvent through a cleaning loop until the sulfur-containing deposits are removed. The process includes optionally repeating purging and introducing cleaning solvent for the mixer unit that remained on line. The process includes optionally returning to operation where the at least two mixers me operational.

[0599] In some embodiments, prior to introducing the cleaning solvent the mixer is drained, then flushed and dried with an inert gas. In some embodiments, the mixer is drained, then flushed and dried with an inert gas prior to being brought back online for normal operations.

[0591] In some embodiments, the process for removing sulfur-containing deposits during the operation of an oxidative dehydrogenation reactor complex includes operating a chemical complex including in cooperative arrangement: at least two mixers for piemixing an oxygen containing gas and a lower alkane containing gas to produce a mixed feedstock stream, at least one oxidative dehydrogenation reactor, and a feedline connecting each of the at least two mixers to the at least one oxidative dehydrogenation reactor, in which the feedlines are fitted with sprayers to introduce cleaning solvent to internal walls of the feedline. The at least two mixers are connected by the feedline in parallel to the at least one oxidative dehydrogenation reactor so that either a first gas mixing unit or a second gas mixing unit is connected to the at least one oxidative dehydrogenation reactor during normal operations, or the at least two mixers are connected in parallel to the at least one oxidative dehydrogenation reactor so that both the first gas mixing unit and the second gas mixing unit are operating and connected to the at least one oxidative dehydrogenation reactor during normal operations. Art oxidative dehydrogenation catalyst, for example, as part of a catalyst material, contained within the at least one oxidative dehydrogenation reactor reacts with the mixed feed stock stream to produce a product stream including the corresponding aikene. The process includes monitoring the pressure within the chemical complex during norma! operation. The process includes introducing cleaning solvent into the feedline through the sprayer to remove sulfur-containing deposits when a pressure drop is observed in the chemical complex. The process includes continuing to monitor the pressure within the chemical complex during operations and while cleaning solvent is being introduced. The process includes stop cleaning solvent flow once the pressure in the chemical complex returns to normal operating levels.

[0592] In some embodiments, additional components or additives may be included in the cleaning solvent. For example, in sotne embodiments, sodium bisulfate is introduced with the cleaning solvent. In some embodiments, sodium bisulfate is added to DMDS and used as the cleaning solvent

[0593] The following examples are merely illustrative and are not intended to be limiting. Unless otherwise indicated, all percentages are by weight.

[0594] EXAMPLES

[0595] In an experiment to find an effective solvent that could dissolve sulfur fouling. The following chemicals were tested experimentally: toluene*, methanol, wash oil* (Refinery Heavy Reformate - Aromatic hydrocarbons), heptane, water, EnviroSol (citms-based solvent/degreaser), dimethyl sulfoxide (DMSO), carbon disulfide sCS ). dimethyl disulfide* (DMDS), and tertiary butyl polysulfide (TBPS).

[0596] CS was found to be a good solvent for removing fouling material at room temperature. However, this solvent presents a safety hazard and has very' high toxicity. Additionally, its effect on the ODH process is unknown.

[0597] Heated toluene was also found to be effective and is a safer solvent than CS . As a result, heated toluene has been recommended for use under extreme circumstances. For example, when the feed vaporizers are blown down the material is pushed into the flare header where it could accumulate and block the line. As this also presents some safety hazards and could lead to a full site shutdown, the plant could inject toluene into the flare and heat the piping using external steam hoses. However, using a heated solvent to remove sulfur-based fouling can lead to precipitation down the line once the solvent cools. The entire line should be warmed up to prevent this issue.

[0598] Due to the above challenges, an alternative solution was sought. It was found in literature that heated DMDS may be a good material to dissolve solid sulfur fouling, as it may not act as a true solvent, but rather as a reactant. Without wishing to be bound by theory, it is believed that the sulfur fouling enters an equilibrium with the DMDS solvent, which allows it to remain in the liquid phase regardless of temperature. It is stated in literature that DMDS is capable of taking up as much as 600 wt. % of elemental sulfur as polysulfides at 80 °C.

[0599] Testing was completed in the iaboratoiy, using the fouling collected from the feed vaporizer and heated DMDS. Approximately 1 g of fouling was submerged in 10 g of DMDS. The mixture was heated to 80 °C, after approximately 3Q minutes the solid material was completely dissolved and separated into a dark black liquid phase and a yellow liquid phase. The material was removed from the heat and left overnight at rootn temperature, upon further inspection, it was verified that the material retnained in liquid fonn and no solid fouling was present. Since a decrease in temperature typically favors precipitation, the vial was then cooled to approximately -60 ^C, C using dry ice. At such low temperatures, it was found that the entire mixture would become a gei-iike soiid. Once the vial was removed from the cooling medium and allowed to return to room temperature the material became a liquid once again

[0600] The majority of the solvents listed above were tested at room temperature for extended periods of time (20 hours for solvents 1 - 6 and 1 hour for sulfur-based solvents 7 - 10) additionally, solvents marked with an asterisk (*) w ere also tested at high temperatures (up to 8Q °C).

[0601] Overall, the majority of the solvents were not effective in dissolving the sulfur-based fouling. The noted exceptions were carbon disulfide at room temperature, toluene at high tempe ratures, and DMDS at high temperatures.

[0602] Effect of DMDS on catalyst performance

[0603] Using a Micro Reactor set up Catalyst long term activity testing using fixed bed reactor platform on Micro Reactor Unit 1 was conducted to test the robustness of the ODH catalysis continuously for 10 days for DMDS effect study. This test was carried out with consistent run condition with temperature of 369 °C and 3000h * gas hourly space velocity (140 seem of 18% (¾/82% ethane) to aim for 25% conversion. In addition, the regeneration condition was 380C with airflow' of 250 seem for 3 hours each time the regeneration is shown on the Figure 23.

[0604] Figure 23 is a long-term microreactor unit (MRU) run with dimethyl disulfide DMDS injections. [0605] Results are shown in Figure 23. The ODH catalyst activity and selectivity dropped only with very high dosage (0.22ml of DMDS injection in lOmms, corresponding to 19.38 wt.-% of DMDS in the feed), which was simulating the conditions of extreme carryover of DMDS to reactor, which would be possible in case of DMDS injection process upset. For normal injection rate (0.21 ml/ 16 hours, corresponding to 0 2249 wt. % of DMDS in the feed), the impact on activity and selectivity was not noticeable.

[0606] Table Cl. Short term MRU run at 365 °C with weight hourly space velocity (WHSV) of 2.90 lr ^! , with the WHS V based on the active phase, and a gas hourly space velocity (GHSV) of 3000 weight hourly space velocity per hour (If ¹ )

[0607] Another short-term run test was conducted with a different batch of ODH catalyst and as shown in Table Cl, the activity and selectivity was not changed due to (DMDS) injection (0.2ml/5hr).

Use of ODH reaction to convert ethane to ethy lene in an ethane / ethylene mixture

[0608] Techniques are provided for an improvement to chemical complexes having an ethane cracker and a C: splitter. There is a marked increase in the availability of ethane and natural gas liquids particularly in North America such as unconventional shale gas. There are a number of proposals to build chemical complexes to crack ethane and produce petrochemical products. In cracking ethane, there are many high-energy steps. Cracking furnaces are energy intensive, as is the downstream separation train as the low and dose molecular weight compounds such as methane, ethane and ethylene need to be separated. Plant managers and engineer tend to be conservative and are unlikely to go to lower energy' oxidative dehydrogenation processes, as they have not been commercially implemented to any extent. Incremental expansion of a cracker or a cooling train is expensive. As plants expand, a method to increase capacity at a reduced cost is to install an oxidative dehydrogenation unit intermediate the cracker and the separation tram or combining an oxidative dehydrogenation reactor with the separation train (e.g., the C2 splitter) Recycled streams containing ethane and ethylene could pass through the oxidative dehydrogenation unit without requiring expansion of the cracker and potentially then pass to the separation stage without putting an undue load on the cooling train.

[0609] Techniques are provided for a chemical complex in which there is an oxidative dehydrogenation process to dehydrogenate etisaue to ethylene intermediate a chemicals cracker (e.g., a steam cracker) and tise associated downstream separation units. This will provide expansion capacity' at reduced operating costs. More particularly, in one aspect the overheads from the C2 splitter could be passed through the oxidative dehydrogenation unit to reduce the ethane content (polish the product stream). In some cases, the upper portion of the rectifying portion of the C2 splitter is used to reduce very' low' amounts of residual ethane in the ethylene. The technology of the present patent application may be applied to a new ethylene-manufacturing site (greenfield development) or could be a retrofit to an existing facility to expand capacity at a minimum cost. [0610] Techniques are provided for a (petrochemical complex composing a steam cracker composing a C2 splitter, the improvement comprising integrating into the complex intermediate the cracker and the C2 splitter a reactor for oxidative dehydrogenation of etisane in mixed stream comprising ethane and ethylene.

[0611] In certain embodiments, the oxidative dehydrogenation unit is integrated with the feed stream to the C2 splitter.

[0612] In certain embodiments, the oxidative dehydrogenation unit is integrated with the overhead stream from the C2 splitter.

[0613] In certain embodiments, the oxidative dehydrogenation unit is integrated with the bottom stream from the C2 splitter.

[0614] In certain embodiments, the oxidative de drogenation unit is integrated with the C ₂ splitter taking a feed from a lower tray from the C2 spliter and reluming the product to a higher tray in the C2 spliter.

[0615] In certain embodiments, the oxidative dehydrogenation unit is integrated with the feed stream to a hydrogenation unit to remove acetylene

[0616] In certain embodiments, the chemical complex is further comprised of one or more unit operations selected from the group consisting of absorption separation of ethane from ethylene, adsorption separation of ethane from ethylene, a high pressure polyethylene plant, a gas phase polyethylene plant, a slurry phase polyethylene plant, a solution phase polyethylene plant, an acetic acid plant, a vinyl acetate plant, an ethylene glycol plant, an ethanol plant, an ethylene halide plant, an ethanol dehydrogenation plant, an acetic acid dehydrogenation plant.

[9617] In certain embodiments, the oxidative dehydrogenation of ethane to ethylene is conducted at a temperature from 250 °C to 600 °C, preferably 300 °C to 550 °C, and a pressure from 0.5 to 100 psig (3.4 to 689.5 kPag) and has a productivity' of not less than 1000 g of olefin per kg of catalyst per hour. [0618] In certain embodiments, the oxidative dehydrogenation reaction has a selectivity of not less than 80% to produce the corresponding olefin.

[0619] In certain embodiments, the oxidative dehydrogenation catalyst is supported on inert porous ceramic membrane selected from oxides of titanium, zirconia, aluminum, magnesium, ytfria, lantana, silica and their mixed compositions to provide from 0.1 to 20 weight % of the catalyst and from 99.9 to 80 weight % of the porous membrane.

[0620] In certain embodiments, the oxidative deh drogenation reactor comprises an outer shell and one or more internal ceramic tubes defining a separate flow passage for ethane down the interior of the tubes and an annular passage between the external shell of the reactor and the ceramic tubes defining a flow path for an oxygen containing gas.

[0621 [ In certain embodiments, the ceramic tube further comprises an internal steel mesh and a external steel mesh.

[0622] In certain embodiments, the chemical complex further comprises an oil-based olefin paraffin absorption unit.

[0623] In certain embodiments, the chemical complex further comprises an adsorption olefin paraffin separation unit.

[0624] In certain embodiments, the adsorbent comprises one or more metals ions in the +1 oxidation state selected from the group consisting of silver and copper, although care must be exercised in the use of these compounds when separating streams containing acetylene due to the potential of forming explosive mixtures. [0625] In certain embodiments, the adsorbent is selected from the group consisting of synthetic or natural zeolites

[0626] In certain embodiments, the adsorbent is selected from the group consisting of ZSM-5, ETS-4, CTS-l, and ion-exchanged ETS-10.

[0627] In certain embodiments, the adsorbent is a metal dithiolene selected from the group of complexes of the formulae: (i) M[S ₂ C ₂ (R ^! R ² )] ₂ , and (ii) M[S ₂ C ₆ (R ³ R ⁴ R ⁶ R ⁷ )] ₂ . M is selected from the group consisting of Fe, Co, Ni, Cu, Pd and Pt, and R ¹ , R ² , R ³ , R ⁴ , R ⁵ , and R ⁶ are independently selected from the group consisting of a hydrogen atom, electron- ithdrawing groups including those that are or contain heterocyclic, cyano, cafboxylate, carboxylic ester, keto, nitre, and sulfonyl groups, hydrocarbyl radicals selected from the group consisting of Ci-C4, alkyl groups, Cs-Cg, alkyl groups, C ₂ -Cg, alkenyl groups and Ce-Cg aryl groups which hydrocarbyl radicals are unsubstituted or fully or partly substituted, preferably those substituted by halogen atoms.

[9628] In certain embodiments, the ethylene halide plant reacts ethylene, optionally in the presence of oxygen, with a halide to produced one or more products selected from the group consisting of ethyl chloride, ethylene chloride, ethylene dichloride, ethyl bromide, ethylene bromide, and ethylene dibromide.

[9629] In certain embodiments, the acetic acid plant oxidizes / hydrates ethylene from the oxidative dehydrogenation process, the steam cracker, or both to produce acetic acid.

[9639] In certain embodiments, acetic acid from the acetic acid plant is reacted with ethylene to produce vinyl acetate. [0631] In certain embodiments, immediately downstream of the oxidative dehydrogenation reactor there is a low-temperature (typically below the temperature of the oxidative dehydrogenation reaction) reactor to consume residual oxygen without consuming more than 3 «'eight % of the ethylene produced.

[0632] In certain embodiments, the fuel for the low-temperature reactor is selected from the group consisting of methane, hydrogen, carbon monoxide, and mixtures thereof and is added to the product stream from the oxidative deh drogenation reactor in an amount sufficient to consume residual oxygen

[0633] In certain embodiments, the low-temperature reactor uses a catalyst, which is a mixture of M O and CuMu Oi wherein the mixture has an empirical formula Cu-Mn _x O _p «'herein x is from 0.1 to 8 and p is a number to satisfy the valence state of the mixed catalyst.

[0634] The Catalyst System There are a number of catalysts described here, which may be used in accordance with the provided techniques. The following catalyst systems may be used individually or in combination. One of ordinary skill in the art would understand that combinations should be tested at a laboratory scale to determine if there are any antagonistic effects w'hen catalyst combinations are used.

[0635] The Support

[0636] There are several ways the oxidative dehydrogenation catalyst may be supported.

[0637] Figure 31 shows a schematic diagram of an oxidative dehydrogenation reactor containing three spaced apart fixed beds of catalyst.

[0638] In Figure 31, the ethane or ethane containing gas 3100 enters the reactor generally shown at 3101 by an inlet 3102 The ethane or ethane containing gas enters the first catalyst bed 3103. Oxygen or an oxygen containing gas 104 flows into a first space between the first catalyst bed 103 and a second catalyst bed 105 The oxygen flows into each bed. The stream of oxygen and partially reacted ethane or ethane containing gas flows into a second bed of catalyst 3105. Further, oxygen or oxygen containing gas 3106 flows into a second space between the second catalyst bed 3105 and a third catal st bed 3107. The reactants continue to react in the third catalyst bed and the resultant stream of ethylene flow s into collector (footer) 3108 and out exit 3109.

[0639] The Membrane

[0640] As noted previously, the support should have a low' surface area, preferably, less than 50 m ² /g, arid more preferably, less than 20 m ² /g. The support may be prepared by compression molding. At higher pressures, the interstices within the ceramic precursor being compressed collapse. Depending on the pressure exerted on the support precursor, the susface area of the support may be from about 20 to 10 m ² /g. The support; will be porous and will have a pore volume from about 0.1 to 3.0 rnl/g, typically, from 0.3 to 1.0 ml/g. The pore size of the ceramic may be small. Preferred pore size (diameter) ranges from about 3 to 10 nm. The small pore diameter is helpful in the ceramic membrane application as it helps maintain the pressure drop across the membrane so that a break in the membrane is readily detected by a sudden change in pressure. Additionally, the small pore diameter promotes a more uniform distribution of the reaction over the entire catalyzed surface of the membrane. That is, if larger pores are used, a majority of the oxygen tends to diffuse through the portion of the ceramic the oxygen containing gas initially contacts. The remaining portion of the ceramic is largely unused. [0641] The ceramic support may be prepared front the ceramic material using conventional techniques. For example, the starting material maty be cleaned, washed and dried (or spray dried) or produced front a sol/gel of the ceramic and w'here necessary groutsd or milled to the appropriate particle size. The powder may be subjected to benefication, such as, acid or base «'ashing to alter the pore size of the ceramic. [0642] The resulting powder is dried or calcined to remove associated water as noted above (water of hydration, etc.) and may be formed into a suitable substrate, preferably, tubular, by, for example, compression molding or isostatic compaction at pressures from about 5 to 200 MPag (725 to 29,000 psig), with or without a binder and sintering at temperatures to fuse the particles (e.g., at temperatures from about 0.5 to 0.75 of the melting temperature of the ceramic material.

[0643] Other techniques maty be used, such as, tape casting or slip casting of slurries and the subsequent “punching of’ the required shape, such as, circular, square or annular, etc. For example, annular sections could be “stacked” to produce a “tube”.

[0644] While a tube is generally considered cylindrical, it could have any cross section shapes, such as, square, rectangular, hexagonal or stars, etc. It the case of a non-cylindrical tube, wall sections could be made by slip casting and then hermetically joining the wall sections together to form a central passage defined by an outer ceramic wall. The joints need to be hermetically sealed to prevent oxygen coming in contact with the ethane feed and forming an explosive mixture. Glass cement or a ceramic cement or slip would be used for this purpose. A hermetic seal also needs to be at the ends of the tube where it enters and exits the reactor or joins to the steel parts of the reactor.

[0645] In some embodiments, once ihe ceramic tube is prepared, the catalyst may be deposited on the surface of the tube in contact with the ethane

[0646] The ceramic membrane may have a thickness from about 0.1 to 10 cm, typically, from 1 to 8 cm, preferably, from 2 to 7 cm.

[0647] While ceramics are strong, they can be brittle. It is preferred to have a supporting structure at least on one side, preferably, the outside of the ceramic tube. Most preferably, there is a support structure on the outside and inside of the tube. The structure should be in the form of a mesh or a web having holes there through to permit the oxygen containing gas to pass through the support and the ceramic to react at the surface of the tube bearing the catalyst. The support may be any material suitable for use at the reactor operating temperatures. From a cost point of view, a steel mesh is likely most cost effective. Preferably, the steel is a stainless steel. The support structure should provide sufficient integrity lo Ihe tube to permit a shutdown of ihe reactor, if the ceramic is breached (e.g., becomes cracked, ere.)

[0648] One or more tubes are then placed inside the reactor. In one embodiment, the reactor is designed to have a plug flow of feedstock (e.g., primarily, ethane) through a passage between ihe reactor shell and the ceramic tube and a flow' of oxygen containing gas through ihe ceramic lube. There are a number of possible arrangements. The reactor could comprise several shorter tubes placed end to end to provide a tube of appropriate length. Alternatively, the design could be similar to a core shell heat exchanger with a number of parallel tubes through which the oxygen containing gas is passed with and an enclosed shell providing a passage between the external wall of the reactor and the ceramic tubes defining a flow' path for the ethane. The flow' paths might be reversed (ethane on the interior and oxygen on the exterior of the tube).

[0649] An embodiment of a membrane (ceramic tube) oxidative dehydrogenation reactor is shown in Figure 6 and is described in more detail by the text associated with Figure 6. Another embodiment in which ethane or ethane gas enters the reactor is shown in Figure 7 and is described in more detail by the text associated with Figure 7. [065(1] The flows of the reactants may be concurrent or counter current (e.g., ethane up the outside of the tube and oxygen down the inside of the tube).

[0651] The feed to the reactor comprises two separate flows to opposite sides of a tube. In one embodiment, one flow, preferably, to the internal surface of the tube is an oxygen containing gas, which is selected from the group consisting of oxygen, mixtures comprising from 100 to 21 vol. % of oxygen and from 0 to 79 vol. % of one or tnore inert gases. Some inert gases may be selected from the group consisting of nitrogen, helium and argon and mixtures thereof. Preferably, the oxygen containing gas is air as it provides for a much simpler plant ope ratio a

[0652] The second flow, in some embodiments to the outside of the tube comprises one or more, CVCe, preferably (¼·€ _¨ paraffins, most preferably, pure or undiluted ethane or an ethane containing gas. Most preferably, the ethane should have a purity greater than 90 vol. %, preferably, greater than 95 vol %, most preferably, greater than 98 vol. %. However, it may be possible to operate with more dilute paraffin feeds, typically, comprising at least 60 vol. %, most preferably, not less than 80 vol. %of ethane and less than 40 vol. %, most preferably, less than 20 vol. % of one or more gases selected from the group consisting of methane, nitrogen, helium, argon and mixtures thereof. Preferably, the ethane containing gas is undiluted ethane as it provides for a much simpler plant operation and beter productivity (space-time yield).

[0653] The ratios of the gas components will be a function of the method of operating the reaction to reach either the complete consumption of oxygen, or complete consumption of ethane, or both. The further separation will include separation of ethylene from unreacted ethane or admixed gases (methane, CO _? ., inert gases, oxy gen). The oxygen containing gas flow rate has to be large enough to provide sufficient oxygen to the catalyst to provide the oxy gen needed for the oxidative dehydrogenation reaction. In one embodiment, in the ceramic membrane mode, the hydrocarbon stream passes over the oxidative dehydrogenation catalyst, optionally containing one or more metal oxides capable of releasing oxygen to the oxidative dehydrogenation catal st. The feed rate of oxygen gas should be sufficient to keep the catalyst active but low' enough to minimize carryover of oxygen into product olefin (ethylene). One can calculate the ratio of oxy gen to paraffin based on the stoichiometry of the reaction. However, the reaction will also be affected by the take up and release rate of the oxygen to and from the catalyst, because oxygen is fed to the opposite side of the membrane and is supplied to the active mixed oxide catalyst through tire porous ceramic membrane. The rate of oxygen supply is regulated by the pressure differential (PR) from the oxygen side of the ceramic varying typically from 0.05 to 0 5 atm. Typically, the molar ratio of hydrocarbon (paraffin) to oxygen feed may range from i ; 1 to 3 : ! , preferably, from 1.5:1 to 2.5:1. Given the foregoing, one of ordinary skill in the art will be able to determine the preferred ratio and flow' rates of the two gas flows for the ceramic membrane mode. The shutdown of the oxygen flow' results in fast but reversible loss of the ethane conversion.

[0654] Oxygen Scavenging

[0655] The amount of oxygen that is entrained in the product ethylene stream should be minimized for further processing. However, there will likely be some small amount of oxygen in ihe product stream. It is highly desirable that the oxygen be removed from the product stream prior to further processing of the product stream immediately downstream of the oxidative dehydrogenation reactor may be a low temperature reactor to consume residual oxygen without consuming more than about 3 wt. % of tise ethylene produced. Tints low temperature reactor, typically, uses a catalyst, which is a mixture of MroOs and CuMroO·;, the mixture having an empirical formula Cu-Mn _x O _p, where x is from O. i to 8, and p is a number to satisfy the valence state of the mixed catalyst. The low temperature oxygen scavenging reactor operates at temperatures less than or equal to 400 °C, typically from 100 °C to 400 °C. The fuel for the low temperature reactor may be selected from the group consisting of methane, hydrogen, CO and mixtures thereof which may be either added to or present in the paraffin 1 ^' eed stream or added to the product stream from the oxidative dehydrogenation reactor in an amount sufficient to consume residual oxygen. In some embodiments, the oxygen scavenger, sometimes referred to as an afterburner, may be followed by a number of other process steps including a water wash, CO removal, product separation which may include the typical C splitter or other means to separate ethylene from ethane One such embodiment is shown in Figure 24.

[0656] Separation of the Product Stream

[0657] The ethylene, preferably , after passing through the oxygen scavenger and a drier, may be fed to a C splitter downstream of the cracker to separate ethylene and ethane.

[0658] There are a number of options to combine a C splitter and an oxidative dehydrogenation unit.

[0659] Figure 24 is a schematic diagram of a conventional C splitter 2400 (cryogenic distillation tower).

Feed 2420, a mixture predominantly of ethylene and ethane, is fed to the column 2421. An overhead stream of ethylene 2422 leaves the top of the column 2421 and passes through a condenser 2423 to a reflux drum 2424 and a pump 2425 The condensed and re-pressurized stream is split into an ethylene product stream 2427 and a high purity stream 2426 that is fed back to the upper trays of the splitter 2421. At the bottom of the splitter 2421, a stream of ethane 2428 passes through a pump 2429 to two heaters 2430 and 243 i and is ready for further processing, such as, recycle to the cracker. Towards the bottom of the C spliter 2421 , a stream of ethane 2432 is taken and passed through a reboiler 2433 and recy cled back to the spliter 2421 This is considered a base case against which the provided techniques may be evaluated.

[0660] Figure 25 is a schematic diagram of an embodiment 2500 of a C splitter integrated with an oxidative dehydrogenation unit at the overhead stream (ethylene product stream). In Figure 25, a feed 2541, predominantly of ethylene and ethane, is fed to splitter 2542. A relatively pure stream of ethylene 2543 exits the top of the C splitter 2542. A portion of the stream 2543 is fed to a condenser 2547, a reflux drum 2548 and a pump 2549 and fed back to the C spliter 2542. The remaining portion of product stream 2543 is feed to a heater 2544 and then to the oxidative dehydrogenation unit 2545 resulting in a stream 2546 of ethylene and traces of CO At the botom of spliter 2542, near or at the last tray, a stream 2550 of ethane and co-produets is taken A portion of the product 2550 is passed through a reboiler 2554 and the vaporized stream 2550 is recycled to the C splitter 2542. The other portion of the ethane product stream is fed to a pump 2551 and then through heaters 2552 and 2553 and the ethane stream 2550 is ready for further processing.

[0661 ] Figure 26 is an example of an embodiment 2600 of an oxidative dehydrogenation reactor integrated with the botom product stream (ethane) from a C spliter. In Figure 26, a feed 2660 of ethylene and ethane is fed to the C spliter 2661. The overhead stream 2662 largely ethylene is fed to a condenser 2663, a reflux drum 2664 and a pump 2665 A portion of the product ethylene stream 2662 is fed to the C splitter 2661 and a portion of the stream is available for further mixing with the product stream of oxidative dehydrogenation reactor 2670 integrated with the bottom stream from the C spliter 2661 At the botom of the C splitter 2661, a relatively pure stream of ethane 2666 is fed to pump 2667 and heaters 2668 and 2669. The stream is then fed to oxidative dehydrogenation unit 2670 and the resulting stream of ethylene 267 i is combined with overhead stream 2662 to form product ethylene stream 2672. Near the bottom of the C2 spliter, above where stream 2666 is taken off, a stream of ethane 2673 is passed through a reboiler 2674 and fed back to the splitter 2661.

[0662] Figure 28 is an embodiment 2800 showing integration of an oxidative dehydrogenation unit within the C splitter. In figure 28, a stream of ethylene and ethane 2880 is fed to C2 splitter 2881. The overhead stream of ethylene 2882 is led to a condenser 2883, reflux dram 2884 and pump 2885. A portion of stream 2882 is fed back to the C2 splitter. A portion of the ethy lene stream 2886 is available for downstream processing (e.g., polymerization to polyethylene, conversion to acetic acid, vinyl acetate). Towards the middle of the C2 splitter 2880, a mixed stream 2887 of ethylene and ethane is withdrawn. The stream passes through a heater 2888, and, depending on the pressure of the stream, a pressure reduction device 2889, for example, a turbo-expander. Stream 2887 then passes through oxidative dehydrogenation unit 2890. The product stream 289!, having a higher ethylene content tisan stream 2887, then passes through a compressor 2892 and chiller 2893 and is fed back to the C2 splitter 2881. At the botom of the C2 splitter 2881, a relatively pure stream of ethane 2894 is removed and fed to pump 2895 and heaters 2896 and 2897. The resulting stream 2894 is then ready for further processing (e.g., acetic acid). Stream 2898 is fed to a reboiler 2899 and returned to the splitter 2881.

[0663] Figure 29 shows an embodiment 2900 where the oxidative dehydrogenation unit is integrated into the feed from the cracker to the C2 splitter. In this embodiment, the ethylene and residual ethane product stream 29100 from the cracker is split at valve 2901. A portion of the feed 29100 is fed to oxidative dehydrogenation unit 2902. The resulting stream, which is higher in ethylene, is fed through a compressor 2903 and then a condenser 2904 and to the C2 splitter 2905. At the top of the C2 splitter 2905, an o verhead stream 2906 of high purity ethylene is fed to a condenser 2907, reflux drum 2908, pump 2909 and back to C2 splitter 2905 A portion of the ethylene stream 2906 is available for downstream processing. At the bottom of the C _? splitter 2905, a relatively pure stream of ethane 2910 is removed and fed to pump 2911 and heaters 2912 and 2913. The resulting stream 2910 is then ready for further processing (e.g., acetic acid). Stream 2914 is fed to a reboiler 2915 and relumed to the splitter 2905.

[0664] Figure 30 is a schematic diagram of an embodiment 3000 in which an oxidative dehydrogenation unit is integrated downstream of the acetylene hydrogenation unit. In Figure 30, a feed 3000 predominantly comprising about 60 mole % ethylene and 40 mole % ethane from the cracker passes through a heaters 3001 and 3002 The feed 3000 then passes to parallel hydrogenation units 3003 and 3004 to produce a stream 3007 having an acetylene content less than about 1 ppm. The stream 3007 passes through chiller 3005 to a greets oil knock out drum 3006 There ase two lines from the knockout dram 3006. One line goes through valve 3015 to two driers (3013, 3014). The other line passes through valve 3008 to oxidative dehydrogenation unit 3009. Stream 3010 from the oxidative dehydrogenation unit 3009 lias a higher ethylene content than the stream from the hydrogenation units 3003 and 30Q4. Stream 3010 passes through a compressor 3011 and a cooler 3012 and is mixed with the stream going to driers 3013 and 3014. By controlling valves 3015 and 3008, the amount of feed to the oxidative dehydrogenation unit 3009 may be controlled from 0 to 100%.

[0665] Separation means alternatively or in addition to a C2 splitter can be implemented.

[0666] One method of separation of a product stream of ethylene and ethane is by absorption. The gaseous product stream comprising primarily ethane and ethylene may be contacted in a counter current flow with a heavier paraffinic oil, such as, mineral seal oil or medicinal white oil at a pressure up to 800 psig (about 5.5xl0 ³ kPag) and at temperatures from about 25 °F to 125 °F (about -4 °C to about 52 °C). The ethy lene and lower boiling components are not absorbed into the oil. The ethane and higher boiling components are absorbed into the oil. The ethylene and lower boiling components may then be passed to tire C splitter. The absorption oil may be selectively extracted with a solvent, such as, furfural, dimethyl formamide, sulfur dioxide, aniline, nitrobenzene, and other known solvents to extract any heavier paraffins.

[0667] Another separation method is an adsorption method. The adsorbent preferentially adsorbs one of the components in the product stream. The adsorption method typically comprises a train of two or more adsorption units so that when a unit has reached capacity the feed is directed to an alternate unit while the fully loaded unit is regenerated typically by one or more of a change in temperature o r pressure or both.

[0668] There is a significant amount of art; on the separation of ethylene atsd ethane using silver or copper ions in their +1 oxidation state. The olefins are preferentially absorbed into a complexing solution that contains the complexing agent selected from silver (I) or copper (I) salts dissolved in a solvent. Some silver absorbents include silver nitrate, silver fluoroborate, silver fluorosilicate, silver hydroxyfiuoroborate, and stiver trifiuoroacetate. Some copper absorbents include cuprous nitrate; cuprous halides such as cuprous chloride; cuprous sulfate; cuprous sulfonate; cuprous caiboxydates; cuprous salts of fluorocafboxylic acids, such as, cuprous trifiuoroacetate and cuprous peifluoroacetate; cuprous fluorinated aceiyiacetonate; cuprous he.xafhioroacetylacetonate; cuprous dodecylbenzenesulfonate; copper-aluminum halides, such as, cuprous aluminum tetrachloride; CuAlCltyCT ; CUAIC ₂ H ₅ CI ₃ ; and cuprous aluminum cyanotrichloride. If the product stream has been dried prior to contact with the liquid absorbent, the absorbent should be stable to hydrolysis.

The complexing agent preferably is stable and has high solubility in the solvent. After one absorbent solution is substantially loaded, the feed of product stream is switched to a further solution. The solution of absorbent, which is fully loaded, is then regenerated through heat or pressure changes or both. This releases the ethylene. [0669] As noted previously, care needs to be taken in using these types of materials to avoid detonations.

[0670] In some cases, supports such as zeolite 4A, zeolite X, zeolite Y, alumina and silica, may be treated with a copper salt, to selectively remove carbon monoxide and/or olefins from a gaseous mixture containing saturated hydrocarbons (i.e., paraffins), such as, ethane and propane.

[0671] Similarly, copper salts and silver compounds may be used, arid are supported, alternatively, on silica, alumina, MCM-41 zeolite, 4 A zeolite, carbon molecular sieves, polymers such as Amber!y st-35 resin, and alumina to selectively adsorb olefins from gaseous mixtures containing olefins and paraffins. Both kinetic and thermodynamic separation behavior was observed and modeled. The adsorption of the olefin takes place at pressures from 1 to 35 atmospheres, preferably, less than 10 atmospheres, most preferably, less than 2 atmospheres at temperatures from 0 to 50 °C, preferably from 25 to 50 °C and the desorption occurs at pressures from 0.01 to 5 atmospheres, preferably, 0.1 to 0.5 at temperatures from 70 °C to 200 °C, preferably, from 100 °C to 120 °C.

[0672] In certain embodiments, the adsorbent may be a physical adsorbent selected from the group consisting of natural and synthetic zeolites without a silver or copper salt.

[0673] In general, the adsorbent may be alumina, silica, zeolites, carbon molecular sieves, etc. Typical adsofbents include alumina, silica gel, carbon molecular sieves, zeolites, such as, type A and type X zeolite, type Y zeolite, etc. The preferred adsorbents are type A zeolites, and the most preferred adsorbent is type 4A zeolite. [0674] Type 4A zeolite, i.e., the sodium form of type A zeolite, has an apparent pore size of about 3.6 to 4 Angstrom units. Tints adsorbent provides enhanced selectivity atsd capacity its adsorbing ethylene from ethylene- ethane mixtures and propylene from propylene-propane mixtures at elevated temperatures. This adsorbent is most effective for implementation with the provided techniques when it is substantially unmodified, i.e., when it has only sodium ions as its exchangeable cations. However, certain properties of the adsorbent, such as, thermal and light stability, may be improved by partly exchanging some of the sodium ions with other cations (other than silver or copper). Accordingly, it is within the scope of a preferred embodiment to use a type 4A zeolite in which some of the sodium ions attached to the adsorbent are replaced with other metal ions, provided that the percentage of ions exchanged is not so great that the adsorbent loses its type 4A character. Among the properties that define type 4 A character are the ability of the adsorbent to selectively adsorb ethylene fro ethylene-ethane mixtures and propylene from propylene-propane gas mixtures at elevated temperatures, and to accomplish this result without causing significant oligomerization or polymerizatio of the alkenes present in the mixtures. In general, it lias been determined that up to about 25% (on an equivalent basis) of the sodium ions in 4A zeolite can be replaced by ion exchange with other cations without divesting the adsorbent of its ty pe 4A character. Cations that may be ion exchanged with the 4A zeolite used in the alkene -alkane separation include, among others, potassium, calcium, magnesium, strontium, zinc, cobalt, manganese, cadmium, aluminum, cerium, etc. When exchanging other cations for sodium ions it is preferred that less than about 10 percent of the sodium ions (on an equivalent basis) be replaced with such other cations. The replacement of sodium ions may modify the properties of the adsorbent. For example, substituting some of the sodium ions with other cations may improve the stability of the adsorbent.

[0675] A particularly preferred zeolite is ZSM-5.

[0676] In addition to zeolites, there are a number of titanium homologues referred to as ETS compounds.

[0677] A large pore diameter titanosilicate designated “ETS- 10” may be used. In contrast to ETS-4 and

CTS-1 (referenced below), the large pore titanosilieate material, ETS- 10, which lias pore diameters of about 8 A, cannot kinetically distinguish light olefins from paraffins of the same carbon number. However, high degrees of selectivity may be possible for the separation of ethylene from ethane using as prepared ETS-10 zeolites. Na- ETS-10 may be capable of selectively adsorbing ethylene from a mixture of ethylene and ethane under thermodynamic conditions, even at ambient temperature. Although, selectivity for ethylene adsorption using Na- ETS-10 has been measured to be high at ambient temperature, the adsorption isotherms for ethylene and ethane had highly rectangular shapes consistent with a low-pressure swing capacity. Consequently, Na-ETS-10 is not readily applicable to pressure swing absorption processes (PSA), at least at lower or ambient temperatures.

[0678] Howe ver, cationic modificatio of as prepared Na-ETS- i 0 provides an adsorbent for the PSA separation of olefins and paraffins having the same number of carbon atoms, at ambient temperatures. The mono-, di- and tri-valent cations are selected from the group 2-4 metals, a proton, ammonium compounds, and mixtures thereof. Some specific non-limiting examples of mono-, di, or tri-valent cations that can be implemented with the provided techniques include, Li , K ⁺ , Cs ⁺ , Mg ^{2 t} , Ca ^2t , Sr ²¹ , Ba ²⁺ , Sc· ¹¹ , Y ³⁺ , La ³⁺ , Cu ⁺ , Zn ²⁺ , Cd ^2t , Ag ⁺ , Au ⁺ , H ⁺ , NHA, and NKV where R is an alkyl, aiyl, alkylaiyl, or aiylalkyl group. The cationic modifiers are generally added to unmodified Na-ETS- 10 in the form of a salt or an acid The anionic counterion associated with the cationic modifier is not specifically defined, provided that it does not adversely affect the modification (i.e., cation exchange) reactions. Suitable anions include but are not limited to acetate, carboxylase, benzoate, bromate, chlorate, perchlorate, clioriie, citrate, nitrate, nitrite, sulfates, and halide (F, Cl, Br, I) and mixtures thereof. Suitable acids include inorganic and organic acids, with inorganic acids being preferred. [0679] Heat treatment of ETS-4 may give a controlled pore volume zeolite material, dubbed “CTS-1” which is a highly selective absorbent for olefin/paraffin separations. The CTS-1 zeolite, which has pore diameters from about 3-4A, selectively adsorbed ethylene from a mixture of ethylene and ethane through a size exclusion process. The pore diameter of CTS-1, allow ed diffusion of ethylene, while blocking diffusion of ethane which was too large to enter the pores of the CTS-1 zeolite, thereby providing a kinetic separation. The CTS- 1 adsorbent was successfully applied to a PSA process in which ethylene or propylene could be separated from ethane or propane, respectively.

[06811] The above adsorbents may be used in pressure swing adsorption units. Typically, the range of absolute pressures used during tise adsorption step can be from about 10 kPa to about 2,000 kPa, (about 1 5 to about 290 pounds per square inch (psi)) preferably from about 50 kPa to about 1000 kPa (from about 7 2 to about i 45 psi) The range of pressures used during the release of adsorbate (i.e , during the regeneration step) can be from about 0.01 kPag to about 150 kPag (about 0.0015 to about 22 psig), preferably, from about 0.1 kPag to about 50 kPag (about 0.015 to about 7.3 psig). In general, the adsorption step can be carried out at from ambient temperatures to above about 200 “C, preferably less than 150 °C, most preferably, less than 100 °C, provided that the temperatures do not exceed temperatures at which chemical reaction of the olefin, such as, a oligomerization or polymerization takes place.

[0681] Anothe r class of adsorbents is ionic liquids Olefins and paraffins can be separated using ionic liquids of the formula a metal dithiolene selected from the group of complexes of the formulae: 2 where M is selected from the group consisting of Fe, Co, Ni, Cu, Pd and Pt; and R ¹ , R ² , R ³ , R ⁴ , R ⁵ , and R ⁶ are independently selected from the group consisting of a hydrogen atom, electron-withdrawing groups including those that are or contain heterocyclic, cyano, carboxylate, carboxylic ester, keto, nitro, and sulfonyl groups, hydrocaibyl radicals selected from the group consisting of alkyl groups, Cs-Ce, alkyl groups, Cj-C» alkenyl groups and Ce-Cg aryl groups which hydrocafbyl radicals are unsubstituted or fully or partly substituted, preferably those substituted by halogen atoms. The ionic liquid may be used with a non-reactive solvent or co solvent. The solvent may be selected from the group of conventional aromatic solvents, typically toluene. Adsorption pressures may range from 200 psig to 300 psig (1.3xl0 ³ to 2x10’ kPag), preferably, below 250 psig ( 1.7x10 ’ kPag) and adsorption temperatures may range from ambient to 200 °C, preferably, below 150 °C, and the olefin may be released from the ionic liquid by one or more of lowering the pressure by at least 50 psig (3.4xl0 ² kPag) and increasing the temperature by not less than 15 °C

[0682] Downstream Unit Operations in the Complex

[0683] The complex may comprise one or more unit operations using ethylene, ethane, or both as a feed stream.

[0684] The further unit operations may be one or more of the following processes individually or in combination: a high pressure polyethylene plant; a gas phase polyethylene plant; a slurry phase polyethylene plant; a solution phase polyethylene plant; an acetic acid plant; a vinyl acetate plant; an ethylene glycol plant; an ethanol plant; an ethylene halide plant; an ethanol dehydrogenation plant; and an acetic acid dehydrogenation plant.

[0685] Ethylene Polymerization

[0686] The ethylene eats be polymerized. There are a number of well-known methods for polymerizing ethylene.

[0687] The process can be a high-pressure process. Typically, the pressures range from about 80 to 3 iO MPag (e.g., about 11 ,500 psig to about 45,000 psig) preferably from about 200 to 300 MPag (about 30,000 psig to about 43,500 psig) and the temperature ranges from 130 °C to 350 °C, typically, from 150 °C to 340 °C. The supercritical ethylene together with one or more of initiators, chain transfer agent and optional comonomers are fed to a high-pressure reactor. A non-limiting example of a high-pressure reactor is a tubular reactor. Tubular reactors may have a length from about 200 m to about 1500 in, and a diameter from about 20 mm to about 100 mm. The residence time is generally quite short, in the order of seconds to less than 5 minutes.

[0688] Solution and slurry polymerization processes are fairly well known in the art. These processes are conducted in the presence of an inert hydrocarbon solvent/dilirent typically a C4-12 hydrocarbon that may be unsubstituted or substituted by a C alkyl group, such as, butane, pentane, hexane, heptane, octane, cyclohexane, methylcyclohexane, or hydrogenated naphtha. An alternative solvent is IsoparE (C«-Ci2 aliphatic solvent, Exxon Chemical Co.).

[0689] The polymerization may be conducted at temperatures from about 20 to about 250 °C. Depending on the product being made, this temperature may be relatively low', such as, from 20 °C to about 180 °C, typically, from about 80 °C to 150 °C, and the polymer is insoluble in the liquid hydrocarbon phase (diluent) (e.g., a slurry' polymerization). The reaction temperature may be relatively higher from about 180 °C to 250 °C; preferably, from about 180 °C to 230 °C, and the polymer is soluble in the liquid hydrocarbon phase (solvent). The pressure of the reaction may be as high as about 15,000 psig for the older high-pressure processes o r may range from about 15 to 4,500 psig.

[069(1] The polymerization conld be gas phase, either fluidized bed or stirred bed. In the gas phase polymerization of a gaseous mixture comprising from 0 to 15 mole % of hydrogen, from 0 to 30 mole % of one or more Cs-Cg alpha-olefins, from 15 to 100 mole % of ethylene, and from 0 to 75 mole % of an inert gas at a temperature from 50 °C to 120 °C, preferably, from 75 to about 110 °C, and at pressures, typically, not exceeding 3447 kPag (about 500 psig), preferably, not greater than 2414 kPag (about 350 psig).

[0691] Suitable olefin monomers include, but are not limited to, ethylene and C3-C10 alpha olefins, which are unsubstituted or substituted by up to two Ci-Cg alkyl radicals. Illustrative non-limiting examples of such alpha olefins are one or more of propylene, 1-butene, I-pentene, 1-hexene, 1-heptene, 1 -octene, and l-decene. The polymers described here have a wide range of molecular weigh! distribution (Mw/Mn or po!ydispersity). The molecular weight distribution may be controlled from about 2 5 to about 30

[0692] The polyethylene polymers which may be prepared, typically, comprise not less than 60, preferably, not less than 70, most preferably, not less than 80 weight % of ethylene and the balance of one or more C3-C10 alpha olefins, preferably, selected front the group consisting of 1 -butene, 1-hexene and 1 -octene. [0693] The catalyst used in the solution, slurry and gas phase polymerization may be one or more of chromium catalyst (Phillips type catalysts), Ziegler Natta type catalyst, and single site type catalysts including metallocene catalysts, constrained geometry' catalysts, and bulky ligand heleroatom catalyst (e.g„ phosphinimine catalysts), the catalyst are used with one or more activators, such as, aluminum halides, alkyl and oxalkyl compounds or MAO orborales.

[0694] In gas phase and slurry polymerizations, the catalyst and, typically, lire activators are on a support such as alumina or silica.

[0695] Acetic Acid Unit

[0696] Ethylene or ethane or a mixture thereof may be oxidized to product acetic acid, which may be reacted ith further ethylene to produce ethyl acetate, which may then be converted to vinyl acetate. Ethylene recovered in the separation processes noted above may be fed to an oxidation reactor together with oxygen and or water in a weight ratio from 1:0.1 - 250 by weight, such as 1:0.1 - 100 or 1 :0 -150 but preferably in a ratio 1:0.1 -10 by weight in the presence of a supported catalyst. The oxidation reaction of this step may suitably be carried out at a temperature in the range from 100 to 400 °C, typically, in the range 140 °C to 350 °C at atmospheric or superatmospheric pressure, for example, in the range from 5 to 27 barg (50 to 270 kPag). There are a number of catalysts, which may be used as in this type of reaction. Typically, the catalysts compose molybdenum and tungsten with one or more transition metals having an atomic number from 44 to 47 and 77 to 79.

[0697] The resulting acetic acid may be fed to a further oxidation reactor together with ethylene to form ethyl acetate or with ethylene and an oxygen containing gas to form vinyl acetate.

[0698] Acetic acid may also be dehydrogenated to produce ethylene. In this case, acetic acid from other sources, such as, fermentation, could be dehydrogenated to produce ethylene.

[0699] Ethylene Epoxide

[0700] Ethylene oxide is ty pically produced by a direct oxidation process in which ethylene is directly oxidized with air or purified oxygen (95% or greater) over a catalyst, typically, silver silicate, but on occasions elemental silver may be used, on a silica support (or co -precipitated with the silica support). The catalyst may contain activators or chemicals to reduce coking. The reaction occurs at temperature from 100 C to 300 °C, typically, from 140 °C to 250 °C, preferably, less than 200 ^C, C The pressure may be from about 7 psig (about 50 KPag) to about 300 psig (about 2. 1 X10 ³ kPag). It is even more preferable to use a pressure from about 15 psig (about 304 kPag) to about 100 psig (6.9 XiO ² kPag). Typically, the space velocity may range from about 10 hr ¹ to about 15,000 hr ¹ . Preferably, the space velocity is in the range from about 10 hr ¹ to about 6000 hr ¹ . More preferably, the space velocity is in the range from about 50 hr ¹ to about 3000 hr ¹ . U.S. Patent No. 4,845,253 issued July 4, 1989 to Bowman assigned to The Dow Chemicals Company discloses one such process, the contents of winch are herein incorporated by reference.

[0701] Ethylene Glycol Unit

[0702] Ethylene epoxide is an intermediate for a number of downstream derivatives Ethylene epoxide maty be converted to ethylene glycol by reacting ethylene oxide with C<½ in a presence of a catalyst, such as, alkali halides, quaternary ammonium halides, and quaternary phosphonium halides, to produce ethylene carbonate. The ethylene carbonate may be converted to ethylene glycol by reaction with water, typically, less than about 2: 1 weight ratio of water to carbonate in the presence of a base (Na COs).

[0703] Ethylene glycol may be converted into a number of other chemically useful compounds such as PET andPHET.

[0704] Ethanol Unit

[0705] The gas phase direct hydration of ethylene to ethanol may be conducted over a solid catalyst, which is a porous substrate, typically, clay, silica, or alumina impregnated with phosphoric acid. In this gas phase hydration process, it is typical to provide a mole ratio of about 0.4 to 0.8 mole of water per mole of ethylene. In some processes, phosphoric acid Is added to the feed lo make up for catalyst losses during the process. The reaction may be conducted at temperatures front about 235° to 250 °C and at pressures from about 700 psig to 1200 psig ((4 2X10 ³ kPag to about 8.2X10 ³ kPag).

[0706] Ethanol Dehydration Unit

[0707] It will be recognized by those skilled in the art that in jurisdictions where there is a good supply of fermentable organic material (e.g., sugar cane) ethanol can be produced by fermentation and subsequently dehydrated over for example sulphuric acid to produce ethylene.

[0708 [ Ethylene Halide Unit

[0709] The complex can contain a unit operation for the halogenation of ethylene to vinyl chloride or to ethylene chloride (EDC). EDC may be obtained by the direct halogenation or oxy halogenation of ethy lene, optionally, in foe presence of oxygen. The direct halogenation may take place in the gas phase by reaction between ethylene and a gaseous halide (e.g , HC!) in the presence of a catalyst (FeC¾). This is an exothermic reaction and heat needs to be removed from the reactor. In the oxy halogenation process, oxygen and water are also present in the reactor and the catalyst component is Cud

[0710] The following examples are merely illustrative and are not intended to be limiting. Unless otherwise indicated, all percentages are by weight and Portland cement is used unless otherwise specified.

[0711] EXAMPLES

[0712] Example: Base case C splitter (Figure 2)

[0713] Ethylene and ethane may be separated via cryogenic distillation; an exatnple base case of ethylene and ethane separation via ctyogerac distillation is shown in Figure 24. In this exatnple, the Ci spliter feed is a 60% vapour fraction mixture comprised of 60 mole % ethylene and 40 mole % ethane at 1600 kPa. The feed stream enters the column on approximately tray 60, where trays are numbered front the top of the column down. The distillation column contains approximately 100 trays with 80% Murphree tray efficiency in the example shown. At the column pressure, the saturation temperature of pure ethylene is -37 °C and the saturation temperature of ethane is -16 °C. The reflux ratio is 3 6, with a resulting ethylene distillation purity of 99 95 mole % and ethane bottoms purity of 99 5 mole %. The overhead condenser fully condenses the ethylene distillate and reflux, and requires a thermal duty of 35 MW The kettle-type reboiler requires a thermal duty of 28 MTV. [0714] The example process conditions for the base case are summarized in Table D 1.

[0715] TABLE D1 : Example of C spliter base case process conditions

[0716] The base case is developed for an older C spliter in operation at a NOVA Chemicals facility at Joffre, Alberta, Canada. When modeled using AspenTech Aspen Plus® software, the model-predicted production and heat/energy balance is not less than 95% of the actual operation of the plant

[0717] In the following examples, the oxidative dehydrogenation unit was modeled on that of Figure 9, a membrane reactor, using AspenTech Aspen Plus® software.

[0718] Example: Oxidative dehydrogenation integrated with the C spliter overhead stream (Figure 25) [0719] In this example, the oxidative dehydrogenation unit and C splitter operation were modeled. The ethylene product purity was increased to at least 99.9 mole %, more preferably to 99.95 mole %.

[0720] By decreasing the ethylene purity of the overhead stream from 99.95 mole %, as shown in the base case example, to 95 mole % and decreasing the reflux rate accordingly, the ethylene distillate rate can be increased by approximately 6%. This process configuration is shown in Figure 25. The overhead condenser duty for this example case is 25 MTV and the reboiler duty is 27 MW, resulting in a 19% total thermal energy savings compared with the base case. The decrease in overhead condenser duty is due to the decrease in the reflux rate and because the condenser is condensing a smaller mass flow rate of ethy lene product stream as compared to the base case. The auxiliary thermal and pumping duty required to condense and pressurize the ethylene product stream from the oxidative dehydrogenation unit from 100 kPag and 30 C to a saturated liquid at 1560 kPag is approximately 15 MW, therefore, the integrated process requires more energy than the base ease. However, this process configuration allows for debottlenecking of the existing column by increasing the rate of ethylene production per unit feed and allows for an increase in column capacity by decreasing the reflux ratio and the required capacity of the overhead condenser. The results of the modeling are set forth in Table D2.

[0721] TABLE D?,: Example of oxidative dehydrogenation integration with C _? . spliter overhead stream process conditions

[0722] Example: Oxidative dehydrogenation unit integrated with the bottom stream of the C splitter (Figure 26)

[0723] In this example, oxidative dehydrogenation technology is modeled using AspenTeeh Aspen Plus®) to debottleneck an existing C splitter (base case configuration of the C splitter is shown in Figure 2). In this example, the oxidative dehydrogenation unit increases the ethylene purity to at least 99 9 mole %, more preferably to 99.95 mole %.

[0724] By decreasing the ethane purity of the bottoms stream from 99.5 mole %, as shown in the base case example, to 89.5 mole % and decreasing the reflux rate accordingly, thermal duty required in the column can be reduced. An oxidative dehydrogenation unit could be applied to convert ethane in the bottoms stream to ethylene as shown in Figure 26. The overhead condenser duty for tins example case is 33 MW and the reboiler duty is 26 MW, resulting in a 6% total thermal energy savings compared with the base case. The decrease in overhead condenser duty is due to the decrease in the reflux rate and because the condenser is condensing a smaller mass flow rate of ethylene product stream as compared to the base case. The example conditions are summarized in Table D3. This process configuration allows for debotlenecking of the existing column by increasing the rate of ethylene production per unit feed and allows for an increase in column capacity by decreasing the reflux ratio and the required capaci ty of the overhead condenser and botoms reboiler.

[0725] Figure 27 illustrates that even 18 mole % of ethylene in the C splitter botom product does not significantly affect the selectivity of the oxidative dehydrogenation process.

[0726] The energy savings reported in the examples are from the separation area only, and it should also be considered that the ethane converted an oxidative dehydrogenation unit in th s example will not be recycled back to steam cracking furnaces, resulting in up to 40% energy savings in furnace operation as well as increasing the cracking furnaces throughput by up to 40% as a result of converting recycled ethane in an oxidative dehydrogenation unit and not recycling it back to tise furnaces.

[0727] TABLE D3: Example of oxidative dehydrogenation integration with C2 spliter bottoms stream process conditions

[0728] Example: Integration of an oxidative dehydrogenation unit between the stages of the C2 spliter (Figure 28)

[0729] The process of Figure 28 was modeled using the AspenTech Aspen Plus® software to analyze the integrated system behavior.

[0730] In this example, oxidative dehydrogenation technology is used to debottleneck the C2 splitter whereby a sidedraw is taken from the C2 splitter and the ethylene content of this slipstream is increased by at least 25%. This ethylene-enriched sidedraw is returned as a secondaty feed stream to the C2 splitter at a tray above the sidedraw tray.

[0731] In tliis example, approximately 20% of the feed molar flow rate is taken as a sidedraw from a tray with 20 mole % ethylene and 80 mole % ethane composition, and is converted in an oxidative deh drogenation unit. The primary feed is composed of 60 mole % ethylene and 40 mole % ethane. With 50% conversion of ethane to ethylene and 0% conversion of ethylene to other products, the 60 mole % ethylene and 40 mole % ethane sidedraw is compressed and condensed to the column conditions existing on the same-composition tray. The process conditions simulated for this integration example are summarized in Table D4.

[0732] The reflux ratio required to achieve 99.95 mole % ethylene In the overhead product can be decreased by approximately 2%, however, due to the Increased distillate rate in this example, the reflux rate must be Increased by approximately 13% to maintain purity specifications. The boil-up rate required to minimize the ethylene content in the bottoms stream must be increased by approximately 17%. The ethylene distillate mass flow rate could be increased by approximately 15% in this example. The increase in ethylene produced per unit feed into the C2 splitter would result in an increase in ethylene separation capacity in the C2 splitter.

[0733] TABLE D4: Example of oxidative dehydrogenation integration between the stages of the C2 spliter process conditions

107341 Example: Integration of an oxidative dehydrogenation unit with the feed to a C _? splitter (Figure 29)

[0735] In tills example, the process of Figure 29 was modeled using AspenTech Aspen Plus®.

[0736] A slipstream is taken from the primary feed to the C2 splitter and the ethylene content of this slipstream is increased by at least 25%. This ethylene -enriched slipstream is returned as a secondary' feed stream to the C2 splitter at a tray above the primary' feed tray.

[0737] In this example, approximately 20% of the feed molar Slow rate is taken as a slipstream and is hydrogenated in ail oxidative dehydrogenation unit. The Seed is composed of 60 mole % ethylene and 40 mole % ethane. With 50% conversion of ethane to ethylene and 0% conversion of ethylene to other products, the 80 mole % ethylene and 20 mole % ethane oxidative dehydrogenation product stream is compressed and condensed to the column conditions existing on the same-composition tray. The process conditions simulated for this integration example are summarized in Table D5.

[0738] The reflux ratio required to achieve 99.95 mole % ethylene in the overhead product can be decreased by approximately 10%, resulting in negligible increase in the reflux rate required in addition, negligible increase in the boil-up rate is required to minimize the ethylene content in the bottoms stream. The ethylene distillate mass flow rate could be increased by approximately 7% in this example. The increase in ethylene produced per unit feed into the C spliter would result in an increase in ethylene separation capacity in the C2 splitter.

[0739] TABLE D5: Example of oxidative dehydrogenation integration with the feed stream of the C2 splitter process conditions

[0740] Example: Integration of an oxidative dehydrogenation unit base with the acetylene removal unit upstream from driers (Figure 30)

[0741] In tliis example, the feed stream to an oxidative dehydrogenation unit is the product from an acetylene hydrogenation reactor in a steam cracking plant. In this case, the ethane, winch is present in the feed, is dehydrogenated at a conversion of at least 60% more preferably 80% _> , more preferably 99.5%.

[0742] Example: Integration of an oxidative dehydrogenation unit with an oil refiner)'

[0743] In this example, the feed stream to an oxidative dehydrogenation unit is an ethane/ethylene mixed stream from an oil refiner)', which may contain but is not limited to the mixed C2 fraction from FCC, hydrocracking and hydrotreating operations. The ethylene content its this mixture can he from 8 volume % to 80 volume % In this case the ethane, which is present in the feed, is dehydrogenated at a conversion of at least 60% more preferably 80%, more preferably 99.5%.

[0744] Example: Integration of an oxidative dehydrogenation unit with an oil sands/bitumen upgrader [0745] In this example, the feed stream to an oxidative dehydrogenation unit is an ethane/ethylene mixed stream from an oil sands / bitumen upgrader, which may contain but is not limited to the mixed C ₂ fraction from fluid coking, delayed coking and hydrocracking operations. The ethylene content in this mixture can be from 8 volume % to 80 volume %. In this case, the ethane, which is present in the feed, is dehydrogenated at a conversion of at least 60% more preferably 80%, more preferably 99.5%.

[0746] Example: Operation of a membrane oxidative dehydrogenation reactor with bundled membrane tithes (Figure 6)

[0747] In the present example, the membrane reactor consists of a bundle of membrane tubes, wherein the catalyst is loaded inside the tubes as shown in Figure 6. This reactor design reduces the potential of membrane tube damage due to the different thermal expansion coefficients between the reactor vessel wall, catalyst and the membrane material. The internal reactor wall may be coated with ceramic to withstand the high temperature during any potential runaway reaction.

[0748] Three feed conditions are possible: 1. Ethane is preheated, oxygen is not preheated; 2. Oxygen is preheated, ethane is not preheated; 3. Both ethane and oxygen are preheated.

[0749] Example: Operation of a membrane oxidative dehydrogenation reactor with etsclosed bundled membrane tubes (Figure 7)

[0750] The reactor design considered its the present example is the same as the previous exasnple, except that oxygen is supplied to the membranes through individual and separate tubes; each membrane lias its own oxygen tube. This option could reduce the potential for multiple membrane tube damage if one tube is ruptured. The reactor has to be designed in a way that if one membrane is broken and no reaction is occurring, the 0x5' gen level is safely diluted by the reaction products from the other membrane tubes. In addition, an analyzer downstream of the reactor could detect an increase in the product oxygen content and shut down the reactor immediately for inspection of potential membrane damage.

[9751] Example: Operation of a multiple bed oxidative dehydrogenation reactor (Figures 31 and .32) [0752] In the multiple bed reactor approach shown in Figttre 31, oxygen and ethane are supplied to the first bed as either a pre-mixed fluid or the oxygen and ethane are mixed in the reactor inlet. The reactor can operate in either upward flow or downward flow modes of operation. The concentration of oxygen in the mixture supplied to each catalyst bed is such that the mixture is above its upper explosion limit (UEL). The maximum allowable amount of oxygen may be calculated based on the maximum allowable temperature in the reactor in the case of a runaway reaction. Oxygen is supplied at a temperature below the w eight-averaged bed temperature (WABT) and acts as a quench gas.

[9753] Oxygen and hydrocarbons can be mixed together without ignition according to poor art outlined in US20100191005A1. The gas stream 3201 can be filtered to reduce the presence of particles, which can be potential ignition sources. Figure 32 is a schematic diagram of a method of mixing oxygen 3202 and hy drocarbons, where oxygen is supplied inside of a membrane or distributor screen. This screen can be coated with an oxidative dehydrogenation catalyst on the hydrocarbon site, whereby, the membrane is impermeable to gas on the upper surface 3200 Oxygen can permeate the membrane and mix with hydrocarbons on the membrane surface 3203.

[0754] Example: Ethane oxidative dehydrogenation on a lab scale in two catalyst beds operation

[0755] In the present example, two lab-scale catalyst beds in series have been used to demonstrate oxidative dehydrogenation of ethane. Both catalyst beds are 0.2 cm’ in volume. The first catalyst bed is charged with 281 mg of catalyst and the inlet gas volume flow' rate is 900 cnrVhr and is composed of 77.3 mol % ethane and 22.3 mol % oxygen. The second catalyst bed is charged with 290 mg of catalyst; the inlet gas for the second catalyst bed consists of the whole product from the first catal st bed and 300 cniVhr of oxygen. The second catalyst bed inlet and outlet component mass flow rates are summarized in Table D6.

[0756] TABLE D6: Example ethane oxidative dehydrogenation on a lab scale in two catalyst bed operation, second catalyst bed results

[0757] Techniques are provided for an improved method for making a catalyst for the oxidative deh drogenation of lower alkanes to lower alkenes. Multicomponent metal oxide catalysts for the oxidative dehydrogenation of alkanes are known. Such catalysts are typically made by mixing solutions of metals and then precipitating the metal oxide “mixture” from the solution and calcining it. As a result, the catalysts are heterogeneous mixtures of various metal oxides and phases and may include some highly active species but also some species, which have a significantly lower activity. Applicants have found that by treating the precipitated metal oxides with a controlled amount of hydrogen peroxide prior to calcining the activity of the catalyst is improved.

[0758] In the specification, the temperature at which there is an approximate 25% conversion of ethane to ethylene is determined by plotting a graph of conversion to ethylene against temperature typically with data points below and above 25% conversion. Then a plot of the data is prepared or tire data is fit to an equation and tire temperature at which there is a 25% conversion of ethane to ethylene is determined. In some instances in die examples, the data had to be extrapolated to determine die temperature at which 25% conversion occurred.

[0759] In the specification, the phase selectivity at 25% conversion is determined by plotting the selectivity as function of temperature. The data is then plotted on a graph of selectivity against temperature or fit to an equation. Then having calculated the temperature at which 25% conversion occurs one can determine either from the graph or from the equation the selectivity at that temperature. Any other catalysts described herein may be used instead of, or in addition to, the catalyst described below. Accordingly, the conversion value may be selected to be a different number, such as 35 %.

[0760] In some embodiments, the calcined catalysts have the formula: Mo _1.0 V _0.22-033 Te _0.10-0.i6 Nb _0.15-0.19 O _d as determined by PIXE, where d is a number to satisfy the valence of the oxide. In some embodiments, the molar ratio of Mo:V in the calcined cataly st is from 1: 0.22 to 1:0.33. In some embodiments the molar ratio of Mo:V in the calcined catalyst is from 1: 0.22 to 1:0.29. In some embodiments, the molar ratio of Mo:V in the calcined catalyst is from 1 :022 to 1 :0.25. In some embodiments, the molar ratio of Mo:Te in the calcined catalyst is greater than 1 :0.10 and less than 1:0.16. In some embodiments, the molar ratio of Mo:Te in the calcined catalyst is from 1 :0.11 to 1: 0.15

[07611 The catalyst precursor may be prepared by mixing solutions or slurries (suspensions) of oxides or salts of the metallic components.

[0762] In some embodiments, the precursor may be prepared by a process comprising the following steps: i) forming an aqueous solution of ammonium heptamolybdate (tetrahydrate) and telluric acid at a temperature from 30 °C to 85 °C and adjusting the pH of the solution to 6.5 to 8.5, for example from 7 to 8, or for example from 7.3 to 7.7 (in one embodiment preferably with a nitrogen-containing base to form soluble sal ts of the metals); is) preparing a aqueous solution of vanadyl sulpliate at a temperature from room temperature to 80 °C (for example 50 °C to 70 ^C C, or for example 55 °C to 65 °C); iii) mixing the solutions from steps i) and ii) together; iv) slowly (dropwise) adding a solution of niobium monoxide oxidate (NbCXCiQrHti) to the solution of step iii) to form a slurry; and v) heating the resulting slurry in an autoclave under an inert atmosphere at a temperature from 150 °C to 190 °C for not less titan 10 hours. [0763] In some embodiments, the slurry from step v) is filtered, washed with deionized water, and dried for a time from 4 to 10 hours at a temperature from 70 to 100 °C.

[0764] In some embodiments, following step i) one or more of the following steps may be incorporated in the process: a) evaporating the aqueous solvent to obtain a solid; b) drying the solid at a temperature from 80 °C to 100 °C; and c) redissolving the solid in water at a temperature from 40 °C to 80 °C (for example 50 °C to 70 °C, or for example 55 °C to 65 °C).

[0765] in some embodiments, following step ii) the solutions are cooled to a temperature from 20 °C to 30 °C.

[0766] In some embodiments, as a part of step vi) the solution is cooled to a temperature from 20 °C to 30 °C.

[0767] In some embodiments, the precursor may be made by a process comprising: i) forming an aqueous solution of ammonium heptaniolybdate (tetrahydrate) and telluric acid at a temperature from 30 °C to 85 ^C C and adjusting the pH of the solution to 7.3 to 7.7 (for example 7.4 to 7.5) with a nitrogen-containing base to form soluble salts of the metals; ii) evaporating the aqueous solvent to obtain a solid; iii) drying the solid at a temperature from 80 °C to 100 °C; iv) redissolving the solid in water at a temperature from 40 °C to 80 ^C C (for example 50 ^C C to 70 °C, or for example 55 °C to 65 °C); v) preparing a aqueous solution of vanadyl sulphate at a temperature from room temperature to 80 °C (for example 50 °C to 70 °C, or for example 55 °C to 65 °C); vi) cooling the solutions from steps iv) and v) to a temperature from 20 to 30 °C; vii) mixing the cooled solutions from step vi together; viii) slowly (dropwise) adding a solution of niobium monoxide oxalate (NbC CbOtiTb) to the solution of step vii) to form a (brown) slurry; ix) heating the resulting slurry in an autoclave under an atmosphere free of oxygen at a temperature from 150 °C to 190 °C for not less than 10 hours; x) cooling the autoclave to room temperature and filtering and washing with deionized water the resulting solid; and si) drying the washed solid for a time from 4 to 10 hours at a temperature from 70 to 100 °C.

[0768] In some embodiments, the (catalyst) reactor may be lined with a coating selected from stainless steel, silica, alumina coating and polytetrafiuoroefhylene, preferably poiyietrafluoroethyiene (TEFLON) seeded with catalyst having a 25% conversion to ethylene at 420 °C or less and a selectivity to ethylene of not less than 90%. [0769] The seed catalyst may be a catalyst having the empirical formula (measured by PIXE): Mo1 0V0 22- o:;3Teojo-oj6Xbo 15-0 isOa, where d is a number to satisfy the valence of the oxide and having not less than 75 wt % of a crystalline component of the formula (TeO)o ₃₉ (Mo _3.52 Vi _. o ₆ Nbo _.42 )0] ₄ as determined by XRD.

[0770] In some embodiments, the (catalyst precursor) reactor may be lined with a coating of a fully fluorinated ethylene propylene polymer (FEP) seeded with a catal st having a 25% conversion to ethylene at 420 °C or less and a selectivity to ethylene of not less than 90%.

[0771] In some embodiments, the seed catalyst has the empirical formula (measured by PIXE) Mo1 _. 0V0 _. 22- ossTeo _.i o-o _.i eNbo is-o isO _d w'here d is a number to satisfy' the valence of the oxide and having not less than 75 wt. % of a crystalline component of the formula (TeO)o _. 39(Mo3 _. 52Vi _.0 6Nbo _. 42)Oi4 as determined by XRD.

[0772] The seed catal st loadings may range from 1 to 15 wt. % of the surface of tire reactor (e.g. steel. Teflon, or FEP). [0773] In some instances, the (catalyst precursor) reactor contains particulates of stainless steel, silica, alumina, and polytetrafluoroethylene seeded with a catalyst having a 25% conversion to ethylene at 420 °C or less and a selectivity to ethylene of not less than 90%.

[0774] in some embodiments, the seed catalyst has the empirical formula (measured by PIXE) Mo1.0V0.22- 033Te0.10-0.i6Nb0.15-0.19Od where d is a number to satisfy the valence of the oxide and having not less than 75 wt. % of a crystalline component of the formula (TeO)o _. 39(Mo3 52Vi o6 bo42)Oi4 as determined by XRD.

[0775] The particulates may be (irregular such as flakes, granules, globules, filaments etc. or regular such as spheres, elliptical, rods (stirring bare), rectangular prisms (both right and non-right), pentagonal prisms, pyramids etc.)

[0776] The seed catalyst loadings on die particulates may range from 1 to 15 wt. % of the particulates.

[0777] In some circumstances, it may be easier to replace particulates on which die seed catalyst lias been depleted with new' seed particles having an appropriate loading of seed particles than to replenish die seed coating on the interior surface of the catalyst reactor.

[0778] In some embodiments, the catalyst produced from a hydrothermal reactor seeded with catalyst having a 25% conversion to ethylene at 420 °C or less and a selectivity to ethylene of not less than 90% has the empirical formula as determined by PIXE, M01V0.34 -0.39Te0.09-0.14Nb0 _. 14-0 _.i 6O _d where d is a number to satisfy the valence of the oxide.

[0779] The peroxide treatment may take place at atmospheric pressure and room temperature (e.g. from 15 °C to 30 °C) to about 80 °C, in some instances from 35 °C to 75 °C in other instances from 40 °C to 65 °C. The peroxide has a concentration from 10 to 30 w't. %, in some instances form 15 to 25 wt. %. The treatment time may range from 1 to 10 hours, in some cases from 2 to 8 hours, in other cases from 4 to 6 hours.

[0780] The catalyst precursor is treated with the equivalent of from 0.3 - 2.8, in some embodiments from 0.3- 2.5 mL of a 30 wt. % solution of aqueous ¾(¾ per gram of precursor. The treatment should be in a slurry (e.g. the precursor is at least partially suspended) to provide an even distribution of ί i ·() _. · and to control the temperature rise. For post calcination treatment with H2O2 there is a sudden delayed violent reaction with H2O2. The process is an instantaneous reaction, which is more controlled and safer

[07811 The treated catalyst precursor is then subject to calcining to produce the active oxidative dehydrogenation catalyst. The treated precursor may be calcined in an inert atmosphere at a temperature from 200 °C to 600 °C for a time from 1 to 20 hours. The purge gases used for calcining are inert gases, including one or more of nitrogen, helium, argon, CO2 (preferably high purity > 90%), tire gases or mixture containing less than 1 vol.~% hydrogen or air, at 200-600 °C, or for example at 300-500 °C. The calcining step may take from 1 to 20, in some instances from 5 to 15 in other instances from about 8 to 12 hours, for example about 10 hours. The resulting mixed oxide catalyst is a friable solid, in some embodiments the solid is insoluble in water. In some embodiments, the calcined product has a bulk density from 1.20 to 1.53 g/ec. This bulk density is based on how much 1.5 rnL of pressed and crushed catal st weighs.

[0782] The resulting oxidative dehydrogenation catalyst is heterogeneous. It has an amorphous component and a crystalline component. The elemental analysis of the catalyst may be determined by any suitable technique. One useful technique is Particle Induced X-Ray Emission analysis (PIXE) From a PIXE analysis of the catalyst precursor prior to treatment and after treatment with H ₂ 0 ₂ it is determined that the empirical molar ratio of Mo to V decreases, for example, from 1:0.33 to 1:0.40 to from 1: 0.22 to 1:0.33, in some instances from 1.0:0.22 to 1.0:0.25 compared to a calcined material which has not been treated with hydrogen peroxide. Further it is found that, in some embodiments, the molar ratio of Mo:Te is tightened and increased (over the base catalyst) from a range from 1: 0.03 to 1: 0.13 to greater than 1:0.10 and less than 1:0.16, in some instances from 1.0:0.11 to 1:0 to 0.15 compared to a calcined oxidative dehydrogenation catal st which has not been so treated.

[0783 g The catalyst has one or more crystalline components and an amorphous component. The crystalline component may be analyzed using X-Ray diffraction (XRD). There are a number of suppliers of X-Ray diffractometers including Rigaku Shimadzu. Olympus and Bruker. A powder sample is irradiated with X-Rays. The X-Rays given off from the sample pass through a diffraction grid and are collected in a goniometer (recorder). The results may be analyzed using a computer program (for example one provided by tire instrument supplier) and compared to a database (International Center for Diffraction Data ICDD) using a computer to determine the composition of the crystalline phase(s).

[0784] The 20 X-Ray diffraction pattern has a ratio of peak height at 20 from 0 to 20° to maximum peak height of less than 15%, in some instances less than 10%.

[Q785] The crystalline phase of the catalyst is also heterogeneous. The X-Ray diffraction results may be analy zed by computer programs to identify various likely crystalline species and their relative amounts compared to the structures in a database (e.g , deconvoluted).

[0786] In some embodiments, the crystalline phase includes the following crystalline species:

(Moo _. eNbo.rrVo.is/sOn, T eOo i (M00.73 V c Nbo.or/rO?, (TeO)o.39(Mo3.52Vi.o6Nbo.42)Oi4, V _1.1 M00.9O5; M04V6O25, and VOM0O4.

[0787] X-Ray diffraction analysis of the precursor and the calcined catalyst shows treatment results in a change in the composition of the crystalline phase. The treatment can increase the phase of the crystalline component having the empirical formula (TeO)o39(Mos 5 ₂ V i _. c _t eXbo _/ o) to not less than 75 wt %, in some instances not less than 85 wt. %, in some instances not less than 90 wt. %, in some instances not less than 95 wt. % of the of the cr stalline phase.

[0788] In some embodiments, the phase of the crystalline component having the empirical formula TeOo _. 7i(Moo _. 73Vo.2Nbo.o7)309 is present in an amount of from about 2.4 to 12 wt %, in some embodiments the phase is present in amounts less than about 8 wt. %, in further embodiments less than 3.5 wt. %.

[0789] The calcined catal st product is a dry friable product, which in some embodiments is insoluble in water. If required the catal st may be subject to a sizing step, such as grinding, to produce a desired particle size. Depending on how tire catalyst is to be used the particle size may be different. For example for spray drying with a support Site particle size may range from about 5 to 75 pm, in some cases from 10 to 60 pm. For use in a bed in unsupported form, the particles may lave a size from about 0.1 to 0.5 mm in some instances from 0.2 to 0.4 mm. [0790] The feed to the oxidative deh drogenation reactor includes oxygen in an amount below the upper explosive/flammability limit. For example for ethane oxidative deh drogenation, the oxygen will be present in an amount of, for example, not less than about 16 mole %,or for example, about 18 mole %, or for example from about 22 to 27 mole %,or 23 to 26 mole %. It is desirable not to have too great an excess of oxygen as this may reduce selectivity arising from combustion of feed or final products. Additionally too high an excess of oxygen in the feed stream may require additional separation steps at the downstream end of the reaction.

[0791 J To maintain a viable fluidized or moving bed, the mass gas flow rate through the bed must be above the minimum flow required for fluidization, and, for example, from about 1.5 to about 10 times U _mf , for example , from about 2 to about 6 times U _nif . U _m r is used in the accepted form as the abbreviation for the minimum mass gas flow' required to achieve fluidization, C. Y. Wen and Y. H. Yu, "Mechanics of Fluidization", Chemical Engineering Progress Symposium Series, Vol. 62, p. 100-111 (1966). In some instances, the superficial gas velocity required ranges from 0.3 to 5 m/s.

[0792] The reactor may also be a fixed bed reactor.

[0793] The oxidative dehydrogenation process comprises passing a mixed feed of ethane and ox gen at a temperature less than 420 °C in some instances less than 410 °C, in some instances less than 400 °C, in some instances less than 390 °C, in some instances less than 380 °C, in some instances as low' as 375 °C at a gas hourly space velocity of not less than 500 hr ¹ , for example, not less than 1000 hr ¹ , or for example, not less than 2800 hr ¹ , or for example, at least 3000 hr ¹ through one or more beds and a pressure from 0.8 to 1.2 atmospheres comprising passing the mixture over the oxidative deh drogenation catalyst. In some embodiments, the oxidative deh drogenation reactor operates at temperatures below' 400 °C, or for example from 375 °C to 400 °C.

[0794] The outlet pressure from the reactor may be from 105 kPag (15 psig) to 172.3 kPag (25 psig) and the inlet pressure is higher by the pressure drop across the bed which depends on a number of factors including reactor configuration, particle size in the bed and the space velocity. The pressure drop may be below' 689 kPa (100 psi), for example, less than 206.7 kPa (30 psi).

[0795] The residence time of one or more alkanes in the oxidative dehydrogenation reactor is from 0.002 to 20 seconds.

[0796] The Oxidative Dehydrogenation Processes

[0797] The catalyst described here may be used with a fluidized bed or a fixed bed exothermic reaction. The fixed bed reactor is a tubular reactor and in further embodiment, the fixed bed reactor comprises multiple tubes inside a shell (e.g. a shell and tube heat exchanger type construction). In a further embodiment, the fixed bed reactor may comprise a number of shells in series and/or parallel. The reactions may involve one or more dehydrogenation steps including oxidative dehydrogenation, and hydrogen transfer steps including oxidative coupling of a hydrocarbon.

[0798] In some embodiments, these reactions are conducted at temperatures from about 375 °C up to about 410 °C, at pressures from about 100 to 21,000 kPag (15 to 3000 psig), at, for example, an outlet pressure from rom 105 kPag (15 psig) to 172.3 kPag (25 psig), in the presence of an oxidative dehydrogenation catalyst. The hydrocarbon stream may contain a range of compounds including C2-C4 aliphatic hydrocarbons.

[0799] In some embodiments, the reactions include die oxidative coupling of aliphatic hydrocarbons, for example C ₁ -C ₄ aliphatic hydrocarbons particularly methane (e.g. when the ethane stream contains sortie methane) and the oxidative dehydrogenation of C2-C4 aliphatic hydrocarbons. Such reactions may be conducted using a mixed feed of hydrocarbons, in some embodiments methane or ethane or both and 0x5' gen in a volume ratio from 70:30 to 95:5 at a temperature less than 420 °C, for example less than 400 °C at a gas hourly space velocity of not less than 280 hr ¹ , in some embodiments not less than 500 hr ¹ , or not less than 1000 hr ¹ , or for example not less titan 2800 hr ^] , or for example at least 3000 hr ¹ , and a pressure from 0.8 to 1.2 atmospheres. In some embodiments, the process may have an overall conversion of from about 50 to about a 100%, or from about 75 to 98% and a selectivity to ethylene of not less titan 90%, in some instances not less than 95%, in further embodiments not less titan 98%. In some cases the temperature upper control limit is less titan about 400 °C. in some embodiments less titan 385 °C. [08001 The resulting product stream is treated to separate etliyiene from the rest of the product stream, which may also contain co-products such as acetic acid, and un-reaeted feed, which is recycled back to the reactor.

[0801] Separation

[0802] The product stream front the reactor should have a relatively low content of ethane less, than 20 wt. %, in sortie cases less titan 15 wt. % in some cases less than 10 wt. %. Additionally, the product stream should have a low content of by products such as water, carbon dioxide, and carbon monoxide, for example cumulatively in a range of less than 5 wt. %, or for example less than 3 wt. %.

[0803] The feed and by products may need to be separated from the product stream. Some processes may use so-called dilute ethylene streams. For example if the product stream does not contain too much ethane, for example less than about 15 vol. % the stream may be used directly without further purification in a polymerization reactor such as a gas phase, sluriy or solution process.

[0804] The most common technique would be to use a cryogenic C _? splitter.

[0805] Other known ethylene/ethane separation techniques could also be used including adsorption (oil, ionic liquids and zeolite).

[0806] Figure 33 is a schematic diagram of the fixed bed reactor unit 3300 used for the oxidative dehydrogenation reaction. The reactor 3300 was a fixed bed stainless steel tube reactor having a 2 mm (¾”) outer diameter and a length of 117 cm (46 inches). The reactor 3300 is in an electrical furnace sealed with ceramic insulating material. There are seven thermocouples in the reactor 3300 indicated at numbers 1 through 7 in Figure 33. Thermocouples are used to monitor the temperature in that zone of the reactor. Thermocouples #3 and #4 are also used to control the heating of the reactor bed. The feed flow's from the top to the bottom of the reactor. At the inlet, there is a ceramic cup 3308 to prevent air drafts in the reactor. Below the ceramic cup 3308 is a layer of quartz wool 3309. Below the layer of quartz wool 3309 is a layer of caialyfically inert quartz powder. Below the quartz powder is the fixed bed 3310 comprising catalyst. Below the fixed bed 3310 is a layer of quartz powder 3311, a layer of quartz wool 3312, and a ceramic cup 3313. At the exit of the bed was a gas analyzer to determine the composition of the product stream. The GHSV was 2685 hr ¹ and the pressure was ambient.

[0807] For the examples, tire bed temperature was taken as an average of the temperatures from thermocouples #2, #3 and #4. The feed stream was assumed to have the same temperature as the bed. [0808] EXAMPLES

[0809] Baseline Experiments

[0810] Two different base line catalysts were prepared in geographically separated laboratories.

[0811] Laboratory 1 used a TEFLON ^® lined reactor for the hydrothermal treatment, which had on its surface crystals of effective catalyst prepared previously.

[0812] The formation of the pre catalyst in glassware procedure was as follows:

[0813] (NHrieMoeTeGir.xiLO (6.4069 g) was added to 20 mL of distilled water in a 100 mL glass beaker and stirred on a warm water bath (80 °C). VOSO4 x ¾0 (3.6505 g) was dissolved in 10 mL of distilled water in a 50 mL beaker at room temperature. The VOSO4 solution was poured into the (NILLMoeTeC^ solution mid a brown solution resulted immediately (Solution El).

[9814] HsPSIbOi C Gib] 7.5 H ₂ 0 (2.3318 g) was dissolved in 10 mL of warm water and added under air atmosphere to the Solution El. A dense dark brown-gray colored slurry formed, which was stirred for 10 minutes under air atmosphere.

[9815] Hydrothermal Treatment

[9816] Samples of the slurry' were then heated in an autoclave having a TEFLON liner seeded with prior produced catalyst under an inert atmosphere at a temperature from 150 °C to 190 °C for not less than 10 hours. The slurry was filtered, washed with deionized water, and dried for a time from 4 to 10 hours at a temperature from 70 °C to 100 °C.

[9817] The base line catalyst subsamples wore immediately calcined.

[9818] Two samples were prepared in this manner.

[9819] One sample (Sample EA) had a 25% conversion of ethane to ethylene at about 370 °C and a selectivity' at this temperature of 98%.

[9820] The second sample (Sample EB) had a 25% conversion of ethane to ethylene at about 354°C and a selectivity' at this temperature of 99%.

[0821 ] This show's even with catalyst seeds in the hydrothermal treatment there is variability'. (This may be due to differences in the seed crystals.)

[9822] Three subsamples of precatalyst EA and four subsamples of precatalyst EB were treated with various amounts of ¾0 ₂ then calcined and then used to oxidatively dehydrogenate a mixture of 78% ethane and 22% oxygen at a flow rate of 600 cmVh.

[0823] The results are set forth in Table El below. [0824] TABLE El

[0825] Treatment of a catalyst having 25% conversion at temperatures below 420 °C and high selectivity with

30% hydrogen peroxide in amounts from 0 3 to 2.8 mL per gram of catalyst does not provoke a measurable performance loss

[0826] Second Laboratory

[0827] Precatalyst Preparation

[0828] (NH-dsMosTeOar. HaO (6 4 g) was dissolved in 20 ml. of water in a 100 ml. round-bottomed flask with the aid of a warm water bath. The clear solution was cooled to room temperature VOSO ₄ x3 47 H ₂ 0 (3.4 g) was dissolved in 10 mL of water in a 30 mL beaker (also with the aid of a warm water bath). The blue solution formed was cooled to room temperature. The VOSO ₄ solution was poured into the ( DeMoeTeCfo solution. The beaker was rinsed with water (2x0 5 mL), and the rinsing solution was added to the flask. A brown solution formed was bubbled with nitrogen and was stirred for 10 minutes. An aqueous solution of Hr^NbCXibCbb] (0.3431 mmol/g solution, 13.29 g, 4.56 rnrnol of Nb) was added slowly to the above brown solution with a pipette (in~2.5 minutes). A dull red stone colored shiny formed, which was stirred with bubbling of N ₂ for about 10 minutes.

[0829] Hydrothermal Treatment in an unseeded TEFLON lined reactor

[0830] The slurry was transferred to 60 mL autoclave having a clean Teflon liner, which was degassed and refilled with N ₂ (ambient pressure). The autoclave was heated with a heating sleeve with the content magnetically stirred (300 rpm). The mixture was heated at 175 °C internal temperature for 48 hours. The autoclave was cooled to room temperature and the content was filtered and washed with 500 mL of water. The cake was dried at 90 °C overnight, ground, and sieved through 2.50 micron sieve. The purple solid was calcined at 600 °C (0 level in nitrogen stream: 0.4 ppm). This catalyst appeared brown in color after calcining.

[0831] The catalyst was tested as above.

[0832] The ODH reaction was carried out at temperatures up to 420 °C to avoid auto ignition temperature of the feed gas. The conversion was low and the graph of conversion as a function of temperature had to be extrapolated linearly to get a rough estimate of the temperature at which there was 25% conversion. The estimated temperature at which there was 25% conversion was 635 °C. This would not be commercially viable as it is significantly above the auto ignition temperature of a feed gas comprising 82% ethane and 18% oxygen.

[0833] Seeded TEFLON lined reactor

[0834] At the second laboratory, the precatalyst was prepared using the following general procedure.

[0835] The procedure to prepare catalyst w'as as follows.

[0836] A slurry prepared as above, was poured into a 300 mL autoclave having a TEFLON liner. The reactor was dedicated and not washed between hydrothermal treatments and it had residual crystals of catalyst made during its prior use. The autoclave was closed. The headspace was purged of oxygen with N ₂ (20 psig). After purging, the valve was closed and the autoclave was put in an oven at 23°C. The temperature was raised to 175°C and held without stirring at tins temperature for 50 hours. The autoclave was taken out of the oven and cooled to room temperature. The pressure of the autoclave was released through a water bubbler. The autoclave was opened. The solid was filtered, rinsed with 500 mL of water and was dried at 80 °C overnight. The brown solid (6.19 g) was loaded in a quartz boat and was calcined under a slow stream (30 rnL/uiin) of purified nitrogen (RT to 60G°C, 4 hours, 600°C kept for 2 hours). The solid obtained was a black powder, which was ground and sieved through a 250 micron sieve (5.94 g). The resulting solid was loose (fluffy).

[Q837] The catalyst was tested in the ODH reactor using the above conditions

[Q838] From the experiments, the temperature at which there was 25% conversion to ethylene ranged from 370 °C to 383 °C and a selectivity at these temperatures was greater than 90%. This is a narrow range consideri ng the heterogeneous nature of the catalyst and the complexity of crystalline phases and consistently below the auto ignition temperature of the feed.

[0839] The nature of the nucleation sites was not clear. It is believed if the sites comprise cataly st having a 25% conversion below 400 °C and selectivity to ethylene at this temperature above 95% the resulting catalyst has a higher probability of the resulting catalyst having these properties

[0840] A number of samples of catalyst prepared as above in a seeded TEFLON lined hydrothermal reactor were subject to XRD analysis as described in the examples below to determine the crystalline phases in the catalyst. The results are presented in Table E2.

[0841] TABLE E2

[0842] This shows that even if the reactor wall (TEFLON liner or steel) in the hydrothermal reactor is seeded with catalyst there can be a significant variability in the final catalyst. [0843] The samples of calcined catalyst obtained from a seeded reactor have the following empirical formula as determined by PIXE:

[0844] MO _I VO.34 -0 39Te0.09-0.14Nb0 14-0 i6Od where d is a number to satisfy the valence of the oxide. The samples had a 25% conversion at a temperature from 372 °C to 383 °C and a selectivity to ethylene at these temperatures from 93 to 96%.

[0845] Second Laboratory

[0846] A series of catalysts were prepared in a clean glassware reactor and subject to hydrothermal treatment in a stainless steel reactor without a TEFLON liner and without any catalyst seeding.

[0847] General Reaction Step

[0848] (NFDeMoeTeOii.xITiO (19.2086 g, 15.15 mmol, 1.00 molar equivalents) was dissolved in 60 mL of distilled water in a 500 mL round-bottomed flask with the aid of a warm water bath. The resulting clear and colorless solution was allowed to cool to room temperature. VOSO4 x3.47 H ₂ 0 (10.2185 g, 62.59 mmol, 4.13 molar equivalents) was dissolved in 25 mL of distilled water in a 30 mL beaker with die aid of a warm water bath. The resulting clear blue solution formed was cooled to room temperature.

[0849] The VOSO4 solution was poured into the (NiDgMogTeC^ solution and a brown solution resulted immediately. The beaker which contained the VOS04 solution was rinsed with two i mL aliquots of water and these rinsings were added to the flask. The resulting b rown solution was stirred under addition of bubbling nitrogen for 15 minutes. Aqueous EbpSibiXCaOils] (0.3420 nmol-Nb gisotaton), 39.9737 g^imon}, 13.68 mmol(Nb), 0.903 molar equivalents) was added slowly (dropwise over seven minutes) under N ₂ bubbling to the brown solution via a pipette. A dull purple colored slurry formed, which was stirred with bubbling of N ₂ for 10 minutes.

[0850] General Hydrothermal Treatment Step

[0851] The slurry was poured to a 600 mL bare steel autoclave, which contained a TEFLON stir bar. The autoclave was closed and the atmosphere inside of the autoclave was evacuated (vacuum) and filled with N ₂ (30 psig from bulk nitrogen line) 10 times, followed by an additional 10 repeats of purging with N ₂ (30 psig from bulk nitrogen line) and releasing of N ₂ pressure (positive pressure relief) to a water bubbler. The autoclave was left under ambient pressure of N ₂ atmosphere and the vessel was sealed using a needle valve on the autoclave head.

[0852] The autoclave was put into a heating blanket setup, where the heat is controlled by heat controller via thermocouples inside and outside the autoclave. The heating blanket and autoclave were wrapped in thermal insulating ceramic fiber tape to ensure proper insulation. The temperature was raised to 173 ^® C over the period of an hour and the reaction was let to proceed, with the addition of Stirling, at this temperature for 48 hours.

[0853] The autoclave was then cooled to room temperature slowly without stirring. Once cooled, the excess pressure that built up during the process of the reaction inside the autoclave was release through a water bubbler and (he autoclave was opened. The solid (deep purple color) was filtered, rinsed with approximately 300 mL of distilled water (filtrate vibrant blue color) and was dried in an oven at 90 °C overnight.

[0854] General Calcination Step

[0855] The dried catalyst solids were ground using a mortar/pestle and sieved through a 250 micron porosity sifter. The less than 0.25 micron particle size dark purple solid was loaded in a quartz boat and the boat was placed into glass furnace tube, which is used for calcination. To ensure the exclusion of air during the calcination, the setup was purge under nitrogen. The calcination proceeded under a slow stream (30 inL/min) of purified nitrogen (vent through water bubbler) under the following conditions: RT to 600°C in 4 hours and held at 600°C for 2 hours. The solid obtained was a black powder, which was ground and sieved through a 250 micron sieve resulting in a powder that w'as loose and fluffy.

[0856] The catalysts were tested as above. The temperature at which 25% conversion (either measured or lineally extrapolated) ranged from 380 to 504°C, This was a broad spread in 25% conversion temperature as there was no obvious difference between the preparations. Of the five samples two had a 25% conversion temperature below 400 °C. which was felt to be a “reasonable” ceiling temperature for a large scale commercial ODH reactor. [0857] These examples further illustrate the variability in manufacturing catalysis having a 25% conversion below 400 °C absent catalyst seed having the desired properties (temperature at which there is 25% conversion less than 400 °C and a selectivity for ethylene at this temperature of greater titan 90%).

[0858] In the literature it is known (Catal sis Communications 2005, 6, 215-220; Catalysis Communications 2012, 21, 22-26; Journal of Catalysis 2012, 285, 48-60) to treat ODH catalyst post calcining with hydrogen peroxide to improve performance.

[0859] In the second lab a sampled of the catalyst prepared as above catalyst was calcined at 600°C for from 2 to 4 hours. The calcined sample was then treated with about 12 to 16 ml, of 30% w/w ¾(½ aqueous solution per gram of catalyst. The reaction was inconsistent in that there was no indication of reaction (e.g. no heat or bubbling) or the incubation period to start the reaction was extremely unpredictable (e.g. 20 minutes to 3 hours) and when it started the reaction was extremely fast (in seconds) and violent (potentially explosive).

[0860] The addition of H ₂ 0 ₂ post calcination of the catalyst is not a commercially viable route to cataly st preparation due to the complication described above and safety implications [0861] First Laboratory

[0862] In the first lab portions of the baseline catalyst precursor prepared with a seeded TEFLON liner were treated with up to 5.6 ml of 30% w/w H ₂ 0 ₂ aqueous solution per gram of precursor prior to calcining. The treatment of the precursor resulted in an immediate, controlled and observable reaction (bubbling and mild heating, which never exceeded about 60°C). The treated precursor was then calcined in the normal manner.

[0863] The catalysts were then tested in the ODH reactor.

[0864] Treatment causes a minor variation of the selectivity' (between 99% and 98%) up to the peroxide amount 3.5 cc, and only the use of a larger H ₂ G ₂ excess (5.6 cc) provokes a measurable selectivity loss.

[0865] Figure 34 is a plot of the temperature at which there is a 25% conversion of ethane to ethylene against the volume of 30%H ₂ O ₂ for 1.41 g of a catalyst having a temperature at which there is 25% conversion of less than 240 °C and a selectivity to ethylene of greater titan 95% prepared at foe first laborator .

[0866] Figure 35 is a plot of foe selectivity for conversion to ethylene at the temperature at which there is a 25% conversion to ethylene against the volume of 30% H ₂ 0 ₂ for 1.41 g of catalyst having a temperature at which there is 25% conversion of less titan 240 °C and a selectivity to ethylene of greater than 95% prepared at the first laboratory. [0867] These plots show the volumes of 30% ¾(¾ per 1.4 g of catalyst at which catalyst having a temperature at which there is 25% conversion of less than 240 °C and a selectivity to ethylene of greater than 95% is relatively uncompromised up to about 5.6 mL of 30% H2G2 per 1.4 g of catalyst (i.e. 0.30 - 2.8 mL ¾(¾ of a 30% solution per gram of catalyst).

[0868] In the Second Laboratory

[0869] Portions of the baseline catalyst precursor prepared as above and treated in a stainless reactor without a TEFLON liner and without seeding were treated with up to from 0.35 to 1.42 mL of 30% w/w H2G2 aqueous solution per gram of precursor prior to calcining. The treatment of the precursor resulted in an immediate, controlled and observable reaction (bubbling and mild heating, which never exceeded about 60°C).

[0870] A series of PIXE characterizations of the base line catalyst and catalyst treated in accordance with the provided techniques were obtained from laboratory two.

[0871] Typical base line untreated catal st had a PIXE characterization set forth below:

(MOl.00V0.37-0.39 Ten.03-0. OsNbo.14-0. i5pe0.003 -0,006)07.89-9.07

[0872] In base line catalysts treated in an unlined hydrothermal reactor, small amounts of iron and chromium were detected. The iron ranged from a minimum of 0.0026 to a maximum of 0.0416 moles / er mole of Mo. The chromium ranges from 0.000 to 0.0065 moles per mole of Mo.

[Q873] For catalyst treated in accordance with the provided techniques prior to calcining, the PIXIE analysis was (Moi.ooVo.2s-o.29Teo.i3Nbo.i5-o.i6Feo.oo8)08.n.

[0874] It is believed that these amounts of iron and chromium in catalyst of the above noted base structure do not contribute to the oxidative dehydrogenation characteristics of the catalyst.

[0875] Hydrogen peroxide treatment of a catalyst having a temperature at which there is 25% conversion of less than 420 °C and a selectivity to ethylene of greater than 95% prepared in laboratory 2.

[0876] Sample E1A

[0877] A catalyst precursor was prepared in the above manner and treated in a stainless steel hydrothermal reactor without a TEFLON liner and without seeding with a catalyst having a temperature at which there is 25% conversion of less than 240 °C and a selectivity to ethylene of greater than 95% was treated with hydrogen peroxide. [0878] 5.4672 g of the crude purple catalyst precursor was used for hydrogen peroxide treatment. The catalyst precursor was added to a 400 mL beaker, containing a stir bar, and 20 mL of distilled water was added to create a dark slum . The slurry ^? was agitated through stirring and 4 mL H2O2 (30% w/w in H ₂ 0; ratio of 1.41 go» a : 1 mLu202) was added all at once and vigorous bubbling and heat resulted. The reaction self-heated and bubbled and changed from dark purple slurry to a black slurry. The reaction was stirred and allowed to proceed for 5 minutes before w ^? ork up. The solid was filtered, rinsed with approximately 100 mL of water and was dried in an oven at 90 °C overnight to produce 4.4296 g of grey precursor for calcination step. The sample was calcined as above.

[0879] Sample E1C

[0880] 5.5679 g of the crude purple catalyst precursor was treated in the same maimer as Example 1 A except that the reaction was allowed to proceed for 2 hours before work up. Some minor bubbling was observed to be arising from the reaction even after a 2 hour reaction time. The solid w ^? as filtered (filtrate color was yellow), rinsed with approximately 100 nxL of water and was dried in an oven at 90 °C overnight to produce 4 6231 g of vibrant purple precursor for calcination step. The resulting sample was calcined as above.

[0881] Sample E1B

[0882] A sample of the precursor prepared in a glass flask as above was not treated and calcined as above. [0883] The samples were then used in the oxidative dehydrogenation of ethylene.

[0884] The results of the oxidative dehydrogena tion test are set out in Table E3.

[08851 TABLE E3

[0886] Treatment of a precursor for a catalyst having a temperature at which there is 25% conversion of less than 420 °C and a selectivity to ethylene of greater than 95% with i nxL of 30% ¾(¾ per 1.4 g of catalyst precursor does not adversely affect the catalyst.

[0887] The samples were then subject to XRD analysis using a Rigaku Ultima X-Ray Diffractometer, 285 mm radius theta/theta goniometer; D/teX-ULTRA High Speed Detector; and ASC-48 Automatic Sample changer. The software used was Data Acquisition Rigaku “Standard Measurement” application; Analysis software MDI Jade 2010 version 2,6.62014; and the comparative database was ICDD PDF-4+ 2014 (with 354,264 inorganic data patterns).

[0888] TABLE E4

[0889] Table E4 suggests that it is desirable to increase the content of the (TeO)o (Mo «V i Nb)o )O _M phase and reduce the content of the (TeO)o. (Moo.? _?, Vo.zNbjo.sn^O _? phase.

[0890] A further sample of catalyst having a temperature at which there is 25% conversion of less than 420 °C and a selectivity to ethylene of greater than 95% was tested. [0891] Example E2 A

[0892] 7.0552 g of the erode purple catalyst precursor was treated in the same manner as Example El A except that the reaction was allowed to proceed for 20 minutes before work up. The solid was filtered, rinsed with approximately 100 mL of water and was dried in an oven at 90 °C overnight to produce 5.8907 g of black precursor for calcination step.

[0893] Example E2B

[0894] The baseline catalyst was not treated.

[0895] The results of the oxidativ e dehydrogenation test me set out in Table E5.

[0896] TABLE E5

[0897] Treatment of a precursor for a catalyst having a temperature at which there is 25% conversion of less than 420 °C and a selectivity to ethylene of greater than 95% with 1 mL of 30% H Chper 1.4 g of catalyst precursor does not adversely affect the catalyst.

[0898] The samples w'ere then subject to XRD analysis as above.

[0899] TABLE E6

[0900] The treatment with H2O2 increases the relative proportion of the phase having the structure (TeO)o.39(Mo3.5 ₂ Vi.o6 ₂ Nb)o. ₄₂ )Oi ₄ and improves die performance of the catalyst.

[0901] Examples of treating a catalyst, which does not have a temperature at which there is 25% conversion of less than 420 °C and a selectivity to ethylene of greater than 95% with ¾(¾ [0902] Example E3B

[0903] A sample of tire catalyst precursor which was calcined without treatment with H ₂ 0 ₂ was tested in tire oxidative dehydrogenation reactor. This was the above catalyst that had an estimated temperature for 25% conversion of 504 °C.

[0904] Example E3A

[0905] 5.9354 g of the erode purple catalyst precursor for the untreated sample was treated with hydrogen peroxide. The catalyst precursor was added to a 250 mL round bottom flask, containing a stir bar, and 20 mL of distilled water was added to create a dark slurry. The slurry' was agitated through stirring and 8.5 mL ¾<¼ (30% w/w in 3¾0; ratio of 0.705 g ₀ o _H : 1 mL _H 202) was added all at once and vigorous bubbling and heat resulted. The reaction seif-heated and bubbled and changed from dark purple slum' to a black slurr . The reaction wns stirred for 2 hours before work up. The dark purple solid was filtered, rinsed with approximately 100 mL of water and was dried in an oven at 90 °C overnight to produce 3.7494 g of grey solid for calcination step.

[0906] Example E3C

[0907] 4.9755 g of the crude purple catalyst precursor was treated as E3A above except 1.75 mL H ₂ 0 ₂ (30% w/w in Rio; ratio 2.82 goon : 1 mL _H2 02) was added all at once and less bubbling and heat resulted. The reaction slurry remained dark purple. The reaction was stirred for 2 hours before work up. The dark purple solid was filtered, rinsed with approximately 100 mL of water and was dried in an oven at 90 °C overnight to produce 3.8326 g of grey solid for calcination step.

[0908] The samples were then tested in the oxidative dehydrogenation reactor. The results are shown in Table E7.

[0909] TABLE E7

[0910] Treatment of a precursor for a catalyst having a temperature at which there is 25% conversion of greater than 420 °C and a selectivity to ethylene of less than 95% with 1 mL of 30% ¾(¼ per 07 to 28 g of catalyst precursor significantly improves the catalyst.

[0911] The samples were then subject to XRD analysis as above.

[0912] TABLE E8

[0913] The da ta suggests that increasing the content of the phase (TeO)o39CM03 52 V i.o ₆₂ Nb)o. ₄₂ )Oi ₄ significantly increases the activity and selectivity of the catalyst.

[0914] Example E4 /Treatment of the mother liquor with H2G2 without filtration

[0915] A sample of precursor was prepared as above. A portion was used as a base line reference (without treatment with H2O2). Then 4.96 g of the crude purple catalyst precursor and aqueous mother liquor (-500 mL) from hydrothermal treatment was added to a 250 mL round bottom flask, containing a stir bar to create a dark slurry. The dark slurry was kept under nitrogen atmosphere. The sluny was agitated through stirring and 3.6 mL ¾<¾ (30% w/w in H2O ; ratio of 1.39 go _DH : 1 mL _H 202) was added all at once and no apparent vigorous bubbling and heat resulted. The reaction changed from dark purple slurry to a black slurry. The reaction was stirred for 3 hours before work up. The dark purple solid was filtered rinsed with approximately 200 rrrL of water and was dried in an oven at 90 °C overnight to produce 3.3720 g of grey solid for calcination step.

[0916] The catalysts were tested for activity in the oxidative dehydrogenation reactor. The results are set forth in Table E9.

[Q917] TABLE E9

[0918] Treatment of a precursor with 1 mLmo2 (30 wt %) per 1.4 g of precursor prior to separation from the reactor prior to drying improves the activity of the calcined dehydrogenation catalyst.

[0919] Example E5: 100 g Sample

[0920] A number of samples of catalysts (approximately 40, 40, and 20 g) were combined into a 5 L round botom flask and 400 ml. of distilled water was added to create a purple slurry. A 100 mL dropper funnel was attached to the flask and 39 ml. of H2O2 (30 % wt/wt; -2.82 goon / 1 mLELo 2) was added slowly over 16 minutes dropwise to the stirring slurry. The slurry changed from dark purple to black in color. The solids were filtered, rinsed with DI water and dried at 90 °C in an oven over night. The solids were then ground with a motor and pestle and seized through a 250 micron porosity sifter to collect 101.7 g of a loose and fluffy powder for calcination.

[0921 ] All of the powder was loaded into a quartz tube, winch acted as the boa t, with some space above to allow' gas flow. The quart tube boat was placed inside a larger quartz tube and placed into a unit for calcination. The calcination uni t had been thoroughly purged under nitrogen, both bulk and purified to ensure a sufficiently anaerobic environment for calcination. Purified nitrogen flowed over the sample at 150 standard cubic centimeters per minute. The sample was heated from room temperature to 600°C in 4 hours and held at 600°C for 4 hours and cooled to room temperature in 4 hours.

[0922] A small approximately 2 g sample of the resulting 100 g of catalyst was screened in the oxidative dehydrogenation reactor as described above and it had 25 % conversion at 376.5 °C and selectivity to ethylene at this conversion of 97%.

[0923] Disclosed here are methods for controlling the carbon dioxide output from an ODH process. A method includes introducing, into at least one ODH reactor a gas mixture of a lower alkane, oxygen and carbon dioxide, under conditions that allow production of the corresponding alkene and smaller amounts of various by-products. For multiple ODH reactors, each reactor contains the smite or different ODH catal st, provided at least one ODH catalyst is capable of using carbon dioxide as an oxidizing agent. In some embodiments, steam or other inert diluents may also be introduced into the reactor as part of the gas mixture. In some embodiments, the amount of carbon dioxide leaving the reactor is subsequently monitored. If the amount of carbon dioxide output is below a desired level then the amount of steam introduced into the reactor can be decreased. If the amount of carbon dioxide output is above the desired level then the amount of steam introduced into the reactor can be increased. Alternatively, the feed volumetric ratio of oxygen to lower alkane that is added to the at least one ODH reactor can be increased to decrease the carbon dioxide output, or the feed volumetric ratio of oxygen to lower alkane that is added to the at least one ODH reactor can be decreased to increase the carbon dioxide output.

[0924] In some embodiments, the lower alkane is ethane, and the corresponding alkene is ethylene.

[0925] In some embodiments, at least one ODH reactor is a fixed bed reactor. In some embodiments, at least one ODH reactor is a fixed bed reactor that includes heat dissipative particles within the fixed bed. In some embodiments, the heat dissipative particles have a thermal conductivity that is greater than the catalyst. In some embodiments, at least one ODH reactor is a fluidized bed reactor

[0926] In some embodiments, at least one ODH catalyst is a mixed metal oxide catalyst. In further embodiments, at least one ODH catalyst is a mixed metal oxide of the formula: Mo _a V _¾ Te _c Nb _</ Pd _* 0/, wherein a, b, c, d, e and f are the relative atomic amounts of the elements Mo, V, Te, Nb, Pd and O, respectively; and when a = 1 , b = 0.01 to 1.0, c = 0 to 1.0, d = 0 to 1 .0, 0.00 < e < 0.10, and f is a number to satisfy the valence state of the catalyst. [0927] Techniques are provided for oxidative dehydrogenation (ODH) of lower alkanes into corresponding alkenes. Lower alkanes are intended to include saturated hydrocarbons with from 2 to 4 carbons, and the corresponding alkene includes hydrocarbons with the same number of carbons, but with a single double carbon to carbon bond. While any of the lower alkanes can be converted to their corresponding alkenes using the methods disclosed here, the provided techniques can be implemented on the ODH of ethane, producing its corresponding alkene, ethylene.

[0928] The ODH Process

[0929] ODH of lower alkanes comprises contacting a mixture of a lower alkane and oxygen in an ODH reactor with an ODH catalyst under conditions that promote oxidation of the lower alkane into its corresponding alkene. Conditions within Site reactor are controlled by rite operator and include, but are not limited to. parameters such as temperature, pressure, and flow rate. Conditions will vary and can be optimized for a particular Sower alkane, or for a specific catalyst, or whether an inert diluent is used in the mixing of the reactants.

[0930] Use of an ODH reactor for performing an ODH process consistent with the provided techniques falls within the know'ledge of the person skilled in the art. For best results, the oxidative deh drogenation of a lower alkane may be conducted at temperatures from 300 °C to 370 °C, or from 300 °C to 365 °C, or from 330 °C to 360 °C, at pressures from 0.5 to 100 psig (3.447 to 689.47 kPag), or from 15 to 50 psig (103.4 to 344.73 kPag), and the residence time of the lower alkane in the reactor may be from 0.002 to 30 seconds, or from 1 to 10 seconds. While the ODH process is likely to occur at temperatures less than 370 °C, in some embodiments, it is not expected to be efficient or commercially viable.

[0931] In some embodiments, die process Isas a selectivity for the corresponding alkene (ethylene in the case of ethane ODH) of greater than 85%, or greater than 90%. The flow of reactants and inert diluent can be described in any number of ways known in the art. In some embodiments flow is described and measured in relation to the volume of all feed gases (reactants and diluent) that pass over the volume of the active catalyst bed in one hour, or gas hourly space velocity (GHSV). The GHSV can range from 500 to 30000 h ¹ , or greater than 1000 h ^! or greater than 1000 h ^-1 but up to 30000 h ¹ . The flow rate can also be measured as weight hourly space velocity (WHSV), which describes the flow in terms of the weight, as opposed to volume, of the gases that flow over the weight of the active catalyst per hour. In calculating WHSV the weight of the gases may include only the reactants but may also include diluents added to the gas mixture. When including the weight of diluents, when used, the WHSV may range from 0.5 h ¹ to 50 h ¹ , or from 1.0 to 25.0 h ¹ .

[0932] The flow of gases through the reactor may also be described as the linear velocity of the gas stream (m/s), which is defined in the art as the flow rate of the gas stream/cross-sectional surface area of the reactor/void fraction of the catalyst bed. The flow rate generally means the total of the flow rates of all the gases entering the reactor and is measured where the oxygen and alkane first contact the catalyst and at the temperature and pressure at that point. The cross-section of the reactor is also measured at the entrance of the catalyst bed. The void fraction of the catalyst bed is defined as the volume of voids in the catalyst bed/total volume of the catalyst bed. The volume of voids refers to the voids between catalyst particles and does not include the volume of pores inside the catalyst particles. The linear velocity' can range from 5 cm/sec to 1500 cin/sec, or from 10 cm/sec to 500 cm/sec.

[0933] The space-time yield of corresponding alkene (productivity) in g/hour per kg of the catalyst should be at least 100 or above, or greater than 500, or greater than 1500, up to a maximum of 10000 at 350 to 370 °C. in some embodiments, the productivity of the catalyst will increase with increasing temperature until the selectivity' is sacrificed.

[0934] ODH Catalyst

[0935] Any of the ODH catal sts known in the art are suitable for use with the methods disclosed here. When choosing a catalyst, a skilled user would appreciate that catalysts can vary with respective to selectivity and activity. For some embodiments of ODH of ethane, mixed metal oxides are die catalyst of choice as they can provide high selectivity to ethylene without significant loss in activity. Example catal sts are those of the formula:

Mo«ViTe _c Nb,iPd _e ^Q _/ wherein a, b, e, d, e and f are the relative atomic amounts of the elements Mo, V, Te, Nb, Pd and O, respectively; and when a = 1, b = 0.01 to 1.0, e = 0 to 1.0, d = 0 to 1.0, 0.00 < e < 0.10, and f is a number to satisfy the valence state of the catalyst.

[0936] ODH Reactor

[0937] Any of the known reactor ty pes applicable for the ODH of low'er alkanes may be used with the methods disclosed here. In some embodiments, the methods may be used with conventional fixed bed reactors. In a typical fixed bed reactor reactants are introduced into the reactor at one end, flow past an immobilized catalyst, products me formed and leave at the other end of Site reactor. Designing a fixed bed reactor suitable for the methods disclosed herein can follow techniques known for reactors of this type. A person skilled in the art would know which features are required with respect to sltape and dimensions, inputs for reactants outputs for products, temperature and pressure control, and means for immobilizing tire catal st.

[0938] Also contemplated are the use of inert non-catalytic heat dissipative particles within one or more of the ODH reactors, in some embodiments. In some embodiments, Site heat dissipative particles are present within the bed and comprise one or more non catalytic inert particulates having a melting point of at least about 100 °C in some embodiments, at least 250 °C in further embodiments, and in further embodiments at least 500 °C aud up to a maximum of 600 °C above the temperature upper control limit for the reaction, 600 °C, a particle size in range of 0.5 to 75 mm, in some embodiments 0.5 to 15 nun, in further embodiments in range of 0.5 to 8 mm, in further embodiments in the range of 0.5 to 5 m and a material thermal co nductivity of greater than 30 W/niK (watts/meter Kelvin) within the reaction temperature control limits. In some embodiments, the particulates are metals alloys and compounds having a thermal conductivity of greater than 50 W/mK (watts/meter Kelvin) within the reaction temperature control limits. Some suitable metals include silver, copper, gold, aluminum, steel, stainless steel, molybdenum, and tungsten.

[0939] The heat dissipative particles may have a particle size typically from about 1 to 15 mm. In some embodiments, the particle size may be from about 1 mm to about 8mm. The heat dissipative particles may be added to the fixed bed in an amount from 5 to 95 wt. %, in some embodiments 30 to 70 wt. %, in other embodiments 45 to 60 wt. % based on the entire weight of the fixed bed. The particles are employed to potentially improve cooling homogeneity and reduction of hot spots in the fixed bed by transferring heat directly to the walls of the reactor. [0940] Also contemplated is the use of a fluidized bed reactor, where typically, the catalyst is supported by a porous structure, or distributor plate, located near a bottom end of the reactor and reactants flow through at a velocity sufficient to fluidize the bed (e.g. the catalyst rises and begins to swirl around in a fluidized manner). The reactants are converted to products upon contact with the fluidized catalyst and subsequently removed from the upper end of the reactor. Design considerations include shape of the reactor and distributor plate, input and output, and temperature and pressure control, all of which would fall under knowledge of the person skilled in the art.

[0941 ] Techniques are provided for using a combination of both fixed bed and fluidized bed reactors each with the same or different catalyst. The multiple reactors may be in series, or parallel configuration, design of which Mis within the knowledge of the worker skilled in the art. [0942] Flammability Limits

[0943] Safety of the process is a primary concern. For that reason, mixtures of a lower alkane with oxygen should, in some embodiments, comprise ratios that fall outside of the flammability envelope. In some embodiments, a ratio of alkane to oxygen may fall outside the upper flammability envelope in this instance, the percentage of oxy gen in the mixture is less than 30 vol. %, less than 25 vol. %, or less than 20 vol. %.

[0944] In embodiments with higher oxygen percentages, alkane percentages may be adjusted to keep the mixture outside of the flammability envelope. While a person skilled in the art would be able to determine an appropriate level it is recommended that the percentage of alkane is less 40 vol. %. For instance, in an example where the mixture of gases prior to ODH comprises 20 vol. % oxygen and 40 vol. % alkane, die balance is made up with an inert diluent, such as one or more of nitrogen, carbon dioxide and steam. In some embodiments, the inert diluent should exist in die gaseous state in the conditions within the reactor and should not increase the flammability of the hydrocarbon added to the reactor, characteristics that a skilled worker would understand when deciding on which inert diluent to employ. Inert diluent can be added to eidier of die lower alkane containing gas or die oxygen containing gas prior to entering the ODH reactor or may be added directly into the ODH reactor.

[0945] In some embodiments, mixtures that fall within the flammability envelope may be employed, for example, in instances where the mixture exists in conditions that prevent propagation of an explosive event. That is, the flammable mixture is created within a medium where ignition is immediately quenched. For example, a user may design a reactor where oxygen and the lower alkane are mixed at a point where they are surrounded by flame arresting material. Any ignition would be quenched by the surrounding material. Flame arresting materials include but are not limited to metallic or ceramic components, such as stainless steel walk or ceramic supports. In some embodiments, oxygen and lower alkane can be mixed at a low temperature, where an ignition event would not lead to an explosion, then introduced into the reactor before increasing the temperature. The flammable conditions do not exist until the mixture is surrounded by the flame arrestor material inside of the reactor.

[0946] Carbon Dioxide Output

[0947] Carbon dioxide can be produced in the ODH reaction as a by-product of oxidation of the lower alkane. Carbon dioxide can also be added into the ODH reactor when used as an inert diluent. Conversely, carbon dioxide may be consumed when it acts as an oxidant for the dehydrogenation reaction. The carbon dioxide output is therefore a function of the amount of carbon dioxide added and produced minus that consumed in the oxidative process. In some embodiments, the disclosed methods control the degree to which carbon dioxide acts as an oxidizing agent so as to impact the overall carbon dioxide output coming off the ODH reactor.

[0948] Measuring the amount of carbon dioxide coming off the ODH reactor can be done using any means known in the art. For example, one or more detectors such as GC, IR, or Raman detectors, are situated immediately downstream of the reactor to measure the carbon dioxide output. While not required, the output of other components may also be measured. These include but are not limited to the amounts of ethylene, unreacted ethane and oxygen, and by-products such as acetic acid. In addition, it should be noted that depending on the chosen metric for carbon dioxide output, the output levels of the other components, for example ethane, may actually be required. [0949] Carbon dioxide output can be stated using any metric commonly used in the art. For example, the carbon dioxide output can be described in terms of mass flow' rate (g/min) or volumetric flow rate (cmVinin). In some embodiments, normalized selectivity can be used to assess the degree to which carbon dioxide is produced or consumed. In that instance the net mass flow rate of CO2 — the difference between the mass flow rate of CO2 entering and leaving the ODH reactor — is normalized to the conversion of ethane, in essence describing wha t fraction of ethane is converted into carbon dioxide as opposed to ethylene, or other by-products such as acetic acid. A carbon selectivity of 0 indicates that the amount of carbon dioxide entering the reactor is the same as the carbon dioxide output. In other words, the process is carbon dioxide neutral. A positive carbon dioxide selectivity alerts a user that carbon dioxide is being produced, and that any oxidation of carbon dioxide that is occurring is insufficient to offset that production resulting in the process being carbon dioxide positive.

[9950] When output of carbon dioxide or other components produced such as acetic acid and carbon monoxide, are described in terms of normalized product selectivity, the calculation is performed according to the formula:

Net mass flow rate of X min)

Molecular weight of X (g X /mol X)

Selectivity (wt%) mass How rate of converted C½H _e (g €½H _S / min) Mol. Equiv. of X

*

Molecular weight of C ₂ H ₃ mol C ₂ H ₆ (g C ₂ Hg / mol C _£ H _§ )

[Q951] When output of carbon dioxide where X is the product that is being assessed, the net mass flow rate refers to flow in g/min for X or ethane entering the reactor minus the flow rate exiting the reactor, and molar equivalent (Mol. Equiv.) refers to the amount of X, in moles, that reacts completely with one mole of ethane. Selectivity is refesxed to as a wt. % despite the fact the calculation results in converting wt. % to a molar percentage of carbon atoms because weight flow rate is the measurement that Is used in the calculation

[0952] One potential advantage of the provided techniques is the possibility of carbon dioxide negative process. In this instance, carbon dioxide is oxidized at a higher rate than it is produced and shows a negative carbon selectivity. The ODH process may be producing carbon dioxide, but the degree to which carbon dioxide is consumed while acting as an oxidizing agent offsets any production that is occurring. Many industrial processes, in addition to ODH, produce carbon dioxide which must be captured or flared where it contributes to the emission of greenhouse gases. Using a carbon dioxide negative process the excess carbon dioxide from other processes may be captured and used as the inert diluent in the ODH process under conditions where there is negative carbon selectivity. Another advantage, then, is the ability to reduce the amount of carbon dioxide produced in the ODH process in combination with other processes, such as thermal cracking. In addition, oxidation of carbon dioxide is endothermic and by increasing the degree to which carbon dioxide acts as an oxidizing agent, heat produced from ODH of ethane is partially offset by oxidation of carbon dioxide, reducing the degree to which heat must be removed from the reactor in some embodiments, when acting as an oxidizing agent, carbon dioxide can produce carbon monoxide, which can be captured and used as an intermediate in production of other chemical products, such as methanol or formic acid.

[09531 Addition of steam

[09541 The amount of steam added to die ODH process affects the degree to which carbon dioxide acts as an oxidizing agent. In some embodiments, steam may be added directly to the ODH reactor, or steam may be added to the individual reactant components — the lower alkane, ox gen, or inert diluent — or combinations thereof, and subsequently introduced into the ODH reactor along with one or more of the reactant components. Alternatively steam may be added indirectly as water mixed with either die lower alkane, ox gen or inert diluent, or a combination thereof, with the resulting mixture being preheated before entering the reactor. When adding steam indirectly as water, the preheating process should increase the temperature, so that the water is entirely converted to steam before entering the reactor.

[Q955] Increasing the amount of steam added to a reactor increases the degree to which carbon dioxide acts as an oxidizing agent. Decreasing the amount of steam added to the reactor decreases the degree to which carbon dioxide acts as an oxidizing agent. In some embodiments, a user monitors the carbon dioxide output and compares it to a predetermined target carbon dioxide output. If the carbon dioxide output is above the target a user can then increase the amount of steam added to the ODH process. If the carbon dioxide output is below' the target a user can decrease the amount of steam added to the ODH process, provided steam lias been added. Setting a target carbon dioxide output level is dependent on the requirements for the user. In some embodiments, increasing the steam added will have the added effect of increasing the amount of acetic acid and other by-products produced in the process. A user that is ill equipped to separate out larger amounts of acetic acid from the output of the ODH may prefer to reduce steam levels to a minimum, while a user that desires a process that consumes carbon dioxide may prefer to maximize the amount of steam that can be added. In some embodiments, the amount of steam added to one or more reactors is up to 60 wt. %. It should be noted that using wt. % to describe the amount of steam, or other components, added as part of the feed is a true wt. %, meaning it is the mass flow rate of the component divided by the total mass flow' of all feed components multiplied by 100. This is different than the use of wt % to describe product selectivity.

[0956] The effect of adding steam on the carbon dioxide output is more pronounced at lower temperatures.

For example, at temperatures ranging from 300 °C to 340 °C Site carbon dioxide selectivity may change from 1 wt. % to 20 w't. %, depending upon the change in steam added to the reaction. On the other hand, at higher temperatures, ranging from 350 °C to 370 °C, the change in carbon dioxide selectivity may range from 0.25 wt. % to 1.5%. [0957] In some embodiments, where reaction temperatures are less than 340 °C, changing the amount of steam added to the reactor by at least 20 wt. % results in a change in carbon dioxide output, measured as normalized product selectivity, of at least I wt. %.

[0958] In some embodiments, where reaction temperatures are less than 340 °C, and the amount of steam added to the reactor is from 0 wt. % to about 20 wt. % and is increased to from about 35 wt. % to about 60 w't. % the result is an absolute decrease in carbon dioxide output, measured as normalized product selectivity, of from 2.5 wt.

% to 15 wt. %.

[0959] In some embodiments, where reaction temperatures ate less titan 340 °C, and the amount of steam added to the reactor is from 35 wt. % to about 60 wt. % and is decreased to from about 20 wt. % to about 0 wt. % the result is an absolute increase in carbon dioxide output, measured as normalized product selectivity, of from 2.5 wt. % to 15 w't. %.

[0960] In some embodiments, where reaction temperatures are greater than 350 ^C C, and the amount of steam added to the reactor is from about 0 wt. % to about 10 wt. % and increased to from about 40 wt. % to about 60 wt. % and results in an absolute decrease in carbon dioxide output measured as normalized product selectivity, of from about 0.5 wl. % to 1.0 wt. %.

[Q961] In some embodiments, where reaction temperatures are greater than 350 °C, and the amount of steam added to the reactor is from about 40 wt. % to about 60 wt. % and decreased to from about 0 wt. % to about 10 wt.

% and results in an absolute increase in carbon dioxide output, measured as normalized product selectivity, of from about 0.5 wl. % to 1.0 wt. %.

[0962] When using two or more reactors it is contemplated that a user may choose to control carbon dioxide output in only one, or less than the whole complement of reactors. For example a user may opt to maximize carbon dioxide output of an upstream reactor so that the higher level of carbon dioxide can comprise part of the inert diluent for the subsequent reactor. In this scenario, addition of steam to the first reactor would be minimized while in the second reactor the addition of steam could be maximized to promote use of carbon dioxide as an oxidant. The carbon dioxide produced in the first reactor can act as both an inert diluent and as an oxidant in the second reactor. Maximizing carbon dioxide output upstream minimizes the amount of inert diluent that would need to be added to the stream prior to the next reactor.

[0963] There is no requirement for adding steam to an ODH process, as it is one of many alternatives for the inert diluent. For processes where no steam is added, the carbon dioxide output is maximized under the conditions used with respect to ethane, oxygen and inert diluent inputs. Decreasing the carbon dioxide output is then a matter of adding steam to the reaction until carbon dioxide output drops to the desired level In embodiments where oxidative dehydrogenation conditions do not include addition of steam, and the carbon dioxide output is higher than the desired carbon dioxide target level, steam may be introduced into the reactor while keeping relative amounts of the main reactants and inert diluent— lower alkane, oxygen and inert diluent — added to the reactor constant, and monitoring the carbon dioxide output, increasing the amount of steam until carbon dioxide decreases to the target level. [0964] In some embodiments, where carbon dioxide is not added as a diluent it is unlikely that a carbon dioxide negative process will occur. However, a carbon dioxide neutral process can be achieved by increasing steam added so that any carbon dioxide produced in the oxidative dehydrogenation process can then be used as an oxidizing agent such that there is no net production of carbon dioxide. Conversely, if a user desires net positive carbon dioxide output then the amount of steam added to the process can be reduced or elimina ted to help maximize carbon dioxide production. As the carbon dioxide levels increase there is potential to reduce oxygen consumption, as carbon dioxide is competing as an oxidizing agent. The skilled person would understand that using steam to increase the degree to which carbon dioxide acts as an oxidizing agent can impact ox gen consumption. The implication is that a user can optimize reaction conditions with lower oxygen contributions, which may assist in keeping mixtures outside of flammability limits.

[0965] Relative Volumetric Oxygen/Ethane Ratio

[0966] The relative volumetric oxygemethane ratio added to the ODH process can also impact the degree to which carbon dioxide acts as an oxidizing agent. Increasing the amount of oxygen added relative to the amount of ethane added decreases the carbon dioxide selectivity The degree which carbon dioxide selectivity changes is dependent upon the change in the relative volumetric ratio of oxygemethane added to the reactor and whether an inert diluent is included in the input stream. The effect is more pronounced in the absence of inert diluent, which, for safety reasons, limits the amount of oxy gen added to no more than 30 vol. %, or no more than 20 vol. %, in the absence of diluent. It is conceivable to use a much higher vol. % (¾, but in order to remain outside the flammability limits the corresponding amount of ethane would be restricted to levels below about 3 vol. %.

[0967] The relative volumetric oxygemethane ratio is determined by dividing the volume % of oxygen fed to the ODH process by the volume % of ethane added. For example, a gas mixture of 20 vol. % oxygen, 40 vol. % inert diluent, and 40 vol. % of carbon dioxide has a relative volumetric oxygen: etliane ratio of 0.5. The relative volumetric oxygemethane ratios should be specified such that the feed mixture containing oxygen, ethane and optional heat removal diluent remains outside the flammable envelope. In embodiments where there is an absence of inert diluent the relative volumetric oxygemethane ratios may fall between 0.1 and 0.45. In embodiments where inert diluent is present the oxygemethane ratio may range from 0.1 to 20

[0968] Altering the ratio of oxygemethane can be accomplished by keeping vol. % of either oxygen or ethane constant and then reducing or increasing the vol. % of either oxygen or ethane and then increasing or reducing the vol. % of inert diluent added to the process by an equivalent amount in some embodiments, the vol. % of oxy gen added is kept constant while the etliane vol. % is adjusted with corresponding adjustments to the vol. % of inert diluent added. When air is used as the source of oxygen the vol. % is adjusted to reflect the composition of air where oxygen is -21 vol. % and nitrogen is about -78%. The contribution of nitrogen would be used to calculate the vol.

% of inert diluent added to the reaction.

[0969] In some embodiments, changing the relative volumetric ratio of oxygen: etliane can be done by reducing the vol. % of either oxygen or etliane and increasing by a similar vol. % the one of oxygen and etliane that was not reduced, while keeping die vol. % of inert diluent added constant. [0970] In some embodiments, the amount of oxygen added is about 20 vol. % and the amount of ethane added ranges from 80 vol. % to 15 vol. % with corresponding ranges of inert diluent added to the ODH process ranging from 0 to 65 vol. %. Within these ranges the relative volumetric oxygemethane ratio ranges from 0.25 to about 1.33. [0971] In some embodiments, the amount of inert diluent added to the ODH process is from about 40 vol. % to about 55 vol. % and the oxygemethane ratio is about 0.30.

[0972] The effect of altering relative volumetric oxygemethane ratio added to the ODH process on carbon dioxide output, measured as carbon dioxide selectivity, can be a change in carbon dioxide selectivity of up to 5 wt. %. In some embodiments, the change in carbon dioxide selectivity is about 2.5 wt. %. In other embodiments, the change in carbon dioxide selectivity is about 1.0 wt. %.

[0973] Carbon dioxide negative

[0974] An aspect of the provided techniques is the ability of an operator to tailor conditions to promote oxidation of carbon dioxide so that the overall process is either carbon dioxide neutral or even carbon dioxide negative. By including carbon dioxide as or part of the inert diluent a net carbon dioxide negative process can be followed. This would allow using captured carbon dioxide from a process that produces carbon dioxide, minimizing the need to flare or convert the captured carbon dioxide. For example, a process of ODH of ethane results in a product stream that includes unreacted ethane, ethylene, water and one or more of carbon dioxide, acetic acid, and carbon monoxide. The wide variety of products necessitates separation downstream of the reactor. Acetic acid and water are removed using a quench tower, while carbon dioxide can be removed via a combination of an amine wash tower and a caustic tower. The remaining ethane and ethylene can be separated using a splitter so that the ethane can be recycled to the ODH reactor and the relatively pure ethylene can be used in downstream applications most notably polymerization using any known catalyst to make polyethylene. For example, the ethylene produced can be used to make low density polyethylene (LDPE), linear low density polyethylene (LLDPE), high density polyethylene (HOPE) and the lowest density products, elastomers and plastoniers using methods known in the art. [0975] The carbon dioxide removed by the amine wash tower would normally be flared off, contributing to the emission of greenhouse gases. The carbon dioxide can be used as inert diluent in the ODH process where the amount of steam added and relative volumetric oxygemethane ratio added are adjusted accordingly to promote oxidation of the carbon dioxide added in some embodiments, captured carbon dioxide from an ODH process separation train is used as inert diluent and the amount of steam added to the ODH process is adjusted so that carbon dioxide output is neutral or negative. A person skilled in the ait would appreciate that operating under carbon dioxide negative conditions cannot continue endlessly without external supply of carbon dioxide. As the supply of captured carbon dioxide approaches zero the operator can reduce the amount of steam added under the ODH process is carbon dioxide neutral.

[0976] In some embodiments, techniques are provided for a continual carbon dioxide negative process where carbon dioxide is supplied from an industrial process, such as thermal cracking. In this instance, the opportunity exists for reducing the amount of carbon dioxide that is to be flared under normal operating conditions for the industrial process. In this embodiment the operator maximizes the amount of steam added to the reaction and the relative volumetric oxygemethane ratio added to the reactor to decrease carbon dioxide selectivity so that added carbon dioxide added from an industrial process is almost entirely consumed.

[0977] Examples

[0978] Example G1

[0979] The effect of altering the amount of steam injected into an ODH process on the carbon dioxide output was assessed using two fixed bed reactors, connected in series. The catalyst present in each of the reactors was a mixture of several batches of a mixed metal oxide catalyst of the formula: Mo1 _. 0V0 _. 30-0 50Te0 _. 10-0 _. 20Nb0 _.i 0-0 _. 20O _d , where the subscripts represent the range of atomic amounts of each element, relative to Mo, present in the individual batches, and d represents the highest oxidation state of the metal oxides present in the catal st. The catalyst was extruded with 6.8 wt. % of T1O2. Furthermore, the catalyst was diluted physically with Denstone® 99 Alumina pow'der with weight ratio of catalyst to diluent of 2.1. Denstone® 99 Alumina consists mainly (99 wt. %) of alpha structure AI2O3. Ethane, carbon dioxide, and oxygen were premixed before addition of water, followed by preheating with the entire composition being fed to the first of the two reactors. The preheating step is necessary to ensure the water added is converted to steam before injection into the reactor. Output from the first reactor was sent directly into the second reactor without addition of new' reactants. For each reactor, the temperature was held in the range of 334 °C to 338 °C at ambient pressure. The process was run continuously over a period of three days.

[Q980] The relative amounts of ethane, carbon dioxide, and oxygen remained the same while the flow rate of steam added to reactor was altered. The relative amounts of ethane, carbon dioxide, and oxygen added to the first reactor were 33. 54, and 13 respectively. The gas hourly space velocity (GHSV) was kept constant at 610 h ^'1 . Flow rates of reaction ethane, carbon dioxide and oxygen were altered accordingly to maintain GHSV at 6IOI1 ^{' 1} after altering the amount of steam added to reactor.

[0981] Steam was added indirectly as water with the ethane, carbon dioxide and oxygen mixture. The amount of water added to the mixture before entering the first reactor wns varied, starting with no w'ater and increasing in increments up to a flow rate of l.OcmVmin For each flow rate of wrater added to the mixture, a corresponding weight % of steam in the total feed mixture w¾s calculated. Table G1 show's the effect that changing the amount of steam added to the reactor had on output of carbon dioxide, carbon monoxide, and acetic acid.

[0982] Results listed in Table G1 are averaged from two or more experimental mas at each of the prescribed conditions. The results show' that increasing the flow rate of water added to the mixture and corresponding increase in the weight % of steam added to the reactor leads to a decrease in the carbon dioxide selectivity. A carbon dioxide negative process was seen when the water was added at a flow rate of 1.0 cmVmin (Example Gl-5), which corresponds to 39 weight % of steam added. In addition, reverting back to no steam added (Example Gl-6) followed by increasing to 39 weight % (Example Gl-7) resulted in the carbon dioxide selectivity going positive back to negative. Table G1 reveals that changes in carbon dioxide selectivity me more pronounced when the levels of steam added, when reactor temperature is below' 340°C, is changed from below' 20% to 35 wt. % up to the maximum of 40 wt. %. Some of the data points shown in Table G1 are presented in the plot of Fig. 36 (Experiments Gl-6, Gl-2, G 1- 3, 01-4, Gl-7). [0983] Finally, it should be noted that increasing the steam results in a higher production of acetic acid and also is accompanied by a higher conversion rate of ethane.

[0984] Table G1 - Normalized product selectivity' of ODH products in response to changes in steam added to the reactor.

Temp - 334-338 °C; GHSV - 610 h ^l ; Vol ratio 0 :C I i . - 0.4

[0985] Example G2

[0986] A second experiment was conducted using the same reactor configuration from Example G1 but under different operating conditions. The catalyst was a mix of several batches as described for Example Gl, and for comparison included a freshly mixed catalyst and a mixed catalyst 8 months after being used intermittently. The relative volumetric amounts of ethane, carbon dioxide, and oxy gen added to the first reactor were 42, 37, and 21 respectively. Note the higher volumetric feed ratio of O /CVHg compared to Example Gl. In addition, the gas hourly space velocity (GHSV) was higher, and kept constant at 1015 h ¹ , with reaction temperature being held from between 321 °C to 325 °C. Similar to Example Gl flow rates of ethane, carbon dioxide and oxygen were altered accordingly to maintain GHS V at 1015 hr ¹ after altering the amount of water added. The corresponding steam content added to tire first reactor was changed from 0 wt. % to 16 wt. %.

[0987] The results of Example G2 are shown in Table G2. As the catalyst ages, selectivity towards the production of by-products, most notably CO2, generally increases, with a concomitant decrease in ethylene selectivity. This can be seen by comparing experiment G2-1 with experiment G2-2, where experiment G2.-1 corresponds to the fresh catalyst and experiment G2.-2 corresponds to the 8 month old catalyst. Originally, the catalyst showed 91% selectivity to C _? ¾ and a negative CO _? selectivity of -1.0. After 8 months, selectivity to € _? ¾ dropped to 89% and CO _? selectivity moved into positive territory at 5.0. Experiment G2.-2 shows that the method is also effective with an older catalyst, as increasing weight % of steam added to reactor from 0 to 16 weight % results in a drop in CO _? selectivity to 3.0 from 5.0 (Experiment G2-3). This decrease is in good agreement with the observed trend in Example Gl. [0988] Table G2 - Normalized product selectivity of ODH products using higher feed ratio of () /( ^' l i _¬ ana with fresh versus used catalyst.

Temp - 321-325 °C; GHSV - 1015 h ^’1 ; Vol ratio 0 ₂ :C ₂ H ₆ - 0.5

[0989] Example G3

[0990] A third experiment was conducted using the same reactor configuration as the previous examples but only using the second reactor in the series and under variable feed volume ratios of oxygen to ethane. The catalyst used was a mixed metal oxide catalyst of the formula: Mox.oVo.3 ₇ Teo.23Nbo 14O4.97 and was extruded with ~55 wt. % of Versal 250 in balance mixed metal oxide. Three relative volumetric amounts of oxygen and ethane were tested, including 16 vol. % <¾: 38 vol. % C ₂ ¾, 19 vol. % 0 ₂ : 36 vol. % C ₂ H ₆ , and 21 vol. % 0 ₂ : 33 vol. % € ₂ H ₆ , which correspond to 0 ₂ :C ₂ H ₆ volumetric ratios of 0.4, 0.5, and 0.6, respectively. The relative volumetric amount of C0 ₂ added was maintained at 46 vol. %, the gas hourly space velocity (GHSV) was kept constant at 1111 lr ^! , the reaction temperature was held between 359 °C and 360 °C, and reactions were performed at ambient pressure. No steam was added to the reaction.

[0991] The results of Example G3 are shown in Table G3. As the volumetric ratio of oxygemethane is increased the selectivity towards the production of C0 ₂ decreases. This effect is accompanied by slight increases to selectivity towards ethylene and carbon monoxide, while acetic acid selectivity remains unchanged. Example G3 demonstrates that altering volumetric ratio of oxygemethane added to the reactor while keeping other parameters unchanged, can decrease the selectivity to carbon dioxide. This effect is also demonstrated by comparing Examples Gi and G2, specifically experiment numbers Gl-1 and G2-1 where no steam was added, in that the carbon selectivity was lower in experiment number G2-1 where a higher volumetric ratio of oxygemethane was added to the reactor.

[0992] Table G3 - Normalized product selectivity of ODH products in response to variations of volumetric feed ratio of 0 ₂ /€ ₂ H ₆ at elevated temperature and without the addition of steam.

Temp - 359-360 °C; GHSV - 1110 hr ¹ ; Steam added - 0 vol. %

[0993] Example G4

[0994] A fourth experiment was conducted using the same reactor configuration as the previous examples and similar to Example G1 but using a higher volumetric ratio of oxyge ethane (0.5) added to the reactor a higher GHSV (1111 h ¹ ), and a higher temperature of 360 °C. The weight % of steam added to the reactor was changed from 0 wt. % to 40 wt. % while keeping the relative volumetric amount of CO ₂ steam added (46 vol. %) constant. The results are presented in Table G4 and demonstrate that at higher temperatures, flow' rates and volumetric ratio of oxygemethane increasing the amount of steam added to the reactor from 0 wt. % to 40 wt. % decreases the C0 ₂ selectivity. In tins example, the C<¼ selectivity decreased from 6.0 wt. % to 5.3%. This decrease is lower than what is seen when operating at a lower temperature, low' flow rate (GHSV), and lower relative volumetric ratio of oxygemethane added to the reactor.

[0995] Table G4 - Normalized product selectivity of ODH products in response to changes in steam added to the reactor at higher temp., GHSV, and vol. ratio Ctyeihane.Temp - 360 °C; GHSV - 1111 h-. Vol ratio ()·:( ·H. - 0.5

[0996] Techniques are provided for a process for production of olefins. In particular, the process involves an oxidative dehydrogenation process of one or more C ₂ -C ₄ alkanes to one or more C ₂ -C ₄ alkenes with differing feed ratios of ethanol to steam. The oxidative deh drogenation product distribution can be altered by co-feeding different ratios of ethanol to steam, in the range of 0.01 to 0.50 ethanoksteam to an oxidative dehydrogenation reactor. Increasing the ethanol to steam ratio in the reported range was found to: increase ethylene yield; decrease acetic acid yield; increase, or decrease, or cause no effect on CO or C0 ₂ yield; and have negligible effect on ethane conversion. The feed ethanol is converted to carbonaceous products without negatively effecting the catalyst activity.

[0997] As mentioned, the oxidative dehydrogenation product distribution can be altered by co-feeding different ratios of ethanol to steam, in the range of 0.01 to 0.50 ethanol :stearn to an oxidative dehydrogenation reactor. Increasing the ethanol to steam ratio in the reported range was found to: increase ethylene yield; decrease acetic acid yield; increase, or decrease, or cause no effect on CO or COi yields; and have negligible effect on ethane conversion. The feed ethanol is converted to carbonaceous products without negatively affecting the catalyst activity. In an embodiment, the present techniques provides a process for increasing C2-C4 alkene yield from an oxidative dehydrogenation reactor system, the oxidative deh drogenation reactor system including at least one oxidative dehydrogenation reactor, the oxidative dehydrogenation reactor containing at least one bed of mixed metal oxide catalyst, a feed stream to the oxidative dehydrogenation reactor including not less than 20 vol. % of one or more C2-C4 alkanes, the process including: forming an alkanol stream comprising a C2-C4 alkanol and steam and an optiomsl heat removal diluent gas wherein the volumetric ratio of C2-C4 alkanol to steam is in the range of 0.01 to 0.50; heating the alkanol stream to a temperature above its dew point; and co-feeding the alkanol stream with the feed stream to the oxidative dehydrogenation reactor, wherein the C2-C4 alkene yield is increased by at least 1 g alkene/(kgcat-hr).

[Q998] “Alkanol stream ^” refers to a stream comprising a C2-C4 alkanol and steam. “C2-C4 alkanol” refers to one of ethanol, n-propanol, isopropanol, n-butanoJ, sec-butanol, isobutanoi or tert-butanol, ora combination thereof. [0999] The feed stream to the ODH reactor comprises not less than 20 vol. % of one or more C2-C4 alkanes, up to 30 vol. % oxygen, with or without a heat removal diluent gas which is selected from inert gases, N ₂ , C02, and steam

[1000] It lias now been found that the addition of an alkanol stream comprising a C2-C4 alkanol and steam and an optional heat removal diluent gas to the ODH reactor can provide a substantial improvement over adding only the feed stream to the ODH reactor. The alkanol stream can be fed to the ODH reactor as a separate feed, or the alkanol stream can be added to the feed stream which feeds the ODH reactor, or the alkanol stream can be split between a separate feed and adding to the feed stream both of which feed the ODH reactor.

[J 0011 The C2-C4 alkanol can be ethanol, butanol, or propanol, or a combination of ethanol and butanol, ethanol and propanol, butanol and propanol, or a combination of all three alkanois. In all cases, the alkanol stream comprises water in the form of steam. The volumetric ratio of C2-C4 alkanol to steam is in the range of 0.01 to 0.50, in the range of 0.01 to 0.25, or in the range of 0.01 to 0.10.

[1002] The optional heal removal diluent gas in the alkanol stream can be comprised of N ₂ , C0 ₂ , or another heat removal diluent gas, or a mixture of heat removal diluent gases.

[1003] The alkanol stream should be above its dew point temperature prior to being fed to the ODH reactor. The temperature of the alkanol stream can be increased by heating the alkanol stream above its dew point. If at least part of the alkanol stream is added to the feed stream which feeds the ODH reactor, the temperature of the combined alkanol stream and feed stream which feeds the ODH reactor should be above its dew point temperature. f 1004] When an alkanol stream was added to an ODH reactor producing ethylene, the addition of an alkanol stream was found to both increase ethylene yield, and decrease the acetic acid yield. The addition of an alkanol stream was found to have no effect on either CO or CO2 yield, and have negligible effect on ethane conversion. The addition of an alkanol stream was found to have a negligible effect on the ODH catalyst activity.

[1005] The C2-C4 alkanol added to the ODH reactor system is ethanol, propanol, or butanol, or a mixture of two or all three of these molecules. The C2-C4 alkanol reacts with oxygen in the ODH reactor system to produce a carboxylic acid in some embodiments, the C2-C4 alkanol is preferably ethanol. The ethanol can react with oxygen to form acetic acid. The C2-C4 alkanol can also dehydrate in the ODH reactor system to form a C ₂ -C ₄ alkene. In some embodiments, the C ₂ -C ₄ alkanol is ethanol, and the C ₂ -C ₄ alkene is ethylene.

[1006] The ODH reactor has an ODH catalyst. The ODH catalyst in a second reactor may be Site same catalyst type as the ODH catalyst in a first reactor. The catalyst may be a low temperature catalyst that includes molybdenum (Mo), vanadium (V). tellurium (Te), niobium (Nb), and oxygen (Q) wherein the molar ratio of Mo to V is from 1 :0.12 to 1:0.49, the molar ratio of Mo to Te is from 1 :0.01 to 1 :0.30, the molar ratio of Mo to Nb is from 1 :0.01 to 1:0.30, and oxygen is present at least in an amount to satisfy the valency of any present metals. The molar ratios of Mo, V, Te, Nb can be determined by inductively coupled plasma mass spectrometry (ICP-MS). The catalyst may be a low temperature catalyst which provides for an ODH reaction at less than 450 °C, less than 425 °C, or less than 400 °C.

[1007] A Fixed Bed Reactor Unit (FBRU), shown as simplified reactor set up 100 in FIG. 37, was used to conduct experiments on manipulating the ODH product distribution by means of co-feeding different ratios of ethane to steam. The apparatus consisted of two fixed bed tubular reactors in series. Each reactor was wrapped in an electrical heating jacket and sealed with ceramic insulating material. Each reactor was a SS316L tube which had an outer diameter of F and 34” in length. In these experiments, ethane, ethylene, carbon dioxide, oxygen, nitrogen and liquid feed (water, alcohol or a mixture thereof) were fed separately (on as-needed basis) and pre-mixed and heated prior to the reactor inlet 3718 with the indicated composition (given in each experiment), and a temperature of greater than or equal to 220 °C. The flow' passed from the upstream reactor to the downstream reactor at stream 3719, and the product stream exited the downstream reactor at stream 3720. Both reactors were controlled at the same reaction temperature. The temperature of each of tire reactors were monitored using corresponding 7-point thermocouples shown by #l-#7 in the upstream reactor, and #8-#14 in the downstream reactor. The highest temperature between thermocouple points was used for controlling the reactor temperature using the corresponding back pressure regulator that controlled the pressure and boiling temperature of water inside the desired reactor wa ter jacket 3714. it is noteworthy that only thermocouple points #3 to #6 in the upstream reactor and #9 to #12 in the downstream reactor were located in the reactor bed, and the reaction temperature for each reactor was being reported as an average of these points.

[1008] The catalyst bed 3715 consisted of one weight unit of catalyst to 2.33 units of weight of SS 316 particles, and the total weight of catal st in each reactor was 150 g for example 4. The catalyst bed 3715 consisted of one weight unit of catalyst to 2.14 units of weight of Denstone® 99 powder, and Site total weight of catal st in each reactor was 143 g for examples Hi-I-13 and H5. The catalyst composition was measured by PIXIE analysis had the formula MoVo _. 3o-o _. 4oTeo _.i o-o2oNbo _l o-oroOx, in which X can be calculated based on the highest oxidation state of the metals present in this catalyst. The rest of the reactor, below and above the catalyst bed was packed with quartz powder 3716 and secured in place with glass wool 3717 on the top and the bottom of the reactor tube to avoid any bed movement during the experimental runs.

[1009] The present invention will be described by reference to the following examples. The following examples are merely illustrative of the invention and are not intended to be limiting.

[1010] In the following examples, it was found that when an alkanol stream was added to an ODH reactor producing ethylene, the addition of an alkanol stream was found to both increase ethylene yield, and decrease acetic acid yield. The addition of an alkanol stream was found to have no effect on either CO or C0 ₂ yield, and have negligible effect on ethane conversion. The addition of an alkanol stream was found to have a negligible effect on the ODH catalyst activity.

[1011] Example Hi: Changing ODH product distribution by varying the feed volumetric ratio of ethanol to steam

[1012] In order to explore the effect of feed volumetric ratio of ethanol to steam on ODH product distribution, the feed mixture of C2H5OH-H2O-C2H6-O2 was fed into the FBRU reactors (FIG. 37) with fixed GHSV, reaction temperature and feed volumetric ratio of €2¾:q2 (as shown in FIG. 38, Table la). The catalyst bed was diluted with diluted with Denstone© 99 alumina. The feed volumetric ratio of ethanol to steam was varied in the range of 0.01 to 0.10 by injecting different concentrations of ethanol in balance water into the reactor feed.

[1013] The catalyst activity is reported in FIG. 38, Table lb. Based on the results it can be seen that an increase in feed volumetric ratio of ethanol to steam led into the following trends: increase in ethane to ethylene selectivity? and yield; decrease in ethane to acetic add selectivity and yield; minor increase or negligible effect on ethane to CO and CO selectivity and yield; negligible change in ethane conversion; and 100% conversation of ethanol.

[1014] Based on the observed trends, the following observations can be made: ethane ODH product distribution can be manipulated towards generating more ethylene and less acetic acid by co-feeding different feed ratios of ethanol to steam; and portion of the co-fed ethanol can successfully be converted to ethylene.

[1015] Note that < 0.9 vol. % methanol impurity was observed in the injected liquid feed. Due to small quantity of this compound, the impurity was not reported and has a negligible effect on the catalyst activity calculation given in FIG. 38, Table lb. For all of the experiments, ail of the ethanol was converted to carbonaceous products. It was identified that these products are ethylene and acetic acid.

[1016] Example H2: Baseline before and after Example HI to show no catalyst deactivation [1017] The catalyst baseline activity ^? in the ODH reactors (FIG. 37) was tested prior to conducting the experiments in Example HI, and afterwards, to investigate whether presence of ethanol and steam caused arty catalyst deactivation. The catalyst activity results me shown in FIG. 39. Table 2, which shows that presence of ethanol and steam in tire ODH feed slightly increases the catalyst activity.

[1018] Example H3 : Removal of residual O2 in the last stage of an ODH reactor by reacting it with ethanol, and ethanol is mainly converted to ethylene and acetic acid f 1019] In order to explore the effect of ethanol in removing the residual 0 ₂ from the last stage of a ODH reactor, the feed composition of 8.96 vol. % C ₂ ¾-71.86 vol. % C2H4-O.49 vol. % G _2~ 0.49 vol. % €0 ₂ -17.43 vol. % H ₂ O-O.78 vol. % ethanol at GHSV of 648 h ^! and reaction temperature of 151-153 °C and inlet reactor pressure of 13.3-13.9 psig was fed to the ODH reactors (FIG. 37) for a duration of 29 hr:45 nun. The catalyst bed was diluted with Denstone® 99 alumina. The feed and product gas compositions as a function of elapsed time are shown in FIG. 40, Table 3a. Note that 0.5 vol. % methanol impurity was observed in the feed composition. Due to the small quantity' of this compound, the impurity was not reported and has no effect on the catalyst activity calculation. From this table: all of the <¾ present in the feed has been consumed; ethane conversion to carbonaceous products was found to be negligible; ethanol was partially converted to ethylene and acetic acid; and the C0 ₂ content in the feed and product stream remained almost unchanged.

[1020] The catalyst activity towards ethanol conversion to ethylene and acetic acid has been reported in FIG. 40. Table 3b.

[1021] Example H4: Effect of ethanol addition in the presence of ODH catalyst and SS 316 diluent

[1022] In order to understand the effect of feed ethanol on ODH product yield and distribution, two experiments were conducted using ODH catalyst diluted with stainless steel 316. During the experiments the reactor (FIG. 37) operating conditions remained unchanged, while the feed composition was varied to reflect the presence of 0 vol. % to 0.8 vol. % feed ethanol. The fixed and varied parameters are shown in FIG. 41, Table 4a. The ODH product yield and distribution for these experiments are reported in FIG. 41, Table 4b. The following observations can be made: addition of feed ethanol compared to no feed ethanol led into minor or negligible increase in ethane conversion; and addition of feed ethanol compared to no feed ethanol led into increase in ethylene yield, and decreases in the yields of acetic acid, CO and C0 ₂ .

[1023] Example H5 : Effect of ethanol addition with ODH catalyst and Denstone® 99 Alumina diluent

[ 024] In order to understand the effect of feed ethanol on ODH product yield and distribution over ODH cataly st diluted with Denstone® 99 alumina, two experiments were conducted. During the experiments the reactor (FIG 37) operating conditions remained unchanged, while the feed composition w'as varied from 0 vol. % to 0.8 vol. % feed ethanol. The details of the fixed and varied parameters are shown in FIG. 42, Table 5a. The ODH product yield and distribution for these experiments are reported in FIG. 42, Table 5b. The following observations can be made: addition of feed ethanol compared to no feed ethanol led to minor or negligible increases in ethane conversion; and addition of feed ethanol compared to no feed ethanol led to increases in yields of ethylene, CO, and C0 ₂ , and a decrease in acetic acid yield.

[1025] Based on the observations from Examples H4 and H5, ODH catalyst additives (Denstone® 99 alumina, and stainless steel 316), in the presence of feed e thanol versus no feed ethanol, ethylene yield increased, and acetic acid yield decreased, while negligible increase in ethane conversion was observed. Use of Denstone® 99 alumina caused an increase in CO and C0 ₂ yields, whereas use of stainless steel 316 cause a decrease in CO and C0 ₂ yield. Tins implies that both of the mentioned additives can improve ethylene yield in the ODH process while feed ethanol is present. f 1026] Example H6: Comparing the effect of ethanol addition with ODH catalyst in the presence of Denstone® 99 Alumina diluent versus SS 316 diluent

[1027] In order to explore the effect of feed volumetric ratio of ethanol to steam on ODH product distribution as a function of the two different ODH catalyst beds, namely ODH catal st physically mixed with Denstone® 99 alumina, and ODH catalyst physically mixed with stainless steel 316, the fixed feed mixture of C2H5OH-H2O-C2H6- O2 was fed into the reactors (FIG. 37) once with each of the two mentioned catalysts. During these experiments,

GHS V, reaction temperature and reaction pressure were kept unchanged, as shown in FIG. 43, Table 6a, ethanol was added to the reactor by injecting the appropriate concentration of ethanol in balance water into the reactor feed. The catalyst activity is reported in FIG. 43, Table 6b. By comparing the catalyst activity results over ODH catalyst mixed with stainless steel 316 to the catalyst activity results over ODH catal st mixed with Denstone® 99 alumina, (he following observations can be made: in (he presence of stainless steel 316 (compared to in presence of Denstone®

99 alumina), ethane conversion was increased; and in presence of stainless steel 316 (compared to in presence of Denstone® 99 alumina), yield towards C2H4 and CH3COOH increased, while yield toward CO and CO2 remained almost unchanged.

[1028] Limiting acetic acid production in ODH process. An oxidative dehydrogenation process to convert ethane to ethylene is disclosed, in which the oxidative dehydrogenation process includes acetic acid in the feed to the reactor is described. The production of acetic acid is limited during the oxidative dehydrogenation process to convert ethane to ethylene. The process of oxidative dehydrogenation includes feeding ethane and oxygen into a reactor where contact with a catalyst leads to conversion of the ethane into ethylene and acetic acid. By including acetic acid in the feed the amount of acetic acid produced is limited and the ratio of ethylene produced to ethane consumed increases. It is an object of the provided techniques to limit the degree to which acetic acid is produced in the oxidative dehydrogenation of ethane by including acetic acid along with ethane and oxygen in the initial feed to the process.

[1029] In one aspect, oxidative dehydrogenation of ethane is provided. More specifically, included are contacting a feed stream comprising ethane, oxygen acetic acid, and optionally an inert diluent, with an oxidative dehydrogenation catalyst in a reactor under oxidative dehydrogenation conditions to produce a product stream comprising ethylene, unreacted ethane, water, and acetic acid, wherein the concentration of acetic acid in the feed stream is from 0.01 to 5 vol. % of the feed stream.

[1030] In another aspect, included is a downstream separation that separates the product stream into a liquid stream comprising water and acetic acid and a gas stream comprising ethylene, ethane, carbon dioxide and carbon monoxide. In some embodiments, the liquid stream includes less titan 5 vol. % of other oxygenates. At least a portion of the liquid stream can be recycled to the reactor as part of the feed stream.

[1031] Techniques are provided for oxidative dehydrogenation of ethane into ethylene, under conditions that limit production of acetic acid. Typically, oxidative deh drogenation of ethane involves feeding a gas stream of ethane and oxygen, and optionally an inert diluent, into an oxidative dehydrogenation reactor comprising an oxidative dehydrogenation catal st. Contact of the ethane and oxygen with the oxidative dehydrogenation catalyst results in the formation of ethylene and various by-products, including acetic acid. The production of by-products such as acetic acid results in a need for costly downstream separation of the acetic acid from the target product ethylene. Reducing the amount of acetic acid produced can reduce the complexity and size of acetic acid downstream separation infrastructure.

[1032] The provided techniques, in one aspect, seeks to limit the production of acetic acid in a process of oxidative dehydrogenation of ethane, by including acetic acid in the feed stream. It was surprising found that by including from 0.01 vol. % up to 5 vol. % of acetic acid in the feed steam that the formation of acetic acid was partially or completely suppressed. In addition, by using acetic acid in the feed stream there is potential to increase ethylene yield and or ethane conversion. In another aspect, acetic acid produced in the oxidative dehydrogenation of ethane process can be recovered and recycled for addition to the feed stream. Typically, the feed stream comprises ethane and oxygen, and optionally an inert diluent. The feed to the reactor comprises a dry feed of one or more Ci- Ci alkanes, typically ethane; an oxygen containing gas, generally air or a synthetic air a mixture of air and nitrogen; carbon dioxide either as a component of air or as a recycle stream. The feed to the reactor may also comprise acetic acid from an external source or from a rec cle from another unit in the process such as the separator (or quench tower), or both.

[1033] Figure 44 is a schematic diagram in which the ODH reactor 44200 can be fed by stream 44100 which contains a dry feed of one or more C2-C1 alkanes, typically ethane; an oxygen containing gas generally air or a synthetic air a mixture of air and nitrogen, and carbon dioxide. The product from the ODH reactor 44200 is separated into a vapor 44400 and a liquid 44500 stream in the separator 44300. Acetic acid can be fed into the ODH reactor 44200 via stream 44600.

[1034] Figure 45 is a schematic diagram in which the ODH reactor 44200 can be fed by stream 44100 which contains a dry feed of one or more CY C alkanes, typically ethane; an oxygen containing gas, generally air or a synthetic air a mixture of air and nitrogen, and carbon dioxide. The product from the ODH reactor 44200 is separated into a vapor 44400 and a liquid 44500 stream in the separator 44300. Acetic acid can be fed into the ODH reactor 44200 via stream 44600 and also via a portion of the liquid product stream 45700 which has been recycled from the separator 44300.

[1035] The equipment to conduct the reaction also includes upstream of the reactor a vaporizer to vaporize any water and acetic acid from the separator (or quench tower). The mixture of water and acetic acid (the effluent from the separator or quench tower) may comprise from 0.01 to 30 vol. % of acetic acid, water and (the balance) less than 5 vol. % of oxygenates (e.g. CO, CO2, etc.). In some embodiments the stream from the separator (or quench tower) comprises less than 10 vol. % of acetic acid, in some embodiments less than 5 vol. % acetic acid.

[1036] The effluent from the separator (or quench tower) may be diluted with water to reduce the concentration of acetic acid in the effluent to less than 10 vol. % of acetic acid. Generally, not all of the effluent from the separator (or quench tower) is required to be recycled to the ODH reactor inlet. The unused effluent may be sent for further processing such as upgrading to glacial acetic acid, or vinegar, etc. Typically foe effluent feed to the ODH reactor provides 0.1 to 5 vol. % of acetic acid in foe gas vapor feed to the ODH reactor; an OiiCiHe in a volume ratio (mole ratio) from 0.5 to 0.7 and I¾0 and CO2 in a volume fraction so that foe feed composition is outside (above or below) the flammable envelope. In some instances, the feed to the reactor comprises from 2/57/16/8/17 vol. %to 3/29/26/14/28 vol. % of acetic acid H O/C Hs/Oz/CO,.

[1037] The vaporizer is operated at temperatures to vaporize the effluent feed. As the reactor operates at temperature of not less than about 300 °C the feed to the ODH reactor is typically from 150 °C - 250 °C, and a small portion of the reactor is used to heat the feed to the catalyst bed inlet temperature, e.g. from 320 °C to 350 °C, or from 330 °C to 340 °C. This portion of the bed is typically loaded with heat conductive non-catalytic material. The stream from the vaporizer does not pass through the mixer and is fed directly to the inlet of the ODH reactor.

[1038] The feed gas plus the gas from the vaporizer passes through the catalyst bed where some of the ethane is converted to ethylene. The product stream from the separator (or quench tower) of the ODH reactor may be further subject to a number of treatments as outlined above.

[1039] Catalysts

[1040] There are a number of catalysts which may be used in die oxidative dehydrogenation of lower alkanes to their corresponding alkenes such as ethylene and propylene.

[1041] Typically, the oxidative dehydrogenation catalysts comprise a mixed metal oxide catalysts selected from the group consisting of:

[1042] catalysts of the formula:

Mo _fl ViT e _r Nbr _f PdcO/ wherein a b, c, d, e and f are the relative atomic amounts of the elements Mo, V, Te, Nb, Pd and O, respectively; and when a = 1, b = 0.01 to 1.0 preferably O.lto 0.4, c = 0 to 1.0 preferably 0.1 to 0.3, d = 0 to 1.0 preferably 0.1 to 0.3, 0.00 < e < 0.10 preferably from 0.03 to 0.1 and f is a number to satisfy the valence state of the catalyst;

[1043] catalysts of the formula:

N i _g A*B ,D _/ 0/ wherein: g is a number from 0.1 to 0.9, preferably from 0.3 to 0.9, most preferably from 0.5 to 0.85, most preferably 0.6 to 0.8; h is a number from 0.04 to 0.9 preferably 0.4 to 0.6; i is a number from 0 to 0.5, preferably from 0.01 to 0.3; j is a number from 0 to 0.5 preferably from 0.01 to 0.3; and f is a number to satisfy the valence state of the catalyst; A is selected from the group consisting of Ti, Ta, V, Nb Hf W, Y, Zn, Zr, Si and Al, preferably V, Nb, Hf, and W, or mixtures thereof; B is selected from the group consisting of La, Ce, Pr, Nd, Sin, Sb, Sn, Bi, Pb, ΊΊ, In, Te, Cr, Mn, Mo, Fe, Co, Cu, Ru, Rh, Pd, Pt, Ag, Cd, Os, Ir, Au, Hg, preferably La, Ce, Sb, Sn, Bi, In, Te, Cr, Mn, Mo, Fe, Co, Cu, Ru, Rh, Pd, Pi, Os, and Ir, and mixtures thereof; D is selected from the group consisting of Ca, K, Mg, Li, Na, Sr, Ba, Cs, and Rb, preferably Ca, K, Mg, Sr, Ba, Cs, and Rb, and mixtures thereof; and O is oxygen;

[1044] catal sts of the formula: ooEtG/O _/ wherein: E is selected from the group consisting of Ba, Ca, Cr, Mn, Nb, Ta, Ti, Te, V, W and mixtures thereof; G is selected from the group consisting of Bi, Ce, Co, Cu, Fe, K, Mg, V, Ni, P, Pb, Sb, Si, Sn, Ti, U. and mixtures thereof; a = i; k is 0 to 2, preferably 0.2 to 0.6; 1 = 0 to 2. preferably 0.2 to 0.6. with the proviso that the total value of 1 for Co, Ni, Fe and mixtures thereof is less than 0.5; and f is a number to satisfy the valence state of the catalyst; [1045] catalysts of the formula: V _ra Mo„Nb ₀ T %,Me _? 0/ wherein: Me is a metal selected from the group consisting ol ^' Ta, Ti, W, Hf, Zr, Sh and mixtures thereof; m is from 0.1 to 3, in some instances from 0.5 to 1; n is from 0.5 to 1.5, in some instances from 0.5 to 1; o is from 0.001 to 3 in some instances from 0.01 to 1; p is from 0.001 to 5 in some instances from 0.01 to 1; q is from 0 to 2 in some instances from 0.03 to 1; and f is a number to satisfy the valence state of the catalyst; and [1046] catalysts of the formula:

Mo nC,U;Z _k M _n O/ wherein: X is at least one of Nb and Ta; Y is at least one of Sb and Ni; Z is at least one of Te. Ga, Pd, W, Bi and A1 , in some embodiments Te, Pd, W, and Bi; M is at least one of Fe, Co, Cu, Cr, Ti, Ce. Zr, Mn, Pb, Mg, Sn, Pt, Si, La, K, Ag and In in some instances Fe, Co, Cu, Cr, Ti, Ce, Zr, Mn, Mg, Sn, Pt, La, Ag and In ; a=1.0 (normalized); r = 0.05 to 1.0 in some embodiments 0.05 to 0.5; s = 0.001 to 1.0 in some embodiment 0.01 to 0.4; t = 0.001 to 1.0 in some embodiment from 0.01 to 0.4; u 0.001 to 0.5 in some embodiments 0.01 to 0.03; v 0.001 to 0.3 in some embodiments from 0.01 to 0.2; and f is a number to satisfy the valence state of the catalyst.

[1047] Although some embodiments of catalyst that may be used in the process are shown above, any number of other catalysts may be used in addition to, or instead of, these catalysts. This includes any of the other catalysts described herein.

[1048] EXAMPLES

[1049] A Fixed Bed Reactor Unit (FBRU) was used to conduct experiments on manipulating the ODH product distribution by means of co-feeding different ratios of ethane to steam. An example of the apparatus is shown in Figure 37.

[1050] Example II. Baseline Case

[1051] In order to explore the effect of feed steam (in absence of acetic acid) on ODH catalyst activity and product distribution experiments were conducted at fixed reaction and feed operating conditions while the wet feed composition was changed as shown in Table II . Note that all the data reported in Table II are the average of two or more experimental data. The “Liquid Water” column refers to water at ambient temperature and pressure that was added to the inlet of the reactor to generate the desired wt. % of steam in the gaseous/vapor feed effluent as reflected in the column under the name of “Steam added (wt. %)”.

[1052] Table 11 : ODH product distribution at different flow rate of feed water [1053] The experimental conditions were kept constant: relatively constant value of WHSV 0.55 -0.76 h ¹ (GHSV = 610 h ¹ ), reaction temperature 334 °C - 338 °C, reaction pressure is at ambient pressure, and dry feed gas composition of C2H6/O2/CO2 of 33/14/54 (vol. %). The experimental condition that was varied was the quantity of water in the GDH feed which was changed in the range of 0 - 1 ciriVmin. At each water flow rate, tire reactors were operated for 1 - 3 days. The reactors were operated without interruption as the water flow ⁷ rate was being changed [1054] The results are shown in Table 11. Looking at the results, it can be seen that tire increase in the flow rate of feed water from 0 - 1 cm ³ /min (experiment II -1 to 11-5 in Table II), led to an increase in ethane conversion and acetic acid selectivity/yield.

[1055] Example 12. 2 vol. % Acetic Acid Injection to Display Product Flexibility

[1056] In order to explore the effect of feed of acetic acid on ODH catalyst activity and product distribution, two experiments were conducted at the following fixed reaction and feed operating conditions while the wet feed composition was changed as shown in Table 12.

[1057] Table 12: Wet Feed Composition of the Conducted Experiments

[1058] in these experiments, the gas hourly space velocity was GHSV = 648b ^'1 , reaction temperature = 320 °C-321 °C, reaction operating pressure ~ atmospheric, dry feed composition of C2H6/O2/CQ2 = 38/21/41 vol. %. The catalyst activity results are shown in Table 13.

[1059] Table 13 : Catal st Activity for the Conducted Experiments

[1060] The presence of 2 vol. % acetic acid (or 0.03 volume ratio of acetic acid to water) in the feed led to partial suppression of acetic acid formation, no change to ethylene yield, and no change in ethane conversion. This finding supports ODH product flexibility, as by changing the ratio of feed acetic acid to water the main result is suppression of formation of acetic acid in the process.

[1061] The catalyst baseline activity was tested prior to conducting experiment 12-1 and afterwards to investigate whether presence of feed steam-acetic acid (during Experiment 12) could case any catalyst deactivation. The catalyst activity results are shown in Table 14 which shows that presence of acetic acid-steam in the ODH feed had no effect on catalyst activity.

[1062] Table 14 ; Cataly st Baseline Activity Prior to and After Experiment 12-1

[1063] Example 13, 3 voL % Acetic Acid Injection to Display Product Flexibility

[1064] In order to explore the effect of feed of acetic acid on ODH catalyst activity and product distribution, two experiments were conducted at the following fixed reaction and feed operating conditions while the wet feed composition was changed as shown in Table 15.

[1065] Table 15 ; Wet Feed Composition of the Conducted Experiments

[1066] In these experiments, the gas hourly space velocity was GHSV = 64810, reaction temperature = 324 °C, reaction operating pressure ~ atmospheric, dry feed composition of C2 JO2/CO2 - 38/21/41 vol. %. The catalyst activity results are shown in Table 16. f 1067] Table 16: Catalyst Activity for the Conducted Experiments

[1068] The presence of 3 voi. % acetic acid (or 0.1 volume ratio of acetic acid to water) in the feed led to: complete suppression of acetic acid formation, almost no change to ethylene yield, and an increase in ethane conversion. This finding reports ODH product flexibility by means of changing the ratio of feed acetic acid to water mainly suppresses formation of acetic acid in the process.

[1069] The catalyst baseline activity was tested prior to conducting Experiment 13-1 and afterwards to investigate whether the presence of feed stream-acetic acid during Experiment 13 could cause any catalyst deactivation. The catalyst activity results are shown in Table 17, which show's that presence of acetic acid-steam in the ODH feed had no effect on catalyst activity.

[1070] Table 17 : Catalyst Baseline Activity Prior to and After Experiment 3- 1

[1071] Experiment 14 - AspenPlus™ Simulation Data with and without Acetic Acid in tire Feed

[1072] AspenPlus™ simulations were used to compare an ODH plant using (1) no acetic acid in the feed, and

(2) acetic acid fed to the reactor up to 0.3 kg per kg ethane. The resulting data is shown in Error! Reference source not found.6. The simulation results show' the following trends: ( 1) there was less acetic acid in the product than the feed at a ratio of acetic acid in the feed to ethane in the feed of 0.1 kg/kg and higher; and (2) there was a higher ratio of eth lene produced to ethane consumed as the amount of acetic acid in the feed to ethane in the feed was increased. [ 1073] Some embodiments include product diversification in the oxidative dehydrogenation (ODH) of ethane to ethylene to generate coproducts, such as ethanol and acetaldehyde, in addition to the product ethylene.

Techniques are provided for low -temperature reactions (e.g., below 370 °C) that produce ethylene from ethane and that may generate byproducts, such as carbon monoxide (CO), carbon dioxide (CO2), and acetic acid. Low- temperature ODH catalyst may be utilized. In embodiments, at least two ODH reactors in series are employed. Additional ODH reactors may be included in parallel. In some embodiments, a second ODH reactor operates in series downstream of a first ODH reactor in certain implementations, the two reac tors in series may have the same or similar type of ODH catalyst.

[1074] For implementations of two reactors (each having ODH catal st) in series, tire first reactor may produce ethylene and the second reactor may convert sortie of the ethylene produced in the first reactor to ethanol and acetaldehyde as coproducts of the ethylene. Production of these coproducts may enhance profitability of the ODH reactor system that produces eth lene. The balance of ethylene production ethanol production, and acetaldehyde production (and other coproduct production) by tire ODH reactor system may be adjusted depending on market need. Other coproducts may include, for example, acetic acid that is as a byproduct in the ODH conversion of ethane to ethylene in the first reactor. Acetic acid may be produced from ethylene in the first reactor. Acetic acid may also be produced in the second reactor.

[1075] The block flow diagram of the ODH system depicted in FIG. 47 is an embodiment of the present techniques. As indicated by FIG. 47, a product effluent stream from a first reactor (having ODH catalyst) can be sent to a second reactor (having ODH catalyst) to convert carbonaceous material (e.g., ethylene) from the first reactor into ethanol and acetaldehyde in the second reactor.

[1076] For example, in the second reactor, ethanol may be formed by ethylene hydration: [1] C2H4 + ¾0 C2¾0 as a bulked chemical reaction. Acetaldehyde may be formed in the second reactor by ethanol oxidative dehydrogenation in presence of CO2 as a mild oxidant: [2] ChHgO + CO2 Ά C2H4O + H2O + CO as a bulked chemical reaction. In certain implementations, the amount of production of ethanol (e.g , via reaction [1]) and acetaldehyde (e.g., via reaction [2]) versus the amount of ethylene production may be adjusted depending, for example, on market need and other factors. Moreover, it may be beneficial that oxygen (O2) not be present in second reactor to avoid oxidation of ethanol and acetaldehyde to acetic acid in the second reactor. Lastly, the residence time in the second reactor to generate ethanol via the aforementioned reaction [1] (ethylene hydration) and acetaldehyde via the aforementioned reaction [2] (ethanol oxidative dehydrogenation) may be similar or at least within an order of magnitude of the residence time to generate ethylene from ethane via the ODH reaction in the first reactor.

[1077] FIG. 47 is an ODH coproduction system 4700 having an ODH reactor system 4702 with a first ODH reactor 4704 and a second ODH reactor 4706 in series. The reactors 4704, 4706 are labeled as “ODH” reactors because the reactors 4704, 4706 have ODH catalyst and because ODH reactions may occur in the reactors 4704, 4706. Reactions other than ODH may also occur in the reactors 4704. 4706. The “oxidative deh drogenation” or “ODH” reaction may refer to Site combination of endothermic dehydration of an alkane and exothermic oxidation of hydrogen. [1078] In some cases, the two ODH reactors 4704, 4706 in series may operate at similar operating conditions of temperature and pressure. However, in other instances, the operating temperature in the first ODH reactor 4704 is different than the operating temperature in the second ODH reactor 4706. The operating pressure in the second ODH reactor 4706 may be less than the operating pressure in the first ODH reactor 4704 to facilitate flow of effluent from the first ODH reactor to the second ODH reactor.

[1079] The first ODH reactor 4704 and the second ODH reactor 4706 may each be a fixed-bed reactor (e.g., a tubular fixed-bed reactor), f!uidized-bed reactor, ebullated bed reactor, or heat-exchanger type reactor, and so on.

The reactors or reactor s stem generally utilize a heat-transfer fluid for removing heat from the reactor for temperature control of the reactor. The heat transfer (cooling) medium can be, for example, oil, molten salt, or pressurized water.

[1080] For a fixed-bed reactor, reactants may be introduced into the reactor at one end and flow past an immobilized catal st. Products are formed and an effluent having the products may discharge at the other end of the reactor. The fixed-bed reactor may have one or more tubes (e.g., ceramic tubes) each having a bed of catalyst and for flow of reactants (e.g., ethane in first reactor 4704, ethylene in second reactor 4706) and products (e.g., ethy lene in first reactor 4704, ethanol and acetaldehyde in second reactor 4706). The tubes may include, for example, a steel mesh. Moreover, a heat-transfer (e.g., cooling) jacket adjacent the tirbe(s) may provide for temperature control of the reactor. The aforementioned heat transfer fluid may flow' through the reactor jacket.

[18811 In other embodiments, the first ODH reactor 4704 and the second ODH reactor 4706 are each a fluidized bed reactor. In implementations, a fluidized bed reactor may have a support for the ODH catalyst. The support may be a porous structure or distributor plate and disposed in a bottom portion of the reactor. Reactants may flow' upward through the support at a velocity to fluidize the bed of ODH cataly st (e.g., the catalyst rises and begins to swirl around in a fluidized manner). The reactants are converted to products upon contact with the fluidized cataly st. An effluent having products may discharge from an upper portion of the reactor. The fluidized bed reactor may have a heat-transfer (e.g., cooling) tubes, jacket, or heat pipes to facilitate temperature control of the reactor.

The aforementioned heat transfer fluid may flow' through the reactor tubers or jacket. Lastly, the fluidized bed reactor can be (1) a non-circulating fluidized bed, (2) a circulating fluidized bed with regenerator, or (3) a circulating fluidized bed without regenerator.

[1082] As indicated, the first reactor 4704 lias an ODH catalyst 4708 and the second reactor 4706 also has an ODH catalyst 4710. The ODH catalyst 4710 in the second reactor 4706 may be the same catalyst type as the ODH catalyst 4708 in the first reactor 4704. In some implementations, the catalyst 4708, 4710 is a low' temperature catalyst that includes molybdenum, vanadium, tellurium, niobium, and oxygen wherein tire molar ratio of molybdenum to vanadium is from 1:0.12 to 1 :0.49, the molar ratio of molybdenum to tellurium is from 1:0.01 to 1:0.30, the molar ratio of molybdenum to niobium is from 1:0.01 to 1:0.30, and oxygen is present at least in an amount to satisfy the valency of any present metal oxides. The molar ratios of molybdenum, vanadium, tellurium, and niobium can be determined by inductively coupled plasma mass spectrometry (ICP-MS). The catalyst may be low temperature in providing for an ODH reaction at less than 370 °C, less than 360 °C, or less than 350 °C. [1083] The catalyst 4708, 4710 may be a low-temperature ODH catalyst available from NOVA Chemicals Corporation having headquarters in Calgary, Canada in particular, this implementation of catalyst is a mixed metal oxide having the formula Mo _a V _b Te _c Nb _d Pd _e O _f , where a, b, c, d, e, and f subscripts are relative atomic amounts of the elements Mo, V, Te, Nb, Pd, O, respectively. When a=l, then fori).01 to 1.0, c=0.01 to 1.0, d=0.01 to 1.0, O.OO<e<0.1O, and f is a number to satisfy the valence state of the catalyst. This catalyst may provide for the ODH reaction to occur at a temperature of less than 350 °C or less than 370 °C. in addition, the ODH catalyst 4708, 4710 may be, for example, a fixed bed or in a fluidized bed.

[1084] In operation, the first ODH reactor 4704 receives a feed 4712 that may be one or more streams. The feed 4712 may include ethane, oxygen, diluent, and so forth. The diluent may be, for example, carbon dioxide or nitrogen. Other diluents are applicable.

[1085] The first reactor 4704 converts ethane to ethylene in an ODH reaction via the ODH catalyst 4708. Byproducts may include water, carbon monoxide (CO), and carbon dioxide (CO2). A byproduct may be acetic acid (C H4O2) formed by the conversion of ethane or ethylene to acetic acid. For example, ethylene may be oxidized to acetic acid as in the bulked chemical reaction C2H4 + O2 C4H4O2. This bulked chemical reaction may involve C2IT4 + ¾0 -> C2H5OH followed by C2H5OH + 0 ₂ - C4H4O2 + H ₂ 0. The byproduct acetic acid formed in the first reactor 4704 may be a coproduct.

[1086] The effluent 4714 from the first ODH reactor 4704 may include ethylene, CO, CO _? _. water, acetic acid, and unreacted ethane. The effluent 4714 may be labeled as a product effluent of the first reactor 4704 because the effluent 4714 has products, such as ethylene and acetic acid, from the first reactor 4704. The effluent 4714 may be processed to remove oxygen (O2) from the effluent 4714 prior to introduction of the effluent 4714 to the second ODH reactor 4706. A reason to remove 0 ₂ may be to promote the production of ethanol and acetaldehyde in the second ODH reactor 4706. In particular, the removal of O2 from the effluent 4714 may beneficially deprive the second ODH reactor 4706 of O2 and thus avoid oxidation of ethanol (and acetaldehyde) via O2 into acetic acid in the second reactor 4706.

[1087] The oxidative dehydrogenation of ethanol to acetic acid using 0 ₂ (e.g., C ₂ ¾0 + O2 fo C2H4O + H ₂ 0) (as in the first reactor 4704) may have faster kinetics compared to selective oxidation of ethanol and acetaldehyde to acetic acid using CO2 (e.g., C ₂ H ₄ 0 + C0 ₂ □ C2H4O2 + CO) (as in the second reactor 4706). 0 ₂ is generally a stronger oxidizer than C0 ₂ . Therefore, when 0 ₂ is present in the reactor, the 0 ₂ may oxidize ail (or most of) the ethanol and acetaldehyde to acetic acid. Thus, the presence of 0 ₂ may hinder sustained formation of ethanol and acetaldehyde (intermediate between acetic acid and ethanol) by consuming (oxidizing) most of ethanol and acetaldehyde to acetic acid. There is less probability to produce and discharge ethanol and acetaldehyde from the second reactor 4706 if 0 ₂ is present in the second reactor 4706.

[1088] Therefore, the ODH reactor system 4702 includes an 0 ₂ removal system 4716 that receives the effluent 4714 from the first ODH reactor 4704. One implementation of an 0 ₂ removal system 4716 is to remove 0 ₂ by reacting 0 ₂ in the effluent 4714 with CO in the effluent 4714 to form additional CO2 than entering in the effluent 4714. The 0 ₂ removal system may include a vessel or reactor having a selective oxidation catalyst that provides for Site reaction of CO with 0 ₂ to give C0 ₂ . The selective oxidation catal st may be, for example, a silver-cerium (IV) oxide silica (Ag-CeO /SiCh). Other implementations of an (¾ removal system (e.g., a vessel having a separation membrane, etc.) are applicable.

[1089] The (¾ removal system 4716 may discharge a processed effluent 4718 that is the first-reactor effluent 4714 without the (¾ (removed by the Cb removal system 4716). The processed effluent 4718 may include ethylene, CO, CO2, water, acetic acid, and unreacted ethane. The processed effluent 4718 may be sent to the second ODH reactor 4706. At least two reactions may occur in second ODH reactor 4706: (a) hydration of ethylene to ethanol as in [1] C2H4+H2O-» CiHgO; and (b) oxidative dehydrogenation of ethanol to acetaldehyde as in [2] TTHbO + CC¾ C2H4O + H2O + CO. An example of a third reaction may be the oxidation of acetaldehyde to acetic acid as in [3]

( i f ,0 + ( ()· ···> C2H4O2 + CO.

[1090] The effluent 4720 from the second ODH reactor 4706 may include ethane, ethylene. CO, CO2, water, acetic acid, ethanol, and acetaldehyde. The effluent 4720 may be labeled as a product effluent of the second reactor 4706 because the effluent 4720 lias products from the second reactor 4706. The products in the effluent 4720 discharged from the second reactor 4706 may include ethylene, acetic acid, ethanol, or acetaldehyde, or any combination thereof, and additional products.

[1091] The effluent 4720 from the second ODH reactor 4706 may be fed to a liquid product scrubber 4722. Some of the processed effluent 4718 from the O2 removal system 4716 may bypass 4724 the second ODH reactor 4706 (e.g., through a bypass 4724 conduit) and sent to the product scrubber 4722. The percent of the processed first- reactor effluent 4718 sent directly to the scrubber 4722, (bypassing 4724 the seco d reactor 4706) may range from 0% to 100%.

[1092] As discussed, the second ODH reactor 4706 may convert a portion of the ethylene to ethanol, acetaldehyde, and acetic acid. The ethanol, acetaldehyde, and acetic acid may be referred to as oxygenated products. The split ratio of the processed first-reactor effluent 4718 going to the second ODH reactor 4706 versus going directly to the liquid product scrubber 4722 may be adjusted depending, for example, on the desired selectivity towards the oxygenated product versus the desired selectivity towards ethylene. Such may be influenced by market need or other considerations.

[1093] The liquid product scrubber 4722 may be a vessel (e.g , column, tower, etc.) to separate liquid products from gas products. The liquid product scrubber 4722 may be a separation vessel, such as a quench tower, a spray tower, a venture scrubber, or a packed tower, and the like. In certain implementations with the scrubber 4722 as a packed tower, water may be led to the top portion of tower vessel as a shower internal to the tower to scrub off the liquid product mixture in the packed tower. The water can be, for example, demineralized water, process water, water having other components such as acetic acid, and water of varying quality.

[1094] A packed tower (a packed-bed column) may be a vessel (chamber) having a bed of packing. The packing or packing material may be various shapes, such as Raschig rings, spiral rings, or Berl saddles and provide surface area for liquid-gas contact. The packing may be held in place for example, by wire mesh retainers and supported by in a bottom portion of the packed-bed scrubber. In operation, scrubbing liquid may be introduced above the packing and flow's down through the packing bed. In vertical designs (packed towers), a gas stream may flow' up tlie chamber (countercurrent to Site liquid). f 1095] A spray tower (or spray chamber) may be a cylindrical vessel made of steel or plastic and having nozzles that provide for spraying liquid into the vessel. The flow of inlet gas and inlet liquid may be in the opposite direction (countercurrent flow). In certain implementations, several nozzles may be place along the spray tower at different heights for the introduction of liquid spray.

[1096] A liquid product discharge 4726 from the scrubber 4722 may include ethanol, acetaldehyde, acetic acid, and water. A gas product discharge 4728 from the scrubber 4722 may include ethylene, CO, CO2, and ethane. In some implementa tions, the liquid product discharge 4726 is from a bottom portion of the scrubber 4722, and the gas product discharge 4728 is from art upper portion of tire scrubber 4722.

[1097] The liquid product discharge 4726 may be processed in a liquid-product separation system 4730 that separates liquid coproducts. The separation system 4730 may include one or more distillation columns, liquid-to- liquid extraction columns and the like, to separate the liquid products. The arrow 4732 represents multiple discrete product streams from the separation system 4730. The multiple separate product streams may include (1) an ethanol stream, (2) an acetaldehyde stream, and (3) an acetic acid stream. These streams may be labeled as coproducts of die ethylene production in the ODH reactor system 4700.

[1098] As mentioned, gas product discharge 4728 from the liquid product scrubber 4722 may include ethylene, CO, CO _?. , and ethane. The gas product discharge 4728 from the scrubber 4722 may be sent to a gas- product separation train 4734. The separation train 4734 may include, for example, multiple columns (e.g., distillation coiumn(s), extraction columns, scrubber columns, etc.). Two or more of the columns may be disposed in series to form a train of columns. In some examples, the separation train 4734 is a series of at least three distillation columns.

[1099] The arrow 4736 represents multiple discrete streams that discharge from the separation train 4734. The multiple separate streams may include (1) an ethane stream, (2) an ethylene stream, (3) a CO stream, and (4) a CO2 stream. In certain implementations the ethane stream may be recycled as feed to the first ODH reactor 4704, the ethylene stream is product that can be subjected to additional processing or consumed, and the CO2 may be sent to a CO2 capture or conversion plant (unit), and the like.

[1100] FIG. 48 is a method 4800 of coproduction in an ODH system that converts etliane to ethylene. The method involves at least two reactors in series. Each reactor has an ODH catalyst. The two reactors may each be a fixed-bed reactor. A fixed-bed reactor may have a cylindrical tube(s) filled with catalyst pellets as a bed of catalyst. In operation, reactants flow through the bed and are converted into products. The catal st in the reactor may be one large bed, several horizontal beds, several parallel packed tubes, or multiple beds in their own shells, and so on. [1101] In other implementations, the two reactors may each be a fluidized bed reactor. A fluidized bed reactor may be a vessel in which a fluid is passed through a solid granular catalyst (e.g., shaped as spheres or particles) at adequate velocity to suspend the solid catalyst and cause the solid catalyst to behave as though a fluid. In some implementations, the first reactor as a fluidized bed reactor lias a recirculating mode of operation in circulation between reactor and regenerator of the fluidized bed reactor-regenerator system. In contrast for some implementations, the second reactor may generally not employ such recirculation, as the recirculation may promote excessive production of acetic acid. If the second fluidized-bed reactor is equipped with recirculation, the recirculation valves may be fully closed in certain instances to be operational as the second reactor for producing a desired balance of coproducts.

[1102] At step 4802, the method 4800 includes feeding ethane, oxygen, and a diluent to the first reactor having ODH catalyst. The feed may be provided through one or more conduits. The oxy gen may be introduced into the first reactor separate from the ethane. The oxygen may be fed to the reactor by feeding air to the reactor. The diluent may be, for example, carbon dioxide or nitrogen.

[1103] At step 4804, the method 4800 includes contacting ethane with the ODH catalyst in presence of oxygen in the first reactor to dehydrogenate ethane to ethylene and to form carbon dioxide carbon monoxide, water and acetic acid. The acetic acid can be characterized as a byproduct and labeled as a coproduct of the produced ethylene.

[1104] At step 4806, the method 4800 includes discharging a first-reactor effluent (a product effluent of the first reactor) from the first reactor to the second reactor or to a packed column (as described with respect to step 4814). A portion of the first-reactor effluent may be sent to the second reactor and the remaining portion of the first- reactor effluent may be sent to the packed column. The fi rst-reactor effluent may be sent to the packed column to avoid conversion of ethylene in the first-reactor effluent to coproducts in the second reactor.

[1105] The first-reactor effluent includes ethylene, carbon dioxide, carbon monoxide, water, acetic acid, and imreacted ethane. The first-reactor effluent may be processed in route to the second reactor or to the packed column. For instance, the method may remove oxygen from the first-reactor effluent, as discussed below with respect to block 4808. The first-reactor effluent may discharge through an oxygen separation vessel (see block 4808) to the second reactor or to the packed column.

[1106] At step 4808, the method 4800 includes removing oxygen from the first-reactor effluent. In some embodiments, the method 4800 includes removing oxygen from the first-reactor effluent upstream of the second reactor and the packed column. Thus, the first-reactor effluent fed to the second reactor (and to the packed column) may be processed first-reactor effluent having oxygen removed.

[1107] In some implementations, the removal of oxygen from the first-reactor effluent may involve flowing the first-reactor effluent through an oxygen-separation vessel. In some examples, the oxygen is removed via a selective oxidation catalyst in the oxygen-separation vessel as an oxygen-separation reactor. For instance, the removal of oxygen may be by consuming oxygen in the reaction of oxygen with carbon monoxide present to give carbon dioxide. The selective oxidation catalyst may be, for example, a silver-cerium (IV) oxide silica (Ag- CeCVSiOi) or a copper/zinc/zirconium oxide (CuZnZrO _x ), and the like. Other implementations of an O2 removal system (e.g, a vessel having a separation membrane, etc.) are applicable.

[1108] As mentioned, some or all of the processed first-reactor effluent (having O2 removed) may discharge from the 0 ₂ removal s stem to the packed column and thus bypass the second reactor. The percent of the processed first-reactor effluent sent to the packed tower (bypassing the second reactor) may be modulated and range from 0% to 100%.

[1109] At step 4810. the method 4800 includes contacting the processed first-reactor effluent with the ODH catalyst in the second reactor to form ethanol and acetaldehyde. In some instances the method 4800 may include adding carbon dioxide to the second reactor to favor production of acetaldehyde over production of ethanol in the second reactor. Additional acetic acid may also be formed. The ethanol and acetaldehyde (and acetic acid) may be coproducts along with the ethylene produced in the first reactor.

[1110] Without (¾ in the second reactor, there may be no measurable conversion of ethane to ethylene in the second reactor. A trace conversion of ethane to ethylene may occur in tire second reactor, for example, due to any oxy gen present on the catalyst surface.

[Ill 1 J The reaction kinetics in the second reactor may favor production of acetaldehyde over acetic acid or may favor production of acetic acid over acetaldehy de. depending on the reaction conditions. Reaction conditions that may be adjusted included temperature, pressure. GHS V, feed composition, and the like. In the example discussed below, more acetic acid was produced than acetaldehyde in the laboratory ODH reactor performing as a second reactor in a series of two ODH reactors.

[1112] As discussed, reactions that may occur in Site second reactor include: [1] C ₂ H ₄ +¾0-^ € ₂ ¾0 (ethanol); [2] C ₂ H ₆ 0 + C0 ₂ C ₂ 1¾Q (acetaldehyde) + H ₂ 0 + CO; and [3] C ₂ H ₄ 0 + C0 ₂ □ C ₂ H ₄ 0 ₂ (acetic acid) +

CO. In certain implementations, C0 ₂ may be added to the second reactor to favor production of acetaldehyde over production of ethanol in the second reactor. Conversely, in some implementations, some CO? is removed from the first-reactor effluent to favor production of ethanol over acetaldehyde in the second reactor.

[1113] The addition of more C0 ₂ (e.g., increasing concentration of C0 ₂ in the first-reactor effluent feed) may increase the rate of formation of acetaldehyde and acetic acid in the second reactor. The sensitivity of the ( 1) reaction of ethanol/C0 ₂ to acetaldehy de and (2) reaction of acetaldehyde/C0 ₂ to acetic acid in relation to change in feed C0 ₂ concentration may depend on reaction conditions. In one example, the ODH reactor system includes a conduit to add C0 ₂ to the effluent from the first reactor flowing to the second reactor. In another example, a conduit is utilized to add C0 ₂ directly to the second reactor. However, in other examples, no C0 ₂ is added to either the first- reactor effluent or to the second reactor but instead the amount of C0 ₂ fed to the second reactor is the amount of C0 ₂ that discharges in the first-reactor effluent from the first reactor

[1114] As for removal of C0 ₂ from the first-reactor effluent, reactions conditions in the second reactor may determine if less C0 ₂ in the feed to the second reactor would hinder formation of acetaldehyde (and acetic acid) in the second reactor. A C0 ₂ removal unit (if employed) may be disposed between the first reactor and the second reactor to remove C02 from the first-reactor effluent flowing to the second reactor. The C0 ₂ removal unit may be, for example, an amine tower or caustic tower. In some implementations, liquid in the first-reactor effluent may be removed before subjecting the first-reactor effluent to the C0 ₂ removal unit.

[1115] At step 4812, die mediod 4800 include discharging a second-reactor effluent (a product effluent of the second reactor) from the second reactor. For example, the second-reactor effluent may be discharged to a separation column, such as a column having packing to separate liquid from gas. The second-reac or effluent includes ethylene ethanol, and acetaldehyde as coproducts. The second-reactor effluent may also typically include acetic acid as a coproduct. Further, the second-reactor effluent may include carbon dioxide, carbon monoxide, water, unreacted ethane, and so on. [1116] At step 4814 the method 4800 includes separating the second-reactor effluent into a liquid stream and a gas stream. The liquid stream may include ethanol, acetaldehyde, acetic acid, and water. The gas stream may include ethylene, ethane, carbon dioxide, and carbon monoxide. In implementations, the separating may involve separating the second-reactor effluent into the gas stream and the liquid stream in a packed column, discharging the gas stream overhead from the packed column, and discharging the liquid stream from a bottom portion of the packed column. The packed column may be labeled, for example, as a scrubber or liquid scrubber. The packed column may ^¬ be the packed-tower embodiment of the liquid product scrubber 4722 depicted in FIG. 47.

[1117| The method 4800 may include maintaining an operating temperature of the first reactor and an operating temperature of the second reactor at less than 370 °C, less than 360 °C, or less than 350 °C. As for operating pressure the reactor inlet pressure for each reactor may be less than 80 pound per square inch gauge (psig), or less titan 70 psig. The reactor inlet pressure for each reactor may in the range of 1 psig to 80 psig, or in the range of 5 psig to 75 psig. Other operating conditions of the reactor in the embodiment of the reactor as a tubular fixed-bed reactor may be gas hourly space velocity (GHSV) in the range of 200 hour ¹ to 10,000 hour ¹ and a linear velocity range of the feed through the reactor of at least 5 centimeters per second (cm/sec). The ODH catalyst in the first reactor and the ODH catalyst in the second reactor may each be a mixed metal oxide having formula Mo _a V _b Te _c Nb _d Pd _e O _f , where a, b, c, d, e. and f subscripts are relative atomic amounts of elements Mo, V, Te, Nb, Pd, O, respectively, and whena=l, b=0.01 to 1.0, c=0.01 to 1.0, d=0.01 to 1.0 0.00<e<0.10, and f is a number to satisfy valence state of the catalyst.

[1118] Example

[1119] This example is given only as an example and not meant to limit the present techniques. Ethanol and acetaldehyde were synthesized over ODH catalyst from feed resembling a first-reactor effluent (e.g., 4718 discussed above) from a first ODH reactor (e.g., 4704). A laboratory ODH reactor simulating an implementation of the second ODH reactor 4706 received the feed and was utilized for the synthesis that produced ethanol and acetaldehyde.

[1120] For simplification, CO and acetic add were not included in the feed. The liquid in the feed was water as noted in Table II below. The gas in the feed was ethane, ethylene, O2, and CO2, as given in Table J2 below

[11211 Table J 1 gives the liquid product from the synthesis in the laboratory ODH reactor receiving the feed of water, ethane, ethylene, 0 _2, and C0 ₂ listed collectively in Tables II and 2. Table 12 gives the gas product from the synthesis in the laboratory ODH reactor. Again, the laboratory ODH reactor acted as a second ODH in a series of two ODH reactors.

[1122] FIG. 49 is the ODH reactor system 4900 having the ODH reactor 4902 employed to perform the example synthesis over ODH catalyst that produced the ethanol and acetaldehyde from the feed resembling a first- reactor effluent from a first ODH reactor. The ODH reactor 4902 is a large-scale laboratory reactor that is a continuous tubular fixed-bed reactor. As indicated, the ODH reactor 4902 simulated a second ODH reactor of two ODH reactors disposed in series. The ODH reactor 4902 may be characterized as an implementation of the second ODH reactor 4706 previously described.

[1123] The ODH reactor 4902 is a tubular fixed-bed reactor having two tubes disposed in series and with each tube having a fixed bed of catalyst. The two tubes are constructed of Type 316L stainless steel. The two tubes can each be characterized as two respective tubular fixed-bed reactors but the two tubes were operated collectively in the Example as a single tubular fixed-bed reactor.

[1124] The two tubes each have a heat transfer jacket that receives circulating oil from a closed-loop oil bath for heating or cooling of the two tubes to maintain a desired temperature in the reactor. In the Example, the reactor temperature was maintained in the range of 309 °C to 312 °C. The temperature of the two tubes was monitored with thermocouples as temperature sensors.

[1125] The GDH catalyst in the two fixed beds is Mo1.0V0.37Te0.23Nb0.14Od-4.97 with the numerical subscripts indicating molar ratios of the molybdenum vanadium, tellurium, niobium, and ox gen, as determined by inductively coupled plasma mass spectrometiy (ICP-MS). The ODH catalyst is available from NOVA Chemicals Corporation having headquarters in Calgary' . Canada. The inside diameter of each tube is 2.1 centimeters (cm) giving a catalyst- bed diameter of 2.1 cm. The catalyst-bed height in the first tube is 128.5 cm. The catalyst bed height in the second tube is 135 cm. The two catalyst beds are diluted with a diluent powder. The mass ratio of diluent powder to catalyst is 1.22. The diluent powder is Versa!™ Alumina V-25G manufactured by Honeywell UOP having headquarters in Des Plaines. Illinois, United States of America The mass of catalyst in the two beds combined is 171 grams. The mass of the diluent in the two beds combined is 209 grams. In preparation for addition to the reactor beds, the catalyst and diluent powder (alumina) are mixed together as powders and then extruded together into a cylindrical shape. The cylindrical extrudate shape has a diameter of about 1.7 millimeter (mm) and a length in the range of 2 mm to 10 mm.

[1126] A combined gas feed 4904 of ethane, ethylene, and C(¾ was fed to the inlet of the ODH reactor 4902 from respective gas cylinders of ethane, ethylene, and CO2. The available pressure of the gas cylinders provided motive force for flow' of the combined gas feed 4904 into the ODH reactor 4902 A respective flow valve associated with each gas cylinder gave the desired flow' rate of each gas component. Water 4906 as the liquid feed was introduced into the gas feed 4904 flowing to the reactor 4902. A pump provided the motive force for flow' of the water 4906 into the reactor 4902.

[1127] The inlet pressure at the reactor 4902 was 19 psig due to the hydraulic backpressure generated by flow of the feed gas through the reactor beds downstream condenser, and associated piping. The linear velocity' of the feed through the reactor 4902 was 132 cin/sec. The GHSV was 717 hour ¹ (h ^s ). The GHSV is the ratio of the volumetric flow' rate of the gas feed 4904 plus vapor feed generated from evaporation of liquid feed 4906 at the inlet of reactor at standard conditions for temperature (0 °C) and pressure (101 kilopascals) to the combined volume of the two fixed-beds of catalyst. The weight hourly space velocity (WHSV) was 0.7 Ir ¹ . The WHSV is the ratio of the weight (mass) flow rate of the total feed (gas feed 4904 plus w'ater 4906) to the combined weight (mass) of the two fixed-beds of catalyst. The diluent powder is not counted (included) in the basis for the volume or weight of the catal st bed. The weight and volume of the ODH active phase catalyst (mixed metal oxide) is the basis in the GHSV and WHSV calculations.

[1128] The product effluent 4908 discharged from the reactor 4902 to a condenser 4910. The cooling medium in tiie condenser 4910 was distilled water. The condenser 4910 is a shell-and-tube heat exchanger with product gas on the tube side and the distilled water on die shell side. A product gas stream 4912 discharged from the condenser 4910 to a vent system 4914. A sample syringe was utilized to collect a gas sample 4918 of the product gas stream at a sample point downstream of the condenser 4910. A liquid product stream 4920 discharged from the condenser 4910 to a liquid collection system 4922. A liquid sample 4924 of the liquid product stream 4920 was obtained. [1129] Tables J1 and 12 give the feed to the ODH reactor 4902 as an implementation of the second reactor 4706 (see FIG. 47). Again, the ODH reactor 4902 acted as a second ODH in a series of two ODH reactors. Table II gives the liquid feed (water 4906). Table 12 gives the gas feed (gas 4904). The total feed is the combination of the liquid feed (Table Jl) and the gas feed (Table 48). The total feed is a first-reactor effluent-type feed without CO and acetic acid. Table 11 gives the liquid product from die synthesis in the ODH reactor 4902 receiving die total feed given collectively in Tables H and 12. Table 12 gives the gas product from die synthesis in the ODH reactor 4902.

[1130] Table Jl. Liquid Feed Composition and Liquid Product Composition

[1131] Table 12. Diy Gas Feed Composition and Gas Product Composition

‘The wet feed gas composition (mol %) was C2H6/C2H4/CO2/H2O = 10.0/69.3/7.3/13.4

[1132] As can be seen in Tables Jl and J2, the product (effluent) from the ODH reactor 4902 acting as a second reactor included acetaldehyde ethanol acetic acid, and ethylene. In implementations, the amount of coproducts, such as acetaldehyde and ethanol, produced may be increased by recirculating a portion of the product effluent back through the reactor.

[1133] An embodiment is a method of coproduction in the production of ethylene. The method includes contacting ethane with an ODH catalyst in presence of oxygen in a first reactor to dehydrogenate ethane to ethylene and to form carbon dioxide and water. The contacting of the ethane with the ODH catalyst in the presence of oxygen in the first reactor may also form carbon monoxide and acetic acid. The method includes discharging a first-reactor effluent from the first reactor to a second reactor. The first-reactor effluent includes ethylene, carbon dioxide, and water. The first-reactor effluent may also include carbon monoxide acetic acid, and unreacted ethane. The method may include removing oxygen from the first-reactor effluent upstream of the second reactor. The first-reactor effluent is contacted with an ODH catalyst in the second reactor to fonn ethanol and acetaldehyde. A second-reactor effluent is discharged from the second reactor. The second-reactor effluent includes ethylene, ethanol, and acetaldehyde. The ethylene, ethanol, and acetaldehyde may be coproducts. The second-reactor effluent may also include carbon dioxide, carbon monoxide, water, and acetic acid. Acetic acid may be a coproduct. The method may include maintaining an operating temperature of the first reactor and an operating temperature of the second reactor at less than 370 °C. The method may include adding carbon dioxide to the second reactor to favor production of acetaldehyde over production of ethanol.

[11341 The method may include separating the second-reactor effluent into a liquid stream and a gas stream wherein tire liquid stream includes ethanol, acetaldeh de, acetic acid, and water, and wherein the gas stream includes ethylene, ethane, carbon dioxide and carbon monoxide. The separating may involve separating the second- reactor effluent into the gas stream and the liquid stream in a packed column, discharging the gas stream overhead from the packed column, and discharging the liquid stream from a bottom portion of the packed column. Lastly, the ODH catal st in the first reactor and tire ODH catalyst in the second reactor may each be a mixed metal oxide having formula Mo _a V _b Te _c Nb _d Pd _e O _f , where a, b. c, d, e, and f subscripts are relative atomic amounts of dements Mo, V. Te Nb Pd, O, respectively, and whena=l, b=0.01 to TO, c=0.01 to TO, d=0.01 to 1.0,

0.00<e<0.10, and f is a number to satisfy the valence state of the catalyst. The variable f may be a number to satisfy at least the valence state of the corresponding elements in the catalyst.

[1135] Another embodiment is a method of coproduction in a sy stem that converts ethane to ethylene. The method includes: feeding ethane, oxygen, and a diluent to a first reactor; contacting ethane with an ODH catalyst in the first reactor to produce ethylene, acetic acid, carbon dioxide, and water; discharging from the first reactor a first- reactor effluent haying ethylene, acetic acid, carbon dioxide, and water; removing oxygen from the first-reactor effluent; contacting the first-reactor effluent with an ODH catalyst in a second reactor to produce ethanol and acetaldehyde; and discharging from the second reactor a second-reactor effluent having ethylene, ethanol, acetaldehyde, acetic acid, carbon dioxide, and water. The contacting of the ethane with the ODH catalyst in the first reactor may produce carbon monoxide and therefore, the first-reactor effluent may Slave carbon monoxide. The second-reactor effluent may also Slave carbon monoxide. The contacting of the first-reactor effluent with the ODH catalyst in the second reactor may produce acetic acid. An operating temperature of the first reactor and an operating temperature of the second reactor are less than 370 °C. In certain implementations, the first reactor and the second reactor each are a fixed-bed reactor having the ODH catalyst in a fixed bed.

[1136] Yet another embodiment is a method of coproduction in the conversion of ethane to ethylene. The method includes contacting ethane with an ODH catalyst in presence of oxy gen at a temperature of less than 370 °C in a first ODH reactor to generate ethylene, acetic acid, carbon dioxide, and water. The method includes discharging from the first ODH reactor a first-reactor effluent having ethylene, acetic acid, carbon dioxide, water and unreacted ethane through an oxygen-separation vessel to a second ODH reactor. The method may include removing oxygen from the first-reactor effluent flowing through the oxygen-separation vessel via a selective oxidation catalyst in the oxygen-separation vessel. The method includes contacting die first-reactor effluent with an ODH catalyst at a temperature of less dan 370 ^C C in the second ODH reactor to generate ethanol and acetaldehyde. The method includes discharging from the second ODH reactor a second-reactor effluent having ethylene, ethanol, acetaldehyde, acetic acid, carbon dioxide, and water to a separation column. Ethylene, ethanol, acetaldehyde, and acetic acid in the second-reactor effluent may be coproducts. The first ODH reactor and the second ODH reactor may each be a fluidized bed reactor having the first ODH catalyst and the second ODH catalyst, respectively, in operation as a fluidized bed of catalyst. In other implementations, the first ODH reactor and the second ODH reactor are each a tubular fixed-bed reactor.

[1137] Yet another embodiment is a system to produce ethylene and coproducts. The system has a first reactor with a first ODH catalyst to dehydrogenate ethane into ethylene and form carbon dioxide, carbon monoxide, and water. The first reactor may be a fixed-bed reactor having the first ODH catalyst as a fixed bed of catalyst. The first reactor may have a cooling jacket to maintain mi operating temperature of the first reactor less than 370 °C. A first- reactor effluent from the first reactor includes ethylene, unreacted ethane, carbon dioxide, carbon monoxide, and water. The first reactor may also form acetic acid and, therefore, the first-reactor effluent may include acetic acid. The first reactor may form acetic acid by oxidation of ethylene and ethanol in presence of oxygen. The system includes a second reactor having a second ODH catalyst to form ethanol and acetaldehyde from the first-reactor effluent. An oxygen-separation vessel may be operationally disposed between the first reactor and the second reactor to remove oxygen from the first-reactor effluent. In some implementations, the oxygen-separation vessel may have a selective oxidation catalyst to facilitate removal of oxygen from the first-reactor effluent. The second reactor may be a fixed-bed reactor having the second ODH catalyst as a fixed bed of catalyst. The second reactor may have a heat- transfer (e.g., cooling) jacket to maintain an operating temperature of the second reactor less than 370 °C. The second reactor discharges a second-reactor effluent having ethylene, ethanol, and acetaldehyde as coproducts. The second reactor may form acetic acid from the first-reactor effluent. The second reactor may form acetic acid by oxidation of ethanol and acetaldehyde in presence of carbon dioxide. The second-reactor effluent may include acetic add, carbon dioxide, water, and carbon monoxide. Acetic acid may be a coproduct. ft 138] The system may include a separation column to separate the second-reactor effluent into a gas stream and a liquid stream, wherein the gas stream includes ethylene, ethane, carbon monoxide, and carbon dioxide, and wherein the liquid stream includes ethanol, acetaldehyde, acetic acid, and water. In certain implementations the separation column includes: packing to separate the second-reactor effluent into the gas stream and the liquid stream; an overhead outlet on an upper portion of the separation column to discharge the gas stream; and a botoms outlet on a lower portion of the separation column to discharge the liquid stream. Lastly, the first ODH catalyst and the second ODH catalyst may each include a mixed metal oxide having formula Mo _a V _b Te _c Nb _d Pd _e Or, where a, b, c, d, e, and f subscripts are relative atomic amounts of elements Mo, V, Te, Nb, Pd, O, respectively, and when a=l, b=0.01 to 1.0, c=0 to 1.0, d=0 to 1.0, 0.00<e<0.10, and f is a number to satisfy valence state of the catalyst.

[1139] Yet another embodiment is a system to produce ethylene and coproducts. The system includes a first ODH reactor (e.g., a tubular fixed-bed reactor, a fluidized bed reactor, etc.) to receive ethane and oxygen. The first ODH reactor may also receive hydrogen. In some instances, the first ODH reactor receives diluent, such as nitrogen or carbon dioxide or other diluent. The first ODH reactor has a first ODH catalyst to convert ethane into ethylene, generate carbon dioxide and water and convert ethylene into acetic acid. The first ODH reactor may also form carbon monoxide. The system includes an oxygen-separation vessel to receive a first-reactor effluent from the first ODH reactor and remove oxy gen from the first-reactor effluent. The system includes a second ODH reactor (e.g., a tubular fixed-bed reactor, a fluidized bed reactor, ebullated bed reactor, etc.) to receive the first-reactor effluent from the oxygen-separation vessel. The second reactor has a second ODH catalyst to generate ethanol and acetaldehyde from the first-reactor effluent at an operating tempera ture of less than 450 °C and discharge a second-reactor effluent having ethylene, ethanol, acetaldehyde, and acetic acid as coproducts. The second-reactor effluent may include carbon dioxide, carbon monoxide, ethane, and water. The system may include a packed column to receive the second-reactor effluent. In some implementations, the packed column discharges an overhead stream having ethylene, ethane, carbon monoxide, and carbon dioxide, and discharges a bottoms stream having ethanol, acetaldeh de, acetic acid, and water.

[1140] In some embodiments, the ethylene produced using die systems and methods for coproduction in the production of ethylene described herein, or any of the processes or complexes described herein, can be used to make various olefin derivatives. Olefin derivatives include, but are not limited to, polyethylene, polypropylene, ethylene oxide, propylene oxide, polyethylene oxide, polypropylene oxide, vinyl acetate, vinyl chloride, acrylic esters (e.g., methyl methacrylate), thermoplastic elastomers, thermoplastic olefins, and blends and combinations thereof. The polyethylene can be produced using any suitable polymerization process and equipment. Suitable ethylene polymerization processes include, but are not limited to gas phase polyethylene processes, high pressure polyethylene processes, low pressure polyethylene processes, solution polyethylene processes slurry polyethylene processes and suitable combinations of the above arranged either in parallel or in series. Further, the polyethylene can include homopoly mers of ethy lene or copoly mers of ethylene and a-olefins, resulting in high density polyethylene (HDPE), medium density polyethylene (MDPE), low density polyethylene (LDPE), linear low density polyethylene (LLDPE), and very low density polyethylene (VLDPE).

[1141] Techniques described here provide a one pass process to oxidatively dehydrogenate lower paraffins (alkanes, for example, n-alkanes) to produce alpha olefins. The techniques include the oxidative dehydrogenation of one or more alkanes selected from the group consisting of ethane and propane and mixtures thereof in the presence of a supported catalyst. The process indudes(a) passing through an oxidative dehydrogenation reactor containing a fluidized bed of the catalyst the one or more alkanes and oxygen at a temperature from 250 °C to 370 °C, a pressure from 3.447 to 689.47 kPag (0.5 to 100 psig) or from 103.4 to 344.73 kPag (15 to 50 psig) and a residence trine of the one or more alkanes in the reactor from 0.002 to 10 seconds, and reducing the catalyst, the catalyst having an average residence time in the dehydrogenation reactor of less than 30 seconds. The process includes (b) feeding the reduced catalyst to a regeneration reactor and passing a stream of air optionally with additional nitrogen at a temperature from 250 °C to 400 °C and pressures from 3.447 to 689.47 kPag (0.5 to 100 psig) [for example, from 103.4 to 344.73 kPag (15 to 50 psig)] through the bed to oxidize the catalyst. The process includes (c) passing the oxidized catal st back to the oxidative dehydrogenation reactor wherein the amount of oxygen in the feed to the reactor is below the upper flammability limit for Hie feed. The conversion of alkane to alkene is not less than 50 mol % per pass and the selectivity for the conversion of alkane to alkene is not less than 0.9 on molar basis. [1142] In some embodiments, the process comprises passing the product stream through one or more oxygen scavenging reactors. Reactors can be operated in parallel, where one is being oxidized and another is being reduced to lower oxidation state of the metals in die catalyst.

[1143] In some embodiments, oxygen scavenging reactors use the same catalyst used in oxidative dehydrogenation reactors.

[1144] in some embodiments, the oxidative dehydrogenation reactor comprises a riser and the regeneration reactor is a separate fluidized bed reactor, the regeneration reactor being connected with the riser to flow' oxidized catalyst back to the riser (e.g. CFB [circulating fluidized bed] type reactor).

[1145] In some embodiments, the top of the riser comprises a distributor system to improve temperature control in the reactor [to minimize combustion of the alkane feed and] to maintain the overall selectivity of the reactor above 90 mol %.

[1146] In some embodiments, one or more of low temperature steam and atomized water is/are passed into the catalyst flow into Site riser to cool the catalyst to control die heat balance of the oxidative dehydrogenation reactor. [1147] In some embodiments, there is a downcomer between the oxidative dehydrogenation reactor and the regeneration reactor to flow reduced catalyst from the oxidative dehydrogenation reactor to the regeneration reactor. [1148] In some embodiments, low temperature steam [counter current to the flow of catalyst through the downcomer] is passed to strip entrained alkane feed and product.

[1149] In some embodiments, air or a mixture of air and nitrogen is passed through the regeneration reactor in an amount to substantially extract the oxygen from the air or a mixture of air and nitrogen and generating a gas product stream comprising not less than 85 vol.-% of nitrogen.

[1150] In some embodiments, a portion of the oxygen reduced effluent stream is collected from the regenerator reactor and optionally cooled and recycled to the regenerator reactor.

[1151] In some embodiments, a CO promoter is added to the regenerator reactor.

[ 152] In some embodiments, the alkene product is separated from the oxidative dehydrogenation reactor from water and oxygenates iu the product stream from the oxidative dehydrogenation unit.

[1153] In some embodiments, unused nitrogen is passed from the effluent stream from the catalyst regenera tion reactor to a site integrated unit operation using nitrogen as a part of the feedstock.

[1154] In some embodiments, two or more fixed bed reactors are used as oxygen scavengers having piping and valves so that the feed to the fluidized bed oxidative dehydrogenation reactor passes through one or more of the fixed bed reactors having an oxidative dehydrogenation catalyst rich in surface oxygen which the effluent is oxidized to depleted the catalyst of oxygen, and passing the product stream through one or more of the fixed bed reactors having an oxidative dehy drogenation catalyst depleted of surface oxygen, to remove residual oxygen from the product by reaction and switching the flow of product stream to reactors to oxygen depleted reactors and the flow' of feed stream to oxygen rich reactors.

[1155] In some embodiments, the site integrated unit operation is selected from an ammonia plant and an acrylonitrile plant, urea plant and, an ammonium nitrate plant. [1156] In some embodiments, the residence time of the catalyst in the oxidative dehydrogenation reactor is less than 30 seconds (for example, less than 10 seconds or less than 5 seconds).

[1157] In some embodiments, the residence time of the catalyst in the regeneration reactor is less than 3 minutes.

[1158] in some embodiments, the ratio of residence time of the catal st in the regenerator to the residence time of the catalyst in the oxidative dehydrogenation catalyst is not less than 3.

[1159] in some embodiments, the product stream from the oxidative dehydrogenation reactor and at least a portion of the effluent stream from the regenerator reactor are passed through separate steam generators to recover heat.

[1160] In some embodiments, the product stream from the oxidative dehydrogenation reactor is cooled and passed through a coluimi to separate combustion products from alkene.

[1161] In some embodiments, the product stream from the oxidative dehydrogenation reactor is cooled and passed through art amine unit to remove CCb.

[1162] In some embodiments, the support is selected from the group consisting of silicon dioxide, fused silicon dioxide, aluminum oxide, titanium dioxide, zirconium dioxide, thorium dioxide, lanthanum oxide, magnesium oxide, calcium oxide, barium oxide, tin oxide, cerium dioxide, zinc oxide, boron oxide, yttrium oxide. [1163] In some embodiments, the alkane is ethane.

[1164] In some embodiments, the conversion to ethylene in the oxidative dehydrogenation reactor is greater than 60 mol %.

[1165] In some embodiments, in the oxidative dehydrogenation reactor the selectivity to ethylene is greater than 75 mol %.

[1166] The feed to the oxidative dehydrogenation reactor includes oxygen in an amount below the upper explosive/ignition limit. For example for ethane oxidative dehydrogenation, typically the oxygen will be present in an amount of not less than about 5 mole-%, in some cases not less than about 18 mole-%, for example from about 22 to 27 mole-%, or 23 to 26 mole-%. It is desirable not to have too great an excess of oxygen as this may reduce selectivity arising from combustion of feed or final products. Additionally too high an excess of oxygen in the feed stream may require additional separation steps at the downstream end of the reaction.

[1167] The process will be described in conjunction with Figure 50 which schematically illustrates a circulating fluidized bed reactor.

[1168] In some embodiments, the reactor system 5000 comprises a fluidized bed oxidative dehydrogenation reactor 5003 and a regenerator reactor 5004. The fluidized bed oxidative dehydrogenation reactor riser 5002 and the regeneration reactor 5004 are joined by a downcomer 5010 which conducts clean oxidized supported catalyst from the regenerator reactor 5004 to the oxidative dehydrogenation reactor riser 5002. Each of the fluidized bed oxidative dehydrogenation reactor riser 5002, the fluidized bed 5006 in the deh drogenation reactor 5003 and the regeneration reactor 5004 contain fluidized bed of catalyst particles 5005, 5006, and 5007 respectively. In the oxidative dehydrogenation reactor riser 5002 and the regeneration reactor 5004 above fluidized catalyst beds 5005 and 5007 are disengagement zones 5008 and 5009, respectively. [1169] The inlet 50 i 1 to downcomer 50 iO is attached to the regenerator reactor 5004 generally at a point between about 1/3 to 2/3 the height of the fluidized bed 5007. The downcomer 5010 enters the bottom of the oxidative dehydrogenation reactor riser 5002. The reactor riser 5002 extends up into the dehydrogenation reactor 5003 above the fluid level of the fluidized bed 5006 (typically 1/3 to 2/3 of height). The reactor riser 5002 flares to form an inverted cone disperser 5012 to provide a disengagement zone for catalyst from the product. Optionally, a disperser plate 5013 may be used above the cone 5012. The disperser may have a shape other than an inverted cone; however, care must be taken to ensure a substantially uniform gas flow around the disperser.

[1170| The oxidative dehydrogenation reactor operates at temperatures below 370 °C typically from 350 °C to 450 °C. pressures from 3.447 to 689.47 kPag (0.5 to 100 psig) or from 103.4 to 344.73 kPag (15 to 50 psig) and a residence time of the one or more alkanes in the oxidative deh drogenation reactor riser 5002 from 0.002 to 20 seconds.

[1171] Flared section 5012 of tire riser should be sufficiently broad to cause the catal st particles to drop in the fluidized bed zone 5006, the disperser plate 5013 should be high enough to minimize catalyst attrition.

[1172] Port 5014 in the riser 5010 permits the introduction of one or more of low temperature steam and atomize water (e.g. a mist), at a temperature at least about 25 °C desirably 50 °C lower than the temperature of the oxidative dehydrogenation reactor. In some embodiments steam has a temperature from about 200 °C to about 400 °C, in further embodiments the temperature may be from about 300 °C to 350 °C. The steam cools the catalyst coming from the regenerator reactor 5004 and also removes any entrained or absorbed impurities (e.g. ethylene or air). The atomized water may have temperature from 50 °C to 75 °C on introduction to port 5014. All or part of the atomized water can be recycled water from the reaction product.

[1173] Port 5015 at or towards the base of the oxidative dehydrogenation reactor riser 5002 is an inlet for the hydrocarbon feed typically high purity- ethane mixed with oxygen or an oxygen containing gas. The hydrocarbon feed and oxygen could be combined proximate and upstream of the oxidative dehydrogenation reactor. As this is a fluidized bed reactor it is necessary' that the upward flow of hydrocarbon feed and oxygen containing gas be sufficiently well distributed to fluidize the bed of catalyst particles to minimize hot spots.

[1174] The process may be used to generate ethylene from relatively pure feedstock.

[1175] The ethane individually should comprise about 95 w4 % of ethane (for example, 98 wt. % of ethane) and not more than about 5 wt. % of associated hydrocarbons such as methane. The feed can be oxygen having a relatively high purity, in some embodiments above 90% purity, in further embodiments greater titan 95% purity. While air may be used as a source for oxygen it may give rise to downstream separation issues.

[1176] In some embodiments, the reactor may be used to replace an ethane / ethylene splitter or off-gas from refinery or other hydrocarbon processing process in which case the feedstock can comprise from 10 - 80 voi. % ethylene and balance ethane.

[1177] To maintain a viable fluidized bed. the mass gas flow' rate through the bed must be above the minimum flow' required for fluidization, from about 1.5 to about 10 times L and in some cases, from about 2 to about 6 times U _mf . U _f is used in the accepted form as the abbreviation for the minimum mass gas flow required to achieve fluidization. Typically the superficial gas velocity required ranges from 0.3 to 5 m/s. [1178] At the upper end of the oxidative dehydrogenation reactor, below disengagement zone 5009 is post 5016 which permits the spent cataly si stream to settle and leave the reactor. At the top of the reactor 5003 , there are cyclones 5017 to remove any catalyst fines, which were not settled in disengagement zone 5009.

[1179] The average residence time of the supported catalyst in the oxidative dehydrogenation reactor riser 5002 is less titan about 30 seconds in some cases less than 15 seconds in some cases from 1 to 6 seconds. The port 5016 connects downcomer 5018 with the oxidative dehydrogenation reactor 5003 and the regeneration reactor 5004. Port 5019 in the downcomer 5018 is positioned proximate the regeneration reactor 5004. Port 5019 allows the introduction of steam at a temperature from about 300 °C to 370 °C, in some embodiments from 350 °C to 360 °C to flow' counter current to the stream of spent catalyst to remove entrained feedstock and product. In some cases the steam may also burn of surface coke on the catal st particles. The flow rate of the steam in the downcomer should be sufficiently low' to prevent the supported catalyst from being pushed back into disengagement zone 5009.

[1180] The regeneration reactor is also a fluidized bed reactor. Port 5020 at the bottom of Site regeneration reactor permits air and in some cases rec cled cooled nitrogen back into Site reactor. The regeneration reactor is typically operated at temperatures from 250 °C to 370 °C and pressures from 3.447 to 689.47 kPag (0.5 to 100 psig or from 103.4 to 344.73 kPag (15 to 50 psig). The residence time of the supported catalyst in the regeneration reactor is less than 3 minutes. Typically the ratio of the residence time of the catalyst in the regenerator reactor to the residence time in the oxidative dehy drogenation reactor is not less than 3.

[1181] Port 5021 on the upper portion of the regenerator reactor 5004, above the fluidized bed of supported catalyst particles permits the off gas to leave the reactor. There may be cyclones as described for the oxidative dehy drogenation reactor 5003 in the upper section of the regenerator reactor 5004 to remove any cataly st fines from nitrogen product stream. Air and optionally nitrogen which may be cooled are passed through the regeneration reactor. The oxygen is substantially taken from the air. The off gas will comprise from 85 vol. % to 100 vol. % of nitrogen.

[1182] The above description of the circulating fluidized bed reactor has been largely schematic. There may be various valves, filters, etc.at the ports. The selection of appropriate valves would be well known to those skilled in the art. Similarly there may be suitable fans and compression means used to force gases through the system. The selection of appropriate fans or compressors or expanders for cooling would be known to one of ordinary skill in the art.

[1183] It is desirable to recover as much energy as possible from the oxidative deh drogenation reaction and the regeneration reaction. The ethylene product and the co-products (e.g. COi and CO) from the oxidative reactor are fed to separate steam generators to generate steam. Part of the steam may be recycled back to the process. The steam could be injected in the riser to cool the catal st particles. The steam could also be injected into the downcomer to bum off any coke and to entrain any absorbed or adsorbed feed or products.

[1184] The oxygen containing stream passing through Site regenerator is substantially depleted of oxygen on exit front the reactor (e.g. the exit stream comprises not less than about 90 vol. % of nitrogen). If nitrogen is also used as a component of the feed stream a part of the product stream may be rec cled to the inlet for the regeneration reactor. The portion of the product stream from the regeneration reactor may be subject to one or more cooling or refrigeration steps to maintain the heat balance in the regenerator. In some embodiments a CO promoter may be added to the regenerator to minimize the heat release in regenerator (i.e. reduce/control C0 ₂ production).

[1185] The process should be operated to have a conversion of not less than 50 mol % (to ethylene) and a selectivity of not less than 90 mol %, and in some cases, greater than 95 mol % to ethylene.

Separation of Product Streams:

[1186] The stream 5022 exiting the dehydrogenation reaction comprises ethylene, water (vapor - steam) and a small amounts of ethane, unconsumed oxygen and off gases typically CO and COi. The issue of separation needs to be considered in the context of the intended use for the ethylene.

[1187] There are a number of processes which may use dilute ethylene such as polymerization processes. However, tliis approach needs to be balanced with the effect of polar molecules such as CO and CO2 and oxygen on the catal st used for tire polymerization. It may be preferable to separate the polar molecules prior to separation of ethylene and ethane. The polar molecules may be separated by an adsorption bed such as a zeolite bed. In the simplest embodiment, depending on the ratios of the components tire bed could be regenerated mid all the components fed to a burner to burn the CO. However, at a chemical complex there are other unit operations which could use CO as a feed (various carboxylic acid and anhydride processes (acetic acid, meth acrylic acid and maleic anhydride). If there is a significant amount of CO and CO2 the components could be separated. There are a number of well-known methods in the art to separate CO2 and CO. The stream would be cooled and washed and then passed through an adsorber such as activated carbon (to remove impurities from the CO?) or a liquid amine separator, a liquid carbonate separator, or a caustic tower to absorb the CO2. CO could be separated by a number of techniques. Depending on the volume a vacuum separation method using activated carbon as an adsorbent may be suitable a membrane separation may be suitable and adsorption on copper ions (on a suitable support) may be suitable.

[1188] Oxygen Removal - Fixed Bed:

[1189] In some embodiments, there may be two or more fixed bed reactors having an oxidative dehydrogenation catalyst which releases or takes up oxygen are used as scavengers to accommodate the product flow' out of the circulating fluidized bed oxidative dehydrogenation reactor. The fixed bed reactors have piping and valves so that the feed to the fluidized bed oxidative dehydrogenation reactor passes through one or more of the fixed bed reactors having a catalyst containing oxygen which is consumed or given up. Tliis is not so much of an issue with the pre-reactor operating in oxidative dehydrogenation mode since any excess alkane not dehydrogenated in the pre-reactor will be converted in the fluidized bed oxidative dehydrogenation reactor. The key issue is depleting the catalyst in the fixed bed reactor of oxygen. The piping and valves flow the product stream through one or more of the fixed bed reactors having oxidative dehydrogenation catalysts which are depleted of oxygen. The depleted fixed bed catalyst scavenges oxygen from the product stream. As noted the valves and piping of the streams can be operated so that feed streams flow through the oxygenated fixed bed catalyst reactor and the product stream flow's through one or more of the oxygen depleted fixed beds catalyst reactors.

[1190] The oxidative dehydrogenation catalyst containing oxygen may have ox gen as lattice oxygen, adsorbed oxygen or adsorbed oxygen on the cataly st, the support or both. The oxidative dehydrogenation cataly st depleted of oxygen l as a reduced oxygen content, for example, about 60% less oxygen in the catalyst and suppost as lattice oxygen, adsorbed oxygen or adsorbed oxygen on the catalyst, the support or both.

[1191] In some embodiments, an oxygen sensor is included at the exit of the fluidized bed oxidative deh drogenation reactor. Additionally, there should be an oxygen sensor at the exit for the dehydrogenated product from each fixed bed reactor to determine the oxy gen level leaving the product leaving that fixed bed reactor. When the oxygen level rises at the dehydrogenated product outlet of the fixed bed reactor operating in scavenger mode it indicates the catalyst have substantially taken up reactive oxygen (and may be returned to use as a pre-reactor). The amount of reactive oxygen uptake by the oxygen depleted catal st in the pre-reactor operation in oxygen scavenging or chemisorption mode should be not less than about 1.5%, typically about 2% of the total oxygen in tire catal st (this will also correspond to the amount of reactive ox gen available for release from die catalyst in the pre-reactors in oxidative deh drogenation mode).

[1192] One mode for operation using three pre-reactors is illustrated schematically in Figures 8, 9, and 10 (in which lllce parts have like numbers) and die table below. In Figures 8, 9, and 10 the valves are not shown. The main reactor configuration is the same however the switching of the valves causes the pre-reactor, scavenger reactor and the guard reactor to appear to “switch ^” places. One pre-reactor operates as such and converts part of the feed stream to ethy lene. One oxy gen depleted pre-reactor acts as a primary' oxygen scavenger or chemisorption reactor and a second pre-reactor (also oxygen depleted acts as a guard or secondary' oxygen scavenger or chemisorption reactor). In an alternate embodiment the oxygen may be separated from the product stream using cryogenic methods. However, this adds both capital and operating costs to the process.

[1193] Operation:

[1194] The above scavenging process is more fully described in Canadian Patent application 2,833 822 filed Nov. 21, 2013, tiie text of which is herein incorporated by reference.

[1195] Residual gases from the downcomer would also be subject to the same separation techniques to recover them.

[1196] As noted above it may not be necessary to separate the ethane from the ethylene at this stage however, if desired there are a number of techniques that may be used.

[1197] The most common techniques would be to use a ayogenic C>; splitter. Other separation techniques include the following

[1198] One method of separation of the product stream is by absorption. The gaseous product stream comprising primarily ethane and ethylene may be contacted in a counter current flow' with a heavier paraffinic oil such as mineral seal oil or medicinal white oil at a pressure up to 800 psig ( about 5.5xl0 ^J kPag) and at temperahixes from about 25 °F to 125 °F (about -4 °C to about 52 °C). The ethylene and lower boiling components are not absorbed into the oil. The ethane and higher boiling components are absorbed into the oil. The ethylene and low'er boiling components may then be passed to the C _? splitter if required. The absorption oil may be selectively extracted with a solvent such as furfural, dimethyl formamide, sulfur dioxide, aniline, nitrobenzene, and other known solvents to extract any heavier paraffins. Tins process is more fully described in U.S. Patent 2,395,362 issued May 15, 1945 to Welling, assigned to Phillips Petroleum Company, the contents of which are herein incorporated by reference. [1199] Another separation method is an adsorption method. The adsorbent preferentially adsorbs one of the components in the product stream. The adsorption method typically comprises a train of two or more adsorption units so that when a unit has reached capacity' the feed is directed to an alternate unit while the fully loaded unit is regenerated typically by one or more of a change in temperature or pressure or both.

[1200] There is a significant amount of art on the separation of eth lene and ethane using silver or copper ions in their +1 oxidation state. The olefins are preferentially absorbed into a complexing solution that contains the complexing agent selected from silver (1) or copper (I) salts dissolved in a solvent. Some silver absorbents include silver nitrate silver fluoroborate, silver fluorosilicate, silver hydroxyfluoroborate, and silver trifluoroacetate. Some copper absorbents include cuprous nitrate; cuprous halides such as cuprous chloride; cuprous sulfate; cuprous sulfonate; cuprous carboxylates; cuprous salts of fluorocaibox lic acids, such as cuprous trifluoroacetate and cuprous perfluoroacetate; cuprous fluorinated acetylacetonate; cuprous hexafluoroacetylacetonate; cuprous dodecylbenzenesulfonate; copper-aluminum halides, such as cuprous aluminum tetrachloride; CuAlCIfrCh; CUA1C ₂ H5CI3; and cuprous aluminum cyano trichloride. If the product stream has been dried prior to contact with the liquid adsorbent, the absorbent should be stable to hydrolysis. The complexing agent can be stable and can have a high solubility in the solvent. After one adsorbent solution is substantially loaded the feed of product stream is switched to a further solution. The solution of adsorbent which is fully loaded is then regenerated through heat or pressure changes or both. This releases the ethylene.

[1201] In an alternative to the solution process supports such as zeolite 4A, zeolite X, zeolite Y, alumina and silica, may be treated with a copper salt, to selectively remove carbon monoxide and/or olefins from a gaseous mixture containing saturated hydrocarbons (i.e paraffins) such as ethane and propane.

[1202] Similarly, of copper salts and silver compounds supported alternatively on silica, alumina, MCM-41 zeolite, 4A zeolite, carbon molecular sieves, polymers such as Amherlyst-35 resin, and alumina may be used to selectively adsorb olefins from gaseous mixtures containing olefins and paraffins. Both kinetic and thermodynamic separation behavior was observed and modeled. The adsorption of the olefin takes place at pressures from 1 to 35 atmospheres, for example, less than 10 atmospheres or less than 2 atmospheres at temperatures from 0 °C to 50 °C, preferably from 25 °C to 50 °C and the desorption occurs at pressures from 0.01 to 5 atmospheres, for example 0.1 to 0.5 at temperatures from 70 °C to 200 °C or from 100 °C to 120 °C.

[1203] In some embodiments, the adsorbent may be a physical adsorbent selected from the group consisting of natural and synthetic zeolites without a silver or copper salt.

[1204] In general, the adsorbent may be alumina, silica, zeolites, carbon molecular sieves, etc. Typical adsorbents include alumina, silica gel, carbon molecular sieves, zeolites, such as type A and type X zeolite, type Y zeolite, etc. The adsorbents can be of type A zeolites, for example, type 4A zeolite.

[1205] Type 4A zeolite, i.e. the sodium form of type A zeolite, has an apparent pore size of about 3.6 to 4 Angstrom units. This adsorbent provides enhanced selectivity and capacity in adsorbing ethylene from ethylene- ethane mixtures and propylene from propylene -propane mixtures at elevated temperatures. This adsorbent is most effective when it is substantially unmodified, i.e. when it Isas only sodium ions as its exchangeable cations.

However, certain properties of the adsorbent, such as thermal and light stability, may be improved by partly exchanging some of the sodium ions with other cations (other than silver or copper). Accordingly, in some embodiments a type 4 A zeolite in which some of the sodium ions attached to the adsorbent are replaced with other metal ions is used, provided that the percentage of ions exchanged is not so great that the adsorbent loses its type 4A character. Among the properties that define type 4A character are the ability of the adsorbent to selectively adsorb ethylene from ethylene-ethane mixtures and propylene from propylene-propane gas mixtures at elevated temperatures, and to accomplish this result without causing significant oligomerization or polymerization of the aikenes present in the mixtures. In general, it has been determined that up to about 25 percent (on an equivalent basis) of the sodium ions in 4A zeolite can be replaced by ion exchange with other cations without divesting the adsorbent of its type 4A character. Cations that may be ion exchanged with the 4A zeolite used in the alkene-alkane separation include, among others, potassium, calcium, magnesium, strontium, zinc, cobalt, manganese, cadmium, aluminum, cerium, etc. When exchanging other cations for sodium ions it can be preferred that less than about 10 percent of the sodium ions (on an equivalent basis) be replaced with such other cations. The replacement of sodium ions may modify the properties of the adsorbent. For example, substituting some of the sodium ions with other cations may improve the stability of the adsorbent.

[1206] In some embodiments, the zeolite is ZSM-5.

[1207] In addition to zeolites there are a number of titanosilicate homologues referred to as ETS compounds. [1208] However, cationic modification of as prepared Na-ETS-10 provides ao adsorbent for the PSA separation of olefins and paraffins having the same number of carbon atoms, at ambient temperatures. The mono-, di- and fii-valent cations are selected from the group 2-4 metals a proton, ammonium compounds and mixtures thereof. Some specific non-limiting examples of mono-, di-, ortri-valent cations that can be used include, Li*, K*, Cs*, Mg ² *, Ca ² *, Sr ^2* , Ba ² *, Sc ³ *, Y ³ *, La ^3* , Cu ^* , Zn ² *, Cd ² *, Ag*, Au ^* , H*, NH ₄ ^* . and NR ₄ * where R is an alkyl, aryl, alkylaryl, or axy!alkyl group. The cationic modifiers are generally added to unmodified Na-ETS-10 in the form of a salt or an acid. The anionic counterion associated with the cationic modifier is not specifically defined, provided that is does not adversely affect the modification (i.e. cation exchange) reactions. Suitable anions include but are not limited to acetate, carboxylate, benzoate, bromate, chlorate, perchlorate, chorite, citrate, nitrate, nitrite, sulfates, and halide (F, Cl, Br, I) and mixtures thereof. Suitable acids include inorganic and organic acids.

[1209] Heat treatment of ETS compounds can result in a controlled pore volume zeolite material, dubbed “CTS-1” which is a highly selective absorbent for oiefin/paraffin separations. The CTS-1 zeolite, which lias pore diameters of from about 3-4A, selectively adsorbed ethylene from a mixture of ethylene and ethane through a size exclusion process. The pore diameter of CTS-1, allowed diffusion of ethylene, while blocking diffusion of ethane which was too large to enter the pores of the CTS-1 zeolite, thereby providing a kinetic separation. The CTS-1 adsorbent was successfully applied to a PSA process in which ethylene or propylene could be separated from ethane or propane respectively.

[1210] The above adsorbents may be used in pressure swing adsorption units. Typically, the range of absolute pressures used during the adsorption step can be from about 10 kPato about 2,000 kPa. (about 1.5 to about 290 pounds per square inch (psi) or from about 50 kPa to about 1000 kPa (from about 7.2 to about 145 psi). The range of pressures used during the release of adsorbate (i.e. during the regeneration step) can be from about 0.01 kPag to about 150 kPag (about 0.0015 to about 22 psig) or from about 0.1 kPag to about 50 kPag (about 0.015 to about 7 3 psig). In general, the adsorption step can be carried out at from ambient temperatures to above about 200 °C, less than 150 °C, or less than 100 °C, provided that the temperatures do not exceed temperatures at which chemical reaction of the olefin, such as a oligomerization or polymerization takes place.

[1211J Another class of adsorbents is ionic liquids. Olefins and paraffins can be separated using ionic liquids of the formula a metal dithiolene selected from the group of complexes of the formulae:

(i) M[S ₂ C _{2 (} R ^' )[.·: and

(ii) \I|S _. · C. iR ^: R ' l R ;!-. wherein M is selected from the group consisting of Fe. Co, Ni, Cu. Pd and Pt; and R ^! , R ² , R ³ , R ⁴ , R ⁵ , and R ⁶ are independently selected from the group consisting of a hydrogen atom, electron-withdrawing groups including those that are or contain heterocyclic, cyano, carboxyiate, carboxylic ester, keto, nitro, and sulfonyl groups, hydrocarbyl radicals selected from tire group consisting of Ci-6, alkyl groups, Cs-8, alkyl groups, C2-8, alkenyl groups and Ce-s aryl groups which hydrocarbyl radicals are unsubstituted or fully or partly substituted (in some embodiments, substituted by halogen atoms). The ionic liquid may be used with a non-reactive solvent or co solvent. The solvent may be selected from the group conventional aromatic solvents, typically toluene. Adsorption pressures may range from 200 psig to 300 psig (i 3x!0 ³ to 2xl0 ³ kPag) or below 250 psig (i 7xl0 ³ kPag) and adsorption temperatures may range from ambient to 200 °C or below 150 °C, and the olefin may be released from the ionic liquid by one or more of lowering the pressure by at least 50 psig (3.4xi0 ² kPag) and increasing the temperature by not less than 15 °C. f 1212] The nitrogen from the regeneration rector, not recycled to the regeneration reactor could be used in a number of downstream unit operations. Potential downstream unit operations include an ammonia plant, an acrylonitrile plant, a urea plant and an ammonium nitrate plant

[1213] The catalyst used in the following experiments was of the formula:

MOaVbNbcTeeOd where a is from 0.90 to 1.10, b is from 0.25 to 0.4, c is from 0 to 0.3, e is from 0 to 0.3, and d is a number to satisfy the valence state of the mixed oxide catalyst.

[1214] The reactor used in Site experiments consisted of a quartz tube reactor. The sample size w ^; as typically about 0.5 cm ³ , 0.17 g. The particle size for the catalyst was 0.2 - 0.7 nmi.

[1215] The reactor was initially operated in a regeneration (oxidation of the catalyst) mode. The reactor was heated to a temperature from 355 °C to 397 °C in air for 30 minutes. Then the gas flow was switched to a mixture of 75 voL-% ethane and 25 voL-% ox gen. The flow rate of the mixture of ethane and oxygen was varied over 300/600/1200 cut ³ (Sip) per hour. The reaction took place during tlie first minute of the passage of the reactants over the oxidized catalyst bed. The catalyst bed was then reoxidized and then a mixture of ethane and oxygen were passed over the oxidized catalyst. The gas leaving the reactor was analyzed to measure the residual oxygen and the amount of ethane, ethylene and by-products in the product gas.

[1216] Expert me nt #K 1 :

[1217] (Air <--- gas mixture [75%C ₂ H ₆ +25%0 ₂ ]) 355 °C

[1218] Figures 51a, 51b, 52a, and 52b demonstrate a time dependence of the ethane and 0 ₂ conversion as well as the selectivity of ethylene formation upon the gradual reduction of the pre-oxidized catalyst by the reaction mixture supplied at 600 cc/hrat two different temperatures.

[1219] As one can see (Figs. 5 la, 5 lb, 52a, and 52b), all transient processes take place during the first operation minute in our testing conditions. The effect is not pronounced at — 355 °C, only a slight increase of the conversion without any selectivity loss can be noted (Figs. 51a and 5 lb). The same effect of the conversion rise becomes much stronger at 400 °C, but in this case it is accompanied by a substantial loss of the selectivity due to additional formation of C(¾ (Figs. 53a and 53b). It is necessary to mention that the more actively occurring process at 400 °C is accompanied by measurable self-heating of the catalyst layer (~5-6 °C measured on the w'all of the reactor). Some contribution of undesirable complete oxidation in the gas phase cannot be excluded

[1220] Experiment K2:

[ 1221] Experiment K 1 w ⁷ as repeated at 398 °C.

[1222] Comparing experiments K! and K2, the maximum conversion was increased to above 70%.

[1223] Experiment K3:

[1224] For clarification, an additional test was carried out with varied gas flow ⁷ rates (300 and 1200 cc/min) using same condition as experiments Ki and K2, except that the flow rate was 1200 cm3/h. Results obtained are presented in Figures 53a, 53b, 54a, and 54b.

[1225] The data obtained (Figs. 52a~54b), shows that the selectivity is related to the feed flow rate (space velocity). Reduction of the gas flow ⁷ rate down to 600 h ^'1 causes a temporary ⁷ drop of the selectivity down to 75% (Fig 54b) Again, the process is accompanied by considerable self-heating of the catalytic layer after the gas switch to the reaction mixture (-6-7 °C measured on the wall of the reactor). The selec tivity curves are summarized and compared in Fig. 55. It is interesting to note at short residence time, despite the high conversion, very little gas phase oxy gen is consumed. So, the contribution of undesirable complete oxidation with temporary heating of the catalyst bed becomes more and more pronounced upon tire rise of the contact time (Fig. 55). At the same time, an increase of the gas flow rate up to 2400 h ^'1 permits us to avoid a considerable contribution of total oxidation (Fig. 55).

[1226] For quantitative comparison of the conversion data, all the results obtained a t flow rates differing by a factor of 2 are presented in Fig. 56 using an absolute scale (i.e., as a function of the amount of the ethane fed through the reactor). All three curves look quite similar (Fig. 56). It is evident that the increased starting conversion of ethane (70-80%) is caused by the presence of extra-oxygen stored in the pre-oxidized catalyst, and Site transient process shown in Figure 56 is related with the gradual loss of this additional oxygen. The results obtained permit us to calculate the amount of the “reactive” lattice ox gen involved in the reaction during die transient process. Depletion of oxygen from die catalyst is the same for all three tests and can be evaluated as -1% from die total lattice oxygen of our Mo-V-Te-Nb-O _x catalyst.

[1227] Experiment K4:

[1228] Periodical redox cycle (pure 0 ₂ <-> pure C ₂ H ₆ ): reference testing

[1229] For clarification this experiment was done in die absence of oxygen in die etliarie stream under die same other conditions (i.e. flow rate and temperature). In tills test, the catal st charge placed into a quartz reactor was heated to at given temperature (354 °C or 397 °C) in pure 0 ₂ flow, kept for 30 min. then the gas flow (600 cm ³ /h) was switched to pure C ₂ H ₆ . and the sample of the outgoing gas was analyzed after a given time. After reoxidation of the catalyst for 30 min measurements were repeated several times with varied time interval, and resulting response curves of products were received (up to 3 min). Figures 57a, 57b, 58a, and 58b demonstrate a time dependence of ethane conversion and residual 0 ₂ content as well as selectivity of ethylene formation upon the catalyst gradual reduction by the ethane at two different temperatures.

[1230] Transient processes take place during 1-2 minutes in our testing conditions (Figs. 57a, 57b, 58a, and 58b). Reaction is accompanied by a measurable selectivity? loss. Effect is quite pronounced even at -350 °C (Fig

57b) and becomes stronger at 400 °C (Fig. 58b). It is important to note that reaction is accompanied by a measurable self-heating of the catalyst layer (4-8 °C measured on the outer wall of the reactor). Back switch to 0 ₂ flow? for the catalyst reoxidation is also accompanied by some catalyst heating (3-4 °C). In addition, this heating seems to be non- uniform but moving throughout the layer during reaction. Taking into account that this over-heating of the catalyst is considerably stronger inside the catalyst bed the role of non-isothennal conditions provided by switch between two pure gases could be important.

[1231] The above examples also illustrates that the conversions and selectivity using a pulse mode of GDH are not as effective as provided techniques.

[1232] Techniques are provided for an oxidative dehydrogenation process using circulating bed reactor (similar to a fluidized bed catalyst cracker (FCCj) providing good yields of olefin product at high selectivity. f 1233] Incorporating into a fixed bed reactor for an exothermal reaction having a catalyst supported on a support having a thermal conductivity typically less than 30 W/mk within the reaction temperature control limits heat dissipa tive particles having a thermal conductivity of at least 50 W/mk less titan 30 W/mk within the reaction temperature control limits helps control the temperature of the reactor bed The Catalyst

[1234] Techniques are provided that are suitable for use with any fixed bed reactor in which there is a desire to have a better control over the heat flow within the fixed bed and also the transfer of heat into or out of the bed. Since the inert non catalytic heat dissipative particles present in Hie bed have a thermal conductivity of greater titan 50. in some embodiments 100, in further embodiments 150, still further embodiments 200, W/mK (watts/meter Kelvin) within the reaction temperature control limits, the inert non catalytic heat dissipative particles may transfer heat directly to the walls of the reactor improving the cooling homogeneity (or heating if the wall are heated) and reduction of hot spots in the fixed bed.

[1235] The reactions may comprise one or more of oxidative cracking, isomerization, oxidative coupling, oxidative dehydrogenation, hydrogen transfer, polymerization and desulphurization of a hydrocarbon or any other exothermic reaction. In some embodiments, the reaction is oxidative dehydrogenation of a C alkane or the oxidative coupling of a C _M alkane. These last two reactions are of concern as the feed comprises a hydrocarbon and oxygen. If the ratio of oxygen to hydrocarbon exceeds the lower flammability (explosive) limit and the reaction temperature of the bed exceeds the ignition temperature of the mixture there is a certainty of an undesired outcome. [1236] In some methods of carrying out such reactions, the reactant stream is diluted with steam or an inert gas such as nitrogen to keep the reactive mixture below the lower flammability (explosive) limit. This type of approach tends to reduce the per pass conversion of the reactants and product stream needs to be separated, typically using some type of unit like a C _. splitter which is energy' intensive and greenhouse gas producing.

[1237] Another approach is to operate such reactions above the lower flammability (explosive) limit but at a temperature below the auto-ignition temperature of the feed. In such a method of operation it is critical to have a uniform temperature within the bed (i.e no hot spots) and to have a good control over the removal of heat from the fixed bed. Hydrothermal synthesis for preparation of mixed metal oxide catalysts is known in the art

[1238] Generally a hydrothermal synthesis step is used for preparation of the catalyst prior to addition of the Pd compound. Compounds containing elements Mo, V, Nb, and Te and a solvent are mixed to form a first admixture. The first admixture is then heated in a closed vessel for from 24 to 240 hours. One useful solvent for the hydrothermal synthesis of the first admixture is water. Any water suitable for use in chemical syntheses can be utilized, and includes, without limitation, distilled water, de-ionized,) water. The amount of solvent used is not critical.

[1239] Preparation of the admixture is not limited to addition of all compounds of Mo, V, Nb, and Te at the same time prior to heat treatment in a first closed vessel. For example, the Mo and Te compounds may be added first, followed by the V compound and eventually the Nb compound. For a further example, the process may be reversed in that the Te and Nb compounds are combined followed by addition of a mixture of the Mo and V compounds. Other sequences of addition would be apparent to a person skilled in the art. Sequence and timing of addition is not limited by these examples.

[1240] In some embodiments, the first admixture is heated at a temperature of from 100°C to 200°C. In some embodiments, the first admixture is heated at a temperature from 130°C to 190°C. In some embodiments, the first admixture is heated at a temperature from 160°C to 185°C.

[1241] Following hydrothermal synthesis of the first four components of the catalyst the first insoluble material is recovered from the first closed vessel. At this point, the first insoluble material may be dried prior to a fust calcining in order to remove any residual solvent. Any method known in the art may be used for optional diving of the first insoluble material, including, but not limited to, air drying, vacuum drying, freeze drying, and oven drying.

[1242] In some embodiments, the first insoluble material may be subjected to peroxide washing prior to optionttl drying and prior to a first calcining. The peroxide washing treatment may take place at atmospheric pressure and room temperature (e.g. from 15 °C to 30 °C) to about 80 ^C C, in some cases from 35 °C to 75 °C, in some cases from 40 °C to 65 °C and the peroxide has a concentration from 10 to 30 wt. %, in some cases from 15 to 25 wt. %, and a time from 1 to 10 hours, in some cases from 2 to 8 hours, in some cases from 4 to 6 hours.

[1243] The first insoluble material is treated with the equivalent of from 1.3 to 3.5 mis of a 30 wt. % solution of H2O2 per gram of precursor. The treatment should be in a slurry' (e.g. the precursor is at least partially suspended) to provide an even distribution of I-fiO _? and to control the temperature rise. For post calcination treatment with ¾<¼ there is a delayed violent reaction with H2O2. The process is an instantaneous reaction which is more controlled and safer.

[1244] Methods for calcination are well known in the art. The first calcining of the first insoluble material is conducted in a second closed vessel with an inert atmosphere. The second closed vessel for the calcination may be a quartz tube. The inert atmosphere may include any material that does not interact or react with the first insoluble material. Examples include, without limitation, nitrogen, argon, xenon, helium or mixtures thereof. In some embodiments, the inert atmosphere comprises gaseous nitrogen.

[1245] Calcination methods for preparation of mixed metal oxide catalysts vary in the art. Variables include the time, temperature range, the speed of heating, use of multiple temperature stages, and the use of an oxidizing or inert atmosphere. The speed of heating is not critical and may range from between O.PC/minute to around 10°C/minute. In addition, the inert gas may be present statically or may be passed over the catalyst at flow rates where the loss of catalyst is minimized, i.e. carryover out of bed.

[1246] In some embodiments, the time for the first calcining ranges from I hour to 24 hours. In some embodiments, the time for the first calcining ranges from 3 hours to 15 hours. In some embodiments, the time for the first calcining ranges from 4 hours to 12 hours.

[1247] In some embodiments, the first calcining takes place in an inert atmosphere at a temperature front 500 °C to 700 °C. In some embodiments, the first calcining takes place in an inert atmosphere at a temperature from 550 °C to 650 °C. In some embodiments, the first calcining takes place in an inert atmosphere at a temperature of front 580 °C to 620 °C. The resulting calcined product is suitable as an oxidative dehydrogenation catalyst. f 1248] In some embodiments following the first calcining, the first calcining product is mixed with a Pd component to form a second admixture. For these aspects, the addition of a Pd component to the catalyst is only effective in increasing the activity of the catalyst, without significantly decreasing the selectivity', depending on the me thod for addition and the nature of the Pd compound used. The addition of the Pd compound must be performed following the first calcining of the first insoluble material containing the four components Mo, V, Te, and Nb. In some embodiments the Pd compound, in the form of an aqueous solution, is added dropwise to the first calcining product until saturation and the mixture forms a paste in some embodiments, the Pd component and the first calcining product are mixed in an aqueous solution to form a slurry. In sortie embodiments, the aqueous solution is water. Any water suitable for use in chemical syntheses can be utilized, and includes without limitation distilled water and de-ionized water. The amount of solvent used is not critical.

[1249] The amount of Pd component added, either in dropwise fashion or in a slimy, will correspond roughly with 0.044 mmolp _d /go _DH catalyst to yield a final relative atomic amount of Pd, represented by the subscript e in the formula Mo _a VbTe _c NbdPdeOf, between 0.001 and 0.1.

[1250] The nature of the Pd compound used must be free of halogens. One useful Pd component is tetra-amine Pd nitrate, chemically represented by the formula [(NHfoPdXNOah.

[1251] Before the second calcining the second admixture, the product may be dried using any method known in the art, including, but not limited to air drying, vacuum drying, freeze drying, and oven drying.

[1252] The second calcining is performed under conditions and follows the same limitations as those applicable to the first calcining. The resulting second insoluble material is retrieved from the second closed vessel and can be used directly as a catalyst for ODH, using conditions where the only atmospheric components exposed to the catal st are oxygen and ethane. The ratios of oxygen and ethane and the temperature used for the ODH process are such that the upper explosive limit is not triggered. The ability' to perform ODH using this catalyst whereby there is no dilution of the reactants with nitrogen or other inert gas or water confers a commercial advantage as costly downstream processes for the removal of excess oxygen or any unwanted byproducts are not required or are limited in nature.

[1253] The Heat Dissipative Particles for the Fixed Bed

[1254] The heat dissipative particles for the fixed bed comprises one or more non catalytic inert particulates having a melting point at least 30 °C, in some embodiments at least 250 °C, in further embodiments at least 500 °C above the temperature upper control limit for the reaction, a particle size in range of 0.5 to 75 mm, in some embodiments in the range of 0.5 to 15 mm, in some embodiments in the range of 0.5 to 8 mm, in some embodiments in the range of 0.5 to 5 mm and a thermal conductivity' of greater than 30 W/rnK (watts/meter Kelvin) within the reaction temperature control limits in some embodiments, the particulates are metals alloys and compounds having a thermal conductivity of greater than 50 W/inK (watts/meter Kelvin) within the reaction temperature control limits. Sortie suitable metals include silver copper, gold, aluminum, steel, stainless steel, molybdenum, and tungsten.

[1255] The heat dissipative particles may have a particle size typically front about 1 to 15 mm. In some embodiments the patficie size may be from about i mm to about 8mm. The heat dissipative particles may be added to the fixed bed in an amount from 5 to 95 wt. %, in some embodiments 30 to 70 wt. %, in other embodiments 45 to 60 wt. % based on the entire weight of the fixed bed.

[1256] The Processes

[1257] The provided techniques may be used with any fixed bed exothermic reaction. In some embodiments, the fixed bed reactor is a tubular reactor, and in some embodiments, the fixed bed reactor comprises of multiple tubes inside a shell (e.g. a shell and tube heat exchanger ty pe construction). In some embodiments, the fixed bed reactor may comprise a number of shells in series and/or parallel. The reactions may involve one or more of cracking, isomerization, dehydrogenation including oxidative dehydrogenation, h drogen transfer including oxidative coupling and desulphurization of a hydrocarbon.

[1258] Typically these reactions are conducted at temperatures from about 200 °C up to about 850 °C at pressures from about 80 to 21,000 kPag (about 12 to 3000 psig) in the presence of a catalyst. The hydrocarbon stream may contain a wide range of compounds including C1.20 aliphatic, or aromatic hydrocarbons.

[1259] In some embodiments, tire reactions are the oxidative coupling of aliphatic hydrocarbons typically C1.4 aliphatic hydrocarbons particularly methane and the oxidative dehy drogenation of C2.4 aliphatic hydrocarbons. Such reactions may be conducted using a mixed feed of hydrocarbon, in some embodiments methane or ethane and oxygen i n a volume ratio from 70:30 to 95:5 at a temperature less than 420 °C at a gas hourly space velocity of not less than 280 hr ¹ , in some embodiments not less than 1000 hr ¹ , in some embodiments not less than 2000 hr ¹ and a pressure from 80 to 1000 kPa (0.8 to 1.2 atmospheres). Typically, the process may have an overall conversion of from about 50 to about a 100%, typically from about 75 to 98% and a selectivity to ethylene of not less than 90%, in some instances not less than 95%, in further embodiments not less than 98%. In some cases, the temperature upper control limit is less than about 400 °C, in some embodiments less than 385 °C.

[1260] The resulting product stream is treated to separate ethylene from the rest of the product stream which may also contain co-products such as acetic acid, and un-reacted feed which is recycled back to the reactor.

[1261] Additionally, the product stream should have a low content of carbon dioxide, and casbon monoxide, and acetic acid, generally cumulatively in a range of less than 10 wt. % or less than 2 wt. %.

[ 262] There are up to four competing reactions for oxidative dehydrogenation.

Reaction L 1 CM f. + 0.5 (> <-> C ₂ H ₄ + I TO (DH1 = - 105 kJ/Mole (/ ! !,}

Reaction L2 C ₂ ¾ + 2.5 C¾ < 2CO + 3H ₂ 0 (DH2 = - 862 kl/Mole C ₂ ¾)

Reaction L3 C ₂ H ₆ + 3.5 C¾ <->2Ci¾ + 3H ₂ 0 (DH3 = - 1430 kJ/Mole C ₂ H ₆ )

Reaction L4 C ₂ H ₆ + 1.5 C¾ < C ₂ H ₂ 0 ₂ + H ₂ 0 (DH4 = - 591 kJ/Mole C ₂ H ₆ )

[1263] From a temperature / heat control point of view, if a catalyst preferentially leads to reaction Li, there is a lower potential for a thermal runaway.

[1264] The feed and by products may need to be separated from the product stream. Some processes may use so called dilute ethylene streams. For example, if the product stream does not contain too much ethane, for example less than about 15 vol. % the stream may be fed directly without further purification to a polymerization reactor such as a gas phase, slimy or solution reactor. f 1265] The most common separation technique would he to use a cryogenic C2 splitter. Other known ethylene/ethane separation techniques could also be used including adsorption (oil, ionic liquids and zeolite).

[1266] In the non-limiting examples, the catalysts were prepared by a hydrothermal process as described above.

[1267] The catalyst in Example LI had the empirical formula: (MoiooVoaeTeo nNbo _. niOi _. s? as determined by XRD.

[1268] For the comparative example, the catalyst was not treated with hydrogen peroxide. For Example L 1 , the sample comprise a mixture of five catalyst samples treated with per oxide. The catalyst for the comparative example lias a slightly higher propensity to oxidize feed to C0 _2.

[1269] In the examples, the fixed bed reactor unit used is schematically shown in Figure 33. The reactor was a fixed bed stainless steel tube reactor having a 2 mm (¾?) outer diameter and a length of 117 cm (46 inches). The reactor is in an electrical furnace sealed with ceramic insulating material. There ate seven thermocouples in the reactor indicated at numbers // 1 through #7. Thermocouples are used to monitor the temperature in that zone of the reactor. Thermocouples #3 and #4 are also used to control the heating of the reactor bed. The feed flows from the top to the bottom of the reactor. At the inlet there is a ceramic cup 3308 to prevent air drafts in the reactor. Below' the ceramic cup is a layer of quartz wool 3309. Below the layer of quartz wool is a layer of catalytically inert quartz powder. Below the quarts powder is the fixed bed 3310 comprising catalyst and diluent. Below' the fixed bed is a layer of quartz powder 3311, a layer of quartz wool 3312 and a ceramic cup 3313. At the exit of the bed was a gas analyzer to determine the composition of the product stream. The fixed bed comprised 28.83 g of catalyst and 3.85 g of diluent (32.86 g total weight % of diluent 11.7 wt. % of total bed.). The GHSV was 2685 hr-1 and the pressure was ambient.

[1270] For the examples, the bed temperature was taken as an average of the temperatures from thermocouples #2, #3, and #4. The feed stream was assumed to have the same temperature as the bed. A stoichiometric reactor block was run using the above temperature conditions using Aspen Plus simulation to calculate the overall beat release of the reactions.

[1271] Comparative Example

[1272] The heat dissipative particles in this example were quartz particles having a mean particle size of 568 micrometers. The reaction temperature (bed temperature) increased to 355 °C and then there was a thermal reaction mu away. The overall conversion to ethylene was 19% and the selectivity' to ethylene was 93%. The calculated heat duty of the reactions was calculated to be a heat release of -26.28 kJ/hr. At this time there was a rapid drop in oxygen content in the product stream and a fast thermal reaction runaway began. The reaction was quenched with nitrogen.

[1273] Example Li

[1274] The heat dissipative particles were 316 Stainless Steel particles having a mean particle size of 568 micrometers. The weight % of diluent was Site same as for Example LI. As the steel is denser tliart quartz tins resulted in a lower volume % of diluent in the bed. These conditions were believed to tend toward a thermal runaway. The reactor was operated to maintain an overall conversion of 19% with a selectivity to ethylene of 89%. The calculated overall heat of reaction was -31.13 kj/hr. The temperature of the bed rose to 372 °C. No runaway reaction was observed. The stainless steel diluent permitted a better release of heat through the reactor walls to control the reaction.

[1275] Example LI shows the bed temperature did not rise above 372 °C while in the comparative example the bed temperature approached 355 °C followed by a thermal reaction run away. Example LI shows dissipation in the heat of reaction.

[1276] Techniques are provided to control/dissipate the heat generated from the oxidative dehydrogenation reaction.

[1277] Techniques described herein provide a system for stable homogenous mixing of gases. More specifically the system manages the mixing of a hydrocarbon gas, such as ethane, with a gaseous oxidant in a maimer to avoid ignition. Additionally, some embodiments are directly applicable to use in the catal tic oxidative dehydrogenation of ethane into eth lene.

[1278] Provided herein is a gas mixer for the desirable mixing of a h drocarbon containing gas with a gaseous oxidant. The gas mixer and method for mixing described includes a closed mixing vessel where bubbles of gas injected at the bottom of the vessel are mixed during their rise to the top of the vessel forming a homogeneous mixture that ca stably be removed. This simple design and method allows for managed mixing of gases and is applicable to avoid excess reaction in catalytic oxidative processes such as oxidative dehydrogenation of paraffins [1279] Mixing of a hy drocarbon containing gas with a gaseous oxidant can be associated with potential ignition. When the ratio of hydrocarbon to oxygen within a mixture is within the flammability envelope, ignition can cause an excess reaction. For processes where mixing of hydrocarbons with oxygen is implemented, avoiding ignition or undesired reaction may be involved. While processes for catalytic oxidation of hydrocarbons, for example ODH of paraffins into olefins, are typically performed with mixtures of hydrocarbons and oxygen that are outside the flammability envelope, there remains potential for ignition when the mixture is heterogeneous and includes hotspots where the ratio of hydrocarbon to oxygen is within flammability limits. This is particularly true when the components are in the initial phase of mixing before homogeneity is achieved

[1280] It is commonly known that for ignition to occur in this context, there must be both a mixture of hydrocarbon and oxygen within the flammability envelope and an ignition event. An ignition event may take the form of entrained particles present within either the hydrocarbon or gaseous oxidant striking a metallic surface within the mixing apparatus and creating a spark. If the spark occurs in a region near where streams of the hydrocarbon gas and gaseous oxidant meet and have not reached homogeneity outside flammability limits, ignition may result. With that in mind, when mixing hydrocarbons and oxygen, a focus may be to decrease the chances of an ignition causing spark or maximizing the rate of mixing to shorten the window when ignition may occur due to the existence of heterogeneous pockets of unfavorable hydrocarbon/oxygen compositions. Additional options me applicable, as disclosed herein.

[1281] A relatively simple gas mixer for stable operation of mixing a hydrocarbon containing gas with a gaseous oxidant is disclosed herein. The two gases are introduced directly, and separately, into the bottom of a closed mixing vessel where they form bubbles that are mixed as they rise to the top of the vessel. Mixing results in a homogeneous swarm of bubbles that exit the liquid as non-flammable compositions of hydrocarbon and oxygen, useful for applications involving catalytic oxidation of hydrocarbons, including the oxidative deh drogenation (GDH) of ethane into ethylene.

[1282] Provided is a gas mixer for mixing a hydrocarbon containing gas and a gaseous oxidant comprising a closed mixing vessel flooded with a non-flammable liquid and having a top end, a bottom end, internal mixing means, injection and removal points for non-flammable liquid, inputs for each of the hydrocarbon containing gas and the gaseous oxidant at or near the bottom of the closed mixing vessel, and an outlet at the top of the closed mixing vessel for site removal of a substantially homogeneous mixture of hydrocarbon and oxygen in a ratio that corresponds to the relative amounts of hydrocarbon containing gas and gaseous oxidant introduced into the bottom of the closed mixing vessel.

[1283] Provided is a method for mixing a hydrocarbon containing gas with a gaseous oxidant comprising introducing a hydrocarbon containing gas and a gaseous oxidant into die bottom of a closed mixing vessel flooded with a non-flammable liquid and in a ratio that once combined Mis outside of the flammability envelope and recovering a homogeneous mixture of the previously introduced hydrocarbon containing gas and gaseous oxidant. [1284] Provided is a process for the oxidative dehydrogenation of ethane to ethylene comprising the managed mixing of an ethane containing gas with a gaseous oxidant in a closed mixing vessel flooded with a non-flammable liquid, passing the mixture of the ethane containing gas and the gaseous oxidant through a heat exchanger to raise the temperature to at least 200 °C, introducing the heated mixture into an ODH reactor containing an ODH catalyst removing ethylene, unconverted ethane and the various byproducts that include CO, CO? ¾0, CH ₃ COOH, minimal hydrocarbons and possibly unreacted O2 from the ODH reactor and directing them to a quench tower to remove CH3COOH and H ₂ 0. The unreacted 0 ₂ may be removed in an 0 ₂ removal vessel prior to being sent to an amine wash. The 0 ₂ removal vessel can be placed prior or after the quench tower. In the amine wash, C0 ₂ is removed. Methane may then be removed by passage through a demethanizer. The remaining mixture, containing unreacted ethane, ethylene and other hydrocarbons may be recycled back to go through the process again, starting with mixing with a gaseous oxidant in the closed mixing vessel.

[1285] Gas Mixer

[1286] A schematic representation of an embodiment of the gas mixer is shown in Figure 59. The gas mixer

5901 comprises a closed mixing vessel 5910 having a top end 5909 and a bottom end 5907. The dosed mixing vessel 5910 is flooded with a non-flammable liquid, the choice of which depends on the application for which the mixed gas is to be used. Non-flammable liquid may be added to the closed mixing vessel 5910 via a nozzle or inlet

5902 located at the top end 5909, while non-flammable liquid may be removed from the outlet 5903 located at the bottom end 5907.

[1287] Construction of the mixing vessel 5910 can be accomplished with a variety of materials including stainless steel, carbon steel, and any other material chemically compatible witiithe hydrocarbon to be mixed. Furthermore, the lining of mixing vessel 5910 may be coated with a spark suppressing material such as Teflon, sapphire, or oxide-based ceramic liners or the like. f 1288] Hydrocarbon containing gas may be introduced into the closed mixing vessel 5910 through the hydrocarbon containing gas supply nozzle 5904, while the gaseous oxidant may be introduced via gaseous oxidant supply nozzle 5905. The hydrocarbon containing gas supply nozzle 5904 and the gaseous oxidant supply nozzle 5905 cooperate with the closed mixing vessel 5910 in a way so that introduction of the gases directly into the nonflammable liquid occurs at or near the bottom end 5907 of the closed mixing vessel 5910. For the purposes of this disclosure, the term “nozzle” refers simply to the point where contact between the gases and the non-flammable liquid within the closed mixing vessel 5910 first occurs, and can include any means known within the art. While not essential, the hydrocarbon containing gas supply nozzle 5904 and the gaseous oxidant supply nozzle 5905 are ideally orientated such that streams of the hydrocarbon containing gas and the gaseous oxidant impinge upon one another immediately upon entering the mixer. The introduced gases rise and are mixed through mixing zone 5908 and are available for removal idler exiting the non-flammable liquid at the top of die closed mixing vessel 5910 through the mixed gas removal line 5906.

[1289] As the term suggests, non-flammable liquids used to flood the closed mixing vessel 5910 must not be flammable. That is, the non-flammable liquid must not be capable of igniting or burning. Suitable non-flammable liquids include water, ethylene glycol, silicon oils, and carbon tetrachloride. One embodiment comprises water as the non-flammable liquid. While any non-flammable liquid may be used with the various embodiments disclosed herein, it is important to consider that mixed gas removed from the gas mixer 5901 will comprise the hydrocarbon containing gas, gaseous oxidant, and in some instances carryover of non-flammable liquid. For this reason selection of a non-flammable liquid must consider any potential effects the carry over may have on downstream applications. Catalysts used for oxidative reactions may be sensitive to catalytic poisoning by specific non-flammable liquids that are carried over in a gaseous state.

[1290] The temperature, along with the pressure, play a role in determining what fraction of the nonflammable liquid may enter the gaseous state, joining the hydrocarbon and oxygen gas present in bubbles that are mixing and rising to the top end of the closed mixing vessel 5910. The temperature and pressure can be controlled to minimize the carryover of non-flammable liquid into the gas mixture leaving through mixed gas removal line 5906. Temperature control using a heater, within or without the closed mixing vessel 5910, is contemplated for use with the provided techniques. In one embodiment the closed mixing vessel 5910 is temperature controlled using a heater that is external to the closed mixing vessel 5910. in another embodiment the closed mixing vessel 5910 is temperature controlled using a heater that is located within the closed mixing vessel 5910.

[1291] In some instances it may be desirable, for recycling purposes, to include a secondary hydrocarbon containing gas supply nozzle or product supply nozzle 5915. For example, some oxidative reactions are not as efficient as others and may include conversion rates below an acceptable level. In those cases, it may be desirable to send a product line containing product and unreacted hydrocarbon back to start the oxidative reaction process again, with the intent of maximizing conversion of the starting hydrocarbon — the hydrocarbon originally mixed in the gas mixer before passage through an oxidative process. The product stream, similar to and containing unreacted starting hydrocarbon, would need to be mixed with oxidant before entering the reactor. If the product contained in the product stream is more reactive to oxygen than tire starting hydrocarbon, it would be safer to introduce the product stream into the reactor at a point where the oxygen is already partially mixed and diluted. To this end the secondary hydrocarbon containing gas supply nozzle 5915 should be at a position distant from the gaseous oxidant supply nozzle 5905. The position of the secondary hydrocarbon containing gas supply nozzle 5915 is not critical, provided it is in a position where the oxygen present in the closed mixing vessel 5910 has begun mixing with the hydrocarbon containing gas, and there is sufficient residence time for the product gas to mix thoroughly with the added oxygen and hydrocarbon containing gases ideally, the position of the secondary hydrocarbon containing gas supply nozzle is near a point equidistant from the gaseous oxidant supply nozzle 5905 and the point where mixed gas removal line 5906 leaves the top end 5909 of the closed mixing vessel 5910. The secondary hydrocarbon containing gas supply nozzle 5915 may also be used as mi additional input location for the introduction of the hydrocarbon containing gas. In one embodiment, there is a secondary' hydrocarbon containing gas supply nozzle 5915 for introducing a product stream from an oxidative process or additional hydrocarbon containing gas into tire closed mixing vessel 5910 at a point distant from gaseous oxidant supply nozzle 5905.

[1292J In instances where there is recycling of an oxidative process such that a product line is fed back to tire gas mixer 5901 for introduction into the closed mixing vessel 5910 via the secondary hydrocarbon containing gas supply nozzle 5915, it is contemplated that heat from the product line may be used in temperature control of the closed mixing vessel 5910. The heat provided from an oxidative process, for example ODIT, may be used in this fashion and would therefore assist in reducing the cost associated with providing heat through an internal or external heater. In another embodiment, the closed mixing vessel 5910 is temperature controlled using heat from a product line leaving an exothermic oxidation process.

( ί 293] Internal mixing

[1294] The efficiency of mixing of the gases within zone 5908 is dependent upon, among other things, the residence time and the frequency of interactions between bubbles of gas. In other words, how' often do bubbles collide break and reform together, permitting mixing of the gas compositions from each of the bubbles which combine to form a homogeneous mixture. While mixing can occur naturally given sufficient time, it is not likely that a homogeneous mixture will be produced without internal mixing where collisions between bubbles are promoted. Without internal mixing the vessel would need to be of such height as to be not economically feasible. Means for promoting mixing are well knowm in the art and include use of a static mixers, random packing, structured packing, and impellers.

[1295] Static mixers promote mixing by creating a multitude of tortuous pathwa s that increase the distance that bubbles need to travel to reach the top of the vessel and consequently static mixers act partly by increasing the residence time in addition, the pathways comprise limited space that results in an increased probability that bubbles collide and ultimately mix to combine their gaseous contents. In an embodiment, the internal mixing means comprises a static mixer.

[1296] Random and structured packing act similar to static mixers in that they provide for increased residence time and probability of interaction between bubbles by creation of a plethora of winding pathway s. Random packing involves filling at least a part of the closed mixing vessel 5910 with a packing material that comprises objects 6012 of varying shape and size (Figure 60 A) that create random pathways for the bubbles to follow as they rise to the top (see dashed arrow in Figure 60 A). An example of commonly used random packing is glass beads of varying diameter. In another embodiment, the internal mixing means comprised a packed bed.

[1297] Structured packing also increases residence time and probability' of contact between bubbles, but differs from random packing in that the structured packing has an ordered arrangement so that most of the pathways are of a similar shape and size (see dashed arrows in Figure 60B). For example, use of corrugated metal plates 6013 (Figure 60B) provides a structured, as opposed to random, array of pathways. In another embodiment, the internal mixing means comprises a structured bed. Figures 60A and 60B are provided as simplified examples for random and structured packing and should not be seen to limit the present disclosure in any way. In addition, while not shown in the figures, random and structured packing are supported within the gas mixer using means known in the art.

[1298] Power driven mixers, which can promote interactions by creating flow within the vessel, can be used. Impellers include a rotating component 6014 (direction of rotation shown by solid circular arrow), driven by a motor, that may force Site non-flammable liquid, and associated bubbles of gas, to the outside wall and away from tiie center of rotation. Impellers can create axial flow or radial flow depending upon design, and can be further subtyped as propellers, paddles, or turbines (see dotted arrows on either side of the gas mixer in Figure 60C). Furthermore, the position of the impeller may be subject to change through vertical mo vement throughout the mixing zone 5908. Motor driven pumping of an impeller further improves mixing. In another embodiment the closed mixing vessel 5910 further comprises an impeller.

[1299] Similar to the closed mixing vessel 5910, the internal mixing means, whether a static mixer, random or structured packing, or ao impeller may be comprised of any material that is chemically compatible with the hy drocarbon to be mixed.

[1300] The shape and design of the closed mixing vessel 5910 impacts the residence time. The overall shape of the vessel is not critical, but the distance between where the gas enters and exits the mixing zone 5908 is. The point of first contact between the gases and the water in the closed mixing vessel 5910 should be a distance from the top that allows for a residence time that permits complete mixing before removal. In another embodiment, the ideal entry point is near the bottom of the vessel. Where the lines containing the gas enter the vessel is not important, provided the nozzle — the point where the gas contacts the water in the vessel — is in the position w'here residence time is sufficient. For example, in Figure 61 the lines feeding the hydrocarbon containing gas and the gaseous oxidant enter the closed mixing vessel 5910 near the top end 5909 but the nozzles 6104 and 6105 introduce the gases near the botom end 5907.

[1301 ] Another consideration for the optimum mixing of the gases is the surface area over which the gases are dispersed. A larger surface area of dispersion promotes better mixing. While injection through a single inlet is feasible, provided sufficient residence time, more thorough mixing occurs when a larger number of smaller bubbles are dispersed over a larger surface area. Having multiple hydrocarbon containing gas supply nozzles and multiple gaseous oxidant supply nozzles allows each of the gases to be introduced in multiple locations. Conversely a single nozzle may comprise multiple exit points where gas can enter the vessel, effectively dispersing the gas over a greater surface area compared to dispersion from a nozzle with a single exit point. In an embodiment, at least one of the hydrocarbon containing gas supply nozzle 6104 and the gaseous oxidant supply nozzle 6105 comprise a sparger. f 1302] In another embodiment, the hydrocarbon containing gas supply nozzle 6204 and the gaseous oxidant supply nozzle 6205 are arranged as spargers in the form of concentric rings. For example, in Figures 62A and 62B the hydrocarbon containing gas supply nozzle 6204 comprises a circular sparger that is larger than and surrounds the gaseous oxidant supply nozzle 6205, which is also circular and fits inside the hydrocarbon containing gas supply nozzle 6204. Furthermore, the exit points for tire hydrocarbon containing gas and the gaseous oxidant from their respective nozzles are arranged such that the streams of gas impinge on one another, initiating mixing as early as possible after introduction into the mixer (see Figure 62 A). The arrangement of the gas supply nozzles is not limited to examples provided here. A series of concentric rings, with alternating oxygen and hydrocarbon containing gas supply nozzles, is also contemplated.

[1303] Emergency Shutdown

[1304] Another embodiment relates to emergency shutdown procedures common to oxidative reaction processes. It is well known that when undesirable conditions occur in an oxidative reaction process an emergency shutdown procedure can be initiated to limit damage to equipment, reduce likelihood of personal injur . and prevent or minimize release of chemicals into the surrounding environment. Known emergency shutdown procedures include the cessation of adding reactants while at the same time providing a flow of an inert material, such as nitrogen, to the reaction zone to displace the reactants from the reactor.

[1305] In some embodiments, it is contemplated that for an additional safety component an inert material inlet, located near the top end and above the liquid level, may be included for the introduction of a flow of an inert material. In addition a suppression outlet leading to any known explosion suppression system may be included near the top end of the gas mixer. When an unsafe operating condition is detected at any point in the oxidative process, flow of an inert material through the inert material inlet can be initiated while the suppression outlet can be opened. These events can be coordinated with a reduction or termination of the hydrocarbon and oxidant reactants. The end result is that any mixed gases within the mixer are displaced to the explosion suppression system or to downstream components of the oxidative process. The flow of inert material acts as diluent and promotes movement in a single direction so that backflow of materials from the oxidation reactor into the gas mixer are prevented

[ \ 306] In an embodiment, the gas mixer further comprises an inert material inlet, located near the top end of the gas mixer, for introducing an inert material into the gas mixer above the level of the non-flammable liquid, and a suppression outlet for removing gaseous mixtures, located near the top end of the gas mixer and leading to an explosion suppression system.

[1307] Method for mixing a hydrocarbon containing gas and a gaseous oxidant

[1308] Techniques are provided for applications that require the mixing of a hydrocarbon containing gas with a gaseous oxidant. It is well known that gaseous compositions containing a hydrocarbon and oxygen in ratios that Ml within the flammability envelope are potentially hazardous. An ignition event, such as a spark, can ignite the mixture and potentially lead to an expiosion. While applications that require mixing of hydrocarbons and oxygen normally do so with ratios that are safe and not susceptible to ignition there are moments during initial mixing where heterogeneous pockets of unfavorable hydrogen/oxygen compositions exist and may ignite if a spark occurs. f 1309] Techniques are provided for a method for mixing a hydrocarbon containing gas with a gaseous oxidant that is simple, and sale in that ignition events are unlikely to occur. The method comprises introducing, separately and simultaneously, a hydrocarbon containing gas and a gaseous oxidant directly into a closed mixing vessel having a top end and a bottom end and flooded with a non-flammable liquid, in close proximity to the bottom end, allowing the bubbles of gas to mix while surrounded by the non-flammable liquid, and removing from the top of the vessel, after mixing is complete, a homogeneous mixture of the hydrocarbon containing gas and the gaseous oxidant in a ratio that is outside of the flammability envelope.

[1310] In some embodiments, the amount of the gases introduced into the bottom end of the closed mixing vessel 5910 will result in a final composition that comprises a ratio of hydrocarbon containing gas to gaseous oxidant that is outside of the flammability envelope. The chosen ratio will depend on the nature of the gases and the application for which the mixture will be used. For example, for an ODH application, Site ratio of ethane to oxygen chosen will depend on whether under the proposed ODH reaction conditions the ratio is above the higher explosive limit or below' the low'er explosive limit. In comparison the ratio of ethylene to oxygen added to the reactor would be different because ethylene is more reactive than ethane. The temperature of the ODH process to be employed must also be taken into consideration as higher temperatures correspond to a smaller window of non-flammable ratios of ethane to oxygen. For example, a molar ratio of about 80:20 ethane to oxygen for catalytic ODH would fall above the upper flammable limit, while a ratio of about 1.5:98.5 ethane to oxygen would fall below the lower flammable limit, with each ratio safe enough in that ignition events would not lead to an explosion or flame propagatio under ODH reaction conditions. Ratios falling between that — 50:50 for example — would be potentially flammable.

[1311] The next consideration after determining the desired final ratio of hydrocarbon to oxygen is determining the flow rate at which each gas is added to the bottom of the closed mixing vessel 5910. The flow rate of the gases and the corresponding pressure would need to be higher than the pressure of the non-flammable liquid in the closed mixing vessel 5910 In the absence of a pressure differential, the gases cannot enter the closed mixing vessel 5910 and consequently the mixing zone 5908. Furthermore, if the pressure of the non-flammable liquid is higher than the line containing the gas to be introduced there may be, in the absence of a one way valve, flow' back of non-flammable liquid into the gas supply lines. This should be avoided.

[1312] When determining flow rates, the skilled worker must correlate the flow' rates with the pressure and temperature used within the closed mixing vessel 5910. The conditions within the closed mixing vessel 5910 are chosen to reflect the amount of carryover of non-flammable liquid into the gas mixture removed through mixed gas removal outlet 5906. The flow' rate of the incoming gases must be sufficient to allow entry into the non-flammable liquid at the predetermined temperature and pressure.

[1313] As a further safety precaution, the gaseous oxidant can be diluted with non-flammable liquid prior to entry into the closed mixing vessel 5910. The prior dilution of the gaseous oxidant permits the saturation of incoming oxygen molecules with molecules of the non-flammable liquid that discourage ignition events igniting any hydrocarbons that interact with the oxygen during the early stages of mixing. Dilution of die gaseous oxidant with non-flammable liquid can be accomplished by directing a non-flammable liquid line into tire gaseous oxidant line prior to the gaseous oxidant nozzle. Non-flammable liquid present within the closed mixing vessel 5910 that is ejected via outlet 5903 may be suitable for this purpose, provided this non-flammable liquid passes through a filter to remove particulate matter prior to introduction into the gaseous oxidant line. In one embodiment, the gaseous oxidant is diluted with non-flammable liquid prior to introduction into the closed mixing vessel 5910.

[1314] The choice of gas mixer and associated design of the closed mixing vessel should consider the factors discussed above. The gas mixer must allow for a residence time that allows complete, or near complete, mixing to create a homogeneous composition of gas where there are no potentially unsaid pockets of gas with undesirable ratios of hydrocarbon to oxygen.

[1315] The final consideration is the removal of the mixed gas from the top of the closed mixing vessel, which can be accomplished with any variety of means for removal well known in the art

[1316] Oxidative Dehy drogenation

[1317] Oxidative dehydrogenation of paraffins to olefins is an alternative to the costly energy intensive and environmentally unfriendly thermal cracking method currently used. In ODH, a stream of one or more alkanes are passed over a catalyst in the presence of oxygen, to produce corresponding olefins and a variety of byproducts that can be removed in downstream processing steps. Since in ODH the conversion of paraffins to olefins is assisted by a catalyst the required operating temperatures are significantly lower than the temperature required for thermal cracking. In addition, for conversion of ethane to ethy lene, under some conditions, ODH provides for higher conversion and selectivity rates. Despite these advantages ODH is not employed commercially due to the risk of thermal runaway of the reaction and consequential explosion. This risk is due to the requirement for mixing a hy drocarbon containing gas with oxygen or a gaseous oxidant.

[1318] Provided herein is a process for the oxidative dehydrogenation of a paraffin to a corresponding olefin. More specifically, provided herein is a process for oxidative dehydrogenation of ethane into ethylene comprising mixing of ethane and oxygen in a ratio that falls outside of the flammability envelope in a closed mixing vessel, passing the mixture of ethane and oxygen through a heat exchanger to raise the temperature to at least 200 °C, introducing the heated mixture into an ODH reactor containing an ODH catalyst to produce ethylene, carbon monoxide, carbon dioxide, water, acetic acid minimal hydrocarbons, and, possibly, unreacted O2. The residual products are directed through a quench tower to remove water and acetic acid, then through downstream processing units, such as a catalytic bed to remove unreacted oxygen, an amine wash to remove carbon dioxide, and a demethanizer to remove methane.

[1319] By using the gas mixer and method of mixing a hydrocarbon containing gas and a gaseous oxidant discussed above the inherent risks of catalytic ODH are minimized. The mix of ethane and oxygen entering the reactor is outside the flammability envelope so that thermal runaway and subsequent explosion is not likely. Furthermore, by premixing the gases a user can ensure consistent conv ersion due to the homogeneous nature of the ethane, oxygen mix.

[1320] Examples

[1321] In a mixture of ethane and oxygen, at 25°C, the upper explosive limit (UEL), defined by a ratio in mole %, of ethane to oxygen, is approximately 62.18 to 32.81 — where the balance is in the form of vaporous water— -or 1 90 to 1. A model of a theoretical mixer where temperature and pressure are independently controllable was used to predict that a ratio of ethane to oxygen at the UEL can be maintained over a range of temperature and pressure. At each temperature and pressure there is a fraction that includes water that leaves the liquid phase and enters the vapor phase. By controlling temperature and pressure the carryover of vaporous water can be manipulated. Table N! shows the fraction of the mixture that is water, ethane and oxy gen, at various temperatures and pressure. As the temperature and pressure are increased the fraction of water within the gaseous phase decreases. Furthermore, at the temperature and pressures listed the UEL is not breached i.e., the mole % of oxygen does not exceed 62.81.

113221 TABLE Nl: Calculated Mole Fraction of Ethane, Oxygen, and Water as a Function Of The Temperature And Pressure

[ 1323] Techniques are provided for oxidative dehydrogenation (ODH) of lower alkanes into corresponding alkenes. See, e.g. FIGS. 2-3 and associated description. More specifically, techniques are provided for a chemical complex for ODH that includes an oxygen separation module. The present disclosure seeks to further improve economic efficiency by reducing the amount of costly oxygen required during steady state operations. Furthermore, the use of ODH by-products to drive oxygen separation limits the amount of carbon dioxide that would be released into the atmosphere.

[1324] Techniques are provided an oxidative dehydrogenation reaction complex with oxygen separation integration that is economically feasible and amenable to large scale commercial production of olefins from lower alkanes.

[1325] Techniques are provided for a chemical complex comprising, in cooperative arrangement: at least one ODH reactor, comprising an ODH catalyst, for accepting a lower alkane and oxygen, and optionally an inert diluent, and for releasing a product stream comprising unconverted lower alkane, a corresponding aikene, and possibly (optionally) carbon oxides, oxygenates, and water, a quench tower for removing oxygenates and water from the product stream; an amine wash tower for removing carbon dioxide from the product stream; a drier for removing residual water; a distillation tower for removing C2/C2 ₊ hydrocarbons from the product stream to produce an overhead stream enriched with C _:i hydrocarbons; and an oxygen separation module for accepting atmospheric air and die overhead stream to produce an oxygen enriched stream that can iced the ODH reactor.

[1326] In some embodiments, the chemical complex further comprises a flooded gas mixer for premixing the oxy gen and the lower alkane prior to introduction into the oxidative dehydrogenation reactor.

[1327] In some embodiments, the chemical complex further comprises a heat exchanger immediately downstream of each at least one ODH reactor and upstream of the quench tower.

[13281 In some embodiments, the chemical complex further comprises a caustic wash tower immediately downstream of the amine wash.

[1329] In some embodiments, the oxidative deh drogenation reactor uses a catalyst selected from the group consisting of:(i) catalysts of the formula: Mo _a V _b Te _c NbaPd _e O _f, where a, b, c, d, e and f are the relative atomic amounts of die elements Mo. V, Te, Nb, Pd and O, respectively, and when a = 1 (normalized), b = 0.01 to 1.0, c = 0.0 to 1.0, d = 0 to 1.0, 0.00 < e < 0.10 and f is dependent on the oxidation state of the other elements, i.e., f is a number to satisfy the valence state of the catalyst; (it) cataly sts of the formula: Ni _g A _h BjP _j <¼ where g is a number from 0.1 to 0.9, from 0.3 to 0.9, from 0.5 to 0.85, or from 0.6 to 0.8, h is a number from 0.04 to 0.9, i is a number from 0 to 0.5, j is a number from 0 to 0.5 f is a number to satisfy the valence state of the cataly st, A is selected from the group consisting of Ti, Ta, V, Nb, Hf, W, Y, Zn, Zr, Si and A1 or mixtures thereof, B is selected from the group consisting of La, Ce, Pr, Nd, Sm, Sb, Sn, Bi Pb, Tl, In, Te, Cr, Mn Mo, Fe, Co, Cu Ru, Rh, Pd, Pt, Ag, Cd, Os, If, Au Hg and mixtures thereof, D is selected from the group consisting of Ca, K, Mg, Li Na, Sr, Ba, Cs, and Rb and mixtures thereof, and O is oxygen; (iii) catalysts of the formula: Mo _a E _k Gi<¾ where E is selected from the group consisting of Ba, Ca Cr, Mn, Nb, Ta, Ti, Te V, W and mixtures thereof, G is selected from the group consisting of Bi, Ce Co, Cu, Fe, K, Mg, V, Ni, P Pb, Sb, Si, Sn, Ti, U, and mixtures thereof, and when a = I (normalized), k is 0 to 2, 1 = 0 to 2, with the proviso that the total value of 1 for Co, Ni, Fe and mixtures thereof is less tlian 0.5, and f is a number to satisfy' the valence state of the catalyst; fiv) catalysts of the formula: V _m Mo _n Nb ₀ Te _p Me _q O _{f ,} where Me is a metal selected from the group consisting of Ta, Ti W, Hf, Zr, Sb and mixtures thereof, m is from 0. 1 to 3 ; n is from 0.5 to 1.5, 0 is from 0.001 to 3, p is from 0.001 to 5, q is from 0 to 2, and f is a number to satisfy the valence state of the catalyst; andtv) catalysts of the formula: Mo _a V _r X _s Y _| Z„M _v O _f . where X is at least one of Nb and Ta, Y is at least one of Sb and Ni; Z is at least one of Te, Ga, Pd, W, Bi and Al, M is at least one of Fe, Co, Cu, Cr, Ti, Ce, Zr, Mn, Pb, Mg, Sn, Pt, Si, La, K, Ag and In, and w'hen a = 1.0 (normalized), r = 0.05 to 1.0, s = 0.001 to 1.0, t = 0.001 to 1.0, u = 0.001 to 0.5; v = 0.001 to 0.3, and f is a number to satisfy tire valence state of the catalyst.

[1330] In some embodiments, the ODH catalyst is supported.

[1331 ] In some embodiments, the ODH reactor comprises a fixed bed reactor.

[1332] In some embodiments, the ODH reactor comprises a fluidized bed reactor.

[1333] In some embodiments, the ODH reactor comprises an outer shell and one or more internal ceramic tubes defining a separate flow' passage for ethane dowui the interior of the tubes and an annular passage between the external shell of the reactor and the ceramic tubes defining a flow' path for an oxygen containing gas. f 1334] In some embodiments, the ceramic tubes further comprise an internal steel mesh and an external steel mesh.

[1335] In some embodiments, the C2/C2 ₊ hydrocarbons fraction leaving the distillation tower is directed to a splitter.

[1336] in some embodiments, the distillation tower is capable of separating the C2/C2 ₊ hydrocarbons fraction into lower alkane and its corresponding alkene, where the corresponding alkene may be withdrawn from the distillation tower through a side outlet and the lower alkane may be withdrawn through a bottom outlet and directed back to Site QDH reactor.

[1337] In some embodiments, the oxygen separation module is tubular and the oxygen transport membrane comprises an inner tube that is within an outer shell wherein a retentate side comprises the annular space between the inner tube and outer shell and a permeate side is the space within the inner tube.

[1338] In some embodiments, the chemical complex further comprises a fuel enhancement line upstream of the oxygen separation module for addition of supplemental combustible feedstock into the overhead stream feeding the oxygen separation module.

[1339] In some embodiments, the oxygen separation module further comprises means for heating the oxygen separation module to a temperature not less than 850°C.

[1340] In some embodiments, the chemical complex further comprises means for heating the overhead stream prior to entry into the oxygen separation module.

[13411 In some embodiments, the ODH reactor comprises more than one ODH reactor, each containing the same or different ODH catalyst, connected in series wherein the product stream from each reactor, except the last ODH reactor, is fed into a downstream ODH reactor

[1342] Techniques are provided for a process for ODH using the chemical complex described above.

[1343] In some embodiments, there is provided a process for the oxidative dehydrogenation of a lower alkane into a corresponding alkene using the chemical complex described previously, the process comprising: i) contacting, in one or more oxidative dehydrogenation reactors and optionally in the presence of an inert diluent, an oxygen containing gas and a lower alkane containing gas with an oxidative dehydrogenation catalyst to produce a product stream comprising unreacted lower alkane and its corresponding alkene, inert diluent, unreacted oxygen, and possibly (optionally) one or more of: carbon oxides, including carbon dioxide and carbon monoxide, oxygenates, including but not limited to, one or more of acetic acid and maleic acid, and water; ii) passing the product stream from i) through a quench tower to remove oxygenates and water from the product stream; iii) passing the product stream from ii) through an amine wash to remove carbon dioxide from the product stream; iv) passing the product stream from iii) through a drier to remove residual water; v) passing the product stream from iv) through a distillation tower to separate an overhead stream comprising Cl hydrocarbons and inert diluent from C2/C2+ hydrocarbons; vi) passing tire overhead stream from v) and optionally a combustible fuel into one or both of the retentate side and the permeate side of an oxygen separation module; vii) introducing air into the retentate side resulting in combustion of combustible Cl hydrocarbons and/or the combustible fuel generating heat to raise temperature of oxygen transport membrane to at least 850°C permitting unreacted oxygen from tire air to cross the oxygen transpost membrane into permeate side producing oxygen enriched gas on permeate side and leaving oxygen depleted air on the retentate side; viii) expelling the oxygen depleted air and combustion products through exhaust of the oxygen separation module; ix) removing oxygen enriched gas and combustion products through the outlet of the oxygen separation module; and x) directing the oxygen enriched gas back to i) as, or part of, the oxygen containing gas introduced into the one or more ODH reactors.

[1344J In some embodiments, the process further comprises that the oxygen enriched gas leaving the oxygen separation module comprises at least 50 vol. % oxygen. il345j In some embodiments, the process further comprises that the oxygen enriched gas leaving the oxygen separation module comprises less than 5 vol. % inert diluent.

[1346] In some embodiments, the process further comprises passing the product stream from i) through a heat exchanger before passing through the quench tower in ii).

[1347] In some embodiments, the process further comprises passing the product stream from iii) through a caustic wash before passing through the drier in iv).

[1348] In some embodiments, the process further comprises adding a low pH compound to the quench tower during remo val of water and oxygenates from the product stream from i).

[1349] In one aspect, there is a chemical complex useful for ODH, and, in another aspect there is described a process for ODH that may be performed in the chemical complex outlined in the first aspect. Lower alkanes are intended to include saturated hydrocarbons with from 2 to 6 carbons, and the corresponding alkene includes hydrocarbons with the same number of carbons, but with a single double carbon to carbon bond. For ethane, ethylene is its corresponding alkene.

[1350] Various tools commonly used for chemical reactors, including flowmeters, compressors valves and sensors for measuring parameters such as temperature and pressure can be used. It is expected that the person of ordinary skill in the art would include these components as deemed necessary for operation or for compliance with legal obligations related to safety' regulations.

[1351] ODH Reactor

[1352] Any of the known reactor types applicable for the ODH of hydrocarbons can be used. Particularly suited for use are conventional fixed bed reactors. In a typical fixed bed reactor, reactants are introduced into the reactor at one end, flow past an immobilized catalyst, products are formed and leave at the other end of the reactor. Designing a suitable fixed bed reactor may follow techniques known for reactors of this type. A person skilled in the art would know ⁷ which features are required with respect to shape and dimensions, inputs for reactants, outputs for products, temperature and pressure control, and means for immobilizing the catalyst.

[i 353] In some embodiments, the ODH reactor comprises a fixed bed reactor.

[1354] A fluidized bed reactor can be used. These types of reactors are also well known. Typically, the catalyst is supported by a porous structure, or distributor plate, located near a bottom end of the reactor, and reactants flow through at a velocity sufficient to fluidize the bed (e.g., the catalyst rises and begins to swirl around in a fluidized maimer). The reactants are converted to products upon contact with tire fluidized catalyst and subsequently removed from the upper end of the reactor. Design considerations include shape of the reactor and distributor plate, input and output, and temperature and pressure control, all of which would fall under knowledge of the person skilled in the art

[1355] In some embodiments, the ODH reactor comprises a fluidized bed reactor.

[1356] Multiple ODH reactors, either in series or in parallel, can be used. A swing bed type reactor is also envisioned in an embodiment. In tins instance, parallel beds are alternatively exposed to a hydrocarbon feed comprising mainly hydrocarbons with optional residual oxygen, or an oxygen feed that is hydrocarbon free. The oxy gen feed is directed to one reactor to re-oxidize a spent catalyst while simultaneously the hydrocarbon feed is passed through the other bed containing active oxidized catalyst allowing ODH to occur. A valve configuration allows swinging the oxygen and hydrocarbon feeds between the two beds to regenerate the oxidized catal st in one bed while ODH is occurring in the other bed. Use of multiple reactors, including ODH reactors in either a parallel, series, or swing bed type arrangement is well known in the art.

[1357] In some embodiments, the ODH reactor comprises multiple inlets for introduction of an oxygen containing gas. In this embodiment, oxygen addition is distributed in a staged manner throughout the reactor, limiting peak temperature increases by leveling oxygen concentration through the height or length of the reactor. [1358] In some embodiments, the ODH reactor comprises an outer shell and one or more internal ceramic tubes defining a separate flow passage for ethane down the interior of the tubes and an annular passage between the external shell of the reactor and the ceramic tubes defining a flow path for an oxygen containing gas.

[1359] In some embodiments, the ceramic tubes further comprise an internal steel mesh and an external steel mesh.

[1360] By-product removal

[1361] Oxidative dehydrogenation of alkanes inevitably produces not only corresponding alkenes, but other by-products as well. Depending on the conditions, including the catalyst type, the levels of by-products present downstream can range from minimal (less than 2 vol. %), to significant (greater than 2 vol. %). Even at minimal levels by-products are undesirable as they may interfere with downstream applications where the produced alkene is utilized. For ODH of lower alkanes, for example, ethane, the most common by-products include carbon oxides, including carbon monoxide and carbon dioxide, oxygenates, and water.

[1362] Typically, the separation of oxygenates and water from an ODH reactor product stream is achieved by using a quench tower. In the present disclosure, the term “oxygenates” refers to by-products of the oxidative dehydrogenation process that contain carbon, hydrogen, and oxygen, and include, but are not limited to, ethanol, acetaldehyde, acetic acid, acrylic acid, and maleic acid. While the primary' purpose of a quench tower is the cooling of a gaseous product stream, there is a secondary benefit for tire purposes of the provided techniques. Cooling of the gaseous product line after leaving the reactor promotes condensation of water and oxygenates which can then be separated from tire components that remain in the gaseous phase, namely tire lower alkane, its corresponding alkene, and any carbon oxides. Some quench towers involve the spraying of water, or other liquid in which oxygenates me soluble, from the top of the tower onto the product stream entering from the bottom of the tower. Contact with water promotes cooling and ultimately condensation of the heavier components slated for removal. f 1363] In some embodiments, a product stream containing unconverted alkane, corresponding alkene, residual oxygen and by-products are passed through a quench tower to remove water and oxy genates. The remainder is passed on for the next step of purifica tion. Techniques of this nature have been thoroughly developed and are commonplace in the prior art. The person skilled in the ait would understand how to integrate a quench tower into the chemical complex described in this disclosure.

[1364] Multiple quench towers can be used. Where multiple ODH reactors are employed, it is preferred that each ODH reactor is followed by a quench tower, especially in instances where the reactors are in series. In this setting, oxygenates and water are removed before the remainder, optionally supplemented with additional oxygen, is passed on to the next ODH reactor in the series. In a parallel arrangement, the product streams fro t the parallel reactors may be combined before introduction into a quench tower.

[1365] Another common and well known separation method applicable for use with the provided techniques is the use of alkylamines, referred to herein as amines, in a scrubber to remove carbon dioxide front gaseous compositions. Carbon dioxide present in a gas is absorbed by aqueous amine solution which can then be separated from the remaining gaseous components. The amine is stripped of carbon dioxide by heating above 100°C and recycled to continue the process, while water from the stripper vapor is condensed leaving relatively pure carbon dioxide. The carbon dioxide, highly concentrated, can be captured and sold, or, alternatively it can be recycled back to act as an Inert diluent for the lower alkane and oxygen containing gases when introduced into the ODH reactor. This is one advantage of the provided techniques. Carbon dioxide produced in the process can be captured instead of being flared where it contributes to greenhouse gas emissions. This becomes more relevant with the addition of the oxygen separator which also produces carbon dioxide.

[1366] Amine scrubbing lias been used, especially in the petrochemical industry, for over sixty' years. There is a plethora of prior art and common knowledge that is available to the person skilled in the art for design and operation of an amine scrubber (referred to as amine wash tower) for use with the provided techniques.

Consideration of the type of amines used in the process requires special attention. Amines used vary in their ability' to remove oxygen and in their tendency to promote the formation of degradation products. For example, monoethanolamine (MEA) is commonly used and is capable of removing a high percentage of carbon dioxide even at low concentrations, but can also react with the carbon dioxide to form degradation products. Tins results in lower carbon dioxide capture and a reduction of available amines for subsequent absorption cycles.

[1367] The stream leaving tire amine wash tower comprises unconverted lower alkane, corresponding alkene, and carbon monoxide, and possibly methane as a contaminant present in the original hydrocarbon feedstock and/or produced in the reactor inert diluent other than carbon dioxide, if used, may also be present in the stream leaving the amine wash tower. The stream leaving the amine wash tower will also likely contain water — carryover from the amine wash tower — that should be removed via a dryer prior to directing the stream to a distillation tower. This is essential when cryogenic distillation is emplo ed as any water present in the stream may freeze in tire distillation tower, causing problems related to plugging and fouling of the tower. Dehydration of gaseous compositions using a dryer is well known in the art. Methods include but are not limited to, absorption using a sorbent such as triethyleneglycol (TEG), adsorption with at least two solid desiccant containing adsorption beds, and condensation. The product stream can contain less than 50 ppm of water, Jess than 25 ppm of water, or Jess than 10 ppm of water before being passed on to the next stage.

[1368] After removal of water, the product stream can be further separated into an overhead stream and a C2/C2 ₊ hydrocarbons stream using a distillation tower. The overhead steam comprises mainly Ci hydrocarbons (hydrocarbons with only one carbon), comprising mostly carbon monoxide but with the possibility of smaller amounts of methane, and inert diluent if used. The C2/C2 ₊ hydrocarbons stream would comprise the unconverted lower alkane and its corresponding alkene, and any additional hydrocarbons (hydrocarbons containing 2 or more carbons), that were present as impurities in the original hydrocarbon feedstock added to the ODH reactor. Using a distillation tower for separation of Ci hydrocarbons and C2/C2 ₊ hydrocarbons is well known in the art, and employs heating and cooling of gases in the presence of trays which capture condensed species. The spacing and number of trays dictate the degree of separation.

[1369] The distillation tower can comprise an upper outlet for removal of the overhead stream and a lower outlet for removal of the remainder, including the higher weight C2/C2 ₊ hydrocarbons. The overhead stream is directed toward the nest step in the disclosed chemical complex, i.e. the oxygen separation module. The C2/C2- hydrocarbons can then be directed to a C ₂₊ splitter to separate the lower alkane from its corresponding alkene. The lower alkane can be fed back to the ODH reactor, and the corresponding alkene, the target product, can be captured and employed for use in a variety of applications that depend on the nature of the alkene. For example, if the desired product is ethylene then use in synthesis of polyethylene would be appropriate.

[1370] As mentioned, the degree of separation capable within a distillation tower is dependent upon the number of trays within the unit. The most common method involves cryogenic distillation so the nature of the species targeted for separation and their relative volatilities plays a role. For example, the relative volatility' of ethylene to ethane is quite small. As a result, a tower designed to separate the two species would need to be tall and include a large number of trays. The difference in relative volatilities between C2/C2 ₊ hydrocarbons and Ci hydrocarbons is significant enough that a smaller tower with fewer trays would suffice. A person skilled in the art would understand from this relationship that a smaller tower would be sufficient to separate out carbon monoxide and methane (Ci hydrocarbons), from the unconverted lower alkane and its corresponding alkene. However if separation of the lower alkane with the corresponding alkene is also desired then a much larger tower would be needed. In that case, the tower would include another outlet, or side out where the corresponding alkene may be withdrawn from the distillation tower. Also contemplated is the separation of the lower alkane and corresponding alkene in a separate unit, after removal of the lower alkane and corresponding alkene from the distillation tower. Specifically, a splitter, which is well known in the art, may be used. In some embodiments, the stream of C2/C2 ₊ hydrocarbons leaving the distillation tower is directed into a splitter.

[1371 ] In some embodiments, the distillation tower comprises an outlet for removal of the overhead stream and an outlet for removal of the C2/C2 ₊ hydrocarbons stream. In some embodiments, the distillation tower comprises a side outlet for removal of alkenes.

[1372] Oxygen Separation Module f 1373] An embodiment of the oxygen separation module comprises a sealed vessel with two compartments, separated by a temperature dependent oxy gen transport membrane, as shown in Figures 63 A, 63B, 63C, and 63D. The two compartments are the retentate side 6317 and the permeate side 6318. That the membrane is temperature dependent means that at a critical temperature the membrane will selectively allow oxygen to pass through from one side to the other. The oxygen separation module also comprises at least two inlets, air input 6320 for introducing atmospheric air into the retentate side 6317 and the o ther for introducing overhead stream 6316 into either of the retentate side 6317 or the permeate side 6318, or both retentate side 6317 and permeate side 6318. Finally, there are two outputs from the oxygen separation module. There is exhaust 6321 for removal of oxygen depleted air and combustion products from the retentate side, and an outlet for removal of oxygen enriched gas and possibly combustion products from the permeate side into oxygen enriched bottom line 6322. The oxygen emiched gas, and possibly combustion products, may be recycled back as or part of the oxygen containing gas introduced into the ODH reactor.

[1374] In some embodiments, the oxygen transport membrane 6319 is a tube and fits inside a larger tube 6327 which forms the outer wall of oxygen separation module 6306. The annular space between the larger tube 6327 and oxygen transport membrane 6319 corresponds to the retentate side, while the space within oxygen transport membrane 6319 corresponds to the permeate side. Material suitable for construction of the outer wall 6327 include those resistant to temperatures that exceed 850°C and approach 1000°C, selection of which falls within the knowledge of the skilled worker. Optionally, a flow controlling means 6326 (FIG. 63D) may be included that allows for flow into both sides (retentate side and permeate side) at varying levels. For example an operator may operate the flow controlling means 6326 to choose what portion of the flow from overhead stream 6316 enters retentate side 6317 and what portion enters permeate side 6318. Depending upon conditions, an operator may switch between the two sides, allow equivalent amounts to enter each side or bias the amount directed to one of the two sides.

[1375] The present disclosure contemplates the inlet for the overhead stream entering the oxygen separation module 6306 into either of the permeate side (Fig. 63 A) or the retentate side (Fig. 63C). Tins disclosure also contemplates the use of a valve for switching between directing the overhead stream to the retentate side or the permeate side. Tins would allow an operator to choose which of the sides, permeate or retentate, that the overhead stream is directed to.

[1376] The overhead stream can be introduced into both the retentate side and permeate side simultaneously. This includes the ability to alter the relative amount of overhead stream which is entered into each side. For example, an operator may choose to permit 80% of the overhead stream to enter into the retentate side and only 20% to the permeate side, or vice versa. To be clear, the amount of the overhead stream that enters either side, permeate or retentate, can range from 0% to 100%, with the fraction for each side totaling 100%. Precision valves that can control the flow sent to either side are well known in die art, and include, without limitation, solenoid valves, ball valves, or a combination of a backpressure needle valve and solenoid valve.

[1377] The oxygen transport membrane component of the oxygen separation module selectively allows passage of oxygen when the membrane reaches a critical temperature. Membranes of this nature are known. Specifically, a Mixed Ionic-Electronic Conducting (MIEC) membrane can be used. Movement of oxygen across die membrane is driven by an oxygen partial pressure gradient, moving from the high oxygen partial pressure side to the low oxygen partial pressure side. To get the oxygen to move to the permeate side, a skilled operator would understand that the partial pressure of oxygen on the retentate side would need to be increased to the point where it equals or exceeds the partial pressure of oxygen on the permeate side. For example, if oxygen on the permeate side is close to 100% of the volume at a pressure of the 3 atm, then the pressure on the retentate side would need to be increased to at least 5 atm when atmospheric air is added and contains approximately 21 % oxygen by volume. Alternatively, the pressure on the permeate side could be reduced to levels at or below 0.2 atm using a vacuum driven process.

[13781 Also contemplated in the design of the oxygen separation module is die ability to add a sweep gas, such as, steam or carbon dioxide, to the permeate side to dilute oxygen that crosses over from the retentate side. The effect of the sweep gas is die lowering of the oxygen partial pressure on die permeate side to drive diffusion of oxygen from die retentate side. A result of this configuration is a much lower percentage of oxygen within die oxygen enriched bottom line 6322, as it is diluted by the sweep gas. Theoretically the oxygen percentage could drop well below 10%. However, if water is the sweep gas, then a heat exchanger downstream of oxygen separation module 6306 can be used to remove the water following condensation, increasing the relative amount of oxygen in the fine. If carbon dioxide is used then an operator can determine the amount required to produce the desired oxygen level in the oxygen enriched bottom line 6322. By altering the amount of sweep gas an operator can control how much oxygen is present in the line as it leaves the oxygen separation module. A person skilled person in the art would understand this relationship and would be familiar with using a sweep gas and with using means for controlling the pressure in a sealed vessel such as, the type contemplated for the oxygen separation module described here.

[1379] It is well known that oxygen flux across the membrane is dependent upon the thickness of the membrane. A thin membrane allows oxygen to cross more quickly titan a thick membrane. A membrane comprised of a single layer, or monolithic type membrane, may be reduced in thicknesses in the range of 0.1 to 0 2 mM to allow' greater oxygen flux. However, these thicknesses are not practical due to susceptibility to mechanical instability. If a monolithic membrane is to be used, thicknesses below' 0.2 mm are not recommended. Other known membrane configurations include asymmetric membranes where a very thin conducting layer is supported on both sides by a porous structure. This allows a user to employ very thin membranes that allow' higher oxygen flux without sacrificing stability it is not essential to use any particular membrane structure provided the oxygen flux across the membrane is sufficient. The oxygen transport membrane can have an oxygen flux within the range of 300 to 1500 l/hr*m ² , from 500 to 1300 l/hr*m ² , or from 700 to 1000 l/hr*m ² .

[1380] Theoretically, the oxygen transport membrane can reach 850°C due to the exothermic nature of combustion of the Ci hydrocarbons present in the overhead stream. However, in instances where the Ci hydrocarbons (as the sole source of feedstock for combustion) are insufficient to reach the required temperature, combustible fuel can be added to the oxygen separation module to include an independent means for heating the oxygen separation module, including the oxygen transport membrane. For instance a separate line may add a combustible fuel, for example, methane, either into the overhead stream before entering the oxygen separation module, or directly into the oxygen separation module. Alternatively a heat exchanger or other means may be employed to heat the module to the required temperature. It can be beneficial that when using a heat exchanger or other means for heating that heat is distributed evenly throughout the module. The overhead stream can be heated just upstream of the oxygen separation module.

[1381] During start-up of the chemical complex, the oxygen transport membrane may not be at the required temperature. As a result, oxygen from the injected air cannot pass into the permeate side. In this instance, it can be beneficial to direct the overhead stream solely into the retentate side so that combustion on that side can contribute to increasing the temperature of die oxygen transport membrane to the point where oxygen can cross. When at steady state and the temperature of die ox gen transport membrane exceeds 850°C, the overhead stream may be directed to either side because oxygen can freely pass and permit combustion such that heat is continuously generated. Alternatively during startup, other means, such as a heat exchanger, may be used to heat the membrane. [1382] ODH Process

[1383] Techniques are provided for use of the ODH reactor as described. For best results, die oxidative dehydrogenation of a lower alkane may be conducted at temperatures from 300 °C to 550 °C, from 300 °C to 500 °C, or from 350 °C to 450 °C, at pressures from 0.5 to 100 psig (3.447 to 689.47 kPag) or from 15 to 50 psig (103.4 to 344.73 kPag), and the residence time of the lower alkane in the reactor is typically from 0.002 to 30 seconds, or from 1 to 10 seconds. For some catalyst formulations described herein, different temperature regimes may be used, due to changes in stability.

[1384] The lower alkane containing gas is ideally of a purity greater than 95 vol. % (for example, 98 vol %). In some embodiments, the process includes the addition of an ethane containing gas of purity of 95 vol. % or 98 vol. %.

[1385] The process can have a selectivity' for the corresponding alkene (ethylene in the case of ethane ODH) of greater than 95 mol %, preferably, greater than 98% The gas hourly space velocity (GHSV) will he from 500 to 30000 h ¹ (for example, greater than 1000 h ¹ ). The space-time yield of corresponding alkene (productivity) in g/hour per kg of the catalyst is not less than 50, greater than 1500, greater than 3000, or greater than 3500 at 350 °C to 370 °C. It should be noted that the productivity of the catalyst will increase with increasing temperature until the selectivity is sacrificed.

[1386] When the lower alkane is ethane, the specificity of conversion to ethylene should be not less than 50 mol %, greater titan 90 mol %, or at least 95 mol %.

[1387] Safety of the process is a primary' concern. For that reason, mixtures of a lower alkane with oxygen should comprise ratios that fall outside of the flammability envelope. A ratio of alkane to oxygen may fall outside the upper flammability envelope. In this instance the percentage of oxygen in the mixture is not greater than 30 vol. %, not greater than 25 vol. %, or not greater titan 20 vol. %.

[1388] With higher oxygen percentages, it is imperative to choose alkane percentages that keep the mixture outside of Site flammability envelope. While a person skilled in tire art would be able to determine an appropriate level it is recommended that the percentage of oxygen not exceed 40 vol. %. For instances where the mixture of gases prior to ODH comprises 10 vol. % oxygen and 15 vol. % alkane, the balance must be made up with an inert diluent, such as nitrogen, carbon dioxide, or steam. The inert diluent should exist in the gaseous state in the conditions within the reactor and should not increase the flammability of the hydrocarbon added to the reactor, characteristics that a skilled worker would understand when deciding on which inert diluent to employ. Inert diluent can be added to either of the lower alkane containing gas or the oxygen containing gas prior to entering the ODH reactor or may be added directly into the ODH reactor.

[1389J While the volume ratios of lower alkane to oxygen that do not equal or approximate 3 : 3 can be used, it is preferable that the addition of each is close to 1:1. The reason for this is that the goal is for 100 mol % conversion, with minimal unreacted alkane and oxygen leaving the ODH reactor. When the components are added in an unbalanced ratio the presence of oxygen or unreacted alkane in the product stream is inevitable. In one embodiment the product stream leaving the ODH reactor contains less than 5 mol % unreacted lower alkane, preferably less than 2.5 mol %, most preferably less than 1 mol %. In another embodiment, the product stream leaving the ODH reactor contains less titan 2 mol % oxygen preferably less titan 1.5 mol % oxygen, most preferably less titan 1 mol % oxygen.

[1390] The ratio of oxygen to lower alkane added to the ODH reactor may also effect the composition and contribution of by-products to the product stream leaving the ODH reactor. Excess oxygen may oxidize the corresponding alkene to a carboxylic acid. For example, ethylene produced in the ODH reactor may be further oxidized to acetic acid. Depending upon the desired product this may be desirable. A skilled operator would understand ho w changing the ratio of added gases, in combination with ODH catalyst selection, alters the products present in the stream leaving the ODH reactor.

[1391] Removal of by-products such as oxygenates, for example acetic acid, is routine for operators skilled in these types of processes. The quench tower, which is primarily used to reduce the temperature of tire product stream, may be used to isolate oxygenates and water produced in the ODH reactor. The cooling of the product stream results in condensation of oxygenates at a much higher temperature than the dew point of the alkanes or the corresponding alkene gases. By taking advantage of this difference operators may capture the condensed products and allow the gaseous remains to move onto the next step in the separation of by-products from the product stream. Captured oxygenates may be used in other well-known downstream processes. For example, in ODH of ethane to ethylene, the ethylene may be further oxidized to acetic acid, which may be reacted with ethylene to produce vinyl acetate or other oxygenates.

[1392] Lo ' pH compounds can be added to the quench tower which has the effect of improving removal of oxygenates. In the absence of addition of low' pH compounds it is possible that not all oxygenates will undergo condensation within the quench tower in this case, any gaseous residual oxygenates may be passed on to the next stage. Addition of a low pH compound, such as sodium bicarbonate, may promote conversion of oxygenates into compounds with a higher dew point increasing die likelihood of condensation.

[1393] Removal of carbon dioxide from the product stream, in combination with the oxygen separation module, is one of the advantages. Carbon dioxide produced in the oxygen separation module, due to combustion on the permeate side of the oxygen transport membrane, can be captured instead of being released to the atmosphere. The oxygen enriched gas and associated combustion products that are recycled back re-enter the chemical complex so that any carbon dioxide present can be isolated in the amine wash. Furthermore, the carbon dioxide isolated by the amine wash can be recycled back to the ODH reactor, where it can be used as the inert diluent.

[1394] While ODH does not produce significant amounts of carbon dioxide, it does produce carbon monoxide, which ordinarily would be flared into the atmosphere when the opportunity to convert the carbon monoxide to value added chemicals is not feasible at the manufacturing site. The techniques provided here can allow for the combustion of the carbon monoxide in a s stem that captures the resulting carbon dioxide and shuttles it back through the ODH chemical complex where it can be captured.

[1395] It should be noted that, theoretically, removal of oxygenates and carbon dioxide prior to oxygen separation is not absolutely essential. It is conceivable to pass the product stream from the ODH reactor directly to an oxygen separation module. However, in this instance the target alkene would be subjected to combustion and lost, defeating the purpose of the ODH reaction. It is necessary to separate the target alkene prior to ox gen separation. The techniques provided can include separation of unconverted alkane and corresponding alkene from the lighter Ci h drocarbons using a cryogenic distillation process. The presence of oxygenates, such as acetic acid, and carbon dioxide would severely impact the function of a cjyogenic distillation process. For this reason, the removal of oxygenates and carbon dioxide should be considered when implementing the provided techniques.

[1396] The amine wash results in additio n of water into the product stream, which must be removed prior to distillation. As previously discussed dehydration of gaseous compositions falls within the common general knowledge of those skilled in the art.

[1397] Distillation of gaseous products and separation of components is also well known in the art. The skilled worker would know how to use a distillation tower to separate Ci hydrocarbons from C2/C2 ₊ hydrocarbons. [1398] The process of ODH as it relates to oxygen separation may vary, mostly dependent upon the temperature of the oxygen transpost membrane. When the oxygen transport membrane is below the temperature at which oxygen can selectively pass through the preferred embodiment is to direct the overhead stream into the retentate side, where atmospheric air is introduced. In tins situation the oxygen within the air is present for the combustion of the Ci hydrocarbons present in the overhead stream. An operator must judge on whether the degree to which tins combustion raises the temperature of the oxygen transport membrane is significant enough for selective oxygen transport to occur. If it is insufficient, meaning the temperature does not surpass 850°C, regardless of the amount of Ci hydrocarbon gas flow'ing into the module, then additional combustible fuel may be added. For example, adding methane to the overhead stream may be sufficient to reach the required temperature.

[1399] Provided enough combustion is occurring with addition of combustible fuel and the temperature of the membrane is above 850°C then the combustible fuel or the overhead stream may be directed into the permeate side. The reason this is possible is that since the membrane is hot enough, oxygen can pass through and act on the C i hydrocarbons present in the overhead stream and added to the permeate side, releasing heat so as to maintain the membrane in an oxygen transport able mode. Where the overhead stream is directed to depends on the desired degree of oxygen separation. When directed to the retentate side, combustion results in production of water and carbon dioxide, which cannot pass through and are therefore ejected through the exhaust. In this inode, it is not possible to capture the carbon dioxide produced in the chemical complex described. There are other modes for capture that may he involved, but are not integrated into the ODH chemical complex. The oxygen that passes in this configuration is unaccompanied by the combustion products and therefore is of very high purity. In some embodiments, the overhead stream is directed to the retentate side and the oxygen enriched stream comprises at least 95 vol. % oxygen, preferably 98 vol. % oxygen, most preferably 99 vol. % oxygen.

[1400] In the alternative, the overhead stream may be directed into the permeate side. In this setting the oxy gen transport membrane must be at the required temperature in this case the Ci hydrocarbons within the overhead stream and added to the permeate side are subjected to combustion with the oxygen crossing the membrane. Any unreacted ox gen and the combustion products are mixed before leaving. As a result the oxygen is diluted and die oxygen enriched stream contains a lower degree of ox gen. The degree of oxygen dilution may also be significantly increased when a sweep gas is employed, even approaching levels below 10 vol. %. In some embodiments, the overhead stream is directed to the permeate side and die ox gen enriched stream comprises at least 20 vol. % oxygen, preferably 55 vol. % oxygen, most preferably 90 vol. % oxygen, with the balance comprising carbon dioxide and water, and possibly inert diluent.

[1401] Optimization of the process requires an operator to understand that the side to which the overhead stream is directed will impact on the fate of carbon dioxide produced and the degree to which carbon dioxide contributes to the oxygen enriched gas directed back. Since carbon dioxide is a suitable inert diluent for dilution of the lower alkane and oxygen containing gases it is expected that an operator may adjust the ratio of the overhead stream entering into the retentate side relative to the permeate side so as to produce an oxygen enriched gas with a desired level of carbon dioxide. Ideally, the level will be adjusted so that when combined with carbon dioxide isolated by the amine wash the total amount will equal the amount required for dilution of the lo wer alkane and oxygen containing gases while at the same time minimizing the amount of carbon dioxide released into the atmosphere after ejection from the oxygen separation module exhaust.

[1402] In some embodiments, the entirety of the carbon dioxide isolated in the amine wash is recycled back as inert diluent and the ratio of the overhead stream entering the retentate side relative to the permeate side is altered to allow for production of oxygen enriched gas with a degree of carbon dioxide that when mixed with carbon dioxide from the amine wash falls within the levels required for a safe mixture with the lower alkane containing gas.

[ i 403] This provided techniques will further be described by reference to the following examples. The following examples are merely illustrative of the provided techniques and are not intended to be limiting. Unless otherwise indicated, all percentages are by weight.

[1404] Examples

[1405] A kinetic model was developed using Aspen Plus V8.6 software and used in simulations to predict production rates of various species following oxidative hydrogenation of ethane using a fixed bed reactor. The simulation data for various feed compositions and temperatnre/pressure conditions was compared to experimental data collected from ODH of ethane using a fixed bed reactor under the conditions described below.

[1406] ODH Reaction conditions

[1407] ODH of ethane was performed was performed using two bench scale fixed bed tubular reactors, in series constructed with SS316L stainless steel tubes having an optical diameter of 1.0” and a length of 34”. Each reactor was wrapped in an electrical heating jacket and sealed with ceramic insulating material. Temperature within the reactors was monitored using a 7 point thermocouple. The catalyst bed in each reactor comprised 150 g in total weight of catalyst having the formula MoV _0.40 Nb _{0 io} Te _0.14 O, with relative atomic amounts of each component, relative to a relative amount of Mo of 1, shown in subscript. The catal st bed was secured in place by packing quartz powder above and below the bed, with glass wool on the top and bottom of the reactor tube to prevent bed movement during experimental runs.

[1408] For experimental runs the reaction pressure was ~1 bar with flow through the reactor having a weight hourly space velocity (WHSV) of 0.68 hr ¹ . Three separate runs were completed with temperatures of 316°C (Run 01), 332°C (Run 02), and 336°C (Run 03). A feed composition of ethane, ethylene, oxygen, and carbon dioxide in weight percent (wt. %) of 43/0/22/35. premixed, was introduced into reactor 01. The mole fraction of each component produced in the reactors was measured downstream of reactor 02. The results for each experimental run were compared to results obtained with the same conditions using the kinetic model (see Table 01). There is excellent agreement between the experimental runs and the simulation, as seen by the R squared values which range from 0.9933 to 0.9961.

[1409] TABLE 01 : Comparison of experimental data (EXP) on fixed bed reactor unit with kinetic model simulation data (SIM)

[1410] Chemical complex simulation

[1411] Following confirmation that tire kinetic model provides reasonable estimates for the values of products produced in the ODH reactor, further simulations were conducted on a speculative chemical complex comparable to the provided techniques but without the final oxygen separation module. An ethane ODH simulation was performed with a target ethylene production rate of 20 kg/hr, using of a flow rate, in kg/hr, of ethane entering the ODH reactor 202 of approximately 23.38 kg hr, (stream 218) and of oxygen (stream 216) of approximately 16.72 kg/hr, diluted in a stream of carbon dioxide flowing at a rate of 28.61 kg/hr (also stream 216). The simulation took into account that ethane feedstock is frequently contaminated with trace amounts of propane and methane. Simulated results of the composition for various components within each stream are shown in Table 02, along with the temperature and pressure at the points where measurements were taken. For the purposes of clarity, trace elements were not included (e.g., hydroxide radicals). The results fall within the ranges expected under the simulated ODH conditions. The flow rate of carbon dioxide coining out of the amine wash tower 208 in carbon dioxide bottom outlet 230 of 29.63 kg/hr would be sufficient to provide enough inert diluent flowing into the ODH reactor 202.

[1412| TABLE 02: Stream properties at locations within the chemical complex

[1413] Oxygen separation module simulation

[1414] To ascertain whether the combustible Ci hydrocarbons present in overhead stream 238 coming off of the distillation tower 212 would be sufficient feedstock to raise the temperature of the oxygen transport membrane to the required temperature, an additional model was created using Aspen Plus V8.6. The Rstoic block in Aspen was used to simulate the adiabatic combustion process. Development of the model assumed that the oxygen transport membrane can permeate <¾ from the retentate side to the permeate side with 100% selectivity only when the membrane module operating temperature is > 850°C. The values of the various components present in overhead stream 238 from the previous simulation were used in the model (e.g. 0.20 kg/hr carbon monoxide and 0.35 kg/hr methane). Four scenarios were modeled, differing in whether additional combustible fuel was added via line 250 (Figures 2 and 3) and to which side of the membrane the overhead stream was directed. In Case OO, no additional combustible fuel was added to overhead stream 238, 100% of which was added to the retentate side. Case 01 - a flow rate of 2, 1 kg/hr of methane was added via line 250 to the overhead stream 238, which was directed to the retentate side. Case 02 - a flow of rate of 2.1 kg/hr of methane was added via line 250 to the overhead stream 238, which was directed to the permeate side. Case 03 - a flow rate of 2.1 kg/hr of methane was added via line 250 to the overhead stream 238, 50% of which was directed to the retentate side and 50% was direc ted to the permeate side. The simulation using these conditions was used to estimate the temperature within the oxygen separation module and the rates of flow for the various components coming out of the oxygen separation module via oxygen enriched bottom line 227, shown in Table 03.

[1415] TABLE 03: Simulation results for oxygen separation module

[1416] Results of the simulation show that under conditions where no methane is added via line 250 to overhead stream 238, there are not enough combustible Ci hydrocarbons present in the overhead stream 238 for raising the temperature within the oxygen separation module to above 850°C. The resulting temperature of 568°C corresponded with no 0 ₂ crossing, the membrane and ultimately leaving the oxygen separation module via line oxygen enriched bottom line 227. In Cases 01 -03, the addition of methane at a flow rate of 2.1 kg/hr via line 23 w¾s sufficient to raise the temperature of the module to 850°C. All three eases result in 17 kg/hr of 0 ₂ leaving the module, which exceeds the amount of 16.7 kg/hr entering ODH reactor 202 in the earlier simulation. Ou a final note, it is not surprising that Cases 02 and 03 provided for carbon dioxide iu oxygen enriched botom line 227, as those cases included sending the overhead stream into the permeate side. Carbon dioxide is produced by combustion of the methane added to the permeate side. The carbon dioxide present in oxygen enriched bottom line 227 can act as inert diluent in the reaction in the ODH reactor 202.

[1417] A process, a system, and an apparatus for separation of an oxygenate from a stream is provided. More specifically, a stream comprising the oxygenate is introduced to a quench tower along with a caustic outlet stream comprising a metal salt. Contact between the oxygenate and the metal sail results in conversion of a portion of the oxygenate into a derivative salt. The derivative salt and unconverted oxygenate are condensed by quenching and substantially removed from the quench tower as an oxygenate outlet stream. The gaseous components of the stream, minus a substantial portion of the oxygenate, are removed from the quench tower as a quench outlet gas stream. [1418] In one aspect, a method for separation of an oxygenate from a stream is provided. More specifically, the stream comprising the oxygenate is introduced to a quench tower along with a caustic outlet stream comprising a me tal salt. The streams are quenched with addition of water and contact between the oxygenate and the metal salt during quenching facilitates conversion of the oxygenate into a derivative salt. An oxygenate outlet stream comprising a substantial portion of the derivative salt and at least a substantial portion of unconverted oxygenate with steam is removed from the quench tower. A quench outlet gas stream, comprising gaseous components present in the stream, is also removed from the quench tower.

[1419] In yet another aspect, an apparatus for separation of an ox genate from a stream is provided. More specifically, the apparatus comprises a quench tower comprising a quench inlet, a quench outlet a metal salt inlet and an oxygenate outlet. The quench inlet is configured to receive the stream comprising the oxygenate. The metal salt inlet is configured to receive into the quench tow'er a caustic outlet stream comprising a metal salt, allowing contact of the caustic outlet stream with the stream. The quench outlet is suitable for removing a quench outlet stream and the oxygenate outlet is suitable for removing an oxygenate outlet gas stream comprising at least a substantial portion of a derivative salt formed by contact of the oxygenate with the metal salt and at least a substantial portion of the unconverted oxygenate.

[1420] In yet another aspect a system for separation of an oxygenate from a stream is provided. More specifically, the system comprises a quench tower configured to receive a stream comprising the oxygenate and a caustic outlet stream comprising a metal salt resulting in contact of the oxygenate with the metal salt and conversion of a portion of the oxygenate into a derivative salt, quench the stream and the caustic outlet stream, remove at least a substantial portion of a derivative sait and at least a substantial portion of the unconverted oxygenate, and remove an quench outlet gas stream comprising gaseous components of the stream.

[1421 ] In one aspect, a method is provided to convert a lower alkane to an alkene. More specifically, an input stream comprising oxygen and the lower alkane is introduced to an oxidative dehydrogenation (ODH) reactor. The input stream may include an inert diluent to control flammability. At least a portion of the lower alkane is converted to file alkene in the ODH reactor and an ODH outlet stream comprising the alkene, unconverted alkane, an oxygenate, a carbon-based oxide, and water or steam, is produced. The ODH outlet stream and a caustic outlet stream comprising a metal salt are in introduced to a quench tower and quenched. Contact in the quench tower between the oxygena te and the metal salt facilitates conversion of a portion of the oxygenate into a derivative salt. A quench outlet gas stream comprising at leas t a substantial portion of the alkene and at least a substantial portion of the carbon-based oxide with unconverted alkane is removed from the quench tower, as is an ox genate outlet stream comprising at least a substantial portion of the unconverted oxygenate and at least a substantial portion of die derivative salt with water. The quench outlet gas stream is introduced to a caustic wash tower and contacted with a caustic agent in the caustic wash tower to form a metal salt that is removed from the caustic tower and may recycled and used as part of the caustic outlet stream introduced into die quench tower with the ODH oudet stream. f 1422] In another aspect, an apparatus is provided for oxidative dehydrogenation (ODH) of a lower alkane to an alkene. More specifically, the apparatus comprises an ODH reactor, a quench tower, a caustic wash tower, and a return line. The ODH reactor comprises an ODH inlet and an ODH outlet. The ODH inlet is suitable for transporting an ODH inlet stream comprising the lower aikane, oxygen, and optionally feed diluent into the ODH reactor. The ODH outlet is suitable for transporting an ODH outlet stream comprising the alkene, unconverted alkane, an oxy genate, and a carbon-based oxide, water and optionally inert diluent. The quench tower comprises a quench inlet, a quench outlet, a metal salt inlet, and an oxygenate outlet. The quench inlet is in fluid communication with the ODH outlet to receive the ODH outlet stream. The quench outlet is suitable for transporting a quench outlet gas stream comprising at least a substantial portion of the alkene and at least a substantial portion of the carbon-based oxide with unconverted alkane and optionally inert diluent. The oxygenate outlet is suitable for transporting an oxygenate outlet stream comprising at least a substantial portion of the oxygenate and a derivative salt and water. The caustic wash tower comprises a wash inlet, a wash outlet, a caustic inlet, and a caustic outlet. The wash inlet is in fluid communication with the quench outlet to receive the quench outlet stream. The caustic outlet is suitable for transporting a caustic outlet stream comprising a metal salt. The return line is in fluid communication with the caustic outlet to receive the caustic outlet stream and output the caustic outlet stream into the metal salt inlet of the quench tower.

[1423] In another aspect, a system is provided for oxidative dehy drogenation (ODH) of a lower alkane. More specifically, the system comprises an ODH reactor, a quench tower, a caustic wash tower, and a return line. The ODH reactor is configured to receive an input stream comprising oxygen and the lower aikane, an optionally inert diluent. The ODH reactor is configured to produce an ODH outlet stream comprising an alkene, unconverted aikane, an oxygenate, steam and a carbon-based oxide. The quench tower is configured to receive and quench the ODH outlet stream and a caustic outlet stream comprising a metal salt, contact the oxygenate with the metal salt to convert a portion of the oxygenate to a derivative salt, remove an oxygenate outlet stream comprising at least a substantial portion of the unconverted oxygenate and at least a substantial portion of the derivative salt and produce a quench outlet gas stream comprising at least a substantial portion of the alkene, unconverted alkane and at least a substantial portion of the carbon-based oxide. The caustic wash tower is configured to receive the quench outlet gas stream and contact a substantial portion of the carbon-based oxide from the quench outlet stream with a caustic agent to form a caustic outlet stream comprising a metal salt. The return line is configured to direct the caustic outlet stream into the quench tower and contact the caustic outlet stream with the ODH outlet stream to form a derivative salt from the metal salt and the oxygenate. The oxygenate outlet stream comprises a substantial portion of the derivative salt.

[1424] Many chemical production processes can have a co-product of an oxygenate such as, for example, ethanol, acetic acid, acrylic acid, maleic acid, and maleic anhydride. A quench tower is typically used for removing the oxygenate from a process stream. In the quench tower a quenching agent can condense the oxygenate in the process stream while an unreacted hydrocarbon and a carbon-based oxide or a sulfide, and an inert diluent, such as nitrogen, can be in a gas state. This can enable separation of the condensed oxygenate from the gaseous components. In some quenching processes, the oxygenate can be diluted to a low concentration that may be insufficient for subsequent applications. f 1425] The oxygenate can require purification and/or further processing in order to generate a product sufficient for subsequent applications. For example, water may have to be removed from the oxygenate to increase the concentration of the oxygenate. Separation of the oxygenate from water can increase the complexity of a quench tower and/or a separation vessel due to the small thermal (e.g., boiling point) separation between the oxygenate and the water. In various examples, a mixture of oxygenate and water can be azeotropic. The separation vessel may employ a large column, a high quantity of stages, a high reflux ratio, and a high energy demand to separate an azeotropic mixture of oxygenate and water. i 1426] In the petrochemical industry . a process stream can be treated with a caustic agent in order to remove a contaminant. For example, during the processing of gasoline, kerosene, and liquified petroleum gas (LPG), sulfides and organic acids are removed by treatment with a caustic agent such as sodium hydroxide. In an ethane cracking process carbon dioxide can be removed using a caustic agent. The treatment can comprise reacting the caustic agent with the contaminate to form a different product which can be removed from the process stream. For example, reacting gaseous hydrogen sulfide with a solution of caustic sodium hydroxide can produce water and sodium hydrogen sulfide which can be removed in the liquid state with the water. In the case of an ethane cracker, carbon dioxide can be removed from the process stream by conversion of the carbon dioxide to sodium bicarbonate in the caustic tower.

[1427] Upon reacting the caustic agent with the contaminate, the caustic agent becomes consumed (e.g., spent). The spent caustic may be undesirable and may require disposal which can be costly and increase complexity of the chemical production process. For example, the spent caustic can be sold to pulp and paper manufacturers which may require hauling of the spent caustic to a different facility. Spent caustic can also be disposed by deep well injection, incineration, and/or neutralized by wet air oxidation. These disposal processes can require additional energy, cost, and complexity in the chemical production process.

[1428] Converting the spent caustic to a marketable product which can remove the oxygenate from the process stream can lower energy requirements, cost, and complexity of a chemical production process. Thus, a method, a system and an apparatus are provided which can enhance the purification of the oxygenate and reduce energy requirements for the purification. More specifically, a stream comprising the oxygenate can be introduced to a quench tower and the oxy genate can be removed from the stream. A caustic outlet stream comprising a metal salt can be introduced to the quench tower. The stream can be contacted with the caustic outlet stream to form a derivative salt from the metal salt and the oxygenate. A quench outlet stream can be produced in the quench tower and an oxygenate outlet stream comprising at least a substantial portion of the oxygenate and at least a substantial portion of the derivative salt can be produced in the quench tower.

[1429] Oxidative dehydrogenation (ODH) can couple the endothermic dehydrogenation of an alkane with the strongly exothermic oxidation of hydrogen. For example, ODH of an alkane can comprise contacting an alkane and oxygen in an ODH reactor with an ODH catalyst under reaction conditions (e.g., temperature, pressure, flow rate, etc.) that can promote oxidation of the alkane into the corresponding alkene. The corresponding alkene includes hydrocarbons with the same number of carbons as the alkane used in the ODH reactor, but with the addition of one carbon to carbon double bond. For example, utilizing ODH, ethane can be converted to ethylene, propane can be converted to propylene, and butane can be converted to butylene.

[1430] Any ODH catalyst known in the art can be suitable for use with the provided techniques. For example, an ODH catalyst containing a mixed metal oxide can be used. Additionally, reaction conditions can be controlled to adjust the selectivity and yield of the ODH reactor products. As known in the art, conditions will vary and can be optimized for a particular alkane, for a specific catalyst, a select product, and/or a particular inert diluent. A coproduct of an ODH reaction can be an oxygenate which may need to be removed from the process stream and the ODH process may generate spent caustic.

[ 14311 Thus, in various examples, a method, a system and an apparatus are provided for converting a lower alkane to an alkene. An input stream comprising oxygen and tire lower alkane, and optionally feed diluent can be introduced to an ODH reactor. At least a portion of the lower alkane can be converted to the alkene in the ODH reactor and an ODH outlet stream comprising the alkene, unconverted alkane and an ox genate, and a carbon-based oxide can be produced. The ODH outlet stream can be introduced to a quench tower and Site oxygenate can be removed from the ODH outlet stream. A quench outlet stream compri sing at least a substantial portion of the alkene and at least a substantial portion of the carbon-based oxide can be produced in the quench tower. Additionally, an oxygenate outlet stream comprising at least a substantial portion of the oxygenate can be produced in the quench tower. The quench outlet stream can be introduced to a caustic wash tower. The quench outlet stream can be contacted with a caustic agent in the caustic wash tower to form a caustic outlet stream comprising a metal salt. The caustic outlet stream can be introduced to the quench tower. The ODH outlet stream can be contacted with the caustic outlet stream to form a derivative salt from the metal salt and the oxygenate. The oxygenate outlet stream can comprise a substantial portion of the derivative salt.

[1432] Referring to FIG 64, illustrated is a flow diagram of a non-limiting example of a system 6400 to convert an alkane to an alkene. As illustrated, an ODH reactor 6402 and a quench tower 6404 can be in operative communication. For example, an ODH outlet 6402b of the ODH reactor 6402 can be in fluid communication with a quench inlet 6404a of the quench tower 6404 via an ODH outlet line 6410. Additionally, a quench outlet 6404c of the quench tower 6404 can be in fluid communication with a wash inlet 6406a of the caustic wash tower 6406 via a quench outlet line 6414. Accordingly, the ODH reactor 6402 can he in fluid communication with the caustic wraslt tower 6406 via the quench tower 6404.

[1433] The ODH reactor 6402 can comprise an ODH inlet 6402a winch can be configured to receive an ODH inlet stream from an ODH inlet line 6408 and can be suitable to transport the ODH inlet stream into the ODH reactor 6402. The ODH inlet stream can comprise a gaseous mixture of a lower alkane and oxygen. In various examples, the ODH inlet stream additionally can include at least one of a carbon-based oxide, a sulfide, steam, and an inert diluent, as described herein. In various examples, the ODH inlet stream can comprise another hydrocarbon such as, for example, methane. The inert diluent can comprise, for example nitrogen, helium or argon among others. In various examples, the carbon-based oxide can comprise at least one of carbon dioxide and carbon monoxide. The concentration of the ox gen and the lower alkane within the mixture in the ODH inlet stream and the temperature and pressure of the ODH inlet stream can be adjusted such that the mixture can be outside of the flammability limits of the mixture. In various examples, the lower alkane is in a gas state. Invasions examples, the earhon-based oxide is in a gas state. In various examples, the sulfide is in a gas state.

[1434] in various examples, there may be multiple ODH inlet lines configured to introduce the ODH inlet stream to the ODH reactor 6402. For example, each reactant (e.g., lower alkane, oxygen, steam, carbon-based oxide, and inert diluent) may be added directly to the ODH reactor 6402, each in separate inlet lines (not shown). Alternatively, one or more reactants may be pre-mixed and added in more than one inlet line. In various example, reactants may be mixed together prior to the ODH reactor 6402 and subsequently introduced into the ODH reactor in a common ODH inlet. In various examples, steam may be added indirectly as water mixed with an additional reactant and the resulting mixture can be preheated before entering the ODH reactor 6402. When adding steam indirectly as water, the preheating process can increase the temperature of the mixture so that the water can be substantially converted, and in various examples fully converted, to steam before entering the ODH reactor 6402. [1435] The ODH reactor 6402 can include a catalyst capable of catalyzing the conversion of the reactants within die ODH inlet stream to products such as, for example, an alkene and an oxygenate and in various examples, a carbon-based oxide. The catalyst may be, for example, a mixed metal oxide catalyst, many varieties of which have been described in the art. In various examples, the products may additionally include water.

[1436] As known in the art, the catalyst composition, the composition of the ODH inlet stream, and reaction conditions within the ODH reactor 6402, such as temperature and pressure, can be adjusted in order to promote selectivity, as desired, of a product. For example, the ratio of the lower alkane to oxygen can be outside of the upper flammability limit of the mixture. In various examples, the o xygen concentration in the ODH inlet stream can be in a range of 0.1 % to 30 % by volume of the ODH inlet stream, and in some examples range from 0.1% to less than 30 % by volume, less than 25 % by volume, or less than 20 % by volume. In various examples, the lower alkane concentration in the ODH inlet stream can range from 0.1 % to 80 % by volume of the ODH inlet stream, and in some examples range from 0 1% to less than 50 % by volume or less than 40 % by volume.

[1437] In various examples increasing the steam concentration in the ODH inlet stream can increase the amount of oxygenate produced relative to the alkene produced in the ODH reactor 6402 In various examples, reducing the steam concentration in the ODH inlet stream can decrease the amount of oxygenate produced relative to the alkene produced in the ODH reactor 6402. The concentration of steam in the ODH inlet stream can be in a range of 0.1 % to 90 % by volume of the total ODH inlet stream 6408, and in some examples range from 0.1% to less titan 40 % by volume, or less titan 25 % by volume. In various examples, the concentration of the stream in the ODH inlet stream can be at least 1 % by volume. In some examples, the ODH inlet stream can comprise 20 % oxygen by volume, 40 % lower alkane by volume, and the balance being steam, carbon dioxide, and/or an inert diluent. In some examples, the ODH inlet stream can comprise 10 % oxygen by volume, 15% lower alkane by volume, and the balance being steam, carbon dioxide, and/or an inert diluent.

[1438] In various examples, the ODH process has a selectivity for the corresponding alkene (e.g., ethylene in Site case of ethane ODH) of greater than 95 mol % such as, for example greater than 98 mol %. The gas hourly- space velocity (GHSV) within the ODH reactor 6402 can be from 500 to 30000 h ¹ and in some examples die GHSV within the ODH reactor 6402 can be greater than 1000 h ^'1 . In various examples the space-time yield of corresponding alkene (e.g., productivity) in grams(g)/hour per kilogram (kg) of the catalyst can be at least 100 such as, for example, greater than 500, greater titan 1000, or greater than 2000, at an ODH reactor temperature of, for example, 300 °C to 370 °C. In various examples, the productivity of the catalyst can increase with increasing temperature in the ODH reactor 6402 until the selectivity of the alkene decreases.

[1439] The provided techniques can include use of an ODH reactor for performing an ODH reaction. In various examples, the reaction can be conducted at temperatures in a range of 300 °C to 370 °C such as, for example, 300 °C to 360 °C, or 300 °C to 350 °C. In various examples, the reaction can be conducted at pressures in a range of 0.5 pounds per square inch gauge (psig) to 100 psig (3.447 to 689.47 kPag) such as, for example, 15 psig to 50 psig (103.4 to 344.73 kPag). In various examples, the lower alkane can have a residence time in the ODH reactor 6402 in a range of 0.002 seconds (s) to 30 s, or from 1 s to 10 s.

[1440] The products of the ODH reaction can leave the ODH reactor 6402 through the ODH outlet 6402b in an ODH outlet stream. The ODH outlet 6402b can be configured to receive the ODH outlet stream and can be suitable to transport the ODH outlet stream 6410 out of the ODH reactor 6402 and into tire ODH outlet line 6410. In various examples, in addition to the products, the ODH outlet stream can include usireacted components from the ODH inlet stream such as, for example, lower alkane, carbon-based oxide, oxygen, steam, inert diluent, and combinations thereof. In various examples, the temperature of the ODH outlet stream can be in a range of 300 °C to 370 °C, such as for example, 300 °C to 360 °C, and in certain examples 300 °C to 350 °C.

[1441] Any of the kno w n reactor types applicable for the ODH of an alkane may be used with the provided techniques. For example, a fixed bed reactor, a fluidized bed reactor, or combinations thereof can be used for the ODH reactor 6402. In a typical fixed bed reactor reactants are introduced into the reactor at an inlet and flow past an immobilized catalyst. Products are formed and leave through the outlet of the reactor. A person skilled in the art would understand which features are required with respect to shape and dimensions of the reactor inputs for reactants, outputs for products, temperature and pressure control, and means for immobilizing the catalyst.

[1442] In a typical fluidized bed reactor, the catalyst bed can be supported by a porous structure or a distributor plate and located near a lower end of the reactor. Reactants flow through the fluidized bed reactor at a velocity sufficient to fluidize the bed (e.g., the catalyst rises and begins to swirl around in a fluidized manner). The reactants can be converted to products upon contact with the fluidized catalyst and the reactants are subsequently removed from an upper end of the reactor. A person of ordinary skill in the art would understand which features are required wi th respect to shape and dimensions of the reactor, the shape and size of the distributor plate, the input temperature, the output temperature, the reactor temperature and pressure, inputs for reactors, outputs for reactants, and velocities to achieve fluidization.

[1443] In various examples, there may be multiple ODH reactors connected in series or in parallel. Each ODH reactor may be the same or different. For example, each ODH reactor can contain the same or different ODH catalyst. In various examples, the multiple ODH reactors can each be a fixed bed reactor, can each be a fluidized bed reactor, or the multiple ODH reactors can be combinations of fixed bed reactors and fluidized bed reactors. [1444] Regardless of the configuration of the ODH reactor 6402, the ODH outlet 6402b can be in fluid communication with the quench inlet 6404a of the quench tower 6404 via the ODH outlet line 6410 to direct the ODH outlet stream to the quench tower 6404. The quench inlet 6404a can be configured to receive the ODH outlet stream from the ODH outlet line 6410 and can be suitable to transport the ODH outlet stream into the quench tower 6404. In various examples, the quench inlet 6404a can be configured to receive a product stream and can be suitable to transport the product stream into the quench tower 6404. The product stream can comprise at least one of a hydrocarbon, such as, for example, an alkane or an alkene, and an organic alcohol, such as, for example, ethanol. [1445J The quench tower 6404 can comprise a flash dram, an oxygenate scrubber, the like, or combinations thereof. The quench tower 6404 can be configured to quench the components in the ODH outlet stream and remove at least a substantial portion of the alkene from the ODH outlet stream. In various examples, the quench tower 6404 can facilitate the removal of oxygenate and water from Site ODH outlet stream. The quench tower 6404 can produce a quench outlet stream comprising at least a substantial portion of the alkene from the ODH outlet stream and in various examples at least a substantial portion of the caibon-based oxide from the ODH outlet stream. In various examples, the quench outlet stream can comprise additional components from the ODH outlet stream such as, for example, a portion of tire oxygen, a portion of the oxygenate, a portion of the inert diluent a portion of the steam and a portion of the imreacted alkane. In various examples, the quench outlet stream is in a gas state. The quench outlet stream exits the quench tower 6404 through the quench outlet 6404c. The quench outlet 6404c can be configured to receiv e the quench outlet stream and can be suitable to transport the quench outlet stream out of the quench tower 6404 into the quench outlet line 6414.

[1446] The quench tower 6404 can produce an oxygenate outlet stream comprising at least a substantial portion of the o xygenate from the ODH outlet stream and in some examples, a derivative salt as discussed herein. In various examples, the oxygenate outlet stream can comprise additional components from the ODH outlet stream such as, for example, a substantial portion of the water (e.g., steam), lower alkane, alkene. oxygen, and carbon-based oxide. The oxygenate outlet stream can exit the quench tower 6404 through an oxygenate outlet 6404b of the quench tower 6404. The oxygenate outlet 6404b can be configured to receive the oxygenate outlet stream and can be suitable to transport the oxygenate outlet stream out of the quench tower 6404 into the oxygenate outlet line 6412.

[1447] In various examples, the quench tow'er 6404 can be in operative communication with a caustic wash tower 6406. The quench outlet 6404c can be in fluid communication with the wash inlet 106a of the caustic wash tower 6406 via the quench outlet line 6414 to direct the quench outlet stream to the caustic wash tower 6406. The wash inlet 6406a can be configured to receive the quench outlet stream from the quench outlet line 6414 and can be suitable to transport the quench outlet stream into the caustic wash tower 6406.

[1448] The caustic wash tower 6406 can comprise the wnsh inlet 6406a, a wash outlet 6406c, a caustic inlet 6406d, and a caustic outlet 6406b. The caustic inlet 6406d can be configured to receive a caustic agent stream comprising a caustic agent from a caustic agent line 6420 and can be suitable to transport the caustic agent stream into the caustic wash tow'er 6406. The caustic agent can comprise a hydroxide, such as, for example at least one of sodium hydroxide potassium hydroxide, and ammonia hydroxide. In various examples, the caustic agent stream includes water or any other suitable component.

[1449] The caustic w'ash tow'er 6406 can be configured to contact the caustic agent stream with the quench outlet stream. In various examples comprising a carbon-based oxide comprising carbon dioxide, the caustic agent can react with carbon dioxide and/or sulfide in the quench outlet stream to form a metal salt. The metal salt may be, for example, at least one of a sulfide and a carbonate. The carbonate can comprise at least one of sodium bicarbonate, potassium carbonate, and ammonium bicarbonate. The sulfide can comprise hydrogen sulfide in various examples, the metal salt can be water soluble. The reaction can remove at least a substantial portion of the carbon-based oxide (e.g., carbon dioxide), and in various examples the sulfide (e.g., hydrogen sulfide), from tire quench outlet stream and produce a wash outlet stream and a caustic outlet stream. For example, the reaction of sodium hydroxide and carbon dioxide is shown in Scheme PI.

[14501 Scheme PI

C0 ₂ + NaOH ® NaHC0 ₃

[1451] The wash outlet stream can comprise unreacted components from the quench outlet stream. The wash outlet 6406c can be configured to receive the wash outlet stream and can be suitable to transport the wash outlet stream out of the caustic wash tower 6406 into the wash outlet line 6416.

[1452] The caustic outlet stream can comprise a substantial portion of the metal salt and in some examples, at least one of wa ter, caustic agent, and oxygenate. In various examples, the caustic outlet 6406b can be configured to receive the caustic outlet stream and can be suitable to transport the caustic outlet stream into a return line 6418. The return line 6418 can be configured to receive the caustic outlet stream and output the caustic outlet stream into a metal salt inlet 64Q4d of the quench tower 6404. In various examples, the caustic outlet stream can comprise a spent caustic stream.

[1453] In various examples, the caustic outlet stream may be produced by various suitable processes. For example, the caustic outlet stream can be produced by a cracking process such as ethylene cracking, a refinery process such as mercaptan oxidation, a paper manufacturing process, a soap manufacturing process, a detergent manufacturing process a food manufacturing process, any oilier suitable caustic producing process and combinations thereof. In various examples, a storage vessel can store caustic waste and the storage vessel can comprise a storage vessel outlet (not shown) suitable to output the caustic outlet stream into the metal salt inlet 6404d. Accordingly, the method, the sy stem, and the apparatus according to the present disclosure are not limited to ODH processes and the method, the system, and the apparatus according to the present disclosure can be used with other suitable processes.

[1454] In various examples, the caustic waste stream and the ODH outlet stream can be separately and/or concomitantly introduced into the quench tower 6404.

[1455] The quench tower 6404 can be configured to contact the caustic outlet stream with the ODH outlet stream. In various examples, the quench tower 6404 can be configured to react the caustic outlet stream with the ODH outlet stream to form a derivative salt and in various examples, a carbon-based oxide and or a sulfide, from the metal salt and the oxygenate. In various examples, the quench tower 6404 can react the metal salt with the oxygenate, and in some examples, with water and caustic agent, to form the derivative salt and the carbon-based oxide and/or sulfide. The derivative salt can comprise an acetate an acrylate, and a malonate. For example, the acetate can comprise at least one of sodium acetate, potassium acetate, and ammonium acetate. The acrylate can comprise at least one of sodium acry late, potassium acrylate, and ammonium acrylate. The malonate can comprise at least one of sodium malonate, potassium malonate, and ammonium malonate. In various examples, the derivative salt can he water soluble. As an example, the reaction of sodium bicarbonate and the oxygenate to form sodium acetate, carbon dioxide, and water is illustrated by the reaction in Scheme P2.

[1456] Scheme P2

Na.liCO _: + CH ₃ COOH C0 ₂ + H ₂ 0 + NaC ₂ H ₃ 0 ₂

[1457] In various examples, the mole ratio of the metal salt in the caustic outlet stream to oxygenate in the ODH outlet stream can be in a range of 0.8:1 to 1.2:1 such as for example 1 : 1. In various examples, the mole ratio of the metal salt in the caustic outlet stream to oxygenate in the ODH outlet stream can be greater than 1 : 1 such as, for example, 2:1.

[1458] In various examples, the quench tower 104 can be configured to maintain a pH in a range of 2 to 12 such as, for example, 4 to 11, 4 to 7, or 7 to 13. in various examples, the quench tower 104 can be configured to maintain a pH in a range of a pKa of the oxygenate to a pKa of the metal salt in order to facilitate the formation of the derivative salt in various examples, the oxygenate comprises acetic acid having a pKa of 4.7 and sodium bicarbonate having a pKa of 10.3. In various examples, the pH is measured in a mixture of water, oxy genate, and metal salt.

[1459] The quench outlet stream can comprise a substantial portion of the carbon-based oxide in the quench tower 6404 from the ODH outlet stream. In various examples, the carbon-based oxide from ODH outlet stream can pass through the quench tower 6404 substantially unreacted. The oxygenate outlet stream can comprise the oxygenate, the derivative salt, and water. Adding the caustic outlet stream to the quench tower can decrease the amount of oxygenate and increase the amount of derivative salt in the quench outlet stream. The decrease in oxygenate in the quench outlet stream can be a result of the conversion of the oxygenate to the derivative salt. The conversion of the ox genate to the derivative salt can facilitate the removal of the oxygenate from the ODH outlet stream and limit the oxygenate from exiting the quench tower 6404 in the alkene outlet stream.

[1460] The quench tower 6404 can be a single stage or multiple stages. For example, referring to FIG. 65, illustrated is a flow _' diagram of a non-limiting example of a system 6500 comprising a multistage quench tower. As illustrated, the ODH outlet line 6410 can be in fluid communication with a first heat exchanger (HX) inlet 6522a of a first HX 6522. The first HX inlet 6522a can be configured to receive the ODH outlet stream from the ODH outlet line 6410 and can be suitable to transport the ODH outlet stream into the first HX 6522. The first HX 6522 can be configured to adjust the temperature of the ODH outlet stream. For example, the first HX 6522 can cool the ODH outlet stream to a temperature of less than 200 °C such as, for example, less than 100 °C, less than 50 °C, less than 40 °C, and in some examples the first HX 6522 can cool the ODH outlet stream to a temperature of 20 °C to 80 °C. In various examples, the first HX 6522 can cool the ODH outlet stream to a temperature which induces condensation of the oxygenate such as, for example, a temperature less than or equal to the boiling point of the oxygenate and/or a temperature that reduces the vapor pressure of the oxygenate. The first HX 6522 can be any HX as known in the art. For example, the first HX 6522 can be a standalone HX separate from a quench tower. In various examples, the first HX 6522 can be an integrated HX that is part of a quench tower. f 1461] The temperature adjusted ODH outlet stream can exit the first HX 222 through a first HX outlet 6522b as a first HX outlet stream. The first HX outlet 6522b can be configured to receive the first HX outlet stream and can be suitable to transport the first HX outlet stream out of the first HX 6522 into the first HX outlet line 6536.

[1462] The first HX outlet 6522b can be in fluid communication with a separator inlet 6538a of a separator 6538 via the first HX outlet line 6536 to direct the first HX outlet stream to the separator 6538. The separator 6538 can comprise a vapor-liquid separator such as, for example, a flash drum. The separator inlet 6538a can be configured to receive the first HX outlet stream from the first HX outlet line 6536 and can be suitable to transport the first HX outlet stream into the separator 6538.

[1463] The separator 6538 can be configured to condense the oxygenate and produce a condensate outlet stream substantially comprised of liquid and an alkene outlet stream substantially comprised of gas. The condensate outlet stream can comprise oxygenate from the first HX outlet stream. In various examples, the condensate outlet stream can comprise at least 80 % oxygenate by weight such as, for example, at least 90 % oxygenate by weight, at least 95 % oxygenate by weight, or 80 % to 100 % oxygenate by weight. In various examples, the condensate outlet stream can additionally comprise water from the first HX outlet stream.

[1464] A condensate outlet 6538b of the separator 6538 can be configured to receive the condensate outlet stream and can be suitable to transport the condensate outlet stream out of the separator 6538 and into the condensate line 6542. An alkene outlet 6538c of the separator 6538 can be configured to receive the alkene outlet stream and can be suitable to transport the alkene outlet stream out of the separator 6538 into the alkene outlet line 6540.

[1465] In various examples, a second HX 6524 can be provided in fluid communication with the separator 6538. For example, the alkene outlet line 6540 can be in fluid communication with a second HX inlet 6524a of the second HX 6524. The second HX inlet 6524a can be configured to receive the alkene outlet stream from the alkene outlet line 6540 and can be suitable to transport the ODH outlet stream into the second HX 6524. The second HX 6524 can be configured to adjust the temperature of the alkene outlet stream. For example the second HX 6524 can cool the alkene outlet stream to a temperature of less than 170 °C such as, for example, less than 100 °C, less than 50 °C, less than 40 °C, and in some examples, the second HX 6524 can cool the alkene outlet stream to a temperature of 20 °C to 80 °C. In various examples, the second HX 6524 can cool the ODH outlet stream to a temperature which induces condensation of the oxygenate such as, for example, a temperature less than or equal to the boiling point of the oxygenate and/or a temperature that reduces the vapor pressure of the oxygenate. The second HX 6524 can be any HX as known in the art. For example, the second HX 6524 can be a standalone HX separate from a quench tower. In various examples, the second HX 6524 can be an integrated HX that is part of a quench tower.

[1466] The temperature adjusted ODH outlet stream can exit the second HX 6524 through a second HX outlet 6524b as a second HX outlet stream. The second HX outlet 6524b can be configured to receive the second HX outlet stream and can be suitable to transport the second HX outlet stream out of the second HX 6524 into the second HX outlet line 6556.

[1467] The second HX outlet 6524b can be in fluid communication with a quench inlet 6504a of a quench tower 6504 via the second HX outlet line 6556 to direct the second HX outlet stream to the quench tower 6504. The quench inlet 6504a can be configured to receive the second HX outlet stream from the second HX outlet line 6556 and can be suitable to transport the second HX outlet stream into the quench tower 6504.

[1468] A metal salt inlet 6504d of the quench tower 6504 can be configured to receive the caustic outlet stream from the return line 6418 and can be suitable to transport the caustic outlet stream into the quench tower 6504. The quench tower 6504 can be configured to contact the caustic outlet stream with the second HX outlet stream. In various examples, the quench tower 6504 can be configured to react the caustic outlet stream with the second HX outlet stream to form a derivative salt and in various examples, a carbon-based oxide and/or sulfide. In various examples the quench tower 6504 can react the metal salt with the oxygenate, and in some examples, with water and caustic agent, to form the derivative salt and carbon-based oxide and/or sulfide. The caustic outlet stream can enable more efficient removal of the oxygenate from the alkene outlet stream. Removing more oxygenate from tiie alkene outlet stream can lengthen the operational life of downstream equipment that can be forded by formation of a derivative salt from the oxygenate.

[1469] The quench tower 6504 can be configured to quench the components in tire second HX outlet stream and remove at least a substantial portion of the alkene from the second HX outlet stream. In various examples, the quench tower 6504 can facilitate the removal of oxygenate and water from the second HX outlet stream. The quench tower 6504 can produce a quench outlet stream comprising at least a substantial portion of the alkene and at least a substantial portion of the carbon-based oxide from the second HX outlet stream. In various examples, the quench outlet stream can comprise additional components from the second HX outlet stream such as for example oxygen, oxygenate inert diluent, water (e.g., steam), and unreacted alkane. The quench outlet stream exits the quench tower 6504 through the quench outlet 6504c. The quench outlet 6504c can be configured to receive the quench outlet stream and can be suitable to transport the quench outlet stream out of the quench tower 6504 into the quench outlet line 6414.

[1470] The quench tower 6504 can produce a derivative salt outlet stream comprising at least a substantial portion of the oxygenate from the second HX outlet stream and/or at least a substantial portion of the derivative salt formed by the quench tower 6504. In various examples, the derivative salt outlet stream can comprise additional components from the second HX outlet stream such as, for example a substantial portion of the water lower alkane, alkene, oxygen, and carbon-based oxide. The derivative salt outlet stream exits the quench tower 6504 through an oxygenate outlet 6504b of the quench tower 6504. The oxygenate outlet 6504b can be configured to receive the derivative salt outlet stream and can be suitable to transport the derivative salt outlet stream out of the quench tower 6504 into the oxygenate outlet line 6512.

[1471 ] In various examples, the oxygenate in the oxygenate outlet stream, the condensate outlet stream, or derivative salt outlet stream may be subject to further processing. For example, referring to FIG. 66, the oxygenate can be separated from the derivative salt in a separation vessel 6626. FIG. 66 is a flow' diagram of a non-limiting example of a sy stem 6600 comprising the separation vessel 6626. As illustrated the separation vessel 6626 has a separation inlet 6626a, a first separation outlet 6626b. and a second separation outlet 6626c. The separation inlet 6626a can be configured to receive the oxygenate outlet stream from oxygenate outlet line 6412 and may be suitable to transport the oxygenate outlet stream into the separation vessel 6626. In various examples, the separation inlet 6626a can be configured to receive at least one of the derivative salt outlet stream from the oxygenate outlet line 6512 and the condensate outlet stream from the condensate line 6542 and may be suitable to transport the respective stream(s) into the separation vessel 6626.

[1472] The separation vessel 6626 can separate the oxygenate from the derivative salt and, in various examples, the separation vessel 6626 can separate the oxygenate from water. The presence of the derivative salt in the separation vessel 6626 can enhance the separation of oxygenate from the water. For example, the deriva tive salt and oxygenate may disassociate and/or react with water to form a derivative salt ion (e.g., CH ₃ COO ) and an acid (e.g., H ₃ 0 ⁺ , Na ⁺ ) . Since the derivative salt and oxygenate can form a common ion, an increase in the concentration of one of the derivative salt and oxygenate can affect the other. For example, the reactions of sodium acetate (C ₂ H ₃ Na0 ₂ ), acetic acid (CILCOOH), bicarbonate ion (HCOc ), carbon dioxide (C0 ₂ ), and water (H ₂ 0) is illustrated in Scheme P3.

[1473] Scheme P3

[1474] As illustrated in Scheme P3, sodium acetate can form an acetate ion which can affect the equilibrium reaction of acetic acid and water. For example, the sodium acetate can cause the equilibrium reaction of acetic acid and water to have a higher preference for the separate species of acetic acid and water than an aceta te ion and an acid relative to without the presence of acetate.

[1475] The separation vessel 6626 can comprise various equipment known to those of ordinary' skill in the art. For example, the separation vessel 6626 can comprise an extraction tower, a packed column, a sieve-tray column, a spray column a KARR column, a rotating disc contactor, a stirred cell extractor, a rectification tower a stripper, and combinations thereof. In various examples, the separation vessel 6626 can comprise a liquid-liquid extractor. Accordingly, the derivative salt in the oxygenate inlet stream can increase the efficiency of the separation vessel 6626 and can facilitate efficient separation of die oxygenate from water.

[1476] The separation vessel 6626 can produce a second separation outlet stream comprising a substantial portion of the oxygenate from the oxygenate outlet stream. In various examples, the second separation outlet stream can comprise additional components from the oxygenate outlet stream such as, for example, water. In various examples, the second separation outlet stream can comprise at least 80 % oxygenate by weight such as, for example at least 90 % oxygenate by weight at least 95 % oxygenate by weight, or 80 % to 100 % oxygenate by weight. The second separation outlet stream can exit the separation vessel 6626 through the second separation outlet 6626c of the separation vessel 6626. The second separation outlet 6626c can be configured to receive the second separation outlet stream and can be suitable to transport the second separation outlet stream out of the separation vessel 6626 into the second separation outlet line 6628.

[1477] The separation vessel 6626 can produce a first separation outlet stream comprising a substantial portion of the derivative salt from the oxygenate outlet stream and in various examples, a substantial portion of the water from the oxygenate outlet stream. In various examples, the first separation outlet stream can comprise at least 10 % derivative salt by weight such as, for example, at least 30 % derivative salt by weight, at least 50 % derivative salt by weight, or 30 % to 70 % derivative salt by weight In various examples, the first separation outlet stream can comprise at least 5 % water by weight such as, for example, at least 10 % water by weight, at least 25 % water by weight, or 15 % to 50 % water by weight The first separation outlet stream can exit the separation vessel 6626 through the first separation outlet 6626b of tire separation vessel 6626 The first separation outlet 6626b can be configured to receive the first separation outlet stream and can be suitable to transport the first separation outlet stream out of the separation vessel 6626 into the first separation outlet line 6630

[1478] The separation vessel 6626 can be configured with a recycle line 6632 in fluid communication with the fust separation outlet line 6630 and/or first separation outlet 6626b. The recycle line 6632 can be configured to recycle a portion of the derivative salt from the first separation outlet stream to the separation vessel 6626 via the recycle inlet 6626d. The recycle line 6632 can be configured to receive a portion of the first separation outlet stream and can be suitable to transport a recycle stream to a recycle inlet 6626d of tire separation vessel 6626. The recycle inlet 6626d can be configured to receive the recycle stream and can be suitable to transport the recycle stream into the separation vessel 6626. For example, the recycle stream can comprise a portion of the derivativ e salt from the first separation outlet stream, and in various examples, a portion of the water from the first separation outlet stream. [1479] The recycle line 6632 can be configured to recycle the derivative salt from the first separation vessel outlet stream until a select concentration of derivative salt is achieved in the separation vessel 6626. In various examples and referring to Figures 64 and 66, the return line 6418 can enable additional generation of derivative salt in the quench tower 6404 which would flow to the separation vessel 6626 through the oxygenate outlet line 6412 to increase the concentration of derivative salt in the separation vessel 6626

[1480] In various examples, a supplemental salt stream can be added to the separation vessel 6626. In various examples, the supplemental salt can comprise ethyl acetate.

[1481] Referring to FIG. 67, in various examples, an oxygen remover 6744 can be disposed intermediate the ODH reactor 6402 and the quench tower 6404. FIG. 67 is a flow diagram of a non-limiting embodiment of a system 6700 comprising an oxygen remover 6744. As illustrated, the oxygen remover 6744, comprising a remover inlet 6744a and a remover outlet 6744b, can be provided in fluid communication with the ODH reactor 6402 (FIG. 64) via ODH outlet line 6410 and the quench tower 6404 via remover outlet line 6746. The remover inlet 6744a can be configured to receive the ODH outlet stream and can be suitable to transport the ODH outlet stream into the oxygen remover 6744. The oxygen remover 6744 can remove a substantial portion of the oxygen in the ODH outlet stream and produce a remover outlet stream comprising the ODH outlet stream with the substantial portion of the oxygen removed. The oxygen remover 6744 can be of various designs as known in the art. The remover outlet 6744b can be configured to receive the remover outlet stream and can be suitable to transport the remover outlet stream out of the oxygen remover 6744 into the remover outlet line 6746. The quench inlet 6404a of the quench tower 6404 can be configured to receive the remover outlet stream.

[1482] Referring to FIG. 68, in various examples, an amine tower 6848 can be disposed intermediate the quench tower 6404 and the caustic wash tower 6406. FIG. 68 is a flow diagram of a non-limiting example of a system 6800 comprising an amine tower 6848. As illustrated, the amine tower 6848, comprising an amine tower inlet 548a and an amine tower outlet 6848b, can be provided in fluid communication with the quench tower 6404 (FIG. 64) via quench outlet line 6414 and the caustic wash tower 6406 via amine tower outlet line 6850. The amine tower inlet 6848a can be configured to receive the quench outlet stream and can be suitable to transport the quench outlet stream into the amine tower 6848. The amine tower 6848 can remove a substantial portion of carbon dioxide in the quench outlet stream and produce an amine tower outlet stream comprising the quench outlet stream with the substantial portion of the carbon dioxide removed. The amine tower 6848 can be of various designs as known in the art. i 14831 The amine tower outlet 6848b can be configured to receive Site amine tower outlet stream and can be suitable to transport the amine tower outlet stream out of the amine tower 6848 into the amine tower outlet line 6850. The wash inlet 6406a of the caustic wash tower 6406 can be configured to receive the amine tower outlet stream from die amine tower outlet line 6850.

[1484] Having a high efficiency oxygenate removal prior to the amine tower 6848 can limit, and in some examples prevent, amine degradation to presence of the oxygenate in the amine tower 6848. For example, the oxygenate can form heat stable salts with amine in the amine tower 6848 which can degrade the efficiency and shorten the operational life of the amine tower 6848.

[1485] Referring to FIG. 69, in various examples, a polymerization reactor 6952 can be in fluid communication with the caustic wash tower 6406 via the wash outlet line 6416. FIG. 69 is a flow diagram of a non- limiting example of a system 6900 comprising a poly merization reactor 6952. As illustrated, the polymerization reactor 6952, comprising a polymerization inlet 6952a and a poly merization outlet 6952b, can be provided in fluid communication with the caustic wash tower 6406 via the wash outlet line 6416. In various examples, a demethanizer and C _l splitter (not shown) may be disposed in the wash outlet line 6416 between the caustic wash tower 6406 and the polymerization reactor 6952. The polymerization inlet 6952a can be configured to receive the ODH outlet stream and can be suitable to transport the ODH outlet stream into the polymerization reactor 6952. The polymerization reactor 6952 can produce a polymer from the alkene and produce a polymerization outlet stream comprising the polymer. In various examples, the polymer comprises at least one of polyethylene, polypropylene, and polybutylene. The pol merization reactor 6952 can be of various designs as known in the art. The polymerization outlet 6952b can be configured to receive the pol merization outlet stream and can be suitable to transport the pol merization outlet stream out of the polymeriza tion reactor 6952 into the polymerization outlet line 6954.

[1486] Concentrations of the components wi thin the sy stem can be measured any a t point in the process using any means known in the art. For example, a detector such as a gas chromatograph, an infrared spectrometer, and a Raman spectrometer can be disposed downstream or upstream of ODH reactor 6402, quench tower 6404, caustic wash tower 6406, separator 6538, separation vessel 6626. oxygen remover 6744, amine tower 6848, and polymerization reactor 6952.

[1 87] In various examples, the ODH inlet stream 6408 can comprise mixtures that fall within the flammability limits of the components. For example, the mixture may exist in conditions that prevent propagation of an explosive event. In these examples, the flammable mixture can be created within a medium where ignition can be immediately quenched. In various examples oxygen and the lower alkanes can he mixed at a point where they are surrounded by a flame arresting material. Thus, any ignition can he quenched by the surrounding material. Flame arresting material includes, for example, metallic or ceramic components, such as stainless steel walls or ceramic supports in various examples, oxygen and lower alkanes can be mixed at a low temperature, where an ignition event may not lead to an explosion, then the mixture can be introduced into the ODH reactor before increasing the temperature. Therefore, the flammable conditions may not exist until the mixture can be surrounded by the flame arresting material inside of the reactor.

[14881 In various examples, the olefins produced using an ODH reactor or any of the processes or complexes described herein can be used to make various olefin derivatives utilizing a polymerization reactor. Olefin derivatives include, but are not limited to, polyethylene polypropylene, ethylene oxide, propylene oxide polyethylene oxide, polypropylene oxide vinyl acetate, vinyl chloride, acrylic esters (e.g., methyl methacrylate), thermoplastic elastomers, thermoplastic olefins, blends thereof, and combinations thereof [1489J In various examples, ethylene and optionally ct-olefins can be produced in an ODH reactor, or any of the processes or complexes described herein and are used to make polyethylene utilizing a polymerization reactor. The polyethylene made from the ethylene and optional a-olefins described herein can include homopolymers of ethylene, copolymers of ethylene and ct-olefins, resulting in HOPE, MDPE, LDPE, LLDPE and VLDPE.

[1490] The polyethylene produced using the ethylene and optional ct-olefins described herein can be produced using any suitable polymerization process and equipment. Suitable ethylene polymerization processes include, but are not limited to gas phase polyethylene processes, high pressure polyethylene processes, low pressure polyethylene processes, solution polyethylene processes, slurry polyethylene processes and suitable combinations of the above arranged either in parallel or in series.

[1491] Techniques are provided for a process for converting a lower alkane to an alkene. The process can include introducing an input stream comprising oxygen and the lower alkane to an ODH reactor 6402. In various examples, the input stream additionally can include at least one of a carbon-based oxide, steam, and an inert diluent. At least a portion of the lower alkane can be converted to the alkene in the ODH reactor 6402. In various examples, the alkane can comprise ethane and the alkene comprises ethylene. In various examples, the alkane can comprise propane and the alkene comprises propylene. In various examples, the alkane comprises butane and the alkene can comprise butylene. An ODH outlet stream comprising the alkene, an oxygenate, and a carbon-based oxide may be produced. In various examples, the ODH outlet stream can comprise at least one of a sulfide, water, an unreacted alkane, oxygen, and an inert diluent.

[1492] The ODH outlet stream to can be introduced to a quench tower 6404 and the oxygenate can be removed from the ODH outlet stream in the quench tower 6404 to produce a quench outlet stream comprising at least a substantial portion of the alkene and at least a substantial portion of the carbon-based oxide. Additionally, the quench tower 6404 can produce an ox genate outlet stream comprising the at least a substantial portion of the oxygenate. f 1493] In various examples, the ODH outlet stream can be introduced to an oxygen remover 6744 prior to the quench tower 6404. Oxygen can be removed from the ODH outlet stream in the oxygen remover 6744 and the ODH outlet stream can be introduced to the quench tower 6404 after the oxygen remover 6744.

[1494] The quench outlet stream can be introduced to a caustic -wash tower 6406. The quench outlet stream can be contacted with a caustic agent to form a caustic outlet stream comprising a metal salt. In various examples, the quench outlet stream is contacted with tire caustic agent in the caustic wash tower 6406.

[1495] in various examples, the quench outlet stream can be introduced to an amine wash tower 6848 prior to the caustic wash lower 6406. A substantial portion of the carbon-based oxide can be removed from the quench outlet stream. The quench outlet stream with the substantial portion of the carbon-based oxide removed can be introduced to the caustic wash tower 6406.

[1496] The caustic outlet stream can be introduced the quench tower 6404 and die ODH outlet stream can be contacted with die caustic outlet stream to form a derivative salt and in various examples a carbon-based oxide and a sulfide. In various examples, the ODH outlet stream is contacted with the caustic outlet stream in die quench tower 6404. The oxygenate outlet stream can comprise a substantial portion of the derivative salt. In various examples, the pH of the quench to wer 6404 can be maintained in a range of 2 to 12 such as, for example, 4 to 11, 4 to 7, or 7 to 11. In various examples, the pH of the quench tower 6404 can be maintained in a range of a pKa of the oxygenate to a pKa of the metal salt.

[1497] In various examples, the oxygenate outlet stream can be introduced to a separation vessel 6626. The oxygenate can be separated from the derivative salt within he oxygenate outlet stream. A second oxygenate outlet stream comprising a substantial portion of the oxygenate from the oxyge nate outlet stream can be produced. A separation outlet stream comprising a substantial portion of the derivative salt from the oxygenate outlet stream can be produced. In various examples, a portion of the separation outlet stream can be recycled to the separation vessel 6626. In various examples, a supplemental salt can be introduced to the separation vessel 6626 such as, for example, ethyl acetate

[1498] In various examples, the ODH outlet stream can be separated into a first intermediate stream and a second intermediate stream. The first intermediate stream can comprise at least a substantial portion of the oxygenate from the ODH outlet stream. The second intermediate stream can comprise at least a substantial portion of the alkene from the ODH outlet stream. The second intermediate stream can contact the caustic outlet stream to form the derivative salt and in various examples a carbon-based oxide and/or a sulfide.

[1499] In various examples, olefin derivatives can be produced from the alkene.

[1500] Techniques are provided for an alternative use for die caustic waste stream which limits, and in some examples, can eliminate a need to dispose of the caustic waste stream. Additionally, the reuse of the caustic waste stream can introduce a useful product of derivative salt which can aid in oxygenate separation from the quench outlet stream and purification of oxygenate in the separation vessel. The efficient removal of the ox genate from rite quench outlet stream can lengthen the operational life of downstream equipment such as protecting the amine tower against fouling and amine solution degradation. Moreover, the derivative salt can be sold. Furthermore, the efficient purification of the ox genate can create a marketable product such as, for example, glacial acetic acid. f 1501] The method, system, and apparatus according to the present disclosure may comprise other suitable process equipment such as, for example, a compressor and a pump.

[1502] EXAMPLES

[1503] Computational modeling of a liquid-liquid separation vessel using ASPEN Plus® version 8.6 chemical process simulation software, commercially available from Aspen Technology, Inc. Bedford, Massachusetts, was used to demonstrate the increase in concentration of a dilute oxygenate stream using the method described. The model simulates the effect of temperature, mass flow rate and composition of the oxy genate outlet stream on the composition of the separation outlet stream and the second oxygenate outlet stream. The compositions chosen for each example reflect compositions that may be present in an oxygenate outlet stream that is produced from a quench tower downstream from an oxidative dehydrogenation of ethane process. Ox genate outlet streams from an ethane ODH process typically comprises dilute acetic acid where the acetic acid mass fraction ranges from 1% to 5%, but in some instances may reach 25%. The oxygenate outlet stream may also comprise trace levels of carbon oxides, such as carbon dioxide. Addition of a caustic outlet stream comprising a metal salt into the quench tower may have the effect of reducing the mass fraction of acetic acid in the oxygenate outlet stream.

[1504] Example P I

[1505] For Example PI, input levels represent compositions of carbon dioxide, water acetic acid, and sodium acetate (as a sodium ion and acetate ion) representative of an oxygenate outlet stream coming directly from the quench tower with no additional sodium acetate added (via a recycle line). The total mass flow rate was set at 6980 kg/hr and at a pressure of 185.7 kPa gauge and a temperature of 40°C. The simulation results revealed a mass flow rate of the separation outlet stream of 5436 kg/hr and a mass flow rate of the second separation outlet stream was 1545 kg/hr.

[1506] Example P2

[ 507] For Example P2, input levels represent compositions of carbon dioxide, water, acetic acid, and sodium acetate (as a sodium ion and acetate ion) for the oxygenate outlet stream that includes additional sodium acetate (added via a recycle line). The total mass flow rate was set at 55982 kg/hr and at a pressure of 465 kPa gauge and a temperature of 65°C. The simulation results revealed a mass flow rate of the separation outlet stream of 54917 kg/hr and a mass flow rate of the second separation outlet stream was 975 kg/hr.

[1508] Example P3

[1509] For Example P3, input levels represent compositions of carbon dioxide, water, acetic acid, and sodium acetate (as a sodium ion and acetate ion) for the oxygenate outlet stream that includes additional sodium acetate (added via a recycle line). The total mass flow rate was set at 61014 kg/hr and at a pressure of 465 kPa gauge and a temperature of 52°C. The simulation results revealed a mass flow rate of the separation outlet stream of 60537 kg/hr and a mass flow rate of the second separation outlet stream was 477 kg/hr. [1510] Table PI

[1511] As shown in Table PI, all examples show a significant separation of sodium acetate from acetic acid, resulting in a much more concentrated and purer solution of acetic acid. The second oxygenate outlet stream in each example show's no detectable sodium acetate. Addition of excess sodium acetate to the oxygenate outlet stream via recycling of the separatio n outlet stream increased mass flow rate without increasing the mass fraction of acetic acid in the second oxygenate outlet stream. However, in both Examples P2 and P3, the mass fraction of acetic acid in the oxy gen outlet stream was significantly lower compared to Example PI. This demonstrates that recycling of the separation outlet stream would be helpful for oxy genate outlet streams where oxy genate levels are lower, such as in instances where the quench tower comprises a first stage and a second stage. Removal of a substantial portion of the acetic acid in the first stage results in the stream going to the second stage having a significantly lower mass fraction of acetic acid. These results show' that even in this instance the acetic acid level of the second oxygenate outlet stream is over 90% (mass fraction).

[1512] Techniques are provided for oxidative dehydrogenation (ODH) of lower alkanes into corresponding alkenes. In some embodiments, the carbon dioxide output levels from an ODH process are controlled.

[ 1513] A method of converting one or more alkanes to one or more alkenes that includes a) providing a first stream containing one or more alkanes and oxygen to an oxidative dehydrogenation reactor; b) converting at least a portion of the one or more alkanes to one or more alkenes in the oxida tive dehydrogenation reactor to provide a second stream exiting the oxidative dehydrogenation reactor containing one or more alkanes, one or more alkenes, oxygen, carbon monoxide and optionally acetylene; and c) providing the second stream to a second reactor containing a catalyst that includes a group 11 metal to convert a least a portion of the carbon monoxide to carbon dioxide and reacting the acetylene. f 1514] In some embodiments disclosed herein, the degree to which carbon monoxide is produced during the ODH process can be mitigated by converting it to carbon dioxide, which can then act as an oxidizing agent. The process can be manipulated so as to control the output of caibon dioxide from the process to a desired level. Using the methods described herein a user may choose to operate in caibon dioxide neutral conditions such that surplus carbon dioxide need not be flared or released into the atmosphere.

[1515] Disclosed herein are methods for mitigating caibon monoxide and/or acetylene formation in an ODH process and controlling the caibon dioxide output from the ODH process. Aspects of the methods include introducing, into at least one ODH reactor a gas mixture of a lower alkane, oxygen and carbon dioxide, under conditions that allow production of the corresponding alkene and smaller amounts of various by-products. For multiple ODH reactors, each reactor contains the smite or different ODH catal st, provided, in some embodiments, that at least one ODH catalyst is capable of using carbon dioxide as an oxidizing agent. In some embodiments steam or other inert diluents may also be introduced into the reactor as part of Site gas mixture. In some embodiments the amount of carbon dioxide leaving the reactor is subsequently monitored. If the amount of carbon dioxide output is below a desired level then the amount of steam introduced into the reactor can be increased. If the amount of carbon dioxide output is above the desired level then the amount of steam introduced into the reactor can be decreased. [1516] In some embodiments, the lower alkane is ethane, and the corresponding alkene is ethylene.

[1517] In some embodiments, at least one ODH reactor is a fixed bed reactor. In some embodiments, at least one ODH reactor is a fixed bed reactor that includes heat dissipative particles within the fixed bed. In some embodiments the heat dissipative particles have a thermal conductivity that is greater than the catalyst. In alternative embodiments, at least one ODH reactor is a fluidized bed reactor.

[1518] In some embodiments, at least one ODH catalyst is a mixed metal oxide catalyst. In particular embodiments, at least one ODH catalyst is a mixed metal oxide of the formula: Mo _a V _& TeJNTb _a Pd _e O;, wherein a, b, c, d, e and f are the relative atomic amounts of the elements Mo, V, Te, Nb, Pd and O, respectively; and when a = 1 , b = 0.01 to 1.0, c = 0.01 to 1 .0, d = 0.01 to 1 .0, 0.00 < e < 0.10 and f is a number to satisfy the valence state of the catalyst.

[1519] In some embodiments, at least one ODH catalyst is a mixed metal oxide of the formula:

M06.25-7.25 30 where d is a number to satisfy the valence of the oxide.

[1520] Various embodiments relate to oxidative dehydrogenation (ODH) of lower alkanes into corresponding alkenes. Lower alkanes are saturated hydrocarbons with from 2 to 4 carbons, and the corresponding alkene includes hydrocarbons with the same number of carbons, but with one carbon to carbon double bond. While any of the lower alkanes can be converted to their corresponding alkenes using the methods disclosed herein, one particular embodiment is the ODH of ethane, producing its corresponding alkene, ethylene.

[1521 ] Carbon Dioxide Output

[1522] Carbon dioxide can be produced in the ODH reaction as a by-product of oxidation of die alkanes and recycled from the oxidation of carbon monoxide. Carbon dioxide can also be added into the ODH reactor when used as an inert diluent. Conversely, carbon dioxide may be consumed when it acts as an oxidant for the deh drogenation reaction. The carbon dioxide output is therefore a function of the amount of carbon dioxide added and produced minus that consumed in the oxidative process. In some embodiments, the disclosed methods control the degree to which carbon dioxide acts as an oxidizing agent so as to impact the overall carbon dioxide output coming off the ODH reactor.

[1523] Measuring the amount of carbon dioxide coining off the ODH reactor can be done using any means known in the art. For example, one or more detectors such as GC, IR, or Rahman detectors, are situated immediately downstream of the reactor to measure the carbon dioxide output. While not required, the output of other components may also be measured. These include but ate not limited to the amounts of ethylene, unreacted ethane, carbon monoxide and oxygen, and by-products such as acetic acid. In addition, it should be noted that depending on Site chosen metric for carbon dioxide output, the output levels of the other components for example ethane, may actually be required.

[1524] Carbon dioxide output can be stated using any metric commonly used in the art. For example, the carbon dioxide output can be described in terms of mass flow' rate (g/min) or volumetric flow rate (crnVmin). In some embodiments, normalized selectivity can be used to assess the degree to which carbon dioxide is produced or consumed. In that instance, the net mass flow' rate of C(¾ — the difference between the mass flow rate of C(¾ entering and leaving the ODH reactor — is normalized to the conversion of ethane, in essence describing what fraction of ethane is converted into carbon dioxide as opposed to ethylene, or other by-products such as acetic acid.

A carbon selectivity of 0 indicates that the amount of carbon dioxide entering the reactor is the same as the carbon dioxide output. In other words, the process is carbon dioxide neutral. A positive carbon dioxide selectivity alerts a user that carbon dioxide is being produced, and that any oxidation of carbon dioxide that is occurring is insufficient to offset that production, resulting in the process being carbon dioxide positive which may result in a lower selectivity for the olefin.

[1525] In some embodiments, product selectivity for carbon dioxide is less than about 10 wt. %, in some cases less than about 7.5 wt. % and in other cases less than about 5 wt. %. The product selectivity for carbon dioxide can be any of the values or range between any of the values recited above.

[1526] In some embodiments, the total amount of carbon dioxide in the stream exiting the one or more ODH reactors can be essentially the same as the total amount of carbon dioxide in the stream entering the one or more ODH reactors.

[1527] In this instance, essentially the same means that the difference between the amount of carbon dioxide in the stream exiting the ODH reactors is within 2 weight percent (+ 2 wt. %) of the amount of carbon dioxide entering the ODH reactors in some embodiments, tire amount of carbon dioxide in the stream exiting the ODH reactors can be about +5 wt. %, in some cases about +7.5 wt. % and in other cases about +10 wt. % and can be about -5 wt. %. in some cases about -7.5 wt. % and in oilier cases about -10 wt. % of the amount of carbon dioxide in the stream entering the ODH reactors. The difference between the amount of carbon dioxide in the stream exiting the ODH reactors and the amount of carbon dioxide entering the ODH reactors can be any value or range between any of the values recited above. f 1528] In some embodiments, the methods and apparatus disclosed herein provide the possibility of a carbon dioxide negative process. In this instance, carbon dioxide is oxidized at a higher rate than it is produced and shows a negative carbon selectivity. The GDH process may produce carbon dioxide, but the degree to which carbon dioxide is consumed while acting as an oxidizing agent offsets any production that is occurring. Many industrial processes, in addition to ODH, produce carbon dioxide which must be captured or flared where it contributes to the emission of greenhouse gases. When using a carbon dioxide negative process, the excess carbon dioxide from other processes may be captured and used as the inert diluent in the ODH process under conditions where there is negative carbon selectivity. An advantage then is the ability to reduce the amount of carbon dioxide produced in die ODH process in combination with other processes, such as thermal cracking. In addition, oxidation of carbon dioxide is endothermic and by increasing the degree to which carbon dioxide acts as an oxidizing agent, heat produced front ODH of ethane is partially offset by oxidation of carbon dioxide, reducing the degree to which heat must be removed from the reactor. In some embodiments, when acting as an oxidizing agent, carbon dioxide can produce carbon monoxide which can be captured and used as an intermediate in production of other chemical products, such as methanol or formic acid.

[1529] In embodiments of a carbon dioxide negative process the total amount of carbon dioxide in the stream exiting the one or more ODH reactors is less the total amount of carbon dioxide in the stream entering the one or more ODH reactors. In this instance, the difference between the amount of carbon dioxide in the stream exiting the ODH reactors is less than about 1 wt. % in some circumstances less than about 2 wt. %, in other circumstances less than about 3 wt. %, in some cases less than bout 5 wt. %, in other cases less than about 7.5 wt. % and in some situations less than about 10 wt. % and can be higher as a non-limiting example less than about 20 wt. % less than the amount of carbon dioxide in the stream entering the ODH reactors. The difference between the amount of carbon dioxide in the stream exiting the ODH reactors and the amount of carbon dioxide entering the ODH reactors can be any value or range between any of the values recited above

[1530] The ODH Process

[1531] ODH of alkanes includes contacting a mixture of one or more alkanes and oxygen in an ODH reactor with an ODH catalyst under conditions that promote oxidation of the alkanes into their corresponding aikene. Conditions within the reactor are controlled by the operator and include, but are not limited to, parameters such as temperature, pressure, and flow rate. Conditions will vary and can be optimized for a particular alkane, or for a specific catalyst, or whether an inert diluent is used in the mixing of the reactants.

[1532] An ODH reactor can be used for performing an ODH process consistent with tire provided techniques. For best results, the oxidative dehydrogenation of one or more alkanes may be conducted at temperatures from 300 °C to 450 °C, or from 300 °C to 425 °C, or from 330 °C to 400 °C, at pressures from 0.5 to 100 psig (3.447 to 689.47 kPag), or from 15 to 50 psig (103.4 to 344.73 kPag), and the residence time of the one or more alkanes in the reactor may be from 0.002 to 30 seconds, or from 1 to 10 seconds.

[1533] In some embodiments, the process hits a selectivity for the corresponding aikene (ethylene in the case of ethane ODH) of greater titan 95%, or for example, greater than 98%. The gas hourly space velocity (GHSV) can be from 500 to 30000 h ¹ , or greater than 1000 Ir ¹ . In some embodiments, the space-time yield of corresponding alkene (productivity) in g/hour per kg of the catalyst can be at least 900 or above, or greater than 1500, or greater than 3000, or greater than 3500, at 350 to 400 °C In some embodiments, the productivity of the catalyst will increase with increasing temperature until the selectivity is decreased.

[1534] ODH Catalyst

[1535] Any of the ODH catalysts known in the art are suitable for use in the methods disclosed herein. Non- limiting examples of suitable oxidative dehydrogenation catalyst include those containing one or more mixed metal oxides selected from:

[1536] catalysts of the formuia:

Mo _a V,Te _f Nb _rf Pd,Or where a, b, c, d, e and f are the relative atomic amounts of the elements Mo, V, Te, Nb, Pd and O, respectively; and when a = 1, b = 0.01 to 1.0, c = 0.0 to 1.0, d = 0 to 1.0, 0.00 < e < 0.10 and f is a number to satisfy the valence state of the catalyst;

[1537] caddy sts of the formuia: where g is a number from 0.1 to 0.9, in many cases from 0.3 to 0.9, in other cases from 0.5 to 0.85, in some instances 0.6 to 0.8; h is a number from 0.04 to 0.9; i is a number from 0 to 0.5; j is a number from 0 to 0.5; and f is a number to satisfy the valence state of the catalyst; A is chosen from Ti. Ta. V, Nb, Hf, W. Y, Zn, Zr, Si and A1 or mixtures thereof; B is chosen from La, Ce, Pr, Nd, Sm, Sb, Sn, Bi, Pb, Tl, In, Te, Cr, Mn, Mo, Fe, Co, Cu, Ru, Rh, Pd, Pt. Ag, Cd, Os, Ir, Au, Hg, and mixtures thereof: D is chosen from Ca, K, Mg, Li, Na, Sr, Ba, Cs, and Rb and mixtures thereof; and O is oxygen;

[1538] catalysts of the formula:

Mo _a E.cG,0/ where E is chosen from Ba Ca, Cr, Mn, Nb Ta, Ti, Te, V, W and mixtnres thereof; G is chosen from Bi, Ce, Co,

Cu, Fe, K Mg, V, Ni P, Pb, Sb, Si, Sn, Ti U, and mixtures thereof; a = i ; k is 0 to 2; 1 = 0 to 2, with the proviso that the total value of 1 for Co, Ni, Fe and mixtures thereof is less than 0.5; and f is a number to satisfy the valence state of the catalyst;

[1539] catal sts of the formula: tyMo _^ Nb/fe _p Me ty where Me is chosen from Ta, Ti, W, Hf, Zr, Sb and mixtures thereof; m is from 0.1 to 3; n is from 0.5 to 1.5; o is from 0.001 to 3; p is from 0.001 to 5; q is from 0 to 2; and f is a number to satisfy the valence state of the catalyst; and

[1540] catal sts of the formula:

Mo _a V _; X _S Y ,Z,,M,O _/ where X is at least one of Nb and Ta; Y is at least one of Sb and Ni; Z is at least one of Te. Ga, Pd, W, Bi and Al; M is at least one of Fe, Co, Cu, Cr, Ti, Ce, Zr, Mn, Pb, Mg, Sn, Pt, Si, La, K, Ag and In; a=1.0 (normalized); r = 0.05 to 1.0; s = 0.001 to 1.0; t = 0.001 to 1.0; u = 0.001 to 0.5; v = 0.001 to 0.3; and f is a number to satisfy the valence state of the catal st. ί 1541] When choosing a catalyst, those skilled in the art can appreciate that catalysts may vary with respective to selectivity and activity. Some embodiments ofODH of ethane use mixed metal oxide catalysts that can provide high selectivity to ethylene without significant loss in activity. Non-limiting example catalysts are those of the formula:

Mq _a Vi.T e _c Nb _d Pd _f Or wherein a, h, c, d, e and f are the relative atomic amounts of the elements Mo, V, Te, Nb, Pd and O, respectively; and when a = 1, b = 0.01 to 1.0, c = 0 to 1.0, d = 0 to 1.0, 0.00 < e < 0.10 and f is a number to satisfy the valence state of the catalyst; il542j a mixed metal oxide having the empirical formula:

M06.5-7 0V3O1J where d is a number to satisfy the valence of the oxide; and a mixed metal oxide having the empirical formula:

1n1q6 25-7 25 V O where d is a number to satisfy the valence of the oxide.

[1543] In some embodiments, the catalyst may be supported on/agglomerated with a binder. Some binders include acidic, basic or neutral binder slurries ofTi(¼, Zr O ₂ AI ₂ O ₃ , AIO(OH) and mixtures thereof. Another useful binder includes M^Os. The agglomerated catalyst may be extruded in a suitable shape (rings, spheres, saddles etc.) of a size typically used in fixed bed reactors. When the catalyst is extruded, various extrusion aids known in the art can be used. In some cases, the resulting suppo rt may have a cumulative surface area of less than 35 m ² /g as measured by BET, in some cases, less than 20 nf/g, in other cases, less than 3 nri/g. and a cumulative pore volume from 0.05 to 0.50 em ³ /g.

[1544] ODH Reactor

[ 545] Any of the known reactor types applicable for the ODH of alkanes may be used with the methods disclosed herein. In some embodiments the methods may be used with conventional fixed bed reactors. In a typical fixed bed reactor reactants are introduced into the reactor at one end, flow' past an immobilized catalyst, products are formed and leave at the other end of the reactor. Designing a fixed bed reactor suitable for the methods disclosed herein can follow' techniques known for reactors of this type. A person skilled in the art would know which features are required with respect to shape and dimensions, inputs for reactants, outputs for products, temperature and pressure control, and means for immobilizing the catalyst.

[1546] In some embodiments, the use of inert non-catalytic heat dissipative particles can be used within one or more of the ODH reactors. In various embodiments, the heat dissipative particles are present within the bed and include one or more non catalytic inert particulates having a melting point at least 30 °C, in some embodiments at least 250 °C, in further embodiments at least 500 °C above the temperature upper control limit for the reaction; a particle size in the range of 0.5 to 75 rmn, in some embodiments 0.5 to 15, in further embodiments in the range of 0.5 to 8. in further embodiments in the range of 0.5 to 5 mm; and a thermal conductivity of greater than 30 W/iriK (watts/meter Kelvin) within the reaction temperature control limits. In some embodiments the particulates are metal alloys and compounds having a thermal conductivity of greater than 50 W/ ^' mK (watts/meter Kelvin) within tire reaction temperature control limits. Non-limiting examples of suitable metals that can be used in these embodiments include, but are not limited to, silver, copper, gold, aluminum, steel, stainless steel, molybdenum, and tungsten. [1547] The heat dissipative particles can have a particle size of from about 1 mm to about 15 mm. In some embodiments, the particle size can be from about 1 mm to about 8 mm. The heat dissipative particles can be added to the fixed bed in an amount from 5 to 95 wt. %, in some embodiments from 30 to 70 wt. %, in other embodiments from 45 to 60 wt. % based on the entire weight of the fixed bed. The particles are employed to potentially improve cooling homogeneity' and reduction of hot spots in the fixed bed by transferring heat directly to the walls of the reactor.

[1548] Additional embodiments include the use of a fluidized bed reactor, where the catalyst bed can be supported by a porous structure, or a distributor plate, located near a bottom end of the reactor and reactants flow through at a velocity sufficient to fluidize die bed (e.g. die catalyst rises and begins to swirl around in a fluidized maimer). The reactants are converted to products upon contact with the fluidized catalyst and the reactants are subsequently removed from the upper end of the reactor. Design considerations those skilled in the art can modify and optimize include, but are not limited to, the shape of the reactor, the shape and size of the distributor plate, the input temperature, the output temperature, and reactor temperature and pressure control [1549] Some embodime nts include using a combination of both fixed bed and fluidized bed reactors, each with the same or different OD3T catalyst. The multiple reactors can be arrayed in series or in parallel configuration, the design of which falls within tire knowledge of the worker skilled in the art.

[1550] Oxy gen/Alkane Mixture

[1551] Safety of the ODH process is a primary' concern. For that reason, in some embodiments, mixtures of one or more alkanes with oxygen should be employed using ratios that fall outside of the flammability' envelope of the one or more alkanes and oxygen. In some embodiments, the ratio of alkanes to oxygen may fall outside the upper flammability' envelope. In these embodiments, the percentage of oxygen in the mixture can be less than 30 wt. %, in some cases less than 25 wt. %, or in other cases less than 20 wt. %, but greater than zero.

[1552] In embodiments with higher oxygen percentages, alkane percentages can be adjusted to keep the mixture outside of the flammability' envelope. While a person skilled in the art would be able to determine an appropriate ratio level, in many cases the percentage of alkane is less than about 40 wt. % and greater than zero. As a non-limiting example, where the mixture of gases prior to ODH includes 20% oxygen and 40% alkane, the balance can be made up with an inert diluent. Non-limiting examples of useful inert diluents in this embodiment include, but are not limited to, one or more of nitrogen, carbon dioxide, and steam. In some embodiments, the inert diluent should exist in the gaseous sta te at the conditions within the reactor and should not increase the flammability' of the hydrocarbon added to the reactor, characteristics that a skilled worker would understand when deciding on which inert diluent to emplo . The inert diluent can be added to either of the alkane containing gas or the oxygen containing gas prior to entering the ODH reactor or may be added directly into the ODH reactor.

[1553] In sortie embodiments, the volumetric feed ratio of ox gen to ethane (O2/C2IT0) provided to die one or more ODH reactors can be at least about 0.3, in some cases at least about 0.4, and in other cases at least about 0.5 and can be up to about 1, in sortie cases up to about 0.9, in other cases up to about 0.8, in some instances up to about 07 and in oilier instances up to about 06 The volumetric feed ratio of oxygen to etliane can be any of the values or range between any of the values recited above.

[1554] In some embodiments, mixtures that fail within the flammability envelope may be employed, as a non- limiting example, in instances where the mixture exists in conditions that prevent propagation of an explosive event. In these non-limiting examples, the flammable mixture is created within a medium where ignition is immediately quenched. As a further non-limiting example, a user may design a reactor where oxygen and the one or more alkanes are mixed at a point where they are surrounded by a flame arresting material. Any ignition would be quenched by the surrounding material. Flame arresting materials include, but are not limited to, metallic or ceramic components, such as stainless steel walls or ceramic supports. In some embodiments oxygen and alkanes can be mixed at a low- temperature, where an ignition event would not lead to an explosion, then introduced into the reactor before increasing the temperature. The flammable conditions do not exist until the mixture is surrounded by the flame arrestor material inside of the reactor.

[1555] Carbon Monoxide Output

[1556] Carbon monoxide can be produced in the ODH reaction as a by-product of oxidation of the one or more alkanes. The carbon monoxide output is a function of the amount of carbon monoxide produced in the oxidative process.

[1557] Measuring the amount of carbon monoxide coming off the ODH reactor can be done using any means known in the art. For example, one or more detectors such as GC, 1R or Rahman detectors, are situated immediately downstream of the reactor to measure the carbon monoxide output. While not required, the output of other components may also be measured. These include, but are not limited to, the amounts of ethylene, unreacted ethane, acetylene, carbon dioxide and oxygen, and by-products such as acetic acid.

[1558] Carbon monoxide output can be stated using any metric commonly used in the art. For example, the carbon monoxide output can be described in terms of mass flow' rate (g/min) or volumetric flow' rate (cmVniin) In some embodiments, normalized selectivity can be used to assess the degree to which carbon monoxide is produced or consumed. In that instance the net mass flow rate of CO — the difference between the mass flow rate of CO leaving the ODH reactor — is normalized to the conversion of etliane, in essence describing what fraction of ethane is converted into carbon monoxide as opposed to ethylene, or other by-products such as acetic acid.

[1559] Many industrial processes, in addition to ODH, produce carbon monoxide which must be captured or flared where it contributes to the emission of greenhouse gases. Using the carbon monoxide mitigation steps disclosed herein converts most, if not all, carbon monoxide resulting from the ODH process to carbon dioxide. An advantage then is the ability to reduce or eliminate the amount of carbon monoxide produced in the ODH process in combination with other processes, such as thermal cracking.

[1560] Acetylene Output

[1561 ] Acetylene can be produced in the ODH reaction as a by-product of oxidation of the one or more alkanes. The acetylene output is a function of the amount of acetylene produced in die oxidative process.

[1562] Measuring the amount of acetylene coming off the ODH reactor can be done using any means known in the art. For example, one or more detectors such as GC, IR, or Rahman detectors, are situated immediately downstream of the reactor to measure the acetylene output. While not required, the output of other components may also be measured. These include but are not limited to the amounts of ethylene, unreacted ethane, carbon monoxide, carbon dioxide and oxygen, and by-products such as acetic acid.

[1563] Acetylene output can be stated using any metric commonly used in the art. For example, the acetylene output can be described in terms of mass flow rate (g/min), volumetric flow rate (cnrVmin) or volumetric parts per million (vppm). in some embodiments, normalized selectivity can be used to assess the degree to which acetylene is produced or consumed. In that instance the net mass flow rate of acetylene — the difference between the mass flow rate of acetylene leaving the ODH reactor-— is normalized to the conversion of ethane, in essence describing what fraction of ethane is converted into acetylene as opposed to ethylene, or other by-products such as acetic acid.

[1564] Using the acetylene mitigation steps disclosed herein reacts most, if not all, acet lene resulting from the ODH process. An advantage then is die ability to reduce or eliminate the amount of acetylene produced in the ODH process in combination with other processes, such as thermal cracking and eliminate downstream unit operations in an ODH -type process.

[1565] Addition of Steam

[1566] The amount of steam added to the ODH process affects the degree to which carbon dioxide acts as an oxidizing agent. In some embodiments steam may be added directly to the ODH reactor, or steam may be added to the individual reactant components —the lower alkane, oxygen, or inert diluent — or combinations thereof, and subsequently introduced into the ODH reactor along with one or more of the reactant components. Alternatively, steam may be added indirectly as water mixed with either the lower alkane, oxygen or inert diluent, or a combination thereof, with the resulting mixture being preheated before entering the reactor. When adding steam indirectly as water the preheating process should increase the temperature so that the water is entirely converted to steam before entering the reactor.

[1567] Increasing the amount of steam added to a reactor increases the degree to which carbon dioxide acts as an oxidizing agent. Decreasing the amount of steam added to the reactor decreases the degree to which carbon dioxide acts as an oxidizing agent. In some embodiments, a user monitors the carbon dioxide output and compares it to a predetermined target carbon dioxide output. If the carbon dioxide output is above the target a user can then increase the amount of steam added to the ODH process. If the carbon dioxide output is below the target a user can decrease the amount of steam added to the ODH process, provided steam lias been added. Setting a target carbon dioxide output level is dependent on the requirements for the user. In some embodiments, increasing the steam added will have the added effect of increasing the amount of acetic acid and other by-products produced in the process. A user that is ill equipped to separa te out larger amounts of acetic acid from the output of the ODH may instead reduce steam levels to a minimum, while a user that desires a process that consumes carbon dioxide may choose to maximize the amount of steam that can be added.

[1568] In some embodiments, the amount of steam added to the one or more ODH reactors can be up to about 50 wt. %, in some circumstances up to about 40 wt. %, in some eases up to about 35 wt. %, in other cases up to about 30 wt. %, and in some instances up to about 25 wt. % and can be zero, in some cases at least 0.5 wt. %, in other cases at least 1 wt. %, in other cases at least 5 wt. %, in some instances at least 10 wt. % and in other instances at least 15 wt % of the stream entering the one or more ODH reactors. The amount of steam in the stream entering the one or more ODH reactors can he any value or range between any of the values recited above.

[1569] In some embodiments, when using two or more ODH reactors a user may choose to control carbon dioxide output in only one, or less than the whole complement of reactors. For example, a user may opt to maximize carbon dioxide output of an upstream reactor so that the higher level of carbon dioxide can be part of the inert diluent for the subsequent reactor. In that instance, maximizing carbon dioxide output upstream minimizes the amount of inert diluent that would need to be added to the stream prior to the next reactor.

[1570] There is no requirement for adding steam to an ODH process as it is one of many alternatives for the inert diluent. For processes where no steam is added, Site carbon dioxide output is maximized under the conditions used with respect to ethane, oxygen and inert diluent inputs. Decreasing the carbon dioxide output can then be a matter of adding steam to the reaction until carbon dioxide output drops to the desired level. In embodiments where oxidative dehydrogenation conditions do not include addition of steam, and the carbon dioxide output is higher titan the desired carbon dioxide target level, steam may be introduced into the reactor while keeping relative amounts of the main reactants and inert diluent — lower alkane, oxygen and inert diluent — added to the reactor constant, and monitoring the carbon dioxide output, increasing the amount of steam until carbon dioxide decreases to the target level.

[1571] In some embodiments, a carbon dioxide neutral process can be achieved by increasing steam added so that any carbon dioxide produced in the oxidative dehydrogenation process can then be used as an oxidizing agent such that there is no net production of carbon dioxide. Conversely, if a user desires net positive carbon dioxide output then the amount of steam added to the process can be reduced or eliminated to maximize carbon dioxide production. As the carbon dioxide levels increase there is potential to reduce oxygen consumption, as carbon dioxide is competing as an oxidizing agent. The skilled person would understand that using steam to increase the degree to which carbon dioxide acts as an oxidizing agent can impact oxygen consumption. The implication is that a user can optimize reaction conditions with lower oxygen contributions, which may assist in keeping mixtures outside of flammability limits

[1572] In some embodiments, the stream exiting the one or more ODH reactors can be treated to remove or separate water and water soluble hydrocarbons from the stream exiting the one or more ODH reactors. In some embodiments, this stream is fed to the second reactor.

[1573] Acetic Acid Removal

[1574] Prior to being fed to the second reactor, the stream exiting the one or more ODH reactors is directed to quench tower or acetic acid scrubber, which facilitates removal of oxygenates, such as acetic acid, and water via a bottom outlet. A stream containing unconverted lower alkane (such as ethane), corresponding alkene (such as ethylene), unreacted oxygen, carbon dioxide, carbon monoxide, optionally acetylene and inert diluent, are allowed to exit die scrubber and are fed to the second reactor.

[1575] The oxygenates removed via the quench tower or acetic acid scrubber can include carbox lic acids (for example acetic acid) aldehydes (for example acetaldehyde) and ketones (for example acetone). The amount of oxygenate compounds remaining in the stream exiting the scrubber and fed to the second reactor will often be zero, i.e, below the detection limit for analytical test methods typically used to detect such compounds. When oxygenates can be detected they can be present at a level of up to about 1 per million by volume (ppmv), in some cases up to about 5 ppmv, in other cases less than about 10 ppmv, in some instances up to about 50 ppmv and in other instances up to about 100 ppmv and can be present up to about 2 vol. %, in some cases up to about 3 vol. %, and in other cases up to about 1,000 ppmv. The amount of oxygenates or acetic acid in the stream exiting the scrubber and fed to the second reactor can be any value, or range between any of the values recited above.

[1576] The Second Reactor

[1577| In some embodiments, the ODH reactor (or reactors) can provide a stream containing at least a small amount of oxygen remaining as reactor effluent. In some embodiments, the oxygen can provide a benefit to the ODH reactor product gas. In some embodiments, when the ODH catalyst is exposed to mi ox gen free reducing environment at elevated temperature it may become permanently degraded. In other embodiments, if Site level of oxygen in the product gas from the ODH reactor contains less than about 1 ppm of oxygen, most, if not all, of the one or more alkanes are converted to one or more alkenes in the inlet portion of the reactor and a large portion of the reactor cataly st bed is not utilized

[1578] In some embodiments, oxygen in the ODH reactor product gas causes serious safety and operational issues in the downstream equipment, as a non-limiting example, at the first compression stage of an ODH process. This process safety consideration presents a need to remove oxygen to a very low or non-detectable level before the product gas is compressed.

[1579] One method used to reduce/eliminate oxygen in the ODH product gas focuses on catalytically combusting a small portion of the ODH product gas to the complete consumption of any residual oxygen. This approach is viable, however, in many cases it is undesirable, because it increases the overall oxygen consumption in the ODH process and, in the non-limiting example of the alkane being ethane, reduces overall process selectivity' toward ethylene.

[1580] Techniques are provided for a process where the ODH reaction can proceed with partial consumption of CO _? (CO2 can act as an oxidizing agent, and be converted to CO), reducing overall oxygen consumption in the process by providing a portion of the required oxygen from CO2. In some embodiments, more oxygen passes through the catalyst bed unconverted when CO2 is provided and acts as an oxidizing agent.

[1581 ] Oxidation of Carbon Monoxide

[1582] In the process, the ODH reactor product stream is fed to the second reactor, which contains a catalyst that includes one or more selected from a group 11 metal, a group 4 metal, a group 7 metal, a group 9 metal, a lanthanide metal, and an actinide metal and/or their corresponding metal oxides capable of converting a t least a portion of the carbon monoxide to carbon dioxide. The carbon dioxide can be recycled to the ODH reactor to act as an oxidizing agent as described above.

[1583] In sortie embodiments, the group 11 metal can be selected front copper, silver gold and combinations thereof. In some embodiments, the group 11 metal is silver or copper.

[1584] In some embodiments, the group 4 metal can be selected from titanium, zirconium, hafnium, rutherfordium and combinations thereof. In some embodiments, the group 4 metal is zirconium. f 1585] In some embodiments, the group 7 metal can be selected from manganese, technetium, rhenium, bohrium and combinations thereof. In some embodiments, the group 7 metal is manganese.

[1586] In some embodiments, the group 9 metal can be selected from cobalt, rhodium, iridium, meitemium and combinations thereof. In some embodiments, the group 9 metal is cobalt.

[1587] in some embodiments, the lanthanide metal can be selected from La, Ce, Pr, Nd, Pm, Sm, Eu, Gd, Tb, Dy, ho, Er, Tm, Yb and combinations thereof. In some embodiments, the lanthanide metal is Cerium.

[1588] in some embodiments, the actinide metal can be selected from Ac, Th, Ps, U, Np, Pu, Am, Cm, Bk, Cf, Es. Fm, Md, No and combinations thereof In some embodiments, the actinide metal is thorium.

[1589] In some embodiments, the second reactor catalyst, in some cases a group 11 metal, is used in conjunction with a promoter. In some embodiments, the promoter is selected from one or more of the lanthanide and actinide metals (as defined above) and their corresponding metal oxides. In some embodiments, the promoter is selected from one or more of the lanthanide metals and their corresponding metal oxides. In some embodiments, the promoter includes cerium and its corresponding metal oxides.

[1590] In some embodiments, the second reactor catalyst, in some cases a group 11 metal, and optional promoter are provided on a support. The support is typically an inert solid with a high surface area to which the second reactor catalyst and optional promoter can be affixed. In many embodiments, the support includes Si, Ge, Sn, their corresponding oxides and combinations thereof.

[1591] In some embodiments, non-limiting examples of suitable second reactor catalysts with optional promoters and supports include Ag/SiO _? , AgCe0 ₂ /Si0 _2, AgZrCK/SiCf, AgCoiO- _' ./SiO _?, Cu/Si0 _2, CuCe0 ₂ /Si0 ₂ , CuZrO /SiO , CuCo Cf/SiCf and combinations thereof

[1592] In some embodiments, non-limiting examples of suitable second reactor catalysts with optional promoters and supports include AgCe0 ₂ /Si0 ₂ , AgZrCVSiO and combinations thereof.

[ 593] In some embodiments, the second reactor catalyst includes silver the optional promoter includes cerium and the support includes Si(¾.

[1594] In some embodiments, the second reactor catalyst includes copper, the optional promoter includes cerium and the support includes Si(¾.

[1595] In some embodiments, when oxidation of carbon monoxide is preferentially desired, the second reactor catalyst includes manganese, the optional promoter includes cerium and the support includes Si0 ₂ .

[1596] In some embodiments, the group 11 metal with optional promoter and optional support can be used in a process where 1) some oxy gen is in the stream leaving the QDH reactor; 2) the temperature in the stream is decreased; 3) the cooled stream is fed to an acetic acid scrubber; 4) the stream from the acetic acid scrubber is fed to reactor 2 as described above, where most or all of the residual Oi is consumed and CO is converted to CO2; and 5) optionally, the C0 ₂ is recycled back to the ODH reactor.

[1597] In some embodiments, the amount of ox gen in the stream leaving the ODH reactor in 1) can be at least about 80 ppm, in some cases at least about 100 ppm, in other cases at least about 150 ppm and in some instances at least about 200 ppm and can be up to about 5 wt. %, in some cases up to about 4 wt. %, in other cases up to about 3 wt. %. in some instances up to about 2 wt. %, in other instances up to about i wt. %, and in particular situations up to about 500 ppm. The amount of oxygen in the stream leaving the ODH reactor in i) can be any of the values or range between any of the values recited above.

[1598] In some embodiments, when there is oxygen in the stream leaving the second reactor (in some instances the amount of oxygen will be undetectable or zero ppm), the amount of oxygen in the stream leaving the second reactor can be at least about 1 ppm, in some cases at least about 2 ppm, in other cases at least about 3 ppm and in some instances at least about 5 ppm and can be up to about 3 wt. %, in some cases up to about 0.9 wt. %, in other cases up to about 0.8 wt. %, in some instances up to about 0.7 wt. %, in other instances up to about 0.6 wt. %, and in particular situations up to about 0.5 wt. %. The amount of oxygen in the stream leaving the second reactor can be any of the values or range between any of the values recited above.

[1599] In some embodiments, the amount of carbon monoxide in the stream leaving the ODH reactor in 1) can be at least about 100 ppm, in some cases at least about 200 ppm, in oilier cases at least about 300 ppm and in some instances at least about 400 ppm and can be up to about 10 wt. %, in some cases up to about 9 wt. %, in other cases up to about 8 wt. %, in some instances up to about 7 wt. %, in other instances up to about 6 wt. %. and in particular situations up to about 5 wt. %. The amount of carbon monoxide in the stream leaving the ODH reactor in 1) can be any of the values or range between any of the values recited above.

[1600] In some embodiments, when there is carbon monoxide in the stream leaving the second reactor (in some instances the amount of carbon monoxide will be undetectable or zero ppm), the amount of carbon monoxide in the stream leaving the second reactor can be at least about 1 ppm, in some cases at least about 2 ppm, in other cases at least about 3 ppm and in some instances at least about 5 ppm and can be up to about 8 wt. %, in some cases up to about 7 wt. %, in other cases up to about 6 wt. %, in some instances up to about 5 wt. %, in other instances up to about 4 wt. %, and in particular situations up to about 3 wt. %. The amount of carbon monoxide in the stream leaving the second reactor can be any of the values or range between any of the values recited above.

[1601] In some embodiments, temperature in the second reactor can be at least about 40, in some cases at least about 45, in other eases at least about 50 and in some instances at least about 55 °C and can be up to about 200 °C, in some cases up to about 150 °C, in some cases up to about 120 °C, in some cases up to about 90 °C, in some eases up to about 85 °C, in some eases up to about 80 °C, in some cases up to about 75 °C and in some cases up to about 70 °C. The temperature of second reactor can be any temperature value or range between any of the temperature values, including a temperature gradient within the second reactor, recited above.

[1602] Acetylene Elimination

[1 03] In the process, the ODH reactor product stream is fed to the second reactor, which contains a catalyst that includes one or more selected from a group 11 metal, a group 4 metal, a group 9 metal, a lanthanide metal, and an ac tinide metal and/or their corresponding metal oxides capable of reacting at least a portion of the acetylene.

[T 604] In some embodiments, the group 11 metal can be selected from copper, silver gold and combinations thereof. In some embodiments, the group 11 metal is silver.

[1605] In some embodiments, the group 4 metal can be selected from titanium, zirconium, hafnium, rutlierfordiimi and combinations thereof. In some embodiments, the group 4 metal is zirconium. f 1606] In some embodiments, the group 9 metal can be selected from cobalt, rhodium iridium meitemium and combinations thereof. In some embodiments, the group 9 metal is cobalt.

[1607] In some embodiments, the lanthanide metal can be selected from La, Ce, Pr, Nd, Pm, Sm, Eu, Gd, Tb, Dy, ho, Er, Tm, Yb and combinations thereof. In some embodiments, the lanthanide metal is Cerium.

[1608] in some embodiments, the actinide metal can be selected from Ac, Th, Ps, U, Np, Pu, Am, Cm, Bk, Cf, Es, Fm, Md, No and combinations thereof. In some embodiments, the actinide metal is thorium.

[1609] in some embodiments, the second reactor catalyst, in some cases a group 11 metal, is used in conjunction with a promoter. In some embodiments, the promoter is selected from one or more of the iantlianide and actinide metals (as defined above) and their corresponding metal oxides. In some embodiments, the promoter is selected from one or more of die iantlianide metals and their corresponding metal oxides. In some embodiments, the promoter includes cerium and its corresponding metal oxides.

[1 10] In some embodiments, die second reactor catalyst, in some cases a group 11 metal, and optional promoter are provided on a support. The support is typically an inert solid with a high surface area, to which the second reactor catalyst and optional promoter can be affixed. In many embodiments, the support includes Si. Ge, Sn, their corresponding oxides and combinations thereof.

[1611] In some embodiments, non-limiting examples of suitable second reactor catalysts with optional promoters and supports include Ag/SiO?, AgCeO/SiO _z, AgZrCb/SiO?, AgCoiO- _' ./SiO?., Cu/SiO?., CuCeO/SiOi, CuZrCti/SiCb, CuCo Cf/SiCb and combinations thereof.

[1612] In some embodiments, non-limiting examples of suitable second reactor catal sts with optional promoters and supports include AgCeO /SiO?., AgZrOi/SiO? and combinations thereof.

[1613] In some embodiments, the second reactor catalyst includes silver the optional promoter includes cerium and the support includes Si(¾.

[1614] In some embodiments, the second reactor catalyst includes copper, the optional promoter includes cerium and the support includes Si(¾.

[1615] In some embodiments, the group 11 metal with optional promoter and optional support can be used in a process where 1) some acetylene is in the stream leaving the ODH reactor; 2) the temperature in the stream is decreased; 3) the cooled stream is fed to an acetic acid scrubber; 4) the stream from the acetic acid scrubber is fed to reactor 2 as described above, where most or all of the acetylene is consumed and CO is oxidized to C0 ₂ ; and 5) optionally, the CO2 is recycled back to the ODH reactor.

[1616] In some embodiments, when there is acetylene in the stream leaving the ODH reactor (in some instances the amount of acetylene will be undetectable or zero vpprn), the amount of acetylene in the stream leaving the ODH reactor in 1) can be at least about 1 vppm, in some cases at least about 2 vppm, in other cases at least about 5 vppm and in some instances at least about 10 vppm and can be up to about 1000 vppm, in some cases up to about 750 vppm, in other cases up to about 500 vppm, in some instances up to about 400 vppm, in other instances up to about 300 vppm, and in particular situations up to about 300 vppm. The amount of acetylene in the stream leaving tire ODH reactor in 1) can be any of the values or range between any of tire values recited above. f 1617] In some embodiments, the amount of acetylene in the stream leaving the second reactor will be less than the amount entering the second reactor and, in many instances, the stream exiting the second reactor will be substantially free of acetylene.

[1618] In some embodiments, when there is acetylene in the stream leaving the second reactor (in many instances the amount of acetylene will he unde tectable, less than 1 vppm, or zero vppm), the amount of acetylene in the stream leaving the second reactor can be at least about 1 vppm, in some cases at least about 2 vppm, in other cases at least about 3 vppm and in some instances at least about 5 vppm and can be up to about 100 vppm, in some cases up to about 50 vppm, in other cases up to about 25 vppm, in some instances up to about 20 vppm, in other instances up to about 15 vppm, and in particular situations up to about 10 vppm. The amount of acetylene in the stream leaving the second reactor can be any of the values or range between any of the values recited above.

[1619] In some embodiments, temperature in the second reactor can be at least about 40 °C, in some cases at least about 45 °C, in some cases at least about 50 °C and in some cases at least about 55 °C and can be up to about 200 °C, in some cases up to about 150 °C, in some cases up to about 120 °C, in some cases up to about 90 °C, in some cases up to about 85 °C, in some cases up to about 80 °C, in some cases up to about 75 °C and in some cases up to about 70 °C. The temperature of second reactor can be any temperature value or range between any of the temperature values, including a temperature gradient within the second reactor, recited above.

[1620] In some embodiments, the stream from the ODH reactor is cooled to a lower temperature prior to being fed to an acetic acid scrubber (as described below). The temperature of the stream prior to entering the acetic acid scrubber can be at least about 40 °C. in some cases at least about 45 °C, and in some cases at least about 50 °C and can be up to about 90 °C, in some cases up to about 85 °C, in some cases up to about 80 °C, in some cases up to about 75 °C and in some cases up to about 70 °C. The temperature of the ODH reactor product stream fed to an acetic acid scrubber can be cooled to any temperature value or range between any of the temperature values recited above.

[1621] In some embodiments, the configuration described above can allow for the size of the air separation plant to be reduced, as well as improving the life of the ODH catalyst, by allowing it to be exposed to an oxygen containing environment at all times. In additional embodiments the configuration described above can improve the reliability and safety of the ODH reactor and downstream equipment.

[1622] In some embodiments, the net C0 ₂ generation in the process described herein can be optimized to be zero. In these embodiments, the need to flare off any CO2 (with some amount of alkane/alkene) from the (XVrecyie loop as described herein. In these embodiments, the total process yield of alkane to alkene can be improved.

[1623] ODH Complex

[1 24] In some embodiments, the chemical complex (shown in one embodiment schematically in Figure 2) includes, in cooperative arrangement, an ODH reactor 202, a quench tower or acetic acid scrubber 204, a second reactor 206 (as described herein), an amine wash tower 208, a drier 210, a distillation tower 212, and an oxygen separation module 214. ODH reactor 202 includes an ODH catalyst capable of catal zing, in the presence of oxygen which may be introduced via oxygen line 216, the oxidative dehydrogenation of alkanes introduced via alkane line 218. Although second reactor 206 is shown directly after quench tower or acetic acid scrubber 204, it can be placed further downstream. In some eases, the process configuration can be more energy efficient if second reactor 206 is placed after the input stream has been compressed.

[1625] The ODH reaction may also occur in the presence of an inert diluent, such as carbon dioxide, nitrogen, or steam, that is added to ensure the mixture of oxygen and hydrocarbon are outside of flammability limits. Determination of whether a mixture is outside of the flammability limits, for the prescribed temperature and pressure, is within the knowledge of the skilled worker. An ODH reaction that occurs within ODH reactor 202 may also produce, depending on the catalyst and the prevailing conditions within ODH reactor 202, a variety of other products which may include carbon dioxide carbon monoxide, oxygenates, and water. These products leave ODH reactor 202, along with unreacted alkane, corresponding alkene, residual oxygen, carbon monoxide and inert diluent, if added, via ODH reactor product line 220.

[1626] ODH reactor product line 220 is directed to quench tower or acetic acid scrubber 204 which quenches the products from product line 220 and facilitates removal of oxygenates and water via quench tower bottom outlet 222. Unconverted lower alkane corresponding alkene, unreacted oxygen, carbon dioxide, carbon monoxide, and inert diluent added to quench tower 204 exit through quench tower overhead line 224 and are directed into second reactor 206.

[1627] Second reactor 206 contains the group 11 metal with optional promoter and optional support as described above, which causes unreacted oxygen to react with carbon monoxide to form carbon dioxide or, optionally, reacts acetylene to reduce or eliminate it. In second reactor 206, most or all of the unreacted oxygen and acetylene is consumed. All or a portion of the carbon dioxide in reactor 206 can be recycled back to ODH reactor 202 via recycle lines 226 and 227 to act as an oxidizing agent as described above. The remaining unconverted lower alkane, co rresponding alkene, unreacted oxygen (if present), all or part of the carbon dio xide, carbon monoxide (if present), and inert diluent are conveyed to amine wash tower 208 via line 228.

[1628] Any carbon dioxide present in line 228 is isolated by amine wash tower 208 and captured via carbon dioxide bottom outlet 230 and may be sold or, alternatively, may be recycled back to ODH reactor 202 as described above. Constituents introduced into amine wash tower 208 via line 228 other than carbon dioxide, leave amine wash tower 208 through amine wash tower overhead line 232 and are passed through a dryer 210 before being directed to distillation tower 212, where C2/C2 ₊ hydrocarbons are isolated and removed via C2/C2 ₊ hydrocarbons bottom outlet 236. The remainder includes mainly €1 hydrocarbons, including remaining inert diluent and carbon monoxide (if any), which leave distillation tower 212 via overhead stream 238 and is directed to oxygen separation module 214.

[1 29] Oxygen separation module 214 includes a sealed vessel having a retentate side 240 and a permeate side 242, separated by oxygen transport membrane 244. Overhead stream 238 may be directed into either of retentate side 240 or permeate side 242. Optionally, a flow controlling means (for example flow controlling means 6326 shown in Figure 63D) may be included that allows for flow into both sides at varying levels. In that instance an operator may choose what portion of the flow from overhead stream 238 enters retentate side 240 and what portion enters permeate side 242. Depending upon conditions an operator may switch between the two sides, to allow' equivalent amounts to enter each side, or bias the amount directed to one of the two sides. Oxygen separation module 214 also includes air input 246 for the introduction of atmospheric air, or other oxygen containing gas, into the retentate side 240. Combustion of products introduced into retentate side 240, due to the introduction of oxygen, may contribute to raising the temperature of oxygen transport membrane 244 to at least about 850 °C so that oxygen can pass from retentate side 240 to permeate side 242. Components within the atmospheric air, or other oxygen containing gas, other than oxygen, cannot pass from retentate side 240 to permeate side 242 and can only leave oxy gen separation module 214 via exhaust 248.

[1630] As a result of oxygen passing from retentate side 240 to permeate side 242, there is separation of oxygen from atmospheric air. or other oxygen containing gas, introduced into retentate side 240. The result is production of oxygen enriched gas on permeate side 242, which is then directed via ox gen enriched bottom line 227 to ODH reactor 202, either directly or in combination with ox gen line 70 (as shown in Figure 2). When overhead stream 238 is directed into retentate side 240 the degree of purity of oxygen in oxygen enriched bottom line 227 can approach 99%. Conversely, when overhead stream 238 is directed into permeate side 242 the degree of purity of oxygen in oxygen enriched bottom line 227 is lower, with an upper limit ranging from 80% - 90% oxygen, the balance in the fo rm of carbon dioxide, water, and remaining inert diluent, all of which do not affect the ODH reaction as contemplated by the present disclosure and can accompany the enriched oxygen into ODH reactor 202. Water and carbon dioxide can be removed by quench tower 204 and amine wash tower 208 respectively. In some embodiments, some or all of the carbon dioxide can be captured for sale as opposed to being flared where it contributes to greenhouse gas emissions. In other embodiments, when carbon dioxide is used in the ODH process, any carbon dioxide captured in the amine wash can be recycled back to ODH reactor 202.

[1631] Oxygen transport membrane 244 is temperature dependent, only allowing transport of oxygen when the temperature reaches at least about 850 °C. In some embodiments the components in overhead stream 238 by themselves are not capable, upon combustion in the presence of oxygen, to raise the temperature of oxygen transport membrane 244 to the required level. In such embodiments, the chemical complex also includes fuel enhancement line 250, upstream of oxygen separation module 214, where combustible fuel, as a non-limiting example methane, may be added to supplement the combustible products from overhead stream 238.

[1632] In some embodiments, the oxygen separation module 214 is a tube (an example is depicted schematically in Figure 63 A, 63B, 63C, and 63D). For example, the oxygen separation module 214 can be substantially the same as any of the embodiments of the oxygen separation module 6306 shown in Figs. 63 A, 63B, 63C, or 63D. The oxygen transport membrane 244 (membrane 6319 of Figs. 63 A, 63B, 63C, or 63D) can be a tube and can fit inside a larger tube (tube 6327 of Figs. 63 A, 63B, 63C, or 63D) which forms the outer wall of oxygen separation module 214. The annular space between the larger tube 6327 and oxygen transport membrane 6319 corresponds to the retentate side 240, while the space within oxygen transport membrane 6319 corresponds to the permeate side 242. Material suitable for construction of the outer wall include those resistant to temperatures that exceed 850 °C and approach 1000 °C, selection of which falls within tire knowledge of the skilled worker.

[1633] The present disclosure contemplates tire inlet for the overhead stream 238 (stream 6316 of Figs. 63 A. 63B. 63C, or 63D) entering the oxygen transport module 214 into either of the permeate side 242 (an example is depicted schematically in Figure 63 A) or the retentate side 240 (an example is depicted schematically in Figure 63C). In some embodiments, oxygen separation module 214 can have C _L hydrocarbon containing line directed to the retentate side 240. The present disclosure also contemplates the use of a valve (valve 6326 of Fig. 63D) for switching between directing the overhead stream 238 to the retentate side 240 or the permeate side 242 (an example is depicted schematically in Fig. 63D). This would allow an operator to choose which of the sides, permeate 242 or retentate 240, that the overhead stream 238 is directed to.

[1634] In some embodiments, a concent for ODH processes is the mixing of a hydrocarbon with oxygen. Under certain conditions the mixture may be unstable and lead to an explosive event. Accordingly, a hydrocarbon containing gas may be mixed with an oxygen containing gas in a flooded mixing vessel. By mixing in this way pockets of unstable compositions are surrounded by a non-flammable liquid so that even if an ignition event occurred it would be quenched immediately. Provided addition of the gases to the ODH reaction is controlled so that homogeneous mixtures fall outside of the flammability envelope, for the prescribed conditions with respect to temperature and pressure, the result is a safe homogeneous mixture of hydrocarbon and oxygen.

[1635] In some embodiment, there is a flooded gas mixer 302 upstream of ODH reactor 202 (an example is depicted in Fig. 3). In this instance, oxygen line 216 and alkane line 218 feed directly into flooded gas mixer 302. A homogeneous mixture that includes hydrocarbon and oxygen, and optionally an inert diluent, can be introduced into ODH reactor 202 from flooded gas mixer 302 via mixed line 304 (Figure 3). Oxygen enriched bottom line 227 may feed directly into or in combination with oxygen line 216 into flooded gas mixer 302.

[1636] The tempe rature of the conte nts within product line 220 i n a typical ODH process ca reach about 450

°C. It can be desirable to decrease the temperature of the stream before introduction into quench tower or acetic acid scrubber 204 as described previously. In that instance, the present disclosure contemplates the use of a heat exchanger immediately downstream of each ODH reactor 202 and immediately upstream of quench tower 204. [1637] As mentioned previously with reference to Figure 2, in the ODH process configuration depicted in Figure 3, although second reactor 206 is shown directly after quench tower or acetic acid scrubber 204, it can be placed further downstream. In some cases, the process configuration can be more energy efficient if second reactor 206 is placed after the input stream has been compressed.

[1638] In some embodiments, the olefins produced using the one or more ODH reactors, or any of the processes or complexes described herein, can be used to make various olefin derivatives. Olefin derivatives include, but are not limited to polyethylene, polypropylene, ethylene oxide, propylene oxide, polyethylene oxide, polypropylene oxide, vinyl acetate, vinyl chloride, acrylic esters (e.g. methyl methacrylate), thermoplastic elastomers, thermoplastic olefins and blends and combinations thereof.

[1639] In some embodiments, eth lene and optionally a-olefins are produced in the one or more ODH reactors, or any of the processes or complexes described herein, and are used to make polyethylene. The polyethylene made from the ethylene and optional a-olefins described herein can include homopol mers of ethylene, copolymers of ethylene and a-olefins, resulting in HOPE, MDPE, LDPE, LLDPE and VLDPE.

[1640] The pol eth lene produced using Site ethylene and optional a-olefins described herein can be produced using any suitable polymerization process and equipment. Suitable ethylene polymerization processes include but are not limited to gas phase polyethylene processes, high pressure polyethylene processes, low pressure polyethylene processes, solution polyethylene processes, slurry polyethylene processes and suitable combinations of the above arranged either in parallel or in series.

[1641] Various tools commonly used for chemical reactors, including flowmeters, compressors, valves, and sensors for measuring parameters such as temperature, pressure and flow rates can be used. It is expected that the person of ordinary skill in the art would include these components as deemed necessary' for safe operation.

[1642] Techniques are provided for a method of converting one or more alkanes to one or more alkenes. The method includes: providing a first stream including one or more alkanes and oxygen to an oxidative deh drogenation reactor; converting at least a portion of the one or more alkanes to one or more alkenes in the oxidative dehydrogenation reactor to provide a second stream exiting the oxidative deh drogenation reactor including one or more alkanes, one or more alkenes, oxygen, one or both of carbon dioxide and carbon monoxide and optionally acetylene; and providing the second stream to a second reactor containing a catalyst including a group 11 metal and optionally a promoter including Ce0 ₂ . ZrO 2 and combinations thereof supported onSiCb to convert a least a portion of the carbon monoxide to carbon dioxide and reacting any acetylene. This aspect can include any combination of the following features.

[1643] In some embodiments, the one or more alkanes include ethane.

[1644] In some embodiments, the one or more alkenes include ethylene.

[1645] In some embodiments, the first stream can include one or more inert diluents, an oxygen containing gas and a gas containing one or more lower alkanes.

[1646] In some embodiments, the second stream can include one or more unreacted lower alkanes; one or more lower alkenes; oxy gen; one or more inert diluents; carbon dioxide; carbon monoxide; acetic acid; and water.

[1647] In some embodiments, the oxidative dehydrogenation reactor includes a single fixed bed type reactor.

[1648] In some embodiments, the oxidative dehydrogenation reactor includes a single fluidized bed type reactor.

[1649] In some embodiments, the oxidative dehydrogenation reactor includes a swing bed type reactor arrangement and/or a moving bed reactor.

[1650] In some embodiments, the group 11 metal is selected from the group of copper, silver, gold and combinations thereof.

[1651 ] In some embodiments, the group 11 metal is silver or copper.

[1 52] In some embodiments, the catalyst in the second reactor includes Ag/Si0 ₂ , AgCe0 ₂ /Si0 _2,

AgZrCb/SiCb Cu SiCb, CuCe0 ₂ /Si0 ₂ , CuZr0 ₂ /Si0 ₂ , CuCo ₃ 0 ₄ /Si0 ₂ and combinations thereof.

[1653] In some embodiments, an acetic acid scrubber is placed between the oxidative dehy drogenation reactor and the second reactor.

[1654] In some embodiments, the temperature in the second reactor is from 40 to 100 ^C C.

[1655] Techniques are provided for a chemical complex for oxidative dehydrogenation of lower alkanes, the chemical complex includes in cooperative arrangement:

[1656] (Qi) at least one oxidative dehydrogenation reactor, that includes an oxidative deh drogenation catalyst designed to accept, optionally in die presence of an inert diluent, an oxygen containing gas and a lower alkane containing gas, and to produce a product stream that includes the corresponding alkene and possibly one or more of: unreacted lower alkane; oxy gen; inert diluent; carbon oxides, including carbon dioxide and carbon monoxide; oxygenates, including acetic acid, acryiic acid, maleic anhydride and maleic acid; and water;

[1657] (Qii) a quench tower for quenching the product stream and for removing water and soluble oxygenates from the product stream;

[1658] (Qiii) an oxidation reactor for oxidizing carbon monoxide to carbon dioxide and optionally reacting acetylene;

[16591 (Qiv) an amine wash for removing carbon dioxide from the product stream; i 16601 (Qv) a dryer for removal of water from the product stream;

[1661] (Qvi) a distillation tower for removing C2/C2 ₊ hydrocarbons from the product stream to produce an overhead stream enriched with Ci hydrocarbons;

[1662] (Qvii) optionally, a means for introducing a combustible fuel into the overhead stream; and [1663] (Qviii) art oxygen separation module that includes: an oxygen transport membrane housed inside a sealed vessel and having a retentate side and a permeate side; a first inlet for introducing the overhead stream, combustible fuel, or both into the retentate side; a second inlet for introducing the overhead stream, combustible fuel, or both into the permeate side; an air inlet for introducing air into the retentate side; a exhaust for discharge of oxygen depleted air and combustion products from the retentate side; and an outlet for removing oxygen enriched gas and combustion products from the permeate side, where the components in (Qi) through (Qviii) are connected in series in the sequence described, the overhead stream from (Qvi) may be directed into the retentate side, the permeate side, or both the retentate side and the permeate side and the oxygen enriched gas and combustion products from the permeate side may be directed back to (Qi) as or part of the oxygen containing gas introduced into the at least one oxidative dehydrogenation reactor. This aspect can include any combination of the following features.

[ 664] In some embodiments, the chemical complex includes a non-flammable liquid flooded gas mixer for premixing the oxygen containing gas the lower alkane containing gas and inert gases prior to introduction into the at least one oxidative dehydrogenation reactor. In many aspects when the non-flammable liquid inside the complex is water, then the genera ted saturated steam in the overhead of the mixer can also act as an inert diluent (in addition to the fed inert gases).

[1665] In some embodiments, the oxidative dehydrogenation catalyst includes one or more of the mixed metal oxides described in the fourth aspect above.

[1666] In some embodiments, the at least one oxidative dehydrogenation reactor includes a single fixed bed type reactor.

[1667] In some embodiments, the at least one oxidative dehydrogenation reactor includes a single fluidized bed type reactor.

[1668] In some embodiments, the at least one oxidative dehydrogenation reactor includes a swing bed type reactor arrangement. f 1669] In some embodiments, the at least one oxidative dehydrogenation reactor includes more than one oxidative dehydrogenation reactor, each including the same or different oxidative dehydrogenation catalyst, connected in series, and where the product stream from each oxidative dehydrogenation reactor except the last oxidative dehydrogenation reactor in the series is fed into a downstream oxidative dehydrogenation reactor.

[1670] In some embodiments, the at least one oxidative dehydrogenation reactor includes more than one oxidative dehydrogenation reactor connected in parallel and each includes the same or different oxidative dehydrogenation catalyst.

[16711 In some embodiments, the chemical complex includes at least one heat exchanger immediately upstream of the quench tower.

[1672] In some embodiments, the chemical complex includes a caustic wash tower immediately downstream of the amine wash.

[1673] In some embodiments, the C2/C2 ₊ hydrocarbons leave the distillation tower and are directed to a splitter for separation of unreacted lower alkane and corresponding alkene into an unreacted lower alkane stream and a corresponding alkene stream.

[1674] In some embodiments, the distillation tower further provides for separation of the C ₂ /C ₂₊ hydrocarbons portion of the product stream into an unreacted lower alkane stream and a corresponding alkene stream.

[1675] In some embodiments, the unreacted lower alkane stream is directed back to the at least one oxidative deh drogenation reactor as part of the lower alkane containing gas.

[1676] In some embodiments, the oxygen separation module is tubular and the oxygen transport membrane includes an inner tube that is within an outer shell and where the reientate side includes the annular space between the inner tube and outer shell and the permeate side is the space within the inner tube.

[1677] In some embodiments, the oxygen separation module includes an additional inlet into the retentate side, the permeate side, or both, for introduction of combustible fuel into the oxygen separation module.

[1678] In some embodiments, the oxidation reactor contains a catalyst that includes a group 11 metal.

[1679] In some embodiments, the oxidation reactor contains a catalyst that includes a group 11 metal selected from copper, silver, gold and combinations thereof.

[1680] In some embodiments, the oxidation reactor contains a catalyst that includes silver or copper.

[16811 In some embodiments, the oxidation reactor contains a catalyst that contains a catalyst that includes

Ag/SiOi, AgCeG ₂ /Si0 ₂ , AgZr0 ₂ /Si0 _2, Cu/SiC¾, CuCe0 ₂ /SiG ₂ , CuZr0 ₂ /Si0 ₂ , CuCo ₃ 0 ₄ /Si0 ₂ and combinations thereof.

[1 82] In some embodiments, the stream entering the oxidation reactor includes from 100 ppm to 5 wt. % oxygen and optionally from 1 vppm to 1000 vppm acetylene.

[1683] In some embodiments, the stream exiting the oxidation reactor is essentially free of ox gen and acetylene.

[1 84] In some embodiments, the amount of carbon dioxide in the stream entering the oxidation reactor is greater than the amount of carbon dioxide in tire stream exiting the oxidation reactor. f 1685] In some embodiments, the amount of carbon monoxide in the stream entering the oxidation reactor is less than the amount of carbon monoxide in the stream exiting the oxidation reactor.

[1686] In some embodiments, the temperature in the oxidation reactor is from about 40 °C to about 100 °C.

[1687] Techniques are provided for a method of converting one or more alkanes to one or more alkenes in a carbon dioxide negative process that includes (a) providing a first stream containing one or more alkanes, steam, carbon dioxide and oxygen to an oxidative dehydrogenation reactor; (b) converting at least a portion of the one or more alkanes to one or more alkenes in the oxidative dehydrogenation reactor to provide a second stream exiting the oxidative dehydrogenation reactor containing one or more alkanes, one or more alkenes, carbon dioxide and one or more of oxygen and carbon monoxide; where the total amount of carbon dioxide in the second stream is less than the total amount of carbon dioxide in the first stream.

[1688] Techniques are provided for a method of converting one or more alkanes to one or more alkenes that includes (a) providing a first stream containing one or more alkanes and oxygen to an oxidative dehydrogenation reactor; and (b) converting at least a portion of the one or more alkanes to one or more alkenes in the oxidative deh drogenation reactor to provide a second stream exiting the oxidative dehydrogenation reactor containing one or more alkanes, one or more alkenes and one or more of oxygen, carbon monoxide and carbon dioxide; where the volumetric ratio of oxygen to one or more alkanes in the first stream is from 0.3 to 1 and where the product selectivity for carbon dioxide is less than 10 weight percent.

[1689] Techniques are provided for a method of converting one or more alkanes to one or more alkenes that includes (a) providing a first stream containing one or more alkanes, steam, carbon dioxide and oxygen to an oxidative dehydrogenation reactor; and (b) converting at least a portion of the one or more alkanes to one or more alkenes in the oxidative dehydrogenation reactor to provide a second stream exiting the oxidative dehydrogenation reactor containing one or more alkanes, one or more alkenes, carbon dioxide and one or more of oxygen and carbon monoxide; where the amount of carbon dioxide in the second stream is + 10 weight percent of the amount of carbon dioxide in the first stream.

[1690] Techniques are provided for a method of converting one or more alkanes to one or more alkenes that includes (a) providing a first stream containing one or more alkanes and oxygen to an oxidative dehydrogenation reactor; and (b) converting at least a portion of the one or more alkanes to one or more alkenes in the oxidative dehydrogenation reactor to provide a second stream exiting the oxidative dehy drogenation reactor contains one or more alkanes and one or more alkenes; where the amount of carbon dioxide in the second stream is equal to the amount of carbon dioxide in the first stream. Any of these aspects can include any combination of the following features.

[1691 ] In some embodiments, carbon dioxide is recycled to at least one oxidative dehydrogenation reactor.

[1692] In some embodiments, the olefins produced using the one or more ODH reactors, or any of die processes or complexes described herein can be used to make olefin derivatives.

[1693] In some embodiments, the olefin derivatives include, but are not limited to polyethylene polypropylene, ethylene oxide, propylene oxide, polyethylene oxide, polypropylene oxide, thermoplastic elastomers, thermoplastic olefins and blends and combinations thereof. f 1694] In some embodiments, the ethylene and optional a-olefms produced in the one or more ODH reactors, or any of the processes or complexes described herein, is used to make polyethylene.

[1695] In some embodiments, the polyethylene includes one or more of homopolymers of ethylene, copolymers of ethylene and cx-olefins, HOPE, MDPE, LDPE, LLDPE, VLDPE and combinations and blends thereof.

[1696] in some embodiments, the polyethylene is produced using one or more processes including gas phase polyethylene processes, high pressure polyethylene processes, low pressure polyethylene processes, solution polyethylene processes shiny polyethylene processes and suitable combinations of the above arranged either in parallel or in series.

[1697] The examples are intended to aid in understanding the present disclosure, however, in no way. should these examples be interpreted as limiting the scope thereof.

[1698] Examples

[1699] Example Qi

[1700] The effect of altering the amount of steam injected into an ODH process on the carbon dioxide output was demonstrated using two fixed bed reactors, connected in series. The catalyst present in each of the reactors was a mixture of several batches of a mixed metal oxide catalyst of the formula: Moi.oVo. ₃ o-o. ₅ oTeo.io-o. ₂ oNbo.io-o. ₂ oO _d , where the subscripts represent the range of atomic amounts of each element, relative to Mo, present in the individual batches, and d represents the highest oxidation state of the metal oxides present in the catalyst. Ethane, carbon dioxide, and oxygen were premixed before addition of water, followed by preheating with the entire composition being fed to the first of the two reactors. The preheating step was necessary to ensure the water added was converted to steam before injection into the reactor. Output from the first reactor was sent directly into the second reactor without addition of new reactants. For each reactor, the temperature was held in the range of 334-338 °C at ambient pressure. The process was ran continuously over a period of three days.

[ 701] The relative amounts of ethane, carbon dioxide, and oxygen remained the same while the flow rate of steam added to reactor was altered. The relative amounts of ethane, carbon dioxide, and oxygen added to the first reactor were 33, 54, and 13 respectively. The gas hourly space velocity' (GHSV) was kept constant at 610 h ^'! . Flow rates of reaction ethane, carbon dioxide and oxygen were altered accordingly to maintain GHSV at 610h ^! after altering the amount of steam added to reactor.

[1702] Steam was added indirectly as water with the ethane, carbon dioxide and oxygen mixture. The amount of water added to the mixture before entering the first reactor was varied, starting with no water and increasing in increments up to a flow rate of l.OcmVmin. For each flow rate of water added to the mixture, a corresponding weight % of steam in the total feed mixture was calculated. Table Ql shows the effect that changing the amount of steam added to the reactor had on output of carbon dioxide, carbon monoxide, and acetic acid. The output of the components was measured as normalized selectivity, according to the formula:

X selectivity (Wt. %) = net mass flow rate X (g X/hr) / ( where X refers to one of ethylene, CO2, CO, and acetic acid.

[1703] Results listed in Table Q1 were averaged from two or more experimental runs at each of the prescribed conditions. The results demonstrate that increasing the flow' rate of water added to the mixture and corresponding increase in the weight % of steam added to the reactor led to a decrease in the carbon selectivity. A carbon dioxide negative process was seen when the water was added at a flow rate of 1.0 cmVrnin, which corresponds to 39 weight % of steam added in addition, reverting back to no steam added followed by increasing to 39 weight % resulted in the carbon dioxide selectivity going positive back to negative. Finally, it should be noted that increasing the steam resulted in a higher production of acetic acid and also was accompanied by a higher conversion rate of e thane.

[1704] TABLE Ql: Normalized Selectivity of ODH Products in Response to Changes in Steam Added to the Reactor

[1705] Example Q2

[1706] For each of experiment numbers Ql-1 through Ql-7 in Table Ql, acetic acid is removed from the ODH product stream. The remaining stream is fed to a reactor containing AgZr0 ₂ /SiQ ₂ catalyst (particle size less than 5 nrti), prepared by impregnating silica with an aqueous silver nitrate solution, at from 105 to 115 °C. The CO in the stream is reacted with oxygen in the stream to form C0 ₂ . The stream exiting the reactor contains less CO and 0 ₂ and more C0 ₂ than the stream entering the reactor.

[1707] Example Q3

[1708] A second experiment was conducted using the same reactor configuration from Example Ql but under different operating conditions. The catalyst included a mix of several batches as described for Example Ql . and for comparison included a freshly mixed catalyst (fresh) and a mixed catalyst 8 months after being used intermitently used. The relative volumetric amounts of ethane, carbon dioxide, and oxygen added to the first reactor were 42, 37 and 21 respectively. Note the higher volumetric feed ratio of 0 ₂ /C ₂ H ₆ used compared to Example Qi. In addition, the gas hourly space velocity (GHSV) was higher, and kept constant at 1015 h ¹ , with reaction temperature being held from between 321 to 325 °C. Similar to Example Ql, flow rates of ethane, carbon dioxide and oxygen were altered accordingly to maintain GHSV at 1015 h ¹ alter altering the amount of water added. The corresponding steam content added to the first reactor was changed from 0 wt. % to 16 wt. %.

[1709J The results of this experiment, shown in Table Q2, demonstrated that when compared to the fresh catalyst (experiment Q3-1) the used catalyst (experiment Q3-2) displayed an increased selectivity towards the production of by-products, most notably C0 ₂ . with a concomitant decrease in ethylene selectivity. The fresh catalyst showed 91% selectivity to C ₂ ¾ and a negative C0 ₂ selectivity of -1.0. With the used catalyst selectivity to C ₂ ¾ dropped to 89% and CO selectivity moved into positive territory at 5.0. Experiment 3 with the used catalyst demonstrated that the disclosed methods are also effective with a used catal st, as increasing weight % of steam added to reactor from 0 to 16 weight % resulted in a drop in C0 ₂ selectivity to 3.0 from 5.0. This decrease was in good agreement with the observed trend in Example Ql.

[1710] TABLE Q2: Normalized Selectivity of ODH Products using Higher Feed Ratio of 0 ₂ /0 ₂ H ₆ and with Fresh Versus Used Catalyst

[1711] Example Q4

[1712] For each of experiment numbers Q3-I through Q3-3 in Table Q2, acetic acid is removed from the

ODH product stream. The remaining stream is fed to a reactor containing an AgZr0 ₂ /Si0 ₂ catalyst (particle size less than 5 mn), prepared by impregnating silica with arr aqueous silver nitrate solution, at from 105-115 °C. The CO in the stream is reacted with oxygen in the stream to form C0 ₂ .

[1713] The stream exiting the reactor contains less CO and 0 ₂ and more C0 ₂ than the stream entering tire reactor.

[1714] Example Q5

[1715] For each of experiment numbers Q3-1 through Q3-3 in Table Q2, acetic acid is removed from the ODH product stream. The remaining stream is fed to a reactor containing an AgCe0 ₂ /Si0 ₂ cataly st (particle size less than 5 mn), prepared by impregnating silica with an aqueous silver nitrate solution, at from 105-115 °C. The CO in the stream is reacted with oxygen in the stream to form C(¾. The stream exiting the reactor contains less CO and O2 and more €(¾ than the stream entering the reactor.

[1716] Examples Q6-Q10 (CO Selective Oxidation Process)

[1717] Experimental Reactor Unit (ERU) Setup

[1718] The ERU -was used to produce feed gas for evaluating the catalysts. The apparatus (400 in Figure 4) includes a fixed bed tube reactor 402. which is surrounded by two-zone electric heater 404. Reactor 402 in these examples was a 316L stainless steel tube which has an outside diameter of 0.5 inches (about 1.27 cm) and inside diameter of 0.4 inches (about 1 cm) and a length of 14.96 inches (about 38 cm). Two main feed gas lines rvere attached to reactor 402. One line 406 was dedicated for a built nitrogen purge gas. and the other line 408 was connected to a dual solenoid valve, which could be switched from ODH process feed gas (gas mixture of ethane/oxygen/nitrogen at a molar ratio of about 36/18/46) to compressed air when regenerating the catalyst bed 414.

[1719] For safety reasons the unit was programmed in a way that prevented air from mixing with the feed gas. This was accomplished through safety interlocks and a mandatory 15-minute nitrogen purge of the reactor when switching between feed gas 406 and air 412. The flow of gases was controlled by mass flow controllers. A 6-point thermocouple 416 was inserted through reactor 402. which was used to measure and control the temperature within catalyst bed 414. The catalyst was loaded in the middle zone of reactor 402 and located in between points 3 and 4 of thermocouple 416, which were the reaction temperature control points. The remaining 4 points of thermocouple 416 were used for monitoring purposes. Catalyst bed 414 included a one to one volume ratio of catal st to quartz sand, a total of 3 ml. The rest of reactor 402, below and above catalyst bed 414 was packed with 100% quartz sand and the load is secured with glass wool on the top and the bottom of reactor 402. A glass tight sealed condenser 418 was located downstream of the reactor 402 at room temperature to collect water/acidic acid, and the gas product flowed to either vent 420 or sampling loop/vent 422 by a three-way solenoid valve.

[1720] CO Selective Oxidation Catalyst Testing Reactor

[1721] A 3 i 6L stainless steel tube with the following dimensions was used to test CO selective oxidation catalysts: outside diameter: 0.25 inches (about 0.63 cm); wall thickness: 0.028 inches (about 0.07 cm); and catalyst bed height: 2 inches (about 5 cm). The total weight of the catalyst was recorded for each catalyst tested. The flow' of gases was controlled by the mass flow controllers on the ERU (also referred to as the MRU, herein). The product gas from the ERU was directly fed into the CO selective oxidation catalyst testing reactor (“Testing Reactor”). The Testing Reactor was placed in a precision heating oven, in which the temperature was controlled within a range of 0.5 °C. There were no thermocouples inside the reactor catalyst bed itself, and as a result, the oven temperature was recorded as the catalyst testing temperature. The catalyst bed consisted of approximately 1 g of catalyst supported between two layers of glass quartz wool. The effluent from the reactor w'as continuously provided for gas chromatography anal sis.

[1722] AgCe on Silica Catalyst Sample [1723] SYLOPOL ^® 2408 silica (W.R. Grace, surface area: 316 m ² / g, pore volume: 1.54 cc/g, 20 g) was impregnated with a solution (40 ml) of Ce(N0 ₃ ) .6H ₂ 0 (2.80 g) and X. The impregnated silica was dried at 90 °C overnight and w'as calcined in air at 500 °C for 6 hours.

[1724] X = AgNO , 103 mL of 0.1N solution. The solution was concentrated to about 20 ml and mixed with Ce(N0 ₃ ) ₃ .6H ₂ 0. Distilled w'ater was added to make 40 ml.

[1725] The catalyst made was CeAg oxide on silica with Ce0 ₂ : 5 wt. %, Ag: 5 wt. %.

[1726] CuCe on Silica Catalyst Sample

[1727] SYLOPOL 2408 silica (20 g) was impregnated with a solution (40 ml) of Ce(N0 ₃ ) ₃ .6H ₂ 0 (2.80 g) and

Y. The impregnated silica was dried at 90°C overnight and was calcined in air at 500 °C for 6 hours.

[1728] Y= Cu(CH ₃ COO) _2, 3.17 g. The solution was concentrated to about 20 ml and was mixed with Ce(N0 ₃ ) _3. 6H 0. Distilled water was added to make 40 ml.

[1729] The catalyst made was CeCu oxide on silica with Ce0 ₂ : 5 wt. %, Cu: 5 wt. %.

[1730] MnCe on Silica Catalyst Sample

[1731] SYLOPOL 2408 silica (20 g) was impregnated with a solution (40 ml) of Ce(N0 ₃ ) ₃ .6H ₂ 0 (2.80 g) and

Z The impregnated silica was dried at 90 °C overnight and was calcined in air at 500 °C for 6 hours.

[1732] Z— M:OC! ₂ .4H ₂ 0, 4.0 g. The solution was concentrated to about 20 ml and was mixed with Ce(N0 ₃ ) ₃ -6H ₂ 0. Distilled water was added to make 40 ml.

[1733] The catalyst made was CeMn oxide on silica with Ce0 ₂ : 5 wt. %, Mu: 5 wt. %.

[1734] CrCe on Silica Catalyst Sample

[1735] SYLOPOL 2408 silica (20 g) was impregnated with a solution (40 ml) of Ce(N0 ₃ ) _3, 6H ₂ 0 (2.80 g) and W. The impregnated silica was dried at 90 °C overnight and was calcined in air at 500 °C for 6 hours.

[1736] W= Cr(N0 ₃ ) ₃ .9H ₂ 0, 698 g. The solution was concentrated to about 20 ml and was mixed with Ce(X0 ₃ ) _3, 6H ₂ 0 Distilled water was added to make 40 ml.

[1737] The catalyst made was CeCr oxide on silica with Ce0 ₂ : 5 wt. % Cr: 5 wt. %.

[1738] Example Q6

[1739] AgCe on Silica Catalyst Testing

[1740] The ODH process was ran using the ERU and catalyst Mo VOx to provide the feed for tins example.

0.15 g of AgCe catalyst was used for tins test at a gas hourly space velocity of approximately 5000 bfo 0 psig on the reactor outlet at 75 °C process temperature. The results are shown in Table Q3.

[1741] TABLE Q3 f 1742] The data show that the AgCe catalyst demonstrates excellent oxygen removal properties via selective oxidation of CO to C0 ₂ , which can be seen from noticeable reduction of CO in the product gas and increase in all of the other compounds it is noteworthy that acetylene is also fully oxidized and was not detected in the product gas from the Testing Reactor.

[1743] Example Q7

[1744] 1.7 g of AgCe catalyst w'as used for this test, with approximately 3 g of the catalyst present in the hot zone of the reactor lg was the value for the catalyst weight used for the long term test calculations in this example, the ODH catalyst used to produce die feed for this example was a MoVOx catalyst (same as example Q6). In this example, the feed sample to die selective CO oxidation reactor was taken at the beginning and at the end of the test in order to confirm the composition of the feed. The test w ^; as executed at 110 °C process temperature, 0 psig reactor outlet pressure, gas hourly space velocity of approximately 3000 lr ^l . The results are summarized in Table Q4.

[1745] TABLE Q4

[1746] “0 ₂ removed” value was calculated as follows: where: V02 is the value of “C¾ removed” and C is the volumetric concentration of oxygen in the feed and product gasses

[1747] Because the composition of the feed to the selective CO oxidation reactor was changing gradually over Site term of the experiment, accurate values for removed oxygen could only be calculated at the very beginning and at the very' end of the run. The data show that even though the catal st had very stable activity toward acetylene oxidation through the whole duration of the run, the activity toward O2 removal via selective CO oxidation gradually- decreased over the duration of the run. Generally, the amount of CO and O2 in the product stream was less titan in the feed stream and the amount of CO2 in the product stream was greater than the amount in die feed stream.

[1748] Example Q8

[1749] The ODH process of example Q6 was used to provide the feed for this example. 0.35 g of AgCe catalyst, regenerated via oxidation, was used for this test. The test was executed at 110 °C process temperature, 0 psig reactor outlet pressure, gas hourly space velocity of approximately 3000 h ¹ . The results are shown in Table Q5.

[1750] TABLE Q5

[1751] The data show that the AgCe catalyst was successfully regenerated. Acetylene was reduced to undetectable levels and oxygen levels in the product stream were less than in the feed stream.

[1752] Example Q9

[1753] The ODH process of example Q6 was used to provide the feed for this example. 1.22 g of fresh CuCe catalyst was used for this test. The test was executed at 120 °C process temperature, 0 psig reactor outlet pressure, gas hourly space velocity of approximately 3000 h ¹ . The results are summarized in Table Q6.

[1754] TABLE Q6 f 1755] The data show that the CuCe catalyst exhibits similar properties to the AgCe catalyst. CuCe catalyzes the reaction of selective oxidation of CO to C0 ₂ and oxidation of acetylene. However, this catalyst sample did not show any catalyst activity at a temperature of 110 °C, which is noticeably higher than 75 °C, at winch fresh AgCe catalyst revealed significant activity' toward selective oxidation of CO.

[1756] Example Q10

[1757] The ODH process of example Q6 was used to provide the feed for this example. 1.01 g of MnCe catalyst was used for this test. The test was executed at 140 °C process temperature, 0 psig reactor outlet temperature, gas hourly space velocity of approximately 3000 Ir ¹ . The results are summarized in Table Q7.

[1758] TABLE Q7

[1759] The data show that the MnCe cataly st exhibits selective CO oxidation properties, however, it did not demonstrate any activity toward oxidation of acetylene. This catal st sample did not show any catalytic activity at the temperature below 140 °C, which is noticeably higher than 75 °C, at which AgCe catalyst revealed significant activity toward selective oxidatio of CO.

[1760] Example Qil

[17611 This experiment was conducted using the same reactor configuration as the previous example Q3, but only using the second reactor in the series and under variable feed volume ratios of oxygen to ethane. The catalyst used was a mixed metal oxide catalyst of the formula: MoioVo riTeo^Nbo iO s _? and was extruded with ~55 wt. % of VERS AL™ 250 support (UOP LLC) in balance mixed metal oxide. Three relative volumetric amounts of oxygen and ethane were tested including 16 vol. % 0 ₂ : 38 vol. % C ₂ H ₆ , 19 vol. % 0 ₂ : 36 voi. % C¾, and 21 vol. % 0 ₂ : 33 vol. % C ₂ H ₆ , which correspond to 0 ₂ :C ₂ H ₆ volumetric ratios of 0.4, 0.5, and 0.6, respectively. The relative volumetric amount of C0 ₂ added was maintained at 46 vol. %, the gas hourly space velocity' (GHSV) was kept constant at 1111 h ¹ , the reaction temperature was held between 359 °C and 360 °C, and reactions were performed at ambient pressure. No steam was added to the reaction

[1762] The results of fids example are shown in Table Q8. As the volumetric ratio of oxygen: ethane is increased the selectivity towards the production of C0 ₂ decreases. This effect is accompanied by slight increases to selectivity towards ethylene and carbon monoxide, while acetic acid selectivity remains unchanged. Experiment Qll-3 demonstrates that altering volumetric ratio of oxygemethane added to the reactor, while keeping other parameters unchanged, can decrease the selectivity to carbon dioxide. This effect is also demonstrated by comparing Examples Q1 and Q3, specifically experiment numbers Ql-1 and Q3-1 where no steam was added, in that tire carbon selectivity was lower in experiment number Q3-1 where a liigher volumetric ratio of oxygemethane was added to the reactor. [1763] TABLE Q8: Normalized Product Selectivity of QDH Products in Response to Variations of Volumetric Feed Ratio of Oi/CVHe at Elevated Temperature and Without the Addition of Steam Temp - 359-360 °C; GHSV - 1110 h- ^! ; Steam added - 0 voi. %

[1764] Example Q12

[1765] This experiment was conducted using the same reactor configuration as the previous examples and similar to example Q1 but using a higher volumetric ratio of oxygemethane (0 5) added to the reactor a higher GHSV (t i l l h ¹ ), and a higher temperature of 360 °C The weight % of steam added to the reactor was changed from 0 wt. % to 40 wt. % while keeping the relative volumetric amount of C(¾ steam added (46 vol. %) constant. The results are presented in Table Q9 and demonstrate that at higher temperatures, flow rates and volumetric ratio of oxygemethane increasing the amount of steam added to the reactor from 0 wt. % to 40 wt. % decreases the CO2 selectivity. In this example, the C(¾ selectivity decreased from 6.0 wt. % to 5.3%. This decrease is lower than what is seen when operating at a lower temperature, low flow rate (GHSV), and lower relative volumetric ratio of oxygemethane added to the reactor.

[1766] TABLE Q9: Normalized Product Selectivity of ODH Products in Response to Changes in Steam Added to the Reactor at Higher Temp., GHSV, and voi. ratio 0 ₂ :ethane. Temp - 360 °C; GHSV - 1111 h ¹ ; Voi ratio 0.5

[1767] Embodiments relate to an ODH process that includes multiple reactors in series for the ODH of lower alkanes (C He - CgH g) into corresponding alkenes. The process includes to remove residual oxygen from the last of a series of oxidative dehydrogenation reactors by adding a mixture of alcohol and steam is herein provided. Techniques are provided for a process to remove O2 from die ODH reactor or the last ODH reactor assuming that there are multiple reactors in series, by means of adding mixture of a C ₁ -C ₃ alcohol, preferably ethanol, and steam, into this reactor to consume the residual (¾ by reacting it with alcohol, preferably ethanol, to generate the corresponding carboxylic acid, preferably acetic acid. The ODH reactor(s) can be fixed bed, fluidized bed, moving bed, ebulliated bed, shell and tube or tube reactor design. The alcohol concentration may be from 0.5 to 2 vol. % of the alcohol and steam mixture. The reaction operating temperature for this bed is in the range of 150 °C up to desired ODH reaction temperature. Once the O2 is consumed in this reactor, the remaining alcohol (e.g., ethanol) is cataiyticaliy dehydrated to generate an alkene (e.g., ethylene) in the same reactor.

[1768| Some embodiments include a process for removing oxygen from the last ODH reactor assuming multiple reactors in series.

[1769] Techniques are provided for a process for consuming oxygen from an oxidative dehydrogenation reactor for the oxidative dehydrogenation of a lower alkane into a corresponding alkene, the process comprising: flowing a C1-C3 alcohol at a concentration of 0.01-15 vol. % in an inert gas into the last 50% of the oxidative dehydrogenation reactor; passing this oxygen-consuming stream at a temperature of between 140 °C and 370 °C through the oxidative dehydrogenation reactor; such that the oxygen concentration in the effluent from the oxidative deh drogenation reactor is less than 10 ppmv. As used herein, the term inert is determined with respect to flammability. Some inert gases or diluents, such as steam or carbon dioxide, may be used as inert diluents to change the flammability limits of mixtures of oxygen and hydrocarbons, but will participate in the catalytic ODH reactions described herein.

[1770] In some embodiments, is a process wherein the C1-C3 alcohol at a concentration of 0.01-15 vol. % in an inert gas is fed into the last 30% of the oxidative dehydrogenation reactor.

[1771] Techniques are provided for a process wherein the C1-C3 alcohol at a concentration of 0.01 -15 vol. % in an inert gas is fed into the last 10% of the oxidative dehydrogenation reactor.

[1772] Techniques are provided for a process for consuming oxygen from at least two oxidative dehydrogenation reactors in series for the oxidative dehydrogenation of a lower alkane into a corresponding alkene, the process comprising: flowing a CYC ₃ alcohol at a concentration of 0.01-15 vol. % in an inert gas into the effluent of the one or more oxidative dehydrogenation reactors upstream of the final reactor to make an oxygen-consuming stream; passing the oxygen-consuming stream at a temperature of between 140 °C and 370 °C through a final oxidative dehydrogenation reactor; such that the oxygen concentration in the effluent from the final oxidative dehydrogenation reactor is less than 10 ppmv.

[1773] An embodiment includes a process wherein the lower alkane is a C1-C3 alkane.

[1774] An embodiment includes a process wherein the C1-C3 alkane is ethane.

[1775] An embodiment includes a process wherein the alcohol reacts with oxygen to form a carboxylic acid.

[1776] An embodiment includes a process w'herein the C1-C3 alcohol is ethanol.

[1777] An embodiment includes a process w'herein the ethanol reacts with ox gen to form acetic acid.

[1778] An embodiment includes a process w'herein any alcohol remaining after reaction with ox gen reacts to form a lower alkene.

An embodiment includes a process wherein the lower alkene is ethylene. f 1780] An embodiment includes a process wherein at least one of the oxidative dehydrogenation reactors comprises a fixed bed type reactor.

[17 1] An embodiment includes a process wherein at least one of the oxidative dehydrogenation reactors comprises a single fluidized bed type reactor.

[1782] An embodiment includes a process wherein at least one of the oxidative dehydrogenation reactors comprises a moving bed type reactor.

[1783] An embodiment includes a process wherein at least one of the oxidative dehydrogenation reactors comprises an ebulliated bed type reactor.

[1784] An embodiment includes a process wherein the oxidative dehydrogenation reactor comprises a shell and tube reactor design, wherein the cooling media can be introduced on the shell or tube side of the reactor, and is not limited to molten salt, mineral or synthetic oils, pressurized water, etc.

[1785] An embodiment includes a process wherein the oxidative dehydrogenation reactor comprises a tube reactor design.

[1786] An embodiment includes a process wherein the C1-C3 alcohol is at a concentration of 0.05-15 vol. % in an inert gas.

[1787] An embodiment includes a process wherein the oxygen-consuming stream is at a temperature above the dew point of the effluent stream.

[1788] An embodiment includes a process wherein the oxygen-consuming stream is at a temperature of between 140 °C and 200 °C.

[1789] Embodiments of the present techniques include a process (system and method) to remove 0 _? _ from the ODH reactor, or the last ODH reactor assuming that there are multiple reactors in series, by means of adding mixture of a C1-C3 alcohol, preferably ethanol, and steam, into this reactor to consume the residual O2 by reacting it with alcohol, preferably ethanol, to generate the corresponding carboxylic acid, preferably acetic acid. The ODH reactor(s) can be fixed bed, fluidized bed, moving bed, ebulliated bed, shell and tube or tube reactor design. The alcohol concentration may be from 0.5 to 2 vol. % of the alcohol and steam mixture. The reaction operating temperature for this bed is in the range of 150 °C up to desired ODH reaction temperature. Once the O2 is fully consumed in this reactor, the remaining alcohol, preferably ethanol, is catalytically dehydrated to generate an alkene, preferably ethylene, in the same reactor.

[:! 790] It is speculative that the portion of the exothermic heat of reaction for converting ethanol to acetic acid provides heat of reaction for the endothermic reaction of ethanol dehydration to ethylene. As a result, addition of small concentration of ethanol to the last reactor generates relatively small net heal of reaction, which can enable the one tube or adiabatic fixed bed reactor design, as opposed to tube and shell heat exchanger reactor design, which can lead to capital cost savings.

[1791 ] Presence of ethanol and steam in the last reactor bed was found to preserve the catalyst from deactivation in the Ch-free environment. The benefits of the explained O2 removal process are summarized as:

[1792] O2 removal from ODH product sheam to avoid fouling in the separation and compression train downstream of the ODH reactors while preserving the catalyst activity in the last reactor bed. f 1793] 0- ₂ removal from ODH product stream to avoid degradation of the amine system for removing CO2 and

H S into heat-stable amine salts.

[1794] Enables the tube reactor design, as opposed to tube and shell heat exchanger reactor design, for the last ODH reactor, which can lead to capital cost savings.

[1795] Minor increase in ethylene yield in the ODH process by converting portion of the ethanol into ethylene.

[1796] Ethanol can come from multiple sources: acetic acid hydrogenation from ODH itself, bio-sources, ethylene h dration, etc. Inclusion of even small amounts of bioethanol to scavenge trace oxygen produces two useful commercial co-products, acetic acid and ethylene.

[1797] The following examples are merely illustrative and are not intended to be limiting. Unless otherwise indicated, all percentages are by weight.

[1798] Examples

[1799] A Fixed Bed Reactor Unit (FBRU) was used to conduct experiments on residual O2 removal. The apparatus is depicted schematically in Figure 37. The apparatus consisted of two fixed bed tubular reactors in series. Each reactor was wrapped in an electrical heating jacket and sealed with ceramic insulating material. Each reactor was SS316L tube which had an outer diameter of 1” and is 34” in length. In these experiments, ethane, ethylene, carbon dioxide, oxygen, nitrogen were fed separately (on as-needed basis) and pre-mixed prior to the reactor inlet, 3718, with the indicated composition (given in each experiment). The flow passed from the upstream reactor to the downstream reactor at stream 3719, and the product stream exited the downstream reactor at stream 3720. Both reactors were being controlled at the same reaction temperature. The temperature of each of the reactors were monitored using co rrespo nding 7 -point thermocouples shown by # l-#7 in the upstream reactor, and #8-#14 in the downstream reactor. The highest temperature between thermocouple points was used for controlling the reactor temperature using the corresponding back pressure regulator that controlled the pressure and boiling temperature of water inside the desired reactor water jacket, 3714. It is noteworthy that only thermocouple points #3 to #6 in the upstream reactor and #9 to #12 in the downstream reactor were located in the reactor bed, and the reaction temperature for each reactor was being reported as an average of these points.

[ 1800] The catalyst bed, 3715, consisted of one weight unit of catalyst to 2.14 units of weight of Denstone 99

(mainly alpha alumina) powder; total weight of the catalyst in each reactor was 143 g catal st having the formula MoV0 40Hb0 _.i 6 e0 _.i 4O, with relative atomic amounts of each component, rela tive to a relative amount of Mo of 1, shown in subscript. The rest of the reactor, below and above the catalyst bed was packed with quartz powder, 3716, and seemed in place with glass wool, 3717, on the top and the bottom of the reactor tube to avoid any bed movement during the experimental runs. [ 1801] TABLE R1 : ODH Residual O ₂ Removal Example

[1802] TABLE R2: Catalyst activity converting ethanol to ethylene and acetic acid during ODH Residual O ₂ Removal Example

[18031 TABLE R3 : Catalyst activity before and after ODH Residual O ₂ Removal Example

[1804] A method for preventing or removing water soluble fouling located downstream of an oxidative dehydrogenation (ODH) reactor is described. The method employs the introduction of water upstream of fouling locations, either continuously or intermittently, which acts to solubilize and carry away fouling material. The me thod has the advantage of being applicable lor use while an ODH process is ongoing, circumventing the need for a costly shutdowm.

[1805] Provided herein is a method for removing or preventing buildup of water soluble fouling that accumulates dowmstream of an ODH reactor used for oxidative dehydrogenation of lower paraffins into corresponding olefins. In some embodiments a solvent is introduced into the outlet pipe from the ODH reactor upstream of where fouling is likely to develop, primarily in a liquid state and in a manner that promotes annular flow, laminar or turbulent, of solvent along the inner surface of the pipe immediately downstream of the location of introduction. In some embodiments, the solvent introduced into the outlet pipe is water.

[1806] In some embodiments, solvent is introduced via a pipe-in-pipe arrangement, with the outlet pipe from the reactor fitting within a downstream pipe that has a larger diameter. In this instance solvent is introduced by way of the gap between the outlet pipe and the downstream pipe.

[1807| In some embodiments, solvent can be sprayed onto the inner surface of the outlet pipe by an instream atomizer or alternatively by a series of jets strategically placed on and continuous with the inner surface of the outlet pipe.

[1808] In some embodiments, solvent is introduced via a plurality of holes with exits continuous with the inner surface of the outlet pipe.

[1809] In some embodiments, water is introduced via a pipe-in-pipe arrangement in combination with instream atomizers, inner surface jets, a plurality of holes or combinations thereof.

[1810] Disclosed herein is a method for removing of and or preventing buildup of substantially water soluble fouling in the outlet pipe of an oxidative dehydrogenation (ODH) reactor used for converting lower paraffins (C2- G _f ) into corresponding olefins. The term “fouling” including in the worst case “pluggage”, refers to water soluble by-products of an ODH process that solidify after leaving the reactor and have the potential to adhere to and build up on the interior surface of the outlet pipe leaving the ODH reactor. Left unchecked the fouling can limit or even block flow' through the exit pipes, impacting the efficiency of the ODH process. Pluggage refers to complete occlusion of the pipe. Water soluble by-products of the ODH process include oxygenates, such as maleic add. The modifier “substantially water-soluble” refers to the possibility that fouling may also include minor amounts of miscible products such as acetaldehyde and ethanol or other non-soluble particulate entrained within the solid fouling. Lower paraffins refer to paraffins with 2 to 4 carbons. In one embodiment the low'er paraffin is ethane and its corresponding olefin is ethylene. For simplicity' the method is described for use with ethane ODH, but may be applied to use of propane and butane ODH

[1811] An advantage of the disclosed method is that fouling can removed or prevented from forming while the ODH reactor is operational. Operational means during operation, as opposed to periods of shutdown where the reactor is idle and flow of reactant and product gases has stopped. Shutdowns are potentially time consuming and costly. The method disclosed reduces the risk of a need for shutdown, which has significant economic implications. However, the method may also be employed during reactor shutdown in this instance, the volumes and flow rales of solvent introduced would be need to be varied accordingly.

[1812] The method described is intended for use in relation to an ODH process wherein a gas mixture 7001 comprising at least ethane and oxygen is introduced, via one or more inlets 7002, into an ODH reactor 7003 that contains an ODH catalyst (Figure 70). Conditions within the ODH reactor 7003 promote conversion of ethane into ethylene. The oxidative reaction catalyzed by the ODH catalyst may produce a variety of by-products, for example carbon dioxide, acetic acid, and water soluble oxygenates such as nraleie acid. Exit stream 7004 comprising ethylene and by-products and any unreacted ethane and oxygen, if present, exit ODH reactor 7003 via outlet pipe 7005. Temperature of the exit stream 7004 may vary and is controlled by an operator depending upon desired reaction conditions which are tailored for a specific target product profile. Upon exiting the ODH reactor 7003, the exit stream 7004 in the absence of exothermic reaction conditions begins cooling naturally and optionally with the aid of cooling mechanisms, such as heat exchanger 7006, before being subjected to a series of separation steps, starting with for example a quench tower for removal of water and acetic acid.

[1813] Without wishing to be bound by theory, as the temperature of exit stream 7004 drops, water soluble oxygenates, present in the gaseous state within ODH reactor 7003, condense within the gaseous exit stream 7004 as liquid droplets then adhering to timer surface of outlet pipe 7005, eventually solidifying as fouling 7007 in one more locations downstream of ODH reactor 7003 and upstream of where separation steps begin. While the fraction of exit stream 7004 that has water soluble oxygenates is minimal, without intervention fouling 7007 may grow over time as more particles of solid water soluble ox genates adhere to the growing mass. Growth of fouling 7007 may be detected as an atypical pressure drop between die reactor and a location downstream of outlet pipe 7005. Introduction of a solvent at one or more locations (arro ws 7008) upstream of the location of the fouling promotes dissolving of the fouling 7007, which can then be carried by the flow of gases and liquids to the separation steps, where it can be removed along with acetic acid

[1814] The ODH Reactor

[1815] The disclosed method contemplates the use of any of the kno wn reactor types applicable for the ODH of hydrocarbons. In some embodiments, the method disclosed herein employs one or more conventional fixed bed reactors. In a typical fixed bed reactor reactants are introduced into the reactor at one end, flow past or over an immobilized catalyst, products are formed and leave at the other end of the reactor. The reactor may include separate inlets for each of the reactants, or may include a single inlet (similar to Fi gure 70) where reactants are premixed and introduced into the reactor as a homogeneous mixture. A person skilled in the ast would know which features are required with respect to shape and dimensions inputs for reactants, outputs for products, temperature and pressure control and monitoring, and means for immobilizing the catalyst.

[1816] In some embodiments, use of one or more fluidized bed reactors is contemplated. These types of reactors are well known. Typically, the catalyst is supported by a porous structure, or distributor plate, located near a bottom end of the reactor and to which reactants are forced through at a velocity sufficient to balance the weight of the catalyst such that it rises and begins to swirl around in a fluidized maimer. The reactants are converted to products upon contact with the fluidized catalyst and subsequently removed from the upper end of the reactor. Design considerations include shape of the reactor and distributor plate, input and output, and temperature and pressure control and monitoring, all of which would fail under knowledge of the person skilled in the art.

[1817] Multiple ODH reactors, either in series or in parallel, can be used. Use of multiple reactors, including ODH reactors, in either a parallel or series arrangement is well known in the art. Where parallel ODH reactors are emplo ed, the method disclosed herein may be used downstream of each of the ODH reactors, after streams from each of the ODH reactors are combined, or both downstream of each of the ODH reactors and after streams of each of the ODH reactors are combined. f 1818] Where ODH reactors in series are employed fouling would not be expected between ODH reactors as the temperature is not likely to drop far enough such that condensation and freezing of oxygenates can occur. In some embodiments, it may be preferable that introduction of solvent for prevention or removal of fouling is only used downstream of the last ODH reactor in the series. However, if fouling is seen between ODH reactors, the method may be used between ODH reactors in series, but a user should ensure that when using water as the solvent, that water introduced between ODH reactors is either removed prior to the next ODH reactor in the series, or that the temperature of the stream is high enough to convert the water into steam prior to entering the next ODH reactor. Liquid water within an ODH reactor is potentially damaging to the catalyst and the reactor and associated components.

[1819] The ODH Process

[1820] An ODH reactor can be used in accordance to the provided techniques. The oxidative dehydrogenation of ethane may be conducted at temperatures from 250 °C to 550 °C, or front 300 °C to 500 °C, or from 350 °C to 450 °C, at pressures from 0.5 to 100 psig (3.447 to 689.47 kPag), or from 15 to 50 psig (103.4 to 344.73 kPag), and the residence time of the lower alkane in the reactor is, for example, from 0.002 to 30 seconds, or from 1 to 10 seconds. [1821] Any of the ODH catalysts kno wn in the art are suitable for use with the method disclosed herein. When choosing a catalyst a skilled user would appreciate that catalysts can vaiy with respective to selectivity and activity. In one embodiment, mixed metal oxides are employed as they can provide high selectivity to ethylene without significant loss in activity. Example catalysts are those of the formula:

VoMo Nb _r T er _f Me _e O/ wherein: Me is a metal chosen from Ta, Ti, W, Hf, Zr, Sb and mixtures thereof; a is from 0.1 to 3; b is from 0.5 to 1.5; c is from 0 to 3; d is from 0 to 5; e is from 0 to 2; and f is a number to satisfy the valence state of the catalyst. Any other catalysts described herein may be used in addition to or instead of, this catalyst formulation.

[ 822] A wide variety' of combinations for ratios of ethane, oxygen, and optionally inert diluent, can be used in the ODH process. A person skilled in the art would understand that for safety- reasons, it may be preferable to choose those compositions where the ratio of oxygen to ethane, in the presence or absence of inert or substantially inert components, falls outside of the flammability' envelope. This includes ratios either above the upper flammability limit or below the lower flammability limit. A person skilled in die art would know how to determine flammability limits, and whether a particular composition, including those that include an inert diluent, would fall outside of the flammability envelope.

[1 23] Oxygen may be supplied as pure oxy gen, or as a component of a gas mixture such as atmospheric air. Atmospheric air contains nitrogen, which acts as an inert diluent if using pure oxygen and an inert diluent, the inert diluent should exist in the gaseous state in the conditions within the reactor and should not increase the flammability of the hydrocarbon added to foe reactor, characteristics thill a skilled worker would understand when deciding on which inert diluent to employ. Inert diluent can be added to either of the ethane or oxygen, or oxygen containing gas if using prior to entering the ODH reactor or may be added directly into the ODH reactor.

[1824] The reaction conditions and reactant compositions, including ratio of ethane to oxygen, and the presence, absence or nature of inert diluent, along with the catalyst employed can impact the product profile, including the selectivity to ethylene, the conversion rate, and the degree that oxygenates, such as maleic acid, are produced in the ODH reaction. A person skilled in the art would be familiar with how adjusting conditions and components impact the product profile. Detecting the presence of fouling, requiring the need for removal, is achieved by monitoring the pressure drop as exit stream 7004 leaves the reactor. Under normal operating conditions the pressure will drop along with the temperature. When fouling starts to develop it may begin to occlude the pipes downstream of the ODH reactor, resulting in a change in the pressure profile. Pressure upstream of the fouling will be higher than normal while pressure downstream of the fouling will be lower than what occurs under normal circumstances. A complete blockage would likely result in a very large pressure increase upstream of the fouling. A larger than normal pressure drop w'ould be expected when there is higher than normal pressure upstream of the fouling and lower than normal pressure downstream of the fouling.

[1825] Introduction of solvent

[1826] Solvents that can be used in accordance with the described method are those that can dissolve the substantially water soluble fouling, exist in liquid form at the temperature and pressure found in outlet pipe 7005 downstream of the ODH reactor, and do not negatively impact downstream separation steps. In some embodiments, the solvent is water. Water used with the method need not be distilled or deionized, and can include some impurities, provided the impurities are non-reactive with respect to olefins and are not likely to negatively impact downstream processing. Collection water from the bottom of an acetic acid scrubber, which is likely the first separation unit downstream of the ODH reactor, may also be used as the solvent. Collection water from the bottom of an acetic acid scrubber may be described as dilute acetic acid, with the concentration of acetic acid ranging from 0 - 50 wt. %. for example from 1 to 10 wt. %. Usi ng dilute acetic acid has the advantage of recycling water used in the acetic acid scrubber so that an additional water source may not be required. Some embodiments of the disclosed method will be described using water as the solvent. In each instance the solvent need not be water but any other substance that can dissolve the substantially water soluble fouling, exists in primarily liquid state in the prevailing conditions, and does not negatively impact downstream processing.

[1827] Introduction of water downstream of the ODH reactor and upstream of fouling can be done in multiple locations prior to the separation steps. The water introduced is primarily in the liquid state at the temperature and pressure w'here it is introduced. In liquid form water has the capability of dissolving the substantially water-soluble fouling, while as a steam this is not possible. Introduction of water is performed so that an annular flow of water, denoted by dashed curved arrows in Figures 71-74, substantially on the inner surface of the pipe contacts and begins to dissolve fouling 7007, carrying fouling particles 7009 away from the side of the pipes (Figure 71).

[1828] In some embodiments, as shown in Figure 71, a pipe in pipe arrangement 7010 can be used to deliver the w'ater in the desired manner. In this arrangement pipe 7011 coming from and continuous with outlet pipe 7005 or an outlet pipe downstream from a cooling mechanism (e.g. heal exchanger 7006)(Figures 70 and 71) becomes an inner pipe 7012. ending and fitting inside an outer pipe 7013 with a larger diameter. The end of inner pipe 7012 may or may not have a flare 7014. so long as there is a gap 7015, continuous or discontinuous, between the internal surface of outer pipe 7013 and the terminal end of inner pipe 7012, where w'ater can be introduced. The gap is chosen such that the water preferentially films on die internal surface of outer pipe 7013, producing what is often referred to as annular flow. Inner pipe 7012 may also include rifling to produce swirling of the water as it is introduced, promoting complete coverage of the internal surface of outer pipe 7013. The end of inner pipe 7012 should be located in positions upstream of where fouling may develop. Multiple pipe in pipe arrangements 7010 can be used in multiple locations. The outer pipe 7013 of an upstream pipe in pipe arrangement 7010 acts like the inner pipe 7012 for the subsequent pipe in pipe arrangement 7010. A user may choose to taper the diameter of outer pipe 7012 as it approaches the next arrangement, or may choose an inner pipe 7012 with a consistent diameter. In this case, the diameter of each subsequent outer pipe 7013 increases.

[1829| In some embodiments, water may be introduced via mi instream atomizer head 7016 (Figure 72), the atomizer spraying the interior walls from a position within the stream of reactants leaving the ODH reactor. In another embodiment, water may be introduced using of multiple jets 7017 within the walls of the exit pipe 7011 (Figure 73A and 73B). Jets 7017 may spray an area of the inner wall (indicated by straight dashed arrows) of the pipe opposite the jet, the size of the area depending on the configuration of the nozzle on the end of the jet 7017. The use of multiple jets can be used to ensure coverage of the imier wall of the exit pipe 7011, the number required depending on the nozzle and the spray area of the jets (see cross section of pipe with jets in Figure 73B). The jets may be in a concentric circle, o r can be staggered with some jets further upstream of subsequent jets.

[1830] In some embodiments, water (indicated by curved dashed arrows) can be introduced via a plurality of holes or perforations 7018 on the inne r surface of the pipe where the water is introduced (Figure 74 A and 74B). The holes would be designed to allow' water to leak or seep into the pipe, including an option for closing the holes and for controlling the flow of water through the holes to promote annular flow once inside the pipe. Designs of this type fall within the kno wledge of the perso skilled in the art. Perforations or holes 7018 can be spaced in a uniform pattern or may dispersed at various locations lengthwise along exit pipe 7011.

[1831] The embodiments for introduction of water are not intended to solely be used in isolation. In some embodiments, introduction of water can be done at one or more locations, with each location employing one or more ways for introducing water into the pipe. For example, a pipe in pipe arrangement may be used in conjunction with an instream atomizer, or perforations, or multiple jets. In using multiple ways for introduction of water into the pipe a user must be mindful that increasing the amount of introduced water may dilute the stream more than when using a single -way of introducing water.

[1832] The introduction of water into pipes downstream of the reactor wherein there is a gaseous flow creates a multiphase flow' arrangement, made up of the gaseous exit stream 7004 and introduced liquid water. A person skilled in the art would be familiar with a variety of mechanisms for introducing water at appropriate flow rates such that an annular flow of water develops substantially at the inner surface of the pipe (i.e. on or near the surface of the pipe as depicted in figures 71, 72, 73a and 74a). Factors to consider when determining flow' ra tes of water to be introduced include the flow rate, temperature, and pressure of exit stream 7004 coining off Site ODH reactor 7003. Flow rate of the gaseous exit stream 7004 will be dependent on the size of the ODH reactor and can range front 2 L/inin for bench scale up to 80,000 L/min for commercial scale reactors. The temperature and pressure will depend upon reactor configuration and reaction conditions discussed previously. Both temperature and pressure, in the absence of fouling, beginning falling immediately' after exiting the reactor. ί 1833] Choosing flow rates should also take into account the degree to which the water begins to dilute the exit stream. A user needs to determine the appropriate flow rate that permits formation of an annular stream without adding too much water that will need to be removed downstream. In some embodiments it may be preferable for removal of water to be done during a quench step that also separates out acetic acid produced. Introduction of copious amounts of water into the exit stream will add to that used in the quench step, diluting the acetic acid further. A user that further concentrates the acetic acid after separation may wish to limit the amount of water added for preventing fouling, so as to reduce the degree to which the acetic acid needs to be concentrated.

[1834] Water is introduced into the pipe under conditions of gaseous flow through the pipe that will result in annular flow of liquid. Formation of an annular flow of liquid for a given gaseous flow rate is a function of gaseous and liquid densities, viscosities, surface tension, and flow rate. For more information on factors related to multiphase flow see Chapter 2, “Gas-Liquid Transport in Ducts”, of the Multiphase Flow Handbook by Clayton T. Crowe,

2006.

[1835] In limited circumstances a brief slug flow could be used. Brief means a short enough period of time to not disrupt reactor operations, such that the reactors needs to be shut down. In this instance, the volume of water introduced into the pipe is high enough to form “slugs” of liquid uatei. preceded and followed by pockets of gaseous flow'. As the slugs move downstream fouling is dislodged from the interior surface of the pipe by the passing slug of water, and dissolves as the slug moves along. If introduction of water at higher volumes associated with slug flow' is maintained then pressure within the reactor may increase to a point where reaction conditions are not ideal. This should be avoided. In addition, downstream separation units must be capable of withstanding any sudden pressure increase resulting from interaction with voluminous slugs of water.

[1836] Choosing a location for introduction of water requires consideration of where fouling develops or is likely to develop. These locations are generally expect to occur where the temperature aud pressure are conducive to water soluble oxygenates forming solid particulate that can adhere to the inner surface of the piping connecting the reactor with downstream separation components. For instance, the melting point of maleic acid is 135°C at ambient pressure. Using a temperature profile of the exit stream as it leaves the reactor would provide a .guide where water can be introduced, Ideally at a location where the temperature exceeds the melting point of maleic acid. However, a user would understand that locations should be chosen where the introduced water is unlikely to boil to a significant degree or the water dissolves the fouling faster than it boils in some embodiments, the introduced water remains substantially in the liquid state until it dissolves the fouling. The temperature of the water after introduction will likely increase, depending on location of introduction and temperature of water at location of introduction, peak and then decline as the stream gets further downstream. The location of introduction may therefore include locations above the boiling point of water, provided the flow' rates allow' that the peak temperature of introduced water to remain below' boiling allowing it to remain in the liquid state for the time required to dissolve the foulant.

[1837] Intermittent or consistent operation

[1838] The method disclosed herein may be emplo ed continuously or intermittently. That is, water may be introduced at one or more locations continuously during operation of the ODH reactor. In this instance a user may opt for the lowest possible flow rate of w'ater to minimize the degree to which the product stream is diluted. Continuous operation may be effective for preventing fouling from forming in the first place as the interior surface of the piping downstream of the reactor and prior to separation components is constantly covered with a film of flowing water. The fouling particles are not subjected to conditions where adherence can be maintained for a period long enough that subsequent particles can adhere to the previous particles which would allow the mass to grow. [1839] Alternatively, a user may opt to employ a method when it is deemed prudent The operating conditions may be such that water soluble byproduc ts are produced at such a low level that accumula tion of fouling takes a long time to develop. When a user determines that the pressure drop from the reactor is outside a normal range expected in the absence of fouling foe method may be used and water can be introduced into or more locations of the user’s choosing. When the pressure drop returns to normal introduction of water into the stream can be ceased. Pressure transducers placed in intervals downstream of the reactor can pinpoint where fouling is accumulating by highlighting those sections where there is an abnormal pressure drop. A user may then opt to introduce water at a location just upstream of die fouling, or multiple locations. If the pressure drop does not correct over time a user may then increase die amount of water added, or the number of locations upstream where water is added. Alternatively, a user may opt for introduction of water such that slug flow is used. Finally, if given sufficient time and the pressure drop does not improve it may indicate that a problem involving fouling that is not water soluble is occurring.

[1840] EXAMPLE

[1841] Two fixed bed reactors, connected in series, were used in an ODH process where ethane, ethylene, carbon dioxide, and oxygen, in ratios of 11-93/0-80/0-8/0-8 vol. %, respectively, were premixed before being fed to the first of the two reactors. The weight hourly space velocity was within the range of 0.65 to 2.70 h ¹ . Output from the first reactor was sent directly into the second reactor without addition of new reactants. For each reactor, the temperature was held in the range of 300-337°C at ambient pressure. The process w¾s nin continuously over a period of foil) _' three days, with feed composition varying within the stated ranges. An abnormal pressure profile was detected immediately downstream of the second reactor and prior to a downstream condenser. The process was stopped dire to triggering of the high pressure alarm for the reactor and the outlet line from the second reactor was disconnected. Upon inspection fouling was detected, the foul ant occluding approximately 40% of the cross section of the pipe. Analysis of the foulant using GC-MS identified maleic acid/anhydride as the main probable component(s)(>90%) with smaller amounts of acetic acid, and a trace amount of 1 ,2-benzene dicarboxylic acid. The fouling was cleared almost immediately upon passing water through the pipe. The results show that foulant can be removed with water and that a person skilled in die art familiar with multiphase flo ' arrangements would be able to design inputs for introduction of water into piping downstream of an ODH reactor in the fashion described by the metiiod disclosed herein.

[1842] A method is provided for the neutralization of Cr(VI) and removal of Cr( VI) residue from petrochemical plant components. A neutralizing solution, containing a reducing agent for conversion of Cr(VI) to Cr(III), is applied to the components, incubated for an appropriate duration, and Cr(VI) residue, comprising mainly Cr(III), is removed from die component. Cleaned components can be re-installed, and decontaminated waste and or catal st particles may be safely disposed of. Testing for presence of Cr(VI) prior to neutralization confirms need to perform the method, while testing following neutralization may highlight need for additional neutralization before re-installment of components or disposal of waste material.

[1843] Techniques are provided for clean-up of petrochemical plant infrastructure. More specifically, techniques are provided for the removal of Cr(VI) deposits on petrochemical plant components, including for example furnaces, reactors, and catalysts, by converting the deposits to Cr(IIl) by exposure to a neutralizing compound, and then removing the Cr(lll) residue.

[1844] Chromium (Cr) is a metallic element that can exist in several valence states, ranging from -2 up to +6. In nature, chromium exists almost exclusively in the divalent Cr(III) and hexavalent Cr(VI) oxidation states, while in industry', the oxidation states Cr(III) and Cr( VI) are accompanied by metallic or elemental chromium, Cr(0), as the most commonly found species of chromium. Of particular concern is the presence of hexavalent chromium, which can be found in compounds such as chromium oxide C1O ₃ , salts of chromate or dichromate (e.g. Na2Cr04, K ₂ C1O ₄ , Na2Cr ₂ 07, K ₂ C1 ₂ O ₇ ). Cr(Vl) contamination on industrial infrastructure is a common problem.

[1845] The method disclosed herein provides a method for neutralization and removal of Cr( VI) deposits that accumulate in or on petrochemical plant components. The method described entails applying a neutralizing solution to deposits in or on the petrochemical plant components, leaving neutralizing solution in contact with the deposits for an incubation period and subsequently removing the deposits. The method disclosed contemplates use for a variety of petrochemical plant components, such as the external surfaces of transfer line exchangers, the tube sheet of a convection section, and decoking drums, and removable components and solid waste such as bolts, catalyst particles and coke dust.

[1846] Neutralizing solution applicable for use with the disclosed method includes a reducing agent capable of reducing Cr(VI) to Cr(III) in a reasonable time period and does not react with the materials of the petrochemical plant components. The reducing agent must be soluble in a suitable solvent, and present in the neutralizing solution in concentrations from 0.5M to 3.5M. In a preferred embodiment, the neutralizing solution comprises the reducing agent sodium thiosulfate dissolved in water or equal parts water and ethylene glycol.

[1847] Application of neutralizing solution may be applied topically to external susfaces by means of spraying with an atomizer, wiping with a sponge soaked with neutralizing solution, or combinations thereof. For removable components and solid waste, application of neutralizing solution may be accomplished by submerging the components in a suitable container. Sealed compartments, including reactors, may be flooded with neutralizing solution.

[1848] In some embodiments, the neutralizing solution is allowed to incubate in contact with target deposits or whole components for periods ranging from 30 minutes up to 3 days. The duration chosen depends on the degree of Cr(VI) contamination, the strength of the neutralizing solution and the manner of application. Minimal contamination on exposed surfaces may benefit from topical application using a stronger neutralizing solution for a shorter time period. In contrast, reactors may be flooded with a weaker neutralizing solution, albeit for a longer time, dependent upon the expected levels of Cr( VI) accumulation.

[1849] Removal of Cr( VI) residue after neutralization may be achieved by wiping exposes surfaces or using vacuum pressure for desiccated residue particles. For submerged and flooded components, the Cr(VI) residue is thought to remain in the neutralizing solution, with removal of the neutralizing solution effecting removal of the deposits.

[1850] A potential use for the disclosed method is for catalyst removal from ODH fixed bed reactors. Commercialization of an ODH process, as an alternative to steam cracking, will require periodic removal of catalysts and associated packing materials. Accumulation of Cr(Vl) on porous catalyst supports and stainless steel packing material may be problematic because proposed mechanisms for catalyst removal includes possibility for dispersion of catalyst particles due to dry' conditions and possibility of static. The proposed method of Hooding the reactor with neutralizing solution prior to catalyst removal serves to neutralize dangerous Cr( VI) particles and minimize airborne dispersion by adding moisture to catalyst particles.

[1851] In some embodiments, exposed surfaces of petrochemical plant components can be tested for tire presence of Cr(VI) prior to neutralization, confirming the need for the method disclosed. Furthermore, neutralized surfaces may be tested following the procedure. Finally, catalyst particles and solid waste materials may also tested following neutralization. Continued presence of Cr(VI) alerts a user that continued or repeated neutralization following the method described herein should be employed before disposal of catalyst particles or solid waste materials.

[1852] The method disclosed herein is useful for removing Cr(VT) deposits on petrochemical plant components, the method comprising applying a neutralizing solution to deposits on components that are suspected of containing Cr(VI), incubating the neutral zation solution on or around the deposits, and finally removing the neutralized deposits. The phrase “Cr(VI) deposits” is meant to encompass materials where Cr(VI) can occur. Cr(VI) containing compounds include chrome oxides and chromates of barium, calcium, potassium, and sodium. Without being bound by theory, formation of CriVI) containing compounds is thought to be due to conversion of chromium Cr(IIT), present in stainless steel as ferrochrome, when exposed to alkali metals or alkali metal salts. For example, reaction of potassium or sodium with chrome oxides results in formation of Cr(VI). In addition, use of some cleaning chemicals containing any of potassium, sodium, calcium or barium when applied to petrochemical plant components can lead to formation of Cr(VI).

[1853] The target deposits for the method disclosed are those that exist in loose form on plant component surfaces or as part of components, such as catalyst particles, that are commonly removed and discarded as solid waste. Loose form is meant to include particles that are exposed, and represent a health hazard as they are amenable to contact with skin or capable of being inhaled. This includes deposits that are dislodged easily during cleaning and can form dust particles that can contact exposed skin or are capable of being inhaled by workers performing cleaning. The deposits are not necessarily visible with the linked eye. The method is not intended for application to Cr(VT) found deep within plant structures that is not susceptible to removal and does not represent an imminent health risk.

[1854] Plant components

[1855] Plant components that are susceptible to formation of Cr( VI) deposits include those components constructed with metal alloys which contain chromium. Examples include, but are not limited to, cracking facility components such as the tube sheet of a convection section, exterior surface of a transfer line exchanger, and decoking drums. Solid waste and porous media, such as support materials for catalysts used in fixed bed ODH reactors, also represent sources of potential €r( VI) deposits. Removal and disposal of such ma terials presents a health hazard when Cr(VI) is not neutralized.

[1856] Catalyst particles in chemical reactors may also experience Cr(Vl) accumulation over time. For example, Cr(VI) accumulation in oxidative deh drogenation (ODH) reactors due contamination of feedstock, such as ethane, can result when contaminating chromium undergoes conversion to Cr(VI) while inside the reactor and is then deposited into the porous surfaces of materials used to support ODH catalysts in addition, in a ty pical fixed bed ODH reactor, catalyst particles are surrounded by packing material tlsat may include flash arresting stainless steel balls, another potential source for accumulation of Cr(VI). Routine replacement of catalyst particles and packing material may expose workers to risk of exposure as unloading the catalyst particles and packing materials may result in airborne dispersion of Cr( VI) deposits.

[1857] Accumulation of Cr(Vl) in catalyst particles is not necessarily limited to ODH reactors but may occur in any reactor where catalyst particles include porous structures where Cr(VI) can be deposited mid accumulate over time. As with other components, there must be a source of chromium either as part of the component or introduced as a contaminant that is fed with a feedstock or other reactant into the reactor. The conditions must also exist where conversion is possible. For example, extreme temperatures and the presence of chemical reactants, such as metal salts or detergents. Finally, the type of reactor is not limiting either. It is possible that swing and fluidized bed type reactors may also experience accumulation of Cr(VI) deposits. Testing of components for the presence of Cr(VI) can be performed to indicate the desirability for using the method described herein.

[1858] Neutralizing solution

[1859] The neutralizing solution comprises an reducing agent that can reduce Cr(VI) to Cr(III) under typical atmospheric conditions existing during cleaning and within a reasonable time period. The reducing agent should be soluble in water or other suitable solvent, not generally react with petrochemical plant components, and result in products of reduction that are not toxic and are easily disposed of. In an embodiment of the disclosed method, the neutralizing solution comprises sodium thiosulfate dissolved in water. Sodium thiosulfate, dissolved in water as a 1 M solution, is slightly basic and would therefore, in contrast to acidic compounds, not promote metal corrosion of petrochemical plant components. The residue formed from reduction of Cr(Vi) after exposure to sodium thiosulfate is non-hazardous, comprising mainly sulfates of sodium, calcium and potassium, and is easily disposed as it is readily soluble in water.

[1 60] In some embodiments, the neutralizing solution comprises sodium thiosulfate dissolved as a 3M solution in a mixture of water and another sui table solvent, miscible with water. For example, a mixture of water and ethylene or propylene glycol may be used. This mixture allows a user to employ the disclosed method at temperatures lower titan if only water is used as the solvent. For obvious reasons, the method disclosed cannot be performed at temperatures below the freezing point of the neutralizing solution, and is preferably performed at temperatures at least 5°C higher than the freezing point of the neutralizing solution. Using only water as the solvent means tire method is limited to being performed at temperatures above G ^B C. In contrast a mixture of water and ethylene or propylene glycol may have a freezing point of around -40°C, depending on the ratio of water to ethylene or propylene glycol. Theoretically, the method could be employed around -40°C using a mixture such as tins however unlikely it is that a maintenance crew would be exposed to such a harsh temperature A skilled user understands that the ratio of wa ter to another miscible solvent impacts the freezing point. For example, while ethylene glycol concentrations of around 50% drop the freezing point to around ~40°C, higher titan 50% start to raise the freezing point.

[1861 J The concentration of reducing agent in the neutralizing solution may range from 0.1M to 3.5M.

Choosing a concentration requires consideration of the degree of Cr(VI) accumulation, the method of application of the neutralizing solution, and incubation time period preference. Higher concentrations are preferable when Cr(VI) accumulation is higher, there are limitations on volumes that can be applied, and the desirability for a shorter incubation period. Cr(VI) accumulation may be higher when intervals between removal are longer, or when conditions promote formation of Cr( VI). such as consistent exposure of materials with high levels of chrome to high levels of oxidizing material. Volume limitations may apply in situations where application of neutralizing solution is primarily topical in nature as opposed to instances where components can be submerged or flooded. In most commercial petrochemical facilities downtime is discouraged for financial reasons. If removal of Cr(VI) deposits requires shut down then a higher concentration would be preferred as it minimizes incubation period.

[1862] Application of neutralizing solution

[1863] Application of neutral zation solution is preferably applied so as to cover as much of the target deposits as possible. In an embodiment, exposed and accessible deposits on external surfaces of petrochemical plant components may be sprayed with neutralizing solution by means of an atomizer. The method disclosed also contemplates manual application with materials such as a sponge, that is soaked with neutralizing solution. Ideally, application leaves as little of the target deposits exposed and unavailable for neutralization. Applicatio n by spraying may also be employed when unloading components, such as catalyst particles. In this case, the components or catalyst particles are sprayed as they are unloaded. The components may be directed to an unloading vessel during unloading. It is contemplated that in addition to spraying during the unloading, the unloading vessel contains neutralizing solution, so that components are submerged in neutralizing solution upon exiting the reactor.

[1864] In some embodiments, for smaller components that are removable, application may be done by submerging, in a suitable container, removed components in neutralizing solution. Smaller components are meant to encompass components that can be handled by hand and without assistance from machinery'. This includes hut is not limited to, filters, coke dust, or solid waste. The smaller components, once removed, may be submerged in a weaker neutralizing solution, such as 0.5M sodium thiosulfate. The suitable container should be leak-proof, be non-reactive with the reducing agent and large enough for the components so that application can cover as much of the component as possible ideally, the size of the container will allow complete submersion of the components. However, the method contemplates submersion of only a part of the component, with the unsubmerged component subjected to the method at a later lime if need be. The method disclosed contemplates submersion of catalyst particles from a fixed bed ODH reactor to ensure neutralization of Cr(VI) before disposal. The volume of catalyst particles and availability of leak-proof containers of sufficient size may necessitate neutralizing catal st particles in batches. f 1865] For components that are compartmentalized or sealed application of neutralization solution can be achieved by Hooding the compartment or sealed space. Drums and fixed bed reactors represent examples where flooding may be employed. For fixed bed reactors, during shutdown the reactor outlet is closed and neutralizing solution can be added to flood the reactor. Flooding application of neutralizing solution may use a lower concentration of reducing agent because there is complete coverage of target deposits. In addition, a lower concentration is desirable due to larger volume of neutralizing solution required, compared to application via sponging or spraying with an atomizer.

[1866] Application may also include repetitive addition of neutralizing solution to target deposits. Tills may be required when single applications do not result in complete neutralization of contaminating Cr(VT). Spraying may be performed intermittently or on a consistent basis before removing neutralized deposits. This is also relevant for situations where applied neutralizing solution may drain away from the target deposits. For example, tubes or pipes where neutralizing solution sprayed on drips off, regardless of location where solution is applied. For compartment type components, including fixed bed ODH reactors, previously applied neutralizing solution may be drained and the compartment re-flooded with fresh neutralizing solution.

[1867] A combination of multiple methods can be implemented. For instance, target deposits on the outside surface of a transfer line exchanger may be sprayed, followed by wiping with a sponge soaked in neutralizing solution. For flooded applications access to deposits may be limited such that flooding is the only feasible method of application.

Incubation

[1868] Once applied the neutralizing solution needs to incubate so as to allow sufficient time for conversion of Cr(VT) to Cr(III). Time required for complete conversion will depend on the levels of Cr(VI) and the concentration of the reducing agent in the neutralizing solution. The incubation period may range from 30 minutes up to 3 days. Incubation periods over three days are inadvisable due to potential for reducing agent to negatively impact components. Shorter periods are useful for routine periodic cleaning where Cr(VI) accumulation is expected to be minimal. Cataly st replacement and reactor unpacking for fixed bed reactors are generally infrequent, scheduled at most once a year. As a result, Cr(VT) may be more prevalent in fixed bed reactors due to longer time periods between neutralization and removal. In that case, a longer incubation period is recommended. For more established and prevalent Cr(VI) accumula tion 24-36 hours may be required

[1 69] A user must consider the atmospheric conditions in which the method disclosed herein is to be used. In colder climates the reduction of Cr(Vl) may be slowed requiring longer incubation times or use of a higher concentration of reducing agent. Higher concentration of reducing agent may be limited by the temperature. For example, solubility of sodium thiosulfate at 20°C is 70.1 g/'ml (-4.44 M) but drops with reductions in temperature. With respect to use on exterior petrochemical plant components the disclosed method should not be employed at temperatures below the freezing point of the neutralizing solution. Using a mixture of ethylene glycol and water for the solvent may lower the freezing temperature. Ideally, the method disclosed herein is performed in seasons where average daily lows are at least 10°C or higher. For use in a reactor, there is no seasoned limitations provided the reactor includes a temperature control and can be heated to at least 10°C. f 1870] A person skilled in the art would understand that a volume of neutralizing solution with a particular concentration of reducing agent has an absolute limit on the amount of Cr(VI) that that volume can neutralize. For example, 3 moles of sodium thiosulfate are required to neutralize 1 mole of chromium oxide. Increasing incubation period will be to no effect if all the sodium thiosulfate has been oxidized. If the amount of neutralizing solution used is insufficient for complete neutralization a skilled user would appreciate that additional applications of neutralizing solution would be required.

[1871] Removal of deposits

[1872] Following incubation neutralized deposits can be removed by the most convenient means available.

For topical applications the deposits can be wiped away with a dry or wet towel or sponge. Alternatively, the deposits can be washed away using a pressure wash or other suitable means. Where neutralizing solution was applied and sufficient time lias passed such that the remaining residue is desiccated such that a dry' powder is all that remains, then negative pressure using a vacuum may be employed. For removable components and solid waste that were submerged in neutralizing solution, the solution can be drain and the components removed from the container. For compartmentalized components the neutralizing solution can be drained and the compartment ri nsed. Regardless of the method for removal, a user should be confident that complete neutralization lias occurred before disposing of targeted deposits. Experience with the method will inform a user when it is likely that the chosen neutralizing solution and associated incubation period are sufficient for complete neutralization. However, periodic testing of treated deposits should be undertaken to confirm the efficiency of the method. Discarding deposits with Cr(VI) still present, even at low levels, may pose a serious health risk.

[1873] For fixed bed reactors the neutralizing solution can be drained. Removal of the catalyst particles and packing material that includes flash arresting steel balls can proceed following draining of neutralizing solution. Typically, for removal of catalyst particles and packing material the bottom end of the reactor is opened and the components are allowed to drop away. Fixed bed ODH reactors are generally void of moisture, such that after shut down the catalyst and packing material are dry, and when allowed to drop away may disperse into the immediate surroundings Cr(VI) containing deposits. Removing the catalyst particles after flooding the reactor with neutralizing solution, reduces risk of dispersion of dry' catalyst and packing materials. The wet condition of the catalyst and packing material following application of neutralizing solution is not conducive to formation of or dispersion of airborne particles. Even if neutralization of Cr( VI) is not complete, any particles that contain Cr(VI), even at very low levels, will not likely be dispersed into the immediate surroundings.

[1874] Testing

[1 75] The method disclosed herein may also include confirming the presence or absence of Cr(VI) both before and after neutralization and deposit removal. 3M ChromateCheck™ swabs may be used for testing surfaces prior to and after neutralization. Quantofix® Chromate Test strips can be used for testing solutions used to neutralize removable components subsequent to neutralization.

[1876] In mi embodiment of the disclosed method, exterior surfaces of petrochemical plants may be wiped with a ChromateCheck™ swab. ChromateCheck™ swabs are self-contained units with glass ampules with reactive materials held inside the swab barrel, and a fiber tip. The swab reactive materials are mixed by shaking, followed by activation by squeezing crash points on the barrel until the reactive liquid reaches the fiber tip. The swab is then passed over the surface to be tested, allowing contact with the reactive liquid. The tip is then observed for a color change indicating the presence of Cr(Vl). Color changes to pink or purple confirm the presence of Cr(Vl). Neutralization following the disclosed method may proceed for surfaces that test positive for Cr(VT). The degree of contamination can be qualitatively assessed with darker purple coloring indicating higher levels of Cr( VI) relative to the lighter pink colors. The degree of contamination may inform selection of the concentration of reducing agent and the incubation period for neutralization. Higher contamination levels may benefit with stronger neutralizing solution and or longer incubation times. il877j Petrochemical plant components frequently include buildup of other deposits, such as oxidation products, coke and grease. Swabbing sections that results in the fiber tip turning black do not indicate the presence or absence of Cr( VI). A strong purple color change may be masked by the black color. If possible, a user should check multiple regions, including regions where surfaces appear to be clear of oxidation products, coke and grease. Regions with yellowish coloring may indicate presence of Cr( VI). Testing by swabbing is a quick way of assessing contamination by Cr(\T), but should not be considered conclusive. Further testing using more sensitive methods, such as X-ray diffraction (XRD), may also be employed.

[1878] In some embodiments, surfaces that have been subjected to neutralization may be assessed for Cr(VI) following removing of deposits. The same surfaces assessed prior to neutralization may be swabbed in the same fashion using a ChromateCheck™, or similar method. Presence of Cr(VI) following a first round of neutralization indicates that further rounds of neutralizing solution application and incubation should be performed. If qualitative assessment suggests that Cr(VE) levels are still high then a stronger neutralizing solution or longer incubation period, or both, should be employed.

[1879] The deposits removed after neutralization may also be tested for the presence of Cr(VI) Disposal of contaminated deposits presents a safety hazard. Assessing contamination using QuantoFix® Chromate test strips is recommended. The removed deposits may be dissolved in water. The resulting solution is then tested with test strips producing a color change similar to that produced by the ChromateCheck™ swabs. If Cr(VI) is present in the deposits then the surfaces should be subjected to another round of neutralization. In addition, the deposit containing solution may also be subjected to neutralization by addition of a stronger neutralizing solution. Disposal of deposits should not be performed until testing indicates the absence of Cr(Vi).

[1880] For components that are submerged in neutralizing solution, the QuantoFix® Chromate test strips can be employed prior to draining the solution. If a purple coloring does not develop, the solution can be drained and tire components or solid waste removed from the container. The components can be cleaned and reinstalled. The solid waste can be safely disposed. If Cr(Vl) persists, then neutralizing solution can be replenished with fresh neutralizing solution or a longer incubation period can be utilized. A longer incubation period may not be viable if all reducing agent in the solution Isas been oxidized.

[1881] In the case of catalyst particles in fixed bed ODH reactors it may not be possible to test for tire presence of Cr(VI) prior to removal. Testing may be limited to the removed particles after incubation following flooding foe reactor with neutralizing solution. Presence of Cr(VI) in the removed catalyst particles indicates a need for a stronger neutralizing solution or longer incubation period for the next scheduled catalyst removal The removed particles can be subjected to further neutralization prior to disposal.

[1882] Examples

[1883] Testing of a variety locations of a steam cracking furnace revealed elevated levels of Cr(Vl) in some locations. In some instances testing was done by swabbing deposits directly with a 3M ChrornateCheck swab, and in other instances samples of deposits were collected in the form of dust or other loose particles and tested by dissolving in water and assessed using Quantolix Chromate test strips. In addition, swab testing of supported GDH catalyst that was active for 12 months also revealed presence of Cr(VI).

[1884] Bolts from a transfer line exchanger were chosen for demonstration of the method described herein. Testing of the bolts by swabbing confirmed the presence of Cr(VI). The swabs from four bolts displa ed a strong pink/purpie color, while a control swab showed no purple coloring. Neutralizing solutions were prepared as 1M solutions of four sulfur containing compounds. The pH of each solution was measured (Table U 1). The solutions were sprayed on separate Cr(VI) positive bolts and allowed to incubate for 24 hours. After 24 hours the applied neutralization solution had dried and was retested for the presence of Cr(VI) in multiple locations on each bolt. The most effective neutralizing solution was 1M sodium thiosulfate, which removed most but not all of the contamination. The other chemicals also displayed a propensity to neutralize Cr(VI), but not to the same degree as sodium thiosulfate. A second experiment using 3M sodium thiosulfate resulted in complete Cr(VI) neutralization after 24 hours.

[1885] Table U 1 : Neutralizing agent comparison -

[1886] This application claims the benefit of the earlier filing date of Canadian application serial number 3008612 filed on June 18, 2019. The contents of Canadian application serial number 3008612 are incorporated herein by reference in their entirety.

[1887] Techniques are provided for removing mixed oxide catalysts used to dehydrogenate one or more Ci- j paraffins to corresponding alkenes from one or more reactors and connected piping. [1888] Oxidative dehydrogenation catalysts including mixed oxides of Mo, V, Nb, Te, and optionally a promoter may be dissolved in aqueous solutions of oxalic acid. This permits the removal of catalyst and catalyst residues from reactors for the oxidative dehydrogenation of paraffins and particularly ethane.

[1889] Techniques are provided for a method to remove from one or more vessels and associated piping a bed of a catalyst selected from the group consisting of:

[1890] i) catalysts including Moo.s-uoVo.i-iNbo _.i -o _. rTeo _. oi-o _. nX _t wO _d where X is selected from Pd, Sb Ba, Al, W, Ga, Bi, Sn, Cu, Ti, Fe, Co, M, Cr, Zr, Pt, €a, and oxides and mixtures thereof, and d is a number to satisfy the valence of the catal st; and

[1891] ii) catalysts of the formula Mo _? \¾M) _e T6 _e( ¾. wherein a is from 0.75 to 1.25, preferably from 0.90 to 1.10; b is from 0.1 to 0.5, preferably from 0.25 to 0.4;

[1892] c is from 0 to 0.5, preferably from 0.1 to 0.35; e is from 0 to 0.35 preferably from 0.1 to 0.3, and d is a number to satisfy die valence state of the mixed oxide catal st,

[1893] optionally on an alumina support typically including alumina in the form of AI2O3 or A1(0)0H or combination thereof including contacting the catalyst with from 10 to 100 niL of not less than a 0.5 molar solution typically not less than 1 M up to the solubility limit of oxalic acid in an aqueous solution at the temperature of treatment per g of catalyst at a temperature from 20 °C up to the boiling temperature of a saturated solution, such as greater than 60 °C or greater than 80 °C, for a period of time for at least 1 hour aud in some cases 20 or more hours. [1894] The treatment can be carried out using agitation (e.g., cyclic pumping of the solution through reactor tubes).

[1895] In some embodiments, there is provided any of the above embodiments wherein the catalyst forms fouling on the piping associated with the reactor.

[1896] In some embodiments, there is provided any of the above embodiments wherein the piping is steel or stainless steel.

[1897] In some embodiments, there is provided any of the above embodiments wherein the catalyst is a bed in one or more reactors.

[1898] In some embodiments, there is provided any of the above embodiments wherein the catalyst is supported on alumina.

[1899] In some embodiments, valuable metals including Pd, Pt, and An are separated from the solution of catalyst and optional support by one or more suitable means including filtration, precipitation, and floatation.

[1900] The catalyst may include a mixture of metal oxides having a composition:

[1901 ] M01Vo.1-1Nbo _. o-1Teo.o-o _. 2Xo.2O _d where X is selected from Pd, Sb Ba, Al, W, Ga, Bi, Sn, Cu, Ti, Fe, Co, Ni, Cr, Zr, Ca, and oxides and mixtures thereof, and d is a number to satisfy the valence of the catalyst.

[1902] In some embodiments, the catalyst includes catal sts of the formula Mo _a V _b Nb _c Te _e O _d wherein: a is from 0.75 to 1.25, such as from 0.90 to 1.10; b is from 0.1 to 0.5, such as from 0.25 to 0.4; c is from 0 to 0.5, such as from 0.1 to 0.35 ; e is from 0 to 0.35 such as from 0.1 to 0.3 ; and d is a number to satisfy the valence state of the mixed oxide catal st. [1903] The catalyst may Slave the formula: Mo ₁ V0 ₂₅ -0 ₄₅ Te0 _.10 -0 _.i6 Nb0 ₁₃ -0 _l eXo-mOa wherein d is a number to satisfy the valence of the catalyst.

[1904] In some embodiments, the catalyst may have the formula: MoiVo _. n-o^Nbo-o _. nTeo-o.nX _t wO _d wherein d is a number to satisfy' the valence of the catalyst.

[1905] in some embodiments, the catalyst may have the formula: Mo 1.0 Vo^-o.isNbo.ii-o.nTeo.io-o.nXo^O _d wherein d is a number to satisfy the valence of the catalyst.

[1906] in some embodiments, the catalyst may have the formula: MoiVo^s-oasNbo.is-o.ieTeo.i-o _.i eXo^O wherein d is a number to satisfy the valence of the catalyst.

[1907] In some embodiments, the catalyst may have the formula: MoiVo _.25 -o _.35 Nbo.i ₃ -o.i ₆ Teo.i-o _.i6 Xo- ₂ 0 wherein d is a number to satisfy the valence of the catalyst.

[1908] In some embodiments, the catalyst may have the formula: M01.0 V0.12-0.19Nb0.19-0.20Te0.06-0.07 Xo-iOd wherein d is a number to satisfy the valence of the catalyst.

[1909] In the catalysts, element X may be present in an amount from 0 up to 2 atoms per atom of Mo; in some embodiments, from 0 up to 1 atoms per atom of Mo; in some embodiments, from 0.001 to 1 atom per atom of Mo; in further embodiments, from 0.001 to 0.5 atoms per atom of Mo; in some embodiments, from 0.001 to 0.01 atoms per atom of Mo.

[1910] The catalyst may be prepared, for example, by mixing aqueous solutions of soluble metal compounds such as hydroxides, sulphates, nitrates, halides, salts of lower (C1-5) mono- or di- carboxylic acids and ammonium salts or the metal acid per se. For instance, the catalyst can be prepared by blending solutions such as ammonium metavanadate, niobium oxalate, ammonium molybdate, telluric acid etc. and. in some embodiments, subjecting the resulting solution to a hydrothermal process under an inert atmosphere and heating to a temperature from 140 °C to 190 °C, in some embodiments from 140 °C to 180 °C, in some embodiments from 145 °C to 175 °C for not less than 6 hours, in some Instances not less than 12 hours, in some embodiments up to 30 hours, or more.

[1911] The pressure in the reactor (Pan- reactor or autoclave) may range from 1 to 200 psig (6.89 kPag to 1375 kPag).

[1912] In some embodiments, the pressure in the pressurized reactor is adjusted and maintained from 30 to 200 psig (206 kPag to 1375 kPag), in some embodiments from 55 psig (380 kPag) to 170 psig (1170 kPag) above atmospheric pressure.

[1913] In some embodiments, the pressure in the reactor (autoclave) may be up to about 10 psig (68.9 kPag), preferably from 1 to 8 psig (6.89 kPag to 55.1 kPag), in some embodiments less than 5 psig (34.4 kPag) above atmospheric pressure.

[1914] The pressures in the reactor are maintained using a pressure relief valve. At lower pressures, the pressure may be maintained by passing the off gas through a column of a fluid such as water or a dense fluid (e.g.. mercury). Optionally, there may be a condenser upstream of the reactor outlet. If present, the condenser is operated at a temperature above 0 °C and below reaction temperature. Gaseous product species are vented from the reactor as described above. f 1915] In some embodiments, the solution for the hydrothermal treatment (catalyst precursor) may include small amounts of ¾{¾ from 0.3-2.5 mL of a 30 wt. % solution of aqueous ¾0 per gram of catalyst precursor. |1916] The resulting solution is then dried typically in air at 100-150 °C and calcined in a flow of inert gas such as those selected from the group consisting of N ₂ , He, Ar, Ne, and mixtures thereof at 200-600°C, preferably at 3Q0~500°C. The calcining step may take from 1 to 20 hours, from 5 to 15 hours or about 10 hours. The resulting oxide is a friable solid, typically insoluble in water.

[1917] In some embodiments, the product from the hydrothermal treatment is treated with from 0.3-2.5 mL of a 30 wt. % solution of aqueous ¾(¾ per gram of catalyst precursor. In some embodiments, there may be a double peroxide treatment in the hydrothermal process and subsequent to drying the catalyst. The dried catalyst is then deposited on an alumina support using conventional methods such as a wet impregnation method or spray drying and the like. In some embodiments, from 10 to 95 wt. %, from 25 to 80 wt %, or from 30 to 45 wt. % of the catalyst is bound or agglomerated with from 5 to 90 wt. %, from 20 to 75 wt. %. from 55 to 70 wt. % of a binder selected from the group consisting of acidic, basic or neutral binder slurries of Al _2{ ¾, and AIO(OH), and mixtures thereof.

[ 1918] The catalyst can be loaded into one or more reactors in series or parallel The reactors maybe be fixed or fluidized bed reactors, in some cases similar to FCC type crackers. In some embodiments, the reactor may be a tube shell type reactor with the catalyst loaded into the tube and tube plates above and below the tubes to permit reactants to flow through the catalyst bed.

[1919] The catalyst may be used for the oxidative dehydrogenation (ODH) of a mixed feed including ethane and oxygen in a volume ratio from 70:30 to 95:5 and optionally one or more C alkanes or alkenes and optionally a further oxygenated species including CO and C0 at a temperature less than 370 °C, a gas hourly space velocity of not less than 100 hr ¹ , and a pressure from 0.8 to 7 atmospheres including passing the mixture through the above catalyst. The ODH process can have a selectivity to ethylene of not less than 90%. The gas hourly space velocity of the ODH process is not less than 500 hr ¹ , not less than 1500 hr ¹ , such as 3000 hr ¹ . The temperature of the ODH process can be less than 370 °C, less than 360 °C, or less than 350 °C.

[1920] Depending on the reaction conditions, the product stream can include ethylene, water and one or more of carbon dioxide casfeon monoxide, carbonic acid and acetic acid. The product stream, and particularly water, carbonic acid and acetic acid may leach one or more components from the catalyst or support. Depending on the duration of the reaction this can result in dissolution of the catalyst or its components and the dissolution of the support and deposition of the components or support and salts thereof such as carbonates and acetates on the walls of the reactor and potentially among catalyst particles. This can result in reduced operating efficiency of the reactor system (e.g , reactors and associated piping). This may result in costly down time to clean the catalyst beds, (particularly in tube and shell type reactors) and the associated piping.

[1921] It has been found that the catalyst, support (Al i¾, and AIO(OH)), and associated bi-products (e.g. salts etc.) can be removed from the catalyst beds and associated piping by dissolving or contacting them with from 10 to 100 rnL of a up to the boiling temperature of a saturated solution of not less titan 0.5 of oxalic per 1 to 5 g of catalyst and associated bi-products at a temperature from 20 °C to up to the boiling temperature of a saturated solution preferably greater than 60 °C, such as greater than 80 °C, for a period of time for at least 1 hour in some cases 20 or more hours.

[1922] The amount of catalyst, support, and associated catalyst and support biproducts in the solution of oxalic acid may be determined by a number of conventional means such as analysis of the solution (e.g., FTIR, etc). [1923] The solution of the catalyst, support, and biproducts may be subject first to filtration to remove particulates such as catalyst support, and then the solution may be dried or substantially dried to recover the metal components of the catalyst if the solution contains valuable metals such as Pd, Pt, and Au, they may be separated by suitable means including filtration, precipitation floatation etc. The solution of catal st components may be used as a starter to begin a new hydrothermal treatment. Further, the solution can be analyzed for metallic components and can be adjusted the composition to that require for a hydro thermal treatment and potentially regenerate the catalyst. [1924] Techniques are provided for a method to remove from one or more reactors and associated piping a catalyst selected from the group consisting of: i) a catal st including Moo _.9 -i _.i Vo _.i -iNbo _.i -iTeo _. oi-o _.2 Xo- ₂ 0 _d where X is selected from Pd, Sb Ba, Al, W, Ga,

Bi, Sn, Cu, Ti, Fe, Co. Ni, Cr, Zr, Pt, Ca and oxides and mixtures thereof, and d is a number to satisfy the valence of the catalyst; and

[1925] ii) a catalyst of the formula

Mo _a VbNb _c Te _e Od wherein a is from 0.75 to 1.25; b is from 0.1 to 0.5; c is fromO to 0.5; e is from 0 to 0.35; and d is a number to satisfy the valence state of the mixed oxide catalyst. The catalyst is optionally on an alumina support.

[1926] The method includes contacting the cataly st with from 10 to 100 mL of not less than 0.5 molar solution of an oxalic acid per i to 5 g of catalyst at a temperature from 20 °C to 100 °C for a period of time of not less than 1 hour.

[ 927] In some embodiments, there is provided a method, wherein the treatment is earned out using agitation including cyclic pumping of the solution through reactor tubes

[1928] In some embodiments, there is provided a method, wherein the catalyst and optional support forms fouling on the piping associated with the reactor.

[1929] In some embodiments, there is provided a method, wherein the piping is steel or stainless steel.

[1930] In some embodiments, there is provided a method, wherein the catalyst is a bed in one or more reactors.

[1931] In some embodiments, there is provided a method, wherein the catalyst is supported on alumina (Al(O)OH).

[1932] In some embodiments, there is provided a method, wherein valuable metals including Pd, Pt and

Au are separated from the solution of catal st and optional support by one or more suitable means including filtration, precipitation, and floatation.

[1933] In some embodiments, there is provided a method, wherein the catalyst has Site formula: M01V0.25- oasTeo _. 10-0. _l eNbo _. 13-0. ieXo-o ₂ 0 _d wherein d is a number to satisfy the valence of the catalyst. f 1934] In some embodiments, there is provided a method, wherein the catalyst has the formula: Mo i Vo 12-0 49- Nbo. ₁ -o. ₁₇ Teo. ₁ -o. r _/ Xo-rO _d wherein d is a number to satisfy the valence of the catalyst.

[1935] In some embodiments, there is provided a method, wherein the catalyst has the formula: M01.0 V0.32- o _.49 Nbo _. 14-0. i _? Teo 10-0.17X0-2O ₁ J wherein d is a number to satisfy the valence of the catalyst.

[1936] in some embodiments, there is provided a method, wherein the catalyst has the formula: M0 V0 ₂₅ -0 ₄₅ - Nbo.i ₃ -o _.i6 Teo.i-o.i ₆ Xo- ₂ 0 _d wherein d is a number to satisfy the valence of the catalyst.

[1937] in some embodiments, there is provided a method, wherein the catalyst has the formula: M0 V0 ₂₅ -035- Nbo.i ₃ -o _.i6 Teo.i-o.i ₆ Xo- ₂ 0 _d wherein d is a number to satisfy the valence of the catal st.

[1938] In some embodiments, there is provided a method wherein the catalyst has the formula: Mo 1.0 V0 12- o.isNbo.19-0.2oTeo.06-0.07 wherein d is a number to satisfy the valence of the catalyst

[1939] In some embodiments, there is provided a method wherein X may be present in an amount from 0 up to 1 atoms per atom of Mo.

[1940] In some embodiments, there is provided a method wherein X may be present in an amount from 0.001 to 1 atoms per atom of Mo.

[1941] Example VI

[1942] Ten grams of oxalic acid was dissolved in 100 mL of distilled water. This mixture was heated to 80 °C in a hot water bath for 30 minutes. To this clear, colorless solution was charged 1.0 g of extruded catalyst of the formula M0 ₁ Vo. ₄₃ Teo. ₁₇ Nbo.i ₆ O _d supported on alumina (Al(O)OH) (catalyst figure 75). The mixture was left to sit unstirred in the 80 °C water bath for 30 minutes. After which 98% of the catalyst was dissolved. After 1 h in the 80 °C water bath the catalyst was completely dissolved. The alumina from the extruded catalyst remained undissolved (Figure 76).

[1943] Example V2

[1944] Thirty' grams of oxalic acid was dissolved in 300 mL of distilled water. This mixture was heated to 80 °C in a hot water bath for 30 minutes. To this colorless solution was charged 3.0 g of extruded catalyst of the formula with alumina. The mixture was left stirring at 600 rpm in the 80 °C water bath for overnight. After which the catalyst and alumina (Al(O)OH) w¾s completely dissolved. A clear, blue solution was formed

[1945] Example V3

[1946] The 1/8” stainless steel 316 tube contaminated with catalyst residue (Figure 77) was submerged approximately 1.5 cm below the surface of a mixture of 3.0 g catalyst of the formula M01Vo 3Teo _. 17Nbo _.i 6O _d , 30 g oxalic acid mixture, in 300 mL of distilled water, in an 80 °C. oil bath for 24 h. After 24 h no visible signs of corrosion or etching of the stainless steel was visible (Figure 78).

[1947] These examples show that mixed oxide catalysts comprising M01Vo43Teo _. 17Nbo _.i 6O _d are soluble in aqueous solutions of oxalic acid and that such solutions do not etch or attack stainless steel.

[1948] An air separation unit in an ODH process is provided. Techniques are provided for a method for the regeneration of oxidative deh drogenation catalysts, including integration of nitrogen enriched off-gas, to make regeneration more stable and effective. f 1949] Techniques are provided for a process for regeneration of catalysts used in at least one oxidative dehydrogenation reactor for the oxidative dehydrogenation of a lower alkane into a corresponding alkene, the process comprising: flowing inert gas into the at least one oxidative dehydrogenation reactor in the absence of steam until the temperature inside the reactor is between 280 °C and 380 °C; and flowing regeneration gas at a temperature of between 280 °C and 380 °C comprising dilute air in which the concentration of oxy gen is less titan about 8 vol. % into the at least one oxidative dehydrogenation reactor until the C(¾ concentration in the gas effluent is less than 110% of the CO2 concentration in the regeneration gas, and the O2 concentration in the gas effluent is at least 90% of the 0 ₂ concentration in the regeneration gas.

[1950] In some embodiments, the process is followed by flowing pure air to the at least one oxidative dehydrogenation reactor.

[1951] Techniques are provided for a process for regeneration of catalysts used in at least one oxidative dehydrogenation reactor for Site oxidative dehydrogenation of a lower alkane into a corresponding alkene, the process containing at least one regeneration bed, the process comprising: flowing inert gas into the at least one oxidative dehydrogenation reactor in the absence of steam until the temperature inside the reactor is between 280 °C and 380 °C; flowing regeneration gas to at least one regeneration bed at a temperature ofbe tween 280 °C and 380 °C in which the concentration of ox gen is less than about 8 vol. % into the regeneration bed until the CO2 concentration in the gas effluent is less than 110% of the CO2 concentration in the regeneration gas, and the <¼ concentration in the gas effluent is at least 90% of the 0 _? concentration in the regeneration gas; and flowing pure air to at least one regeneration bed at a temperature of between 280 °C and 380 °C.

[1952] Techniques are provided for a process for regeneration of catalysts used in at least one oxidative dehydrogenation reactor for the oxidative dehydrogenation of a lower alkane into a corresponding alkene, the process containing at least one regeneration bed, the process comprising: flowing a mixture of regeneration gas and pure air in the absence of steam to at least one regeneration bed at a temperature of between 280 °C and 380 °C in which the concentration of oxygen is less than about 8 vol % into the regeneration bed until the CC¾ concentration in the gas effluent is within 1101 of the CO2 concentration in the regeneration gas, and the (¾ concentration in the gas effluent is within 90% of the O2 concentration in the regeneration gas; and flowing pure air to at least one regeneration bed at a temperature of between 280 °C and 380 °C.

[1953] Techniques are provided for a process for regeneration of catalysts used in at least one oxidative dehydrogenation reactor for the oxidative dehydrogenation of a lower alkane into a corresponding alkene, the process containing at least one regeneration bed, the process comprising: flowing inert gas into the at least one oxidative dehydrogenation reactor in the absence of steam until the temperature inside the reactor is between 280 °C and 380 °C; flowing regeneration gas to at least one regeneration bed at a temperature of between 280 °C and 380 °C in which the concentration of oxygen is less than about 8 vol. % into the regeneration bed until the CO2 concentration in the gas effluent is less than 110% of the C0 ₂ concentration in the regeneration gas, and the 0 ₂ concentration in the gas effluent is at least 90% of the 0 ₂ concentration in foe regeneration gas; and flowing dilute air which lias a maximum 0 ₂ concentration of 8 vol. % to at least one regeneration bed at a temperature of between 280 °C and 380 °C. f 1954] Techniques are provided for a process for regeneration of catalysts used in at least one oxidative dehydrogenation reactor for the oxidative dehydrogenation of a lower alkane into a corresponding alkene, the process containing at least two regeneration beds in series, the process comprising: flowing regeneration gas or a mixture of regeneration gas and pure air in the absence of steam to the first regenera tion bed at a temperature of between 280 °C and 380 °C in which the concentration of oxygen is less than about 8 vol. % into the first regeneration bed until the CO _? concentration in the gas effluent is is less than 110% of the Ct¾ concentration in the regeneration gas, and the (.¾ concentra tion in the gas effluent is at least 90% of the (.¾ concentra tion in the regeneration gas; and flowing pure air to at least the second regeneration bed at a temperature of between 280 ^C C and 380 °C.

[1955] Techniques are provided for a process for regeneration of catalysts used in at least one oxidative dehydrogenation reactor for Site oxidative dehydrogenation of a lower alkane into a corresponding alkene, the process containing at least two regeneration beds in series, the process comprising: flowing tire used air stream from the second regeneration bed to the first regeneration bed, or a mixture of the used air stream gas and pure air, in the absence of steam, to the first regeneration bed at a temperature of between 280 °C and 380 °C in which the concentration of oxygen is less than about 8 vol. % into the first regeneration bed until the CO2 concentration in the gas effluent is less than 110% of the CO _? concentration in the regeneration gas, and the O _? concentration in the gas effluent at least 90% of the O _? concentration in the regeneration gas; and flowing pure air to at least the second regeneration bed at a temperature of between 280 °C and 380 °C.

[1956] In some embodiments, the at least one oxidative dehydrogenation reactor comprises a single fixed bed type reactor, including such as tube-in-shell type reactors.

[1957] In some embodiments, the at least one oxidative dehydrogenation reactor comprises a single fluidized bed type reactor.

[ 958] In some embodiments, the at least one oxidative dehydrogenation reactor comprises a swing bed type reactor arrangement.

[1959] In some embodiments, the at least one oxidative dehydrogenation reactor comprises an ebulated bed type reactor arrangement.

[1960] In some embodiments, the at least one oxidative dehydrogenation reactor comprises a rotating bed type reactor arrangement.

[1961 ] In some embodiments, the at least one oxidative dehydrogenation reactor comprises a heat pump type reactor arrangement.

[1962] In some embodiments, the process further comprises more than one oxidative dehydrogenation reactor connected in parallel, with each other oxidative dehydrogenation reactor comprising the same or different oxidative dehydrogenation catalyst.

[1963] In some embodiments, at least one of the oxidative dehydrogenation reactors comprises a fixed bed type reactor.

[1964] In some embodiments, at least one of the oxidative deh drogenation reactors comprises a fluidized bed type reactor. f 1965] In some embodiments, at least one of the oxidative dehydrogenation reactors comprises a swing bed type reactor arrangement.

[1966] In some embodiments, at least one oxidative dehydrogenation reactor comprises a ebulated bed type reactor arrangement.

[1967] in some embodiments, at least one oxidative dehydrogenation reactor comprises a rotating bed type reactor arrangement.

[1968] in some embodiments, at least one oxidative dehydrogenation reactor comprises a heat pump type reactor arrangement.

[1969] In some embodiments, the lower alkane is ethane.

[1970] In some embodiments, at least one of the oxidative dehydrogenation catalysts comprises a mixed metal oxide selected from the group consisting of: i) catalysts of the formula: Mo,A' ^’ iTe _c Nb _a Pd _c wvherein a, b. c, d, e and f are the relative atomic amounts of the elements Mo. V, Te, Nb, Pd and O, respectively; and when a = 1, b = 0.01 to 1.0, c = 0.01 to 1.0, d = 0.01 to 1.0, 0.00 < e < 0.10 and f is a number to satisfy the valence state of the catalyst; ii) catalysts of the formula: Ni _g A$,D _y O/wherem: g is a number from 0.1 to 0.9, preferably from 0.3 to 0.9, most preferably from 0.5 to 0.85, most preferably 0.6 to 0.8; h is a number from 0.04 to 0.9; i is a number from 0 to 0.5: j is a number from 0 to 0.5; and f is a number to satisfy the valence state of the catalyst; A is selected from the group consisting of Ti, Ta, V, Nb, Hf, W, Y, Zn, Zr, Si and A1 or mixtures thereof; B is selected from the group consisting of La, Ce, Pr, Nd, Sm, Sb, Sn, Bi, Pb, Tl, In, Te, Cr, Mn, Mo, Fe, Co, Cu, Ru, Rh, Pd, Pt, Ag, Cd, Os, Ir, An, Hg, and mixtures thereof; D is selected from the group consisting of Ca, K, Mg, Li, Na, Sr, Ba, Cs, and Rb and mixtures thereof; and O is oxygen; iii) catalysts of the formula: Mo _o E*G _/ 0 _/ · wherein: E is selected from the group consisting of Ba, Ca,

Cr. Mn, Nb, Ta, Ti, Te, V, W and mixtures thereof; G is selected from the group consisting of Bi, Ce. Co, Cu, Fe, K, Mg. V, Ni, P, Pb, Sb, Si, Sn, Ti, U, and mixtures thereof; a = 1; k is 0 to 2; 1 = 0 to 2, with the proviso that the total value of 1 for Co, Ni. Fe and mixtures thereof is less than 0.5; and f is a number to satisfy die valence state of the catalyst; iv) catalysts of the formula: V _M Mo„Nb ₀ Te„Me _o O _/ wherein: Me is a metal selected from the group consisting of Ta, Ti, W, Hf, Zr, Sb and mixtures thereof; m is from 0.1 to 3; n is from 0.5 to 1.5; o is from 0.001 to 3; p is from 0.001 to 5; q is from 0 to 2; and f is a number to satisfy the valence state of the catalyst; and v) catalysts of the formula: MoLnCATZ,,M,,O/ wherein: X is at least one of Nb and Ta; Y is at least one of Sb and Ni; Z is at least one of Te, Ga, Pd, W. Bi and Al; M is at least one of Fe, Co, Cu, Cr. Ti, Ce, Zr, Mn. Pb, Mg, Sn, Pt, Si, La, K, Ag and In; a=1.0 (normalized); r = 0.05 to 1.0; s = 0.001 to 1.0; t = 0.001 to 1.0; u = 0.001 to 0.5; v = 0.001 to 0.3; and f is a number to satisfy the valence state of the catalyst.

[1971] In some embodiments, at least one of the oxidative dehydrogenation catalysts comprises a mixed metal oxide selected from the group consisting of the formula: MoiVo i-iNbo _.i -iTeo oi-o 2X0-0 2O _f wherein X is selected from Pd, Sb Ba, A3, W, Ga, Bi, Sn, Cu, Ti, Fe, Co, Ni, Cr, Zr Ca and oxides and mixtures thereof, and f is a number to satisfy the valence state of the catalyst.

[1972] In some embodiments, the ODH operation can be a fixed bed 7903, as shown in Figure 79. The ODH operation can have an air separation process 7901. in this configuration, a nitrogen waste stream 7909 from the air separation process 7901 is fed to the ODH reactor 7903 while alkane is flowing through the reactor when the bed requires regeneration. The nitrogen waste stream from the air purification unit 7909, which contains N2, CO2, etc., hut substantially no ¾0 is fed to the fixed bed 7903 at the operating temperature of 300-330 °C. The flow is maintained until the CO2 concentration in the gas effluent 7910 decreases to the C0 ₂ concentration in the nitrogen waste stream 7909 and the Q ₂ concentration in the gas effluent 7910 increases to tire 0 ₂ concentration in the nitrogen waste stream 7909. The regeneration flow is then switched from tire nitrogen waste stream 7909 to pure air 7912 at the same temperature.

[1973] In an embodiment, in regeneration mode, the pure air stream 7912 is mixed with the nitrogen w'aste stream 7909 to generate the desired 0 ₂ feed concentration, which is < 8 vol. % 0 ₂ . This combined stream 7913 is fed to the fixed bed at 300-330 °C. This embodiment can also be applied to fixed bed ODH operation, as per Figure 79, swing bed ODH operation, as per Figure 80, and to fluidized bed ODH operation, as per Figure 81.

[1974] In an embodiment, in fluidized bed operation. Figure 81, the combined stream 8114 is fed to a fluidized bed regenerator 8104 at 300-330 °C. The fully regenerated catalyst could then be transported via 8116 to the fluidized bed reactor 8103. When the catalyst becomes deactivated, it could then be transported via 8115 from the fluidized bed reactor 8103 to the fluidized bed regenerator 8104.

[1975] In an embodiment, the ODH operation can have two fluidized bed regenerators 8204 and 8205, as shown in Figure 82. The ODH operation can have an air separation process 8201. A nitrogen waste stream 8211 from an air separation unit 8201 substantially free of water (steam) can be fed to the first regeneration bed 8204, which contains a substantially deactivated catalyst at 300-330 °C, which came from the fluidized bed reactor 8203 via 8216 for an appropriate interval. The regeneration flow is then switched from the nitrogen waste stream 8211 to pure air 8214 at the same temperature. In an embodiment, a pure air stream 8214 is optionally mixed with the nitrogen w'aste stream 8211 to generate the desired 0 ₂ feed concentration, which is < 8 vol. % 0 _2, Once partially regenerated, the catalyst can be transported via 8217 to another fluidized bed regenerator 8205. A pure air stream 8218, is fed to the second regeneration bed 8205, which can contain partially deactivated catalyst, at 300-330 °C. Once fully regenerated, the catalyst in fluidized bed regenerator 8205 can be transported via 8219 to the fluidized bed reactor 8203.

[1976] In an embodiment, the ODH operation can contain reactors that are swing bed, ebulliated bed, or any variation of moving bed, as shown in Figure 83. In a swing bed operating mode, in one cycle an ODH reactor 8303 operates at a high conversion mode. A second ODH reactor 8304 operates in a mode ensuring minimum to no residual 0 ₂ in the final ODH product, and maximum conversion of residual ethane. The catalyst in this second ODH reactor 8304 becomes 0 ₂ depleted and rapidly loses its activity. A second regeneration bed 8306 is freshly deactivated ODH catalyst which can be regenerated with the off-gas 8315 fro t a first regeneration bed 8305. This off gas 8315 is N ₂ emiched regeneration stream with 0 ₂ < 8 vol. % containing substantially no water or steam. Air 8313 can be fed to the first regeneration bed 8305 at the operating temperature of 300-320 °C for an appropriate interval. The flow is maintained until the CQ> concentration in the gas effluent 8315 decreases to the C<¾ concentration in the air stream 8313 and the G ₂ concentration in the gas effluent 8315 increases to the 0 ₂ concentration in the air stream 8313. The first regeneration bed 8305 can contain partially regenerated ODH catalyst which is not prone to thermal runaway and can he safely regenerated with air. When finished a cycle as described, catalyst from the second ODH reactor 8304 can be considered fully deactivated and oxygen depleted, whereas catalyst in the first regeneration bed 8305 can be considered fully regenerated and oxygen saturated. The sequence can then be changed, for example, 8303 becomes 8304. 8304 becomes 8306, 8306 becomes 8305, and 8305 becomes 8303, and the cycle can restart. This configuration of reactors is well known as a lead/guard reactor operation cycle.

[1977] Regeneration with temperature constant is described more fully as either: a) temperature between 300 °C and 380 °C, pressure between >0 and 15 psig, or b) temperature between 250 ^C C and 380 °C, pressure between 15 and 100 psig.

[1978] 0 ₂ concentration in the feed gas to the catalyst bed is 8 vol. % or less for a given time period, the balance being inert gas comprising, for example, C0 ₂ , N ₂ , etc. This concentration of 0 ₂ is sufficient to maintain stable regeneration temperatures. Following the regeneration, the feed gas has an 0 ₂ concentration of 0.2 - 35 vol.

%, preferably 2 - 30 vol. %, most preferably 5 - 22 vol. %, the balance being inert gas comprising, for example,

CO _¾ N ₂ , steam, etc. The 0 ₂ concentration can be ramped up, or stepped up, or held constant, until the C0 ₂ concentration in the gas effluent decreased to the C0 ₂ concentration in the regeneration feed gas, or the 0 ₂ concentration in the gas effluent increased to at least 90% of the 0 ₂ concentration in the regeneration feed gas, or both. The coolant for the reactor can be molten salt, steam, oil, or some other cooling means.

[1979] The following examples are merely illustrative and are not intended to be limiting. Unless otherwise indicated, all percentages are by weight.

[1980] EXAMPLE

[1981] A Fixed Bed Reactor Unit (FBRU) was used to conduct regeneration of the ODH catalyst. The apparatus is shown in Figure 37. The apparatus consisted of two fixed bed tabular reactors in series. Each reactor was wrapped in an electrical heating jacket and sealed with ceramic insulating material. Each reactor was SS316L tube which had an outer diameter of 1” and is 34” in length in these experiments, ethane, ethylene, carbon dioxide, oxygen, nitrogen were fed separately (on as-needed basis) and pie-mixed prior to the reactor inlet 3718 with the indicated composition (given in each experiment). The flow passed from the upstream reactor to the downstream reactor at stream 3719, and the product stream exited the downstream reactor at stream 3720. Both reactors were being controlled at the same reaction tempera ture. The temperature of each of the reactors were monitored using corresponding 7-point thermocouples shown by ίί1 P in the upstream reactor, and #8-#14 in the downstream reactor. The highest temperature between thermocouple points was used for controlling the reactor temperature using tire corresponding back pressure regulator that controlled the pressure and boiling temperature of water inside Site desired reactor water jacket 3714. It is noteworthy that only thermocouple points #3 to #6 in Site upstream reactor and #9 to #12 in the downstream reactor were located in the reactor bed, and the reaction temperature for each reactor was being reported as an average of these points.

[1982] The catalyst bed 3715 consisted of one weight unit of catalyst to 2.14 units of weight of Denstone 99

(mainly alpha alumina) powder, total weight of the catalyst in each reactor was 143 g catalyst having the formula Mo Vo.roNbo.ieTeo _. 14O, with relative atomic amounts of each component, relative to a relative amount of Mo of 1, shown in subscript. The rest of the reactor, below' and above the catalyst bed was packed with quartz powder 3716 and secured in place with glass wool 3717 on the top and the bottom of the reactor tube to avoid any bed movement during die experimental runs. For experimental runs the reaction pressure was ~1 bar with flow through the reactor having a weight hourly space velocity (WHSV) of 1.02 Ir ¹ .

[1983] Once the experiments were completed, the catalyst was regenerated using various techniques. The results of the regeneration examples are presented in Table Wl. Regen #1 resulted in an unsuccessful regeneration process treatment due to not diluting the regeneration air at the start-up of this process. Regens #2 - #5 all showed an increase in ethane conversion post-regeneration treatment compared to pre-regeneration treatment. Ethane-to- ethylene selectivity remained unchanged implying an increase in catalyst activity towards ethylene formation via the ODH reaction.

[1984] TABLE W 1 : ODH Catalyst Regeneration Examples

[1985] Used QDH catalyst may be unloaded from an ODH reactor for disposal. During the unloading and handling of the catalyst, the catalyst may be exposed to the ambient environment. Certain embodiments may perform an elevated-temperature air treatment (e.g., 150 °C or greater) on the used ODH catalyst in the ODH reactor prior to unloading the catalyst from the reactor. The air treatment of the catalyst in the ODH reactor may promote catalyst deactivation and displacement (or reaction) of hydrocarbon associated with the catalyst before unloading the catalyst from the ODH reactor.

[1986] In one aspect, a method is provided to convert a lower alkane to an alkene. More specifically, an input stream comprising oxygen and the lower alkane is provided to an oxidative dehydrogenation (ODH) reactor. At least a portion of the lower alkane is converted to the alkene in the ODH reactor and an ODH outlet stream comprising the alkene, an oxygenate, and a carbon-based oxide is produced. The ODH outlet stream is cooled and at least a portion of the oxygenate is condensed. The alkene is separated from the oxygenate to produce an alkene outlet stream and an oxygenate outlet stream. The alkene outlet stream comprises at least a substantial portion of the alkene and at least a substantial portion of the carbon-based oxide. The oxygenate outlet stream comprises at least a substantial portion of the condensed oxygenate.

[1987] In another aspect, an apparatus is provided for oxidative dehydrogenation (ODH) of a lower alkane to an alkene. More specifically, the apparatus comprises an ODH reactor, a means for cooling, and a flash dram. The ODH reactor comprises an ODH inlet and an ODH outlet. The ODH inlet is suitable for transporting an ODH inlet stream comprising the lower alkane, oygen, and, optionally, an inert diluent, such as steam, carbon dioxide, nitrogen, argon, or helium, among others, into the ODH reactor. The ODH outlet is suitable for transporting an ODH outlet stream comprising the alkene, unconverted alkane, an oxygenate, water in the form of steam, and a carbon- based oxide. The means for cooling is suitable for cooling the ODH outlet stream and condensing at least a portion of the oxygenate. The flash dram comprises a drum inlet, an oxygenate outlet, and an alkene outlet. The drum inlet is in fluid communication with the ODH outlet, to receive the cooled ODH outlet stream. The flash drum is suitable for separating the condensed oxygenate from gaseous alkene and gaseous carbon-based oxide. The alkene outlet is suitable for transporting an alkene outlet stream comprising at least a substantial portion of the alkene and at least a substantial portion of the catbon-based oxide. The oxygenate outlet is suitable for transporting an oxygenate outlet stream comprising at least a substantial portion of foe condensed oxygenate. f 1988] In another aspect, a system is provided for oxidative dehydrogenation (ODH) of a lower alkane. More specifically, the system comprises an ODH reactor, a means for cooling, and a flash drum. The ODH reactor is configured to receive an input stream comprising oxygen and the lower alkane. The ODH reactor is configured to produce an ODH outlet stream comprising an alkene, unconverted alkane, an oxygenate, water in the form of steam, and a carbon-based oxide. The means for cooling is configured to cool the ODH outlet stream to produce a cooled ODH outlet stream. The flash dram is configured to separate the alkene from the oxygenate to produce an alkene outlet stream and an oxygenate outlet stream. The alkene outlet stream comprises at least a substantial portion of the alkene and at least a substantial portion of the carbon-based oxide. The oxygenate outlet stream comprises at least a substantial portion of the condensed oxygenate.

[1989] A process, a system, and an apparatus are provided for converting a lower alkane to an alkene. Oxygen and the lower alkane are provided to an ODH reactor to convert at least a portion of the lower alkane to an alkene. An ODH stream comprising the alkene, an oxygenate, steam, and a carbon-based oxide is produced. The bulk of the oxygenate is removed from the ODH outlet stream by non-dilutive cooling, with residual oxygenate being removed using dilutive quenching with a carbonate. Subsequently, separation of the carbon-based oxide from the alkene is achieved using a caustic tower, which also produces spent caustic in the form of a carbonate, which is then used as the carbonate for dilutive quenching. Dilutive quenching using a carbonate allows conversion of the oxygenate to an acetate which can then be used to simplify separation of the oxygenate from water.

[1990] In some examples, techniques are provided for separation of an ODH product from a process stream. A product of an ODH reaction can be an oxygenate such as, for example, ethanol, acetic acid, acrylic acid, maleic acid, and maleic anhydride. The oxygenate can require purification and/or further processing in order to generate a marketable product. For example, water may have to be removed from the oxygenate. Separation of the oxygenate from water can increase the complexity _' of a quench tower and/or a separation tower due to the small thermal (e.g., boiling point) separation between the oxygenate and the water. In various examples, a mixture of oxygenate and water can be azeotropic. The separation tower may employ a large column a high quantity of stages, a high reflux ratio, and a high energy demand to separate an azeotropic mixture of oxygenate and water.

[1991] Thus, a method, a system, and an apparatus are provided which can enhance the purification of the oxygenate and reduce energy requirements for the purification. More specifically, a method, a system, and an apparatus are provided for converting a lower alkane to an alkene. An input stream comprising oxygen and the lower alkane can be provided to an ODH reactor. At least a portion of the lower alkane can be converted to the alkene in the ODH reactor and an ODH outlet stream comprising the alkene, unconverted alkane, an oxygenate, water in the form of steam, and a carbon-based oxide can be produced. The ODH outlet stream can then be cooled to promote condensation of at least a substantial portion of the oxygenate and a portion of the steam. The ODH outlet stream can then be subjected to a means for liquid-gas separation to produce a first oxygenate outlet stream comprising at least a substantial portion of the condensed oxygenate and water and an alkene outlet stream comprising at least a substantial portion of the alkane at least a substantial portion of the carbon-based oxide and any remaining oxygenate. Using condensation and liquid-gas separation for removing a substantial portion of the oxygenate from the ODH outlet stream is non-dilutive as no additional components are added to the ODH outlet stream. f 1992] The alkene outlet stream can be provided to a quench tower and remaining oxygenate can be removed from the alkene outlet stream. The quench tower, as known to a person skilled in the art, includes addition of a quench agent, usually water, and is therefore dilutive as an additional component is added to the stream. A first quench outlet stream comprising at least a substantial portion of the alkene, unconverted alkane, inert diluent, and at least a substantial portion of the carbon-based oxide can be produced in the quench tower. Additionally, a second quench outlet stream comprising at least a substantial portion of the remaining oxygenate can be produced in the quench tower. The first uench outlet stream can be provided to a caustic wash tower. The first quench outlet stream can be contacted with a hydroxide in the caustic wash tower to form a caustic outlet stream comprising a carbonate. The caustic outlet stream can be provided to the quench tower. The alkene outlet stream can be contacted with the caustic outlet stream to form an acetate. The second quench outlet stream can comprise a substantial portion of the acetate.

[1993] Referring to FIG. 84, illustrated is a flow diagram of a non-limiting example of a system 8400 to convert an alkane to an alkene. As illustrated, an ODH reactor 8402. a flash tower 8403, a quench tower 8404, and a caustic wash tower 8406 can be in operative communication. For example, an ODH outlet 8402b of the ODH reactor 8402 can be in fluid communication with a flash tower inlet 8403a of flash tower 8403 via an ODH outlet line 8410. Additionally, a flash tower outlet 8403b of flash tower 8403 can be in fluid communication with a quench inlet 8404a of the quench tower 8404 via alkene outlet line 8410a. Additionally, a quench outlet 8404c of the quench tower 8404 can be in fluid communication with a wash inlet 8406a of the caustic wash tower 8406 via a quench outlet line 8414. Accordingly, the ODH reactor 8402 can be in fluid communication with the caustic wash tower 8406 via the flash tower 8403 and the quench tower 8404.

[1994] The ODH reactor 8402 can comprise an ODH inlet 8402a which can be configured to receive an ODH inlet stream from an ODH inlet line 8408 and can be suitable to transport the ODH inlet stream into the ODH reactor 8402. The ODH inlet stream can comprise a gaseous mixture of a lower alkane and oxygen. In various examples, the ODH inlet stream additionally can include at least one of a carbon-based oxide steam, and an inert diluent. The inert diluent can comprise, for example, nitrogen, carbon dioxide, steam, argon, helium, and methane. In various examples, the carbon-based oxide can comprise at least one of carbon dioxide and carbon monoxide. The concentration of the oxygen and the lower alkane within the mixture in the ODH inlet stream and the temperature and pressure of the ODH inlet stream can be adjusted such that the mixture can be outside of the flammability limits of the mixture.

[1995] In various examples, there may be multiple ODH inlet lines configured to provide the ODH inlet stream to the ODH reactor 8402. For example, each component {e.g., lower alkane, oxygen, steam, carbon-based oxide, and inert diluent) may be added directly to the ODH reactor 8402, each in separate inlet lines (not shown). Alternatively, one or more components may be pie-mixed and added in more than one inlet line. In various example, components may be mixed together prior to the ODH reactor 8402 and subsequently introduced into the ODH reactor in a common ODH inlet. In various examples, steam may be added indirectly as w'ater mixed with an additional reactant and the resulting mixture can be preheated before entering tire ODH reactor 8402. When adding steam indirectly as water, the preheating process can increase the temperature of the mixture so that the water can be substantially converted to steam before entering the GDH reactor 8402.

[1996] The ODH reactor 8402 can include a catalyst capable of catalyzing the ODH of the reactants within the ODH inlet stream to products such as, for example, an alkene, a carbon-based oxide, water, and an oxygenate. The catalyst may be, for example, a mixed metal oxide catalyst.

[1997] The catalyst composition, temperature and pressure of the ODH reactor 8402, and the composition of the ODH inlet stream can be adjusted in order to vaiy the composition of products as known by one of ordinary skill in the art. For example, the ratio of the lower alkane to oxygen can be outside of the upper flammability limit of the mixture. In various examples, the oxygen concentration in the ODH inlet stream can be in a range of 0.1 % to 30 % by volume of the ODH inlet stream, and in some examples range from 0.1% to less titan 30 % by volume less titan 25 % by volume or less titan 20 % by volume. In various examples, the lower alkane concentration in the ODH inlet stream can range from 0.1% to 80 % by volume of the ODH inlet stream, and in some examples range from 0.1% to less Ilian 50 % by volume or less than 40 % by volume.

[1998] In various examples increasing the steam concentration in the ODH inlet stream can increase the amount of oxy genate produced relative to the alkene produced in the ODH reactor 8402. In various examples, reducing the steam concentration in the ODH inlet stream can decrease the amount of oxy genate produced relative to the alkene produced in the ODH reactor 8402. The concentration of steam in the ODH inlet stream can be in a range of 0.1 % to 90 % by volume of the total ODH inlet stream 8408, and in some examples range from 0.1% to less than 40 % by volume, or less than 25 % by volume. In various examples, the concentration of the stream in the ODH inlet stream can be at least 1 % by volume. In various examples, the ODH inlet stream can comprise 20 % oxygen by volume, 40 % lower alkane by volume, and the balance being steam, carbon dioxide, and/or an inert diluent. In some examples, the ODH inlet stream can comprise 10 % oxygen by volume, 15% lower alkane by volume, and the balance being steam, carbon dioxide, and/or an inert diluent.

[1999] In various examples, the ODH process lias a selectivity for the corresponding alkene (e.g., ethylene in the case of ethane ODH) of greater than 95% such as, for example, greater than 98%. The gas hourly space velocity (GHSV) within the ODH reactor 8402 can be from 500 to 30000 h ^! and in some examples the GHSV within the ODH reactor 8402 can be greater than 1000 h ¹ . In various examples, the linear velocity' within the ODH reactor 8402 can be from 10 cin/s to 500 cm/s. In various examples, the weight hourly space velocity' (WHSV) within the ODH reactor can be from 2.1 to 25 h ^’1 . In various examples, the space-time yield of corresponding alkene (e.g., productivity) in grams(g)/hour per kilogram (kg) of the catalyst can be at least 100 such as, for example, greater than 1000, greater than 2000, or greater than 3500, at an ODH reactor temperature of, for example, 350 °C to 450 °C. In various examples, the productivity of the catalyst can increase with increasing temperature in the ODH reactor 8402 until the selectivity of the alkene decreases.

[2000] An ODH reactor can be used for performing an ODH reaction consistent with the provided techniques. In various examples, the reaction can be conducted at temperatures in a range of 250 °C to 450 ^C C such as, for example. 300 °C to 370 °C, 300 °C to 350 °C, or 300 ^C C to 325 °C. In various examples the reaction can be conducted at pressures in a range of 0.5 pounds per square inch (psig) to 100 psig (3.447 to 689.47 kPag) such as, for example, 15 psig to 50 psig (103.4 to 344.73 kPag). In various examples, the lower alkane can have a residence time in the ODH reactor 8402 in a range of 0.002 seconds (s) to 30 s, or from 1 s to 10 s.

[2001] The products of the ODH reaction can leave the ODH reactor 8402 through the ODH outlet 8402b in an ODH outlet stream. The ODH outlet 8402b can be configured to receive the ODH outlet stream and can be suitable to transport the ODH outlet stream 8410 out of the ODH reactor 8402 and into the ODH outlet line 8410. In various examples, in addition to the products, the ODH outlet stream can include unreacted components from the ODH inlet stream such as, for example, lower alkane, carbon-based oxide, oxygen, steam, inert diluent, and combinations thereof. In various examples, the temperature of die ODH outlet stream can be in a range of 100 °C to 450 °C, such as for example. 300 °C to 425 °C, and in certain examples 330 °C to 400 °C.

[2002] Any of the known reactor types applicable for the ODH of an alkane may be used with the provided techniques. For example, a fixed bed reactor, a fluidized bed reactor, or combinations thereof can be used for the ODH reactor 8402. In a typical fixed bed reactor, reactants are introduced into the reactor at an inlet and flow past an immobilized catalyst. Products are formed and leave through Site outlet of the reactor. A person skilled in the art would understand which features are required with respect to shape and dimensions of the reactor, inputs for reactants, outputs for products, temperature and pressure control, and means for immobilizing the catalyst.

[2003] In a typical fluidized bed reactor, the catalyst bed can be supported by a porous structure or a distributor plate and located near a lower end of the reactor. Reactants flow through the fluidized bed reactor at a velocity sufficient to fluidize the bed (e.g. , the cataly st rises and begins to swirl around in a fluidized manner). The reactants can be converted to products upon contact with the fluidized catalyst and the reactants are subsequently removed from an upper end of the reactor. A person of ordinary skill in the art would understand which features are required with respect to shape and dimensions of the reactor, the shape and size of the distributor plate, the input temperature, the output temperature, the reactor temperature and pressure, inputs for reactors, outputs for reactants, and velocities to achieve fluidization.

[2004] In various examples, there may be multiple ODH reactors connected in series or in parallel. Each ODH reactor may be the same or different. For example, each ODH reactor can contain the same or different ODH catalyst. In various examples, the multiple ODH reactors can each be a fixed bed reactor, can each be a fluidized bed reactor, or the multiple ODH reactors can be combina tions of fixed bed reactors and fluidized bed reactors. [2005] Regardless of the configuration of the ODH reactor 8402, the ODH outlet 8402b can be in fluid communication with the flash tower inlet 8403a of the flash tower 8403 via ODH outlet line 8410 to direct the ODH outlet stream to the flash tower 8403. The ODH outlet stream is subjected to a cooling means 8405 prior to teaching flash tower inlet 8403a or within flash tower 8403. The flash tower outlet 8403b can be in fluid communication with the quench inlet 8404a of the quench tower 8404 via the aikene outlet line 8410a to direct the alkene outlet stream to the quench tower 8404. The quench inlet 8404a can be configured to receive the aikene outlet stream from the alkene outlet line 8410a and can be suitable to transport the aikene outlet stream into the quench tower 8404.

[2006] The cooling means 8405 can be any means that cools the ODH outlet stream after it leaves the ODH reactor. This can include using a sufficiently long ODH outlet line 8410 that allows file ODH outlet stream to cool to a temperature where the oxygenate begins to condense before reaching flash tower 8403. In some embodiments, the most preferable cooling means 8405 comprise a heat exchanger, use of which is well known within the art. In some embodiments, the cooling means 8405 are an integral part of the Hash tower 8403. In some examples, the flash tower may be surrounded by a cooling jacket that cools the ODH outlet stream as it enters flash tower 8403. in another example, cooling tubes are arranged within the space inside flash tower 8403. In some embodiments, a heat exchanger in combination with integrated cooling means are used to cool to ODH outlet stream.

[2007] The cooling means can cool the ODH outlet stream to a temperature of less than 200 °C such as, for example, less than 100 °C, less than 50 °C, less than 40 °C, and in some examples, the cooling means 8405 can cool the ODH outlet stream to a temperature of 20 °C to 80 °C. In various examples, the cooling means 8405 can cool the ODH outlet stream to a temperature which induces condensation of the oxygenate such as. for example, to a temperature less than or equal to tire boiling point of Site oxygenate and/or a temperature that reduces the vapor pressure of the oxygenate. The lower the temperature, without going below a temperature that results in freezing of the water or oxygenate, tire greater the degree of condensation, which would be understood by a person skilled in the art.

[2008] Flash tower 8403 can comprise a flash tower, or any other means that provides for gas-liquid separation. Use of flash towers is well known. At least a substantial po rtion of the oxygenate and water, in the form of steam, within the ODH outlet stream may be in a liquid state after being subjected to cooling means 105 and may exit flash tower 8403 through a first oxygenate outlet 8403c, as a first oxygenate outlet stream, and into the first oxygenate outlet line 8411. In various examples, the first oxygenate outlet stream can comprise at least 0.5 mol% oxygenate, such as, for example, at least 2.0 mol% oxygenate, at least 5 mol%, or 0.5 mol% to 15 mol% oxygenate. The first oxygenate outlet stream can additionally comprise water of from 80 mol% to 99.5 mol%.

[2009] A portion of the oxygenate and water within the ODH outlet stream, including liquid and gaseous forms, may leave the flash tower 8403 as part of the alkane outlet stream. These portions are referred to as remaining oxygenate and remaining water, respectively.

[2010] In various examples, the alkene outlet stream may be subject to a second cooling means prior to or as an integral part of quench tower 8404. The second cooling means can be configured to adjust the temperature of the alkene outlet stream, for example, by cooling to a temperature of less than 200 °C such as, for example, less than 100 °C, less than 50 °C, less than 40 °C, and in some examples, the second cooling means can cool the alkene outlet stream to a temperature of 20 °C to 80 °C. In various examples, the second cooling means can cool the alkene outlet stream to a tempera ture which induces condensa tion of the remaining oxygenate such as, for example, a temperature less titan or equal to the boiling point of the remaining oxygenate and/or a temperature that reduces the vapor pressure of the remaining oxygenate. The second cooling means can use any means known in the art. For example, the second cooling means can be a standalone heat exchanger separate from a quench tower in various examples,

Site second cooling means can be an integrated heat exchanger that is part of a quench tower. In further examples the second cooling means may include a combination of standalone heat exchanger and an integrated heat exchanger.

[2011] The quench tower 8404 can comprise a quench tower, an oxygenate scrubber the like, or combinations thereof. The quench tower 8404 can be configured to quench the components in the alkene outlet stream and remove at least a substantial portion of the oxygenates from the aikene outlet stream. In various examples, the quench tower 8404 can facilitate the removal of remaining oxygenate and water from the aikene outlet stream. The quench tower 8404 can produce a first quench outle t stream comprising a t least a substantial portion of the aikene and at least a substantial portion of the carbon-based oxide from the aikene outlet stream. In various examples, the first quench outlet stream can comprise additional components from the aikene outlet stream such as, for example, a portion of the oxygen, a portion of the oxygenate, a portion of the inert diluent, a portion of the steam, and a portion of the unreacted alkane. The first quench outlet stream exits the quench tower 8404 through the quench outlet 8404c. The quench outlet 8404c can be configured to receive the first quench outlet stream and can be suitable to transport the first quench outlet stream out of the quench tower 8404 into the quench outlet line 8414.

[2012] The quench tower 8404 can produce a second quench outlet stream comprising at least a substantial portion of any remaining oxygenate present in the aikene outlet stream and in some examples, an acetate as discussed herein. In various examples, the second quench outlet stream can comprise additional components from the aikene outlet stream such as, for example, a substantial portion of tire remaining water (e.g., steam), as well as lower alkane, aikene, oxygen, and carbon-based oxide. The second quench outlet stream can exit the quench tower 8404 through second quench outlet 8404b of the quench tower 8404. The second quench outlet 8404b can be configured to receive the second quench outlet stream and can be suitable to transport the second quench outlet stream out of the quench tower 8404 into the second quench outlet line 8412.

[2013] The quench tower 8404 also comprises a carbonate inlet 104d for providing a quenching agent such as water via quench agent line 8418a. ora carbonate solution via return line 8418. Use of quench towers is well known. A person skilled in the art would understand that quenching condenses and dilutes the oxygenate. The result is that seco nd quench outlet stream compri ses a lower moi% of the oxygenate compared to the first oxygenate outlet stream. When using a carbonate solution as the quench agent the effect is more pronounced as a portion of the oxygenate is converted to an acetate as wall be desesibed

[2014] In some examples, the quench agent is provided to the quench tower at a temperature of less than 200 °C such as, for example, less than 100 °C, less than 50 °C, less than 40 °C, and in some examples the quench agent is provided to the quench tower at a temperature of 20 °C to 80 °C. In various examples, the quench agent can be provided to the quench to a temperature which induces condensa tion of the remaining oxygenate such as, for example, a temperature less than or equal to the boiling point of the remaining oxygenate and/or a temperature that reduces the vapor pressure of the remaining oxygenate.

[2015] The quench outlet 8404c can be in fluid communication with the wash inlet 8406a of the caustic wash tower 8406 via the quench outlet line 8414 to direct the first quench outlet stream to the caustic wash tower 8406. The wash inlet 8406a can be configured to receive the first quench outlet stream from the quench outlet line 8414 and can be suitable to transport the first quench outlet stream into the caustic wash tower 8406.

[2016] The caustic wash tower 8406 can comprise the wash inlet 8406a, a wash outlet 8406c, a caustic inlet 84Q6d, and a caustic outlet 8406b. The caustic inlet 8406d can be configured to receive a hydroxide stream comprising a hydroxide from a hydroxide line 8420 and can be suitable to transport the hydroxide stream into the caustic wash tower 8406. The hydroxide may be, for example, an aqueous solution of at least one of sodium hydroxide, potassium hydroxide, and ammonia hydroxide. In various examples, the aqueous solution comprises at least 0.5 mol% hydroxide, such as, for example, at least 1.0 mol% hydroxide, at least 1.25 mol%, or 0.5 mol% to 1.75 mol% hydroxide.

[2017] The caustic wash tower 8406 can be configured to contact the hydroxide stream with the first quench outlet stream. In various examples, where the carbon-based oxide comprises carbon dioxide, the hydroxide can react with carbon dioxide in the first quench outlet stream to form a carbonate. The reaction can remove at least a substantial portion of the carbon-based oxide (e.g., carbon dioxide) from the first quench outlet stream and produce a wash outlet stream and a caustic outlet stream. The carbonate may be, for example, at least one of sodium bicarbonate potassium carbonate, and ammonium bicarbonate. For example, fire reaction of sodium hydroxide and carbon dioxide is shown in Scheme Yl.

Scheme Yl

C0 ₂ + NaOH ® NaHC0 ₃

[2018] The wash outlet stream can comprise unreacted components from the first quench outlet stream. The wash outlet 8406c can be configured to receive the wash outlet stream and can be suitable to transport the wash outlet stream out of the caustic wash tower 8406 into the wash outlet line 8416.

[2019] The caustic outlet stream can comprise a substantial portion of the carbonate and in some examples, at least one of water, hydroxide, and oxygenate. The caustic outlet 8406b can be configured to receive the caustic outlet stream and can be suitable to transport the caustic outlet stream into tire return line 8418. The return line 8418 can be configured to receive the caustic outlet stream and output tire caustic outlet stream into a carbonate inlet 84Q4d of tire quench tower 8404.

[2020] In various examples, the caustic outlet stream can comprise at least 0.5 mol% carbonate, such as, for example, at least 2.0 mol% carbonate, at least 5 mol%, or 0.5 rnol% to 15 mol% carbonate. The caustic outlet stream can additionally comprise water of from 80 mol% to 99.5 mol%.

[2021] The quench tower 8404 can be configured to contact the caustic outlet stream with the alkene outlet stream. In various examples the quench tower 8404 can be configured to react the caustic outlet stream with the alkene outlet stream to form an acetate. In various examples, the quench tower 8404 can react the carbonate with the oxygenate and in some examples, with water and hydroxide, to form the acetate. The acetate can comprise at least one of sodium acetate, potassium acetate, and ammonium acetate. As an example, the reaction of sodium bicarbonate and the oxygenate to form sodium acetate is illustrated by the reaction in Scheme Y2.

Scheme Y2

NaHC0 ₃ + CH _j COOH <® C0 ₂ + H ₂ 0 + NaC ₂ H ₃ 0 ₂

[2022] In various examples, the mole ratio of the carbonate in the caustic outlet stream to oxygenate in the alkene outlet stream can be in a range of 0.8:1 to 1.2:1 such as for example, 1:1. In various examples, the mole ratio of the carbonate in the caustic outlet stream to oxygenate in the alkene outlet stream can be greater than 1:1 such as, for example, 2:1.

[2023] In various examples, the quench tower 8404 can be configured to maintain a pH in a range of 2 to 12 such as, for example, 4 to 7. In various examples, the quench tower 8404 can be configured to maintain a pH in a range of a pKa of the oxygenate to a pKa of the carbonate in order to facilitate the formation of the acetate. In various example, the oxygenate comprises acetic acid having a pKa of 4.7 and sodium bicarbonate having a pKa of 6.4.

[2024] The first quench outlet stream can comprise a substantial portion of the carbon-based oxide produced in the quench tower 8404 and from the alkene outlet stream. The second quench outlet stream can comprise the oxy genate, the acetate, and water. Adding the caustic outlet stream to the quench tower can decrease the amount of oxy genate and increase the amount of acetate in the first quench outlet stream. The decrease in oxygenate in the first quench outlet stream can be a result of the conversion of the oxygenate to the acetate. The conversion of die oxygenate to the acetate can facilitate die removal of the oxygenate from the alkene outlet stream and limit the oxygenate from exiting the quench tower 8404 in the first quench outlet stream.

[2025] In various examples, the second quench outlet stream can comprise at least 0.1 mol% ox genate, such as. for example, at least 0.5 mol% oxygenate, at least 1 mol%, or 0.1 mol% to 5 mol% oxygenate. The second quench outlet stream can additionally comprise at least 0.25 mol% acetate such as, for example, at least 1.75 mol% oxygenate, at least 5 moi%, or 0.25 mol% to 15 mol% acetate water of from 80 mol% to 99.5 mol%.

[2026] In various examples, the oxygenate in the first oxygenate outlet stream and the second quench outlet stream may be subject to further processing. For example, referring to FIG. 85, the oxy genate can be separated from the acetate in a separation tower 8526. FIG. 85 is a flow diagram of a non-limiting example of a system 8500 comprising the separation tower 8526. As illustrated, the separation tower 8526 has a separation inlet 8526a, a first separation outlet 8526b, and a second separation outlet 8526c. The separation inlet 8526a can be configured to receive the first oxygenate outiet stream from first oxygenate outlet line 8411 and the second quench outlet stream from the second quench outline line 8412 and may be suitable to transpo rt at least one of the first oxy genate outlet stream and the second quench outiet stream into the separation tower 8526.

[2027] The separation tower 8526 can separate the oxygenate from the acetate and, in various examples, the separation tower 8526 can separate the oxygenate from water. The presence of the acetate in the separation tower 8526 can enhance the separation of oxygenate from the water. For example, the acetate and oxygenate may disassociate and/or react with water to form an acetate ion (e g , ClhCOCT ) and an acid (e.g., H ₃ 0 ⁺ , Na ⁺ ) . Since the acetate and oxygenate can form a common ion, an increase in the concentration of one of the acetate and oxygenate can affect the oilier. For example, tire reactions of sodium acetate (C ₂ H ₃ Na0 ₂ ), acetic acid (CH ₃ COOH), bicarbonate ion (! iCO _;! ). carbon dioxide (C0 ₂ ), and water (H ₂ 0) is illustrated in Scheme Y3.

Scheme Y3

2 H ₂ 0 + C0 ₂ <® HC0 ₃ + H ₃ 0 ⁺

C ₂ H ₃ Na0 ₂ ® ( or ®) CH ₃ COO ^~ + Na +

[2028] As illustrated in Scheme Y3, sodium acetate can form an acetate ion which can affect the equilibrium reaction of acetic acid and water. For example the sodium acetate can cause the equilibrium reaction of acetic acid and water to Slave a higher preference for the separate species of acetic acid and water than an acetate ion and an acid relative to without the presence of acetate.

[2029] The separation tower 8526 can comprise various equipment known to those of ordinary skill in the art. For example, the separation tower 8526 can comprise an extraction tower, a packed column, a sieve-tray column, a spray column, a KARR column, a rotating disc contactor, a stirred cell extractor, a rectification tower, a stripper, and combinations thereof in various examples, the separation tower 8526 can comprise a liquid-liquid extractor. Accordingly, the acetate in the oxy genate inlet stream can increase the efficiency of the separation tower 8526 and can facilitate efficient separation of the oxygenate from water.

[2030] The separation tower 8526 can produce a second separation outlet stream comprising a substantial portion of the oxygenate from at least one of the first oxygenate outlet stream and the second quench outlet stream.

In various examples, the second separation outlet stream can comprise additional components from at least one of the first oxygenate outlet stream and the second quench outlet stream, such as, for example, water. In various examples, the second separation outlet stream can comprise at least 80 mol% oxygenate such as, for example, at least 90 mol% oxygenate, at least 95 mol% oxygenate, or 80 mol% to 100 mol% oxygenate by weight. The second separation outlet stream can exit the separation tower 8526 through the second separation outlet 8526c of the separation tower 8526. The second separation outlet 8526c can be configured to receive the second separation outlet stream and can be suitable to transport the second separation outlet stream out of the separation tower 8526 into the second separation outlet line 8528.

[2031 J The separation tower 8526 can produce a first separation outlet stream comprising a substantial portion of the acetate from the second quench outlet stream and in various examples, a substantial portion of the water from the second quench outlet stream. In various examples, the first separation outlet stream can comprise at least 10 % acetate by weight such as, for example, at least 30 % acetate by weight, at least 50 % acetate by weight, or 30 % to 70 % acetate by weight. In various examples, the first separation outlet stream can comprise at least 5 % water by weight such as, for example, at least 10 % water by weight, at least 25 % water by weight, or 15 % to 50 % water by weight. The first separation outlet stream can exit the separation tower 8526 through the first separation outlet 326b of the separation tower 8526. The first separation outlet 8526b can be confi gured to receive the first separation outlet stream and can be suitable to transport the first separa tion outlet stream out of the separation tower 8526 into the first separation outlet line 8530.

[2032] The separation tower 8526 can be configured with a recycle line 8532 in fluid communication with the first separation outlet line 8530 and/or first separation outlet 8526b. The recycle line 8532 can be configured to recycle a portion of the acetate from the first separation outlet stream to the separation tower 8526 via the recycle inlet 8526d. The recycle line 8532 can be configured to receive a portion of the first separation outlet stream and can be suitable to transport a recycle stream to a recycle inlet 8526d of the separation tower 8526. The recycle inlet 8526d can be configured to receive the recycle stream and can be suitable to transport Site recycle stream into the separation tower 8526. For example, the recycle stream can comprise a portion of the acetate from the first separation outlet stream, and in various examples, a portion of the water from the first separation outlet stream. [2033] The recycle line 8532 can be configured to recycle the acetate from the first separation tower outlet stream until a select concentration of acetate is achieved in the separation tower 8526. in various examples and referring to FIGs. 84 and 85, the return line 8418 can enable additional generation of acetate in the quench tower 8404 which would flow to the separation tower 8526 through the second quench outlet line 8412 to increase the concentration of acetate in the separation tower 8516.

[2034] in various examples, a supplemental acetate stream can be added to the separation tower 8526. In various examples, the supplemental acetate can comprise ethyl acetate.

[2035] In various examples, an oxygen remover 8644 can be disposed at any point intermediate the ODH reactor 8402 and the caustic wash tower 8406. FIG. 86 is a flow diagram of a non-limiting embodiment of a system 8600 comprising an oxygen remover 8644 when situated intermediate the flash tower 8403 and the quench tower 8404. As illustrated, the oxygen remover 8644, comprising a remover inlet 8644a and a remover outlet 8644b, can be provided in fluid communication with the flash tower 8403 (FIG. 84) via alkene outlet line 8410a and the quench tower 8404 via remover outlet line 8646. The remover inlet 8644a can be configured to receive the alkene outlet stream and can be suitable to transport the alkene outlet stream into the oxygen remover 8644. The oxygen remover 8644 can remove a substantial portion of the oxygen in the alkene outlet stream and produce a remover outlet stream comprising the alkene outlet stream with the substantial portion of the oxygen removed. The oxygen remover 8644 can be of various designs as kno wn in the art. The remover outlet 8644b can be configured to receive the remover outlet stream and can be suitable to transport the remover outlet stream out of the oxygen remover 8644 into the remover outlet line 8646. The quench inlet 8404a of the quench tower 8404 can be configured to receive the remover outlet stream.

[2036] In another embodiment, the oxygen remover 8644 can be situated downstream the ODH reactor 8402 and upstream the flash tower 8403. In another embodiment, the oxygen remover 8644 can be situated down stream the quench tower 8404 and upstream the caustic tower 8406. The oxygen remover precedes the caustic wash tower 8406, or an amine tower (described below), as residual oxygen within the first quench outlet stream may affect operability ^r of the caustic wash tower, or amine tower. Minimizing the level of oxygen within the first alkene outlet stream may also be achieved by altering ODH reaction conditions, as would be apparent to a person skilled in the art.

[2037] Referring to FIG. 87, in various examples, an amine tower 8748 can be disposed intermediate the quench tower 8404 and the caustic wash tower 8406. FIG. 87 is a flow diagram of a non-limi ting example of a system 8700 comprising an amine tower 8748. As illustrated, the amine tower 8748, comprising an amine tower inlet 8748a and an amine tower outlet 8748b, can be provided in fluid communication with the quench tower 8404 (FIG. 84) via quench outlet line 8414 and the caustic wash tower 8406 via amine tow¾r outlet line 8750. The amine tower inlet 8748a can be configured to receive the first quench outlet stream and can be suitable to transport the first quench outlet stream into the amine tower 8748. The amine tower 8748 can remove a substantial portion of carbon dioxide in the quench outlet stream and produce an amine tower outlet stream comprising Site first quench outlet stream with the substantial portion of Site carbon dioxide removed. The amine tower 8748 can be of various designs as known in tire art. [2038] The amine tower outlet 8748b can be configured to receive the amine tower outlet stream and can be suitable to transport the amine tower outlet stream out of the amine tower 8748 into the amine tower outlet line 8750. The wash inlet 8406a of the caustic wash tower 8406 can be configured to receive the amine tower outlet stream from the amine tower outlet line 8750.

[2039] Having a high efficiency oxy genate removal prior to the amine tower 8748 can limit, and in some examples prevent, amine degradation to presence of the oxygenate in the amine tower 8748. For example, the oxy genate can form heat stable salts with amine in the amine tower 8748 which can degrade the efficiency and shorten die operational life of the amine tower 8748.

[2040] Referring to FIG. 88, in various examples, a polymerization reactor 8852 can be in fluid communication with the caustic wash tower 8406 via the wash outlet line 8416. Various oilier process units may be placed between die caustic wash tower 8406 and the polymerization reactor, including a diyer, a demethanizer, and a C-2 splitter, among others. FIG. 88 is a flow diagram of a non-limiting example of a system 8800 comprising a pol merization reactor 8852. As illustrated, the polymerization reactor 8852, comprising a polymerization inlet 8852a and a polymerization outlet 8852b, can be provided in fluid communication with the caustic wash tower 8406 via the wash outlet line 8416. The polymerization inlet 8852a can be configured to receive the ODH outlet stream and can be sui!able to transport the ODH oudet stream into the poly merization reactor 8852. The poly merization reactor 8852 can produce a polymer from the alkene and produce a polymerization outlet stream comprising the polymer. In various examples, the polymer comprises at least one of polyethylene, polypropylene, and polybutylene. The polymerization reactor 8852 can be of various designs as known in the art. The polymerization outlet 8852b can be configured to receive the polymerization outlet stream and can be suitable to transport the polymerization outlet stream out of the polymerization reactor 8852 into the polymerization outlet line 8854.

[2041] Concentrations of the components within the system can be measured any at point in the process using any means known in the art. For example, a detector such as a gas chromatograph, an infrared spectrometer, and a Raman spectrometer can be disposed downstream or upstream of ODH reactor 8402, quench tower 8404 caustic wash tower 8406, separation tower 8526, oxygen remover 8644, amine tower 8748, and polymerization reactor 8852.

[2042] In various examples, the ODH inlet stream 8408 can comprise mixtures that fall within the flammability limits of the components. For example, the mixture may exist in conditions that prevent propagation of an explosive event. In these examples, the flammable mixture can be created within a medium where ignition can be immediately quenched. In various examples, oxygen and the lower alkanes can be mixed at a point where they are surrounded by a flame arresting material. Thus, any ignition can be quenched by the surrounding material. Flame arresting material includes, for example, metallic or ceramic components, such as stainless steel walls or ceramic supports. In various examples, oxygen and lower alkanes can be mixed at a low temperature, where an ignition event may not lead to an explosion, then the mixture can be introduced into the ODH reactor before increasing the temperature. Therefore, the flammable conditions may not exist until the mixture can be surrounded by the flame arresting material inside of the reactor. [2043] In various examples, the olefins produced using an ODH reactor, or any of the processes or eonipiexes described herein, can be used to make various olefin derivatives utilizing a pol merization reactor. Olefin derivatives include, but are not limited to, polyethylene, polypropylene, ethylene oxide, propylene oxide, polyethylene oxide, polypropylene oxide, vinyl acetate, vinyl chloride, aciylie esters (e.g., methyl methacrylate), thermoplastic elastomers, thermoplastic olefins, blends thereof, and combinations thereof.

[2044] In various examples, ethylene and optionally a-olefins can be produced in an ODH reactor, or any of the processes or complexes described herein, and are used to make polyethylene utilizing a polymerization reactor. The polyethylene made from the ethylene and optional a-olefins described herein can include homopolymers of ethylene, copolymers of ethylene and a-olefins, resulting in HOPE, MDPE, LDPE, LLDPE and VLDPE.

[2045] The polyethylene produced using the ethylene and optional a-olefins described herein can be produced using any suitable pol merization process and equipment. Suitable ethylene polymerization processes include, but are not limited to gas phase polyethylene processes, high pressure polyethylene processes, low pressure polyethylene processes solution polyethylene processes, slurry polyethylene processes and suitable combinations of the above arranged either in parallel or in series.

[2046] A process for converting a lower alkane to an alkene can include providing an input stream comprising oxygen and the lower alkane to an ODH reactor 8402. In various examples, providing the input stream to an ODH includes providing oxygen and the lo wer alkane via separate streams. At least a portion of the lower alkane can be converted to the alkene in the ODH reactor 102. In various examples, the alkane can comprise ethane and the alkene comprises ethylene. In various examples, the alkane can comprise propane and the alkene comprises propylene. In various examples, the alkane comprises butane and the alkene can comprise butylene. An ODH outlet stream comprising the alkene, water in the form of steam, an oxygenate, and a carbon-based oxide may be produced. In various examples, the ODH outlet stream can comprise at least one of an unreacted alkane and oxygen.

[2047] The ODH outlet stream is cooled to allow condensation of a portion of the oxygenate and a portion of the steam. The cooled ODH outlet stream is provided to a flash tower 8403, or other means for gas-liquid separation, and at least a substantial portion of the condensed oxygenate and at least a substantial portion of the condensed steam are removed from the ODH outlet stream to produce an alkene outlet stream comprising at least a substantial portion of the alkene and at least a substantial portion of the carbon-based oxide and a first oxygenate outlet stream comprising at least a substantial portion of the condensed oxygenate and at least a substantial portion of the condensed steam In various examples, the alkene outlet stream can comprise at least one of oxygenate, steam, unreacted alkane, and oxygen. Oxygenate and steam present in the alkene outlet stream, referred to as remaining oxygenate and remaining steam, comprise the portions of the oxygenate and steam within the ODH outlet stream that fail to condense during cooling, or are carried by the gaseous alkene and gaseous carbon-based oxide as the alkene outlet stream exits the flash tower 8403.

[2048] The alkene outlet stream can be provided to a quench tower 8404 and the remaining oxygenate and the remaining steam can be removed from the alkene outlet stream in the quench tower 8404 to produce a first quench outlet stream comprising at least a substantial portion of the alkene and at least a substantial portion of the carbon- based oxide. Additionally, the quench tower 8404 can produce a second quench outlet stream comprising at least a substantial portion of the remaining oxygenate and at least a portion of the remaining steam. Quench towers typically involve the quenching of a gaseous stream with water, or other quench agent, to promote condensation of components within the gaseous stream. The condensed components along with the quench agent fall to the bottom of the tower where they can be removed. The gaseous components rise and can be removed from a location near the top end of the quench tower. The temperature of the quench agent is ideally below that of the condensation point of the component that is targeted for removal. For oxygenates, such as acetic acid, the temperature of the quench agent is, for example, between 20°C and 100°C.

[20491 In various examples, at least one of the ODH outlet stream and the alkene outlet stream can be provided to an oxygen remover 8644 prior to the quench tower 8404. Oxygen can be removed from at least one of the ODH outlet stream and the alkene outlet stream in the oxygen remover 8644 and to reduce the levels of oxygen within at least one of the ODH outlet stream and the alkene outlet stream to from 0 to 5 parts per million (ppm). [2050] The first quench outlet stream can be provided to a caustic wash tower 8406. The quench outlet stream can be contacted with a hydroxide to form a caustic outlet stream comprising a carbonate. In various examples, the first quench outlet stream is contacted with the hydroxide in the caustic wvash tower 8406.

[2051] In various examples, the first quench outlet stream can be provided to an amine wash tower 8748 prior to the caustic wvaste tower 8406. A substantial portion of the carbon-based oxide can be removed from the first quench outlet stream. The first quench outlet stream with the substantial portion of the carbon-based oxide removed can be provided to the caustic waste tower 8406. Use of amines such as diethanolamine, monoethanolamine, methyldiethanolamine, is well known for treating gases to remove carbon-based oxides.

[2052] The caustic outlet stream can be provided the quench tower 8404 where it acts as a quench agent, and the alkene outlet stream can be contacted with the caustic outlet stream to form an acetate. In various examples, the alkene outlet stream is contacted with the caustic outlet stream in the quench tower 8404. The second quench outlet stream outlet stream can comprise a substantial portion of the acetate. In various examples the pH of the quench tower 8404 can be maintained in a range of 2 to 12 such as, for example, 4 to 7. In various examples, the pH of the quench tower 8404 can be maintained in a range of a pKa of the oxygenate to a pKa of the carbonate.

[2053] In various examples, the second quench outlet stream can be provided to a separation tower 8526. The oxygenate can be separated from the acetate within the second quench outlet stream. A second oxygenate outlet stream comprising a substantial portion of the oxygenate from the second quench outlet stream can be produced. An extraction outlet stream comprising a substantial portion of the acetate from the oxygenate stream can be produced. In various examples, a portion of the acetate from the extraction outlet stream can be recycled to the separation tower 8526. In various examples, a supplemental acetate can be provided to the separation tower 8526 such as, for example, ethyl acetate. In various examples, the first oxygenate outlet stream can be provided to the separation tower 8526.

[2054] In various examples, olefin derivatives can be produced from the alkene.

[2055] Techniques are provided for an alternative use for the caustic waste stream which limits, and in sortie examples, can eliminate a need to dispose of the caustic waste stream. Additionally, the reuse of the caustic waste stream can provide a useful product of acetate which can aid in oxygenate separation from the quench outlet stream and purification of oxygenate in the separation tower. The efficient removal of the oxygenate from the quench outlet stream can length the operational light of downstream equipment such as protecting the amine tower against fouling and amine solution degrada tion. Moreover, the acetate can be sold. Furthermore, the efficient purification of the oxy genate can create a marketable product such as, for example, glacial acetic acid.

[2056] EXAMPLES

[2057] Computational modeling of a liquid-liquid separation vessel using equations from Scheme Y3 using ASPEN Plus® version 8.6 chemical process simulation software, commercially available from Aspen Technology', Inc. Bedford, Massachusetts, was used to demonstrate the effect of altering the composition of the feed, as mass fraction, to the separation tower, the feed representing the second quench outlet stream, on the composition of the second oxygenate outlet stream and the composition of the extraction outlet stream. The feed components chosen are typical for an ODH process of ethane, and include water and acetic acid, with trace amounts (not shown) of ethane, ethylene, carbon dioxide. Use of acetate in the quench tower results in sodium acetate contributing to the feed composition.

[2058] Example Y1

[2059] Example Y1 represents a feed composition that corresponds to a second quench outlet stream where the ODH outlet stream was not cooled or subjected to non-dilutive separation prior to quenching in the presence of sodium acetate. The feed was modeled at a total mass flow rate of 6980 kg/hr and at a pressure of 185.7 kPa gauge. [2060] Example Y2

[20611 Example Y2 represents a feed composition that corresponds to a second quench outlet stream where the ODH outlet stream was not cooled or subjected to non-dilutive separation prior to quenching in the presence of sodium acetate. The feed was modeled at a total mass flow' rate of 55891 kg/hr and at a pressure of 465 kPa gauge. [2062] Example Y3

[2063] Example Y3 represents a feed composition that corresponds to a second quench outlet stream where the ODH outlet stream was cooled and subjected to non-dilutive separation prior to quenching in the presence of sodium acetate. The feed was modeled at a total mass flow' rate of 11259 kg/hr and at a pressure of 450 kPa gauge. [2064] Example Y4

[2065] Example Y4 represents a feed composition that corresponds to a second quench outlet stream where the ODH outlet stream was cooled and subjected to non-dilutive separation prior to quenching in the presence of sodium acetate. The feed was modeled at a total mass flow' rate of 9259 kg/hr and at a pressure of 450 kPa gauge. [2066] As shown in Table Y 1, the separation tower, as aided by the sodium acetate, produced the second oxygenate outlet stream comprising at least 85 % by weight of acetic acid. The separation tower produced an extraction outlet stream comprising at least 99% of the sodium acetate. The sodium acetate can be recycled into the separation tower and/or can be a commercially marketable product. Additionally, the acetic acid in the second separation outlet can be a commercially marketable product. [2067] Table Y1

[2068] The feed streams in the examples take into account the effect of cooling prior to the quench tower. One skilled in the tut would recognize that when the ODH outiet stream is not cooled prior to quenching that a larger quantity of water would be required to quench the steam and acetic acid present in the ODH outlet stream. This explains why in example Y2 a feed composition lower in acetic acid titan examples Y3 and Y4 was chosen, despite the fact that in examples Y3 and Y4 a substantial portion of the acetic acid was removed prior to the quench tower.

A smaller quantity of water was required for quenching in examples Y3 and Y4 resulting in a higher mass fraction for those samples.

[2069] The examples demonstrate that acetic acid produced in an ODH process can be isolated in a more concentrated from using non-dilutive separation, and that the remainder can be captured as a more dilute solution. The dilute solution can be treated to increase the concentration to a marketable level, and this treatment is simplified by adding spent caustic in the form of a carbonate to the quenching step.

[2070] A process, a system, and an apparatus are provided for converting a lower alkane to an alkene. Oxygen and a lower alkane are provided to an ODH Reactor. At least a portion of the lower alkane is converted to an alkene and an ODH stream comprising the alkene, an oxygenate, water, and carbon monoxide is produced. The ODH stream is provided to a w'ater gas shift/hydrogenation (WGS/H) reactor including a WGS/H catalyst. The ODH stream is reacted within the WGS/Ή reactor and hydrogen and carbon dioxide are generated from the carbon monoxide and water. At least a portion of the oxygenate and hydrogen are converted to an alcohol. Additionally, the alcohol may be dehydrated to form additional alkene and water.

[2071] In one aspect, a method for converting a lower alkane to an alkene is provided. More specifically, oxygen and the lower alkane are provided to an oxidative dehydrogenation (ODH) reactor. At least a portion of the lower alkane is converted to the alkene in the ODH reactor. An ODH outlet stream comprising the alkene, an oxygenate, water, and carbon monoxide is produced. At least a portion of the ODH outlet stream is provided to a water gas shift/hydrogenation (WGS/H) reactor including a WGS/H catalyst. Carbon monoxide and water present in the ODH outlet stream reacts in the presence of the WGS/H catalyst, to form carbon dioxide and hydrogen. At least a portion of the oxygenate present in the ODH outlet stream reacts with die hydrogen formed to form an alcohol. An alcohol outlet stream comprising at least a substantial portion of the alcohol is produced. In another aspect, at ieast a portion of Site alcohol is dehydrated to form an alkene and water. [2072] In another aspect, an apparatus for ODH of a lower alkane to an aikene is provided. More specifically, the apparatus comprises an ODH reactor and a WGS/H reactor. The ODH reactor comprises an ODH inlet and an ODH outlet. The ODH inlet is suitable for transporting an ODH inlet stream comprising the lower alkane and oxy gen into the ODH reactor. The ODH outlet is suitable for transporting an ODH outlet stream comprising the aikene, an oxygenate, water, and carbon monoxide. The WGS/H reactor comprises a WGS/H inlet, a WGS/Ή outlet, and a WGS/H catalyst. The WGS/H inlet is in fluid communication with the ODH outlet to receive the ODH outlet stream. The WGS H reactor is suitable to generate hydrogen and carbon dioxide from the carbon monoxide and water of the ODH outlet stream. The WGS/H outlet is suitable for transporting an alcohol outlet stream comprising an alcohol.

[2073] In another aspect, a system for ODH of a lower alkane to an aikene is provided. More specifically, Site system comprises an ODH reactor and a WGS/H reactor. The ODH reactor is configured to receive oxygen and the lower alkane. The ODH reactor is configured to produce an ODH outlet stream comprising an aikene, an oxygenate, and a carbon-based oxide. The WGS/H reactor comprises a catalyst. The WGS/H reactor is configured to receive the ODH outlet stream and to generate hydrogen and carbon dioxide from the carbon monoxide and water of the ODH outlet stream. At least a portion of the oxygenate and hydrogen is converted to an alcohol.

[2074] In some examples, techniques are provided for catalytic conversion of an ODH product to an alcohol.

A product of an ODH reaction can be an oxygenate such as, for example, acetic acid, acrylic acid, maleic acid, and maleic anhy dride. The oxygenate can require purification and/or further processing in order to generate a marketable product. For example, water may have to be removed from the oxygenate and an additional material such as, for example, hydrogen may have to be added to the oxygenate to facilitate further processing of the oxygenate. The additional material can add complexity to the process and can create a hazardous operational condition. Thus, a method, a system, and an apparatus are provided which can reduce the amount of purification and further processing required for the oxygenate. More specifically, a method, a system, and an apparatus are provided for converting a lower alkane to an aikene Oxygen and a lower alkane can be provided to an ODH reactor. At least a portion of the lower alkane can be converted to an aikene in the ODH reactor and an ODH outlet stream comprising the aikene, an oxygenate, water, and carbon monoxide with unconverted alkane can be produced. At least a portion of the ODH outlet stream can be provided to a water gas shift/hydrogenation (WGS/H) reactor including a WGS/H catalyst. The ODH outlet stream can be reacted within the WGS/H reactor and hydrogen and carbon dioxide can be generated from the carbon monoxide and water. At least a portion of the oxygenate and hydrogen are converted to an alcohol. An alcohol outlet stream comprising at least a substantial portion of the alcohol can be produced.

[2075] Referring to the Fig. 89, illustrated is a flow diagram of a non-limiting example of a system 8900 to convert an alkane to an aikene. As illustrated, an ODH reactor 8902 and a water gas shift/hydrogenation (WGS/H) reactor 8906 can be in operative communication. For example, an ODH outlet 8902 of the ODH reactor 8902 can be in fluid communication with a water gas shift/hydrogenation (WGS/H) inlet 8906a of the WGS/H reactor 8906 via ODH outlet line 8914.

[2076] The ODH reactor 8902 can comprise an ODH inlet 8902a which can be configured to receive an ODH inlet stream from an ODH inlet line 8908 and can be suitable to transport the ODH inlet stream into the ODH reactor 8902. The ODH inlet stream can comprise a gaseous mixture of a lower alkane and oxygen. In various examples, the ODH inlet stream additionally can include at least one of carbon dioxide, water (e.g., steam), and an inert diluent. The inert diluent can comprise, for example, nitrogen. The concentration of the oxy gen and the lower alkane within the mixture in the ODH inlet stream and the temperature and pressure of the ODH inlet stream can be adjusted such that the mixture can be outside of the flammability limits of the lower alkane.

[2077] In various examples, there may be multiple ODH inlet lines configured to provide the ODH inlet stream to the ODH reactor 8902. For example, each reactant (e.g., lower alkane, oxygen, water (e.g., steam), carbon dioxide, and inert diluent) may be added directly to the ODH reactor 8902, each in separate inlet lines. Alternatively, one or more reactants may be pre-mixed and added in more than one inlet line. In various example reactants may be mixed together prior to the ODH reactor 8902 and subsequently introduced into the ODH reactor 8902 in a conmion ODH inlet line. In various examples, steam may be added indirectly as water mixed with an additional reactant and the resulting mixture can be preheated before entering the ODH reactor 8902. When adding steam indirectly as water, the preheating process can increase the temperature of the mixture so that the water can be substantially converted to steam before entering the ODH reactor 8902.

[2078] The ODH reactor 8902 includes an ODH catalyst capable of catalyzing the ODH of the reactants to products such as, for example, an alkene, carbon monoxide, and an oxygenate. The catalyst may be, for example, a mixed metal oxide catalyst. In various examples, the products may additionally include at least one of carbon dioxide and water.

[2079] The ODH catalyst composition, temperature and pressure of the ODH reactor 8902, and the composition of the ODH inlet stream can be adjusted in order to vary the composition of products as known by one of ordinary skill in the art. For example, the ratio of the lower alkane to oxygen can be outside of the upper flammability limit of the mixture. In various examples, the oxygen concentration in the ODH inlet stream can be in a range of 0 1 % to 30 %by volume of the ODH inlet stream, and in some examples range from 0 1 % to less than 30 % by volume, less than 25 % by volume, or less than 20 % by volume. In various examples, the lower alkane concentration in the ODH inlet stream 8908 can range from 0 1% to 80 % by volume of the ODH inlet stream, and in some examples range from 0.1 % to less than 50% by volume or less than 40 % by volume.

[2080] In various examples, increasing the steam concentration in the ODH inlet stream can increase the amount of oxygenate produced relative to the alkene produced in the ODH reactor 8902. in various examples, reducing the steam concentration in the ODH inlet stream can decrease the amount of oxygenate produced relative to the alkene produced in the ODH reactor 8902. The concentration of steam in the ODH inlet stream can be in a range of 0.1 vol. % to 90 vol. % of the total ODH inlet stream, and in some examples range from 0.1 vol. % to less titan 40 vol. %, or less than 25 vol. %. In various examples, tire concentration of steam in the ODH inlet stream can be at least I vol. %. In various examples, the ODH inlet stream can comprise 20 vol. % oxygen, 40 vol. % lower alkane, and Site balance being water (e.g., steam), carbon dioxide, and/or an inert diluent. In various examples, the ODH inlet stream can comprise 10 vol. % oxygen, 15 vol. % lower alkane and the balance being water (e.g., steam), carbon dioxide, and/or inert diluent. [2081] In various examples, the ODH process has a selectivity for the corresponding alkene (e.g., ethylene in the case of ethane ODH) of grea ter than 95 % such as, for example, greater than 98 %. The gas hourly space velocity (GHSV) within the ODH reactor 8902 can be from 500 to 30000 h ^! and in some examples the GHS V within the ODH reactor 8902 can be greater titan 1000 h ¹ In various examples, the space-time yield of corresponding alkene (e.g., productivity) in grams(g)/hour per kilogram (kg) of the catalyst can be at least 100 such as, for example, at least 1500, at least 3000, or at least 3500, at an ODH reactor temperature of 350 °C to 370 °C. In various examples, the productivity' of the ca talyst can increase with increasing temperature in the ODH reactor 8902 until the selectivity of the alkene decreases.

[20821 An ODH reactor can be used for performing an ODH process consistent with the provided techniques. For best results, the oxidative dehydrogenation of a lower alkane may be conducted at temperatures from 300 °C to 370 °C. from 300 °C to 360 °C. for example, front 300 °C to 350 ^C C, at pressures from 0.5 to 100 psig (3.447 to 689.47 kPag), for example, from 15 to 50 psig (103.4 to 344.73 kPag), and the residence time of the lower alkane in the reactor is typically from 0.002 to 30 seconds, for example, from 1 to 10 seconds.

[2083] In some embodiments, the process has a selectivity for the corresponding alkene (ethylene in the case of ethane ODH) of greater than 85%, for example, greater than 90%. The flow' of reactants and inert diluent can be described in any number of ways known in the art. Typically, flow is described and measured in relation to the volume of all feed gases (reactants and diluent) that pass over the volume of the active catalyst bed in one hour, or gas hourly space velocity (GHSV) at standard temperature and pressure (STP). The GHSV can range from 500 to 30000 h ¹ , for example greater than 1000 h ¹ . The flow' rate can also be measured as weight hourly space velocity (WHSV), which describes the flow' in terms of the weight, as opposed to volume, of the gases that flow' over the weight of the active catalyst per hour. In calculating WHSV the weight of the gases may include only the reactants but may also include diluents added to the gas mixture. When including the weight of diluents, w'hen used, the WHSV may range from 0.5 h ¹ to 50 h ¹ , for example from 1.0 to 25.0 h ¹ .

[2084] The products of the ODH reaction can leave the ODH reactor 8902 through the ODH outlet 8902b in an ODH outlet stream. The ODH outlet 8902b can be configured to receive the ODH outlet stream and can be suitable to transport the ODH outlet stream out of the ODH reactor 8902 into the ODH outlet line 8914. In various examples, in addition to the products, the ODH outlet stream can include unreacted components from the ODH inlet stream such as, for example, lower alkane, carbon monoxide, carbon dioxide, oxy gen, water (e.g., steam), inert diluent, and combinations thereof.

[2085] Any of the known reactor types applicable for the ODH of an alkane may be used with the provided techniques. For example, a fixed bed reactor, a fluidized bed reactor, or combinations thereof can be used for the ODH reactor 8902, In a typical fixed bed reactor, reactants are introduced into the reactor at an inlet and flow past an immobilized catalyst. Products are formed and leave through the outlet of the reactor. A person skilled in the art would know which features are required with respect to shape and dimensions of the reactor, inputs for reactants, outputs for products, temperature and pressure control, and means for immobilizing the catal st. Shell-and-tube type reactors are well known as being applicable for use in ODH reactors, owing to the exothermic nature of the reaction. These reactors are designed with the goal of efficient removal of heat to prevent runaway reactions. [2086] In a typical fluidized bed reactor, the catalyst bed can be supported by a porous structure or a distributor plate and located near a lower end of the reactor. Reactants flow through the fluidized bed reactor at a velocity sufficient to fluidize the bed (e.g. , the catal st rises and begins to swirl around in a fluidized manner). The reactants can be converted to products upon contact with the fluidized catalyst and the reactants are subsequently removed from an upper end of the reactor. A person of ordinary skill in the art would know which features are required with respect to shape and dimensions of the reactor, the shape and size of the distributor plate, the input temperature, the output temperature, the reactor temperature and pressure, inputs for reactors, outputs for reactants, and velocities to achieve fluidization.

[2087] In various examples, there may be multiple ODH reactors connected in series or in parallel. Each ODH reactor may be the same or different. For example each ODH reactor can contain the same or different ODH catalyst. In various examples, the multiple ODH reactors can each be a fixed bed reactor, can each be a fluidized bed reactor or can be combinations of fixed bed reactors and fluidized bed reactors.

[2088] Regardless of the configuration of tire ODH reactor 8902, tire ODH outlet 8902b can be in fluid communication with the WGS/H inlet 8906a of the WGS/H reactor 8906 via the ODH outlet line 8914 to direct the ODH outlet stream to the WGS/H reactor 8906. The WGS/H inlet 8906a can be configured to receive the ODH outlet stream from the ODH outlet line 8914 and can be suitable to transport the ODH outlet stream into the WGS/H reactor 8906.

[2089] In an alternative configuration the ODH outlet stream can be in fluid communication with a separator the separator having an inlet configured to receive the ODH outlet stream from the ODH outlet line 8914 and can be suitable to transport the ODH outlet stream into the separator. The separator may include cooling and quenching of the ODH outlet stream for the purpose of separating out water and oxygenate from the ODH outlet stream, producing an alkene outlet stream and an oxygenate outlet stream. The separator may also have an alkene outlet configured to remove the alkene outlet stream from the separator and an oxygenate outlet for removing the oxygenate outlet stream from the separator. The oxygenate outlet stream would be provided to the WGS/H reactor 8906 through the WGS/H inlet 8906a in place of the ODH outlet stream. In various embodiments, an ODH outlet stream and an oxygenate outlet stream can be provided to the WGS/H reactor 8906 via WGS/H inlet 8906a.

[2090] The separator can be a quench tower, an oxygenate scrubber, a flash drum, the like, or combinations thereof. The separator can be configured to remove at least a substantial portion of the alkene from the ODH outlet stream. The separator can produce an alkene outlet stream comprising at least a substantial portion of the alkene from the ODH outlet stream. In various examples, the alkene outlet stream can comprise additional components from the ODH outlet stream such as for example, a portion of the carbon monoxide, a portion of the carbon dioxide, a portion of the oxygen, a portion of the oxygenate, a portion of the inert diluent, a portion of the water (e.g., steam), and a portion of the unreacted alkane.

[2091 ] The separator produces an oxygenate outlet stream comprising at least a substantial portion of the oxygenate from the ODH outlet stream. In various examples, the oxygenate outlet stream can comprise additional components from the ODH outlet stream such as, for example, a substantial portion of the water, a portion of the carbon monoxide, a portion of the carbon dioxide, and a portion of the oxygen. In various examples the oxygenate outlet stream comprises a molar ratio of water to oxygenate of greater than 1 : 1 such as, for example, 1.1:1, or 2:1. [2092] The temperature of the ODH outlet stream, or oxygenate outlet stream when using a separator, can be adjusted prior to entering the WGS/H reactor 8906. For example, the temperature of the ODH outlet stream prior to entering the WGS/H reactor 8906, or separator, can be at a temperature of at least 40 °C or at least 50 °C such as, for example, 40 °C to 350 °C or 50 °C to 200 °C,

[2093] The WGS/H reactor 8906 can be configured to facilitate a water gas shift (WGS) reaction and a h drogenation reaction. In various examples, combining both the WGS and hydrogenation reactions into a single reactor can eliminate the need to add hydrogen and the need to remove water from the oxygenate outlet stream prior to hydrogenation. For example, the WGS/H reactor 8906 can generate, in sifts, hydrogen for hydrogenation of the oxygenate using the WGS reaction. The WGS/H reactor can comprise a WGS/H catalyst to convert a portion of the carbon monoxide and a portion of the water in the oxygenate outlet stream to carbon dioxide and hydrogen as shown in Scheme Zl.

Scheme Z 1 :

CO + H ₂ 0 <® C0 ₂ + H ₂

[2094] As Scheme Zl indicates, carbon monoxide is a reactant. As a result, when using a separator prior to the WGS/H reactor a user may be required to add carbon monoxide to the oxygenate outlet stream as the levels of carbon monoxide may not be sufficient to convert a desirable amount of the water present in the stream. For tins reason, one example of a preferred configuration includes providing the ODH outlet stream 8914 directly to the WGS/H reactor 8906. without first passing through a separator.

[2095] The WGS/H catalyst can comprise a metal oxide catalyst. The metal oxide catalyst can be non-acidic. For example, the non-acidic catalyst may not contain a chemical species that contains an empty orbital which can be capable of accepting an election. In various examples, the metal oxide catalyst can be a solid catalyst. The metal oxide catalyst can comprise, copper oxide, zinc oxide, aluminum oxide, iron oxide, chromium oxide, magnesium oxide, a noble metal, and ceria. For example, the metal oxide catalyst can comprise at least one of copper, iron, platinum, tin, and chromium. The WGS/H reactor 8906 can operate at a temperature in a range of 100 °C to 500 °C such as for example, 200 °C to 300 °C. The WGS H reactor 106 can operate at a pressure in a range of 100 kilopascals (kPag) to 8375 kPag such as, for example, 100 kPag to 500 kPag. In various examples, the liquid hourly space velocity (LHSV) in the WGS/H reactor 106 can be at least 0.1 h ^'1 such as, for example, a range of 0.3 ti ¹ to 0.7 h ¹ , or at least 2 h ¹ .

[2096] The WGS/H reactor 8906 can contact the hydrogen with the oxygenate and can be configured to perform a hydrogenation reaction. In various examples, the WGS/H catalyst can catalyze the WGS/H reaction and the hydrogenation reaction. For example, the WGS/H catalyst can facilitate conversion of a portion of the hydrogen and a portion of the oxygenate from the oxygenate outlet stream to an alcohol and water. As an example, the hydrogenation of acetic acid is shown in Scheme Z2. In various examples, the alcohol can comprise at least one of ethanol, propanol, and butanol.

Scheme Z2 CH ₃ COOH + 2 H ₂ ® CH ₃ H ₂ OH + H ₂ 0

[2097] The products of the WGS and hydrogenation reaction can exit the WGS/H reactor 8906 through the WGS/H outlet 106b in a WGS/H outlet stream. The WGS/H outlet 8906b can be configured to receive the WGS/H outlet stream and can be suitable to transport the WGS/H outlet stream 8918 out of the WGS/H reactor 8906 into the WGS/H outlet line 8918. Moreover, the WGS reaction can be exothermic. Thus, the temperature of the oxygenate outlet stream can be maintained at a temperature of less than 350 °C such that the WGS H outlet stream 8918 does not reach above a select temperature such as, for example 550 °C.

[2098] Various reactor types applicable for a WGS reaction and/or a hydrogenation reaction can be employed for use with the provided techniques. For example, a fixed bed reactor, a fluidized bed reactor, or combinations thereof can be used for the WGS/H reactor 8906. In various examples, there may be multiple WGS/H reactors connected in series or in parallel. In various examples a shell-and-tube reactor design may be appropriate due to the exothermic nature of the WGS/H reaction.

[2099] The alcohol can be converted to another product such as, for example, an alkene, an ether, an aldehyde, the like, and combinations thereof. For example, if the alcohol comprises ethanol, the ethanol can be converted to ethylene, diethyl ether, acetaldehyde, the like, and combinations thereof.

[2100] Referring to Fig. 90, the system 8900 includes a dehydration reactor 8920. As illustrated, the WGS/H outlet 8906b can be in fluid communication with a dehydration inlet 8920a of the dehydration reactor 8920 via the WGS H outlet line 8918 to direct die WGS/H outlet stream to the dehydration reactor 8920. The dehydration inlet 8920a can be configured to receive the WGS/H outlet stream from the WGS/H outlet line 8918 and can be suitable to transport the WGS/H outlet stream into die dehydration reactor 8920.

[2101] The dehydration reactor can convert at least a portion of the alcohol within the WGS/H outlet stream into an alkene. For example, die dehydration of ethanol can be shown in Scheme Z3.

Scheme Z3

CH ₃ H ₂ OH <® CH ₂ CH ₂ + H ₂ O

[2102] The dehydration reactor 8920 can produce a dehydration outlet stream comprising the alkene generated within die dehydration reactor 8920. The dehydration outlet stream can exit the dehydration reactor 8920 through the dehydration outlet 8920b. The dehydration outlet 8920b can be configured to receive the dehydration outlet stream and can be suitable to transport the dehydration oudet stream out of the dehydration reactor 8920 into die dehydration outlet line 8922. The dehydration reactor 8920 can operate at a temperature of 100 °C to 300 °C such as for example, 175 °C to 275 °C, 150 °C to 250 °C and in some examples, the dehydration reactor 8920 can operate at a temperature of less than 200 °C.

[2103] Since, the ODH process operates at a temperature of 300 °C to 375 °C, the WGS/H reactor 8906 can operate at a temperature of 100 °C to 500 °C, and the dehy dration reactor can operate at a temperature of 100 °C to 300 °C an energy savings can be achieved by combining the processes together. For example, the residual heat in the ODH outlet stream can facilitate the WGS and hydrogenation reactions i the WGS/H reactor 8906. In various examples, the residual heat in the WGS H outlet stream can facilitate the dehydration reaction in the dehydration reactor 8920. In various examples, operating the WGS/H reactor 106 in fluid communication with the ODH reactor 8904 enables a reduction in energy required to produce the alcohol relative to operating the processes separately.

[2104] Concentrations of the components within the system can be measured any at point in the process using any means known in the art. For example, a detector such as a gas chromatograph, an infrared spectrometer, and a Raman spectrometer can be disposed downstream or upstream of ODH reactor 8902, separator 8904, WGS/H reactor 8906, and dehydration reactor 8920.

[2105] In various examples, the WGS/H outlet stream, the dehy dration outlet stream, or bo th may be directed to a separator such as those described. Separation of oxygenate, water, and alcohol may be promoted so that the alkenes within the streams can be isolated for downstream applications.

[2106] In various examples, the ODH inlet stream comprises mixtures that fall within the flammability limits of the components. For example the mixture may exist in conditions that prevent propagation of an explosive event. In these examples the flammable mixture can be created within a medium where ignition can be immediately quenched. In various examples, oxygen and Site lower alkanes can be mixed at a point where they are surrounded by a flame arresting material. Thus, any ignition can be quenched by the surrounding material. Flame arresting material includes, for example, metallic or ceramic components, such as stainless steel walls or ceramic supports. In various examples, oxygen and lower alkanes can be mixed at a low temperature, where an ignition event may not lead to an explosion, then the mixture can be introduced into the ODH reactor 8902 before increasing the temperature. Therefore, the flammable conditions may not exist until the mixture can be surrounded by the flame arresting material inside of the ODH reactor 8902.

[2107] In various examples, the olefins produced using an ODH reactor 8902, or any of the processes or complexes described herein, can be used to make various olefin derivatives utilizing a polymerization reactor. Olefin derivatives include, but are not limited to polyethylene, polypropylene, ethylene oxide, propylene oxide, polyethylene oxide, polypropylene oxide, vinyl acetate, vinyl chloride, acrylic esters (e.g., methyl methacrylate), thermoplastic elastomers, thermoplastic olefins, blends thereof, and combinations thereof

[2108] In various examples, ethylene and optionally a-olefins can be produced in an ODH reactor 8902, or any of the processes or complexes described herein, and are used to make polyethylene utilizing a polymerization reactor. The polyethylene made from the ethylene and optional a-olefins described herein can include homopolymers of ethylene, copolymers of ethylene and a-olefins, resulting in HOPE, MDPE, LDPE, LLDPE and VLDPE.

[2109] The polyethylene produced using the ethylene and optional a-olefins described herein can be produced using any suitable polymerization process and equipment. Suitable ethylene polymerization processes include, but are not limited to gas phase polyethylene processes, high pressure polyethylene processes, low pressure polyethylene processes, solution polyethylene processes, slurry polyethylene processes and suitable combinations of die above arranged either in parallel or in series.

[2110] A process for converting a lower alkane to an alkene can include providing an input stream comprising oxygen and tire lower alkane to an ODH reactor 8902. At least a portion of the lower alkane can be converted to the alkene in the ODH reactor 8902. In various examples, the alkane comprises ethane and the alkene comprises ethylene. In various examples, the alkane comprises propane and the alkene comprises propylene. In various examples, the alkane comprises butane and the alkene comprises butylene. An ODH outlet stream comprising the alkene, an oxygenate, water, and carbon monoxide can be produced. In various examples, the ODH outlet stream additionally includes at least one of an unreacted alkane, carbon dioxide, and oxygen. In various examples, the ODH outlet stream can be substantially free of hydrogen. In various examples, the oxygenate can comprise at least one of acetic acid, aciylic acid, maleic acid, maleic anhydride, and ethanol.

[2111J in various examples, the temperature of the ODH outlet stream can be adjusted by vary ing reaction conditions in the ODH reactor 8902, passing the ODH outlet stream through a heat exchanger, or combinations thereof. In various examples, the ODH outlet stream can have a temperature of 300° C to 375° C.

[2112] The ODH outlet stream can be provided to a WGS/H reactor 8906 comprising a WGS/H catalyst. In various examples the WGS/H catalyst can be non-acidic. In various examples, the WGS/H catalyst can comprise at least one of copper, iron, platinum, tin, and chromium. In various examples, die WGS/H reactor 8906 can be maintained at a temperature of 100 °C to 500 °C. In various examples, tire WGS H reactor 106 can be maintained at a pressure of 100 kPag to 500 kPag.

[2113] The WGS/H reactor 8906 can generate hydrogen, in situ, from the carbon monoxide and water of the ODH outlet stream. At least a portion of the oxygenate and hydrogen can be converted to an alcohol. An alcohol outlet stream co mpri sing at least a substantial portion of the alcohol can be produced. In various examples, the alcohol comprises at least one of ethanol, propanol, and butanol.

[2114] The alcohol outlet stream can be provided to a dehydration reactor and at least a portion of the alcohol in the alcohol outlet stream can be converted to a second alkene. In various examples the second alkene is the same species of alkene produced in the ODH reactor and comprises at least one of ethylene, propylene, and butylene. In various examples the WGS/H reactor and the dehydration reactor may be a single reactor. In this instance, the reactor may include a WGS/H catalyst and a dehydration catalyst, spatially separated so that as components of the ODH outlet stream move through the single reactor contact with the WGS/H catalyst precedes contact with the dehydration catalyst. The heat produced by the reaction with the WGS/H catalyst supports the dehydration catalyst for removing ¾0 from the alcohol to produce an alkene.

[2115] Olefin derivatives can be produced from the alkenes, including the alkene produced in the ODH reactor and the second alkene produced in the dehydration reactor.

[2116] Techniques are provided for a route to convert an alkane into an alkene, reduction of an undesired product, an opportunity to run at a high alkane conversion (e.g., greater than 95%) where an increase in oxy genate selectivity may result and generate an additional marketable product.

[2117] Examples

[2118] The WGS/H reactor 8906 and the dehydration reactor 8920 were computationally modeled using ASPEN Plus® version 8.6 chemical process simulation software. The ODH outlet line 8914 was in fluid communication with the WGS/H inlet 8906a of the WGS/H reactor 8906. The WGS/H outlet 8906b of the WGS/H reactor 8906 was in fluid communication with the dehydration inlet 8920a of the dehydration reactor 8920 via the WGS/H outlet line 8918 The WGS/H reactor 8906 and the dehydration reactor 8920 were modeled as RGibhs reactors and were assigned the EoS property method of SR-POLAR.

[2119] Examine Zl to Z5

[2120] For examples Zi to Z5, the composition of the ODH (O) outlet stream was set with a 1 kg/hr mass flow rate and the O outlet stream comprised acetic acid, water, and carbon monoxide as illustrated in Table Zl. Five different simulations were performed while varying the temperature of the WGS/H reactor and the temperature of the dehydration reactor from 100 °C to 500 °C in 100 °C increments as illustrated in Table Zl The pressure of tire WGS/H reactor and the dehydration reactor were each kept constant at 100 kPag.

[2121] Table Zl:

[2122] As showm in Table Zl, the exothermic nature of the WGS reaction and hydrogenation reactions within the W GS/H reactor caused a substantial portion of the acetic acid, carbon monoxide, and water to be converted to ethanol, carbon dioxide, and hydrogen. Additionall , the dehydration reaction of the ethanol to ethylene in the dehydration reactor was favored at temperatures of less than 300 °C and a high conversion of the carbon monoxide in the WGS/H reactor is thermodynamically favored at temperatures less than 300 °C for these conditions. [2123] Examples Z6 to Z 10

[2124] For examples Z6 to Z10, the composi tion of the O outlet stream was set with a 1 kg/hr mass flow rate and the O outlet stream comprised acetic acid, water, and carbon monoxide as illustra ted in Table Z2, Five different simulations were performed while vary ing the pressure of the WGS/H reactor and the pressure of the dehydra tion reactor from 100 kPag to 500 kPag in 100 kPag increments as illustrated in Table Ll. The tempera ture of the WGS/H reactor and the dehydration reactor were each kept constant at 100 °C

[2125] Table Z2:

[2126] As shown in Table Z2, the exothermic nature of the WGS reaction and hydrogenation reactions within the W GS/H reactor caused a substantial portion of the acetic acid, carbon monoxide, and water to be converted to ethanol, carbon dioxide, and hydrogen. Additionally, the dehydration reaction of the ethanol to ethylene in the dehydration reactor is favored at lower pressures for these conditions. The change in pressure caused insignificants changes in ethanol production.

[2127] Certain embodiments include a multi-stage ODH reactor. The multi-stage ODH reactor may be a vessel having multiple reactors in series, and with each reactor in the vessel hatting a fixed bed of ODH catalyst. The multiple reactors of the ODH reactor vessel may be called reactor stages or “stages” in senes. Each stage of the vessel lias a respective fixed bed of ODH catalyst in the vessel. The number of stages may be, for example, at least three stages or at least four stages. A single vessel may have multiple reactor stages each having a fixed bed of ODH catalyst. The catalyst is generally on the process side of the stages. In embodiments, these reactor stages associated with a single vessel are semi-adiabatic ODH.

[2128] Feed for reaction may be introduced to the first stage in the series. The effluent from the first stage may enter the second stage. The effluent from the second stage may enter the third stage, and so on. As for the feed to the first stage, a lower alkane (e.g., ethane), steam, and oxygen (0 ₂ ) are added to die inlet of die first stage in the series of stages. Carbon dioxide (C<¼) may also be added to die inlet of die first stage. In implementations, steam and 0 ₂ are not added to the remaining stages except for any steam and 0 ₂ received in the effluent from the preceding stage in certain embodiments, external steam and 0 ₂ are added to the enhance of the first stage for reaction in the series and with no inter-stage addition of steam or 0 ₂ for subsequent stages from a source external to the reactor. [2129] Each stage has an independent cooling jacket having an inlet for coolant and an outlet for return coolant. The stages individually have a dedicated respective cooling jacket on the multi-stage ODH reactor vessel for temperature control of each stage. In examples, the coolant may be supplied in parallel to the cooling jackets.

The flow rate (e.g., mass flow rate or volumetric flow rate) of the coolant through the cooling jackets may be modulated to control temperature of the reactor stages. In some examples, coolant flow rate is controlled independently to each stage and thus temperature may be controlled independently for each stage. Heat transfer occurs from the reaction mixture and catalyst on the process side of the discrete stages to the respective coolant in the cooling jackets. The temperature of each stage (e.g., temperature of the respective catalyst bed) may be maintained above a lower threshold (e.g., 200 °C) for the ODH reaction of the lower alkane to the corresponding alkene (e.g., ethylene). The temperature in the reactor beds may be maintained below' an upper threshold. For instance, an upper threshold temperature for each stage can be 375 °C, due to thermal stability' of the catalyst. Acetic acid and easbon monoxide may be produced as byproducts in the stages.

[2130] A feed mixing node may be implemented for mixing 0 ₂ and hydrocarbon for feed to the process side of the first stage. As indicated for certain embodiments, 0 ₂ injections to subsequent stages from sources external to the reactor vessel are not implemented. Intermediate 0 ₂ injections can be costly and complex. Moreover, with no intermediate 0 ₂ injections, the 0 ₂ concentration generally decreases through the series of stages. Therefore, the duty of the heat removal may be less in later stages than earlier stages. Further, as also indicated, intermediate steam injection is not implemented in some embodiments. This may facilitate a lower pressure drop through the series of stages. Lastly, in contrast to isothermal ODH, the present semi-adiabatic operation may give less boiling of the coolant in the coolant jacket at the interface surface of the coolant jacket with the process side.

[2131] Methods for steam generation via an catalytic oxidative dehydrogenation reaction are provided. The following aspects are related to an ODH reactor s stem. The ODH reactor system includes a first reactor having a first ODH catalyst to dehydrogenate a lower alkane to a corresponding alkene at a first temperature and facilitate generation of steam. The first reactor lias a first-reactor jacket for heat transfer. The ODH reactor system includes a second reactor having a second ODH catalyst to dehydrogenate unreacted lower alkane in a first-reactor effluent from the first reactor to the corresponding alkene at a second temperature greater than the first temperature and facilitate generation of steam. The second reactor has a second-reactor jacket for heat transfer. The GDH reactor system includes a third reactor having a third ODH catalyst to dehydrogenate unreacted lower alkane in a second- reactor effluent from the second reactor to the corresponding alkene at a third temperature greater than the second temperature and facilitate generation of steam. The third reactor has a third-reactor jacket for heat transfer.

[2132] An aspect relates to a system for oxidative dehydrogenation. The system includes a first reactor having a first ODH catalyst to dehydrogenate an alkane at a first temperature. The first reactor lias a first-reactor jacket to heat a first heat-transfer fluid flowing through the first-reactor jacket to facilitate generation of steam. The system includes a second reactor having a second ODH catalyst to dehydrogenate unreacted alkane from the first reactor at a second temperature greater than the first temperature. The second reactor Isas a second-reactor jacket to heat a second heat-transfer fluid flowing through the second-reactor jacket to facilitate generation of steam. The system includes a third reactor having a third ODH catalyst to dehydrogenate unreacted alkane from the second reactor at a third temperature greater than the first temperature. The third reactor lias a third-reactor jacket to heat a third heat- transfer fluid flowing through the third-reactor jacket to facilitate generation of steam. The third ODH catal st and the second ODH catalyst are different than the first ODH catalyst.

[2133] An aspect relates to a method of oxidative dehydrogenation. The method includes contacting a feed having a lower alkane with a first ODH catalyst in a first reactor at a first temperature to dehydrogenate the lower alkane into a corresponding alkene and to heat a first heat-transfer fluid flowing through a first-reactor jacket to facilitate generation of steam. The method includes contacting a first-reactor effluent from the first reactor with a second ODH catalyst in a second reactor at a second temperature greater than the first temperature to dehydrogenate luireacted lower alkane from the first-reactor effluent into the corresponding alkene and to heat a second heat- transfer fluid flowing through a second-reactor jacket to facilitate generation of steam. The method includes contacting a second-reactor effluent from the second reactor with a third ODH catalyst in a third reactor at a third temperature greater than the first temperature to dehydrogenate unreacted lower alkane from the second effluent into the corresponding alkene and to heat a third heat-transfer fluid flowing through a third-reactor jacket to facilitate generation of steam.

[2134] Certain embodiments encompass a system and method for oxidative dehydrogenation including a first reactor, a second reactor, and a third reactor. The first reactor has a first ODH catalyst to dehydrogenate an alkane to a corresponding alkene at a first temperature and facilitate generation of steam. The second reactor has a second ODH catalyst to dehydrogenate alkane in a first-reactor effluent to the corresponding alkene at a second temperature that may be greater than the first temperature and facilitate generation of steam. The third reactor has a third ODH catalyst to dehydrogenate alkane in a second-reactor effluent to the corresponding alkene at a third temperature that may be greater than the first temperature or the second temperature and facilitate generation of steam.

[2135] The catalytic ODH reaction is exothermic. Therefore, steam may be generated as a coproduct in utilizing heat from the ODH reaction. The steam production may be characterized as integrated with or within the ODH reactor system. In addition, the usage of the produced steam may be integrated at the site having the ODH reactor system. The produced steam may be utilized in the overall ODH system or in other unit operations or units at the facility having the ODH s stem. The produced steam may also be exported for use by other facilities or sites. [2136] Different qualities or pressures of steam may be generated as a coproduct of the ODH reaction. The term “quality” of the steam may refer to the pressure or type of steam. Typical qualities of steam produced are low pressure steam (e.g., 150 pounds per square inch gauge [psig] or less), medium pressure steam (e.g., in the range of 150 psig to 600 psig), high pressure steam (e.g., 600 psig or greater), or very high pressure steam (e.g., 1500 psig or greater), and so forth. There may be different applications for the steam. The use of the steam by the consumers or customers receiving the steam may depend on the quality or pressure of the steam. In some implementations, higher steam pressures of the produced steam may give more versatility in the integration of the steam within the facility or plant. For instance, high pressure steam is used to power turbines attached to compressors while low pressure steam is typically used for heating puiposes, and the like.

[2137] In some implementations, two or more ODH reactors in series may operate at progressively higher temperature to generate different qualities of steam. The respective operating temperature of the ODH reactors may be increasingly greater along the series of ODH reactors. The second ODH reactor may have a higher operating temperature than the first ODH reactor. The third ODH reactor may have a higher operating temperature than the second ODH reactor, and so on.

[2138] The different reaction temperatures among the respective ODH reactors may be due to utilization of different types or grades of catal sts in the respective ODH reactors. The catalyst in the second ODH reactor may give an ODH reaction at a greater temperature than the catalyst in the first ODH reactor. The catalyst in the third ODH reactor may give ao ODH reaction at a greater temperature than the catalyst in the second ODH reactor, and so on.

[2139] Three ODH reactors are depicted in the ODH reactor systems of FIGS. 90-7. However, the present ODH reactor systems may have only two ODH reactors in series or may have more than three ODH reactors (e.g., four ODH reactors, five ODH reactors, etc.) in series or parallel for the generation of steam. The final ODH reactor in the series may discharge an effluent having a product alkene of the ODH reactor system.

[2140] FIG. 90 is an ODH reactor system 9000 including a first ODH reactor 9002, a second ODH reactor 9004, and a third ODH reactor 9006 operationally disposed in series in the illustrated embodiment, the ODH reactors 9002, 9004, 9006 are tubular reactors having a process side and a cooling jacket. The process side is one or more tubes or conduits for the reaction of the alkane to alkene. The process side has catalyst (e.g., a fixed bed of catalyst) for the conversion of the alkane to the corresponding alkene. The system 9000 flows water as a heat transfer fluid through the jacket side to control the reaction temperature on the process side. The heat or energy acquired by the heat transfer fluid through the reactor jacket may he utilized to generate steam as a coproduct. For the system 9000 in operation, the liquid water is depicted with gray shading in FIG. 90. Such gray shading for liquid water is also utilized in FIGS. 91-97.

[2141] The first ODH reactor 9002 has a process side 9008 having a first catalyst 9010. The second ODH reactor 9004 lias a process side 9012 having a second catalyst 9014. The third ODH reactor 9006 has a process side 9016 having a third catalyst 9018. As indicated, each process side 9008. 9012, 9016 may be one or more conduits or tubes in some examples. The first catalyst 9010, second catalyst 9014, and third catalyst 9018 may each be a fixed bed of catalyst. The first catalyst 9010, second catalyst 9014, and third catalyst 9018 may be the same catalyst type or different respective catalyst types.

[2142 J In operation, the process side 9008 of the first ODH reactor 9002 may receive a feed 9020 having an alkane. The feed 9020 as a hydrocarbon feed may also include oxygen for the ODH reaction. However, the oxygen may be added to the first ODH reactor separate from the feed 9020. The alkane in the feed 9020 may be a lower alkane defined as an alkane (saturated hydrocarbon) having a number of carbons in the range of 2 to 6. The first ODH reactor 9002 may receive the feed 9020 via a conduit coupled (e.g., by a flanged connection) to an inlet of the first ODH reactor 9002 vessel at the process side 9008. The first ODH reactor 9002 may convert the alkane in the feed 9020 to a corresponding alkene in a catalytic reaction via the catalyst 9010 on the process side 9008 of the first reactor 9002. Some of the alkane in the feed 9020 is not converted into the corresponding alkene but remains umeacted. The first ODH reactor 9002 discharges an effluent 9022 having the corresponding alkene and unreacted alkane. In some implementations, tire alkane is ethane and tire corresponding alkene is ethylene. In some implementations, the effluent 9022 may also include acetic acid. In addition, the effluent 9022 may include carbon dioxide, water, and so forth.

[2143] The process side 9012 of the second ODH reactor 9004 may receive (e.g., via a conduit) the effluent 9022 from the first ODH reactor 9002. The second ODH reactor 9004 may convert the uureacted alkane to the corresponding alkene in a catalytic reaction via the catalyst 9014 on the process side 9012 of the second reactor 9004. Some of the unreacted alkane is not converted into the corresponding alkene but remains unreacted. The second ODH reactor 9004 discharges an effluent 9024 having the corresponding alkene and unreacted alkane.

[2144] The process side 9016 of the third ODH reactor 9006 may receive (e.g., via a conduit) the effluent 9024 from the second ODH reactor 9004. The third ODH reactor 9006 may convert the unreacted alkane to the corresponding alkene in a catalytic reaction with the catalyst 9018 on the process side 9016 of the third reactor 9006 The third ODH reactor 9006 discharges an effluent 9026 having the corresponding alkene and any unreacted alkane. The corresponding alkene (e.g., ethylene) may be a product of the ODH reactor system 9000.

[2145] In some embodiments, the catalyst 9010, 9014, 9018 in the reactors 9002, 9004, 9006 (process sides 9008, 9012, 9016) is different, respectively, and may give conversion of the alkane to the corresponding alkene at different temperatures, respectively. The first catalyst 9010 may be different than the second catalyst 9014 and the third catalyst 9018, and the second catalyst 9014 may be different than the third catalyst 9018. In certain embodiments, the third reactor 9006 reaction temperature is greater than the second reactor 9004 reaction temperature, and the second reactor 9004 reaction temperature is greater than the first reactor 9002 reaction temperature.

[2146] The arrangement of the three ODH reactors 9002, 9004, 9006 may be a once-through effluent/feed configuration in that the effluent 9022 from the first ODH reactor 9002 is feed to the second ODH reactor 9004, and the effluent 9024 from the second ODH reactor 9004 is feed to the third ODH reactor 9006. In certain embodiments, oxygen may be injected into the effluent 9022 or the second ODH reactor 9004 to supplement the effluent 9022 with oxygen to account for consumption (depletion) of ox gen by the first ODH reactor 9002. Likewise, oxygen may be injected into the effluent 9024 or third ODH reactor 9006 to supplement the effluent 9024 with oxygen to account for depletion of oxygen by the second ODH reactor 9004. Oxygen may also be led directly to the second reactor 9002 or the third reactor 9004.

[2147] In some implementations, the ODH reactor system 9000 may have a conduit to provide, if desired, feed (e.g., similar or same as feed 9020) or fresh lower alkane (e.g., ethane) to the second ODH reactor 9004 to supplement the effluent 9022 received from the first ODH reactor 9002. This additional feed or fresh alkane may be added to the effluent 9022 or directly to the reactor 9004. Similarly, the ODH reactor system 9000 may have a conduit to provide feed (e.g., similar or same as feed 9020) or fresh lower alkane to the third ODH reactor 9006 to supplement the effluent 9024 received from the second ODH reactor 9004. This additional feed or fresh alkane may be added to the effluent 9024 or directly to the reactor 9006.

[2148] Moreover, any acetic acid in the effluent 9022 from the first ODH reactor 9002 may be combusted into carbon monoxide and carbon dioxide and in the second ODH reactor 9004, depending on the operating temperature of the second ODH reactor 9004. Any acetic acid in die effluent 9024 from the second ODH reactor 9004 may be combusted into carbon monoxide and carbon dioxide in the third ODH reactor 9006, depending on the operating temperature of the third ODH reactor 9006.

[2149] The first ODH reactor 9002 has a jacket 9028, the second ODH reactor 9004 has a jacket 9030. and the third ODH reactor 9006 has a jacket 9032. In this illustrated example of FIG. 90, the heat transfer fluid that flows through the jackets 9028, 9030, 9032 is water (e.g., boiler feedwater). In addition, in this example, the ODH reaction system 9000 includes three flash vessels 9034, 9036, 9038 for the generation of steam. A liquid level (water) may be maintained in the three flash vessels 9034, 9036, 9038.

[2150] Water from a source 9040 (e.g., a vessel) is supplied via a motive device 9042 (e.g., pump) through conduits to the flash vessels 9034, 9036, 9038. One or more control components, such as control valves 9044, 9046, 9048, disposed along the respective conduits may maintain or adjust the amount of water conveyed to the flash vessels 9034, 9036, 9038. The control valves 9044, 9046, 9048 may maintain or modulate the volumetric flow rate or mass flow rate of the supplied water. The water may be demineralized water, steam condensate, or boiler feedwater, and the like.

[21511 Water from the flash vessels 9034, 9036, 9038 as the heat transfer fluid may be provided through conduits to the ODH reactors 9002, 9004, and 9006, respectively, to control temperature on the process side of the reactors. Water may circulate from the respective flash vessel 9034, 9036, 9038 through the jacket 9028, 9030,

9032. The motive force for the circulation may be by thermosiphon. In other examples, a motive device (e.g., pump) is disposed on each circulation loop to provide motive force (e.g., to pump) the water through the jacket.

[2152] The reaction temperature in the reactors 9002, 9004, 9006, may depend on the type of catal st. Thus, the temperature of tire water flowing through Site jackets 9028, 9030, 9032 may be affected by the type of catalyst 9010, 9014, 9018 on the process side of the reactors. Therefore, the pressure of the steam generated in Site flash vessels 9034, 9036, 9038 may be affected by the type of catalyst in Site respective reactor.

[2153] For the first ODH reactor 9002. water 9050 from the first flash vessel 9034 enters and flows through tire first-reactor jacket 9028 to acquire heat from the first-reactor process side 9008. The heated water exits the jacket 9028 as return water 9052 to the flash vessel 9034. The heat acquired by the water promotes flashing of liquid water into steam in the flash vessel 9034. Steam 9054 discharges overhead from the flash vessel 9034 (e.g., into a conduit). The steam 9054 may be a coproduct of the ODH reactor system 9000.

[2154] The pressure of the steam 9054 may depend on the catalyst 9010 in first ODH reactor 9002. In other words, the reaction temperature driven by the catalyst 9010 generally affects the temperature of the return water 9052 discharging from the jacket 9028 to the flash vessel 9034. The temperature of the return water 9052 may affect the pressure at which the liquid water flashes into the steam 9054 discharging from the flash vessel 9034. in examples, the steam 9054 may generally be saturated steam. In certain implementations, Site catalyst 9010 in Site first reactor 9002 provides for a reaction temperature of less Ilian 400 °C. The steam 9054 may be, for example, low- pressure steam at less than 150 psig. Moreover, the amount (and pressure) of steam 9054 may be correlative to the first reactor duty (heat generated) removed via heat of vaporization of the Hashing water in the flash vessel 9034.

The duty or amount of heat generated by the first ODH reactor 9002 may be related to the catalyst 9010 employed in the first reactor 9002 and to the production rate of the corresponding alkene in the first reactor 9002, and so on. [2155] For the second ODH reactor 9004, water 9056 discharges from a bottom portion of the second flash vessel 9036 and flows through the second-reactor jacket 9030 to receive heat from the process side 9012 of the second ODH reactor 9004. The heated water exits the jacket 9030 as return water 9058 to the flash vessel 9036. The heat acquired by the water promotes flashing of liquid water into steam in the second flash vessel 9036. Steam 9060 discharges overhead from the second flash vessel 9036. The steam 9060 may be a coproduct of the ODH reactor system 9000.

[2156] The pressure of the steam 9060 may depend on the catalyst 9014 in the second ODH reactor 9004. The reaction temperature driven by the catalyst 9014 in the second ODH reactor 9004 generally affects the temperature of the return water 9058 discharging from the jacket 9030 to the flash vessel 9036. The temperature of the return water 9058 may affect the pressure at which the liquid water flashes into the steam 9060 discharging from the second flash vessel 9036. Moreover, the amount (and pressure) of the steam 9060 may he related to the heat generated by the reaction in the second ODH reactor 9004, and removed via heat of vaporization of the flashing water in the second flash vessel 9036. The amount of heat generated by the second ODH reactor 9004 may be related to the catalyst 9014 employed in the second reactor 9004 and to the amount (rate) of alkane converted to alkene in tire second reactor 9004, and the like.

[2157] In examples, the steam 9060 may generally be saturated steam in certain implementations, the catalyst 9014 in the second reactor 9004 provides for a reaction temperature of greater than 400 °C (e.g., in the range of 400 °C to 500 °C). In those implementations, the steam 9060 exiting overhead from the second flash vessel 9036 may be, for example, medium pressure steam in the range of 150 psig to 600 psig, or in the range of 250 psig to 400 psig.

The catalyst 9014 providing for a reaction temperature in the range of 400 °C to 500 °C may also give or result in the steam 9060 as high pressure steam at 600 psig or greater. In some embodiments, the steam 160 may be routed through a heat exchanger (e.g., shell-and-tube heat exchanger) to heat the steam 9060 to above saturation temperature. Thus, in certain implementations, the steam 9060 is superheated high-pressure steam. In examples, a heat transfer fluid providing heat in the heat exchanger that superheats the steam 9060 is the effluent 9026 from the third ODH reactor 9006 routed through the heat exchanger.

[2158] For the third ODH reactor 9006, water 9062 flows from the bottom portion of the third flash vessel 9038 through the third-reactor jacket 9032 to accumulate heat from the process side 9016 of the third ODH reactor 9006. The heated water exits the jacket 900032 as return water 9064 to the third flash vessel 9038. The heat acquired by the water promotes flashing of liquid water into steam in the third flash vessel 9038. Steam 9066 exits from an upper portion of the third flash vessel 9038 into a conduit for distribution. The steam 9066 may be a coproduct of the ODH reactor system 9000.

[2159] The pressure of the steam 9066 may depend on tire catalyst in third ODH reactor 9006. The reaction temperature associated with the catalyst 9018 in the third ODH reactor 9006 may drive tire temperature of the return water 9064 discharging from the jacket 9032 to the flash vessel 9038. The temperature of the return water 9064 may determine tire pressure at which the liquid water flashes into the steam 9066 discharging from the flash vessel 9038. [2160] In examples, the steam 9066 may generally be saturated steam. The steam 9066 may be subjected to further processing (e.g., heating) to superheat the steam 9066. In certain embodiments, the catalyst 9018 in the third ODH reactor 9006 provides for a reaction temperature at 500 °C or greater, and the steam 9060 is high pressure steam at 600 psig o r greater.

[2161] The amount (and pressure) of the steam 9066 generated may be correlative with the heat generated in the third ODH reactor 9006 and removed by the jacket water fo r temperature control of the third reactor 9006. The removed heat may give the heat of vaporization for the flashing water in the third flash vessel 9038. The amount of heat generated by the third ODH reactor 9008 may be related to the catalyst 9016 employed in the third reactor 9004, the amount (rate) of unreacted alkane converted to the corresponding aikene in the third ODH reactor 9006, and so forth.

[2162] FIG. 91 is an ODH reactor system 9100 having the three ODH reactors 9002, 9004, 9006 discussed above with respect to FIG 90. The system 9100 receives the feed 9020 having an alkane (e.g., ethane). The first ODH reactor 9002 discharges an effluent 9022 having a corresponding aikene and unreacted alkane to the second ODH reactor. The first ODH reactor 9002 provides both (1) catalytic conversion and (2) the effluent 9022 as a preheated feed to the second ODH reactor 9004. The second ODH reactor 9004 discharges an effluent 9024 having the corresponding aikene and unreacted alkane to the third ODH reactor 9006. The second ODH reactor 9004 provides for both (1) catalytic conversion and (2) the effluent 9024 as a preheated feed to the third ODH reactor 9006. The third ODH reactor 9006 discharges an effluent 9026 having the corresponding aikene and any unreacted alkane. The corresponding alkene (e.g., ethylene) in the third-reactor effluent 9026 may be a product of the ODH reactor system 9100.

[2163] In system 9100, the routing of water as the heat transfer fluid is different than in the system 9000 of FIG. 90. The system 9100 lias the flash vessels 9034, 9036, 9038, which may be the same or similar to those depicted in FIG. 90. However, the flash vessels 9034, 9036, 9038 receive input water from the respective reactor jacket instead of directly from the water source 9040. A cascade flow' of water is employed. In other words, the second-reactor jacket 9030 receives water from the first flash vessel 9034. Thus, sortie heat from the first ODH reactor 9002 system may be utilized by the second ODH reactor 9004 system. The third-reactor jacket 9032 receives water from the second flash vessel 9036. Thus, some heat from the first ODH reactor 9002 system and the second ODH reactor 9004 system may be utilized by the third ODH reactor 9006 system. The mass flow rates of the coproduct steam streams 9054, 9060, 9066 may be affected by the alterna tive flow's of the jacket water shown in FIG. 91 as compared to FIG. 90.

[2164] For the first ODH reactor 9002, water is conveyed from the water source 9040 to the first-reactor jacket 9028 to control temperature of the first-reactor process side 9008. In the illustrated embodiment, a pump 9101 (e.g.. centrifugal pump) provides motive force for flow of Site water. For temperature control of the first reactor 9002, the flow rate of the water may be maintained or modulated via a control valve 9102 disposed on the conduit conveying the water. The water flowing through the first-reactor jacket 9028 acquires heat from the first-reactor process side 9008 and discharges from die jacket 9028 as heated water 9104 to the first flash vessel 9034. In implementations, the set point of the control valve 9102 may be specified in response to the temperature of the process side 9008 or the temperature of the heated water 9104. or both. The heat acquired by the water promotes flashing of liquid water into steam in the first flash vessel 9034.

[2165] Steam 9054 discharges overhead from the first flash vessel 9034 into a conduit for distribution. The steam 9054 may be a coproduct of the ODH reactor system 9100. The pressure of the steam 9054 may depend on the cataly st 9010 in first ODH reactor 9002. as discussed. In one example, the pressure of the steam 9054 is low' pressure steam at 150 psig or less. The temperature of the heated water 9104 (and the amount of heat acquired from the first ODH reactor 9002) may determine the pressure at which the liquid water flashes into the steam 9054 discharging from the flash vessel 9034. The amount of heat acquired by the jacket water may be the heat generated by the first ODH reactor 9002 in the ODH catalytic conversion reaction in the first ODH reactor 9002.

[2166] For the second ODH reactor 9004, water 9106 discharges from a bottom postion of the first flash vessel 9034 to the second-reactor jacket 9030. The water 9106 may be conveyed via a pump 9108 (e.g., centrifugal pump). For temperature control of the second reactor 9004, a control valve 9 10 may control the flow' rate of the water 9106. The water 9106 flow's through the second-reactor jacket 9030 and acquires heat from the second-reactor process side 9012. The heated water 9112 exits the jacket 9030 to the second flash vessel 9036. The set point of the control valve 9102 may be set in response to the temperature of the process side 9012 or the temperature of the heated water 9112, or both. The heat acquired by the water promotes flashing of liquid water into steam in the second flash vessel 9036.

[2167] Steam 9060 discharges overhead from the second flash vessel 9036 to be conveyed to users of the steam 9060. The steam 9060 may be a coproduct of the ODH reactor system 9100. The pressure of the steam 9060 may depend on the catalyst 9014 in the second ODH reactor 9004, as discussed in one example, the pressure of the steam 9060 is medium pressure steam in the range of 150 psig to 600 psig. In another example, the pressure of the steam 9060 is high pressure steam at 600 psig or greater.

[2168] The temperature of the heated water 9112 (and the amount of heat acquired from the second ODH reactor 9004) may determine the pressure at which Site liquid water flashes into the steam 9060 discharging from the flash vessel 9036. The amount of heat acquired in Site heated stream 9112 may be at least the heat generated by the second ODH reactor 9004 in the ODH catalytic conversion reaction in the second ODH reactor 9004. The heated water 9112 may also contain heat from the first ODH reactor 9002 system.

[2169] For the third ODH reactor 9006, water 9114 flow's from the bottom portion of the second flash vessel 9036 through the third-reactor jacket 9032 to receive heat from the third-reactor process side 9016. The water 9114 may be transported via a pump 9116 and a control valve 9118. The flow of water 9114 through the jacket 9032 as a heat transfer fluid (jacket water) is for tempera ture control of the third ODH reactor 9006. The heated water 9120 exits the jacket 9032 to the third flash vessel 9038. Optionally, water 9122 as additional jacket water may flow by thermosiphon from a bottom discharge of the flash vessel 9038 to the jacket 9032. The heat acquired by the jacket water promotes flashing of liquid water into steam in the flash vessel 9038.

[2170] Steam 9066 exits from an upper portion of the third flash vessel 9038 into a conduit. The steam 9066 may be a coproduct of the ODH reactor system 9100. The pressure of the steam 9066 may depend on the catalyst 9018 in third ODH reactor 9006, as discussed. In one example, the pressure of the steam 9066 is high pressure steam at 600 psig or greater. The temperature of the heated water 9120 (and the amount of heat acquired from the third ODH reactor 9006) may determine the pressure at which the liquid water flashes into the steam 9066 discharging from the flash vessel 9038. The amount of heat acquired in the heated stream 9120 may be at least the heat generated by the third ODH reactor 9006 in the ODH catalytic conversion reaction in the third ODH reactor 9006. The heated water 9120 may also contain heat from the first ODH reactor 9002 system and the second ODH reactor 9004 system.

[2171] FIG. 92 is an ODH reactor system 9200 having the three ODH reactors 9002, 9004, 9006 discussed above. The first ODFI reactor 9002 performs as a catalytic conversion reactor and also serves as a feed preheater for the second ODH reactor 9004 The second ODH reactor 9004 performs as a catalytic conversion reactor and also serves as a feed preheater for the third ODH reactor 9006.

[2172] Each reactor 9002, 9004, and 9006 may be a reactor vessel having an inlet and an outlet. The inlet may be to a process side of the reactor and the outlet may be from a process side of the reactor. The inlet may be for feed including a hydrocarbon (and oxygen) and the outlet for an effluent including hydrocarbon. The reactor vessel may have a jacket for heat transfer fluid. The jacket may have a jacket inlet to receive heat transfer fluid and a jacket outlet to discharge heat transfer fluid.

[2173] The ODH reactor system 9200 is similar to the ODH reactor system 9100 of FIG. 91, except that the system 9200 does not include the second flash vessel 9036, pump 9116, valve 9118, or tire associated steam 9060 as coproduct stream. The third flash vessel 9038 becomes the second flash vessel in FIG. 92. The amount of the coproduct steam 9066 generated may be increased. As for the jacket water flow, the heated water 9112 that discharges from the second-reactor jacket 9030 flows to the third-reactor jacket 9032. The upstream pump 9108 and control valve 9110 may be sized accordingly. The control valve 9110 may be timed for temperature control of both the second ODH reactor 9004 and the third ODH reactor 9006. The optional thermosiphon flow of water 9122 from the flash vessel 9038 to the jacket 9032 may further facilitate the temperature control in the third reactor 9006.

[2174] The coproduct steam 9066 discharged overhead from the flash vessel 9038 may be heated downstream to superheat the steam 9066. Therefore, the coproduct steam 9066 may be superheated steam. In some implementations, the steam 9066 may be heated in a heat exchanger 9202 (e.g., cross exchanger, shell -and-tube heat exchanger, etc.) by the third-reactor effluent 9026. In one example, the steam 9066 is superheated high-pressure steam at 600 psig or greater.

[2175] FIG. 93 is an ODH reactor system 9300 having die three ODH reactors 9002, 9004, 9006 discussed above. The system 9300 is similar to the ODH reactor system 9200 of FIG. 92, except that steam (with little or no liquid water) is flowed through the third-reactor jacket 9032 as heat transfer fluid.

[2176] The heated wrater 9302 from the second-reactor jacket 9030 is routed to the flash vessel 9038. The heated water 9302 may be liquid water or steam, or a mixture (two-phase flow) of steam and liquid water. For service as a heat transfer fluid, steam 9304 discharges from an upper portion of the flash vessel 9038 through a conduit to the diird-reactor jacket 9032. In examples the sleanr 9304 may be high pressure steam (600 psig or greater) or very high pressure steam (1500 psig or greater). Heated steam exits the third-reactor jacket 9032 into a conduit as coproduct superheated steam 9306 for distribution to users.

[2177] A control valve (not shown) may modulate and control the flow' rate (e.g., mass per time) of the steam 9304 flowing through the jacket 9032 for the temperature control of the third-reactor process side 9016. The control valve (if employed) may be disposed on the jacket 9032 discharge conduit or on the inlet conduit upstream of the jacket 9032. The steam 9306 may be superheated high pressure steam or superheated vesy high pressure steam. [2178] In certain embodiments, the feed 9020 to the first ODH reactor 9002 may be heated (preheated) prior to entry into the first ODFI reactor 9002. For example, the feed 9020 may be routed through a heat exchanger 9310 (e.g., cross exchanger, shell-and-tube heat exchanger, etc.) and heated by the third-reactor effluent 9026 routed through the heat exchanger 9310.

[2179] FIG. 94 is an ODH reactor system 9400 Slaving three reactors that are tire same or similar as the three ODH reactors 9002, 9004, 9006 discussed above. In implementations, the tliree ODH reactors 9002, 9004, 9006 may be fabricated with a heat transfer area (between the process side and jacket) specified based on various factors. Such factors may include the heat transfer fluid that will be utilized, the amount of heat that will be generated by the reactor, the flow' rate of the heat transfer fluid through the jacket, and so forth. For the system 9400 in operation, FIG. 94 depicts the liquid water with gray shading and the iieat transfer fluid with forward slashed lines. Such indications are retained in FIGS. 95-97.

[2180] In the illustrated embodiment of FIG. 94, the heat transfer fluid 9402, 9404, 9406 introduced to the respective reactor jackets 9028, 9030, 9032 may be treated water (e.g., demineralized water, boiler feedwater, etc.), glycol (e.g., ethylene glycol, propylene glycol, etc.), molten salt, or other type of heat transfer fluid. In embodiments with the heat transfer fluid 9402, 9404, 9406 as molten salt, three molten-salt supply systems may be employed to provide molten salt as the heat transfer fluid 9402, 9404, and 9406, respectively.

[2181] Examples of the heat transfer fluid include DOWTHERM™ heat transfer fluids (Dow Chemical Company, Midland, Michigan USA), which may have glycol or synthetic organic compounds generally. Examples of the heat transfer fluid may include DW-Thenn HT products (Huber USA, Gary, North Carolina USA), Syltherm™ silicon fluids (e.g., Sylthenn™ 800) (Dow' Chemical Company. Midland, Michigan USA), and Santolube products (e.g., OS-750™ or OS-124™) (SantoLubes LLC, Spartanburg, South Carolina USA) Of course temperature limitations on these various organic heat-transfer fluids are taken into account

[2182] in certain embodiments, each reactor 9002, 9004, 9006 is associated with a heat exchanger 9408, 9410, 9412 (e.g , shell-and-tube heat exchanger, plate and frame heat exchanger, etc.) that heats water with the heat transfer fluid discharged from the reactor jacket. The water to be heated is pumped via a pump 9042 from a water source 9040 to the three flash vessels 9034, 9036, 9038. Control valves 9044, 9046, 9048 may be disposed on the conduits conveying the water to modulate the respective flow rate of the water to the flash vessels 9034, 9036, 9038. In embodiments, this water is boiler feedwnter. The water is heated in the heat exchanger 9408, 9410, 9412 with the heat transfer fluid (discharged from the reactor jacket) so to flash the water into steam in the flash vessel 9034. 9036. 9038 to generate steam.

[2183] For the first ODH reactor 9002, water 9414 discharges from a bottom portion of the first flash vessel 9034 and is heated in the first heat exchanger 9408. The heated water discharges from the heat exchanger 9408 as return water 9416 to the first flash vessel 9034. This circulatory flow of the water through the heat exchanger 9408 may be by themiosiphon.

[2184] Liquid water in the first flash vessel 9034 flashes into steam 9054 that discharges from the first flash vessel 9034 as coproduct steam. The conditions (e.g., amount, pressure, temperature, etc.) of the steam 9054 may depend on the catalyst 9010 type in the first ODH reactor 9002, the operating (reaction) temperature of the first ODH reactor 9002, the amount of heat generated by the first ODH reactor 9002, and other factors.

[2185] The heat transfer fluid 9402 is heated in the first-reactor jacket 9028 and discharges as heated heat- transfer fluid 9418 to the fi rst heat exchanger 9408. In the first heat exchanger 9408, heat transfer occurs from the heated heat-transfer fluid 9418 to the water 9414. The heat transfer fluid discharges from the heat exchanger 9408 as cooled heat-transfer fluid 9419 to the reactor jacket 9028. In certain implementations, some or all of the cooled heat- transfer fluid 9419 may be returned to the heat-transfer fluid supply system instead of returned to the jacket 9028. [2186] Moreover, a flow' bypass conduit may be provided around the heat exchanger 9408. Thus, a first portion of the heated heat-transfer fluid 9418 may flow through the heat exchanger 9408. A second portion of the heated heat-transfer fluid 9418 bypasses (flow's around) the heat exchanger 9408 through the flow bypass conduit. In examples, the first portion and the second portion may each be in the range of 20 weight percent to 80 weight percent of the heated heat-transfer fluid 9418.

[2187] For the second ODH reactor 9004, the operation of steam genera tion may be similar as with the first ODH reactor 9002 but with the option to generate steam at different pressure. A different pressure steam may be produced, for example, by utilizing a catalyst 9014 in the second ODH reactor 9004 that is different than the catal st 9010 in the first ODH reactor 9002.

[2188] In the steam generation for the second ODH reactor 9004, water 9420 discharges from a bottom outlet of the second flash vessel 9036 and is heated in the second heat exchanger 9410. The heated water exits the heat exchanger 9410 as return water 9422 to the second flash vessel 9034. The motive force for this circulation of water through the second-reactor jacket 9030 may be by thermosiphon. Liquid water in the second flash vessel 9036 vaporizes into steam 9060. The steam 9060 may exit overhead from the second flash vessel 9036 as coproduct steam. The amount, pressure, and temperature of the steam 9060 may depend on the catalyst 9014 type in the second ODH reactor 9004, the operating (reac tion) temperature of the second ODH reactor 9004, the amount of heat generated by the second ODH reactor 9004, and so forth.

[2189J The heat transfer fluid 9404 is heated in the second-reactor jacket 9030 and discharges as heated heat- transfer fluid 9424 from the jacket 9030 to flow through the second heat exchanger 9410. Heat transfer occurs from the heated heat-transfer fluid 9424 to the water 9420. The heat transfer fluid discharges from the heat exchanger 9410 as cooled heat-transfer fluid 9426 to the reactor jacket 9030. in certain implementations, some or ail of the cooled heat-transfer fluid 9426 may be returned to the heat-transfer fluid supply system instead of to the jacket 9030. In addition, in some embodiments, a portion of the heated heat-transfer fluid 9424 may flow around the heat exchanger 9410 via a flow bypass conduit in parallel with the heat exchanger 9410. In examples, the portion routed through the bypass may be at least 20 weight percent of the heated heat-transfer fluid 9424, at least 50 weight percent of the heated heat-transfer fluid 9424, or at least 75 weight percent of the heated heat-transfer fluid 9424. [2190] For the third ODH reactor 9006, the operation of steam generation may be similar as with the first ODH reactor 9002 and the second ODH reactor 9004 but with the option to generate steam at different pressure. A different pressure steam may be generated, for instance, by utilizing a catalyst 9018 in the third ODH reactor 9006 that is different than the catalyst 9010 in the first ODH reactor 9002 and different than the catalyst 9014 in the second ODH reactor 9004.

[21911 To produce steam with the third ODH reactor 9006, water 9428 discharges from a bottom outlet of the third flash vessel 9038 and is heated in the third heat exchanger 9412. The heated water exits the heat exchanger 9412 as return water 9430 to the third flash vessel 9038 via thermosiphon in this example. Liquid water in the third flash vessel 9038 vapo rizes into steam 9066, which may discharge from an upper portion of the third flash vessel 9038 as coproduct steam. The amount pressure, and temperature of the steam 9060 may be correlative with a number of factors. For example, the factors may include the third-reactor catalyst 9018 type the reaction temperature on the third-reactor process side 9016, and the amount of heat generated by the reaction on the process side 9016

[2192] The heat transfer fluid 9406 is heated in the third-reactor jacket 9032 and discharges as heated heat- transfer fluid 9432 to the third heat exchanger 94f 2. Heat transfer occurs from the heated heat-transfer fluid 9332 to the water 9428 in the third heat exchanger 9412. The heat transfer fluid discharges from the heat exchanger 9412 as cooled heat-transfer fluid 9434 to the reactor jacket 9032. Some or all of the cooled heat-transfer fluid 9434 may return to the heat-transfer fluid supply system instead of sent to the jacket 9032. Moreover, in embodiments, a portion of the heated heat-transfer fluid 9432 may flow through a conduit to bypass the heat exchanger 9412,

[2193] FIG. 95 is an ODH reactor system 9500 having three reactors the same or similar as the three ODH reactors 9002, 9004, 9006 discussed above. The three ODH reactors 9002, 9004. 9006 have a first catalyst 9010, a second catalyst 9014, and a third catalyst 9018, respectively. The first catalyst 9010, second catalyst 9014, and third catalyst 9018 may each be a fixed bed of catalyst and be different or same catalyst type with respect to each other. [2194] The system 9500 receives the feed 9020 having an alkane (e.g.. ethane). The first ODH reactor 9002 discharges an effluent 9022 having a corresponding alkene and unreacted alkane. The second ODH reactor 9004 receives the first-reactor effluent 9022 and discharges an effluent 9024 having more corresponding alkene (and less unreacted alkane) than the first-reactor effluent 9022. The third ODH reactor 9006 receives the second-reactor effluent 9024 and discharges an effluent 9026 having the corresponding alkene and any remaining unreacted alkane. The corresponding alkene (e.g., ethylene) in the third-reactor effluent 9026 may be a product of the ODH reactor system 9500.

[2195] The system 9500 utilizes a heat transfer fluid as discussed with respect to FIG. 94. However, in comparison to the system 9400, aspects of the steam generation are altered in system 9500, as indicated below. For instance, in the example of FIG. 95, the second flash vessel 9036 is not employed. i2196| For system 9500, Site water (e.g.. boiler feedwater) to be heated to generate steam is pumped from the water source 9040 via the pump 9042 to the first flash vessel 9034. The control valve 9048 disposed along the conduit conveying the water may maintain or adjust the flow rate of the water to the first flash vessel 9034. Water is fed from the first flash vessel 9034 to both the first heat exchanger 9408 (associated with the first ODH reactor 9002) and the second heat exchanger 9410 (associated with the second ODH reactor 9004). Water 9414 is fed from the first flash vessel 9034 to the first heat exchanger 9408 by thermosiphon. Water 9502 is fed from the first flash vessel 9034 to the second heat exchanger 9410 via a pump 9504. A control valve 9506 disposed along the conduit conveying the water 9502 may maintain or adjust the flow rate of the water 9502 through the second heat exchanger 9410. The flow rate of water 9502 through the second heat exchanger 9410 may be modulated by the control valve 9506 for temperature control of the second reactor 9004.

[2197] The water 9414 in the first heat exchanger 9408 is heated with the heated heat-transfer fluid 9418 from the first reactor jacket 9028. Thus, in this reactor temperature control, the water 9414 receives the heat generated by the first-reactor 9002 reaction. The heated water discharges from the first heat exchanger 9408 as return water 9416 to the first flash vessel 9034. Liquid water in the first flash vessel 9034 flashes into steam 9054 that exits overhead from the first flash vessel 9034 as coproduct steam. In some examples, the steam 9054 is saturated low-pressure steam at 150 psig or less.

[2198] The water 9502 in the second heat exchanger 9410 is heated with the heated heat-transfer fluid 9424 from the second reactor jacket 9030. in other words, for reactor temperature control, the water 9502 receives the heat generated by the second-reactor 9004 reaction. The heated water 9508 discharges from the second heat exchanger 9410 to the flash vessel 9038. The heated water 9508 flowing through the conduit to the flash vessel 9038 may be liquid water or steam, or a mixture thereof (two-phase flow).

[2199] Steam 9510 discharges from an overhead outlet of the flash vessel 9038 is routed through the third heat exchanger 9412 for temperature control of the third ODH reactor 9006. The steam 9510 is heated in the third heat exchanger 9412 with the heated heat-transfer fluid 9432 from the third-reactor jacket 9032. The pressure of the steam 9510 may be different than the pressure of the steam 9054 discharged from the first flash vessel 9034. The pressure may be different due at least in part to utilizing a catalyst 9014 in the second ODH reactor 9004 and a catalyst 9018 in the third ODH reactor 9006 that are different Ilian the catalyst 9010 in the first ODH reactor 9002.

In implementations, the flow rate of the steam 9510 through the heat exchanger 9412 may be modulated by a control valve (not shown) for temperature control of the third reactor 9006. [2200] The heated steam discharges from the third heat exchanger 9412 as superheated steam 9512, which may be a coproduct of die ODH reactor system 9500. In implementations, a control valve (not shown) may modulate the flow' of the steam 9510 (or steam 9512) to control the temperature of tire third reactor 9006. in some examples, the superheated steam 9512 is high pressure steam at 600 psig or grea ter, or very high pressure steam at 1500 psig or greater. In implementations, temperature of the steam 9512 is at least 150 °C above tire saturation temperature, or at least 200 °C above the saturation temperature in one example, tire superheated steam 9512 has a pressure of about 1500 psig and a temperature of at least 500 °C or a t least 510 °C. In another example, the superheated steam 9512 lias a pressure of about 2000 psig and a temperature of at least 530 °C or at least 540 °C.

[220 ί I Referring to FIG. 94 and FIG. 95. a portion of each of the heated heat-transfer fluids 9418, 9424, 9432 that discharge from a reactor jacket may bypass die respective heat exchangers 9408, 9410, 9412. See, for example, tiie dashed line 9514 in FIG. 95. Moreover, a portion of each of die cooled heat-transfer fluids 9419, 9426, 9434 discharging from the respective heat exchangers 9408. 9410. 9412 may be returned to the heat-transfer fluid supply system instead of sent to die reactor jacket. See, for example, the dashed line 9516 in FIG. 95.

[2202] The heat-transfer fluid supply sy stem(s) that provides the heat transfer fluid 9402. 9404, 9406 may have a heat exchanger(s) for temporary' operation to heat the heat transfer fluid 9402, 9404, 9406 supply. The heat exchanger may be put into operation when the three reactors 9002, 9004, 9006 are not performing the catalytic reaction of alkane to aikene and thus are not generating heat. This temporary operation of the heat exchanger may therefore provide heat for steam generation in the ODH reactor system 9400, 9500.

[2203] FIG. 96 is an ODH reactor system 9600 having the three ODH reactors 9002, 9004, 9006 in series each having a catalyst 9010, 9014, 9018. As with the preceding figures, the system 9600 receives the feed 9020 having an alkane (e.g., ethane). The first ODH reactor 9002 converts (via catalyst 9010) some of the alkane to a corresponding aikene (e.g., ethylene). The first ODH reactor 9002 discharges an effluent 9022 having the corresponding aikene and unreacted alkane to the second ODH reactor 9004. The second ODH reactor 9004 converts (via cataly st 9014) some of the unreacted alkane to the corresponding aikene. The second ODH reactor 9004 discharges an effluent 9024 having the corresponding aikene and unreacted alkane to the third ODH reactor 9006. The third ODH reactor 9006 converts (via catalyst 9018) some unreacted alkane to the corresponding aikene. The third ODH reactor 9006 discharges an effluent 9026 having the corresponding aikene and any unreacted alkane. The corresponding aikene in the third-reactor effluent 9026 may be a product of the ODH reactor system 9600.

[2204] The system 9600 includes the three heat exchangers 9408, 9410, 9412 associated with the three ODH reactors 9002, 9004, 9006. The three heat exchangers 9408, 9410, 9412 may be the same or similar as in FIGS. 94 and 95. The three heat exchangers 9408, 9410, 9412 may be configured differently than in FIGS. 94 and 95 because of the different arrangement for source of water flow through the hea t exchangers 9408, 9410, 9412,

[2205] Heat transfer fluid 9402 is supplied to the first-reactor jacket 9028. In the jacket 9028, the heat transfer fluid 9402 receives the heat generated by the ODH catalytic reaction in first ODH reactor 9002. Heated heat-transfer fluid 9418 discharges from the jacket 9028 through heat exchanger 9408 and may return as cooled heat-transfer fluid 9419 to die jacket 9028. [2206] Heat transfer fluid 9404 is supplied to the second-reactor jacket 9030 where the heat transfer fluid 9404 receives the heat generated by the ODH catalytic reaction in second ODH reactor 9004. Heated heat-transfer fluid 9424 discharges from the jacket 9030 through heat exchanger 9410 and may return as cooled heat-transfer fluid 9426 to the jacket 9030.

[2207] Heat transfer fluid 9406 is supplied to the third-reactor jacket 9032 where the heat transfer fluid 9406 receives the heat generated by the ODH catalytic reaction in third ODH reactor 9006. Heated heat-transfer fluid 9432 discharges from the jacket 9032 through heat exchanger 9412 and may return as cooled heat-transfer fluid 9434 to the jacket 9032.

[2208] In the illustrated embodiment, the system 9600 as depicted employs a single flash vessel 9038. Water to be heated to generate steam is supplied to the flash vessel 9038 from a water source 9040 via pump 9042. The water source 9040 may include a vessel holding the water to be supplied. The water may be boiler feedwater. A control valve 9044 on the conduit conveying the water to the flash vessel 9038 may control the flow rate of the water from the source 9040 to the flash vessel 9038.

[2209] Water is provided from the flash vessel 9038 to all three depicted heat exchangers 9408, 9410, 9412. The total flow of water discharging from the bottom po rtion of the flash vessel 9038 gives the water 9414 stream for the inlet to the first heat exchanger 9408, the water 9602 stream for the inlet to the second heat exchanger 9410, and the water 9604 stream for the inlet to the third heat exchanger 9412. The water streams flow through the respective heat exchangers 9408, 9410, 9412 and return to the flash vessel 9038. The motive force for the three circulation loops of water may be thermosiphon or pump.

[2210] The water 9414 is heated in the first heat exchanger 9418 with heat from the heated heat-transfer fluid 9418 discharged from the first-reactor jacket 9028. The water 9602 is heated in the second heat exchanger 9408 with heat from the heated heat-transfer fluid 9424 discharged from the second-reactor jacket 9030. The water 9604 is heated in the third heat exchanger 9412 with heat from the heated heat-transfer fluid 9432 discharged from the third- reactor jacket 9032.

[2211 ] The heat acquired by these water streams that return to the flash vessel promotes (contributes to) the flashing of liquid water in the flash vessel 9038 to generate steam 9606 that may be a coproduct of the ODH reactor s stem 9600. in embodiments, the steam 9606 is high pressure steam at 600 psig or greater. In those embodiments, a portion of the steam 9606 may be diverted and let down in pressure via a control valve 9608 to give low pressure steam 9610 at 150 psig or less. Similarly, a portion of the steam 9606 may be diverted and let down in pressure via a control valve 9612 to give medium pressure steam 9614 in the range of 150 psig to 600 psig or in the range of 250 psig to 400 psig. Both the low pressure steam 9610 and the medium pressure steam 9614 may be coproducts of the ODH reactor system 9600.

[2212] FIG. 97 is an ODFI reactor system 9700 that is the same as the ODH reactor system of FIG. 96, except that a heat exchanger 9702 is incorporated into the system 9700 to superheat the high pressure steam 9606 so that tire coproduct with be superheated steam 9704 (e.g., superheated high-pressure steam at 600 psig or greater). In the illustrated embodiment, heat transfer fluid from third-reactor jacket 9032 is circulated through the heat exchanger 9702 to heat the steam 9606. In other embodiments, the third-reactor effluent 9026 is instead sent through the heat exchanger 9702 as the heat transfer fluid to heat the steam 9606. The heat exchanger 9702 may be a shell-and-tube heat ex ianger or other type of heat exchanger.

[2213] Referring to FIGS. 90-97, the feed 9020 to the first ODH reactor 9002 may he heated (preheated) prior to entry into the first ODH reactor 9002. For instance, the feed 9020 may be routed through a heat exchanger (e.g., cross exchanger, shell-and-tube heat exdianger, etc.) and heated by the third-reactor effluent 9026 routed through the heat exchanger. See, for example, heat exchanger 9310 in FIG. 93. In addition, at least one of the steam 9054, the steam 9060, or the steam 9066 may be heated in a heat exdianger (e.g., cross exchanger, shell-and-tube heat exchanger, etc.) by the third-reactor effluent 9026. The steam 9054, steam 9060, or steam 9066 may be heated to superheat the steam or further superheat the steam. See, for example the heat exchanger 9202 in FIG. 92 that heats the steam 9066 to give superheated steam 9066. The various embodiments for configuration of the ODH reactor system and steam generation discussed with respect to FIGS. 90-97 are given as examples. Other examples for steam generation by the ODH reactor system will be readily apparent to one of ordinary skill in the art with the benefit of the provided techniques.

[2214] Referring to FIGS. 90-97, the catalysts 9010, 9014, 9018 may be the same or different catalyst type. Catalyst types may give different catalytic reaction temperatures in the conversion of the lower alkane to the corresponding alkene. For instance, one catalyst type (labeled, for example, as a low temperature catalyst) may give a catalytic reaction temperature of less than 400 °C. Another catalyst type (labeled, for example, as a medium temperature catalyst) may give a catalytic reaction temperature of at least 400 °C (e.g., in the range of 400 °C to 500 °C). Yet another catalyst type (labeled for example, as a high temperature catalyst) may give a catalytic reaction temperature of at least 500 °C. These catalyst types may be ODH catalyst.

[2215] In some embodiments, the first catalyst 9010 is lory temperature catalyst (e.g reaction temperature less than 400 °C), the second catalyst 9014 is medium temperature catalyst (e.g., reaction temperature of at least 400 °C or in the range of 400 °C to 500 °C), and the third catalyst 9018 is high temperature catalyst (reaction temperature at least 500 °C). In other embodiments, the first catalyst 9010 is lory temperature catalyst and both the second catalyst 9014 and third catalyst 9018 are high temperature catalyst. In yet other embodiments, all three catalysts 9010, 9014, 9018 are high temperature catalyst. Other combinations of catalyst types are applicable among the three catalysts 9010, 9014, 9018.

[2216] The low' temperature catalyst may be conducive for generating lory pressure steam (e.g., less than 150 psig). The medium temperature catalyst may be conducive for generating medium pressure steam (e.g., 150 psig to 600 psig). The high temperature catalyst may be conducive for generating high pressure steam (e.g., at least 600 psig). However, each catalyst type may generate heat in its catalytic reaction that can contribute to steam generation of different stream pressures or qualities.

[2217] An example of the low temperature catal st is a catalyst that includes molybdenum, vanadium, tellurium, niobium, and oxygen wherein the molar ratio of molybdenum to vanadium is from i :Q.12 to 1 :0.49, the molar ratio of molybdenum to tellurium is from 1 :0.0i to 1 :0.30, the molar ratio of molybdenum to niobium is from 1:0.01 to 1:0.30, and oxygen is present at least in an amount to satisfy the valency of any present metal oxides. The molar ratios of molybdenum, vanadium, tellurium, niobium can be determined by inductively coupled plasma mass spectrometry (ICP-MS). Any other catalyst described herein may be used.

[2218] An example of a low temperature catalyst that provides for the GDH reaction to at a temperature of less than 400 °C is a mixed metal oxide having the formula Mo _a V _b Te _c Nb _d Pd _e Or, where a, b, c, d, e, and f subscripts are relative atomic amounts of the elements Mo, V, Te, Mb, Pd, O, respectively. When a=l, thenb=0.Gl to 1.0, 0=0.01 to 1.0, d=0.Gl to 1.0, 0.00<e<0.10, and f is a number to satisfy the valence state of the catalyst.

[2219] Referring to FIGS. 90-93, for certain cases of the first reactor 9002 having a low-temperature ODH catalyst (e.g., providing for an ODH reaction at a temperature of about 400 °C or less), high pressure steam can be generated by the first-reactor 9002 system at tire first flash vessel 9034. However, for other cases of the first reactor 9002 having a low-temperature ODH catalyst, the reactor process side temperature may not be adequate to efficiently drive formation of high pressure steam. In other words, tire available temperature difference (DT) or available log-mean temperature difference (LMTD or ATLM) for between the reactor jacket water and the reactor process side may not be adequate to effectively provide for formation of high pressure steam.

[2220] Referring to FIGS. 94 and 93, in some examples with the first reactor 9002 having a low-temperature ODH catalyst (e.g., providing for an ODH reaction at a temperature of about 400 °C or less) high pressure steam can be generated by the first-reactor 9002 system at the first flash vessel 9034. However, for other examples of the first reactor 9002 having a lo -temperature ODH catalyst the reactor process side temperature may not be adequate to efficiently drive formation of high pressure steam. In other words, the available DT or or LMTD for the temperature difference between the water 9414, 9416 and heal transfer fluid 9418, 9419 (e.g., molten salt) in the heat exchanger 9408 may not be adequate to effectively provide for formation of high pressure steam.

[2221] Referring to FIGS. 90-93, implementations employing water (e.g., boiler feedwater) as the heat transfer fluid may be applicable for reactors 9002, 9004, 9006 of small scale. To implement reactors 9002, 9004, 9006 that are large and operating with water at contemplated pressures and temperatures on the reactor jackets 9028, 9030, 9032 (e.g., for the reactor having medium temperature ODH catalyst or high temperature ODH cataly st) may lead to very large or even impractical wall thicknesses of the reactor jackets 9028, 9030, 9032. Therefore, for large reactors, a heat transfer fluid, such as molten salt, may be beneficial for the reactor jackets in certain instances. See for example, FIGS. 94-97.

[2222] Referring to FIGS. 90-97, the first ODH reactor 9002, the second ODH reactor 9004, and the third ODH reactor 9006 may each be a fixed-bed reactor or tubular fixed-bed reactor. For a fixed-bed reactor, reactants may be introduced into the reactor at one end and flow past an immobilized catalyst. Products are formed and an effluent having the products may discharge at the other end of the reactor. The fixed-bed reactor may have one or more tubes (e.g., ceramic tubes) each having a bed of catalyst and for flow of reactants (e.g., lower alkane or ethane) and products (e.g., corresponding alkene or ethylene). The tubes may include for example, a steel mesh. Moreover, a cooling jacket adjacent foe tube(s) may provide for temperature control of the reactor.

[2223] In oilier embodiments the first ODH reactor 9002, foe second ODH reactor 9004, and the third ODH reactor 9006 may each be a fluidized bed reactor. In implementations, a fluidized bed reactor may have a support for the catalyst. The support may be a porous structure or distributor plate and disposed in a bottom portion of the reactor. Reactants may flow upward through the support at a velocity to fluidize the bed of catalyst (e.g., the catalyst rises and begins to swirl around in a fluidized manner). The reactants are converted to products upon contact with the fluidized catalyst. An effluent having products may discharge from an upper portion of the reactor. A cooling jacket may facilitate temperature control of the reactor.

[2224] Lastly, the temperature referenced for each reactor (e.g., first temperature in the first reactor, second temperature in the second reactor, and third temperature in the third reactor) may be the temperature at which the respective catalyst drives the oxidative deh drogenation and may be labeled as the catalyst temperature. The temperature referenced may be the reactor operating temperature driven by the catalytic oxidative dehydrogenation. The temperature referenced may be the weighted average temperature of the reactor or reactor catalyst bed. e.g.. over the temperature profile from reactor inlet to reactor outlet. The temperature referenced may be or otherwise incorporate reactor peak temperatures, heat-transfer fluid temperature, temperature of steam generated, and so forth. [2225] FIG. 98 is a method 9800 of catalytic oxidative dehydrogenation. The method 9800 may be a method of operating an ODH reactor system having at least two ODH reactors disposed in series. Three ODH reactors are discussed with respect to the method 9800. However, the present methods may accommodate ODH reactors systems having more than three ODH reactors (e.g., four ODH reactors, five ODH reactors, etc.) in series for the generation of steam. The final ODH reactor in the series may discharge an effluent having a product alkene of the ODH reactor system.

[2226] At block 9802, the method includes contacting a feed having a lower alkane with a first ODH catalyst in a first reactor at a first temperature (e.g., less than 400 °C) to dehydrogenate the lower alkane into a corresponding alkene and to heat a first heat-transfer fluid flowing through a first-reactor jacket to facilitate generation of steam. The first catalyst may be disposed as a fixed catalyst bed in the first reactor.

[2227] At block 9804, the method includes contacting a first-reactor effluent from the first reactor with a second ODH catalyst in a second reactor at a second temperature greater than the first temperature to dehydrogenate unreacted lower alkane from the first-reactor effluent into the corresponding alkene and to heat a second heat- transfer fluid flowing through a second-reactor jacket to facilitate generation of steam. The second catalyst may be disposed as a fixed catalyst bed in the second reactor.

[2228] At block 9806, the method includes contacting a second-reactor effluent from the second reactor with a third ODH catalyst in a third reactor at a third temperature (e.g., at least 500 °C) greater than the first temperature to dehydrogenate unreacted lower alkane from the second effluent into the corresponding alkene and to heat a third heat-transfer fluid flowing through a third-reactor jacket to facilitate generation of steam. The third catalyst may be disposed as a fixed catalyst bed in the third reactor.

[2229] For the heat transfer fluid as water, the method may include discharging the first heat-tiansfer fluid from the first-reactor jacket to a first flash vessel and discharging low pressure steam at 150 psig or less from the first flash vessel. See, e.g., FIGS. 90-93. The method may include discharging the second heat-transfer fluid from the second-reactor jacket to a second flash vessel and discharging medium pressure steam in the range of 150 psig to 600 psig from the second flash vessel. See, e.g., FIGS. 90-91. The method may include discharging die third heat- transfer fluid from the third-reactor jacket to a third flash vessel and discharging high pressure steam at 600 psig or greater from the third flash vessel. See, e.g., FIGS. 90-91. The method may include discharging water from the first flash vessel as the second heat-transfer fluid to the second-reactor jacket, and discharging water from the second flash vessel as the third heat-transfer fluid to the third-reactor jacket. See, e.g., FIG. 91. in other embodiments, the me thod may include discharging the second heat-transfer fluid from the second-reactor jacket to the third reactor- jacket as the third heat-transfer fluid, and discharging the third heat-transfer fluid from the third-reactor jacket to a second flash vessel and discharging high pressure steam at 600 psig or greater from the second flash vessel. See, e.g., FIG. 92. In yet other embodiments, the method may include discharging the second heat-transfer fluid from the second-reactor jacket to a second flash vessel, and discharging high pressure steam at 600 psig or greater from the second flash vessel through tire third-reactor jacket as the third heat-transfer fluid to superheat the high pressure steam. See. e.g, FIG. 93.

[2230] At block 9808, Site method includes discharging a third-reactor effluent having the corresponding alkene from the third reactor. The corresponding alkene in the third-reactor effluent may be a product of the ODH reactor system having the third reactor second reactor, and fourth reactor. In some embodiments, the alkane is ethane and the corresponding alkene is ethylene. The first reactor, the second reactor, and the third reactor may each be labeled as an ODH reactor and may each be a tubular fixed-bed reactor. As mentioned, the heat transfer fluid may be water. The heat transfer fluid may be treated water (e.g, demineralized water, boiler feedwater, etc.), synthetic organic compounds, glycol (e.g, ethylene glycol, propylene glycol, etc.), or molten salt.

[22311 The method may include discharging the first heat-transfer fluid from the first-reactor jacket to a first heat exchanger and heating, via the first heat exchanger, a first water with the first heat-transfer fluid from the first- reactor jacket. Similarly, the method may include discharging the second heat-transfer fluid from the second-reactor jacket to a second heat exchanger and heating, via the second heat exchanger, a second water with the second heat- transfer fluid from the second-reactor jacket. Likewise, the method may include discharging the third heat-transfer fluid from the third-reactor jacket to a third heat exchanger and heating, via the third heat exchanger, a third water with the third heat-transfer fluid from the first-reactor jacket. See, e.g, FIGS. 94-97. The method may include discharging the first water as heated from the first heat exchanger to a first flash vessel and discharging low pressure steam at 150 psig or less from the first flash vessel. The method may include discharging the second water as heated from the second heat exchanger to a second flash vessel. See, e.g, FIGS. 94-5. The method may include discharging medium pressure steam in the range of 150 psig to 600 psig from the second flash vessel, and discharging the third water as heated from the third hea t exchanger to a third flash vessel and discharging high steam at 600 psig or greater from the third flash vessel. See, e.g, FIG. 94. The method may include discharging high pressure steam at 600 psig or greater from the second flash vessel as the third water through the third heat exchanger to superhea t the high pressure steam. See, e.g, FIG. 95.

[2232] In sortie embodiments, the method may include discharging the first water as heated by the first heat exchanger to a flash vessel, discharging the second water as heated by the second heat exchanger to the flash vessel, discharging the third water as heated by the third heat exchanger to the flash vessel, and discharging high pressure steam at 600 psig or greater from the flash vessel. See, e.g, FIGS. 96-97. The method may include diverting a portion of Site high pressure steam through a control valve to reduce pressure of the portion to medium pressure steam in a range of 150 psig to 600 psig. The method may include divesting a postion of the high pressure steam through a control valve to reduce pressure of the portion to low pressure steam at 150 psig or less. The method may include superheating the high pressure steam in a heat exchanger with heat from the third heat-transfer fluid or front a third-reactor effluent discharged from the third reactor. See, e.g., FIG. 97. The third-reactor effluent may include the corresponding alkene as a product of the ODH reactor system.

[2233] An embodiment is an ODH reactor system (e.g., FIGs. 90-97) including a first reactor having a first ODH catalyst to dehydrogenate a lower alkane to a corresponding alkene at a first temperature and facilitate generation of steam. The first reactor lias a first-reactor jacket for heat transfer. The ODH reactor s stem includes a second reactor having a second ODH catalyst to dehydrogenate unreacted lower alkane in a first-reactor effluent from the first reactor to the corresponding alkene at a second temperature greater than the first temperature and facilitate generation of steam. The second reactor lias a second-reactor jacket for heat transfer. The ODH reactor system includes a third reactor having a third ODH catal st to dehydrogenate unreacted lower alkane in a second- reactor effluent from the second reactor to the corresponding alkene at a third temperature greater than the second temperature and facilitate generation of steam. The third reactor lias a third-reactor jacket for heat transfer. In implementations, the first ODH catalyst is in a fixed-bed in the first reactor, the second ODH catalyst is in a fixed- bed in the second reactor, the third ODH catalyst is in a fixed-bed in the third reactor, and the third reactor to discharge a third-reactor effluent having the corresponding alkene. In implementations, the first ODH catalyst is different than the second ODH catalyst and the third ODH catalyst, and the second ODH catalyst is different than third ODH catalyst. In some examples, the first temperature is less than 400 °C and the first reactor to facilitate generation of low pressure steam at 150 psig or less, the second temperature is at least 400 °C, and the third temperature is at least 500 °C and the third reactor to facilitate generation of high pressure steam of at least 600 psig or very high pressure steam at 1500 psig or greater.

[2234] The ODH reactor sy stem may include a first flash vessel (e.g., FIGS. 90-93) to receive jacket water from the first-reactor jacket and discharge low pressure steam at 150 psig or less. If so, the ODH reactor system may include a second flash vessel (e.g., FIGS. 90-91) to receive jacket water from the second-reactor jacket and discharge medium pressure steam in a range of 150 psig to 600 psig or discharge high pressure steam at 600 psig or greater. A third flash vessel (e.g., FIGS. 90-91) may receive jacket water from the third-reactor jacket and discharge high pressure steam at 600 psig or greater or discharge very high pressure steam at 1500 psig or greater in other examples, a second flash vessel (e.g., FIG. 92) receives jacket water front the third-reactor jacket and discharges high pressure steam at 600 psig or greater. In yet other examples, a second flash vessel (e.g., FIG. 93) receives jacket water from the second-reactor jacket and discharges high pressure steam at 600 psig or greater (or very high pressure steam at 1500 psig or greater) and discharges the steam through the third-reactor jacket to superheat the steam.

[2235] The ODH reactor s stem (e.g., FIGs. 94-97) may include: a first heat exchanger to heat a first water with heat transfer fluid from the first-reactor jacket; a second heat exchanger to heat a second water with heat transfer fluid from the second-reactor jacket; and a third heat exchanger to heat a third water with heat transfer fluid front the third-reactor jacket. In certain implementations (e.g.. FIG. 94), the ODH reactor system includes: a first flash vessel to receive the first water as heated from the first heat exchanger and discharge low pressure steam at 150 psig or less; a second flash vessel to receive the second water as heated from the second heat exchanger and discharge medium pressure steam in the range 150 psig to 600 psig or high pressure steam at 600 psig or greater; and a third flash vessel to receive the third water as heated from the third heat exchanger and discharge high pressure steam at 600 psig or greater or very high pressure steam at 1500 psig or greater. In some implementa tions (e.g., FIG. 95), the ODH reactor system includes: a first flash vessel to receive the first water as heated from the first heat exchanger and discharge low pressure steam at 150 psig or less; and a second flash vessel to receive the second water as heated from the second heat exchanger and discharge high pressure steam at 600 psig or greater as the third water through the third heat exchanger to superheat the high pressure steam. In other implementations (e.g., FIGS. 96-97), the ODH reactor system includes a flash vessel to receive the first water as heated by the first heat exchanger, the second water as heated by the second heat exchanger, and the third water as heated by the third heat exchanger, and discharge high pressure steam at 600 psig or greater. If so, the system (e.g.. FIGs. 96-97) may include a control valve to reduce pressure of a portion of the high pressure steam to medium pressure steam in a range of 150 psig to 600 psig, and also include a control valve to reduce pressure of a portion of the high pressure steam to low pressure steam at 150 psig or less. The ODH reactor system (e.g., FIG. 97) may have a heat exchanger to superheat the high pressure steam with heat from the heat transfer fluid from the third-reactor jacket or with heat from a third-reactor effluent discharged from the third reactor. The third-reactor effluent may have the corresponding aikene as a product of the ODH reactor system.

[2236] Another embodiment is a system for oxidative deh drogenation, including a first reactor having first ODH catalyst to dehydrogenate an alkane at a first temperature. The alkane may have a number of carbons in a range of two carbons to six carbons. The first reactor has a first-reactor jacket to heat a first heat-transfer fluid flowing through the first-reactor jacket to facilitate generation of the steam. The sy stem includes a second reactor having a second ODH catalyst to dehydrogenate unreacted alkane from the first reactor at a second temperature greater than the first temperature. The second reactor lias a second-reactor jacket to heat a second heat-transfer fluid flowing through the second-reactor jacket to facilitate generation of steam. A third reactor lias a third ODH catalyst to dehy drogenate unreacted alkane from the second reactor at a third temperature greater than the first temperature. The third reactor has a third-reactor jacket to heat a third heat-transfer fluid flowing through the third-reactor jacket to facilitate generation of steam. The third ODH catalyst and the second ODH catalyst are different than the first ODH catalyst. The third ODH catalyst may be different titan the second ODH catalyst. Moreover, the first reactor, the second reactor, and the third reactor may each be a tubular fixed-bed reactor. Lastly, the system for oxidative dehydrogenation may include an ODH reactor sy stem having the first reactor, the second reactor, and the third reactor, and wherein the ODH reactor system to generate high pressure steam at 600 psig or greater, or very high pressure steam at 1500 psig or greater.

[2237] Yet another embodiment is a method of oxidative dehydrogenation. The method (e.g., FIGs. 90-98) includes: contacting a feed having a lower alkane (e.g., ethane) with a first ODH catalyst in a first reactor at a first temperature (e.g., less than 400 °C) to dehydrogenate the lower alkane into a corresponding aikene (e.g., ethylene) and to heat a first heat-transfer fluid flowing through a first-reactor jacket to facilitate generation of steam; contacting a first-reactor effluent from the first reactor with a second ODH catalyst in a second reactor at a second temperature (e.g., at least 400 °C) greater than the first temperature to dehydrogenate unreacted lower alkane from the first-reactor effluent into the corresponding alkene and to heat a second heat-transfer fluid flowing through a second-reactor jacket to facilitate generation of steam; and contacting a second-reactor effluent from the second reactor with a third ODH catalyst in a third reactor at a third temperature (e.g., at least 500 °C) greater than the first temperature to dehydrogenate unreacted lower alkane from the second effluent into the corresponding alkene and to heat a third heat-transfer fluid flowing through a third-reactor jacket to facilitate generation of steam. In examples, the first temperature is less than 400 °C, the second temperature is at least 400 °C (or in the range of 400 °C to 500 °C), and the third temperature is at least 400 °C. The method may include discharging a third-reactor effluent from the third reactor wherein the third-reactor effluent includes the corresponding alkene. In implementations, the first reactor, the second reactor, and the third reactor are each a tubular fixed-bed reactor.

[2238] The method (e.g., FIGs. 90-93 and 98) may include discharging the first heat-transfer fluid from the first-reactor jacket to a first flash vessel (wherein the first heat-transfer fluid is water or primarily water) and discharging low' pressure steam at 150 psig or less from the first flash vessel. If so, the method (e.g., FIGs. 90-91 and 98) may include: discharging the second heat-transfer fluid (water or primarily water) from the seco d-reactor jacket to a second flash vessel; discharging medium pressure steam in the range of 150 psig to 600 psig from the second flash vessel; discharging the third heat -transfer fluid (water or primarily water) from the third-reactor jacket to a third flash vessel; and discharging high pressure steam at 600 psig or greater (or very high pressure steam at 1500 psig or greater) from the third flash vessel. In certain implementations (e.g., FIGs. 91 and 98) the method includes discharging water from the first flash vessel as the second heat-transfer fluid to the second-reactor jacket; and discharging water from the second flash vessel as the third heat-transfer fluid to the third-reactor jacket. In other implementations (e.g. FIGs. 92 and 98), the method includes: discharging the second heat-transfer fluid from the second-reactor jacket (water or primarily water) to the third reactor-jacket as the third heat-transfer fluid; discharging the third heat-transfer fluid from the third-reactor jacket to a second flash vessel; and discharging high pressure steam at 600 psig or greater (or very high pressure steam at 1500 psig or greater) from the second flash vessel. In yet other implementations (e.g., FIGs. 93 and 98), the method include: discharging the second heat- transfer fluid from the second-reactor jacket to a second flash vessel; and discharging high pressure steam at 600 psig or greater (or very high pressure steam at 1500 psig or greater) from the second flash vessel through the third- reactor jacket as the third heat-transfer fluid to superheat the high pressure steam (or very high pressure steam). [2239] In some implementations (e.g., FIGs. 94-98), the method includes: discharging the first heat-transfer fluid from the first-reactor jacket to a first heat exchanger and heating, via the first heat exchanger, a first water with the first heat-transfer fluid from the first-reactor jacket; discharging the second heat-transfer fluid from the second- reactor jacket to a second heat exchanger and heating, via the second heat exchanger, a second water with the second heat-transfer fluid from the second-reactor jacket; and discharging the third heat-transfer fluid from the third-reactor jacket to a third heat exchanger and heating, via the third heat exchanger, a third water with the third heat-transfer fluid from the first-reactor jacket. The method (e.g., FIGs. 94-95 and 98) may include discharging the first water as heated from the first heat exchanger to a first flash vessel, discharging low pressure steam at 150 psig or less from the first flash vessel, and discharging the second water as heated from the second heat exchanger to a second flash vessel If so, the method (e.g., FIGs. 94 and 98) may include discharging medium pressure steam in the range of 150 psig to 600 psig (or high pressure steam at 600 psig or greater) from the second flash vessel, discharging the third water as heated from the third heat exchanger to a third flash vessel, and discharging high pressure steam at 600 psig or greater (or very high pressure steam at 1500 psig or greater) from the third flash vessel. The method (e.g., FIGs.

94 and 98) may include discharging high pressure steam at 600 psig or greater (or very high pressure steam at 1500 psig or greater) from the second flash vessel as the third water through the third heat exchanger to superheat the high pressure steam.

[22401 In implementations (e.g., FIGs. 96-98), the method may include: discharging the first water as heated by the first heat exchanger to a flash vessel; discharging the second water as heated by the second heat exchanger to the flash vessel; discharging the third water as heated by the third heat exchanger to die flash vessel; and discharging high pressure steam at 600 psig or greater (or veiy high pressure steam at 1500 psig or greater) from the flash vessel. The method may include diverting a portion of the high pressure steam (or very high pressure steam) through a control valve to reduce pressure of the portion to medium pressure steam in a range of 150 psig to 600 psig. The method may include diverting a portion of the high pressure steam (o r very high pressure steam) through a control valve to reduce pressure of the po rtion to medium pressure steam in a range of 150 psig to 600 psig to low pressure steam at 150 psig or less. The method (e.g., FIGs. 97-98) may include superheating the high pressure steam in a heat exchanger with heat from the third heat-transfer fluid or from a third-reactor effluent discharged from the third reactor. The third-reactor effluent may include the corresponding aikene as a product of the ODH reactor system.

[2241] Techniques are provided for oxidative dehydrogenation (ODH) of ethane into ethylene with subsequent reaction of at least a portion of the ethylene with acetic acid to provide vinyl acetate monomer. In some embodiments disclosed herein, the degree to which carbon monoxide is produced during the ODH process can be mitigated by converting it to carbon dioxide, which can then act as an oxidizing agent. The process can be manipulated so as to control the output of carbon dioxide from the process to a desired level. Using the methods described herein a user may choose to operate in carbon dioxide neutral conditions such that surplus carbon dioxide need not be flared or released into the atmosphere.

[2242] Disclosed herein are methods lor mitigating carbon monoxide and/or acetylene in an ODH process and controlling the carbon dioxide output from the ODH process. Aspects of the methods include introducing, into at least one ODH reactor a gas mixture of a lower alkane, oxygen and carbon dioxide, under conditions that allow production of the corresponding aikene and smaller amounts of various by-products. For multiple ODH reactors, each reactor contains the same or different ODH catalyst, provided, in some embodiments, the at least one ODH catalyst is capable of using carbon dioxide as an oxidizing agent. In some embodiments steam or other inert diluents may also be introduced into the reactor as part of die gas mixture. In some embodiments die amount of carbon dioxide leaving the reactor is subsequently monitored. If die amount of carbon dioxide output is below a desired level then the amount of steam introduced into the reactor can be increased if die amount of carbon dioxide output is above die desired level then the amount of steam introduced into the reactor can be decreased. [2243] In some embodiments the lower alkane is ethane, and the corresponding alkene is ethylene.

[2244] In some embodiments, the lower alkane is ethane and the corresponding alkene is ethylene. Techniques are provided for methods of converting ethane to ethylene and combining at least a portion of the ethylene with acetic acid to form vinyl acetate. The method according to these aspects includes (a) providing a first stream that includes ethane and oxygen to an oxidative dehydrogenation reactor; (b) converting at least a portion of the ethane to ethylene and acetic acid in the oxidative dehydrogenation reactor to provide a second stream exiting the oxidative dehydrogenation reactor that includes ethane, ethylene, acetic acid, oxygen and carbon monoxide; (c) separating at least a portion of die acetic acid from the second stream to provide an acetic acid containing stream and a third stream that includes ethane, ethylene, oxygen and carbon monoxide; (d) providing the diird stream to a CO Oxidation Reactor containing a catal st that includes a group 11 metal and optionally a promoter that includes CeOi, ZrC>2 and combinations thereof to convert a least a portion of the carbon monoxide to carbon dioxide to produce a fourth stream that includes ethane, ethylene and carbon dioxide; and (e) providing a portion of the foiufh stream and at least a portion of the acetic acid containing stream to a third reactor containing a catalyst that includes a metal selected from the group 10 and group 11 metals and combinations thereof to convert a least a portion of the ethylene and acetic acid to vinyl acetate.

[2245] In some embodiments, at least one ODH reactor is a fixed bed reactor. In some embodiments at least one ODH reactor is a fixed bed reactor that includes heat dissipative particles within the fixed bed. In some embodiments the heat dissipative particles have a thermal conductivity that is greater than the catalyst. In alternative embodiments, at least one ODH reactor is a fluidized bed reactor.

[2246] In some embodiments, at least one ODH catalyst is a mixed metal oxide catal st. In particular embodiments, at least one ODH catalyst is a mixed metal oxide of the formula: MofyCTeJNrb _a Pd _e O/, wherein a, b, c, d, e and f are the relative atomic amounts of the elements Mo, V, Te, Nb, Pd and O, respectively; and when a = I , b = 0.01 to 1.0, c = 0.01 to 1 .0, d = 0.01 to 1 .0, 0.00 < e < 0.10 and f is a number to satisfy the valence state of the catalyst. Any other catalyst described herein may be used in addition to, or instead of, this catalyst formulation. [2247] In other particular embodiments, at least one ODH catalyst is a mixed metal oxide of the formula:

M06.25-7.25V30<j where d is a number to satisfy the valence of the oxide.

[2248] Various embodiments relate to oxidative dehydrogenation (ODH) of lower alkanes into corresponding alkenes. Lower alkanes are saturated hydrocarbons with from 2 to 4 carbons, and the corresponding alkene includes hydrocarbons with the same number of carbons, but with one carbon to carbon double bond. While any of the lower alkanes can be converted to their corresponding alkenes using the methods disclosed herein, one particular embodiment is the ODH of ethane, producing its corresponding alkene, ethylene.

[2249] Carbon Dioxide Output

[2250] Carbon dioxide can be produced in the ODH reaction as a by-product of oxidation of the alkanes and recycled from the oxidation of carbon monoxide. Carbon dioxide can also be added into the ODH reactor when used as an inert diluent. Conversely, carbon dioxide may be consumed when it acts as an oxidant for the deh drogenation reaction. The carbon dioxide output is therefore a function of the amount of carbon dioxide added and produced minus that consumed in the oxidative process. In some embodiments, the disciosed methods control the degree to which carbon dioxide acts as an oxidizing agent so as to impact the overall carbon dioxide output coming off the ODH reactor.

[2251J Measuring the amount of carbon dioxide coming off the ODH reactor can be done using any means known in the art. For example, one or more detectors such as GC, IR, or Rahman detectors, are situated immediately downstream of the reactor to measure the carbon dioxide output. While not required, the output of other components may also be measured. These include but are not limited to the amounts of ethylene, unreacted ethane, carbon monoxide and oxygen, and by-products such as acetic acid. In addition, it should be noted that depending on Site chosen metric for carbon dioxide output, the output levels of the other components for example ethane, may actually be required.

[2252] Carbon dioxide output can be stated using any metric commonly used in the art. For example, the carbon dioxide output can be described in terms of mass flow' rate (g/min) or volumetric flow rate (crnVmin). In some embodiments, normalized selectivity can be used to assess the degree to which carbon dioxide is produced or consumed. In that instance, the net mass flow' rate of C(¾ — the difference between the mass flow rate of C(¾ entering and leaving the ODH reactor — is normalized to the conversion of ethane, in essence describing what fraction of ethane is converted into carbon dioxide as opposed to ethylene, or other by-products such as acetic acid.

A carbon selectivity of 0 indicates that the amount of carbon dioxide entering the reactor is the same as the carbon dioxide output. In other words, the process is carbon dioxide neutral. A positive carbon dioxide selectivity alerts a user that carbon dioxide is being produced, and that any oxidation of carbon dioxide that is occurring is insufficient to offset that production, resulting in the process being carbon dioxide positive which may result in a lower selectivity for the olefin.

[2253] In some embodiments, the methods and apparatus disciosed herein provide the possibility of a carbon dioxide negative process. In this instance, carbon dioxide is oxidized at a higher rate than it is produced and show's a negative carbon selectivity. The ODH process may produce carbon dioxide, but the degree to which carbon dioxide is consumed while acting as an oxidizing agent offsets any production that is occurring. Many industrial processes, in addition to ODH, produce carbon dioxide which must be captured or flared where it contributes to the emission of greenhouse gases. When using a carbon dioxide negative process, the excess carbon dioxide from other processes may be captured and used as the inert diluent in the ODH process under conditions where there is negative carbon selectivity. An advantage then is the ability to reduce the amount of carbon dioxide produced in the ODH process in combination with other processes, such as thermal cracking in addition, oxidation of carbon dioxide is endothermic and by increasing the degree to which carbon dioxide acts as an oxidizing agent, hea t produced from ODH of ethane is partially offse t by oxidation of carbon dioxide, reducing the degree to which heat must be removed from the reactor. In some embodiments, when acting as an oxidizing agent, carbon dioxide can produce carbon monoxide, which can be captured and used as an intermediate in production of other chemical products such as methanol or formic acid. [2254] The ODH Process

[2255] ODH of alkanes includes contacting a mixture of one or more alkanes and oxygen in an ODH reactor with an ODH catalyst under conditions that promote oxidation of the alkanes into their corresponding aikene. Conditions within the reactor are controlled by the operator and include, but are not limited to, parameters such as temperature, pressure, and flow rate. Conditions will vary and can be optimized for a particular alkane, or for a specific catalyst, or whether an inert diluent is used in the mixing of the reactants.

[2256] An ODH reactor can be used for performing an ODH process consistent with the provided techniques. For best results, the oxidative dehydrogenation of one or more alkanes may be conducted at temperatures from 300 °C to 450 °C, or from 300 °C to 425 °C, or from 330 °C to 400 °C, at pressures from 0.5 to 100 psig (3.447 to 689.47 kPag), or from 15 to 50 psig (103.4 to 344.73 kPag), and the residence time of the one or more alkanes in die reactor may be from 0.002 to 30 seconds or from 1 to 10 seconds.

[2257] In some embodiments, die process lias a selectivity for the corresponding aikene (ethylene in die case of ethane ODH) of greater than 95%, or for example greater than 98%. The gas hourly space velocity (GHSV) can be from 500 to 30000 ir ¹ , or greater than 1000 Sr ¹ In some embodiments, the space-time yield of corresponding aikene (productivity) in g/hour per kg of the catalyst can be at least 100 or above, or greater than 1500, or greater than 3000, or greater than 3500, at 350 to 400 °C. In some embodiments, the productivity of the catalyst will increase with increasing temperature until the selectivity is decreased.

[2258] ODH Catalyst

[2259] Any of the ODH catalysts kno wn in the art are suitable for use in the methods disclosed herein. Non- limiting examples of suitable oxidative dehydrogenation catalyst include any of the cataly sts described herein.

[2260] ODH Reactor

[2261] Any of the known reactor types applicable for the ODH of alkanes may be used with the methods disclosed herein. For example, this may include reactors that are fixed bed, swing bed, fluidized bed, recirculating fluidized bed, rotating bed, ebulated bed, moving bed, tubular heat exchanger type, heat pipe or thermal syphon type, or any combinations thereof. In some embodiments, the methods may be used with conventional fixed bed reactors. In a typical fixed bed reactor, reactants are introduced into the reactor at one end, flow past an immobilized catalyst, products are formed and leave at the other end of the reactor. Designing a fixed bed reactor suitable for the methods disclosed herein can follow techniques known for reactors of this type. A person skilled in the art would know' which features are required with respect to shape and dimensions, inputs for reactants, outputs for products, temperature and pressure control, and means for immobilizing the catalyst.

[2262] In some embodiments, the use of inert non-catalytic heat dissipative particles can be used within one or more of the ODH reactors. In various embodiments, the heat dissipa tive particles are present within the bed and include one or more non catalytic inert particulates having a melting point at least 30 °C, in some embodiments at least 250 °C, in further embodiments at least 500 °C above the temperature upper control limit for the reaction; a particle size in range of 0.5 to 75 tmn, in some embodiments 0.5 to 15, in further embodiments in range of 0.5 to 8, in further embodiments in the range of 0.5 to 5 tmn; and a thermal conductivity of greater titan 30 W/mK (watts/meter Kelvin) within the reaction temperature control limits. In some embodiments the particulates are metal alloys and compounds having a thermal conductivity of greater than 50 W/mK (watts/meter Kelvin) within the reaction temperature control limits. Non-limiting examples of suitable metals that can be used in these embodiments include, but are not limited to, silver, copper, gold, aluminum, steel, stainless steel, molybdenum, and tungsten. [2263 J The heat dissipative particles can have a particle size of from about 1 mm to about 15 mm. in some embodiments, the particle size can be from about 1 mm to about 8 mm. The heat dissipative particles can be added to the fixed bed in an amount from 5 to 95 wt. %, in some embodiments from 30 to 70 wt. %, in other embodiments from 45 to 60 wt. % based on the entire weight of the fixed bed. The particles are employed to potentially improve cooling homogeneity and reduction of hot spots in the fixed bed by transferring heat directly to the walls of the reactor.

[2264] Additional embodiments include the use of a fluidized bed reactor, where the catalyst bed can be supported by a porous structure, or a distributor plate, located near a bottom end of the reactor and reactants flow through at a velocity sufficient to fluidize the bed (e.g. the catalyst rises and begins to swirl around in a fluidized manner). The reactants are converted to products upon contact with the fluidized catalyst and the reactants are subsequently removed from the upper end of the reactor. Design considerations those skilled in the art can modify and optimize include, but are not limited to, the shape of the reactor, the shape and size of the distributor plate, the input temperature, the output temperature, and reactor temperature and pressure control [2265] Some embodiments include using a combination of both fixed bed and fluidized bed reactors, each with the same or different OD3T catalyst. The multiple reactors can be arrayed in series or in parallel configuration, the design of which falls within the knowledge of the worke r skilled in the art.

[2266] Oxygen/Alkane Mixture

[2267] Safety of the ODH process is a primary concern. For that reason, in some embodiments, mixtures of one or more alkanes with oxygen should be employed using ratios that fall outside of the flammability envelope of the one or more alkanes and oxygen. In some embodiments, the ratio of alkanes to oxygen may fall outside the upper flammability envelope. In these embodiments, the percentage of oxygen in the mixture can be less than 30 wt %, in some cases less than 25 wt %, or in other cases less than 20 wt. %.

[2268] In embodiments with higher oxygen percentages, alkane percentages can be adjusted to keep the mixture outside of the flammability envelope. While a person skilled in the art would be able to determine an appropriate ratio level, in many cases the percentage of alkane is less titan about 40 wt. %. As a non-limiting example, where the mixture of gases prior to ODH includes 20% oxygen and 40% alkane, the balance can be made up with an inert diluent. Non-limiting examples of useful inert diluents in this embodiment include, but are not limited to, one or more of nitrogen, carbon dioxide, and steam. In some embodiments, the inert diluent should exist in the gaseous state at the conditions within the reactor and should not increase the flammability of the hydrocarbon added to the reactor, characteristics that a skilled worker would understand when deciding on which inert diluent to employ . The inert diluent can be added to either of the alkane containing gas or the oxygen containing gas prior to entering the ODH reactor or may be added directly into the ODH reactor.

[2269] In some embodiments, mixtures that fall within the flammability envelope may be employed, for example, in instances where tire mixture exists in conditions that prevent propagation of an explosive event. In these examples, the flammable mixture is created within a medium where ignition is immediately quenched. As a further example, a user may design a reactor where oxygen and the one or more alkanes are mixed at a point where they are surrounded by a flame arresting material. Any ignition would be quenched by the surrounding material. Flame arresting materials include, but are not limited to, metallic or ceramic components, such as stainless steel walls or ceramic supports. In some embodiments, oxygen and alkanes can be mixed at a low temperature, where an ignition event would not lead to an explosion, then introduced into the reactor before increasing the temperature. The flammable conditions do not exist until the mixture is surrounded by the flame arrestor material inside of the reactor. [22701 Carbon Monoxide Output

[227 ί I Carbon monoxide can be produced in the ODH reaction as a by-product of oxidation of die one or more alkanes. The caibon monoxide output is a function of the amount of carbon monoxide produced in the oxidative process.

[2272] Measuring the amount of carbon monoxide coming off the ODH reactor can be done using any means known in die art. For example, one or more detectors such as GC, IR, or Rahman detectors, are situated immediately downstream of the reactor to measure the caibon monoxide output. While not required, the output of other components may also be measured, such as the amounts of ethylene, unreacted ethane, carbon dioxide and oxygen and by-products (for example, acetic acid).

[2273] Carbon monoxide output can be stated using any metric commonly used in the art. For example, the carbon monoxide output can be described in terms of mass flow rate (g/min) or volumetric flow rate (emVmin) In some embodiments normalized selectivity can be used to assess the degree to which carbon monoxide is produced or consumed. In that instance the net mass flow rate of CO — -the difference between the mass flow rate of CO leaving the ODH reactor — is normalized to the conversion of ethane, in essence describing what fraction of ethane is converted into carbon monoxide as opposed to ethy lene, or other by-products such as acetic acid

[2274] Many industrial processes, in addition to ODH, produce carbon monoxide which must be captured or flared where it contributes to the emission of greenhouse gases. Using the carbon monoxide mitigation steps disclosed herein converts most, if not all, carbon monoxide resulting from the ODH process to carbon dioxide. An advantage then is the ability to reduce or eliminate the amount of carbon monoxide produced in the ODH process in combination with other processes, such as thermal cracking.

[2275] Acetylene Output

[2276] Acetylene can be produced in the ODH reaction as a by-product of oxidation of the one or more alkanes. The acetylene output is a function of the amount of acetylene produced in the oxidative process.

[2277] Measuring the amount of acetylene coming off the ODH reactor can be done using any means known in the art. For example, one or more detectors such as GC, IR, or Rahman detectors, are situated immediately downstream of the reactor to measure the acetylene output. While not required, the output of other components may also be measured, such as die amounts of ethylene, unreacted ethane, carbon monoxide, caibon dioxide and oxygen, and by-products (for example acetic acid).

[2278] Acetylene output can be stated using any medic commonly used in die art. For example, the acetylene output can be described in terms of mass flow' rate (g/min), volumetric flow rate (cnd/rnin) or volumetric parts per million (vppm). In some embodiments, normalized selectivity can be used to assess the degree to which acetylene is produced or consumed. In that instance the net mass flow rate of acetylene — the difference between the mass flow rate of acetylene leaving the GDH reactor — is normalized to the conversion of ethane, in essence describing what fraction of ethane is converted into acetylene as opposed to ethylene, or other by-products such as acetic acid.

[2279] Using the acetylene mitigation steps disclosed herein reacts most, if not all, acetylene resulting from the ODH process. An advantage then is the ability to reduce or eliminate the amount of acety lene produced in the ODH process in combination with other processes, such as thermal cracking and eliminate downstream unit operations in an ODH -type process.

[22801 Addition of Steam

[2281] The amount of steam added to the ODH process affects the degree to which carbon dioxide acts as a reactant. In some embodiments steam may be added directly to the ODH reactor or steam may be added to the individual reactant components — die lower alkane, oxygen, or inert diluent — or combinations thereof, and subsequently introduced into the ODH reactor along with one or more of the reactant components. Alternatively, steam may be added indirectly as water mixed with either the lower alkane, oxygen or inert diluent, or a combination thereof, with the resulting mixture being preheated before entering the reactor. When adding steam indirectly as water the preheating process should increase the temperature so that the water is entirely converted to steam before entering the reactor.

[2282] Increasing the amount of steam added to a reactor increases the degree to which carbon dioxide acts as an oxidizing agent. Decreasing the amount of steam added to the reactor decreases the degree to which carbon dioxide acts as an oxidizing agent. In some embodiments, a user monitors the carbon dioxide output and compares it to a predetermined target carbon dioxide output. If the carbon dioxide output is above the target a user can then increase the amount of steam added to the ODH process. If the carbon dioxide output is below' the target a user can decrease the amount of steam added to the ODH process, provided steam lias been added. Setting a target carbon dioxide output level is dependent on the requirements for the user in some embodiments increasing the steam added will have the added effect of Increasing the amount of acetic acid and other by-products produced in the process. A user that is ill equipped to separate out larger amounts of acetic add from the output of the ODH may instead reduce steam levels to a minimum, while a user that desires a process that consumes carbon dioxide may choose to maximize the amount of steam that can be added. The amount of steam added to the one or more ODH reactors can be up to about 40 wt. %, in some cases up to about 35 wt. %, in other cases up to about 30 wt. %, and in some instances up to about 25 wt. %.

[2283] in some embodiments, when using two or more ODH reac tors, a user may choose to control carbon dioxide output in only one, or less than the whole complement of reactors. For example, a user may opt to maximize carbon dioxide output of an upstream reactor so that the higher level of carbon dioxide can be pari of the inert diluent for the subsequent reactor. In (hat instance, maximizing carbon dioxide output upstream minimizes the amount of inert diluent that would need to be added to the stream prior to Site next reactor.

[2284] There is no requirement for adding steam to an ODH process, as it is one of many alternatives for the inert diluent. For processes where no steam is added, the carbon dioxide output is maximized mider the conditions used with respect to ethane, oxygen and inert diluent inputs. Decreasing the carbon dioxide output is then a matter of adding steam to the reaction until carbon dioxide output drops to the desired level. In embodiments where oxidative dehydrogenation conditions do not include addition of steam, and the carbon dioxide output is higher than the desired carbon dioxide target level, steam may be introduced into the reactor while keeping rela tive amounts of the main reactants and inert diluent — lower alkane, oxygen and inert diluent — added to the reactor constant, and monitoring the carbon dioxide output, increasing the amount of steam until carbon dioxide decreases to the target level. i2285| In some embodiments, a carbon dioxide neutral process can be achieved by increasing steam added so that any carbon dioxide produced in the oxidative dehydrogenation process can then be used as an oxidizing agent such drat there is no net production of carbon dioxide. Conversely, if a user desires net positive carbon dioxide output then the amount of steam added to die process can be reduced or eliminated to maximize carbon dioxide production. As the carbon dioxide levels increase there is potential to reduce oxygen consumption, as carbon dioxide is competing as an oxidizing agent. The skilled person would understand drat using steam to increase die degree to which carbon dioxide acts as an oxidizing agent can impact oxygen consumption. The implication is that a user can optimize reaction conditions with lower oxygen contributions, which may assist in keeping mixtures outside of flammability limits.

[2286] In some embodiments, the stream exiting the one or more ODH reactors can be treated to remove or separate water and water soluble hydrocarbons from the stream exiting the one or more ODH reactors. In some embodiments, this stream is fed to a CO Oxidation reactor.

[2287] Acetic Acid Removal

[2288] Prior to being fed to the CO Oxidation Reactor, the stream exiting the one or more ODH reactors is directed to quench tower or acetic add scrubber, which facilitates removal of oxygenates, such as acetic add, and water via a bottom outlet. A stream containing unconverted lower alkane (such as ethane), corresponding alkene (such as ethylene), unreacted oxygen, carbon dioxide, carbon monoxide, optionally acetylene and inert diluent, are allowed to exit the scrubber and are fed to the CO Oxidation Reactor.

[2289] The oxygenates removed via the quench tower or acetic acid scrubber can include carboxylic acids (for example acetic acid), aldehydes (for example acetaldehyde) and ketones (for example acetone). The amount of oxygenate compounds remaining in the stream exiting the scrubber and fed to the CO Oxidation Reactor will often be zero, i.e, below the detection limit for analytical test methods typically used to detect such compounds. When oxygenates can be detected they can be present at a level of up to about 1 per million by volume (pprnv), in some cases up to about 5 ppmv, in other cases less titan about 10 ppmv, in some instances up to about 50 ppmv and in other instances up to about 100 ppmv and can be present up to about 2 vol. %, in some cases up to about 1 vol. %, and in other cases up to about 1.000 ppmv. The amount of oxygenates or acetic acid in the stream exiting the scrubber and fed to the CO Oxidation Reactor can be any value, or range between any of the values recited above. [2290] The CO Oxidation Reactor

[2291] In some embodiments, the ODH reactor (or reactors) can provide a stream containing at least a small amount of oxygen remaining as reactor effluent. In some embodiments the oxygen can provide a benefit to the ODH reactor product gas. Tu some embodiments, when the ODH catalyst is exposed to an oxygen free reducing environment at elevated temperature, it may become permanently degraded. In other embodiments, if the level of oxygen in the product gas from the ODH reactor contains less than about 1 ppm of oxygen, most, if not all, of the one or more alkanes are converted to one or more alkenes in the inlet portion of the reactor and a large portion of the reactor catalyst bed is not utilized.

[2292] In some embodiments, oxygen in the ODH reactor product gas causes serious safety and operational issues in the downstream equipment, as a non-limiting example, at the first compression stage of an ODH process. This process safety consideration presents a need to remove oxygen to a very low or non-detectable level before the product gas is compressed.

[2293] One method used to reduce/eliminate oxygen in the ODH product gas focuses on catalyticaily combusting a small portion of the ODH product gas to the complete consumption of any residual oxygen. This approach is viable, however, in sortie cases it is undesirable, because it increases the overall oxygen consumption in the ODH process and, in examples in which the alkane is ethane, reduces overall process selectivity toward ethylene.

[2294] Techniques are provided for a process where the ODH reaction can proceed with partial consumption of CO _? (CO _? can act as an oxidizing agent, and be converted to CO), reducing overall oxygen consumption in the process by providing a portion of the required oxygen from CO _? . In many embodiments, more oxygen passes through the catalyst bed unconverted when CO _? is provided and acts as an oxidizing agent.

[2295] Oxidation of Carbon Monoxide

[2296] In the process, the ODH reactor product stream is fed to the CO Oxidation reactor, which contains a catalyst that includes one or more selected from a group 11 metal, a group 4 metal, a group 7, a group 9 metal, a lanthanide metal, and an actinide metal and/or their corresponding metal oxides capable of converting at least a portion of the carbon monoxide to carbon dioxide. The carbon dioxide can be recycled to the ODH reactor to act as an oxidizing agent as described above

[2297] In some embodiments, the group 11 metal can be selected from copper, silver, gold and combinations thereof. In some embodiments, the group II metal is silver or copper.

[2298] In some embodiments, the group 4 metal can be selected from titanium, zirconium, hafnium, rutherfordium and combinations thereof. In some embodiments, the group 4 metal is zirconium.

[2299] In some embodiments, the group 7 metal can be selected from manganese, technetium, rhenium, bohrium and combinations thereof. In some embodiments, the group 7 metal is manganese.

[2300] In some embodiments, the group 9 metal can be selected from cobalt, rhodium, iridium, meitemium and combinations thereof. In some embodiments, the group 9 metal is cobalt.

[2301] In embodiments the lanthanide metal can be selected from La, Ce, Pr, Nd, Pm. Sm, Eu, Gd, Tb, Dy, ho, Er, Trn, Yb and combinations thereof. In some embodiments, the lanthanide metal is Cerium.

[2302] In some embodiments, the actinide metal can be selected from Ac, Th, Ps, U, Np, Pu, Am, Cm, Bk, Cf, Es, Fin, Md, No and combinations thereof. In some embodiments, Site actinide metal is thorium. [2303] In some embodiments, the CO Oxidation reactor catalyst, in some cases a group 11 metal, is used in conjunction with a promoter. In some embodiments, the promoter is selected from one or more of the lanthanide and actinide metals (as defined above) and their corresponding metal oxides. In some embodiments, the promoter is selected from one or more of the lanthanide metals and their corresponding metal oxides. In some embodiments, the promoter includes cerium and its corresponding metal oxides.

[2304] in some embodiments, the CO Oxidation reactor catalyst, in some cases a group 11 metal, and optional promotor are provided on a support. The support is typically an inert solid with a high surface area, to which the CO Oxidation reactor catalyst and optional promotor can be affixed. In some embodiments, the support includes Si. Ge, Sn, their corresponding oxides and combinations thereof.

[2305] In some embodiments, examples of suitable CO Oxidation reactor catalysts with optional promotors and supports include Ag/SiCfe, AgCe02/ ^' Si02, AgZrCVSiCfe, AgCo304/Si02, Cu/SiCh, CuCe02/Si02, CUZ1O2/S1O2, CUC03O4/S1O2 and combinations thereof.

[2306] In some embodiments, examples of suitable CO Oxidation reactor catalysts with optional promoters and supports include AgCe0 ₂ /Si0 _2, AgZi0 ₂ /Si0 ₂ and combinations thereof.

[2307] In some embodiments, the CO Oxidation reactor catalyst includes silver, the optional promoter includes cerium and the support includes S1O2.

[2308] In some embodiments, the CO Oxidation reactor catalyst includes copper, the optional promoter includes cerium and the support includes S1O2.

[2309] In some embodiments, when oxidation of carbon monoxide is preferentially desired, the CO Oxidation reactor catalyst includes manganese, the optional promoter includes cerium and the support includes S1O2.

[2310] In some embodiments, the group i 1 metal with optional promoter and optional support can be used in a process where 1) some oxygen is in the stream leaving the ODH reactor; 2) the temperature in the stream is decreased; 3) the cooled stream is fed to an acetic acid scrubber; 4) the stream from the acetic acid scrubber is fed to reactor 2 as described above, where most or all of the residual O2 is consumed and CO is converted to CO2; and 5) optionally, the CO2 is recycled back to the ODH reactor.

[2311] In some embodiments, the amount of oxygen in the stream leaving the ODH reactor in 1) can be at least about 80 ppm, in some cases at least about 100 ppm, in other cases at least about 150 ppm and in some instances at least about 200 ppm and can be up to about 5 wt. %, in some cases up to about 4 wt. %, in other cases up to about 3 wt. %, in some instances up to about 2 wt. %, in other instances up to about 1 wt. %, and in particular situations up to about 500 ppm. The amount of oxygen in the stream leaving the ODH reactor in 1 ) can be any of the values or range between any of the values recited above.

[2312] In some embodiments , when there is oxygen in the stream leaving the CO Oxidation reactor (in some instances the amount of oxygen will be undetectable or zero ppm), the amount of oxygen in the stream leaving the CO Oxidation reactor can be at least about 1 ppm, in some cases at least about 2 ppm, in other cases at least about 3 ppm and in sortie instances at least about 5 ppm and can be up to about 1 wt. %, in some cases up to about 0.9 wt. %, in otlter cases up to about 0.8 wt. %, in some instances up to about 0.7 wt. %, in oilier instances up to about 0.6 wt. %, and in particular situations up to about 0.5 wt. %. The amount of oxygen in the stream leaving the CO Oxidation reactor can be any of the values or range between any of the values recited above.

[2313] In some embodiments, the amount of carbon monoxide in the stream leaving the ODH reactor in 1) can be at least about 100 ppm, in some cases at least about 200 ppm, in other cases at least about 300 ppm and in some instances at least about 400 ppm and can be up to about 10 wt. %, in some cases up to about 9 wt. %, in other cases up to about 8 wt. %, in some instances up to about 7 wt. %, in other instances up to about 6 wt. %, and in particular situations up to about 5 wt. %. The amount of caibon monoxide in the stream leaving the ODH reactor in 1) can be any of Site values or range between any of the values recited above.

[2314] In some embodiments, when there is carbon monoxide in the stream leaving the CO Oxidation reactor (in some instances the amount of carbon monoxide will be undetectable or zero ppm), the amount of caibon monoxide in die stream leaving the CO Oxidation reactor can be at least about I ppm, in some cases at least about 2 ppm. in other cases at least about 3 ppm and in some instances at least about 5 ppm and can be up to about 8 wt. %, in some cases up to about 7 wt. %, in other cases up to about 6 wt. %, in some instances up to about 5 wt. %, in other instances up to about 4 wt. %, and in particular situations up to about 3 wt. % The amount of carbon monoxide in the stream leaving the CO Oxidation reactor can be any of the values or range between any of the values recited above.

[2315] In some embodiments, temperature in the CO Oxidation reactor can be at least about 40 °C, in some cases at least about 45 °C, in some cases at least about 50 °C and in some caces at least about 55 °C and can be up to about 200 °C, in some cases up to about 150 °C, in some cases up to about 120 °C, in some cases up to about 90 °C, in some cases up to about 85 °C, in some cases up to about 80 °C, in some cases up to about 75 °C and in some cases up to about 70 °C. The temperature of the CO Oxidation reacto r can be any temperature value or range between any of the temperature values, including a temperature gradient within the CO Oxidation reactor, recited above.

[2316] Acetylene Elimination

[2317] In the process, the ODH reactor product stream is fed to the CO Oxidation reactor, which contains a cataly st that includes one or more selected from a group 11 metal, a group 4 metal, a group 9 metal, a lanthanide metal, and an actinide metal and/or their corresponding metal oxides capable of reacting at least a portion of the acetylene.

[2318] In some embodiments, the group 11 metal can be selected from copper, silver, gold and combinations thereof. In some embodiments, the group 11 metal is silver.

[2 19] In some embodiments, the group 4 metal can be selected from titarrium, zirconium, hafnium, rutherfordium and combinations thereof. In some embodiments, the group 4 metal is zirconium.

[2320] In some embodiments, the group 9 metal can be selected from cobalt, rhodium, iridium, meitemium and combinations thereof. In some embodiments, the group 9 metal is cobalt.

[2321 ] In sortie embodiments, the lanthanide metal can be selected from La, Ce, Pr, Nd, Pm, Sin, Eu, Gd, Tb,

Dy, ho, Er, Trn, Yb and combinations thereof. In sortie embodiments, the lanthanide metal is Cerium.

[2322] In sortie embodiments, the actinide metal can be selected from Ac, Th, Ps, U, Np, Pu, Am, Cm, Bk, Cf, Es, Fm. Md, No and combinations thereof. In some embodiments, the actinide metal is thorium. [2323] In some embodiments, the CO Oxidation reactor catalyst, in some cases a group 11 metal, is used in conjunction with a promoter. In some embodiments, the promoter is selected from one or more of the lanthanide and actinide metals (as defined above) and their corresponding metal oxides. In some embodiments, the promoter is selected from one or more of the lanthanide metals and their corresponding metal oxides. In some embodiments, the promoter includes cerium and its corresponding metal oxides.

[2324] in some embodiments, the CO Oxidation reactor catalyst, in some cases a group 11 metal, and optional promotor are provided on a support. The support is typically an inert solid with a high surface area, to which the CO Oxidation reactor catalyst and optional promotor can be affixed. In some embodiments, the support includes Si. Ge, Sn, their corresponding oxides and combinations thereof.

[2325] In some embodiments, examples of suitable CO Oxidation reactor catalysts with optional promoters and supports include Ag/SiCfe, AgCe02/ ^' Si02, AgZrCySiCfe, AgCo304/Si02, Cu/SiCh, CuCe02/Si02, CuZKVSiCh, CUC03O4/S1O2 and combinations thereof.

[2326] In some embodiments, examples of suitable CO Oxidation reactor catalysts with optional promoters and supports include AgCe0 ₂ /Si0 _2, AgZi0 ₂ /Si0 ₂ and combinations thereof.

[2327] In some embodiments, the CO Oxidation reactor catalyst includes silver, the optional promoter includes cerium and the support includes S1O2.

[2328] In some embodiments, the CO Oxidation reactor catalyst includes copper, the optional promoter includes cerium and the support includes S1O2.

[2329] In some embodiments, the group 11 metal with optional promoter and optional support can be used in a process where 1) some acetylene is in the stream leaving the ODH reactor; 2) the temperature in the stream is decreased; 3) the cooled stream is fed to an acetic acid scrubber; 4) the stream from the acetic acid scrubber is fed to reactor 2 as described above, where most or all of the acetylene is consumed and CO is oxidized to CO _? ; and 5) optionally, the CO2 is recycled back to the ODH reactor.

[2330] In some embodiments, when there is acetylene in the stream leaving the ODH reactor (in some instances the amount of acetylene wall be undetectable or zero vppm), the amount of acetylene in the stream leaving the ODH reactor in 1) can be at least about 1 vppm, in some cases at least about 2 vppm, in other cases at least about 5 vppm and in some instances at least about 10 vppm and can be up to about 1000 vppm, in some cases up to about 750 vppm, in other cases up to about 500 vppm, in some instances up to about 400 vppm, in other instances up to about 300 vppm, and in particular situations up to about 300 vppm. The amount of acetylene in the stream leaving the ODH reactor in 1) can be any of the values or range between any of the values recited above.

[23 1 ] In some embodiments, the amount of acetylene in the stream leaving the CO Oxidation reactor will be less Hum the amount entering the CO Oxidation reactor and, in some instances, the stream exiting the CO Oxidation reactor will be substantially free of acetylene.

[2332] In some embodiments, when there is acetylene in the stream leaving the CO Oxidation reactor (in some instances, the amount of acetylene will be undetectable or zero vppm). the amount of acetylene in Site stream leaving the CO Oxidation reactor can be at least about I vppm, in some cases at least about 2 vppm, in some cases at least about 3 vppm and in some cases at least about 5 vppm and can be up to about 100 vppm, in some cases up to about 50 vppm, in some eases up to about 25 vppm, in some eases up to about 20 vppni, in other eases up to about 15 vppm, and in some cases up to about 10 vppm. The amount of acetylene in the stream leaving the CO Oxidation reactor can be any of the values or range between any of the values recited above.

[2333] in some embodiments, temperature in the CO Oxidation reactor can be at least about 40, in some cases at least about 45, in some cases at least about 50 and in some cases at least about 55 °C and can be up to about 200, in some cases up to about 150, in some cases up to about 120, in some cases up to about 90, in some cases up to about 85, in some cases up to about 80, in some cases up to about 75 and in some cases up to about 70 °C. The temperature of CO Oxidation reactor can be any temperature value or range between any of the temperature values including a temperature gradient within the CO Oxidation reactor, recited above.

[2334] As indicated above, when ethylene from the ODH process described herein is used in a YAM process and the ethylene stream contains carbon monoxide, it can poison the YAM catalyst and minimize or prevent the production of vinyl acetate monomer. When used to produce vinyl acetate monomer as described herein, one or more CO Oxidation reactors can be employed.

[2335] As indicated herein, one role of the one or more CO Oxidation reactors is to oxidize at least a portion of, in some cases substantially all, of the carbon monoxide to carbon dioxide in the product stream from the one or more ODH reactors. In this context, oxidizing substantially all of the carbon monoxide to carbon dioxide means that the stream exiting any of the CO Oxidation reactors contains very low amounts, undetectable amounts or no carbon monoxide. When the amount of oxy gen in the stream entering the CO Oxidation reactor is insufficient to oxidize the carbon monoxide in the entering stream, the CO Oxidation reactor can be equipped with a supplementary oxygen feed, which is adapted to provide an oxygen containing gas sufficient to oxidize all of the carbon monoxide to carbon dioxide.

[2336] In some embodiments, monitoring equipment and feed back loops known in the art can be used to adjust the amount of oxygen provided to a CO Oxidation reactor so that the amount of oxygen in the stream entering the CO Oxidation reactor is about equal to the amount required to oxidize the carbon monoxide in entering stream to carbon dioxide

[2337] In some embodiments, when oxygen is provided to the one or more CO Oxidation reactors, the stream exiting the CO Oxidation reactor can contain no carbon monoxide (undetectable or 0 ppm), in some cases no more titan about 10 ppm, in some cases no more than about 20 ppm and in some cases no more than about 30 ppm of carbon monoxide. Additionally, the stream exiting the CO Oxidation reactor can contain up to about 125 ppm, in some cases up to about 100 ppm, in some cases up to about 75 ppm and in some cases up to about 50 ppm of carbon monoxide. The amount of carbon monoxide in the stream exiting the CO Oxidation reactor is low enough to minimize or mitigate any poisoning of the YAM catalyst and can be any of the values or range between any of the values recited above.

[2338] In sortie embodiments, the stream from die ODH reactor is cooled to a lower temperature prior to being fed to an acetic acid scrubber (as described below). The temperature of the stream prior to entering die acetic acid scrubber can be at least about 40, in sortie cases at least about 45, and in some cases at least about 50 °C and can be up to about 90, in some cases up to about 85, in some cases up to about 80, in sortie cases up to about 75 and in some cases up to about 70 °C. The temperature of the ODH reactor product stream fed to an acetic acid scrubber can be cooled to any temperature value or range between any of the temperature values recited above.

[2339] In some embodiments, the configura tion described above can allow for the size of the air separation plant to be reduced, as well as improving the life of the ODH catalyst, by allowing it to be exposed to an oxygen containing environment at ail times. In additional embodiments, the configuration described above can improve the reliability and safely of the ODH reactor and downstream equipment.

[2340] In some embodiments, the net CO2 generation in the process described herein can be optimized to be zero. In these embodiments the need to flare off any C0 ₂ (with some amount of alkane/alkene) from the COi-recyle loop as described herein is minimized. In these embodiments, the total process yield of alkane to alkene can be improved.

[2341] The VAM Process

[2342] In the VAM process, an ethylene containing stream, where substantially all of the carbon monoxide lias been oxidized to carbon dioxide, is reacted with acetic acid from the acetic acid scrubber (the VAM Process Stream) to provide vinyl acetate monomer. In some embodiments, this includes contacting ethylene and acetic acid with an oxygen-containing gas in the presence of a catalyst to produce a product stream that includes vinyl acetate. The product stream can be separated to recover vinyl acetate monomer from the product stream.

[2343] The VAM process stream includes acetic acid and ethylene in a predetermined ratio with water, and may contain ethane, oxygen, nitrogen and the by-products, and carbon dioxide. As indicated above, no or very small amounts (<120 ppm) of carbon monoxide are present. Optionally, an oxygen-containing gas, can be included in the VAM process.

[2344] The catalyst active for the production of vinyl acetate which is used in the VAM process contains a metal selected from the group 10 and group 11 metals and combinations thereof.

[2345] In some embodiments, any suitable catalyst for the VAM process known in the art can be used. Some examples include a shell impregnated catalyst that includes (1) a catalyst support having a particle diameter from about 3 to about 7 mm and a pore volume of about 0.2 to about 1.5 ml per gram; (2) palladium and gold; and (3) from about 3.5 to about 9.5% by weight of potassium acetate wherein the gold to palladium weight ratio in the catalyst is in the range of from about 0.60 to about 1.25. Catalyst containing Pd, K, Mn and Cd as an additional promoter instead of Au may also be used. VAM catalysts containing (1) a catalyst support having a particle diameter of from about 3 to about 7 mm and a pore volume of from about 0.2 to about 1.5 ml/g, an approximately 10% by weight water suspension of the catalyst support having a pH from about 3 to about 9; (2) a palladium-gold alloy distributed in a surface layer of the catalyst support, the surface layer extending less than about 0.5 mm from the surface of the support, the palladium in the alloy being present in an amount of from about 1.5 to about 5.0 g/1 of catal st, and the gold being present in an amount of front about 0.5 to about 2.25 grams per liter of catalyst; and (3) from about 5 to about 60 g/1 of catalyst of alkali metal acetate.

[2346] A shell impregnated catalyst active for the production of vinyl acetate from ethylene, acetic acid and an oxygen containing gas may be used the catalyst contains: (1) a catalyst support having a particle diameter from about 3 to about 7 mm and a pore volume of from about 0.2 to 1.5 ml/g; (2) palladium and gold distributed in the outermost about 1 m thick layer of the catalyst support particles; and (3) from about 3 5 to about 95% by weight of potassium acetate where the gold to palladium weight ratio in the catalyst is in the range of from about 0.6 to about 1 25.

[2347] ODH - VAM Complex

[2348] In some embodiments, the chemical complex (one embodiment shown schematically in Figure 99) includes, in cooperative arrangement, an ODH reactor 9910, a quench tower or acetic acid scrubber 9920, a first CO Oxidation reactor 9930 (as described herein), a vinyl acetate monomer (VAM) reactor 9940, a VAM separation unit 9950 and an inert removal unit 9970. Although first CO Oxidation reactor 9930 is shown directly after quench tower or acetic acid scrubber 9920, it can be placed further downstream. In some cases, the process configuration can be more energy efficient if first CO Oxidation reactor 9930 is placed after the input stream has been compressed.

[2349] ODH reactor 9910 includes an ODH catalyst capable of catalyzing, in the presence of oxygen which may be introduced via oxygen line 9980, the oxidative dehydrogenation of alkanes introduced via alkane line 9990. The ODH reaction may also occur in the presence of an inert diluent, such as carbon dioxide, nitrogen, or steam that is added to ensure the mixture of oxy gen and hy drocarbon are outside of flammability limits. Determination of whether a mixture is outside of the flammability limits, for the prescribed temperature and pressure, is within the knowledge of the skilled worker. An ODH reaction that occurs within ODH reactor 9910 may also produce, depending on the catalyst and the pre vailing conditions within ODH reactor 9910, a variety of other products which may include carbon dioxide, carbon monoxide, oxygenates, and water. These products leave ODH reactor 9910. along with unreacted alkane, corresponding alkene, residual oxygen, carbon monoxide and inert diluent, if added via ODH reactor product line 99100.

[2350] ODH reactor product line 99100 is directed to quench tower or acetic acid scrubber 9920 which quenches the products from product line 99100, and facilitates removal of acetic acid and water via quench tower botom outlet 99110. Unconverted lower alkane, corresponding alkene, unreacted oxygen, carbon dioxide carbon monoxide, and inert diluent added to quench tower 9920 exit through quench tower overhead line 99120 and a portion are directed into first oxidation reactor 9930 via first oxidation line 99130.

[2351 ] First oxidation reactor 9930 contains a group 11 metal with optional promoter and optional support as described above. First oxidation reactor 9930 optionally includes first oxygen line 99140 which can be used to provide an oxygen containing gas to first oxidation reactor 9930. In first oxidation reactor 9930, unreacted oxygen is reacted with carbon monoxide to form carbon dioxide and/or, reacts acetylene to reduce or eliminate it. In first oxidation reactor 9930, most or all of the unreacted oxygen is consumed. The remaining unconverted lower alkane, corresponding alkene, unreacted oxygen (if present), all or part of the carbon dioxide, carbon monoxide (if present), and inert diluent are directed to first ethylene product line 99150 and combined with the acetic acid and water via quench tower bottom outlet 99110 and conveyed to the VAM process.

[2352] In the VAM process, the contents of quench tower bottom outlet 99110 often contains an acetic acid- water mixture with optional traces of ethane/ethylene/CQ/CCb. This can be mixed into effluent stream 99150, which often contains ethane/eihyTene/CO/CCk/Qr with an optional trace of acetic acid and water and can be fed to VAM reactor 9940, where acetic acid and ethylene are combined in the presence of a catalyst active for the production of vinyl acetate. VAM reactor 9940 optionally includes VAM oxygen line 99145 which can be used to provide an oxy gen containing gas to VAM reactor 9940 as described herein. Depending on the scale of the process. VAM reactor 9940 may include either a single reactor or several reactors in parallel or in series. A VAM product stream 99300 that includes vinyl acetate, water, optionally ethane, gaseous by-products and unreacted acetic acid and ethylene is withdrawn from VAM reactor 9940 and is fed to VAM separation unit 9950 where a VAM gaseous stream 99330, which can include ethylene, and optionally ethane together with inert compounds, carbon monoxide and carbon dioxide is separated from water and acetic acid, and which can alternatively be withdrawn overhead via first recycle line 99310 and can be rec cled to and mixed with the contents of quench tower bottom outlet 99110 and provided to VAM reactor 9940 as described above.

[2353] VAM liquid stream 99320, which includes vinyl acetate, water, optionally unreacted acetic acid and optionally high boiling by-products of the process are withdrawn from the base of VAM separation unit 9950 and vinyl acetate is isolated in state of the art equipment not shown. As an example, the contents of VAM liquid stream 99320 can be fed to a distillation column where vinyl acetate and water are removed as an azeotrope and acetic acid and the optional high boiling by-products are removed as a bleed from the base of the distillation column. The water in the overhead stream from the distillation column can be separated from the vinyl acetate in a decanter and a vinyl acetate product stream removed from decanter is purified by conventional means known in the art.

[2354] VAM gaseous stream 99330 is directed to inert removal unit 9970, where carbon dioxide and other inert compounds are separated and directed to outlet stream 99340 and ethylene and optionally ethane and remaining carbon monoxide are directed to ethylene containing stream 99350 and combined into quench tower overhead line 99120.

[2355] In some embodiments, the chemical complex, shown in another embodiment schematically in Figure 100, includes, in cooperative arrangement, an ODH reactor 9910, a quench tower or acetic acid scrubber 9920, a first CO oxidation reactor 9930 (as described herein), a vinyl acetate monomer (VAM) reactor 9940, a VAM separation unit 9950 and an inert removal unit 9970.

[2356] As indicated above, with reference to Figure 99, in the ODH process configuration depicted in Figure i 00, although first CO oxidation reactor 9930 is shown directly after quench tower or acetic acid scrubber 9920, it can be placed further downstream. In some cases, the process configuration can be more energy efficient if first CO oxidation reactor 9930 is placed after the input stream lias been compressed.

[2357] ODH reactor 9910 includes an ODH catalyst capable of catalyzing, in the presence of oxygen which may be introduced via oxygen line 9980, the oxidative dehydrogenation of alkanes introduced via alkane line 9990. The ODH reaction may also occur in the presence of an inert diluent, such as carbon dioxide, nitrogen, or steam, that is added to ensure the mixture of oxy gen and hydrocarbon are outside of flammability limits. Determination of whether a mixture is outside of Site flammability limits for the prescribed temperature and pressure, is within the knowledge of the skilled worker. An ODH reaction that occurs within ODH reactor 9910 may also produce, depending on the catalyst and the prevailing conditions within ODH reactor 9910, a variety of other products which may include carbon dioxide, carbon monoxide, oxygenates, and water. These products leave ODH reactor 9910, along with unreacted alkane, corresponding alkene residual oxygen, carbon monoxide and inert diluent, if added, via ODH reactor product line 99100.

[2358] ODH reactor product line 99100 is directed to quench tower or acetic acid scrubber 9920 which quenches the products from product line 99100, and facilitates removal of acetic acid and water via quench tower bottom outlet 99110. Unconverted lower alkane, corresponding alkene, unreacted oxygen, carbon dioxide, carbon monoxide, and inert diluent added to quench tower 9920 exit through quench tower overhead line 99120.

[2359] Quench tower overhead line 99120 is led to primary' oxidation reactor 99400, which contains a group 11 metal with optional promoter and optional support as described above. In primary oxidation reactor 99400, unreacted oxygen is reacted with carbon monoxide to form carbon dioxide. In primary oxidation reactor 99400, most or all of the umeacted ox gen is consumed. Primary' oxidation reactor 99400 optionally includes first ox gen line 99405 which can be used to provide an oxygen containing gas to primary oxidation reactor 99400. The remaining unconverted lower alkane corresponding alkene, umeacted oxygen (if present) all or part of the carbon dioxide, carbon monoxide (if present), and inert diluent are directed to primary' ethylene product line 99410.

[2360] A portion of the contents of primary' ethylene product line 99410 are directed to first oxidation reactor 9930 via first oxidation line 99130. First oxidation reactor 9930 operates similar to primary' oxidation reactor 99400 and reacts any remaining oxygen with carbon monoxide to form carbon dioxide and/or reduces or eliminates acetylene. The remaining unconverted lower alkane, corresponding alkene, unreacted oxygen (if present), all or part of the carbon dioxide, carbon monoxide (if present), and inert diluent are directed to first ethylene product line 99150 and combined with the acetic acid and water via quench to wer bottom outlet 99110 and conveyed to the VAM process.

[2361] In the VAM process, the contents of quench tower bottom outlet 99110 often contains an acetic acid- water mixture with optional traces of ethane/ethylene/CO/CO _. . This can be mixed into effluent stream 99150, which often contains ethane/ethylene/CO/C VCh with an optional trace of acetic acid and water and can be fed to VAM reactor 9940, where acetic acid and ethylene are combined in the presence of a catalyst active for the production of vinyl acetate. VAM reactor 9940 optionally includes VAM oxygen line 99145 which can be used to provide an oxygen containing gas to V AM reactor 9940 as described herein. Depending on the scale of the process, VAM reactor 9940 may include either a single reactor or several reactors in parallel or in series. A VAM product stream 99300 that includes vinyl acetate, water, optionally ethane, gaseous by-products and umeacted acetic acid and ethylene is withdrawn from VAM reactor 9940 and is fed to VAM separation unit 9950 where a VAM gaseous stream 99330, which can include ethylene, and optionally ethane together with inert compounds, carbon monoxide and carbon dioxide is separated from water and acetic acid, and which can alternatively be withdrawn overhead via first recycle line 99310 and can be recycled to and mixed with the contents of quench tower bottom outlet 99110 and provided to VAM reactor 9940 as described above.

[2362] VAM liquid stream 99320, which includes vinyl acetate, water, optionally unreacted acetic acid and optionally high boiling by-products of the process are withdrawn from the base of VAM separation unit 9950 and vinyl acetate is isolated in stole of the art equipment not shown. As an example the contents of VAM liquid stream 99320 can be fed to a distillation column where vin l acetate and water are removed as an azeotrope and acetic acid and the optional high boiling by-products are removed as a bleed from the base of the distillation column. The water in the overhead stream from the distillation column can be separated from the vinyl acetate in a decanter and a vinyl acetate product stream removed from decanter is purified by conventional means known in the art.

[2363] YAM gaseous stream 99330 is directed to inert removal unit 9970, where carbon dioxide and other inert compounds are separated and directed to outlet stream 99340 and ethylene and optionally ethane and remaining carbon monoxide are directed to ethylene containing stream 99350 and combined into primary ethylene product line 99410.

[23641 In some embodiments, the chemical complex shown in an additional embodiment schematically in Figure 101, includes, in cooperative arrangement, an ODH reactor 9910, a quench tower or acetic acid scrubber 9920, a first CO oxidation reactor 9930 (as described herein), a vinyl acetate monomer (VAM) reactor 9940, a YAM separation unit 9950 and an inert removal unit 9970.

[2365] ODH reactor 9910 includes an ODH catalyst capable of catalyzing, in the presence of oxygen which may be introduced via oxygen line 9980, the oxidative dehydrogenation of alkanes introduced via alkane line 9990. The ODH reaction may also occur in the presence of an inert diluent, such as carbon dioxide, nitrogen, or steam, that is added to ensure the mixture of oxygen and hydrocarbon are outside of flammability limits. Determination of whether a mixture is outside of the flammability limits, for the prescribed temperature and pressure, is within the knowledge of the skilled worker. An ODH reaction that occurs within ODH reactor 9910 may also produce, depending on the catalyst and the prevailing conditions within ODH reactor 9910, a variety of other products which may include carbon dioxide, carbon monoxide, oxygenates, and water. These products leave ODH reactor 9910, along with unreacted alkane, corresponding alkene, residual oxygen, carbon monoxide and inert diluent, if added, via ODH reactor product line 99100.

[2366] ODH reactor product line 99100 is directed to quench tower or acetic acid scrubber 9920 which quenches the products from product line 99100, and facilitates removal of acetic acid and water via quench tower botom outlet 99110. Unconverted lower alkane, corresponding alkene, unreacted oxygen, carbon dioxide, carbon monoxide, and inert diluent added to quench tower 9920 exit through quench tower overhead hue 99120 and a portion are directed into first oxidation reactor 9930 via first oxidation line 99130.

[2367] First oxidation reactor 9930 contains a group 11 metal with optional promoter and optional support as described above. First oxidation reactor 9930 optionally includes first oxygen line 99140 which can be used to provide an oxygen containing gas to first oxidation reactor 9930. In first oxidation reactor 9930, unreacted oxygen is reacted with carbon monoxide to form carbon dioxide and/or reacts acetylene to reduce or eliminate it. In first oxidation reactor 9930, most or all of the unreacted oxygen is consumed. The remaining unconverted lower alkane, corresponding alkene, unreacted oxygen (if present), all or part of the carbon dioxide, carbon monoxide (if present), and inert diluent are directed to first ethylene product line 99150 and combined with Site acetic acid and water via quench tower bottom outlet 99110 and conveyed to the VAM process.

[2368] In the VAM process, the contents of quench tower bottom outlet 99110. which can include acetic, acid, ethylene, and optionally unreacted ethane, unconsumed ox gen-containing gas, water, carbon monoxide, carbon dioxide, and inert compounds are fed to VAM reactor 9940. where acetic acid and ethylene are combined in the presence of a catalyst active for the production of vinyl acetate VAM reactor 9940 optionally includes VAM oxy gen line 99145 which can be used to provide an oxygen containing gas to VAM reactor 9940 as described herein. Depending on the scale of the process, VAM reactor 9940 may include either a single reactor or several reactors in parallel or in series. A VAM product stream 99300 that includes vinyl acetate, water, optionally ethane, gaseous by-products and unreacted acetic acid and ethylene is withdrawn from VAM reactor 9940 and is fed to VAM separation unit 9950 where a VAM gaseous stream 99330, which can include ethylene, and optionally ethane together with inert compounds, carbon monoxide and carbon dioxide is separated from water and acetic acid, which can alternatively be withdrawn overhead via first recycle line 99310 and can be recycled to and mixed with the contents of quench tower bottom outlet 99110 and provided to VAM reactor 9940 as described above.

[2369] VAM liquid stream 99320, which includes vinyl acetate, water, optionally unreacted acetic acid and optionally high boiling by-products of the process are withdrawn from the base of VAM separation unit 9950 and vinyl acetate is isolated in state of the art equipment not shown. As a non-limiting example, the contents of VAM liquid stream 99320 can be fed to a distillation column where vinyl acetate and water are removed as an azeotrope and acetic acid and the optional high boiling by-products are removed as a bleed from the base of the distillation column. The water in the overhead stream from the distillation column can be separated from the vinyl acetate in a decanter and a vinyl acetate product stream removed from decanter is purified by conventional means known in the art.

[2370] VAM gaseous stream 99330 can be directed to inert removal unit 9970, where carbon dioxide and other inert compounds are separated and directed to outlet stream 99340 and ethylene and optionally ethane and remaining carbon monoxide are directed to ethylene containing stream 99350 and combined into quench tower overhead line 99120.

[2371] Quench tower overhead line 99120 is fed to seeondaty oxidation reactor 99500, which contains a group 11 metal with optional promoter and optional support as described above. In seeondaty oxidation reactor 99500, unreacted oxygen is reacted with carbon monoxide to form carbon dioxide. In secondary oxidation reactor 99500, most or all of the unreacted oxygen is consumed. Secondary oxidation reactor 99500 optionally includes oxygen line 99505 which can be used to provide an oxygen containing gas to secondary oxidation reactor 99500

The remaining unconverted lower alkane, corresponding alkene, unreacted oxygen (if present), all or part of the carbon dioxide, carbon monoxide (if present), and inert diluent are directed to primary ethylene product line 99410. [2372] As shown in Figure 102, in the ethylene purification process either of primary ethylene product line 99410 (as shown in Figures 100 and 101) or quench tower overhead line 99120 (as shown in Figure 99) are provided to tire ethylene purification process, winch can include an amine wash lower 99510, a drier 99520, a distillation tower 99530, and an oxygen separation module 9960. In many embodiments, primary ethylene product line 99410 or quench tower overhead line 99120 contain unconverted ethane, eth lene, optionally unreacted ox gen, all or part of the carbon dioxide, optionally carbon monoxide and inert diluent, which are conve ed to amine wash tower 99510.

[2373] Any carbon dioxide is isolated by amine wash tower 99510. and captured via carbon dioxide bottom outlet 99540 and may be sold, or, alternatively, may be recycled back to ODH reactor 9910 via downstream recycle Sine 99550 as shown in Figures 99, 100 and 101. Constituents introduced into amine wash tower 99510, other than carbon dioxide, leave amine wash tower 99510 through amine wash tower overhead line 99560 and are passed through dryer 99520 before being directed to distillation tower 99530 via dryer line 99525, where C2/C2 ₊ hydrocarbons are isolated and removed via C2/C2 ₊ hydrocarbons bottom outlet 99570. The remainder includes mainly Ci hydrocarbons, including remaining inert diluent and carbon monoxide (if any), which leave distillation tower 99530 via overhead stream 99160 and is directed to oxy gen separation module 9960.

[2374J Oxygen separation module 9960 includes a sealed vessel having a retentate side 99170 and a permeate side 99180, separated by oxygen transport membrane 99190. Overhead stream 99160 may be directed into either of retentate side 99170 or permeate side 99180. Optionally, a flow controlling means (an example flow' controlling means 6326 is shown in Figure 63D) may be included that allows for flow into both sides at varying levels. In that instance an operator may choose what portion of the flow from overhead stream 99160 enters retentate side 99170 and what portion enters permeate side 99180. Depending upon conditions an operator may switch between the two sides, to allow equivalent amounts to enter each side, or bias the amount directed to one of die two sides. Oxygen separation module 9960 also includes air input 99200 for the introduction of atmospheric air, or other oxygen containing gas, into retentate side 99170. Combustion of products introduced into retentate side 99170. due to the introduction of oxygen, may contribute to raising the temperature of oxygen transport membrane 99190 to at least about 850 °C so that oxygen can pass from retentate side 99170 to permeate side 99180. Components within the atmospheric air, or other oxygen containing gas, other than oxygen, cannot pass from retentate side 99170 to permeate side 99180 and can only leave oxygen separation module 9960 via exhaust 99210.

[2375] As a result of oxygen passing from retentate side 99170 to permeate side 99180, there is separation of oxygen from atmospheric air or other oxygen containing gas, introduced into retentate side 99170. The result is production of oxygen enriched gas on permeate side 99180 which is then directed via oxygen enriched bottom line 99220 to ODH reactor 9910, either directly or in combination with oxygen line 9980 (as shown in Figures 99, 101 and 102). When overhead stream 99160 is directed into retentate side 170 the degree of purity of oxygen in oxygen enriched botom line 99220 can approach 99% Conversely, when overhead stream 99160 is directed into permeate side 99180 the degree of purity of oxygen in oxygen enriched bottom line 99220 is lower, with an upper limit ranging from 80% - 90% oxygen, the balance in the form of carbon dioxide, water, and remaining inert diluent, all of which do not affect the ODH reaction as contemplated by the present disclosure and can accompany the enriched oxygen into ODH reactor 9910. Water and carbon dioxide can be removed by quench tower 9920 and amine wash tower 99510, respectively. In some embodiments, some or all of the carbon dioxide can be captured for sale as opposed to being flared where it contributes to greenhouse gas emissions in some embodiments, when carbon dioxide is used in the ODH process, any carbon dioxide captured in the amine wash can be recycled back to ODH reactor 9910.

[2376] Oxygen transport membrane 99190 is temperature dependent, only allowing transport of oxygen when die temperature reaches at least about 850 °C. In some embodiments, the components in overhead stream 99160 by themselves are not capable, upon combustion in die presence of oxygen, to raise the temperature of oxygen transport membrane 99190 to die required level. In this embodiment, die chemical complex also includes fuel enhancement Sine 99230, upstream of oxygen separation module 9960, where combustible fuel, as a non-limiting example methane, may be added to supplement the combustible products from overhead stream 99160.

[2377] In some embodiments, the ox gen separation module 9960 is a tube (examples are depicted schematically in Figures 63 A through 63D). The oxygen transport membrane 99190 can be a tube (for example, tube 6319) and can fit inside a larger tube (for example, 6327) which forms the outer wall of oxygen separation module 9960. The annular space between the larger tube and oxy gen transport membrane 99190 corresponds to the retentate side 99170, while the space within oxygen transport membrane 99190 corresponds to the permeate side 99180. Material suitable for construction of Site outer wall include those resistant to temperatures that exceed 850 °C and approach 1000 °C. selection of which falls within the knowledge of the skilled worker.

[2378] The present disclosure contemplates the inlet for the overhead stream 99160 entering the oxygen transport module 9960 into either of the permeate side 99170 (an example is depicted schematically in Figure 63 A) or the retentate side 99180 (an example is depicted schematically in Figure 63B). In some embodiments oxygen separation module 9960 can have a Ci hydrocarbon containing line (for example, line 29) directed to the retentate side 99180. The present disclosure also contemplates the use of a valve (for example, valve 6326) for switching between directing the overhead stream 99160 to the retentate side 99180 or the permeate side 99170 (Figure 63D). This would allow' an operator to choose which of the sides, permeate or retentate, that the overhead stream is directed to.

[2379] In one embodiment there is a flooded gas mixer 99240 (Figure 103) upstream of ODH reactor 9910 (Figures 99, 100 and 101). In this instance oxygen line 9985 and ethane line 9995 feed directly into flooded gas mixer 99240. A homogeneous mixture that includes hydrocarbon and oxygen, and optionally an inert diluent, can be introduced into ODH reactor 9910 from flooded gas mixer 99240 via mixed line 99250 (Figure 103). Oxygen enriched bottom line 99220 may feed directly into or in combination with oxygen line 9985 into flooded gas mixer 99240. Mixed line 99250 can enter ODH reactor 9910 at either or both of oxygen line 9980 and ethane line 9990. [2380] The temperature of the contents within product line 99 i 00 (Figures 99, 101 and 102) in a typical ODH process can reach about 450 °C. It can be desirable to lower the temperature of the stream before introduction into quench tower or acetic add scrubber 9920 as described above. In that instance, the present disclosure contemplates the use of a heat exchanger immediately downstream of each ODH reactor 9910 and immediately upstream of quench tower 9920. Use of a heat exchanger to lower temperatures in this fashion is well known in the art.

[2381] In some embodiments, a concern for ODH processes is the mixing of a hydrocarbon with oxygen. Under certain conditions the mixture may be unstable and lead to an explosive event. U.S. published patent application No. 2018/0009662 (‘662 application) published January 11, 2018, titled “Inherently Safe Oxygen/Hydrocarbon Gas Mixer”, discloses a means to mix a hydrocarbon containing gas with an oxygen containing gas in a flooded mixing vessel. By mixing in this way pockets of unstable compositions are surrounded by a non-flammable liquid so that even if an ignition event occurred it would be quenched immediately. Provided addition of the gases to the ODH reaction is controlled so that homogeneous mixtures fall outside of the flammability envelope, for Site prescribed conditions with respect to temperature and pressure, the result is a safe homogeneous mixture of hydrocarbon and oxygen. The provided techniques may be supplemented with a flooded gas mixer as described in the ‘662 application.

[2382] The vinyl acetate monomer produced using the methods, apparatus and complexes disclosed herein can be polymerized to provide polyvinyl acetate (PVA). Vinyl acetate monomer can also be copolymerized with other monomers to provide various copolymers such as ethylene-vinyl acetate (EVA), vinyl acetate-acrylic acid (VA/AA), polyvinyl chloride - vinyl acetate (PV CA), and vinylpyrrolidone - vinyl acetate (Vp/Va Copolymer). In the above described non-limiting examples, the incorporated vinyl acetate repeat unit can be hydrolyzed to provide the corresponding vinyl alcohol repeat unit. As non-limiting examples, PVA can be partially hydrolyzed to provide a vin l acetate - vinyl alcohol copolymer, completely hydrolyzed to provide polyvinyl alcohol, EVA can be partially h drol zed to provide an ethylene - vinyl acetate - vin l alcohol terpolymer or completely hydrolyzed to provide an eth lene - vin l alcohol copolymer.

[2383] In some embodiments, the vinyl alcohol repeat units can be reacted with buty raldehyde to form vinyl butyral repeat units. When all of the vinyl alcohol units have been reacted to form vinyl butyral repeat units, the resulting polymer is polyvinyl butyral or PVB.

[2384] EXAMPLES

[2385] Example Acl (ODH Process)

[2386] The effect of altering the amount of steam injected into an ODH process on the carbon dioxide output was demonstrated using two fixed bed reactors, connected in series. The catalyst present in each of the reactors was a mixture of several batches of a mixed metal oxide catalyst of the formula: Mo1.0V0.3cc0.50Te0.10-0.20Nb0.10-0.20Od, where the subscripts represent the range of atomic amounts of each element, relative to Mo, present in the individual batches, and d represents the highest oxidation state of the metal oxides present in the catalyst. Ethane, carbon dioxide, and oxygen were premixed before addition of water, followed by preheating with the entire composition being fed to the first of the two reactors. The preheating step was necessary to ensure the water added was converted to steam before injection into the reactor. Output from the first reactor was sent directly into the CO Oxidation Reactor without addition of new reactants. For each reactor, the temperature was held in the range of 334-338 °C at ambient pressure. The process was ran continuously over a period of three days.

[2387] The relative amounts of ethane, carbon dioxide, and oxygen remained the same while the flow rate of steam added to reactor was altered. The relative amounts of ethane, carbon dioxide, and oxygen added to the first reactor were 33, 54, and 13 respectively. The gas hourly space velocity (GHSV) was kept constant at about 610 Irf Flow rates of reaction ethane, carbon dioxide and oxygen were altered accordingly to maintain a gas hourly space velocity of about 610h ^L after altering the amount of steam added to reactor.

[2388] Steam was added indirectly as water with the ethane, carbon dioxide and oxygen mixture. The amount of water added to the mixture before entering the first reactor was varied, starting with no water and increasing in increments up to a flow rate of I.OcmVmin. For each flow rate of water added to the mixture, a corresponding weight % of steam in the total feed mixture was calculated. Table Acl shows the effect that clanging tire amount of steam added to the reactor had on output of carbon dioxide, carbon monoxide, and acetic acid. The output of the components was measured as normalized selectivity, according to the formula: X selectivity (Wt. %) = net mass flow rate X (g X/hr) / ( mass flow rate C ₂ K0 (g C _/ K

CiRo molecular weight (g CJH& molecular weight X

N mol equivalent of com 1 mol C ₂ H6

In this formula, X refers to one of ethylene, CO2. CO, and acetic acid, and N refers to mol. equivalent of the compound X.

[2389] Results listed in Table 1 were averaged from two or more experimental inns at each of the prescribed conditions. The results demonstrate that increasing the flow rate of water added to die mixture and corresponding increase in the weight % of steam added to the reactor led to a decrease in the carbon selectivity. A carbon dioxide negative process was seen when the water was added at a flow rate of 1.0 cmVmin, which corresponds to 39 weight % of steam added. In addition, reverting back to no steam added followed by increasing to 39 weight % resulted in the carbon dioxide selectivity going positive back to negative. Finally, it should be noted that increasing the steam resulted in a higher production of acetic acid and also was accompanied by a higher conversion rate of ethane.

[2390] TABLE Aci : Normalized Selectivity of ODH Products in Response to Changes in Steam Added to the

Reactor

[2391| An acetic acid stream was created by condensing acetic acid from the stream produced from the ODH processes shown in Table Ac I, which also provided an ethylene stream containing the more volatile components. [2392] Examples Ac2-Ac6 (CO Selective Oxidation Process)

[2393] Experimental Reactor Unit (ERU) Setup

[2394] The ERU was used to produce feed gas for evaluating the catalysts. The apparatus (an example 400 is depicted schematically in Figure 4) consisted of fixed bed tube reactor 402, which was surrounded by two-zone electric heater 404. Reactor 402 is a 316L stainless steel tube which had mi outside diameter of 0.5 inches (about 1.25 cm) and inside diameter of 0.4 inches (about 1 cm) and a length of 14.96 inches (about 38 cm). Two main feed gas lines were attached to reactor 402; one line 406 was dedicated for a bulk nitrogen purge gas and the other line 408 was connected to a dual solenoid valve, which could be switched from ODH process feed gas (gas mixture of ethane/ oxygen/Nitrogen at a molar ratio of about 36/18/46) to compressed air when regenerating catalyst bed 414. [2395] For safety' reasons, the unit was programmed in a way that prevents air from mixing with the feed gas. This was accomplished through safely interlocks and a mandatory 15-minute nitrogen purge of the reactor when switching between feed gas 406 and air 412, The flow of gases is controlled by mass flow controllers. A 6-point thermocouple 416 was inserted through reactor 402, which was used to measure and control the temperature within catalyst bed 414. The catalyst was loaded in the middle zone of reactor 402 and located in between points 3 and 4 of thermocouple 416, which were the reaction temperature control points. The remaining 4 points of thermocouple 416 were used for monitoring purposes. Catalyst bed 414 consisted of a 1 : 1 volume ratio of cataly st to quartz sand, a total of 3 nil. The rest of reactor 402, below and above catalyst bed 414 was packed with 100% quartz sand and die load was secured with glass wool on the top and die bottom of reactor 402. A glass tight sealed condenser 418 was located after reactor 402 at room temperature to collect water/acidic acid, and the gas product could flow' to either vent 420 or sampling loop/vent 422 by a dnee-way solenoid valve.

[2396] CO Selective Oxidation Catalyst Testing Reactor

[2397] A 316L stainless steel tube with the following dimensions was used to test CO selective oxidation catalysts:

Outside diameter: 0.25 inches (about 0.63 cm)

Wall thickness: 0.028 inches (about 0.07 cm)

Catalyst bed height: 2 inches (about 5 cm)

[2398] The total weight of the catalyst was recorded for each catalyst that was tested. The flow of gases is controlled by the mass flow controllers on ERU. The product gas from ERU is directly fed in to the CO selective oxidation catalyst testing reactor (“Testing Rector’). The Testing Reactor was placed in a precision heating oven, in which the temperature w'as controlled within less than 0 5 °C. There were no thermocouples inside the reactor catalyst bed itself, as a result, the oven temperature w'as recorded as the catalyst testing temperature. The catalyst bed consisted of approximately 1 g of catalyst supported between two layers of quartz wool. The effluent from the reactor w'as continuously provided for gas chromatography analysis.

[2399] AgCe on Silica Catalyst Sample

[2400] SYLOPOL ^® 2408 silica (W.R. Grace, surface area: 316 m ² /g, pore volume: 3 54 cc/g, 20 g) w'as impregnated with a solution (40 mi) of Ce(N0 ₃ ) ₃ -6H ₂ 0 (2.80 g) and X. The impregnated silica was dried at 90 °C overnight and was calcined in air at 500 °C for 6 hours.

[2401] X = AgNC> ₃ , 103 mL of 0.1N solution. The solution was concentrated to about 20 ml and mixed with Ce(N0 ₃ ) _3. 6H ₂ 0. Distilled water was added to make 40 ml.

[2402] The catalyst made was CeAg oxide on silica with Ce0 ₂ : 5wt. %, Ag: 5wt. %.

[2403] CuCe on Silica Catalyst Sample

[2404] SYLOPOL 2408 silica (20 g) was impregnated with a solution (40 ml) of (Vi n.o! I _. Ό (2.80 g) and Y. The impregnated silica was dried at 90 °C overnight and 'as calcined in air at 500 °C for 6 hours. [2405] Y= CU(CH;;COO) ₂ , 3.17 g. The solution was concentrated to about 20 m3 and was mixed with Ce(N0 ₃ ) ₃ .6H ₂ 0. Distilled water was added to make 40 ml.

[2406] The catalyst made was CeCu oxide on silica with CeG ₂ : 5wt. %, Cu: 5wt. %.

MnCe on Silica Catalyst Sample

[2407] SYLOPOL 2408 silica (20 g) was impregnated with a solution (40 ml) of Ce(N0 ₃ ) ₃ .6H ₂ 0 (2.80 g) and Z. The impregnated silica was dried at 90 °C overnight and was calcined in air at 500 °C for 6 hours.

[2408] Z= MnCl ₂ .4H ₂ 0, 4.0 g. The solution was concentrated to about 20 ml and was mixed with Ce(N0 ₃ ) _3. 6H ₂ 0. Distilled water was added to make 40 ml.

[2409] The catal st made was CeMn oxide on silica with Ce0 ₂ : 5wt. %, Mn: 5wt. %.

[2410] CrCe on Silica Catalyst Sample

[2411] SYLOPOL 2408 silica (20 g) was impregnated with a solution (40 ml) of Ce(N0 ₃ ) ₃ .6H 0 (2.80 g) and W. The impregnated silica was dried at 90 °C overnight and was calcined in air at 500 °C for 6 hours.

[2412] W= Cr(N0 ₃ ) ₃ .9H ₂ 0, 6.98 g. The solution was concentrated to about 20 ml and was mixed with Ce(N0 ₃ ) ₃ .6H ₂ 0 Distilled water was added to make 40 ml.

[2413] The catalyst made was CeCr oxide on silica with Ce0 : 5wt. %, Cr: 5wt. %.

[2414] Example Ac2

[2415] AgCe on Silica Catalyst Testing

[2416] The ODH process was ran using the ERU and catalyst MoVOx to provide the feed for this example.

0.15 g of AgCe catalyst was used for this test at a gas hourly space velocity of approximately 5000 h-1, 0 psig on the reactor outlet, at 75 °C process temperature. The results are shown in Table Ac2 below.

[2417] TABLE Ac2

[2418] The data show that the AgCe catalyst demonstrates excellent oxy gen removal properties via selective oxidation of CO to C0 ₂ , which can be seen from noticeable reduction of CO in the product gas and increase in all of the other compounds it is noteworthy that acetylene is also fully oxidized and was not detected in the product gas from the Testing Reactor.

[2419] Example Ac3

[2420] 1.7 g of AgCe catalyst was used for tins test, with approximately 1 g of the catalyst present in the hot zone of the reactor. 1 g was the value for the catalyst weight used for the long term test calculations. In this example, the ODH catalyst used to produce the feed for this example was a MoVOx based catalyst used as described in Example Ac2. In this example, the feed sample to the selective CO oxidation reactor was taken twice at the beginning and at the end of the test in order to confirm the composition of Site feed. The test was executed at 110 °C process temperature 0 psig reactor outlet pressure, gas hourly space velocity of approximately 3000 h ¹ . The results are summarized in Table Ac3 below.

[2421J TABLE Ac 3

O removed” value is calculated as follows: where V ₀₂ is the value of “CL removed” and C is the volumetric concentration of oxygen in the feed and product gasses

[2422] Because the composition of the feed to the selective CO oxidation reactor was changing gradually over the term of the experime nt, accurate values for removed oxygen could only be calculated at the vesy beginning and at the very end of the run. The data show that even though the catalyst had very stable activity toward acetylene oxidation through the whole duration of the ran, the activity toward O? removal via selective CO oxidation gradually decreased over the duration of the run. Generally, the amount of CO and 0 ₂ in the product stream was less than in the feed stream and the amount of C0 ₂ in the product stream was greater than the amount in the feed stream.

[2423] Example Ac4

[2424] The ODH process of Example Ac2 was used to provide the feed for this example. 0.35 g of AgCe catalyst, regenerated via oxidation, was used for this test. The test was executed at 110 °C process temperature, 0 psig reactor outlet pressure, gas hourly space velocity of approximately 3000 h ¹ . The results are shown in Table Ac4 below.

[2425] Table Ac4

[2426] The data show that the AgCe catalyst was successfully regenerated. Acetylene was reduced to undetectable levels and oxygen levels in the product stream were less than in the product stream.

[2427] Example Ac5

[2428] The ODH process of Example Ac2 was used to provide the feed for tins example. 1.22 g of fresh CuCe catalyst was used for this test. The test was executed at 120 °C process temperature, 0 psig reactor outlet pressure, gas hourly space velocity- of approximately 3000 h ¹ . The results are summarized in Table Ac5 below-.

Table Ac5

[2430] The data show that the CuCe catalyst exhibits similar properties to the AgCe catalyst. CuCe catalyzes the reaction of selective oxidation of CO to CO2 and oxidation of acetylene. However, this catalyst sample did not show any catal st activity at a temperature of 110 °C, which is noticeably higher than 75 °C. at which fresh AgCe catal st revealed significant activity toward selective oxidation of CO.

[2431] Example Ac6

[2432] The ODH process of Example Ac2 was used to provide the feed for this example. 1.01 g of MnCe catal st was used for this test. The test was executed at 140 °C process temperature, 0 psig reactor outlet temperature, gas hourly space velocity of approximately 3000 h ¹ . The results are summarized in Table Ac6 below.

[2433] Table Ac6 [2434] The data show that the MnCe catalyst exhibits selective CO oxidation properties however, it did not demonstrate any activity' toward oxidation of acetylene. This catalyst sample did not show any catalytic activity at the temperature below 140 °C, which is noticeably higher than 75 °C, at which AgCe catalyst revealed significant activity toward selective oxidation of CO.

[2435] Example Ac7 (Vinyl Acetate Monomer Process)

[2436] Catalyst Preparation

[2437] A conventional palladium-gold-potassium acetate catalyst on an alumosilicate support is produced.

The support particles me spherical with a diameter of approximately 5 m and lave a specific surface area of 160 to 175 in ² /g. a bulk density of about 600 g/1 and a total pore volume of about 0.68 cm ³ /g. The concentrations of the impregnating solutions are selected in such a mariner that the finished catalyst contains about 3.3 g palladium, 1.5 g gold and 30 g potassium acetate per liter bulk volume of the catalytic support, which corresponded to a concentration of about 0.55% by weight palladium, about 0.25% by weight gold and about 5% by weight potassium acetate relative to the weight of the support used.

[2438] In a first step the support is first impregnated with a solution of potassium hydroxide. The concentration of the solution of potassium hydroxide is calculated so that after the impregnation a stoichiometric excess of potassium hydroxide on the support of about 620% is present.

[2439] After drying the catalytic supports they are impregnated with an aqueous solution of tetrachloroauric acid and potassium palladium chloride. After 20 hours the insoluble noble-metal compounds are reduced in the aqueous phase with hydrazine hydrate for a period of about 4 hours. Then the catalytic supports are washed free of chloride and dried before they are impregnated with a potassium acetate solution and redried. Before the impregnation with potassium acetate the specific surface area of the catalyst according to DIN 66 132 is about 60-70 m ² /g. Due to the impregnation and activation with potassium acetate the specific surface area of the catalyst declines further to 41 nr /g.

[2440] The CO adsosption of the catalyst before activation is approximately 0.158 ml CO/g catalyst. The particle crush strength of the activated catalyst is about 48N (in radial measuring). The thickness of its outer shell containing noble metal is about 0.3 mm.

[2441] Vinyl Acetate Monomer Synthesis

[2442] The catalyst describer above is used in an oil-heated tubular-flow reactor (reactor length 800 mm, inner diameter 24.8 mm) at normal pressure and a space velocity (GHSV=gas hourly space velocity) of 400 h ¹ with the following gas composition: 76.0% by volume ethylene, 18.0% by volume acetic acid, 6.0% by volume oxy gen. [2443] The reactor temperature is adjusted so that the temperature in the middle of the catalytic bed is between 150° and 160 °C. The resulting vinyl acetate monomer is condensed and recovered in the reactor outlet. [2444] Example 7 ( Vinyl Acetate Monomer Process - Asepn Plus ^® Simulation)

[2445] Aspen Plus simulation (version 9.0, Aspen Technology, Inc. ) was used to simulate an experiment where acetic acid, eth lene, <¼ and N ₂ were fed to a catalytic vinyl acetate reactor. The property method used was WILS-LR. The block flow' diagram of this simulation is shown in Figure 104. Simulation 800 consisted of Mixer block 10410, heater block 820 and VAM reactor 10430. It was assumed that ethylene feed 10440, acetic acid feed 10450 and a postion of O2 feed 10460 were provided from a product stream of a catalytic ODH ethane to ethylene reactor. A portion of (¾ feed 10460, was provided from an external source. N ₂ feed 10470 was assumed to be representing an inert gas added to the total feed mixture to keep it outside of the flammable envelope of the acetic acid-ethylene-oxygen mixture. The following characteristics of each stream were assumed (temperature, pressure, rate, vapor fraction):

Ethylene feed 10440 (120 °C, 243 kPa, 24,968 Kg/hr, 100% vapor fraction)

0 ₂ feed 10460 (120 °C, 243 kPa, 16,749 Kg/hr, 100% vapor fraction)

Acetic acid feed 10450 (128 °C, 243 kPa, 29,426 Kg/hr, 0% vapor fraction)

N ₂ feed 10470 (120 °C. 243 kPa, 24,968 Kg/hr, 100% vapor fraction)

Feed to preheater 10480 (111 °C, 243 kPa. 96,111 Kg/hr, 100% vapor fraction)

Feed to YAM reactor 10490 (160 °C. 228 kPa, 96,111 Kg/hr, 100% vapor fraction)

Product stream 104900 (160 °C, 193 kPa, 96,111 Kg/hr, 100% vapor fraction)

[2446] The energy /mass balance table resulted from this simulation is shown in Table Ac7 below.

[2447] Table Ac7

[2448] The data show that vinyl acetate monomer (VAM) is produced in the process.

[2449] An embodiment is a method that includes (a) providing a stream containing ethane and oxygen to an ODH reactor; (b) converting a portion of the ethane to ethylene and acetic acid in the ODH reactor to provide a stream containing ethane, ethylene, acetic acid, oxygen and carbon monoxide; (c) separating a portion of the acetic acid from the stream to provide an acetic acid stream and a stream containing ethane, ethylene, oxygen and carbon monoxide; (d) providing the stream to a CO Oxidation Reactor containing a catalyst that includes a group 11 metal to convert carbon monoxide to carbon dioxide and reacting acetylene to produce a stream containing ethane, ethylene and carbon dioxide; and (e) providing a portion of the stream and a portion of the acetic acid stream to a third reactor containing a catalyst that includes a metal selected from group 10 and group i 1 metals to produce vinyl acetate.

[2450] Techniques are provided for converting alkanes to alkenes, especially ethane to ethylene, with acetic acid as a byproduct of a first process. Techniques are provided for producing vinyl acetate from the ethylene and acetic acid produced from the first process in a second or serial process. The ethylene can be used to produce homopolymers, copolymers, copolymer compositions and methods of making the same. The vinyl acetate can be used to produce various materials, such as poly vinyl acetate, poly vinyl alcohol, vinyl acetate - vinyl alcohol copolymers, polyethylene-vinyl acetate copolymer, vinyl acetate-acrylic acid copolymer, polyvinyl chloride - vinyl acetate copolymer, vinylpyrroiidone vinyl acetate copolymer, any of the above where at least a portion of the vinyl acetate repeat units are hydrolyzed to vinyl alcohol repeat units, vinyl butyra! and polyvinyl butyral.

[2451] Carbon dioxide may be generated in the conversion of lower alkanes (e.g., ethane) into corresponding alkenes (e.g., ethylene). Carbon dioxide (CO _?. ) is the primary greenhouse gas emitted through human activities. [2452] Methods described herein provide for the catalytic conversion of C0 ₂ in the presence of hydrogen into acetic acid or carbon monoxide. Provided are a system and method for converting C0 ₂ into products by contacting the CO2 with catalyst in the presence of hydrogen in a reactor.

[2453] Techniques are provided for converting C(¾ in the presence of hydrogen (¾) to acetic acid (C2H4O2) or carbon monoxide (CO). A reactor having catalyst (e.g., GDH catalyst) performs the conversion. While certain ODH catalysts can be employed, the CO2 conversion reactor may generally avoid performing an ODH reaction or significant ODH reactions with the ODH catalyst present instead, the ODH catalyst may facilitate reactions (e.g., h drogenation of CQ ₂ ) oilier than ODH, as discussed below. Further, the C0 ₂ conversion can be performed without h drocarbon feed to the reactor. The ODH catalyst(s) employed may be labeled mi “ODH” catalyst because the catalyst can be utilized in other processes to perform ODH to convert lower alkanes to corresponding alkenes. Embodiments of the present reactor may relate to the catalytic hydrogenation of C0 ₂ into acetic acid or carbon monoxide, or both.

[2454] The primary product of the C0 ₂ conversion may be acelic acid in presence of feed water (H ₂ 0) in the reactor. The primary product of the conversion may be CO in absence of feed H ₂ 0 in the reactor. The C0 ₂ in the feed to the reactor can be from an integrated s stem, such as an ethane steam-cracking system, an ODH reactor system that converts ethane to ethylene, and so on. The C0 ₂ conversion reactor that converts C0 ₂ may be labeled as an ODH reactor when the reactor lias ODH catalyst (for the C0 ₂ conversion) and not necessarily because the reactor performs an ODH reaction.

[2455] The ODH catalyst may be a low-temperature ODH catal st that provides for the conversion with reactions at less than 425 °C or less than 400 °C. As discussed below, reactions in the present C0 ₂ conversion reactor that convert the C0 ₂ may include (1) C0 ₂ hydrogenation and (2) w'ater gas shift. The reactions are via a cataly st, which can be an ODH cataly st in certain embodiments

[2456] Advantages of the present techniques may include converting C0 ₂ emissions into value-added products. Implementations may provide opportunity to integrate and convert C0 ₂ emissions from steam cracking systems or ODH reactor systems (that convert a lower alkane to a corresponding alkene) into desirable products, such as acetic acid or CO.

[2457] FIG. 105 depicts a reactor system 10500 having a C0 ₂ converter or C0 ₂ conversion reactor 10502 with a catalyst 10504 for the conversion of C0 ₂ into acetic acid or CO. The catalyst 10504 may be an ODH catalyst. In certain embodiments, the C0 ₂ conversion reactor 10502 can resemble aspects of a conventional ODH reactor that converts ethane to ethylene. The CO conversion reactor 10502 is configured to receive CO _¾ H ₂ , and H ₂ 0 as feed. In contrast, a conventional ODH reactor employing ODH catalyst receives lower alkanes (e.g., ethane), oxygen, and diluent as feed.

[2458] The C0 ₂ converter or conversion reactor 10502 may be a fixed-bed reactor (e.g.. a tubular fixed-bed reactor), a fhiidized-bed reactor, an ebullated bed reactor, or a heat-exchanger type reactor, and so on. The C0 ₂ conversion reactor s stem 10500 may utilize a heat-transfer fluid for controlling temperature of the reactor 10502. The heat-transfer fluid may be employed to add heat or remove heat from the C0 ₂ conversion reactor 10502 or from the reactor system 10500. The heat transfer fluid can he, for example, steam, water (including pressurized or supercritical water), oil, or molten salt, and so forth.

[2459] The reactions collectively in the reactor 10502 to convert CO2 are typically endothermic. Therefore, the heat transfer fluid is a heating medium. For alternate embodiments with the reactions collectively in the reactor 10502 to convert CO2 as exothermic, the heat transfer fluid is a cooling medium. The heating medium and the cooling medium (if employed) may be the same or different ripe of heat transfer fluid. Lastly, in some implementations, the heat transfer fluid may be a cooling medium or heating medium when the reactor 10502 is not in normal operation or is offline or shut down for maintenance activity. In a particular implementation, the heat transfer fluid as a cooling medium may be employed during online regeneration of the catalyst 10504.

[2460] For a fixed-bed reactor, reactants may be introduced into the reactor at one end and flow past an immobilized catal st. Products are formed and an effluent having the products may discharge at the other end of the reactor. The fixed-bed reactor may have one or more tubes (e.g., metal tubes ceramic tubes, etc.) each having a bed of catalyst and for flow of reactants. For tire reactor 10502, the flowing reactants may be CO2, ¾, and optionally H2O. The tubes may include, for example, a steel mesh. Moreover, a heat-transfer jacket adjacent the tube(s) or an external heat exchanger (e.g., feed heat exchanger or recirculation heat exchanger) may provide for temperature control of the reactor. The aforementioned heat transfer fluid may flow through the reactor jacket or external heat exchanger (e.g., sheJl-and-tube heat exchanger).

[24611 In other embodiments, the reactor 10502 is a fluidized bed reactor. In implementations, a fluidized bed reactor may have a support for the ODH catalyst. The support may be a porous structure or distributor plate and disposed in a botom portion of the reactor. Reactants may flow upward through the support at a velocity to fluidize the bed of ODH catal st. The reactants (e.g., CO2, ¾ and optionally 3¾0 for the reactor 10502) are converted to products (e.g , acetic acid or CO in the reactor 10502) upon contact with the fluidized catalyst. An effluent having products may discharge from an upper portion of the reactor. The fluidized bed reactor may have heat-transfer tubers, a jacket, or an external heat exchanger (e.g., feed heat exchanger or recirculation loop heat exchanger) to facilitate temperature control of the reactor. The aforementioned heat transfer fluid may flow' through the reactor tubers, jacket, or external heat exchanger.

[2462] The fluidized bed reactor can be (1) a non-circulating fluidized bed, (2) a circulating fluidized bed with regenerator, or (3) a circulating fluidized bed without regenerator. In the conversion of CO2 to acetic acid or CO, catalyst regeneration may not be typically needed and, therefore, a circulating fluidized-bed platform with regenerator may not be implemented in certain embodiments. However, in practice of embodiments with a reactor for dual purposes such as (1) CO2 conversion and (2) ethane ODH, a circulating fluidized bed with regenerator may be employed. If so, the downcomers to the catalyst regenerator section may be closed during operation of the reactor in the mode of CO2 conversion.

[2463] The catalyst 10504 may be operated as a fixed bed or fluidized bed. In some implementations, the catalyst 10504 in the CO2 conversion reactor 10502 is a first low-temperature ODH catal st that includes molybdenum, vanadium, tellurium, niobium, and oxygen, wherein the molar ratio of molybdenum to vanadium is from 1:0.12 to 1:0.49, the molar ratio of molybdenum to tellurium is front 1:0.01 to 1:0.30, the molar ratio of molybdenum to niobium is from 1 :0.01 to 1:0.30, and oxygen is present at least in an amount to satisfy the valency of any present metal elements. The molar ratios of moly bdenum, vanadium, tellurium, niobium can be determined by inductively coupled plasma mass spectrometry (ICP-MS). The catalyst may be low temperature in providing for CO2 conversion reactions at less titan 425 °C or less than 400 °C. This ODH catalyst can be implemented in the reactor 10502 without ODH reactions but instead with reactions for the CO2 conversion. The catalyst is available from NOVA Chemicals Corporation having headquarters in Calgary, Canada in other embodiments, any of the catalysts described herein may be used in addition to, or instead of this catalyst.

[24641 For example in sortie embodiments, tire catalyst 10504 is a second low-temperature ODH catalyst that is a mixed metal oxide having the formula Mo _a V _f cTe _c Nb _d Pd _e Or, where a, b, c, d, e, and f subscripts are relative atomic amounts of the elements Mo, V, Te, Nb. Pd, O, respectively. When a=l, thenb=G.Ql to 1.0. c=G to 1.0, d=Q to 1.0, 0.00<e<0.10, and f is a number to satisfy the valence state of the catalyst. The number f may be a number to satisfy at least the valence state of tire corresponding elements in the catalyst. This catalyst may provide for the CO2 conversion reaction(s) to occur at a temperature of less titan 400 °C or less than 425 °C. This catalyst is also available from NOVA Chemicals Corporation having headquarters in Calgaty, Canada.

[2465] The CO2 conversion reactor 10502 may have ODH cataly st and may be similar to a conventional ODH reactor that receives ethane, oxygen, and diluent and converts ethane to ethylene. However, the CO? conversion reactor 10502 may be situated and configured to not receive hydrocarbon or ethane for CO2 conversion, but instead to receive CO _? , H _? , and optionally H _? 0 as feed and with a focus to produce acetic acid or CO via the ODH catalyst (and not produce significant amounts of ethylene). The present reactor system 10500 includes conduits (piping) to route CO _? and IT _? (and optionally H _? 0) to a feed inlet nozzle(s) on the CO _? conversion reactor 10502.

[2466] A conduit may route the CO _? from a vessel storing CO _? or from a pipeline or conduit header conveying CO _? . A source of the CO _? may be for example, a steam-cracker furnace system (e.g , from flue gas of a steam cracker furnace) or from an ODH system that converts a lower alkane(s) to a corresponding aikene. For instance, the CO _? source may be an amine tower in the in the steam cracker furnace system or in the conventional ODH system. Thus, the reactor system 10500 may facilitate reduction of C02 emissions associated with those sources. Other sources of CO _? are applicable.

[2467] A conduit may route the H _? to the CO _? conversion reactor 10502 from a vessel storing H _? or from a pipeline or conduit header conveying H _? . The source of H _? may be, for example, a demethanizer distillation column or associated system. Other petrochemical sources of H ₂ , as well as water splitting, etc., are applicable sources of H _? .

[2468] A conduit may route the H _? 0 from, for example, a steam header or steam subheader in implementations, the steam as the entering H _? 0 10510 is low pressure steam at 150 pounds per square inch gauge (psig) or less. The steam may be, for instance, low' pressure steam generated within a conventional ODH s stem or with a steam cracker system. When acetic acid production is favored, a valve on the conduit conveying the H _? 0 10510 (e.g., steam) may be in an open position to allow' die steam 10510 to flow to and enter the CO _? conversion reactor 10502. In implementations liquid water is not added and liquid water does not come in contact with the catalyst bed so to avoid pulverizing die catalyst particles. Instead, steam 10510 may generally be added. [2469] The CO2 conversion reactor 10502 may also be arranged or configured differently than a conventional ODH reactor with respect to reactor temperature control or in the heating or cooling of the reactor. The conversion of ethane to ethylene in an ODH reactor may be generally exothermic. In contrast, the conversion of CO2 to acetic acid or CO in the CO2 conversion reactor 10502 may be endothermic. In some implementations of the reactor 10502 as an endothermic reactor that is a tubular fixed-bed reactor, the C(¾ conversion reactor 10502 may be configured with a preheater heat exchanger or with a reactor vessel jacket receiving a hea ting medium (e.g., steam or oil). Conversely, a conventional ODH reactor that is a tubular reactor may rely on receiving a cooling medium (e.g., oil or molten salt) to the reactor vessel jacket. Other reaction conditions and reactor configurations are applicable.

[2470] In operation, the CQ ₂ conversion reactor 10502 receives feed that includes carbon dioxide 10506 and hydrogen 10508. As indicated, the feed to die reactor 10502 may also include H?0 10510, such as steam. Exemplary reactions in die ODH reactor 10502 include [Adi] and [Ad2j below:

[Adi] reaction Adi (CO2 hydrogenation): 7H ₂ + 5C0 ₂ - 5H ₂ 0 + 3CO + CHfiO?

[Ad2] reaction Ad2 (water gas shift): TI ₂ + C0 _{2 <} H ₂ 0 + CO.

[2471] Acetic acid may be produced in the reactor 10502 by hydrogenation of C0 ₂ reaction as given in a bulked reaction Ad 1 above. CO may be produced in the reactor 10502 by water gas shift reaction as given in reaction Ad2 above. The presence of H ₂ 0 in the feed to the reactor 10502 may suppress the CO formation by pushing the water gas shift reaction (reaction Ad2) back towards C0 ₂ formation. Thus, the presence of H ₂ 0 in the feed may favor the production of acetic acid in reaction Adi . The absence of H ₂ 0 in the feed may favor the production of CO. In either case, C0 emissions may be reduced in source systems that provide C0 ₂ as feed to the reactor 10502 in certain implementations.

[2472] Reaction Adi is a bulked reaction that is the sum of multiple intermediate reactions. There may be as many as five reactions that sum to give the bulk reaction depicted as reaction Adi. The reaction Adi may be labeled as a simplified bulked reaction and with the actual reaction scheme more complex. Furthermore, the bulked reaction Adi does not represent the only bulked reaction that can explain acetic acid generation from H ₂ and C0 _2.

[2473] The effluent 10512 (e.g., product effluent) from the C0 conversion reactor 10502 may discharge to a condenser 10514 in the reactor system 10500. The motive force for flow of the effluent 10512 to the condenser may be by pressure differential between the reactor 10502 and the condenser 10514, and/or by a compressor (e.g., positive displacement or dynamic) disposed along the conduit conveying the effluent 10512 to the condenser, and the like in certain implementations, the flow' of the effluent 10512 may be modulated by the compressor (if employed) or by a control valve (not shown) along the conduit conveying the effluent 10512 to the condenser 10514. The control valve may control the flow rate (e.g., mass flow' rate or volumetric flow rate) of Site effluent 10512. In some embodiments, the control valve (if employed) may function as a backpressure regulator in controlling pressure in the reactor 10502. [2474] The effluent 10512 generally has products from the CO2 conversion reactor 10502 The effluent 10512 may include acetic acid and CO. With presence of ¾0 in the feed to the reactor 10502, the acetic acid may be the main or primary product in the effluent 10512. With absence of 3¾0 in the feed to the reactor 10502, the CO may be the main or primary' product in the effluent 10512. Additional products in the effluent 10512 may include ethane (C2H6) as a third product and ethylene (C2H4) as a fourth product. The effluent 10512 may also include CO2 (unreacted feed), H ₂ (unreacted feed), H ₂ 0 (diluent), and other compounds.

[2475] The condenser 10514 may be an air-cooler heat exchanger, a water-cooler heat exchanger, a quench tower or scrubber column, and so forth. In some implementations, the condenser 10514 is a shell-and-tube heat exchanger. If so, a heat-transfer (cooling) fluid may flow through the shell side and the effluent 10512 flows the tube side. On the other hand, the heat transfer fluid may flow through the tube side and the effluent 10512 flow's through the shell side. In particular implementations, the heat transfer fluid may be water, such as cooling tower water.

[2476] In operation in tire condenser 10514 as a heat exchanger, heat is transferred from the effluent 10512 to the heat transfer fluid. Components in the effluent 10512 may condense due to the heat transfer. The amount of heat transfer and condensation conditions may be affected by the tempe rature and flow rate of the cooling fluid. The effluent 10512 discharging from the condenser 10514 may be separated into liquid components 10516 and gas components 10518. The liquid components 10516 can be acetic acid and ¾0, which can be separated. Acetic acid can be separated from water, for example by azeotropic distillation, liquid-liquid extraction, and other separation techniques. The acetic acid can be sold, such as in glacial or dilute form.

[2477] The gas components 10518 can include CO (main product in absence of H ₂ 0 feed to reactor 10502), C _/ Th (3 ^rd product), C ₂ H ₄ (4 ^,a product), CO2 (imreacted feed), and 1T ₂ (umeacted feed). The gas components 118 may be separated and sold or sent to adjacent systems.

[2478] In implementations, the gas components 10518 may be sent to downstream of an acetic acid scrubber in an ODH system that converts ethane to ethylene. For example, the gas components 10518 may be sent to a separation train downstream of the acetic add scrubber. In other implementations, the gas components 10518 may be sent to downstream of a quench tower in a steam cracker system that converts ethane to ethylene. For example, the gas components 10518 may be sent to a separation train downstream of the acetic acid scrubber. The separation train in either the ODH system or the steam cracker system may include, for instance, an amine tower, a caustic tower (e.g., that removes C0 ₂ ), a demethanizer distillation (e.g., that separates H ₂ /CO/CH ₄ from C ₂ H ₆ /C ₂ H ₄ ), and a C ₂ splitter distillation column (e.g., that separates C ₂ H ₆ from C ₂ H ₄ ). Other configurations of the separation train are applicable.

[2479] Gas components 10518 may be recycled (optional) to the reactor 10502 as recycle 10520 though a recycle conduit. A portion of the gas components 10518 stream may be sent as recycle 10520 to the reactor 10502. The recycle 10520 of the gas components 10518 may increase conversion of C0 ₂ by the reactor 10502 and reactor system 10500. In implementations the recycle 10520 may be added to the feed entering the reactor 10502. The motive force for flow of the recycle 10520 may be via a compressor (e.g., positive displacement or dynamic) disposed along the recycle conduit, or via a compressor upstream of the condenser 10514, and so on. In some applications, the compressor may be low differential pressure blower labeled as a blower. Motive force may be provided without a compressor, such as by an ejector, eductor, jet, injection of motive fluid, and the iike, such as where a carrier fluid is available to operate the devices.

[2480] Lastly, the feed to the CO _? conversion reactor 10502 may be heated via a preheater 10522. In certain embodiments, the preheater 10522 is a heat exchanger, such as a shell-and-tube heat exchanger or other type of heat exchanger. The heating medium can be steam or oil, and the like.

[2481] The source of heat or energy for the preheating can be from an integrated system (e.g., system 10800 in FIG. 108). The source of heat or energy' for the preheating can be, for example, ODH reaction heat from an adjacent ODH reactor system that converts ethane to ethylene. Thus, in those implementations, an adjacent ODH reactor system can provide heat to drive die CO _? conversion reaction in the present reactor system. In other embodiments, a source of heat or energy for the preheating can be, for example from a steam cracking furnace that converts ethane to ethylene. For instance, a stack of tubes can be placed on the top of convection section of the steam cracking furnace to recover heat from furnace off gas to provide heat or energy to die present CO2 conversion reactor system. Other sources of waste heat may be applicable for preheating die feed to die CO2 conversion reactor. [2482] As indicated, both acetic acid and CO may be desired products. CO can be utilized (e.g., combusted) to generate steam. In addition, CO can be mixed with H _? to give sy nthesis gas (syngas). Syngas can then be converted to hydrocarbon-based fuels or methanol. There can be markets for methanol in various industries, such as plastic, automotive, paints and adhesives, construction, and pharmaceutical.

[2483] Acetic acid may be converted to vinyl acetate as a co-monomer for polyvinylchloride copolymer production. Acetic acid may be converted to ethanol via hydrogenation reaction for use of the ethanol, for example, as fuel. Further, acetic acid may be converted to ethylene. For example, ethylene may be produced via a two-stage process of acetic acid hydrogenation followed by ethanol dehydration for use of the ethylene in polyethylene synthesis.

[2484] FIG. 106 is a method 10600 of processing CO2 (e.g., conversion of CO _? ) with ODH catalyst, such as in a reactor system having a reactor with ODH catalyst. The reactor system may be characterized as an ODH reactor system because a reactor in the system employs the ODH catalyst and not necessarily that ODH reactions occur or are performed in the reactor or in the CO2 conversion

[2485] At block 10602, the method includes providing feed having CO _? to the reactor. The feed includes H _? and optionally water (H _? 0). If H _? 0 is fed to the reactor, the H _? 0 may be fed as steam (e.g., low pressure steam less titan less than or equal to 70 psig) to the reactor in certain embodiments. In implementations, the feed may be heated (preheated) in a heat exchanger (e.g., a shell-and-tube heat exchanger) prior to introduction of the feed to the reactor. Alternatively, the preheating of the feed may be performed via an inert bed before the catalyst bed. Other preheating techniques are applicable. In some implementations, the preheating of the feed may facilitate temperature control in Site reactor.

[2486] At block 10604, the method includes contacting the CO _? with the ODH catalyst in the reactor in the presence of the FI _? to convert the CO2 into products, such as acetic acid or CO. Secondary products as byproducts may include, for example, ethane and ethylene. Acetic acid may be produced by hydrogenation of CO _? reaction as given in reaction Adi above. CO may be produced by water gas shift reaction as given in reaction Ad2 above. As discussed, the presence of ¾0 in the feed provided to the reactor 10502 may inhibit the CO formation by moving (reversing) the water gas shift reaction (reaction Ad2) to C0 ₂ formation. Thus, the addition of ¾0 into the feed may promote the production of acetic acid in reaction Adi. The absence of H ₂ 0 in tire feed may promote the production of CO.

[2487] The method may include maintaining an operating temperature of the reactor at less than 425 °C or less than 400 °C. in some implementations, less than 400 °C is implemented to avoid the auto-ignition temperature of acetic acid or to avoid moving equilibrium of the gas shift reaction back to C0 ₂ instead of CO, and the like. i2488| In implementations, the reactor operating pressure may be less titan 80 pound per square inch gauge (psig), less titan 70 psig. or less Ilian 60 psig. The operating pressure may be in a range of 5 psig to 80 psig. Operating pressures outside of this range are applicable.

[2489] Oilier operating conditions of the reactor in embodiments of the reactor as a tubular fixed-bed reactor may be gas hourly space velocity (GHSV), for example, in the range of 100 hour ¹ to 5,000 hour ¹ , or 100 hour ¹ to 10.000 hour ¹ . Linear velocity range of the feed through the reactor may be at least 5 centimeters per second (cm/sec). The linear velocity may be Q/A* e, where Q is the volumetric flow rate of the feed, A is the cross- sectional flow area (based on inner diameter) of the reactor tube, and e is the void space ratio (dimensionless) of the catalyst bed. The void space ratio is the volume of the void space in the catalyst bed divided by the total volume of the catal st bed. The volumetric flow rate of the feed is the volume of the feed passing through the catalyst bed in units of volume per time.

[2490] The reactor operating temperature may be the temperature at which the catalyst (e.g., ODH catalyst) drives C0 ₂ conversion reactions (e.g., C0 ₂ hydrogenation, water gas shift, etc.) in the reactor and which may be maintained by reactor temperature control. The temperature referenced may be the weighted average temperature of the reactor or reactor catalyst bed, e.g., over the temperature profile from reactor inlet to reactor outlet. The reactor operating temperature as referenced may incorporate reactor peak temperatures, and so forth.

[2491] At block 10606, the method includes discharging a product effluent from the reactor to a condenser. The product effluent may include acetic acid and CO. With presence of H ₂ 0 in the feed to the reactor, the acetic acid may be the foremost or majority product in the effluent. Additional products may include CO, ethane, and ethylene but at less amounts than acetic acid. With absence of ¾0 in the feed to the C0 ₂ conversion reactor, the CO may be the foremost or majority product in the effluent. Additional products may include acetic, ethane, ethylene, and methane but at less amounts than CO. In either case (with or without H ₂ 0 in the feed), the effluent may also include unreacted C0 ₂ , unreacted H ₂ , and H ₂ 0.

[2492] The method may include condensing effluent components, such as acetic acid and water, in the condenser. The condenser may be a heat exchanger, such as a shell-and-tube heat exchanger, plate heat exchanger, plate-and-frame heat exchanger, air-cooled heal exchanger (e.g.. finned tube), or other type of heat exchanger. The cooling medium may be, for example, water, air, molten salt, glycol, oil. and so forth.

[2493] At block 10608, the method include discharging a liquid product stream from the condenser and also discharging a gas product stream from the condenser. Alternatively, a single product stream may be discharged from the condenser and the single product stream separated downstream of the condenser into a liquid product stream and a gas product stream.

[2494] The liquid product stream can include acetic acid and H ₂ 0. In implementations, the acetic acid can be separated from the H ₂ 0 and sold. The liquid product stream can also be sent to an ODH reactor system for the conversion of acetic acid to ethylene.

[2495] The gas product stream discharging from the condenser can include CO (primary product in absence of H ₂ 0 feed to the reactor). C ₂ ¾, ( ^' Ή,. unreacted C0 ₂ , and unreacted H ₂ . In embodiments, the gas product stream may be sent, for instance, to an ODH s stem that converts ethane to ethylene. For example, the gas product stream may be introduced into tire ODH s stem (that converts ethane to ethylene) downstream of an acetic acid scrubber in the ODH system. In other embodiments, the gas product stream may be sent to a steam cracker system that converts ethane to ethylene. For example, the gas product stream may be introduced into the steam cracker system downstream of a quench tower in the steam cracker system. In yet other embodiments, the components in the gas product stream may be separated and sent to other systems or sold.

[2496] At block 10610, the method optionally includes recycling a portion of the gas product stream to the reactor. The recycling of a portion (or all) of the gas product stream may reduce C0 ₂ (including unreacted C0 ₂ ) discharged from the reactor system. Moreover, the recycle 10520 of the gas components may increase conversion of C0 ₂ by the reactor system 10500. Optionally, all of the gas product stream may be recy cled, such as in the case with little or no interest to recover CO, ethane, and ethylene from the gas product stream. Such implementations may have a focus or interest in production of acetic acid. Yet, if CO production has value, other implementations may recycle some of the gas product stream and send the remaining portion of the gas product stream for recovery of CO. I-I ₂ , C0 ₂ , etc. In those implementations, the recovered C0 ₂ and ¾ may be utilized in the feed to the reactor in certain instances. Lastly, in certain embodiments where CO, ethane and ethylene content in the reactor effluent are low such that the CO, ethane and ethylene cannot be readily economically recovered then the entire gas product stream may be recycled to the reactor, and free C0 ₂ and H ₂ added for the feed composition to the C0 ₂ conversion reactor.

[2497] Applications of the embodiments of the present reactor system that converts C0 ₂ into acetic add or CO may advance greener ethane-to-eihylene processes by consuming C0 ₂ emissions from those processes. Furthermore, implementations may increase flexibility of steam cracking systems or conventional ODH systems that convert ethane to ethylene to adjust ethylene production versus acetic acid production or CO production. Acetic acid can optionally be converted to other value products, such as ethanol or ethylene, based on market need. CO can optionally be used for syngas generation or utility generation (for example, steam generation). Syngas can then be converted to hydrocarbon-based fuels or methanol.

[2498] An embodiment is a method of processing carbon dioxide with catalyst, including contacting carbon dioxide with the catal st in the presence of hydrogen in a reactor to convert carbon dioxide to acetic acid and carbon monoxide. The method includes discharging a product effluent from the reactor to a condenser heat exchanger. The product effluent includes acetic acid, carbon monoxide, and water. The product effluent may further include ethane, ethylene, unreacted carbon dioxide, or umeacted hydrogen, or any combinations thereof. The method includes condensing the acetic acid and the water in the condenser heat exchanger. The method may include providing a feed having carbon dioxide and hydrogen to the reactor. The method may include heating the feed (e.g., in a heat exchanger) upstream of the reactor. In implementa tions, the feed includes water, and the selectivity of conversion of carbon dioxide in the reactor favors acetic acid over carbon monoxide, ethane, and ethylene. In other implementation, the feed does not include water, and the selectivity of conversion of carbon dioxide in the reactor favors carbon monoxide over acetic acid, ethane, and ethylene.

[2499] Examples

[2500] These examples are given only as an examples and not meant to limit the present techniques. The examples were performed with the reactor system 10700 depicted in FIG. 107. The reactor 10702 in the reactor system 10700 is a large-scale laboratory reactor that is a continuous tubular fixed-bed reactor having ODH catalyst. The reactor 10702 performed as a €(¼ conversion reactor in the Examples.

[2501] FIG. 107 depicts the reactor system 10700 having the reactor 10702 utilized to perform the syntheses over ODH catalyst in the Examples (Examples Adi, Ad2, Ad3, and Ad4) that converted €(½ to acetic acid and CO. The reactor 10702 is a tubular fixed-bed reactor having two tubes disposed in series and with each tube having a fixed bed of catalyst. The two tubes are constructed of Type 316L stainless steel. The two tubes can each be characterized as two respective tubular fixed-bed reactors but the two tubes were operated collectively in the Examples as a single tubular fixed-bed reactor.

[2502] The two tubes each have a heat transfer jacket that receives circulating oil from a closed-loop oil bath for heating or cooling of the two tubes to maintain a desired temperature in the reactor. In the examples, the reactor temperature was maintained at the temperatures noted in Table Adi and Table Ad2 below'. The temperature of the two tubes was monitored with thermocouples as temperature sensors.

[2503] The ODH catalyst in the two fixed beds is Mo 1.0 Vo.37Teo.23Mx>. i40 _d =4.9 _? with the numerical subscripts indicating molar ratios of the moly bdenum, vanadium, tellurium, niobium, and oxygen, as determined by inductively coupled plasma mass spectrometry (ICP-MS). While ICP-MS may not measure oxygen, the oxygen may be determined by the calculation of one hundred percent minus the percentages of Mo, Ve, Te, and Nb. This ODH catalyst is the first low-temperature catalyst described above.

[2504] The inside diameter of each tube is 2.1 centimeters (cm) giving a catalyst-bed diameter of 2.1 cm. The catalyst-bed height in the first tube is 128.5 cm. The cataly st bed height in the second tube is 135 cm. The two catalyst beds are “diluted” with a diluent powder. In other words, the cataly st is diluted with diluent powder during the catalyst shaping (e.g., extrusion, pelletizing, spheretizing, etc.) to give the diluted catalyst bed. The mass ratio of diluent powder to catalyst is 1.22. The diluent powder is Versa!™ Alumina V-250 manufactured by Honeywell UOP having headquarters in Des Plaines, Illinois, United States of America. The mass of cataly st in the two beds combined is 171 grams. The mass of the diluent in the two beds combined is 209 grains. In preparation for addition to tiie reactor beds, the catalyst and diluent powder (alumina) are mixed together as powders and then extruded together into a cylindrical shape. The cylindrical extrudate shape hits a diameter of about 1.7 millimeter (nun) and a length in the range of 2 mm to 10 mm. [2505] A combined gas feed 10704 of CO2 and H ₂ was fed to the inlet of the reactor i0702 from respective gas cylinders of CO2 and ¾. The gas cylinders of C0 ₂ and H ₂ were obtained from Praxair, Inc. (headquarters in Danbury, Connecticut, U SA.) The C0 ₂ content (purity') was greater than 99.9 volume percent. The ¾ content (purity) was greater than 99.999 volume percent. The available pressure of the gas cylinders provided motive force for flow of the combined gas feed 10704 into the reactor 10702. A respective flow valve associated with each gas cylinder gave the desired flow' rate of each gas component. Water 10706 as feed was introduced into the gas feed 10704 flowing to the reactor 10702 in Examples Ad2, Ad3, and Ad4, as noted in Table Ad2 below. The water 10706 was fed as liquid water that evaporated into steam at an inlet portion of the reactor 10702 having an inert bed and upon initial contact with the catalyst bed. Motive force for conveying the water 10706 into the gas feed 10704 and to the reactor 10702 was by a pump. Table Adi below gives the feed composition in volume percent (vol. %) to the reactor 10702.

[2506] The inlet pressure (psig) at the reactor 10702 inlet is given below in Table Adi and Table Ad2. This reactor 10702 inlet pressure is due to the hydraulic backpressure generated by flow' of the feed gas through the reactor 10702 beds, downstream condenser 10710, and associated piping. The gas hourly space velocity (GHSV) and the weight hourly space velocity (WHSV) are given in hour ¹ (Sr ^! ) in Table Adi. The GHSV is the ratio of the volumetric flow' rate of the gas feed 10704 plus vapor feed generated from evaporation of liquid feed 10706 at the inlet of reactor at standard conditions for temperature (25 °C) and pressure (100 kilopascais) to the combined volume of the tw'o fixed-beds of catalyst. The WHSV is the ratio of the weight (mass) flow rate of the total feed (gas feed 10704 plus water 10706) to the combined weight (mass) of the two fixed-beds of catalyst. The diluent powder is not counted (included) in the basis for the volume or weight of the catalyst bed. The weight and volume of the ODH active phase catalyst (mixed metal oxide) is the basis in the GHSV and WHSV calculations.

[2507] The equation for GHSV in the examples to give the GHSV values presented in Table Adi below' is GHSV = Q/V. Q is the volumetric flow rate (also knowm as volume flow' rate, rate of fluid flow, or volume velocity') of the substance or fluid that passes through the catalyst bed in units of volume per time, i.e., is the volume of the substance or fluid that passes through the catalyst bed per unit time. V is the volume of the catalyst bed given in the same units as the volume unit in the volumetric flow' rate. In the examples for the GHSV calculation for determining the GHSV values in Table Adi, V is equal to the volume of the active phase of the catalyst bed. In alternate implementations (not given in Table Adi), the GHSV can be calculated and expressed with V equal to the volume of the catalyst active phase plus the volume of the catalyst diluent in the catal st bed.

[2508] The equation for WHSV in the examples to give the WHSV values presented in Table Adi below is WHSV = m/pi, where m is the mass flow rate of the substance or fluid that passes through the catalyst bed and is in units of mass per time, i.e., is the mass of substance or fluid that passes through the catalyst bed per unit time. The variable m is the trass of the catalyst bed given in the same unite as the mass unit in Site mass flow' rate. In the examples for the WHSV calculation for determining the WHSV values in Table Adi, m is equal to the mass of the active phase of the catalyst bed. In alternate implementations (not given in Table Adi), the WHSV can be calculated and expressed with m equal to the mass of the catalyst active phase plus die trass of die catalyst diluent in die catalyst bed. [2509] The product effluent 10708 discharged from the reactor 10702 to a condenser 10710, which condensed acetic acid and water in effluent 10708. The cooling medium in the condenser 10710 was distilled water. The condenser 10710 is a shell-and-tuhe heat exchanger with product gas on the tube side and the distilled water on the shell side. A product gas stream 10712 discharged from the condenser 10710 to a vent system 10714. A sample syringe was utilized to collect a gas sample 10718 of the product gas stream at a sample point downstream of the condenser 10710. A liquid product stream 10720 discharged from the condenser 10710 to a liquid collection system 10722. A liquid sample 10724 of the liquid product stream 10720 was obtained.

[2510] The C0 ₂ conversion in carbon atom percent (C-atom %) and the selectivity (C-atom %) as normalized given in Table Ad2 are based on analysis of die gas sample 10718 and the liquid sample 10722. Acetic acid was the dominant product when H2O-CO2-H2 feed mixture was utilized. CO was a dominant product when CO2-H2 feed mixture was utilized. When the H2O-CO2-H2 feed mixture was utilized (in all operating conditions), the CO, C2H0, and C2H4 were ranked as the 2 ^nd , 3 ^rd , and 4 ^th dominant products, respectively. Additionally, when die H2O-CO2-H2 feed mixture was utilized, an increase in reactor temperature or reactor inlet pressure gave an increase in acetic -acid selectivity and yield. In those cases, the yield is the C0 ₂ conversion multiplied by the acetic-acid selectivity.

[2511] Table Adi . Reactor operating conditions and feed composition

[2512] Table Ad2. Catalyst Activity and Product Distribution

[2513] Lastly, Aspen Plus ^® software (version 8.6) was utilized to calculate the heat of reaction of Example Ad2. In the simulation, an Aspen Plus ^® RYield reactor block was utilized to simulate the net heat of reaction as 161 kilojoules (KJ) per mole of C0 ₂ converted. This implies that the reaction in Example Ad2 is endothermic. Therefore, a practice may be to provide a preheated feed to this reactor or heat the reactor body to accommodate the net reaction. Aspen Pius ^® software is available from Aspen Technology, Inc. headquartered in Bedfo d, Massachusetts, USA. It should be noted that variations of the balance of reactions in the reactor, including on scale- up to industrial scale, may provide for the net of the reactions in the reactor to be exothermic. Therefore, a practice in some instances may be to cool the reactor, such as by flowing a cooling medium through a heat-transfer jacket of the reactor. Lastly, the reactor in particular implementations might switch between endothermic operation and exothermic operation, depending on the feed composition and operating conditions of the reactor.

[2514J FIG. 108 is integrated systems 10800 including the reactor system 10500 (FIG. 105) that converts CO2 to acetic acid or CO. The integrated systems 10802 also includes a system 10802 that converts ethane to ethylene. The s stem 10802 may be, for example a conventional ODH reactor system or a steam cracker furnace system. The system 10802 includes a component 10804 that performs the conversion of ethane to ethylene. The component 10802 may be an ODH reactor that receives ethane and contacts the ethane on ODH catalyst in the presence of oxygen to convert the ethane to ethylene (e.g., via an ODH reaction). Such an ODH reactor may generate byproducts such as acetic acid, CO, and CO2 but ethylene is the primary' product. The component 10802 may be a steam cracking furnace that receives ethane as feed and converts die ethane to ethylene via steam and high temperature.

[2515] The sy stem 10800 generates CO2 emissions 10806 in the conversion of ethane to ethylene. The CO2 emissions 10806 may be collected as waste or sent to the environment. In the illustrated embodiment, the system 10800 diverts some of the CO2 emissions 10806 as feed 10506 to the €<¾ conversion reactor system 10500. Therefore, the system 10800 reduces CO2 emissions 10806 by diverting a portion of the CO2 emissions to the CO2 conversion reactor system 10500. The source of the CO2 may be, for example, from flue gas of a steam cracker furnace, or from an amine tower in the steam cracker furnace system or in the conventional ODH system.

[2516] As discussed, the C0 ₂ conversion reactor system 10500 may add H ₂ and optionally H ₂ 0 to the feed 10506, as discussed above. The reactor system 10500 may convert the CO2 to acetic acid or CO and other secondary products. In the conversion of the CO>; the reactor system 10500 may provide liquid components 10516 that include water and acetic acid. Acetic acid can be a primary' product If ¾0 is added to the feed 10506. The reactor system 10500 may give gas components 10518 that can include CO (product), C ₂ ¾ (secondary product) C ₂ H ₄ (secondare product), CO2 (unreacted feed 10506), and ¾ (unreacted ¾ added to feed 10506). The CO can be a primary' product if H2O is not added to the feed 10506.

[2517] An embodiment is a me thod of processing carbon dioxide in a reactor system, including contacting carbon dioxide with catalyst in presence of hydrogen in a reactor to convert carbon dioxide to acetic acid, caibon monoxide, ethane, and ethylene. In certain implementations, the reactor is a fixed-bed reactor having tire catalyst (e.g., ODH catalyst) in a fixed bed. The method may include maintaining an operating temperature of the reactor at less titan 425 °C or less than 400 °C. The method includes discharging an effluent from tire reactor to a condenser (a heat exchanger). The effluent includes acetic acid, carbon monoxide, ethane, ethylene water, carbon dioxide, and hydrogen. The method includes condensing the acetic acid and the water in the condenser (e.g., a heat exchanger, quench tower, etc.). The method may include feeding water to the reactor, wherein selectivity of conversion of carbon dioxide in the reactor favors acetic acid over carbon monoxide ethane, and ethylene. On the other hand, the method may include not feeding water to the reactor, wherein selectivity of conversion of carbon dioxide in the reactor favors carbon monoxide over acetic acid, ethane, and ethylene.

[2518] Another embodiment is a system to convert carbon dioxide into products. The system includes a reactor having an ODH catalyst to convert carbon dioxide in the presence of hydrogen into at least acetic acid and carbon monoxide, and discharge a product effluent including at least acetic acid, carbon monoxide, water, and unreacted carbon dioxide. The system includes a condenser (e.g., shell-and-tube heat exchanger) to receive the product effluent and condense acetic acid and water, and discharge a liquid product stream including at least acetic acid and water and a gas product stream including at least carbon monoxide and umeacted carbon dioxide. The system may include a recycle conduit to convey at least a portion of tire gas product stream to the reactor. The system may include a preheater (e.g.. shell-and-tube heat exchanger) to heat feed to the reactor, wherein the feed includes at least carbon dioxide and hydrogen.

[2519] A number of implementations have been described. Nevertheless, it will be understood that various modifications may be made without departing from die spirit and scope of the disclosure.

[2520] A method of converting one or more alkanes to one or more alkenes that includes providing a first stream containing one or more alkanes and oxygen to an oxidative dehydrogenation reactor; converting at least a po rtion of the one or more alkanes to one or more alkenes in the oxidative dehydrogenation reactor to provide a second stream exiting the oxidative dehydrogenation reactor containing one or more alkanes, one or more alkenes, and one or more of oxygen, carbon monoxide and acetylene; and providing the second stream to a second reactor containing a catalyst that includes CuO and ZnO and reacting the second stream to provide a third stream exiting the second reactor containing one or more alkanes, one or more alkenes, and lower or undetectable levels of oxy gen and acetylene compared to the second stream.

[2521] In some embodiments disclosed herein, the degree to which carbon monoxide is produced during the ODH process can be mitigated by converting it to carbon dioxide, which can then act as an oxidizing agent. The process can be manipulated so as to control the output of carbon dioxide from the process to a desired level. Using the methods described herein a user may choose to operate in carbon dioxide neutral conditions such that surplus carbon dioxide need not be flared or released into the atmosphere.

[2522] In some embodiments disclosed herein, the degree to which acetylene is produced during the ODH process can be mitigated by converting it to o ther compounds.

[2523] In some embodiments disclosed herein, the degree to which oxygen is retained in post ODH process streams can be mitigated by converting it to other compounds.

[2524] Disclosed herein are methods for mitigating carbon monoxide and/or acetylene formation in an ODH process and minimizing the amount, if any, of oxygen in post ODH process streams. Aspects of the methods include introducing, into at least one ODH reactor, a gas mixture of a lower alkane, oxygen and optionally carbon dioxide, under conditions that allow production of the corresponding alkene and smaller amounts of various by-products. For multiple ODH reactors each reactor contains the same or different ODH catalyst. In some embodiments a steam containing optional diluents may also be introduced into the reactor as part of the gas mixture. [2525] In some embodiments the lower alkane is ethane, and the corresponding alkene is ethylene.

[2526] In some embodiments, at least one ODH reactor is a fixed bed reactor. In some embodiments at least one ODH reactor is a fixed bed reactor that includes heat dissipative particles within the fixed bed. In some embodiments the heat dissipative particles have a thermal conductivity that is greater than the catalyst. In alternative embodiments, at least one ODH reactor is a fluidized bed reactor.

[2527] in some embodiments, at least one ODH catalyst is a mixed metal oxide catalyst. In particular embodiments, at least one ODH catalyst is a mixed metal oxide of the formula: Mo ₀ ViTe _c Nbr _f Pd , wherein a, b, c, d, e and f are the relative atomic amounts of the elements Mo, V, Te, Mb, Pd and O, respectively; and when a = I, b = 0.01 to 1.0, c = 0.01 to 1.0, d = 0.01 to 1.0. 0.00 < e < 0.10 and f is a number to at least satisfy the valence state of the metals in the catalyst.

[2528] In some embodiments, at least one ODH catalyst is a mixed metal oxide of the formula:

In1q6 25-7 25 VsOei where d is a number to at least satisfy the valence state of the metals.

[2529] Various embodiments relate to oxidative dehydrogenation (ODH) of lower alkanes into corresponding alkenes. Lower alkanes are saturated hydrocarbons with from 2 to 4 carbons, and the corresponding alkene includes hydrocarbons with the same number of carbons, but with one carbon to carbon double bond. While any of the lower alkanes can be converted to their corresponding alkenes using the methods disclosed herein, one particular embodiment is the ODH of ethane, producing its corresponding alkene, ethylene.

[2530] Carbon Monoxide Output

[2531] Carbon monoxide can be produced in the ODH reaction as a by-product of oxidation of the one or more alkanes. The carbon monoxide output is a function of the amount of carbon monoxide produced in the oxidative process.

[2532] Measuring the amount of carbon monoxide coming off the ODH reactor can be done using any means known in the art. For example, one or more detectors such as GC, !R, or Rahman detectors, are situated immediately downstream of the reactor to measure the carbon monoxide output. While not required the output of other components may also be measured. These include but are not limited to the amounts of ethylene, unreacted ethane acetylene, carbon dioxide and oxy gen and by-products such as acetic acid.

[2533] Carbon monoxide output can be stated using any metric commonly used in the art. For example, the carbon monoxide output can be described in terms of mass flow' rate (g/rnin) or volumetric flow' rate (cmVmin). In some embodiments, normalized selectivity can be used to assess the degree to which carbon monoxide is produced or consumed. In that instance the net mass flow rate of CO — the difference between the mass flow rate of CO entering and leaving the ODH reactor — is normalized to the conversion of ethane, in essence describing what fraction of ethane is converted into carbon monoxide as opposed to ethylene, or other by-products such as acetic- acid.

[2534] Many industrial processes, in addition to ODH, produce carbon monoxide which must be captured or flared where it contributes to the emission of greenhouse gases. Using the carbon monoxide mitigation steps disclosed herein converts most, if not all, carbon monoxide resulting from the ODH process to carbon dioxide. An advantage then is the ability to reduce or eliminate the amount of carbon monoxide produced in the ODH process in combination with other processes, such as thermal cracking. In some instances, the carbon dioxide can be captured in the amine wash tower.

[2535] Acetylene Output

[2536] Acetylene can be produced in the ODH reaction as a by-product of oxidation of the one or more alkanes. The acetylene output is a function of the amount of acetylene produced in the oxida tive process.

[2537J Measuring the amount of acetylene coming off the ODH reactor can be done using any means known in the art. For example, one or more detectors such as GC, IR, or Rahman detectors, me situated immediately downstream of the reactor to measure the acetylene output. While not required, the output of other components may also be measured. These include but are not limited to the amounts of ethylene, unreacted ethane, carbon monoxide carbon dioxide and oxygen, and by-products such as acetic acid.

[2538] Acetylene output can be stated using any metric commonly used in tire art. For example, the acetylene output can be described in terms of mass flow rate (g/rnin), volumetric flow rate (cnfVmin) or volumetric parts per million (ppniv). In some embodiments, normalized selectivity can be used to assess the degree to which acetylene is produced or consumed. In that instance the net mass flow rate of acetylene — the difference between the mass flo w rate of acetylene entering and leaving the ODH reactor — is normalized to the conversion of ethane, in essence describing what fraction of ethane is converted into acetylene as opposed to ethylene, or other by-products such as acetic acid.

[2539] Using the acetylene mitigatio steps disclosed herein reacts most if not all, acetylene resulting from the ODH process. An advantage then is the ability to reduce or eliminate the amount of acetylene produced in the ODH process in combination with other processes, such as thermal cracking and eliminate downstream unit operations in an ODH -type process.

[2540] Removal of Carbon Monoxide, Acetylene and Oxygen

[2541] Carbon monoxide, oxygen and acetylene are contaminants that can affect the performance of equipment downstream of the one or more ODH reactors and/or have a negative impact on the purity of the final ethylene product. A reactor placed downstream of the one or more ODH reactors containing a catalyst material that includes CuO and ZnO removes all or part of the carbon monoxide, oxygen and acetylene in the process stream passing through. In some embodiments, the material that includes CuO and ZnO can act as an adsorbent for carbon monoxide, oxygen and acetylene. In other embodiments, the material that includes CuO and ZnO can perform as a selective carbon monoxide oxidation catalyst.

[2542] In some embodiments, after a bed of material that includes CuO and ZnO is depleted of chemosorbed oxygen the material can initiate a chemical reaction whereby oxygen and acetylene are removed or eliminated, without removing carbon monoxide from the process stream. Not being limited by any single theory', it is believed that in this embodiment, CuO and ZnO me reduced to their corresponding elemental metal forms via the reaction. [2543] When the above described reactor containing a catalyst material that includes CuO and ZnO is placed downstream of the one or more ODH reactors, the mode of operation can be beneficial in certain integration options of ODH with different plants where carbon monoxide is a preferred feedstock for downstream plants as compared to carbon dioxide.

[2544] Carbon Dioxide Output

[2545] Carbon dioxide can be produced in the ODH reac tion as a by -product of oxida tion of the alkanes and recycled from the oxidation of carbon monoxide. Carbon dioxide can also be added into the ODH reactor when used as an inert diluent. Conversely, carbon dioxide may be consumed when it acts as an oxidant for the dehydrogenation reaction. The carbon dioxide output is therefore a function of the amount of carbon dioxide added and produced minus that consumed in the oxidative process. In some embodiments, the disclosed methods control the degree to which carbon dioxide acts as an oxidizing agent so as to impact the overall carbon dioxide output coming out of the ODH process.

[2546] Measuring the amount of carbon dioxide coming out of the ODH process can be done using any means known in the art. For example, one or more detectors such as GC, IR, or Rahman detectors, are situated immediately downstream of the reactor to measure the carbon dioxide output. While not required, the output of other components may also be measured. These include but are not limited to the amounts of ethylene, unreacted ethane, carbon monoxide and oxygen and by-products such as acetic acid. In addition, it should be noted that depending on the chosen metric for carbon dioxide output the output levels of the other components, for example ethane, may actually be required.

[2547] Carbon dioxide output can be stated using any metric commonly used in the art. For example, the carbon dioxide output can be described in terms of mass flow rate (g/niin) or volumetric flow rate (cm ³ /min). In some embodiments normalized selectivity can be used to assess the degree to which carbon dioxide is produced or consumed. In that instance, the net mass flow rate of CO _{? —} the difference between the mass flow' rate of CO? entering and leaving the ODH reactor — is normalized to the conversion of ethane, in essence describing what fraction of ethane is converted Into carbon dioxide as opposed to ethylene, or other by-products such as acetic add. A carbon selectivity of 0 indicates that the amount of carbon dioxide entering the reactor is the same as the casfeon dioxide output. In other words, the process is carbon dioxide neutral. A positive carbon dioxide selectivity _' alerts a user that carbon dioxide is being produced, and that any oxidation of carbon dioxide that is occurring is insufficient to offset that production, resulting in the process being carbon dioxide positive which may result in a lower selectivity for the olefin.

[2548] In some embodiments, product selectivity for carbon dioxide is less titan about 10 wt %, in some cases less titan about 7.5 wt. % and in other cases less titan about 5 wt. %. The product selectivity for carbon dioxide can be any of the values or range between any of the values recited above.

[2549] In some embodiments, the total amount of carbon dioxide in the stream exiting the one or more ODH reactors can be essentially the same as the total amount of carbon dioxide in the stream entering the one or more ODH reactors. In tins instance, essentially the same means that the difference between the amount of carbon dioxide in tiie stream exiting die ODH reactors is within 2 volume percent (+ 2 Vol. %) of the amount of carbon dioxide entering the ODH reactors. In some embodiments, the amount of carbon dioxide in the stream exiting the ODH reactors can be about +5 vol. %. in some cases about +7.5 vol. % and in other cases about +10 vol. % and can be about -5 vol. %, in some cases about -7.5 vol. % and in other cases about -10 vol. % of the amount of carbon dioxide in the stream entering the ODH reactors. The difference between the amount of carbon dioxide in the stream exiting the ODH reactors and the amount of carbon dioxide entering the ODH reactors can be any value or range between any of the values recited above.

[2550] In some embodiments, the methods and apparatus disclosed herein provide the possibility of a carbon dioxide negative process. In this instance, carbon dioxide is consumed at a higher rate than it is produced and shows a negative carbon selectivity. The ODH process may produce carbon dioxide, but the degree to which carbon dioxide is consumed while acting as an oxidizing agent offsets any production that is occurring. Many industrial processes, in addition to ODH, produce carbon dioxide which must be captured or flared where it contributes to the emission of greenhouse gases. When using a carbon dioxide negative process the excess carbon dioxide from other processes may be captured and used as the diluent in the ODH process under conditions where there is negative carbon selectivity. An advantage then is the ability to reduce the amount of carbon dioxide produced in the ODH process in combination with other processes, such as thermal cracking. In addition, consumption of carbon dioxide is endothermic and by increasing the degree to which carbon dioxide acts as an oxidizing agent, heat produced from ODH of ethane is partially offset by consumption of carbon dioxide, reducing the degree to which heat must be removed from the reactor. In some embodiments, when acting as an oxidizing agent, carbon dioxide can produce carbon monoxide, which can be captured and used as an intermediate in production of other chemical products, such as methanol or formic acid.

[2551] Acetic Acid Removal

[2552] The stream exiting the one or more ODH reactors can be directed to a quench tower or acetic acid scrubber, in some cases, prior to being fed to the second reactor, which facilitates removal of oxygenates, such as acetic acid, ethanol and water via a bottom outlet. A stream containing unconverted lower alkane (such as ethane), corresponding alkene (such as ethylene), and one or more of unreacted oxygen, carbon dioxide, carbon monoxide, acetylene and inert diluent, are allowed to exit the scrubber and are fed downstream.

[2553] In some embodiments, the stream from the one or more ODH reactors is cooled to a lower temperature prior to being fed to an acetic acid scnibber (as described below). The temperature of the stream prior to entering the acetic acid scnibber can be at least about 40 °C, in some cases at least about 45 °C, and in some cases at least about 50 °C and can be up to about 90 °C, in some cases up to about 85 °C, in some cases up to about 80 °C, in some cases up to about 75 °C and in some cases up to about 70 °C, The temperature of the ODH reactor product stream fed to an acetic acid scnibber can be cooled to any temperature value or range between any of the temperature values recited above.

[2554] The oxy genates removed via the quench tower or acetic acid scrubber can include carboxylic acids (for example, acetic acid), aldehydes (for example, acetaldehyde), alcohols (for example, ethanol) and ketones (for example, acetone). The amount of oxygenate compounds remaining in the stream exiting the scnibber will often be zero, i.e, below the detection limit for analytical test methods typically used to detect such compounds. When oxygenates can be detected they can be present at a level of up to about 1 per million by volume (ppniv), in some cases up to about 5 ppmv, in other cases less titan about 10 ppniv, in some instances up to about 50 ppmv and in other instances up to about 100 ppmv and can be present up to about 2 vol. %, in some cases up to about 1 vol. %, and in other cases up to about 1,000 ppmv. The amount of oxygenates or acetic acid in the stream exiting the scrubber can be any value, or range between any of the values recited above.

[2555] The Second Reactor

[2556] In some embodiments, the ODH reactor (or reactors) can provide a stream containing at least a srnail amount of oxygen remaining as reactor effluent. In some embodiments, the oxygen can provide a benefit to the ODH reactor product gas. In some embodiments, when the ODH catalyst is exposed to an oxygen free reducing environment at elevated temperature it may become permanently degraded. In other embodiments, if the level of oxygen in the product gas from the ODH reactor contains less than about 1 ppmv of oxygen, most, if not all, of the one or more alkanes are converted to one or more alkenes in the inlet portion of the reactor and a large portion of the reactor catal st bed is not utilized.

[2557] In some embodiments, oxygen in the ODH reactor product gas causes serious and operational issues in the downstream equipment, as a non-limiting example at the first compression stage of an ODH process. The presence of oxygen is not only a safety concern, but also an operational concern, as oxygen can result in degradation and/or fouling of downstream equipment, such as the amine system. This process consideration presents a need to remove oxygen to a very low or non-detectable level before the product gas is compressed.

[2558] One method used to reduce/eliminate oxygen in the ODH product gas focuses on catalyticaUy combusting a small portion of the ODH product gas to the complete consumption of any residual oxygen. This approach is viable, however in many cases it is undesirable, because it increases the overall oxygen consumption in the ODH process and, in the non-limiting example of the alkane being ethane, reduces overall process selectivity toward ethylene.

[2559] As described above, a reactor placed downstream of the one or more ODH reactors containing a catalyst material that includes CuO and ZnO removes all or part of the carbon monoxide oxygen and acetylene in the process stream passing through. In some embodiments, the material that includes CuO and ZnO can act as an adsorbent for carbon monoxide oxygen and acetylene. In some embodiments, the material that includes CuO and ZnO can perform as a selective carbon monoxide oxidation catalyst.

[2560] In some embodiments, the amount of oxygen in the stream leaving the one or more ODH reactors can be at least about 80 ppmv, in some cases at least about 100 ppmv, in other cases at least about 150 ppmv and in some instances at least about 200 ppmv and can be up to about 5 vol. %, in some cases up to about 4 vol. %, in other cases up to about 3 vol. %, in some instances up to about 2 vol. %, in other instances up to about 1 vol. %, and in particular situations up to about 500 ppmv. The amount of ox gen in the stream leaving the one or more ODH reactors can be any of the values or range between any of the values recited above.

[2561] In some embodiments, when there is oxygen in the stream leaving the second reactor (in some instances the amount of oxygen will be undetectable or zero ppmv), the amount of oxygen in the stream leaving the second reactor can be at least about 1 ppmv. in some cases at least about 2 ppmv, in some cases at least about 3 ppmv and in some cases at least about 5 ppmv and can be up to about 1 vol. %, in some cases up to about 0.9 vol. %, in some cases up to about 0.8 vol. %, in some cases up to about 0.7 vol. %. in some cases up to about 0.6 vol. %, and in some eases up to about 0.5 vol. %. The amount of oxygen in the stream leaving the second reactor can be any of the values or range between any of the values recited above.

[2562] In some embodiments, the amount of carbon monoxide in the stream leaving the one or more ODH reactors can be at least about 100 ppmv, in some cases at least about 200 ppmv, in some cases at least about 300 ppmv and in some cases at least about 400 ppmv and can be up to about 10 vol. %, in some cases up to about 9 vol. %, in some cases up to about 8 vol. %, in some cases up to about 7 vol. %, in some cases up to about 6 vol. %, and in some cases up to about 5 vol. %. The amount of carbon monoxide in the stream leaving the one or more ODH reactors can be any of the values or range between any of the values recited above.

[2563] In some embodiments, when there is carbon monoxide in the stream leaving the second reactor (in some instances the amount of carbon monoxide will be undetectable or zero ppmv), the amount of carbon monoxide in the stream leaving the second reactor can be at least about 1 ppmv. in some cases at least about 2 ppmv. in some cases at least about 3 ppmv and in some cases at least about 5 ppmv and can be up to about 8 vol. %, in some eases up to about 7 vol. %. in some cases up to about 6 vol. %, in some cases up to about 5 vol. %, in some cases up to about 4 vol. % and in some cases up to about 3 vol. %. The amount of carbon monoxide in the stream leaving the second reactor can be any of the values or range between any of the values recited above.

[2564] In some embodiments, when there is acetylene in the stream leaving the one or more ODH reactors (in some instances the amount of acetylene will be undetectable or zero ppmv), the amount of acetylene in the stream leaving the one or more ODH reactors can be at least about 1 ppmv. in some cases at least about 2 vppm in some cases at least about 5 ppmv and in some cases at least about 10 ppmv and can be up to about 1000 ppmv, in some cases up to about 750 ppmv, in some cases up to about 500 ppmv, in some cases up to about 400 ppmv. in some cases up to about 300 ppmv, and in some cases up to about 300 ppmv. The amount of acetylene in the stream leaving the one or more ODH reactors can be any of the values or range between any of the values recited above. [2565] In some embodiments, the amount of acetylene in the stream leaving the second reactor will be less than the amount entering the second reactor and, in many instances, the stream exiting the second reactor will be substantially free of acetylene.

[2566] In some embodiments, when there is acetylene in the stream leaving the second reactor (in many instances the amount of acetylene will be undetectable, less than 1 ppmv, or zero ppmv), the amount of acetylene in the stream leaving the second reactor can be at least about 1 ppmv, in some cases at least about 2 ppmv, in some cases at least about 3 ppmv and in some cases at least about 5 ppmv and can be up to about 100 ppmv, in some cases up to about 50 ppmv, in some cases up to about 25 ppmv, in some cases up to about 20 ppmv, in some cases up to about 15 ppmv, and in some cases up to about 10 ppmv. The amount of acetylene in the stream leaving the second reactor can be any of the values or range between any of the values reci ted above.

[2567] In some embodiments, temperature in the second reactor can be at least about 100 °C, in some cases at least about 110 °C, in other cases at least about 115 °C and in some instances at least about 120 °C and can be up to about 200 °C, in some instances up to about 190 °C, in other instances up to about 180 °C, in sortie circumstances up to about 175 °C, and in other circumstances up to about 170 °C. The temperature of second reactor can be any temperature value or range between any of the temperature values, including a temperature gradient within the second reactor, recited above.

[2568] In some embodiments, a fixed bed reactor loaded with a catalyst material that includes CuO and ZnO, can be located in three different locations in the ODH process:

[2569] At the ODH reactor outlet, whereby the product from the reactor outlet can be cooled to below 200 °C, before it enters the second reactor.

[2570] At the outlet of the acetic acid scrubber / quench tower, whereby the gaseous feed to the second reactor can be preheated to at least 100 °C

[25711 At the outlet of the first stage compression of the gaseous ODH product, downstream of die acetic acid scrubber, whereby the feed to the second reactor can be, at least, in part preheated by the energy of the compression. [2572] ODH Complex

[2573] In some embodiments, the chemical complex (one embodiment is shown schematically in Figure 109) includes, in cooperative arrangement, an ODH reactor 10910, a quench tower or acetic acid scrubber 10920, a second reactor 10925 (as described herein), an amine wash tower 10930 (which can include a caustic tower) a drier 10940, and a distillation tower 10950. ODH reactor 10910 includes an ODH catalyst capable of catalyzing, in the presence of oxygen which may be introduced via oxygen line 10970, the oxidative dehydrogenation of alkanes introduced via alkane line 10980. Although second reactor 10925 is shown directly after quench tower or acetic acid scrubber 10920, in some instances, it will be more efficiently utilized after the gas stream is compressed, in some cases prior to amine wash tower 10930. Thus, in some cases, the process configuration can be more energy efficient if second reactor 10925 is placed after the input stream lias been compressed.

[2574] The ODH reaction may also occur in the presence of an inert diluent, such as carbou dioxide, nitrogen, or steam, that is added to ensure the mixture of oxygen and hydrocarbon are outside of flammability limits. Determination of whether a mixture is outside of the flammability limits, for the prescribed temperature and pressure, is within the knowledge of the skilled worker. An ODH reaction that occurs within ODH reactor 10910 may also produce, depending on the catalyst and the prevailing conditions within ODH reactor 10910, a variety of other products which may include carbon dioxide carbon monoxide, oxygenates and water. These products leave ODH reactor 10910, along with unreacted alkane, corresponding alkene, residual oxygen, carbon monoxide, acetylene and inert diluent, if added, via ODH reactor product line 10990.

[2575] ODH reactor product line 10990 is directed to quench tower or acetic acid scrubber 10920 which quenches the products from product line 10990 and facilitates removal of oxygenates and water via quench tower bottom outlet 109100. Unconverted low'er alkane, corresponding alkene, unreacted oxygen, carbon dioxide, carbon monoxide, acetylene and inert diluent added to quench tower 10920 exit through quench tower overhead line 109110 and me directed into second reactor 10925.

[2576] Second reactor 10925, which can be variously positioned as described above, contains a catalyst material that includes CuO and ZnO, which removes all or part of the carbon monoxide, ox gen and acetylene. In second reactor 10925, most or all of the imreacted oxygen and acetylene is consumed. The remaining unconverted lower alkane corresponding alkene, imreacted oxygen (if present) all or part of the carbon dioxide, carbon monoxide (if present), acetylene (if present) and inert diluent are conveyed to amine wash tower 10930 via line 109115.

[2577] Any carbon dioxide present in line 109115 is isolated by amine wash tower 10930 and captured via carbon dioxide bottom outlet 109120 and may be sold, or, alternatively, may be recycled back to ODH reactor 10910 as described above. Constituents introduced into amine wash tower 10930 via line 109115, other than carbon dioxide, leave amine wash tower 10930 through amine wash tower overhead line 109130 and are passed through a dryer 10940 before being directed to distillation tower 10950, where C2/C2 ₊ hydrocarbons are isolated and removed via C2/C2 ₊ hydrocarbons bottom outlet 109150. The remainder includes mainly hydrocarbons, including remaining N ₂ or CH used as diluent that is in Site vapor phase and carbon monoxide (if any), which leave distillation tower 10950 via overhead stream 109160.

[2578] In many embodiments, C2/C2 ₊ hydrocarbons bottom outlet 109150 is fed to a C ₂ splitter (not shown) that separates ethane from ethylene.

[2579] Various tools commonly used for chemical reactors, including flowmeters, compressors valves and sensors for measuring parameters such as temperature, pressure and flow _' rates can be used. It is expected that the person of ordinary skill in the art would include these components as deemed necessary for operation.

[2580] Examples

[2581] The ODH catalyst was prepared as follows. A solution of (NilOeMo-iO^di-EO (44.20 g, 35.77 mmol, white solid) in 600 ml, of distilled water was prepared In a 2-L RBF equipped with a magnetic stir bar. A solution of V0S0-!*3.46H ₂ 0 (14.07 g, 62.95 mmol, bright blue solid) in 600 ml, of distilled water was prepared in a i-L beaker equipped with a magnetic stir bar. Both solutions were stirred in a 60 °C water bath until homogeneous. The blue vanadium solution was then added to the clear colorless molybdenum solution. This resulted in a dark purple solution with a fine suspension. Sodium dodecyl sulfate (SDS) (13.57 g, 47.06 mmol, white solid) was added to the reaction mixture. The pusple shiny was left to stir at 60 °C for 1 hour.

[2582] The reaction mixture was transferred to a glass liner, with a total volume of about 1380 mL measured after rinsing. The liner was loaded into a 2-L pressure reactor (Pan· Instrument Company, Moline, IL) and the gap filled with distilled water. The reactor was sealed and the head space evacuated and backfilled with nitrogen gas 1 Ox times. The headspace was left under 15 psig nitrogen gas and sealed. The reactor was transferred to a programmable oven and heated for 24 hours at 230 °C (1-hour ramp to 230 °C, 24-hour cooling ramp back to room temperature). Once cooled to room temperature, the reactor was vented, and the contents filtered using a Buchner funnel and 4 quantitative filter papers. The oily mother liquor was decanted off and the filter papers changed. The filter cake was rinsed with 1250 mL of distilled water. The filtrate was a dark blue color and the product was a charcoal/grey purple color.

[2583] The filter cake was dried in an oven at 90 °C overnight with 15.29 g of product being recovered (37% estimated yield). The uncalcined catal st was broken up with a spatula and then loaded into a programmable muffle furnace. The program was set to ramp over one hour to 280 °C and held there for 9 hours, before cooling back to room temperature naturally. Tins air treated product was ground with mortar and pestle and submitted for CHN analysis. The carbon and nitrogen content was found to be less than 1 wt. %. The material was loaded into a quartz boat and centered in the quartz tube of the QRU furnace. The quartz tube was purged (400 seem) with bulk nitrogen for 8 hours, after which the nitrogen feed was fed through an oxygen scrubbing bed to further purify the nitrogen to less than 4 ppmv oxygen. This ultra-high purity (UHP) nitrogen was purged through the quartz tube overnight. The next morning, the furnace was turned on and heated to 400 °C over a 4-hour ramp. The catalyst was calcined at 400 °C for 2 hours and then cool to ambient temperature naturally.

[2584] 5.0 g of calcined catalyst (92 wt. %) and 0.44 g of beryllium oxide (8 wt. %) were placed into a 100 ruL beaker (92%). About 30 inL of distilled water was added to the beaker and stirred manually. This beaker was placed into an oil bath at about 100 °C and an overhead stirrer was set up so the paddle was just off Site bottom of the beaker. The mixture was stirred at 80 rpm for about 1.5 hours until it formed a paste and the beaker with die paste was placed in an oven at 90 °C to dty overnight. The solid catalyst chunk was broken up with a spatula and the final catalyst was placed in a muffle furnace at 350 °C for 3.5 hours.

[2585] The catal st composition for the second reactor in diese examples is the reduced form of an oxide precursor composition containing 70 wt. % CuO, 20 wt. % ZnO and 10 wt. % Z1O ₂ . For production of this oxide precursor composition, a Cu-Zn-Zr nitrate solution (metal content 15.2 wt. %, Cu : Zn : Zr ratio corresponding to a CuO : ZnO : Zr0 ₂ weight ratio of 7:2: 1) was precipitated with soda solution (20 wt. %) at pH 6.5 and 70 °C. After completion of precipitation, the suspensio n was stirred for a further 120 minutes at pH 6.5 and 70 °C. Next, the solution was filtered, and the filter cake washed free of nitrate with demineralized water and dried at 120 °C. The dried powder was calcined at 300 °C for 240 minutes in a forced air oven.

[2586] The catalyst/beryllium oxide was evaluated in an apparatus depicted in Figure 110. ODH reactor 200 was a lab scale dehydrogenation reactor (ODH micro reactor unit).

[2587] The feed to ODH reactor 110200 included oxygen, ethane and nitrogen at a weight ratio of I8V0I. %/36Voi. %/46Vol % respectively at 8.4 psig and a gas flow rate of 32.8 seem through ¼ inch outside diameter tubing. The ODH catalyst/beryllium oxide described above was placed in ODH reactor 110200 and the gas feed was processed through ODH reactor 110200 at 327 °C. The effluent from ODH reactor 110200 was passed through condenser 110210 where and aqueous solution containing 18 5 wt. % acetic acid was condensed and the remaining effluent gas was passed to second reactor 110220 at 14 psig.

[2588] Second reactor 110220 (piaced in temperature control oven 110230) w'as loaded with 2g of the dried, calcined powder and the effluent gas was contacted with the powder at 150 °C and 8.4 psig to 14 psig at 32.8 seem. The effluent exited second reactor 110220 at ambient pressure and was evaluated in gas chromatograph 110240. [2589] Table SI outlines the results of the test.

[2590] Table SI: Experimental results at oxygen elimination reactor inlet pressure at around between 8.4 -14 psig

[2591J The data in Table SI demonstrate that on the dried, calcined powder catalyst removes 0 ₂ , CO, and acetylene at temperatures of higher than 120 °C. It is also clear form the data shown in Table SI that all the compounds are not being chemosorbed but rather reacted either with oxygen from the catalyst or in the gas stream. The constant presence of oxygen in the feed stream was sufficient to oxidize all of the acetylene, which led to continuous removal of acetylene and O2, even after the catalyst material was depleted of chemosorbed oxygen.

[2592] A number of implementations have been described. Nevertheless, it will be understood that various modifications may be made without departing from the spirit and scope of the disclosure.

[2593] EXAMPLES

[2594] General Procedures

[2595] Reagents

[2596] Reagents purchased from manufacturers were used as received without further purification, unless noted otherwise. All reagents, with the exception of alumina, were purchased from SIGMA ALDRICH ^® . The supplied certificates of analysis for ammonium molybdate ((NH 6M07O24 · 4 H ₂ 0) and vanadium (TV) oxide sulfate hydrate (VOSO4 * 3.46 ¾0 and VOSO4 * 3.36 H ₂ 0) were used to establish the hydrate content for different batches. VOSCb · 3.46 H2O and VOSQi · 3.36 H ₂ 0 (97% purity) were confirmed by ICP-MS analysis. Vanadium and sulfur components were confirmed by KM11O4 titration. (NROeMcwCtyi * 4 H ₂ 0 (81-83% purity) was confirmed for nitrate, aresenate, phosphate, silicate and sulfate specifications by ICP-MS. Heavy metal, magnesium potassium and sodium specifications were determined by ICP-OES. Three different grades of sodium dodecyi sulfate (SDS, SIGMA-ALDRICH ^® ) surfactant were used, but the type of grade had no effect on the synthetic results. The purity of the various grades were 92.5-100.5%, >95%, and > 99.0 % purity. The purity range of the first grade was established by SIGMA-ALDRICH ^® using various methods including: titration (based on totai alkyl sulfate content), purity based via GC (total lauryl sulfate, sodium salt), total water content determined by Karl Fisher titration, and total sodium content. More specifically, the ranges listed above w'ere determined by titration based on the total alkyl sulfate content in the sample.

[2597] Three different types of alumina were investigated. VERSAL™ Alumina V-250 was manufactured from UOP, ALUMAX ^® PB250 was manufactured from PIDC and CATAPAL ^® B Alumina was manufactured from Sasol. [2598] Distilled water was prepared inhouse using a Coming Mega Pure 12A System ACS as distillation apparatus.

[2599] MRU

[2600] The ability of catalysts and catalyst materials described herein to participate in the oxidative dehydrogenation of ethane was tested in a microreactor unit (MRU).

[2601] The MRU included a reactor tube made from SS316L stainless-steel SWAGELOK ^® Tubing, which had an outer diameter of 0.5 inches, an internal diameter of about 0.4 inches, and a length of about 15 inches. A 6- point WIKA Instruments Ltd. K-type thermocouple (as shown in figures 126 and 127) having an outer diameter of 0.125 inches was inserted axially through the center of tire reactor, which was used to measure and control the temperature within the catalyst bed. A room temperature glass tight sealed condenser was located after the reactor to collect water/acidic acid condensates. The gas product flow was allowed to either vent or was directed to a gas chromatography (Agilent 6890N Gas Chromatograph, using CHROMPERFECT ^® - Analysis, Version 6.1.10 for data evaluation) via a sampling loop.

[2602] To prepare catalyst and catalyst materials for testing in the MRU, the catalyst or catalyst material was loaded into a 1-inch round die and pressed with 8 tons of compression force for 10 to 15 seconds of dwelling time. The pressed catalyst or catalyst material was then crushed into small pieces using a mortar and pestle. The crushed catalyst or catalyst material was sieved and a particle sizes between 425 pm and 1 mm were collected to be loaded for testing on the MRU.

[2603] For MRU experiments, the catalyst bed was prepared by Method Af 1 or Method Af2 Method All was employed to test cataly sts and catalyst materials. Method Af2 was employed to test cataly st materials.

[2604] Method Afl. 1.96-3.00 g of catalyst or catalyst material was physically mixed with quartz sand such that the catalyst bed had a total volume of about 3-6 ml,.

[2605] Method Af2. Under this method, the catalyst bed consisted only of catalyst material. Further, the amount of catalyst material was determined based on the amount of catalyst used to prepare the catalyst material. Specifically, the catalyst material was loaded in an amount such that the theoretical amount of (i) catalyst or (ii) catalyst and iron in the catalyst bed was 1.96-3.00 g. For example, if a catalyst material was prepared from 40 wf % catalyst and 60 wt. % alumina, then 4.91 g of catalyst material would be used to prepare the catalyst bed. Likewise, if a catalyst material was prepared from 30 wt. % catalyst, 10 wt. % goethite, and 60 wt. % alumina then 4.91 g of catalyst material would be used to prepare the catalyst bed. The catalyst or catalyst material can be diluted with filler up to 8 mL.

[2606] For both Methods Afl and Af2, the catalyst bed was loaded in the middle zone of the reactor — located between points #2 and #5 of the thermocouple — and the remaining volume of the reactor was packed with quartz sand (see Figures 125 and 126). The load was then secured with glass wool on the top and Site bottom of reactor. [2607] For the MRU testing, a pre-mixed feed gas was fed through the reactor. The pre-mixed feed gas entering the reactor was 36 niol.% ethane, 18 niol.% oxygen, and 46 mol.% nitrogen. Further, the pre-mixed feed gas flow was adjusted by a calibrated mass flow controller to obtain a gas hourly space velocity (GHSV) of about 3,000 hr ¹ , based on the catalyst or catalyst material volume in the catalyst bed as defined by Method Afl or Method Af2.

[2608] The flow rate of the feed gas was about 70 standard cubic centimeters per minute (seem) to about 80 seem (e.g., about 76.1 seem). The catal st bed placed in the reactor tube can include the catalyst or catalyst material and a filler. With reference to the MRU’s catalyst bed, a filler refers to a material that does not participate in the oxidative dehydrogenation of ethane or have other catalytic activity, such as non-selective oxidation under the MRU test conditions. The filler employed was quartz sand. The catal st or catalyst material can be diluted with filler up to 8 rnL. ί2609| The 35% conversion temperature is determined at a weight hourly space velocity (WHSV) of 2.90 h ^! , with the WHSV based on the amount of catal st or die amount of catal st used to prepare the catal st material, and a gas hourly space velocity (GHSV) of about 3,000 h ¹ . Whereby WHS V is defined as mass flow of feed gas to die reactor divided by the weight of the catalyst in the catalyst bed, GHSV is defined as volumetric flow of the reactor feed gas divided by the volume of the catalyst bed.

[2610] Typically, the inlet pressure was in the range of about 1 pound per square inch gauge (psig) to about 2.5 psig and the outlet pressure is in the range of about 0 psig to about 0.5 psig.

[2611] The gas exiting the reactor was analyzed by gas chromatography (Agilent 6890N Gas Chromatograph, using CHROMPERPECT ^® --- Analysis, Version 6.1.10 for data evaluation) to determine the percent of various hydrocarbons (e.g., ethane and ethylene) and, optionally, other gases such as <½, CO _? , and CO and acetylene.

[2612] A catalyst or catalyst material’s 35% conversion temperature was determined as follows. Conversion of the feed gas was calculated as a mass flo w rate change of ethane in the product compared to feed ethane mass flow rate using the following formula:

[2613] In the above equation, C is the percent of feed gas that has been converted from ethane to another product (i.e., ethane conversion) and X is the molar concentration of the corresponding compound in the gaseous effluent exiting the reactor at corresponding temperature. The ethane conversion was then plotted as a function of temperature to acquire a linear algebraic equation. The linear equation for ethane conversion was solved to determine the temperature in which the ethane conversion was 35% (i.e., the 35% conversion temperature).

[2614] Further, the gas exiting the reactor was analyzed by gas chromatography to determine catalyst or catalyst material selectivity to ethylene (i.e., the percentage on a molar basis of ethane that forms ethylene). Selectivity to ethylene was determined using the following equation:

In the above equation, S _Ethyiene is the selectivity to ethylene and X is the molar concentration of the corresponding compound in the gaseous effluent exiting the reactor at corresponding temperature. The selectivity to ethylene w'as determined at the 35% conversion temperature, unless otherwise indicated. As such, after the 35% conversion temperature was determined, the above equation for selectivity was solved using the corresponding values for X _Eihyiene , Xco2, and Xco at the 35% conversion temperature.

[2615] When reported, acetic acid production was determined by running MRU testing long enough to collect an aqueous condensate in the condenser (e.g., 1-3 days). After collecting a sample of the condensate, the sample was submitted for liquid GC analysis (Agilent 6890N Gas Chromatograph, using CHROMPERFECT ^® - Analysis, Version 6.1.10 for data evaluation). To perform the liquid GC analysis, 300-450 mg of liquid sample was transferred to a scintillation vial. Next, 25 mg of isopropanol (IP A) was added as an interned standard. Further, 18-20 inL of distilled ¾0 was added to dilute the sample. Prepared samples were then transferred to GC vials and set in sequence to tested using an auto sampler. The GC analysis was a split injection method with a temperature program and flame ionization detector (FID). Further, a set of 3 calibration standards were run in duplicate for the relative response factor used for calculating acetic acid content in sample.

[2616] ICP-MS

[2617] Inductively Coupled Plasma Mass Spectrometry (ICP-MS), sensitive enough to detect elements in ppb concentration ranges, was one of the analytical techniques used for measuring the elemental composition of catalyst or catalyst materials. ICP-MS analysis was performed on an Agilent 7700X ICP-MS system. Quantitative determination of atoms’ concentration in the samples were determined with the use of an external standard calibration. The calibration curves were constructed after subtracting the reagent blank. Concentrations were given in ng/mg (wt-ppm) or pg/g (wt-ppm) units in this anal sis.

[2618] General ICP-MS Procedure for catalysts containing molybdenum and vanadium.

[2619] Samples were prepared by placing 10 milligram (mg) of catalyst or catalyst material in either 3 milliliters (mL) of 10-15 wt. % NaOCl solution or 3 mL of a 625-35.0 molar (M) NaOH solution. The solution was then heated in an oil bath at 90 °C with rigorous mixing.

[2620] TCP Digestion Method Af 1 :

[2621 ] NaOCl solutions produced a clear, yellow' solution within 5-10 minutes upon heating at 90 °C with vigorous mixing. After mixing the NaOCl solution at 90 °C for an additional 12 hours, a small amount of fine yellow precipitate formed. The precipitate was allowed to settle to the bottom of the mixing vial, and the mother liquor was used for the ICP-MS sample work up and injection. The diluted solution was then further diluted 10-100x using 5% nitric acid and analyzed by ICP-MS.

[2622] ICR Digestion Method Af2 :

[2623] NaOH solutions produced homogenous samples for ICP-MS experiments after 12-24 hours of heating and mixing and resulted in a clear, brown solution. After dissolution was complete, the resulting solution was diluted to a final volume of 60-80 mL with distilled water.

[2624] The diluted solution was then further diluted i0-I00x using 5% nitric acid and analyzed by ICP-MS. [2625] The multi-element scan optimized the instrument parameters to scan for trace (ppb) levels of 50+ elements. The 50+ elements included in the multi-element scan included Fe and Al. When scanning for Mo, V, AL and Fe, higher concentrations of calibration standards were used and the instrument sensitivity was reduced as the elements of interest were found in percent levels. This was done by preparing calibration standards for each of the four elements. These calibration standards were prepared in percentage levels in high concentrations. Normally, calibration standards were prepared in ppm level concentrations. The ICP-MS program was developed such that the four elements Mo, V, Al, and Fe were detected with a high degree of accuracy. The elements that were normally calibrated to ppm level concentrations were excluded as the detector was calibrated for only the four elements at high percentages.

[2626] General ICP-MS procedure for catalyst materials containing molybdenum and vanadium. i2627| ICP Digestion Method AO :

[26281 Catalyst materials including (i) molybdenum, vanadium, oxygen, and iron, and (ii) molybdenum, vanadium, oxygen, aluminum and iron were digested in an oxalic acid or an NaOCl solution to produce a suitable homogenous sample for ICP-MS.

[2629J Approximately 10 mg of catalyst material and 2-3 g of oxalic acid were added to a vial. Then, 2-3 rnL of distilled water was added to create a suspension. The suspension was heated in an oil bath at 90 °C with rigorous mixing. Dissolution of the catalyst generally took 24-72 hours to produce a homogenous, blue solution. After dissolution was complete, the resulting solution was diluted to a final volume of 60-80 ml,. The diluted solution was then further diluted 10-100x using 5% nitric acid and analyzed by ICP-MS.

[2630] CHN Analysis

[26311 Carbon, hydrogen, and nitrogen (CHN) analysis was performed on select samples using a LECO ^® CHN628 Series Determinator, using a combustion technique. A pre-weighed and encapsulated sample was placed in the instrument’s loader where the sample was transferred to the instrument’s purge chamber directly above the furnace, eliminating atmospheric gases from the transfer process. The sample was then introduced to the primary combustion furnace, winch contained only pure oxygen. This results in a rapid and complete oxidation. Carbon, hydrogen, and nitrogen present in the sample are oxidized to form CO2, H2O, and NO _x gases, respectively, and are swept by the oxygen earner through a secondary furnace for further oxidation and particulate removal.

[2632] The combination gases are then collected in a vessel known as a ballast for equilibration. The homogenized gases from the ballast are swept through a 10 enri aliquot loop and, using an ultra-high purity? helium earner gas, on to the detectors. Separate, optimized non-dispersive infrared (ND1R) ceils are utilized for the detection of C0 ₂ and H ₂ 0. The NO _* gases are passed through a reduction tube filled with copper to reduce the NO _* gases to N ₂ and remove any excess oxygen present from the combustion process. The aliquot gas then passes through scrubbers to remove C0 ₂ and H ₂ 0 and onto a thermal conductivity cell (TC) utilized to detect the N ₂ . Prior to every set of samples, the instrument is calibrated using certified standards.

[2633] XRD

[2634] Instrument

[2635] Powder X-Ray Diffraetometry (PXRD) data was collected using a PANalyiical Aeris X-ray diffractometer by SEMx Incorporated. This diffractometer instrument consisted of three basic elements: X-ray tube, sample holder, and X-ray detector. X-ra s were generated in a cathode ray tube (Cu source with Ka radiation = 1.5418 A) with sire resulting X-rays being directed onto the sample. As tire sample and detector are rotated, the intensity of the reflected X-rays is recorded to produce characteristic X-ray spectra. When the incident X-rays reflecting off the sample satisfies the Bragg Equation (hl=2ά sin Q), constructive interference occurs and a peak in intensity occurs (y~axis). X-ray diffractometers were setup such that the sample rotates in the path of the X-ray beams at an angle Q, while the X-ray detector is mounted on an arm to collect the diffracted X-rays and rotates at an angle of 2Q from~5° to 70° (x-axis).

[2636] Qualitative XRD analysis and Rietveld Refinement was performed using HighScore Plus XRD analysis software. The samples were finely ground to reduce particle size and to obtain a uniform mixture. They were then loaded onto the XRD sample holder and die XRD spectrum was acquired. The Rietveld Refinement results were combined with Highscore Plus and EDS results to perform qualitative and quantitative anal sis.

[2637] Amorphous content determination

[2638] The weight percentage of amorphous content was determined by external standard. With an external standard phase, the instrument intensity constant. K-factor, is determined. Corundum was used as the external standard and was measured with the same instrument configuration shortly after the unknown sample was measured. The K-factor approach is described by O’Connor and Raven: 1988, Powder Diffraction, 3 (1), 2-6. For each sample, the weight percentage of the crystalline MoVO _x orthorhombic phase had to be determined in order to assign weight percentages to the amorphous content. The Degree of Crystallinity (DOC) Method, based on the estimation that the total intensity of area contributed to the overall diffraction pattern by each component in the analysis, was used to determine the amount of amorphous phase. The degree of cr stallinity was calculated from the total areas under the defined crystalline and amorphous components from:

DOC ------ Crystalline Area Crystalline Area + Amorphous Area

Where the weight fraction of the amorphous material was calculated from:

W amorphous = 1 DOC

[2639] The Ortho-Mo VO _x phase contributed to the crystalline area and therefore needed to be quantified in order to determine the amorphous area. To compensate for the fact that different materials and backgrounds would have different effects, a sample ofMoVTeNbO _* was used to calibrate some constants needed for the DOC method. Samples containing MoVO _x phases had the ortho-MoVO _x phase weight percentages qualitatively determined using only two elements (Mo and V) based on the MoVTeNbOx calibration.

[2640] Ml phase content determination

[2641] The MoVQ _x orthorhombic phase (also referred to in literature as the Ml phase) was fitted using literature crystal structure data for a different yet cry stallographically analogous compound because the orthorhombic Pba2 crystalline phase was a match. MoV8bO _x XRD simulation: W. Ueda, D. Vitry, T. Kato, N. Watanabe, Y. Endo, 2006, Res. Chem. Informed. 32(3-4), 217-233. Lattice parameters: a = 21.124 A, b = 26.598 A, c = 4.0076 A

[2642] Comparative raw data analysis

[2643] The PXRD raw' data was also analyzed using a Python code through the program Spyder. The code generated overlaid plots. It also analyzed the data by comparing the peak prominence of all the local maxima and generated a plot with peaks meeting an established threshold. Relevant catalyst peaks are highlighted in the plot with vertical lines and the range of the relative peak intensities were provided by the code. The relative peak intensities were calculated as a percentage of the 22.2° 2Q reflection and a maximum and minimum value was created based on the selection of catalysts to be processed by the p thon code [2644] SEM

[2645] Scanning electron microscope (SEM) images were collected using a JSM-IT300LV INTOU CHSCOPE™. Samples were prepared on an aluminum stud with double sided carbon tape.

[2646] SEM-EDS

[2647] Energ -dispersive X-ray spectroscopy (EDS) was conducted using a JEOL JED-2300 DRY SDD EDS detector. Samples were sent to SEMx Incorporated for EDS analysis. The samples were finely ground to reduce particle size and obtain Site uniform mixture. They were then loaded onto EDS stub for analysis by SEM. EDS was used for elemental analysis and surface examination. EDS is a micro-analytical technique that provides a semi- quantitative elemental analysis of the surface of a sample (e.g., the top 1 to 3 microns). The SEM was used to examine the surface morphology at magnifications ranging from 20 to 100,000 times. The EDS instrument detects elements with an atomic number equal or greater than sodium, but also has light element capability, which means that it can also detect carbon, nitrogen, oxygen, and fluorine. The estimated lower detectable limit for any give element generally is between about 0.2 and 0.5 wt. %

[2648] PSD b SEM

[2649] Samples were sent to SEMx Incorporated for particle size analysis using scanning electron microscopy (SEM), model JEOL - JSM300 LV. SEM was used to observe and count the particles in the sample to obtain the Particle Size Distribution (PSD). For the PSD measurements, the SEM instrument took pictures at different magnifications. Measurements were done for 400-800 particles at different magnifications to cover the size range (statistical population). Size was measured by length in micrometers and was measured on the longest dimension of the particles. SEM based PSD is a method for analyzing samples where particles are agglomerated (stuck together) because the analyst can visually see this through the microscope and make the judicious decision to measure the distinct particles rather than the agglomerates. Statistics and analysis were based on total counts measured by SEM. [2650] Pore Volume, BET Surface Area Analysis, and BJH Pore Size Distribution Analysis [26511 Gas adsorption manometry was used for the determina tion of adsorption isotherms of nitrogen at the temperature of liquid nitrogen (approximately 77 K). The amount of gas adsorbed was evaluated by measuring the change in gas pressures. Isothermal nitrogen adsorption processes were measured, and surface areas and volumes were calculated.

[2652] Total pore volume was calculated by nitrogen gas uptake at the relative pressure P/Po = 0.99.

[2653] Brunauer-Emmett-Teller (BET) analysis was applied to quantify' the specific surface area (m ² /g) of the solid samples. BET valuations were performed by multilayer adsorption of nitrogen and measured as a function of relative pressure. Applying BET analysis allows for a quantitative comparison of solids’ surface areas by determining the monolayer capacity from nitrogen multila er adsorption experiment. Monolayer capacity is a representation of total specific surface area and encompasses both the external area and pore area of porous solid. [2654] The Barrett-Joyner-Halenda (BJH) method was used for calculating pore size (A) distributions from experimentally collected adsorption isotherms using the Kelvin model of pore filling (cniVg· A). This technique ciiaracteriz.es pore size distribution independently of external area due to particle size of the sample and can be applied to mesopore and small macropores.

[2655] Nitrogen physisorption experiments were performed on a TiiStar (M1CROMERITICS ^® Instruments) by the University' of Calgary. Samples were analyzed by nitrogen adsorption at 77 K. The as-received samples were loaded into physisorption cells. The samples were degassed at 200 °C for 1 hour prior to the adsorption experiments. [2656] Yield calculations

[2657] Theoretical Yield calculations were based on the weight of each reagent used. The weight of each reagent used in grams was divided by the molecular weight of the compound in grams per mol. For example:

Weight (NH ₄ ) ₆ Mo ₇ 0 ₂₄ · 4H ₂ 0 (g) / Molecular weight (NH ₄ )6Mo ₇ 0 ₂₄ · 4H ₂ 0 (g/mol) - (NH ₄ )6Mo ₇ 0 ₂₄ · 4H ₂ 0 mol T s calculation was performed for Site vanadium as well.

[2658] The theoretical moles of the final product was predicted by assuming that both tire molybdenum and the vanadium have attained the highest oxidation states in the final product. Thus, molybdenum and vanadium formed Mo0 ₃ and V ₂ 0 ₅ respectively.

[2659] The moles of the starting material was used and multiplied by the respective molar equivalents of each of the total oxidized species. The moles were then multiplied by the predicted theoretical weight of the fully oxidized final product in order to get the final theoretical weight of the catalyst. For example:

[2660] Theoretical weight of molybdenum in the fully oxidized state:

(N¾)6MO ₇ 0 ₂₄ · 4H ₂ 0 mol x 7 = 7(Mo0 ₃ mol)

((MoO _Ui ol)) x MO0 ₃ g/mol) = Theoretical weight of fully oxidized molybdenum in grams Theoretical weight of vanadium in the fully oxidized state:

(V0S0 ₄ · 3 46H ₂ 0 mol x ½ ₌ ½ V ₂ 0 ₅ mol)

(V2O5 mol x V2O5 g/mol) = Theoretical weight of fully oxidized vanadium in grams

Total theoretical weight in g = (Theoretical weight of fully oxidized molybdenum in grams) + (Theoretical weight of fully oxidized vanadium in grams)

Percent yield = (Actual measured yield (g) / Theoretical yield (g)) x 100

The percent yield w'as determined by dividing the actual measured yield by the theoretical yield and multiplying by 100

[2661] Synthesis of Catalysts including Mo, V, and O

[2662] Catalysts containing molybdenum and vanadium were synthesized on a small scale (i.e., in a 600 iriL reactor) and on a large scale (i.e., in a 2,000 nxL reactor) from ammonium molybdate ((NH ₄ ) ₆ Mo ₇ 0 _{24 *} 4 H ₂ 0) (SIGMA- ALDRICH ^® ) and vanadium (IV) oxide sulfate hydrate (VOSO4 · 3.x H ₂ 0) (SIGMA- ALDRICH ^® ).

[2663] General small-scale procedure. (NH ₄ ) _& Mo ₇ G ₂₄ · 4 IT ₂ 0 (SIGMA- ALDRICH ^® ) (13.26 g) was weighed into a round bottomed flask. The white solid 'as dissolved in 180 rnL (2.98 ^c 10 ² inol/L) of distilled water with aid of a 60 °C warm water bath and stirring to provide a clear, colorless solution. VOS0 ₄ · 3.x H ₂ 0 (4.22 g) was weighed into a glass beaker. The bine solid was dissolved in 180 rnL (5.20 x 10 ^‘2 mol/L) of distilled water with aid of a 60 °C warm water bath and stirring to provide a clear, blue solution. The blue VOSO4 (aq.) solution was added either all-at-once or dropwise to the colorless (NH, bMo,O _? ^ (aq.) solution to produce a dark purple opaque solution. The resulting solution was stirred for 30 minutes at 60 °C. The pH of the solution was measured with a pH probe to be 2.70 before addition of surfactant.

[2664] After 0-30 minutes, 4.05 g of sodium dodecyl sulfate surfactant (SDS) (SIGMA-ALDR1CH ®) was added as a white powder to the dark molybdenum and vanadium solution to produce an emulsified slurry. The purple-emulsified slurry was stirred for 30 minutes at 60 °C. The pH of the SDS containing slurry was measured to be 2.70.

[2665| After the 30 minutes, the purple slurry was added to a glass or Teflon liner and the liner was added to a stainless-steel high-pressure PARR reactor. The reactor was sealed and purged ten times by alternating with 15 psig (pounds per square inch gauge) nitrogen gas and vacuum, then left under 15 psig nitrogen and sealed. The reactor was then placed into an oven or heating jacket and heated to 220-230 °C. The temperature was estimated through a thermowell filled with Dow' Coming 510 silicon oil in the cases were a heating jacket was employed, or it was the temperature the oven was set to. The reaction was allowed to proceed at 220-230 °C either being stirred at 150 rpin or at a stand-still for 20-48 hours before being allowed to cool to room temperature.

[2666] Once cooled, the reactor was vented, opened, and the resulting spongy-slum ^r was filtered. The solids were washed with water (0.5-3.0 L) until the filtrate ran clear. The solids were optionally washed with one or any combination of ethanol ethyl acetate, and aqueous oxalic acid. Alternatively, the solids were optionally air treated. The washing step was able to remove the oil-like products produced by the decomposition of SDS during the hydrothermal reaction. The effectiveness of the washing step was, optionally, monitored by CHN elemental analysis — and was typically performed until the wt. % carbon was less than 1 wt. %. The washing step was not required when an air calcination was employ ed — which is discussed below. Further, the washed solids were pusple or grey in color. Subsequently, the solids were dried at 90 °C overnight and the next morning the dr' solids were weighed producing a yield of 35-55%. The spongy' solids were optionally ground into powder prior to calcination. [2667] For air calcination the following general procedure was employed. The pre-ealcined catalyst was loaded into a ceramic bowl and placed into a muffle furnace. The muffle furnace was programmed for the following calcination program: ramp to 280 °C over 30 minutes, dwell for 30 minutes at 280 °C, ramp to 400 °C over 30 minutes, dwell for 6 hours at 400 °C, and then cool to ambient temperature naturally. Typically, the calcined catalyst produced a weight yield of 93-96%.

[2668] For nitrogen calcination, the following general procedure w'as employed. The pre-calcined catalyst (catalyst obtained directly after filtration and drying at 90 °C from the hydrothermal reaction) was optionally treated in convective flow air at 280 °C for 3-26 hours to removal any organic residue before calcination. A quartz boat containing the precalcined catalyst was loaded into a quartz tube, which was placed into a split tube furnace. The quartz tube was purged with bulk nitrogen for 8 hours — with the nitrogen feed being passed through an oxygen scrubbing bed to further purify the nitrogen to less than 4 ppm oxygen prior to purging. Subsequently, the pre- calcined catalyst was calcined at 400 °C. The heating program for the furnace was: 4 hours ramp from room temperature to 400 °C dwell at 400 °C for 2 hours and then cool to ambient temperature naturally. Calcined catalyst weight produced a weight yield of 93-96 %.

[2669] Synthesis of Catalysts Af 1.1 -Af 1.6

[2670] Example Af 3 : Synthesis of Catalyst Af 1.3 and Af 1.2

[26 1] (NH ₄ ) ₆ MO _T 0 _{24 *} 4 ¾0 (44.20 g) was added to a 2 L round bottom flask with a magnetic stir bar.

Subsequently, approximately 600 mL of distilled water was added, and the (NH ₄ )bMq7q2 _{4 *} 4 ¾0 was dissolved with the aid of a 60 °C water bath. Next, VOS0 ₄ · 3.46 H ₂ 0 (SIGMA-ALDRICH ^® ) (14.07 g) was placed into a 1 L beaker with a magnetic stir bar. Subsequently, approximately 600 mL of distilled water was added and the VOS0 ₄ · 3.46 H ₂ 0 was dissolved with the aid of a 60 °C water bath. Then, the vanadium solution was added to the round bottom flask containing the molybdenum solution while stirring at 60 °C. Sodium dodecyl sulfate (SDS, SIGMA- ALDRICH ^® , 13.57 g) was added to the round botom flask while stirring at 60 ^C C, and the mixture was stirred for approximately 30 minutes at 60 °C.

[2672] The round bottom flask was then removed from the water bath and allowed to cool before transferring the solution to a 2 L glass liner. The round bottom flask was rinsed with distilled water and the rinse was transferred to the glass liner. The glass liner containing the purple solution w'as inserted into a 2 L PARR reactor. The PARR reactor unit was sealed and subsequently, the PARR reactor headspace w'as pumped and purged ten times with nitrogen. The headspace was left under approximately 15 psig nitrogen and the top valve on the PARR reactor was closed. Subsequently, the PARR reactor was stirred at approximately 150 rpm using an overhead stirrer at a surface temperature from 230 °C to 247 °C for approximately 26 hours. The reaction mixture was then allowed to cool to room temperature (approximately 21 °C) and stirred at approximately 150 rpm for about 3.5 days.

[2673] The reaction mixture was filtered using a Buchner funnel four qualitative filter papers, and an aspirator. The collected solid precalcined catalyst was rinsed with about 1 L of distilled water. Precaution was taken to prevent the filter cake from cracking. The pre-ealeined catalyst w'as then washed with about 1 L of ethanol . The precalcined catalyst material was then dried overnight in an oven at approximately 90 °C. The weight of the precalcined catalyst was 22.15 g.

[2674] Subsequently, a portion of the precalcined catalyst was ground manually using a mortar and pestle.

The precalcined catalyst (5.94 g) w'as placed in a 100 mL beaker and calcined in a muffle furnace under the following conditions: ramp to 280 °C over 30 minutes, hold for 30 minutes at 280 °C, ramp to 400 °C over 30 minutes, hold for 6 hours at 400 °C, off. The catalyst was calcined and cooled to room temperature, yielding about 5.59 g of Cataly si Af 1.1.

[2675] Separately, the remaining 16.2 g of precalcined catalyst was transferred to a 400 rnL beaker. The precalcined catalyst was placed in a muffle furnace and calcined using the same program: ramp to 280 °C over 30 minutes, hold for 30 minutes at 280 °C, ramp to 400 °C over 30 minutes, hold for 6 hours at 400 °C, off. The calcined catalyst (Catalyst Afl.2) was greenish grey with a brownish layer on top.

[2676] For the purposes of catalyst performance testing and characterization, catalysts Af 1.1 and Af 1.2 were considered to be the same. Catalysts Afi.I and Afl.2 were generated using the same calcination procedure except for weight of catalyst calcined in the last step. [2677] Powder XRD data was collected on Catalyst Af 1.1 as per the general XRD procedure discussed above. The plot for Catalyst All.1 is presented in Figures 111 and 112. The raw data PXRD plot was used in establishing the range of the peak intensities presented in Table Af3 and Table AflO.

[2678] Catalyst All .1 was submitted for MRU testing. The results of the MRU testing are presented in Table Af7.

[2679] Catalyst All.2 was analyzed by EDS as per the general procedure SEM-EDS. Results are presented in Table Af2 and Table Af9.

[2680] Catalyst Afl.2 was submitted for SEM imaging. The results are presented in Figure 117.

[268 ί I Bulk density for Catalyst ATI .2 measurements are presented in Table Af6.

[2682] Example Af2 : Sy nthesis of Precaleined Cataly st Af 1.3

[2683] ; \! i i;.\io (> _{! *} 4 H O (SIGMA-ALDRiCH ^® ) (44.30 g) was weighed in a 2 L round bottom flask. The white solid was dissolved in 600 niL of distilled water with the aid of a 60 °C warm water bath and stirring to create a clear, colorless solution. VOSO4 * 3.46 ¾0 (SIGMA-ALDRICH ^® ) (13.99 g) was weighed into a 1 L beaker. The blue solid was dissolved in 600 ml, of distilled water with the aid of a 60 °C warm water bath and stirring to create a clear, blue solution. The blue vanadium solution was added drop-wise via a dropper funnel to the colorless molybdenum solution to produce a black solution. The resulting black solution was stirred for 30 minutes at 60 °C. Immediately after the completion of the addition of vanadium solution, 13.57 g of sodium dodecyl sulfate (SIGMA- ALDRICH ^® ) surfactant was added as a powder to the black molybdenum-vanadium solution to produce a purple slurry with some emulsion present. The purple-emulsified slurry was stirred for 30 minutes at 60 °C.

[2684] After 30 minutes the purple slurry' was added to a glass liner and the glass liner was added to a 2 L PARR high pressure reactor. The PARR reactor had a magnetic stir bar added inside the vessel. The PARR reactor was sealed and purged with 15 psig nitrogen gas and vacuum sequence ten times then left under 15 psig nitrogen. The P ARR reactor was setup in a heating mantle and heated via thermocouples and a temperature control box to 230 °C internal process temperature (temperature controller was set to 242 °C; the internal temperature was about 230 °C after running overnight). The reaction proceeded at 230 °C for 24 hours, with the slum' stirred via magnetic stirring, before being turned off and allowed to cool to room temperature

[2685] Once cooled, the reactor was vented, opened, and the resulting slurry was filtered using a Buchner funnel and 4 layers of qualitative filter paper. The mother liquor was blue. The solids were w'ashed wdth water (1.8 L) until the filtrate ran clear. The washed solids were dark purple in color with some brown oil present. The solids were washed with denatured ethanol (0.5 L) producing a yellow' filtrate. The solids were then additionally washed with water (0.5 L). The resulting grey solids were dried at 90 °C overnight to produce precaleined Catalyst A1Ί.3. Precalcined Catalyst Afl.3 was manually ground and weighed to be 17.82 g.

[2686] Example Af3 : Synthesis of Catalyst Af 1.4

[2687] A 5.29 g portion of precaleined Catalyst Afl.3 was loaded into beaker air calcination via a muffle furnace. The muffled furnace calcination proceeded using the following heating program: heat ramp 10 minutes to 280 °C, dwell for 30 minutes, 20 minutes ramp to 400 °C, dwell 6 hours then cooled. After the air calcination, the solid product was brown and yielded 5.06 g of Catalyst Afl.4. [2688] Catalyst Af 1 4 was then submitted for ICP-MS analysis using the general 1CP-MS procedure described herein using digestion method Af2. The results are presented in Table Af 1.

[2689] Catalyst All.4 was analyzed by EDS as per the general procedure SEM-EDS. Results are presented in Table Af2.

[2690] Powder XRD data was collected as per General Procedure XRD above. The plot for this Catalyst All.4 is presented in Figure 111. The raw data PXRD plot was used in establishing the range of the peak intensities presented in Table Af3.

[26911 Catalyst Afl.4 was analyzed by nitrogen gas adsorption anal sis for pore volume. BET surface area, and BJH pore size distributions anal sis. Pore volume and BET surface area analysis results are presented in Table Af5. BJH analysis results me presented in Figure 121.

[2692] Catalyst Afl.4 was submitted for MRU testing. The results of the MRU testing ate presented in Table Af7 and Table AGO.

[2693] Example Af4: Synthesis of Catal st All.5

[2694] (NH ₄ ) ₆ Mq7q24 · 4 H ₂ 0 (SIGMA-ALDRICH ^® ) (13.26 g) was weighed in a 500 mL round bottomed flask. The white solid was dissolved in 120 mL of distilled water with the aid of a 60 °C warm water bath and stirring to create a clear, colorless solution. VOSQt · 3.46 ¾0 (SIGMA-ALDRICH ^® ) (4.22 g) was weighed into a 500 mL round bottom flask. The blue solid was dissolved in 180 mL of distilled water with the aid of a 60 °C warm water bath and stirring to create a clear, blue solution. The blue vanadium solution was added dropwise to the colorless molybdenum solution to produce a black solution. To the black solution, sodium dodecyl sulfate (SIGMA- ALDRICH ^® ) surfactant was added as a powder to produce a purple shirty with some emulsion present. The purple- emulsified slurry _' wns stirred for 30 minutes at 60 °C.

[2695] After 30 minutes, the purple slurry? wns added to a glass liner and the glass liner was added to a 600 mL PARR high pressure reactor. The PARR reactor wns sealed and purged with 15 psig nitrogen gas and vacuum sequence ten times, then left under 15 psig nitrogen. The PARR reactor was setup with an external heating mantle, thermocouples, and a heat control box for conducting high temperature hydrothermal reaction. The hydrothermal reaction proceeded at 221-224 °C for 25 hours before being turned off and allowed to cool to room temperature. [2696] Once cooled, the reactor was vented, opened, and the resulting slurry was filtered. The filtering was done using a Buchner funnel and 4 layers of qualitative filter paper. The mother liquor was blue. The solids were washed with 0.6 L of water, followed by 0.25 L of denatured ethanol. The solids were dried at 90 °C overnight to yield 6.23 g of precalcined catalyst.

[2697] A 3.19 g portion of the precalcined catalyst was loaded into a beaker for air calcination via a muffle furnace. The muffled furnace calcination proceeded using the following heating program: heat ramp 30 minutes to 280 °C, dwell for 30 minutes, 30 minutes ramp to 400 °C, dwell 6 hours then cooled. After the air calcination, the solid product was brown and yielded 3.06 g of green Catalyst Afl.5.

[2698] Catalyst Afl.5 w ^; as then submitted for ICP-MS analysis using the general ICP-MS procedure described herein using digestion method AO. The results are presented in Table ML [2699] Catalyst Af 1 5 was analyzed by EDS as per the general procedure SEM-EDS. Results are presented in

Table Af2.

[2700] Powder XRD data was collected as per the general XRD procedure discussed herein. The plot for this catalyst is available in Figure 111. The raw data PXRD plot was used in establishing the range of the peak intensities presented in Table Af3.

[2701] Catalyst All .5 was analyzed by nitrogen gas adsorption analysis for pore volume, BET surface area, and BIH pore size distributions analysis. Pore volume and BET surface area analysis results are presented in Table Af5. BJH analysis results me presented in Figure 121.

[2702] Catalyst Afl.5 was then submitted for MRU testing. The results are presented in Table Af7.

[2703] Example Af5: Synthesis of Catal st Afl.6

[2704] ; \ϋ i;L1o C> · 4 H ₂ 0 (SIGMA-ALDRICH ^® ) (44.20 g) was added to a 2 L round bottom flask with a magnetic stir bar. Subsequently, approximately 600 mL of distilled water was added and the (NEEΐbMotOzi · 4 ¾0 was dissolved with the aid of a 60 °C water bath. Next, VOSO4 · 3.46 H ₂ 0 (SIGMA-ALDRICH ^® ) (14.07 g) was placed into a 1 L beaker with a magnetic stir bar. Subsequently , approximately 600 ml of distilled water was added and the VOSO4 * 3.46 IT ₂ 0 was dissolved with the aid of a 60 °C water bath. Then, the vanadium solution was added to the round bottom flask containing the molybdenum solution while stirring at 60 °C. Sodium dodecyl sulfate (SIGMA-ALDRICH ^® ) (13.57 g) was added to the round bottom flask while stirring at 60 °C. and the mixture was stirred for approximately 30 minutes at 60 °C.

[2705] The round bottom flask was then removed from the water bath and allowed to cool before transferring the solution to a 2 L glass liner. The round botom flask was rinsed with distilled water and the rinse was transferred to the glass liner. The glass liner containing the purple solution was inserted into a 2 L PARR reactor. The PARR reactor unit was sealed and subsequently, the PARR reactor headspace was pumped and purged ten times with nitrogen. The headspace was left under approximately 15 psig nitrogen and the top valve on the PARR reactor was closed. Subsequently, the PARR reactor was stirred at approximately 150 rpm using an overhead stirrer at a temperature from 230 °C to 232 °C for approximately 24 hours. The reaction mixture was then allowed to cool to approximately 28 °C and stirred at approximately 150 rpm for overnight.

[2706] The reaction mixture was filtered using a Buchner funnel, 4 qualitative filter papers, and an aspirator. The collected solid precalcined catal st was rinsed with about 1 L of distilled water. Precautions were taken to prevent the filter cake from cracking. The pre -calcined catal st was then washed with about 1 L of ethanol. The precalcined catalyst material was then dried overnight in an oven at approximately 90 °C. The weight precalcined catalyst was 20.07 g.

[2707] Subsequently, the 20.07 g of precalcined catalyst was ground manually using a mortar and pestle. The whole sample of precalcined catalyst was placed in a 400 mL beaker and calcined in a muffle furnace under the following conditions: ramp to 280 °C over 30 minutes, hold for 30 minutes at 280 °C, ramp to 400 °C over 30 minutes, hold for 6 hours at 400 °C, off. The catalyst was calcined and cooled to room temperature, ielding about 19.02 g of Catalyst Afl.6. [2708] Catalyst Af 1.6 was analyzed by EDS as per the general procedure SEM-EDS. Results are presented in

Table Af2.

[2709] Catalyst Afl.6 was submitted for SEM imaging. The results are presented in Figure 116.

[2710] Powder XRD data was collected as per General Procedure XRD above. The plot for this catalyst is available in Figure 111. The raw data PXRD plot was used in establishing the range of the peak intensities presented in Table Af3.

[2711] PSD analysis results are presented in Table Af4.

[2712] Bulk density measurements for Catalyst Afl.6 are presented in Table Af6.

[2713] Catalyst Afl.6 was submitted for MRU testing. The results are presented in Table Af7.

[2714] ICP-MS Analysis of Catalysts Afi.4 and Afl.5.

[2715] The ICP-MS results for Catalysts Afl.2, Afl.4 and Afl.5 are presented in Table ML [2716] Table Ml

[2717] EDS Analysis of Catalysts M ^' l .2 and Ml 4-Afl .6

[2718] The results for the EDS analysis of Catalyst M1.2 and M1.4-M1.6 are presented in Table M2. [2719] Table M2

[2720] PXRD Analysis of Catalyst Ml .4

[27211 Reflection angles and corresponding maximum and minimum relative peak intensity (relative to 22.2°

2Q) for Catalysts Afl.l, and Afl 4-Afl .6 are presented in Table AO. Figure ill presents the plot correlating to peak numbers for Catalysts Afl.l, and MΊ.4-A1Ί.6. [2722] Tabe Af3

[2723] PSD Analysis of Catalysts Af 1 1 and Af 1 6

[2724] Statistical data from the PSD of Catalysts Afl.l/Afl.2, and Afi.6 is presented in Table Af4. [2725] Table Af4 [2726] BET Analysis of Catalysts Afl .4 and Af 1.5

[2727] The results of nitrogen gas adsorption analysis for pore volume and BET surface area analysis for

Catalyst Afl.4 and Catalyst Afl.5 are presented in Table Af5.

[2728] Table AT5

[2729] Bulk density measurements of Catalysts Afl .2 and Afl .6

[2730] Bulk density measurements of Catal sts Afl .2 and Afl .6 are presented in Table Af6.

[2731] Table Af6

[2732] Activity and Selectivity' for Catalysts Afl .1 and Afl .4- Afl .6.

[2733] The MRU results for Catalysts Afl.l and Afl.4-Afl.6 are presented in Table A17.

[2734] Table A17

[2735] Synthesis of Catalyst Materials AG2.1-Af2.4 [2736] Example Af6: Synthesis of Catalyst Material AT2.1 [2737] Catalyst Afl.2 (2.1 g) was placed in a 100 mL beaker CATAPAL ^® B (Sasol) (4.9 g) and distilled water (approx. 25 mL) were then added to the beaker. The reaction mixture was then stirred using an electronic overhead stirrer equipped with Teflon blade agitator. The beaker was then placed in a 100 °C oil bath and the reaction mixture was stirred at approximately 80 rpm for 1 hour. The beaker was then placed in an oven (approx. 90 °C) and allowed to dry for about 2 days. Subsequently, the material -was taken out of the oven and broken up with a spatula and placed in a muffle furnace. The material was then calcined to yield Catalyst Material 2, la (6.32 g) using the program: ramp to 350 °C over 3 hour, hold for 2.5 hours at 350 °C, off

[2738] Catalyst Afl.2 (3.0 g) was placed in a 100 mL beaker. CATAPAL' ^® B (Sasol) (7.0 g) and distilled water (approx. 35 mL) were then added to the beaker. The reaction mixture was then stirred using an electronic overhead stirrer equipped with Teflon blade agitator. The beaker was then placed in a 100 °C oil bath and the reaction mixture was stirred at approximately 80 rpm for 1.5 hours during wliich it took on a paste-like consistency. The beaker was then placed in an oven (approx. 90 °C) and allowed to dry overnight. Subsequently, the material was taken out of the oven mid broken up with a spatula and placed in a muffle furnace at 350 °C for 3 hours yielding Catalyst Material AG. lb (9.08 g).

[2739] Catalyst Materials AG.la and AG. lb wore combined to yield Catalyst Material AG.l (15.40 g).

[2740] Catalyst Material AG.l was analyzed by EDS as per the general procedure SEM-EDS. Results are presented in Table Af9.

[27411 Catalyst Material AG.1 was submitted for SEM imaging. The results are presented in Figure 117. [2742] Powder XRD data was collected as per General Procedure XRD above. The plot for this catalyst is available in Figure 112. The raw data PXRD plot was used in establishing the range of the peak intensities presented in Table AflO.

[2743] PSD analysis results are presented in Table Afl 1.

[2744] Catalyst Material AG.1 was submitted for MRU testing. The results are presented in Table Afl 2 and Table Afl 3.

[2745] Example AG: Synthesis of Catalyst Material AG.2

[2746] Catalyst Afl.2 (2 1 g) was placed in a 100 ml, beaker. CATAPAL ^® B (Sasol) (4.2 g), goethite (SIGMA-ALDRICH ^® ) (0.7 g), and distilled water (approx. 25 mL) were then added to the beaker. The reaction mixture was then stirred using an electric overhead stirrer equipped with Teflon blade agitator. The beaker was then placed in a 100 °C oil bath and the reaction mixture was stirred at approximately 80 rpm for 1 hour during which it took on a paste-like consistency. The beaker was then placed in an oven (approx. 90 °C) and allowed to dry for about 2 days. Subsequently, the material was taken out of the oven and broken up with a spatula and placed in a muffle furnace. The material was then calcined to yield Catalyst Material AG.2a (6.32 g) using the program: ramp to 350 °C over 1 hour, hold for 2.5 hours at 350 ^C C, off.

[2747] Catalyst Afl.2 (3.0 g) was placed in a 100 mL beaker. CATAPAL ^® B (Sasol) (6.0 g), goethite (SIGMA-ALDRICH ^® ) (1.0 g), and distilled water (approx. 35 mL) were then added to the beaker. The reaction mixture was then stirred using an electric overhead stirrer equipped with Teflon blade agitator. The beaker was then placed in a 100 °C oil bath and the reaction mixture was stirred at approximately 80 rpm for 1.5 hours during which it took on a paste-like consistency. The beaker was then placed in an oven (approx 90 °C) and allowed to dry overnight. Subsequently, the material was taken out of the oven and broken up with a spatula and placed in a muffle furnace at 350 °C for 3 hours. To yield Catalyst Material Af2.2b.

[2748] Catalyst Materials Af2.2a and Af2.2b were combined to yield Catalyst Material Af2.2 (9.08 g).

[2749] Catalyst Material Af2.2 was analyzed using the general ICP-MS procedure described herein using ICP digestion method 3. The ICP-MS results for Catalyst material A12.2 are presented in Table Af8.

[2750] Catalyst Material AI2.2 was analyzed by EDS as per the general procedure SEM-EDS. Results are presented in Table AG9.

[275 ί I Catalyst Material A12.2 was submitted for SEM imaging. The results are presented in Figure 117. [2752] Powder XRD data was collected as per General Procedure XRD above. The plot for this catalyst is available in Figure 112. The raw data PXRD plot was used in establishing the range of the peak intensities presented m Table AflO.

[2753] PSD anal sis results are presented in Table Afl 1.

[2754] Catalyst Material AG2.2 was submitted for MRU testing. The results are presented in Table Afl2 and Table ATI 3.

[2755] Example 7: Synthesis of Catalyst Material 2.3

[2756] Catalyst Afl .2 (2.1 g) was placed in a 100 ml beaker. CATAPAL ^® B (Sasol) (2.8 g), goethite (2.1 g) and distilled water (approx. 25 mL) were then added to the beaker. The reaction mixture was then stirred manually. The beaker was then placed in a 100 °C oil bath and the reaction mixture was stirred at approximately 80 rpni for 1 hour during which it took on a paste-like consistency. The beaker was then placed in an oven (approx. 90 °C) and allowed to dry overnight. Subsequently, the material was taken out of the oven and broken up with a spatula and placed in a muffle furnace. The material was then calcined to yield Catalyst Material Af2.3 (6.52 g) using the program: ramp to 350 °C over i hour hold for 2.5 hours at 350 °C, off.

[2757] Catalyst Material Af2.3 was analyzed by ICP-MS using the general ICP-MS procedure described herein using ICP digestion method Af3. The ICP-MS results for Catalyst Material Af23 are presented in Table Af8. [2758] Catalyst Material Af2.3 was analyzed by EDS as per the general procedure SEM-EDS. Results are presented in Table Af9.

[2759] Catalyst Material Af2.3 was submitted for SEM imaging. The results are presented in Figure 117. [2760] Pow'der XRD data was collected as per General Procedure XRD above. The plot for this catal st is available in Figure 112. The raw data PXRD plot was used in establishing the range of the peak intensities presented in Table AflO.

[2761] PSD analysis results are presented in Table Afl 1.

[2762] Catal st Material Af2.3 was then submitted for MRU testing. The results are presented in Table Afl2 and Table Afl3.

[2763] Example Af8: Synthesis of Catalyst Material AG2.4

[2764] Catal st Afl .2 (2.1 g) was placed in a 100 rnL beaker. Goethite (SIGMA-ALDRICH ^® ) (4.9 g), and distilled writer (approx. 25 mL) were then added to the beaker. The reaction mixture was then stirred manually. The beaker was then placed in a 100 °C oil bath and the reaction mixture was stirred at approximately 80 rpm for 0.75 hour during which it took on a paste-like consistency. The beaker was then placed in an oven (approx. 90 °C) and allowed to dry overnight. Subsequently, the material was taken out of the oven and broken up with a spatula and placed in a muffle furnace. The material was then calcined to yield Catalyst Material Af2.4 (6.52 g) using the program: ramp to 350 °C over 1 hour, hold for 2.5 hours at 350 °C, off.

[2765] Catalyst Material A12.4 was analyzed by ICP-MS using the general ICP-MS procedure described herein using ICP digestion method Af3. The ICP-MS results for Catalyst material Af2.4 are presented in Table Af8. i2766| Catalyst Material AT2.4 was analyzed by EDS as per the general procedure SEM-EDS. Results are presented in Table AG9.

[2767] Catalyst Material AW.4 was submitted for SEM imaging. The results are presented in Figure 117. [2768] Powder XRD data was collected as per General Procedure XRD above. The plot for this catalyst is available in Figure 112. The raw data PXRD plot was used in establishing the range of the peak intensities presented m Table AflO.

[2769] PSD analysis results are presented in Table Af 11.

[2770] Catalyst Material Af2.4 was then submitted for MRU testing. The results are presented in Table Afl2 and Table Afl3.

[2771] ICP-MS Analysis of Catalyst Materials Af2.2-Af24

[2772] The ICP-MS results for Catalyst Materials Af2.2-Af2.4 are presented in Table Af8.

[2773] Table Af8

[2774] EDS Analysis of Catalyst ATI .2 and Catalyst Materials Af2.1-Af2.4.

[2775] The EDS analysis of Catalyst Afl .2 and Catalyst Materials Af2.1-Af2.4 are presented in Table Af9. [2776] Tab e Af9

[2777] PXRD Analysis of Catalyst Afl.l and Catalyst Materials Af2.1-Af2.4.

[2778] Reflection angles and corresponding maximum and minimum relative peak intensity (relative to 22.2° 20) for Catalyst Afl.l and Catalyst Materials Af2.1-Af2.4 are presented in Table AflO. See Figure 112 for plot correlating to peak numbers.

[2779] Table AflO

[2780J PSD Analysis of Catalyst Afl.2, and Catalyst Materials Af2.i-Af2.4

[2781 J The PSD analysis of Catalyst Afl.2, and Catalyst Materials Af2.1-Af2.4 are presented in Table Afli.

[2782] Table Mil

[2783] Activity' and Selectivity' for Catalyst Afl.1 and Catalyst Materials Af2 1-Af2.4 [2784] The MRU results for Catalyst Afl .1 and Catalyst Materials Af2.1-Af2.4 are presented in Tables Afl 2 and Afl 3. These catalysts were allowed to run on the MRU for 2-5 days. Table Afl 2 summarizes the catalyst performance based on the MRU loading method and Table Afl 3 summarizes the catalyst composition differences with catalyst performance. The ethane conversion and selectivity to ethylene have been calculated as described above at a specific temperature (348 ± 2 °C). [2785] Table Afl2

[2786] Table Ml 3

* molar formula from EDS measurement Catalyst Afl.l composition is assumed to be the same as catalyst Af 1.2. The reported molar formula is for catalyst Afi 2

J -molar formula from TCP -MS [2787] Synthesis of Catalysts Af3.1 -Af3.18 [2788] Example Af9 : Synthesis of Catalyst Af3.1

[2789] (M ii).Mo O:i · 4 H ₂ 0 (SIGMA-ALDRICH ^® ) (13.26 g) was weighed in a 500 mL round bottomed flask. The white solid was dissolved in 180 mL of distilled water with the aid of a 60 °C warm water bath and stirring to create a clear, colorless solution. VOSCh · 3.46 ¾0 (SIGMA-ALDRICH ^® ) (4.22 g) was weighed into a 250 mL beaker. The blue solid was dissolved in 180 mL of distilled water with the aid of a 60 °C warm water bath and stirring to create a clear, blue solution. The blue vanadium solution was added all-at-once to the colorless molybdenum solution to instantly produce a black solution. The resulting black solution was stirred for 30 minutes at 60 °C. The pH of the solution was measured with a pH probe to be 2.70 before addition of surfactant.

[2790] After 30 minutes, sodium dodeeyl sulfate surfactant (SIGMA-ALDRICH ®) (SDS; 4.05 g) was added as a powder to the black molybdenum-vanadium solution to produce a purple slurry with some emulsion present.

The purple-emulsified slurry was stirred for 30 minutes at 60 °C. The pH of the SDS containing slum was measured to be 2.70.

[2791] After 30 minutes the purple slurry' was added to a glass liner and the glass liner was added to a 600 mL PARR high pressure reactor. The PARR reactor was sealed and purged with 15 psig nitrogen gas a d vacuum sequence ten times then left under 15 psig nitrogen. The PARR reactor was placed into an oven and the ove was heated to 230 °C. The reaction proceeded at 230 °C for 24 hours before being turned off and allowed to cool to room temperature.

[2792] Once cooled, the reactor was vented, opened, and the resulting spongy-sluny was filtered. The filtering was done using a Buchner funnel and 4 layers of qualitative filter paper. The mother liquor w'as blue. The solids were washed with water (05 L) until the filtrate ran clear. The washed solids w'ere dark purple in color with some large dark brown aggregates and brown oil present. The brown aggregates were removed, and the remaining purple spongy' solids were washed with denatured ethanol (0.5 L) producing a yellow filtrate. The solids were dried at 90 °C overnight and the next morning the dry solids were weighed: 568 g.

[2793] The precalcined catalyst was pressed with a spatula into a quartz boat. The quartz boat containing the precalcined catalyst was loaded into a quartz tube, which was placed into a split tube furnace. The quartz tube was purged with bulk nitrogen for 8 hours at 400 seem, before the nitrogen feed was fed through an oxygen scrubbing bed to further purify the nitrogen to less than 4 ppm oxygen. This ultra-high purity (UHP) nitrogen was purged through the quartz tube overnight at 400 seem for 18 hours. Next, the furnace was turned on and heated to 400 °C to calcine the catalyst. The heating program for the furnace was: 4 hours ramp from room temperature to 400 °C, dwell at 400 °C for 2 hours and then cool to ambient temperature naturally.

[2794] The calcined catalyst, Catalyst AD.l, weighed 5.26 g (93% calcination yield). There was no yellow oil observed inside the quartz tube indicating that the SDS oil degradation product was sufficiently removed with ethanol wash for the nitrogen calcination step.

[2795] Cataly st Af3.1 w'as then submitted for MRU testing. The results are presented in Table Af20.

[2796] Example AflO: Synthesis of Catalyst Af3.2 [2797] (\ΉG;.LIoΌ·i * ⁴ H ·0 (SIGMA-ALDRICH ^® ) (13.27 g) and distilled water (120 mL) were added to a 500 mL round bottom flask. V0S0 ₄ · 3.46 H ₂ 0 (SIGMA-ALDRICH ^® ) (4.25 g) and distilled water (120 mL) were placed in a 250 mL beaker. The blue vanadium solution was then slowly poured into the clear and colorless molybdenum solution. Subsequently, SDS (SIGMA-ALDRICH ^® ) (4.09 g) was added to the reaction mixture, which was then allowed stir for 30 min at 60 °C.

[2798] Next, the contents w'ere transferred to a 600 mL glass liner, which was then transferred into a 600 mL PARR reactor and sealed. The PARR reactor was evacuated and backfilled ten times with 15 psig of nitrogen. Subsequently, the PARR reactor was left under 15 psig of nitrogen and heated in an oven at 230 °C for 20 hours after which it was cooled down to room temperature. The contents of the glass liner were filtered through a Buchner funnel using Whatman #4 filter paper. The filtered solid was then rinsed with 500 mL of distilled water and 500 mL of ethanol and then dried in a 90 ^C C oven for 18 hours.

[2799] The precalcined catalyst (4.90 g) was then calcined using the nitrogen calcination described herein to yield Catalyst Af3.2 (4.50 g).

[2800] Example Af 11 : Synthesis of Catalyst Af3.3

[2801] (NH ₄ ) ₆ Mqtq24 * 4 H ₂ 0 (SIGMA-ALDRICH ^® ) (13.26 g) and distilled water (180 mL) were added to a 500 mL round bottom flask. The mixture was stirred at 60 °C for 30 minutes using a warm water bath. VOS0 · 3.46 TI ₂ 0 (SIGMA-ALDRICH ^® ) (4.22 g) and distilled water (180 mL) were placed in a 250 mL beaker. The mixture was stirred at 60 °C for 30 minutes using a warm water bath. The blue vanadium solution was then slo wly poured into the clear and colorless molybdenum solution forming a dark purple solution. Subsequently, SDS (SIGMA- ALDRICH ^® ) (4.07 g) was added to the reaction mixture, which was then allowed stir for 30 minutes at 60 °C.

[2802] The contents were then transferred to a 600 mL glass liner, which was then transferred into a 600 mL PARR reactor and sealed. The PARR reactor was evacuated and backfilled ten times with 15 psig of nitrogen. Subsequently, the PARR reactor was left under 15 psig of nitrogen and heated in an oven at 230 °C for 20 hours after which it was cooled down to room temperature. The contents of the glass liner were then filtered through a Buchner funnel using Whatman #4 filter paper. The filtered solid was rinsed with 500 mL of distilled water and 500 mL of ethanol and then dried in a 90 °C oven for 18 hours to yield Catalyst Ai3.3.

[2803] Example Afl2: Synthesis of Catalyst Af3.4

[2804] (NI L), o (> _. ·, · 4 H ₂ 0 (SIGMA-ALDRICH ^® ) (13.26 g) was added to a 500 mL RBF with a magnetic stir bar. Subsequently, 180 mL of distilled water was added and the (NH ₄ ) _<> Mq7q24 · 4 ¾0 was dissolved using a water bath at approximately 60 °C. Next, VOS0 ₄ · 3.46 ¾0 (SIGMA-ALDRICH ^® ) (4.22 g) was placed into a 250 mL beaker with a magnetic stir bar. Subsequently, approximately 180 mL distilled water was added and the VOS0 ₄ • 3.46 H ₂ 0 was dissolved with the aid of water bath at approximately 60 °C. Then, the vanadium solution was added to tlie round bottom flask containing the molybdenum solution while stirring at 60 °C. Sodium dodecy 1 sulfate (SIGMA-ALDRICH ^® ) (4.07 g) was then added to the round bottom flask containing the vanadium and molybdenum solution while stirring at 60 °C and the mixture was stirred for approximately 30 minutes at 60 °C.

[2805] The solution was then transferred to a 600 mL glass liner and the round bottom flask was rinsed with distilled water and the rinse was transferred to the glass liner. The glass liner containing the dark purple solution was inserted into a 600 ml, PARR reactor. The PARR reactor was sealed. Subsequently, the PARR reactor headspace was pumped and purged ten times with nitrogen. The headspace was left under approximately 15 psig nitrogen. The reactor was placed in an oven at approximately 230 °C for approximately 21 hours and then allowed to cool for approximately 2 hours. Subsequently, the PARR reactor was removed from the oven and allowed to cool for 18 hours.

[2806J The reaction mixture was filtered using a Buchner funnel, 4 qualitative filter papers, and an aspirator. The collected precalcined catalyst was washed with distilled water. The filtrate was initially a dark blue, and the precalcined catalyst was washed until the filtrate was dear. The precalcined catalyst was greyish in color. Then, the precalcined catalyst was washed with ethanol — the filtrate was a light blue at first but transitioned to dear in color as additional ethanol was used for rinsing the solids. The precalcined catalyst material was then dried overnight in an oven at approximately 90 °C. Subsequently, the precalcined catalyst was ground manually using a mortar and pestle producing a light and fluffy precalcined catalyst.

[2807J Next. 4.95 g of the precalcined catalyst was loaded into two quartz boats and placed in a quartz tube. The quartz tube was placed into a split tube furnace. The quartz tube was then purged with bulk nitrogen at approximately 85 seem (0.085 standard liter per minute (slpm)) for 6 hours and then the flow changed to purified nitrogen and the split tube was purged for 18 hours at a flow of 85 seem.

[2808] Next, the temperature program for the split tube was set to: 4 hour ramp to 400 °C, hold for 2 hours at 400 °C; reduced nitrogen flow to approximately to 30 seem. After calcination, the catalyst was allowed to cool to room temperature and then taken out of the split tube, yielding Catalyst Af3.4 (4.65 g).

[2809] Example Afl3: Synthesis of Catalyst Af3.5

[2810] (N¾) ₆ M0 _T O24 * 4 ¾0 (SIGMA-ALDRICH ^® ) (13 26 g) was added to a 500 mL RBF with a magnetic stir bar. Subsequently i 80 mL of distilled water was added and the (NH4)6MO·,-(¾4 * 4 H ₂ 0 was dissolved using a water bath at approximately 60 °C. Next, VOSO4 · 3.463¾0 (SIGMA-ALDRICH ^® ) (4.22 g) was placed into a 250 ml. beaker with a magnetic stir bar. Subsequently, approximately 180 ml, distilled water was added and the VOSO4 • 3.46 H2O was dissolved with the aid of water bath at 60 °C. Then, the vanadium solution was added to the round bottom flask containing the molybdenum solution while stirring at 60 °C Sodium dodecyl sulfate (SIGMA- ALDRICH ^® ) (4.07 g) was then added to the round bottom flask containing the vanadium and molybdenum solution while stirring at 60 °C and the mixture was stirred for approximately 30 minutes at 60 °C.

[2 11] The solution w'as then transferred to a 600 mL glass liner and the round bottom flask w'as rinsed with distilled water and the rinse was transferred to the glass liner. The glass liner containing the dark purple solution was inserted into a 600 mL PARR reactor. The PARR reactor was sealed. Subsequently, the PARR reactor headspace was pumped and purged six times with nitrogen. The headspace was left under approximately 15 psig nitrogen. The reactor was placed in an oven at approximately 230 °C for approximately 21.5 horns and then allowed to cool for approximately 2 hours. Subsequently, the PARR reactor was removed from the oven and allowed to cool overnight. [2812] The reaction mixture was filtered using a Buchner funnel, 4 qualitative filter papers, and an aspirator. The collected precalcined catalyst was washed with distilled water. The filtrate was initially a dark blue and the precalcined catal st was washed until the filtrate was clear. The precalcined catalyst was greyish in color. Then, the precalcined catalyst was washed with ethanol — the filtrate was a light blue at first and changed to clear and colorless with the washes. The precalcined catalyst material was then dried for 18 hours in an oven at approximately 90 °C. Subsequently, the precalcined catalyst was ground manually using a mortar and pestle producing a light and fluffy precalcined catalyst.

[2813] Next, 5.47 g of the precalcined catalyst was loaded into two quartz boats and placed in a quart tube.

The quartz lube was placed into a split tube furnace. The quartz tube was then purged with bulk nitrogen at approximately 85 seem for the remainder of the day and then the flow changed to purified nitrogen and the split tube was purged overnight at a flow of 85 seem.

[2814] Next the temperature program for the split tube was set to: 4 hour ramp to 400 °C, hold for 2 hours at 400 °C; reduced nitrogen flow to approximately to 30 seem. After calcination the catal st was allowed to cool to room temperature and then taken out of the split tube, yielding Catal st Af3.5 (4.97 g).

[2815] Example Afl4: Synthesis of Catalyst AO .6

[2816] i \! i i).Vio O _. · 4 H ₂ 0 (SIGMA-ALDRICH ^® ) (44.20 g) was added to a 2 L RBF with a magnetic stir bar. Subsequently. 600 niL of distilled water was added and the (NH ₄ ) ₆ MQ ₇ 0 ₂₄ · 4 H 0 was dissolved using a 60 °C water bath. Next, VOSO4 · 3.46 H ₂ 0 (SIGMA-ALDRICH ^® ) (14.07 g) was placed into a 1 L beaker with a magnetic stir bar. Subsequently, approximately 600 ml distilled water was added and the VOSO4 * 3.46 H ₂ 0 was dissolved with the aid of a 60 °C water bath. Then, the vanadium solution was added to the round bottom flask containing the molybdenum solution while stirring at 60 °C. Sodium dodecyl sulfate (SIGMA-ALDRICH ^® ) (13.57 g) was then added to the round bottom flask containing the vanadium and molybdenum solution while stirring at 60 °C and the mixture was stirred for approximately 30 minutes at 60 °C.

[2817] After allowing to cool the solution was then transferred to a 2 L glass liner and the round bottom flask was rinsed with distilled water and the rinse was transferred to the glass liner. The glass liner containing the purple solution was inserted into a 2 L PARR reactor. The head unit was sealed. Subsequently, the PARR reactor headspace was pumped and purged ten times with the combination of nitrogen and vacuum. The headspace was left under approximately 15 psig nitrogen. The reaction was stirred at approximately 150 rpm for about 26 hours at temperature from 230 °C to 247 °C (temperature controller was set to 247 °C; the internal temperature was about 238 °C after miming overnight; the set point was then lowered to 239 °C and the internal temperature ranged from 230 °C to 231 °C). After heating, the reaction was allowed to cool overnight while stirring at 150 rpm.

[2818] The reaction mixture was then filtered using a Buchner funnel, filter paper, and an aspirator. Subsequently, the collected precalcined catalyst was washed with distilled water and ethanol. The precalcined catalyst was then dried in an oven at approximately 90 °C for 56 hours to yield 18.1 g of precalcined catal st. Next, a portion of the precalcined catalyst was ground manually with a mortar and pestle.

[2819] The precalcined catalyst was then calcined using the general nitrogen calcination process described herein to yield Catalyst Af3.6.

[2820] Catal st AG3.6 was then submitted for MRU testing. The results are presented in Table Af20.

[2821] Example Af 15 : Synthesis of Catalyst Af3.7 [2822] (NH ₄ )6Mo ₇ q24 · 4 H ₂ 0 (SIGMA-ALDRICH ^® ) (883 g) was weighed in a 500 niL round botomed flask. The white solid was dissolved in 100 mL of distilled water with the aid of a 60 °C warn water bath and stirring to create a clear, colorless solution. VOSCh · 3.46 H ₂ 0 (SIGMA-ALDRICH ^® ) (2.81 g) was weighed into a 250 mL beaker. The blue solid was dissolved in 100 rnL of distilled wa ter wi th the aid of a 60 °C warm wa ter bath and stirring to create a clear, blue solution. The blue vanadium solution was added all-at-once to the colorless molybdenum solution to instantly produce a black solution. To the black solution, 2.70 g of sodium dodeeyl sulfate (SIGMA-ALDRICH ^® ) surfactant was added as a powder to produce a purple slurry with some emulsion present.

The purple-emulsified slurry was stirred for 30 minutes at 60 °C. The purple slurry was removed from the 60 °C warm water bath and allowed to cool for 15 minutes.

[2823] The purple slurry was added to a glass liner and the glass liner was added to a 600 mL PARR high pressure reactor. The PARR reactor was sealed and purged with 15 psig nitrogen gas and vacuum sequence ten times then left under 15 psig nitrogen. The PARR reactor was placed into an oven and the oven w ^; as heated to 230 °C. The reaction proceeded at 230 °C for 20.5 hours before being turned off and allowed to cool to room temperature.

[2824] Once cooled, the reactor was vented, opened, and the resulting spongy-shirty was filtered. The filtering was done using a Buchner funnel and 4 layers of qualitative filter paper. The mother liquor was blue and the solids grey. The solids were washed with water until the filtrate ran clear. The solids were further washed with denatured ethanol. The solids were dried at 90 °C overnight and the next morning the dry solids were weighed: 2.46 g.

[2825] The precalcined catalyst was manually ground using a mortar and pestle and then loaded into a quartz boat. The quartz boat containing the precalcined catal st was loaded into a quartz tube, which was placed into a split tube furnace. The quartz tube was purged with bulk nitrogen for 5 hours at 85 seem, before the nitrogen feed w¾s fed through an oxygen scrubbing bed to further purity the nitrogen to less than 4 ppm oxygen. This ultra-high purity' (UHP) nitrogen was purged through the quartz tube for 18 hours at 85 seem. Next, the UHP nitrogen feed was reduced from 85 seem to 30 seem and the furnace was turned on and heated to 400 °C to calcine the catalyst. The heating program for the furnace w¾s: 4 hours ramp from room temperature to 400 °C, dwell at 400 °C for 2 hours and then cool to ambient temperature naturally. Calcination yielded Catalyst Af3.7.

[2826] Catalyst Af3.7 was then submited for ICP-MS analysis using the general ICP-MS procedure described herein using digestion method Af2. The results are presented in Table Afl4.

[2827] Catalyst Af3.7 was analyzed by EDS as per the general procedure SEM-EDS. The results are presented in Table Afl5.

[2828] Pow'der XRD data was collected as per General Procedure XRD above. The plot for this catalyst is available in Figure 113. The raw data PXRD plot was used in establishing the range of the peak intensities presented in Table Aflo.

[2829] PSD analysis results are presented in Table A1T7.

[2830] Cataly st Af3.7 was anal zed by nitrogen gas adsorption analysis for pore volume, BET surface area, and BJH pore size distributions analysis. Pore volume and BET surface area analysis results are presented in Table Afl8. BJH analysis results are presented in Figure 120. [2831] Catalyst 3.7 was then submited for MRU testing. The results are presented in Table 20.

[2832] Example Afl6: Synthesis of Catalyst AG.8

[2833] (NH ₄ )6MO ₇ 0 ₂ 4 * 4 H ₂ 0 (SIGMA-ALDRICH ^® ) (13.26 g) was weighed in a 500 mL round bottomed flask. The white solid was dissolved in 120 mL of distilled water with the aid of a 60 °C warm water bath and stirring to create a clear, colorless solution. VOSCh · 3.46 H ₂ 0 (SIGMA-ALDRICH ⁽ 1)) (4.28 g) was weighed into a 250 mL beaker. The blue solid was dissolved in 120 mL of distilled water with tire aid of a 60 °C warn water bath and stirring to create a clear, blue solution. The blue vanadium solution was added all-at-once to the colorless molybdenum solution to instantly produce a black solution. To the black solution. 4.11 g of sodium dodecyl sulfate (SIGMA-ALDRICH ^® ) surfactant was added as a powx!er to produce a purple slimy with some emulsion present.

The purple-emulsified shiny was stirred for 30 minutes at 60 °C.

[2834] The purple slurry was added to a glass liner and the glass liner was added to a 600 mL PARR high pressure reactor. The PARR reactor was sealed and purged with 15 psig nitrogen gas and vacuum sequence ten times then left under 15 psig nitrogen. The PARR reactor was placed into an oven and the oven was heated to 230 °C. The reaction proceeded at 230 °C for 24 hours before being turned off and allowed to cool to room temperature. [2835] Once cooled, the reactor was vented, opened, and the resulting spongy-sluny was filtered. The filtering was done using a Buchner funnel and 4 layers of qualitative filter paper. The mother liquor was blue and the solids grey. The solids were washed with water until the filtrate ran clear. The solids were further washed with denatured ethanol until ran clear. The solids were dried at 90 °C overnight and the next morning the dry solids were weighed: 4.32 g.

[2836] The 4.32 g of precalcined catalyst were manually ground using a mortar and pestle and then loaded into a quartz boat. The quartz boat containing the precalcined catalyst was loaded into a quartz tube, which was placed into a split tube furnace. The quartz tube was purged with bulk nitrogen for 5 hours at 85 seem, before the nitrogen feed was fed through an oxygen senibbing bed to further purify the nitrogen to less than 4 ppm oxygen.

This ultra-high purity (UHP) nitrogen was purged through the quartz tube for 18 hours at 85 seem. Next, the UHP nitrogen feed was reduced from 85 seem to 30 scan and the furnace was fumed on and heated to 400 °C to calcine the catalyst. The heating program for the furnace was: 4 hours ramp from room temperature to 400 °C, dwell at 400 °C for 2 hours and then cool to ambient temperature naturally. Calcination yielded Catalyst AG.8.

[2837] Catalyst AG.8 was then submited for ICP-MS analysis using the general ICP-MS procedure described herein using digestion method Af2. The results are presented in Table Afl4.

[2838] Catalyst AG.8 was analyzed by EDS as per the general procedure SEM-EDS. Results are presented in

Table AIT 5.

[2839] Pow'der XRD data was collected as per General Procedure XRD above. The plot for this catal st is available in Figure 113. The raw' data PXRD plot was used in establishing the range of the peak intensities presented in Table Aflo.

[2840] PSD analysis results are presented in Table A1Ί7. [2841] Catalyst Af3 8 was analyzed by nitrogen gas adsosption analysis for pore volume, BET surface area, and BIH pore size distributions analysis. Pore volume and BET surface area analysis results are presented in Table Afl8. BJH analysis results are presented in Figure 120.

[2842J Catalyst AD.8 was then submitted for MRU testing. The results are presented in Table Af20.

[2843J Example Afl7: Synthesis of Catalyst AG.9

[2844J (NH416M07O24 * 4 H2O (SIGMA-ALDRICH ^® ) (8.83 g) was weighed in a 250 mL round bottomed flask. The white solid was dissolved in 120 mL of distilled water with the aid of a 60 °C warm water bath and stirring to create a clear, colorless solution. VOSO4 · 3.46 ¾0 (SIGMA-ALDRICH ^® ) (2.81 g) was weighed into a 250 mL beaker. The blue solid was dissolved in 120 mL of distilled water with the aid of a 60 ^C C warm water bath and stirring to create a clear, blue solution. The blue vanadium solution was added dropwise to the colorless molybdenum solution to produce a brown solution. To the brown solution, 2.71 g of sodium dodecyl sulfate (SIGMA-ALDRICH ^® ) surfactant was added as a powder to produce a purple slurry with some emulsion present.

The purple-emulsified slurry was stirred for 30 minutes at 60 °C.

[2845] The purple slurry was added to a glass liner and the glass liner was added to a 600 ml, PARR high pressure reactor. The PARR reactor was sealed and purged with 15 psig nitrogen gas and vacuum sequence ten times, then left under 15 psig nitrogen. The PARR reactor was placed into an oven and the oven was heated to 230 °C. The reaction proceeded at 230 °C for 2.0 hours before being turned off and allowed to cool to room temperature. [2846] Once cooled, the reactor was vented, opened, and the resulting spongy-sluny was filtered. The filtering was done using a Buchner funnel and 4 layers of qualitative filter paper (Whatman 28310-02.6). The mother liquor was blue and the solids grey. The solids were washed with 0.5 L of distilled water until the filtrate ran clear. The solids were further washed with denatured 0.5 I of ethanol.

[2847] The grey solids were transferred to a 200 mL round bottom flask with 50 ml, of distilled water to create a suspension. In a 100 mL beaker, an aqueous solution of oxalic acid was made using 4 g of anhydrous oxalic acid and 50 mL water. The aqueous oxalic acid solution was added to the suspension contained in the 200 mL round botoms flask. The suspension was left mixing at 500 rpm via a magnetic stir bar in the aqueous oxalic acid solution at 80 °C for 40 minutes, after which, the solution was removed from the heat to cool for 15 minutes. The mixture was filtered through a Buchner funnel using 4 Whatman #4 filter papers. The grey filter cake was rinsed with 500 mL of distilled water. The solids were dried at 90 °C overnight.

[2848] The precalcined catalyst were manually ground using a mortar and pestle and then loaded into a quartz boat. The quartz boat containing the precalcined catalyst was loaded into a quartz tube, which was placed into a split tube furnace. The quartz tube was purged with bulk nitrogen for 5 hours at 85 seem, before the nitrogen feed was fed through an oxygen scrubbing bed to further purify the nitrogen to less than 4 ppm oxygen. This ultra-high purity (UHP) nitrogen was purged through the quartz tube for 18 hours at 85 scan. Next, the UHP nitrogen feed was reduced from 85 seem to 30 seem and the furnace was turned on and heated to 400 °C to calcine the catalyst. The heating program for the furnace was: 4 hours ramp front room temperature to 400 °C, dwell at 400 °C for 2 hours and then cool to ambient temperature naturally. The calcined catal st produced Catalyst AD.9. [2849] Catalyst Af3.9 was then submitted for ICP-MS analysis using the general 1CP-MS procedure described herein using digestion method Af2. The results are presented in Table A1Ί4.

[2850] Catalyst A13.9 was analyzed by EDS as per the general procedure SEM-EDS. Results are presented in Table Afl5.

[2851] Powder XRD data was collected as per General Procedure XRD above. The plot for this catalyst is available in Figure 113. The raw data PXRD plot was used in establishing the range of the peak intensities presented in Table Afl6.

[2852] Catalyst AO.9 was then submitted for MRU testing. The results are presented in Table AGO.

[28531 Catalyst AO .9 was then submitted for FTIR fingerprint. The results are presented in Figure 124.

[2854] Example A1Ί8: Synthesis of Catalyst AO.10

[2855] ; \ϋ i;L1o C> · 4 ICO (SIGMA-ALDRiCH ^® ) (8.83 g) was weighed in a 250 niL round bottomed flask. The white solid was dissolved in 120 mL of distilled water with the aid of a 60 °C warm water bath and stirring to create a clear; colorless solution. VOSCb · 3.46 FI _2Q (SIGMA-ALDRCH ^® ) (2.81 g) was weighed into a 250 mL beaker. The blue solid was dissolved in 120 mL of distilled water with the aid of a 60 °C warm water bath and stirring to create a clear, blue solution. The blue vanadium solution was added dropwise to the colorless molybdenum solution to produce a brown solution. To the brown solution, 2.70 g of sodium dodecyl sulfate (SIGMA- ALDRICH ^® ) surfactant was added as a powder to produce a purple shiny with some emulsion present.

The purple-emulsified sluriy was stirred for 30 minutes at 60 °C.

[2856] The purple slurry was added to a Teflon liner and the Teflon liner was added to a 600 ml PARR high pressure reactor. The PARR reactor was sealed and purged with 20 psig nitrogen gas and vacuum sequence ten times then left under 20 psig nitrogen. The P ARR reactor was placed into an oven and the oven was heated to 230 °C. The reaction proceeded at 230 °C for 20 hours before being turned off and allowed to cool to room temperature. [2857] Once cooled the reactor was vented, opened, and the resulting spongy-sluny was filtered. The filtering was done using a Buchner funnel and 4 layers of qualitative filter paper. The mother liquor was blue and the solids grey. The solids were washed with 05 L of distilled water until the filtrate ran clear. The solids were further washed with denatured 0.5 L of ethanol.

[2858] The grey solids were transferred to a 200 mL round bottom flask with 50 mL of distilled water to create a suspension. In a 100 mL beaker, an aqueous solution of oxalic acid was made using 4 g of anhydrous oxalic acid and 50 mL water. The aqueous oxalic acid solution was added to the suspension contained in the 200 mL round bottoms flask. The suspension was left mixing at 500 rprn via a magnetic stir bar in the aqueous oxalic acid solution at 80 °C for 40 minutes, after which, the solution was removed from the heat to cool for 15 minutes. The mixture was filtered through a Buchner funnel using 4 Whatman #4 filter papers. The grey filter cake was rinsed with 500 mL of distilled water. The solids were dried at 90 ^C C overnight.

[2859] The precalcined catalyst were manually ground using a mortar and pestle and then loaded into a quartz boat. The quartz boat containing the precalcined catalyst was loaded into a quartz tube, which was placed into a split tube furnace. The quartz tube was purged with bulk nitrogen for 5 hours at 85 scan before the nitrogen feed was fed through an oxygen scrubbing bed to further purify the nitrogen to less than 4 ppm oxygen. This ultra-high purity (UHP) nitrogen was purged through the quartz tube for 18 hours at 85 seem. Next, the UHP nitrogen feed was reduced from 85 seem to 30 seem and the furnace was turned on and heated to 400 °C to calcine the catalyst. The heating program for the furnace was: 4 hours ramp from room temperature to 400 °C, dwell at 400 °C for 2 hours and then cool to ambient temperature naturally. The calcined catalyst produced was Catalyst M3.10.

[2860] Catalyst M3.10 was then submitted for ICP-MS analysis using the general ICP-MS procedure described herein using digestion method M2. The results are presented in Table M14.

[2861] Catalyst M3.10 was analyzed by EDS as per the general procedure SEM-EDS. The results are presented in Table Ml 5.

[2862] Powder XRD data was collected as per General Procedure XRD above. The plot for this catalyst is available in Figure 113 The raw data PXRD plot was used in establishing the range of the peak intensities presented in Table Ml 6.

[2863] Catalyst AD .10 was analyzed by nitrogen gas adsorption analysis for pore volume, BET surface area, and BJH pore size distributions anal sis. Pore volume and BET surface area analysis results are presented in Table Afl8. BJH analysis results are presented in Figure 120.

[2864] Catalyst AD.10 was then submitted for MRU testing. The results are presented in Table Af20.

[2865] Example Ml 9 : Synthesis of Catalyst A .1 i

[2866] (NI¾) ₆ MO ₇ 0 ₂₄ · 4 ¾0 (SIGMA -ALDRICH ^® ) (8.83 g) was weighed in a 250 mL round bottomed flask. The white solid was dissolved in 120 mL of distilled water with the aid of a 60 °C warm water bath and stirring to create a clear, colorless solution VOSQt · 3.46 ¾0 (SIGMA-ALDRICH ^® ) (2.81 g) was weighed into a 250 L beaker. The blue solid was dissolved in 120 mL of distilled water with the aid of a 60 °C warm water bath and Stirling to create a clear, blue solution. The blue vanadium solution was slowly poured into the colorless molybdenum solution to produce a brown solution. To the brown solution, 2.17 g of sodium octyl sulfate surfactant (SOS SIGMA-ALDRICH ^® ) was added as a powder to produce a purple slurry. The purple slurry was stirred for 30 minutes at 60 °C.

[2867] The purple slurry was added to a glass liner and the glass liner was added to a 600 mL PARR high pressure reactor. The PARR reactor was sealed and purged with 15 psig nitrogen gas and vacuum sequence ten times, then left under 15 psig nitrogen. The PARR reactor was placed into an oven and the oven was heated to 220 °C. The reaction proceeded at 220 °C overnight before being turned off and allowed to cool to room temperature. [2868] Once cooled, the reactor was vented, opened, and the resulting slurry was filtered. The filtering was done using a Buchner funnel and 4 layers of qualitative filter paper. The solids were washed with 50 mL of distilled water. The solids were further washed with denatured 500 mL of ethanol. The solids were dried at 90 °C overnight and weighed to be 1.82 g.

[2869] The pre-calcined catalyst was manually ground using a mortar and pestle and then loaded into a quartz boat. The quartz boat containing the pre-calcined catalyst was loaded into a quartz tube, which was placed into a split tube furnace. The quartz tube was purged with bulk nitrogen for 5 hours at 85 seem, before tire nitrogen feed was fed through an oxygen scrubbing bed to further purify the nitrogen to iess than 4 ppm oxygen. This ultra-high purity (UHP) nitrogen was purged through the quartz tube for 18 hours at 85 seem. Next the UHP nitrogen feed was reduced from 85 seem to 30 seem and the furnace was turned on and heated to 400 °C to calcine the catalyst. The heating program for the furnace was: 4 hours ramp from room temperature to 400 °C, dwell at 400 °C for 2 hours and then cool to ambient temperature naturally. The calcined catalyst produced was Catalyst Af3.11.

[2870] Catalyst Af3.1 lwas then submitted for ICP-MS analysis using the general 1CP-MS procedure described herein using digestion method Af2. The results are presented in Table Af!4.

[2871] Catalyst A13.13 was anal zed by EDS as per the general procedure SEM-EDS. Results are presented in Table All 5.

[2872| Powder XRD data was collected as per General Procedure XRD above. The plot for this catalyst is available in Figure 113. The raw data PXRD plot was used in establishing the range of the peak intensities presented in Table Afl6.

[2873] Catalyst AO .11 was then submitted for MRU testing. The results me presented in Table AGO.

[2874] Example Af20: Synthesis of Catalyst AO.12

[2875] i \! i i;.\io (> _{! *} 4 IDO (SIGMA-ALDRiCH ^® ) (13.26 g) was weighed in a 500 niL round bottomed flask. The white solid was dissolved in 150 niL of distilled water with the aid of a 60 °C warm water bath and stirring to create a clear, colorless solution. VOSQ · 3.46 ¾0 (SIGMA-ALDRICH ^® ) (4.22 g) was weighed into a beaker. The blue solid was dissolved in 150 ml of distilled water with stirring at room temperature to create a clear, blue solution. The blue vanadium solution was poured slowly to the colorless molybdenum solution to instantly produce a purple solution. The resulting purple solution was stirred at 60 °C for 30 minutes. To the purple solution. 4.07 g of sodium dodecyl sulfate surfactant (SIGMA-ALDRICH ^® ) was added as a powder to produce a purple slurry. The purple shiny was stirred for 30 minutes at 60 °C and then removed from heat and allowed to cool to room temperature

[2876] The purple slurry was added to a glass liner and a magnetic stir bar added. The glass liner was added to a 600 mL PARR high pressure reactor. The PARR reactor was sealed and purged with 15 psig nitrogen gas and vacuum sequence ten times, then left under 15 psig nitrogen. The PARR reactor was setup in a heating mantle and heated via thermocouples and a temperature control box to 230 °C internal process temperature (temperature controller was set to 222 °C; the internal temperature was about 230-231 °C after sunning overnight). The reaction proceeded at 230-231 °C for 24 hours, with the slurry stirred via magnetic stirring, before being turned off and allowed to cool to room temperature.

[2877] Once cooled, the reactor was vented, opened, and the resulting spongy-sluny was filtered. The filtering was done using a Buchner funnel and four layers of qualitative filter paper. The mother liquor was blue and the solids grey. The solids were washed with water (1.5 L) until the filtrate ran clear. The solids were dried at 90 °C overnight and the next morning the dry, grey solids were weighed: 5.16 g.

[2878] The sodium dodecyl sulfate decomposition oils w¾re removed from the solids through combustion (air treatment). The 5.16 g of solids were placed in a muffle furnace open to air at 280 °C for 2 hours. After the 280 °C air treatment, the grey soiid w ^; as weighed to be 4.68 g.

[2879] The 4.68 g of precalcined catalyst were manually ground using a mortar and pestle and then loaded into a quartz boat. The quartz boat was loaded into a quartz tube which was placed into a split tube furnace. The quartz tube was purged with bulk nitrogen for 5 hours at 85 seem, before the nitrogen feed was fed through an oxygen scrubbing bed to further purify the nitrogen to less than 4 ppm oxygen. This ultra-high purity (UHP) nitrogen was purged through the quartz tube overnight at 85 sccrn. The next morning the UHP nitrogen feed was reduced from 85 seem to 30 seem and the furnace was turned on and heated to 400 °C to calcine the catalyst. The heating program for the furnace was: 4 hours ramp from room temperature to 400 °C, dwell at 400 °C for 2 hours and then cool to ambient temperature naturally. The calcined catalyst produced 4.50 g of black Catalyst Af3.12. [2880] Catalyst A13.12 was then submitted for iCP-MS analysis using the general ICP-MS procedure described herein using digestion method Af2. The results are presented in Table Af 14.

[2881} Catalyst Aft.12 was analyzed by EDS as per the general procedure SEM-EDS. Results are presented m Table Afl5.

[2882] Catalyst AD .12 was submitted for SEM imaging. The results ate presented in Figure 118.

[2883] Powder XRD data was collected as per General Procedure XRD above. The plot for Catalyst AD.12 is available in Figure 113. The raw data PXRD plot was used in establishing die range of the peak intensities presented in Table Afl6.

[2884] PSD analysis results are presented in Table Af 17.

[2885] Catalyst AD.12 was analyzed by nitrogen gas adsorption analysis for pore volume, BET surface area, and BJH pore size distributions analysis. Pore volume and BET surface area analysis results are presented in Table Afl8. BJH analysis results are presented in Figure 120.

[2886] Catalyst AD.12 was then submitted for MRU testing. The results are presented in Table ADO.

[2887] Example AD 1 : Synthesis of precalcined Catalyst AD.13

[2888] (N¾)6Mo ₇ 0 ₂ 4 * 4 H ₂ 0 (SIGMA- ALDRICH ^® ) (44.20 g) was added to a 2 L round bottom flask with a magnetic stir bar. Subsequently, approximately 600 mL of distilled water was added and the (MUleMo-O _/ u · 4 H 0 was dissolved with the aid of a 60 °C water bath. Next, VOSO4 * 3.46 H ₂ 0 (SIGMA- ALDRICH ^® ) (14.07 g) was placed into a 1 L beaker with a magnetic stir bar. Subsequently approximately 600 ml, of distilled water was added and the VOSO4 · 3.46 H 0 was dissolved with the aid of a 60 °C water bath. Then, the vanadium solution was added to the round bottom flask containing the molybdenum solution while stirring at 60 °C. Sodium dodecyl sulfate (S1GMA-ALDRICH ^® ) was added to the round bottom flask containing the vanadium and molybdenum solution while stirring at 60 °C and the mixture was stirred for approximately 30 minutes at approximately 60 °C.

[2889] The round bottom flask was then removed from the water bath and allowed to cool for 15 approximately minutes before transferring the solution to a 2 L glass liner. The round bottom flask was rinsed with distilled water and the rinse was transferred to the glass liner. The glass liner containing the purple solution was inserted into a 2 L PARR reactor. The PARR reactor unit was sealed and subsequently, the PARR reactor headspace was pumped and purged ten times with nitrogen. The headspace was left under approximately 15 psig nitrogen and rite top valve on the PARR reactor was closed. Subsequently, the PARR reactor was stirred at approximately 150 rprn using an overhead stirrer at a process temperature of 230 °C (controller temperature set for 248 °C) for approximately 24 hours. The reaction mixture was then allowed to cool to approximately 67 ^C C and stirred at approximately 150 rpmfor 18 hours. [2890] The reaction mixture was filtered warm using a Buchner funnel, 4 qualitative filter papers, and an aspirator. The collected solid precalcined catalyst was rinsed with about 1 L of distilled water to produce precalcined Catalyst A13.13.

[2891] Example AΪ22: Synthesis of Catalyst AG.14

[2892] A portion of precalcined Catalyst AJ3.13 was placed in a 100 rnL beaker and dried in an oven at 90 °C overnight. The dried precalcined catalyst solids, weighing 7.34 g, were manually ground using a mortar and pestle and then loaded into a quartz boat. The quartz boat was loaded into a quartz tube, which was placed into a split tube furnace. The quartz tube was purged with bulk nitrogen for 5 hours at 85 seem, before the nitrogen feed was fed through an oxygen scrubbing bed to further purify the nitrogen to less Ilian 4 ppm oxygen. This ultra-high purity (UHP) nitrogen was purged through Site quartz tube overnight at 85 scan. The next morning the UHP nitrogen feed was reduced from 85 seem to 30 scent and the furnace was turned on and heated to 400 °C to calcine the catalyst.

The heating program for the furnace was: 4 hours ramp from room temperature to 400 °C, dwell at 400 °C for 2 hours and then cool to ambient temperature naturally. After the calcination was complete, large amount of yellow oil were observed throughout the tube. The calcined catalyst produced 4.92 g of Catalyst AG.14. CHN analysis was performed on the calcined catalyst (C% 2.18, H% 0.48, N% 0.21).

[2893] Catalyst AG.14 was then submitted for MRU testing. The results are presented in Table AGO.

[2894] Example Af23 : Synthesis of Catalyst A .15

[2895] A portion of precalcined Catalyst AG.13 was transferred to a Buchner funnel with 4 qualitative filter papers and washed with 500 nil, of ethanol using an aspirator vacuum. The solids were dried at 90 °C overnight and weighed to be 4.74 g.

[2896] The 4 74 g of precalcined catalyst was manually ground using a mortar and pestle and then loaded into a quartz boat. The quartz boat was loaded into a quartz tube, which was placed into a split tube furnace. The quartz tube was purged with bulk nitrogen for 5 hours at 85 seem, before the nitrogen feed was fed through an oxygen scrubbing bed to further purify the nitrogen to less than 4 ppm oxygen. This ultra-high purity (UHP) nitrogen was purged through the quartz tube for 18 hours at 85 seem. Next the UHP nitrogen feed was reduced from 85 scan to 30 seem and the furnace was turned on and heated to 400 °C to calcine the catalyst. The heating program for the furnace was: 4 hours ramp from room temperature to 400 °C, dwell at 400 °C for 2 hours and then cool to ambient temperature naturally. The calcined catalyst produced 4.70 g of Catalyst AG.15. CHN analysis was performed on the calcined catalyst (C% 0.07, H% 0.30, N% 0.13).

[2897] Catalyst AG.15 was then submitted for ICP-MS analysis using the general 1CP-MS procedure described herein using digestion method Af2. The results are presented in Table A1I4.

[2898] Catalyst AG.15 was analyzed by EDS as per the general procedure SEM-EDS. Results are presented in Table Afl5.

[2899] Catalyst AG.15 was submitted for SEM imaging. The results are presented in Figure 118.

[2900] Powder XRD data was collected as per General Procedure XRD above. The plot for tills catalyst is available in Figure 113. The raw data PXRD plot was used in establishing the range of the peak intensities presented in Table Aflo. [2901] PSD analysis results are presented in Table Afl7.

[2902] Bulk density measurements are presented in Table Afl9.

[2903] Catalyst AG.15 was then submitted for MRU testing. The results are presented in Table Af20.

[2904] Example AΪ24: Synthesis of Catalyst AG.16

[2905] A portion of precalcined Catalyst AG .13 was transferred to a 100 nxL beaker and dried at 90 °C overnight. The dried solids were placed in a muffle furnace open to air at 280 °C for 3 hours. After the 280 °C air treatment, the grey solid was -weighed to be 6.49 g.

[2906] The 6.49 g of precalcined catal st was manually ground using a mortar and pestle and then loaded into a quartz boat. The quartz boat wns loaded into a quartz tube which was placed into a split tube furnace. The quartz tube w'as purged with bulk nitrogen for 5 hours at 85 scent, before the nitrogen feed wns fed through an oxygen scrubbing bed to further purify the nitrogen to less than 4 ppm oxygen. This ultra-high purity (UHP) nitrogen was purged through the quartz tube for 18 hours at 85 seem. Next the UHP nitrogen feed was reduced front 85 seem to 30 scent and the furnace was turned on and heated to 400 °C to calcine the catalyst. The heating program for the furnace was: 4 hours ramp from room temperature to 400 °C, dwell at 400 °C for 2 hours and then cool to ambient temperature naturally. The calcination yielded 6.21 g of Catalyst AG.16. CHN analysis was performed on the calcined catalyst (C% 0.54, H% 0.21, N% 0.18).

[2907] Catalyst AG.16 was then submitted for ICP-MS analysis using the general ICP-MS procedure described herein using digestion method Af2. The results are presented in Table Af!4.

[2908] Catalyst AG.16 was analyzed by EDS as per the general procedure SEM-EDS. Results are presented in Table Afl5.

[2909] Catalyst A .16 was submitted for SEM imaging. The results are presented in Figure 118.

[2910] Powder XRD data wns collected as per General Procedure XRD above. The plot for this catalyst is available in Figure 113 The raw data PXRD plot wns used in establishing the range of the peak intensities presented in Table Afl6.

[2911 ] PSD analysis results are presented in Table Af 17.

[2912] Catalyst A .16 was analyzed by nitrogen gas adsosption analysis for pore volume, BET surface area, and BJH pore size distributions analysis. Pore volume and BET surface area analysis results are presented in Table Afl8. BJH analysis results are presented in Figure 120.

[2913] Catalyst A .16 was then submitted for MRU testing. The results are presented in Table AGO.

[2914] Example AGS: Synthesis of Catalyst AG.17

[2915] A 5.1 g portion of precalcined Catalyst Af!.3 wns loaded into a quartz boat. The quartz boat containing the precalcined catalyst was loaded into a quartz tube, which wns placed into a tube furnace. The quartz tube was purged with bulk nitrogen for 8 hours at 400 seem, before the nitrogen feed wns fed through an oxygen scrubbing bed to further purify the nitrogen to less than 4 ppm oxygen. This ultra-high purity (UHP) nitrogen wns purged through the quartz tube overnight at 400 seem. The next morning the furnace wns turned on and heated to 400 °C to calcine the catalyst. The heating program for the furnace was: 4 hours ramp from room temperature to 400 °C, dwell at 400 °C for 2 hours and then cool to ambient temperature naturally. The calcined catalyst. Cataly st AD.17, weighed 4.80 g.

[2916] Catalyst Af3.17 was analyzed by EDS as per the general procedure SEM-EDS. Results are presented in Table Ail 5.

[2917] Powder XRD data was collected as per General Procedure XRD above. The plot for this catalyst is available in Figure 113. The raw' data PXRD plot was used in establishing the range of the peak intensities presented in Table Afl6.

[2918| Catalyst AD.17 was then submited for MRU testing. The results are presented in Table ADO and

Table ADO.

[2919] Example AD6: Synthesis of Catalyst AD.18

[2920] (Nl¾) ₆ Mo ₇ 0 ₂₄ · 4H ₂ 0 (SIGMA-ALDRICH ^® ) (44.2035 g) was charged to a 2 L round bottom flask with 600 mL of distilled water, this was heated and stirred in a 60 °C oil bath until the solution became clear and colorless (5 minutes). VOSCh · 3.36H ₂ 0 (SIGMA-ALDRICH ^® ) (14.0699 g) was charged to a separate 1 L beaker with 600 mL of distilled water. This was heated and stirred using a 60 °C oil bath until the solution turned an electric blue color ( 10 minutes). The electric blue vanadyl sulfate solution was transferred to the ammonium molybdate solution in the 2L round bottom flask, and the resulting purple solution was held at 60 °C. To the 21 round bottom flask was also charged sodium dodecyl sulfate (SIGMA-ALDRICH ^® , 13.5735 g) at 60 °C. The purple solution was stirred in a 60 °C water bath for 30 minutes. Next, the solution was transferred to a 2L glass liner. The final volume was recorded to be 1340 mL. The glass liner was transferred into the 2 L reactor, the gap between the glass liner and the reactor was filled with distilled water. Afterwards, the reactor was sealed, and the headspace purged ten times by filling the reactor with 15 psig N ₂ (g) and evacuated with vacuum. Subsequently 15 psig N ₂ (g) headspace was left in the reactor. Next the reactor was placed in an oven for 24 hours at 230 °C with a 1 -hour ramp time and a 24-hour cooling time. Afterwards, the reactor was vented, and the contents of the reactor were filtered through 4 qualitative filter papers. The dark purple powder collected on the filter paper was rinsed with 1.25 L of distilled water. The dark purple powder w¾s dried in the oven at 90 °C overnight. Following the drying step, the purple powder ( 17.49 g) was air treated in a muffle furnace ( 1-hour ramp to 280 °C and dwell for 26 hours). After the muffle furnace treatment the purple powder (14.39 g) was submitted for CHN analysis (C% 0.87, H% 0.11, N% 0.65). The precalcined catalyst was pressed with a spatula into a quartz boat. The quartz boat containing the precalcined catalyst was loaded into a quartz tube, winch was placed into a split tube furnace. The quartz tube was purged with bulk nitrogen for 8 hours at 400 seem, before the nitrogen feed was fed through an oxygen scrubbing bed to further purity the nitrogen to less than 4 ppm oxygen. This ultra-high purity (UHP) nitrogen was purged through the quartz tube for 18 hours at 400 seem. Next, the furnace was turned on and heated to 400 °C to calcine the catalyst. The heating program for the furnace was: 4 hours ramp from room temperature to 400 °C, dwell at 400 °C for 2 hours and then cool to ambient temperature naturally .CHN analysis was performed on the calcined catalyst (C% 0.45. H% 0.02. N% 0.54).

[2921] Catalyst AD.18 was analyzed by EDS as per the general procedure SEM-EDS. Results me presented in Table Af22.

[2922] Catalyst AD.18 was submitted for SEM imaging. The results are presented in Figure 118. [2923] Powder XRD data was collected as per General Procedure XRD above. The plot for Catalyst All .18 is available in Figure 113. The raw data PXRD plot was used in establishing the range of the peak intensities presented in Table Afl6.

[2924] PSD analysis results are presented in Table Afl7.

[2925] Catalyst A13.18 was then submitted for MRU testing. The results are presented in Tables Af20, Af28 and Af29.

[2926] 1CP-MS Analysis of Catalysts AD.7-AD.12, Af3.15 and Af3.16.

[2927] The ICP-MS results for Catalysts AD.7-AD.12, AD.15 and AD.16 are presented in Table A1Ί4.

[2928] Table AT14: ICP-MS results

[2929] EDS Analysis of Catalysis AD.7-AD.12 and AD.15-AD.18.

[2930] The EDS analysis of Catalysts AD.7-AD.12 and AD.15-AD.18 are presented in Table Afl5.

[2931] Table A 15: EDS Analysis results

[2932] PXRD Analysis of Catalyst Material Af3.7-Af3. 12 and Af3.15-Af3. 18.

[2933] Table Afl6 presents reflection angles and corresponding intensity for Catalysts Af3.7-Af3.12 and Af3.15-Af3.I8 (nitrogen calcined baseline). Figure 113 presents the plot correlating to peak numbers.

[2934] Table Af 16

[2935] PSD Analysis of Catalysts AD.7, AD.8, AD.12, AD.15, and AD.16

[2936] The statistical data from the PSD of Catalysts AD.7, AD.8, AD.12. AD.15, and AD.16 is presented in

Table Afl7.

[2937] Table Af 17

[2938] BET Analysis of Catalysts AD.7, AD.8, AD.10, AD.12, and AD.16.

[2939] The results of nitrogen gas adsorption analysis for pore volume and BET surface area anal sis for Catalysts ADD, AD.8, AD.10, AD.12, and AD.16 are presented in Table Aft 8. [2940] Table Afl8

[2941] Bulk density measurements for Catalyst Af3.15.

[2942] The bulk density measurements for Catalyst Af3 15 are presented in Table Afl9. [2943] Table Af 19

[2944] Activity and Selectivity of for Catalysts Af3.1, Af3.6-Af3.12 and Af3.14-Af3.18.

[2945] The MRU results for Catalysts Af3.1, Af3.6-Af3.12 and Af3.14-Af3.18 are presented in Table Af20.

[2946] Table Af20

[2947] Synthesis of Catal st Materials Af4.1-AF4.9 [2948] Example Af27 : Synthesis of Catalyst Material Af4.1

[2949] Cataly st AC.18 (4.5 g, 32 wt %) was placed in a 100 ml. beaker. Subsequently, 8.44 g (60 wt. %) of

VERSA!.™ 250 alumina (UOP) and goethite (SIGMA-ALDRICH ^® ) (1.12 g, 8 wt. %) were added to the beaker. Distilled water (approx. 40 mL) was added to the beaker and the mixture was stirred manually. The beaker was then placed in a 100 °C silicone oil bath and the reaction mixture was stirred at approximately 85 rpm. The oil bath temperature was increased to 110 °C after about an hour and the reaction mixture was stirred until it took on a pastelike consistency, which took about 1.5 hours. The beaker was then taken out of the oil bath and placed into an oven at 90 °C and allowed to d y overnight. Subsequently, the material was taken out of the oven and placed in a muffle furnace at 350 °C for 2.5 hours and then allowed to cool overnight to room temperature yielding Catalyst Material Af4.1.

[2950] Catalyst Material Af4.1 was analyzed by ICP-MS using the general ICP-MS procedure described herein using digestion method AO. The ICP-MS results for Catalyst Material AIT.1 are presented in Table Af21. [2951] Catalyst Material Af4.1 was analyzed by EDS as per the general procedure SEM-EDS. Results are presented in Table Af22.

[2952] Catalyst Material Af4.1 was submitted for SEM imaging. The results are presented in Figure 119.

[2953] Powder XRD data was collected as per the general XRD disclosed herein. The plot for Catalyst

Material Af4.1 is available in Figure 114. The raw data PXRD plot was used in establishing the range of the perils; intensities presented in Table Af23.

[2954| Catalyst Material Af4.1 was then submitted for particle size distribution analysis by SEM. The results are presented in Table Af24.

[2955] BET Analysis results are presented in Table A£25. BJH analysis results are presented in Figure 122.

[2956] Bulk density measurements are presented in Table Af26.

[2957] Catalyst Material Af4.1 was then submitted for MRU testing. The results are presented in Table Af27.

[2958] Catalyst Material Af4.1 was submitted for FTIR fingerprint analysis. The results are presented in

Figure 123.

[2959] Example Af28: Synthesis of Catalyst Material AF4.2.

[2960] To a 100 mL beaker was charged 4.47 g of Catalyst All.4, 6.705 g ofVERSAL™ 250 alumina, and 20 mL of distilled water. The beaker was then placed directly on a heat plate and the heat plate was set to 80 °C.

Further, au overhead agitator w'as set up with a glass stir rod and 0.5 inch Teflon stir blade, and the slurry was agitated at 100 rpm. After 1 25 hours the shirty turned into a paste. The material was then dried overnight in a 90 °C oven. Subsequently, the material was calcined in air in a muffle furnace at 350 °C for 2 hours with a ramp time of 30 minutes to yield Catalyst Material 4.2.

[2961] Catalyst Material Af4.2 was submited for SEM imaging. The results are presented in Figure 119. [2962] Powder XRD data was collected as per the general XRD procedure described herein. The plot for Catalyst Material Af4.2 is available in Figure 114. The raw data PXRD plot was used in establishing the range of the peak intensities presented in Table AT23.

[2963] BET Analysis results are presented in Table A£25. BJH analysis results are presented in Figure 122. [2964] Bulk density measurements are presented in Table Af26.

[2965] Catalyst Material AG4.2 was submitted for MRU testing. The results are presented in Table AG27.

[2966] Catalyst Material AG4.2 submitted for FTIR fingerprint analysis. The results are presented in Figure 123.

[2967] Example Af29: Synthesis of Catalyst Material AX4.3

[2968] 13.26 g of (NH ₄ ) ₆ Mo ₇ q24 · 4H ₂ 0 was charged into a 500 mL round bottom flask with 180 niL of distilled water. The white powder was dissolved using a magnetic stir bar while being placed in a 60 °C water bath. To a 250 mL beaker with a magnetic stir bar was charged 4.22 g of VOS0 · 3.4ό3¾0 and 180 mL of distilled water. The beaker which contained a magnetic stir bar, was clamped into a 60 °C water bath, and the solution was stirred for 30 minutes at 60 °C to dissolve. After which, the electric blue solution of VOS0 was charged into the round bottom flask containing the (NFLOeMo/O r solution while continuously stirring at 60 °C. Follo wing this addition, 4.07 g of sodium dodecyi sulfate was charged into the round bottom flask. The mixture was stirred for 30 minutes at 60 °C. Subsequently, the solution in the round bottom flask was removed from the water bath for one hour and cooled to room temperature. The room temperature solution was transferred to a 600 mL glass P ARR rector liner. Rinsings from the round bottom flask were charged to the 600 ml, liner. The reactor was sealed and evacuated ten times with vacuum and 15 psig nitrogen. Following, the tenth evacuation and nitrogen back fill the PARR reactor was left under 15 psig of nitrogen and sealed. The PARR reactor was placed in a 230 °C oven for 19 hours and 15 minutes. After which, it was removed from the oven and left to cool for 2 hours. Subsequently, the pressure was released, the reactor was opened, and the contents were filtered through 4 quantitative filter papers using a Buchner funnel. The resulting purple, grey powder was rinsed with distilled water and ethanol. Following the filtration the grey catalyst wras dried in an oven at 90 °C for 90 hours.

[2969] The catalyst was taken out of oven and cooled to room temperature. After which, the catalyst was ground using a mortar and pestle.4.95 g of the light and fluffy catalyst was transferred into two split tube boats, purged with bulk nitrogen gas at 14 scpm for 6 hours. Subsequently, the split tube containing the cataly st was purged wife purified nitrogen at 14 scpm overnight. The temperature program on the split tube was turned on. The split tube hated up to 400 °C over 4 hours, held at 400 °C for 2 hours and then cooled down to room temperature with a flow of 5.5 scpm. The calcination yielded 4.65 g of catal st which was removed from the split tube.

[2970] A portion of the catalyst prepared above (4.0 g), VERBAL™ 250 alumina (UOP) (6.0 g), SUPERFLOC ^® (CYTEC) (0.023 g; polyacrylamide N-100/300), and distilled water (approx. 40 mL) were added to a 100 tnL beaker. The beaker was then placed into an oil bath at approximately 100 °C and the mixture was stirred at approximately 85 rpm until the mixture formed a paste, which took about 1.5 hours. The beaker was then taken out of the oil bath and placed in an oven at approximately 90 °C overnight to dry.

[2971] Next, the material was taken out of the oven and placed into a muffle furnace at 350 °C for 2.5 hours yielding Catalyst Material Af4.3. Catalyst Material Af4.3 was then allowed to cool.

[2972] Catalyst Material Af4.3 was then submitted for MRU testing. The results are presented in Table Af27. [2973] Example Af30: Synthesis of Catalyst Material Af4.4

[2974] To a 500 mL RBF with a magnetic stir bar was charged 13.26 g of (NH ₄ ) ₆ Mo7Q ₂ 4 * 4H ₂ 0 and 180 mL of distilled water. The round bottom flask was clamped into a water bath, and heated to 60 °C. To a separate 250 mL beaker with a magnetic stir bar was charged 4.22 g of VOSO4 · 3.46FFQ and 180 mL of distilled water. The beaker was clamped into a 60 ^C C water bath and the VOSO4 * 3 46H ₂ 0 was stirred to dissolve over the course of 20 minutes. Following that, the electric blue vanadyl solution was charged to the clear and colorless ammonium molybdate solution at 60 °C. To this purple solution was charged 4.07 g of sodium dodecylsulfate wiiiie continuously stirring the solution at 60 °C. This mixture was stirred at 60 °C for 30 minutes. Subsequently, the solution was removed from the water bath, cooled to room temperature and transferred to a 600 mL PARR reactor glass liner. The liner was placed into the PARR reactor and the reactor was sealed. Following this, the PARR reactor was purged with nitrogen gas and evacuated slowly ten times. The PARR reactor was left under 15 psig of nitrogen gas. The PARR reactor was placed in an oven and heated to 230 °C for 22.5 hours. Next the PARR reactor was cooled, the pressure was released and the glass liner containing the catalyst aqueous mixture was removed.

[2975] Subsequently, the purple/grey catalyst aqueous mixture was filtered through a Buchner funnel with 4 qualitative filter papers. A water aspirator was set up to draw vacuum on the catalyst mixture. The catalyst mixture on the filter paper was rinsed with 1.2 L of distilled water and 500 mL of ethanol resulting in a grey colored catalyst filter cake. Following the filtration and rinsing step the catalyst was dried at 90 °C for 18 hours. Following the diving step the light and fluffy catalyst powder was ground using a mortar and pestle. To a calcination boat was weighed 5.47 g of catalyst. This boat was placed into the split tube calcination furnace and purged with bulk nitrogen at 14 slpm for 6 hours. Following that, the split tube furnace was purged using purified nitrogen at 14 slpm for 18 hours. Subsequently, the split tube furnace was ramped up to 400 °C for 4 hours, held at 400 °C for 2 hours and cooled to room temperature under a 5.5 slpm flow' of purified nitrogen. This yielded 4.97 g of a grey catalyst material.

[2976] Catalyst prepared from the description above (4.0 g), VERSAL™ 250 alumina (UOP) (6.0 g), SUPERFLGC ^® (CYTEC) (0.022 g; polyacrylamide N-100/300), and distilled water (approx. 25 mL) were added to a 100 mL beaker. The beaker was then placed into an oil bath at approxima tely 100 °C and the mixture was stirred at approximately 85 rpm until the mixture formed a paste, which took about 1 hour. The beaker was then taken out of the oil bath and placed in an oven at approximately 90 °C overnight to dry.

[2977] Next, the material w'as taken out of the oven and placed into a muffle furnace at 350 °C for 2.5 hours to yield Catalyst Material AG4.4.

[2978] Catal st Material Af4.4 was then submitted for MRU testing. The results are presented in Table AF27. [2979] Example Af31 ; Synthesis of Catalyst Material Af45

[2980] Catalyst AD .6 (3.52 g), VERSAL™ 250 alumina (UOP) (6.6 g), goethite (0.88 g), and distilled water (approx. 40 mL) were added to a 100 mL beaker. Next, the beaker was placed into an oil bath at approximately 100 °C and stirred at approximately 80 rpin until the mixture took on a past-like consistency, winch took approximately

1.5 hours. The reaction mixture was stirred using an electronic overhead stirrer equipped with Teflon blade agitator. The beaker was then taken out of the oil bath and placed in an oven at approximately 90 °C allowed to dry over the weekend. Subsequently, the material was broken up with a spatula and then placed in a muffle furnace at 350 °C for

4.5 hours to yield Catalyst Material Af4.5 (10.14 g).

[29811 Catalyst Material AG4.5 was analyzed by ICP-MS using the general ICP-MS procedure described herein using ICP digestion method AD. The ICP-MS results forCatalyst Material AG4.5 are presented in Table Af21.

[2982] Catalyst Material Af4.5 was then submitted for MRU testing. The results are presented in Table Af27. [2983] Example Af32: Synthesis of Catalyst Material AG4.6

[2984] Catalyst AD.18 (2.75 g), VERSAL™ 250 alumina (UOP) (6.41 g) and distilled water (approximately 20 ml.) were added to a 100 mL beaker which formed a light purple suspension. Subsequently, the beaker was clamped into an oil bath at ap roximately 100 °C and stirred at approximately 106 ppm until the mixture took on a periwinkle purple past-like consistency, which took approximately 1.4 hours. The reaction mixture was stirred using an electronic overhead stirrer equipped with a half inch Teflon™ agitator blade and glass stir shaft. Subsequently the beaker was removed from the oil bath and placed in a 90 °C oven to dry for approximately 18 hours. The light purple powder was crashed up and heated in an air muffle furnace at 350 °C for 2 hours with a ramp time of 0.5 hours. This yielded a light grey Catalyst Material AF4.6 (7.79 g).

[2985] Catalyst Material Af4.6 was analyzed by EDS as per the general SEM-EDS procedure described herein. Results are presented in Table Af22.

[2986] Catalyst Material Af4.6 w'as submitted for SEM imaging. The results are presented in Figure 119. [2987] Powder XRD data was collected as per the general XRD procedure described herein. The plot for Catalyst Material Af4.6 is available in Figure 114. The raw data PXRD plot was used in establishing the range of the peak intensities presented in Table Af23.

[2988] PSD analysis results are presented in Table Af24.

[2989] Catalyst Material Af4.6 wus then submitted for MRU testing. The results are presented in Tables Af28 and Af29.

[2990] Example AD3 : Synthesis of Catalyst Material Af4.7

[2991] Catalyst AD.18 (1.51 g), VERSAL™ 250 alumina (UOP) (1.50 g), goethite (2.09 g), and distilled water (approximately 20 mL) were added to a 100 mL beaker which formed a purple suspension. Next the beaker was clamped into an oil bath at approximately 100 °C and stirred at approximately 100 rpni until the mixture took on a dark green paste-like consistency, which took approximately 4 hours. The reaction mixture was stirred using an electronic overhead stirrer equipped with a half inch Teflon™ agitator blade and glass stir shaft. The beaker was then removed from the oil bath and placed in an oven at approximately 90 °C to dry for approximately 18 hours. Afterwards, the golden-brown powder was crashed up and heated in an air muffle furnace at 350 °C for 2 hours with a ramp time of 0.5 hours. This yielded a merlot red catalyst material, Cataly st Material A14.7 (4.54 g).

[2992] Catalyst Material Af4.7 was analyzed by 1CP-MS using the general ICP-MS procedure described herein using ICP digestion method AG. The ICP-MS results forCatalyst Material Af4.7 are presented in Table Af21.

[2993] Catalyst Material Af4.7 was analyzed by EDS as per the general procedure SEM-EDS. Results are presented in Table AT22.

[2994] Catalyst Material AG4.7 was submitted for SEM imaging. The results are presented in Figure 119. [2995] Powder XRD data was collected as per the general XRD procedure described herein. The plot for Catalyst Material Af4.7 is available in Figure 114. The raw data PXRD plot was used in establishing the range of the peak intensities presented in Table Af23.

[2996] PSD anal sis results are presented in Table Af24.

[2997] Catalyst Material Af4.7 was then submitted for MRU testing. The results are presented in Tables Af28 and Af29.

[2998] Example Af34: Synthesis of Catalyst Material AF4.8

[2999] Catalyst AG.18 (1.26 g), goethite (2.92 g), and distilled water (approximately 20 ml.) were charged to a 10 nil beaker which fo rmed an olive-green suspension. Afterwards, the beaker was clamped into an oil bath at approximately 100 °C and stirred at approximately 110 rpm until the mixture took on an olive-green paste-like consistency which took approximately 1.5 hours. The reaction mixture was stirred using an electronic overhead stirrer equipped with a half inch Teflon™ agitator blade and glass stir shaft. Ne xt the beaker was removed from the oil bath and placed in an oven at approximately 90 °C to dry for approximately 18 hours. Subsequently, the olive- green powder was crashed and heated in an air muffle furnace at 350 °C for 2 hours with a ramp time of 0.5 hours. This yielded a merlot red, orange catalyst material, Catalyst Material Af4.8 (3.50 g).

[3000] Catalyst Material Af4.8 was analyzed by ICP-MS using the general ICP-MS procedure described herein using ICP digestion method AG. The ICP-MS results for Catalyst Material Af4.8 are presented in Table Af21.

[30011 Catalyst Material Af4.8 was analy zed by EDS as per the general SEM-EDS procedure described herein. The results are presented in Table Af22.

[3002] Catalyst Material AG4.8 procedure was submitted for SEM imaging. The results are presented in Figure 119.

[3003] Pow'der XRD data w'as collected as per the general XRD procedure described herein. The plot for this catalyst is available in Figure 114. The raw data PXRD plot was used in establishing the range of the peak intensities presented in Table Af23.

[3004] PSD analysis results are presented in Table Af24.

[3005] Catal st Material AG4.8 was then submitted for MRU testing. The results are presented in Tables AC28 and AG29. [3006] Example .4135; Synthesis of Catalyst Material Af49

[3007] Catalyst A13.18 (3.29 g), goethite (0.82 g), VERSAL™ 250 alumina (UOP) (6.17 g), and distilled water (25 mL) was charged to a 100 mL beaker. Next, the beaker was clamped in an oil bath, an overhead agitator assembly was set up using a half inch stir blade and a glass sir shaft. Following the setup, the oil bath w'as heated to 100 °C and the agitator w'as set to 80 rpm. Subsequently, the suspension was heated and stirred until it became a paste (approximately 1.5 hours). Next, the paste (22.6713 g) was dried in an oven at 90 °C overnight. The resulting powder was air calcined in a muffle furnace at 350 °C for 2.5 hours with a ramp time of 1 hour. After calcination, 9.07 g of Catalyst Material Af4.9 was recovered.

[3008] Catalyst Material Af4.9 was analyzed by EDS as per the general SEM-EDS procedure described herein. The results are presented in Table AT22.

[3009] Catalyst Material AG4.9 was submitted for SEM imaging. The results are presented in Figure 119.

[3010] Powder XRD data was collected as per the general XRD procedure described herein. The plot for this catalyst is available in Figure 114. The raw data PXRD plot was used in establishing the range of the peak intensities presented in Table 23.

[3011] PSD analysis results are presented in Table Af24.

[3012] Catalyst Material Af4.9 was then submitted for MRU testing. The results are presented in Tables Af28 and Af29.

[3013] ICP-MS Analysis of Catalyst Materials Af4.1, Af4.5, Af4.7, and Af4.8

[3014] The results for the ICP-MS analysis of Catalyst Materials Af4.1, Af4.5, Af4.7 and Af4.8 are present in Table Af21.

[3015] Table Af21 [3016] EDS Analysis of Catalyst Ai3.18 and Catalyst Materials Af4.1 and Af4.6-Af4.9

[3017] The EDS analysis of Catalyst Af3.18 and Catalyst Materials AIT.1 , Af4.6-Af4.9 are presented in Table

Af22.

[3018] Table Af22

Table Af22

[3019] PXRD Analysis of Catalyst AD.18 and Catalyst Materials Af4.1-Af4.2, AG4.6-M4.9

[3020] Table Af23 presents the reflection angles and corresponding intensity for Catalyst AD.18 and Catalyst

Materials Af4.1-M4.2, Af4.6-Af4.9. A plot correlating to peak numbers is presented in Figure 114.

[3021] Table AD 3

[3022] PSD Analysis of Catalyst Af3.18 and Cataly st Materials Af4 1, Af4.6-Af4.9

[3023] Table Af24 presents the statistical data from the PSD analysis of Catalyst Af3.18 and Catalyst

Materials Af4. 1 and Af4.6-Af4.9.

[3024] Table Af24

[3025] BET Analysis of Catalyst AD.18 and Catalyst Materials Af4.1-Af4.2, M4.6-AG4.8 [3026] The results of nitrogen gas adsorption analysis for pore volume and BET surface area analysis for Catalyst AB.18 and Catalyst Materials Af4.i~Af4.2, Af4.6-Af4.8 are shown in Table Af25.

[3027] Table Af25

[3028] Bulk density measurements of Catalyst Materials Af4.1 -Af42

[3029] Table Af26 presents the bulk density measurements for Catalyst Material Af4.1 and AG4.2. [3030] Table Af26

[3031] Activity and Selectivity of for Catalyst and Catalyst Materials [3032] Table Af27 presents the MRU results for Catalyst Materials Af4.1-M4.5. [3033] Table Af27

[3034] Table Af28 presents the MRU results for Catalyst AO.18 and Catalyst Materials Af4.6-Af4.9. during first days of MRU testing.

[3035] Table Af28

[3036] Table Af29 presents the MRU results for Catalyst AO.18 and Catalyst Materials Af4.6-Af4.9. after several days of MRU testing, illustrating performance changes arising from tire catal st being exposed to the ethylene conversion process. [3037] Table Af29

*Molar formula as determined by EDS. J -molar formula from ICP-MS

[3038] Example Af36: Synthesis of Catalyst Material AG5.1

[3039] Catalyst Afl.4 (1.75 g) and Catalyst AD.17 (1.93 g) were combined with 3.6823 g of PB 250 alumina (Aiumax)and 1.84 g of goethite in a beaker. The combined solids were dry mixed together then 60 niL of water was added to create a green suspension. The green suspension was heated in die beaker using a 100 °C oil bath and mixed at 80 rpm using a motor-driven overhead stirring Teflon agitator. After 2.5 hours, the water was evaporated to a smearable paste. The sample was dried in the oven at 90 °C overnight, then calcined at 350 °C for 2.5 hours under air in a muffle furnace. The resulting 8.24 g of red solids were Catalyst Material Af5 l

[3040] Catalyst Material Af5.1 was analyzed by EDS as per the general SEM-EDS procedure described herein. The results are presented in Table Af30.

[3041] Powder XRD data was collected as per the general XRD procedure described herein. The plot for Catalyst Material Af5.1 is available in Figure 115. The raw data PXRD plot was used in establishing the range of the peak intensities is presented in Table A131. [3042] Catalyst Material Af5.1 was then submited for MRU testing. The results are presented in Table Af32. [3043] EDS Analysis of Catalyst Material Af5.1

[3044] Table Af30 presents the EDS Analysis results for Catalyst Material Af5.1.

[3045] Table Af30

[3046] PXRD Analysis of Catalyst Material Af5.1

[3047] Table Af31 presents the reflection angles and corresponding intensity for Catalyst Material Af5.1. Figure 115 is the plot correlating to peak numbers.

[3048] Table Af31 [3049] Activity and Selectivity for Catalysts Afl .4 and AC.17, and Catalyst Material Af5.1.

[3050] Table Af32 presents the MRU results for Catalysts Afl.4 and Af3.17 and Catalyst Material Af5.1.

[3051] Table Af32

[3052] Amorphous Content Analysis and XRD Phase Fitting

[3053] Table AT33 present the amorphous content analysis for Catalyst Af 1.1, Afl.4-Afl.6, Af3.7-AD.12, and Af3.15-AT3.18 mid for Catalyst Materials AT4.1, AT4.2, Af4.6-Af4.9, and Af5.1.

[3054] Table Af33

[3055] Microreactor Unit

[3056] The ability of catalysts and catalyst materials described herein to participate in the oxidative dehydrogenation of ethane was tested in a microreactor unit (MRU).

[3057] Setup AgA

[3058] The MRU included a reactor tube made from SS316L stainless-steel SWAGELOK ^® Tubing, which had an outer diameter of 0.5 inches, an internal diameter of about 0.4 inches, and a length of about 15 inches. A 6- point WIKA Instruments Ltd K-type thermocouple having an outer diameter of 0.125 inches was inserted axially through the center of the reactor, which was used to measure and control the temperature within the catalyst bed. A room temperature glass tight sealed condenser was located after the reactor to collect water/acidic acid condensates. The gas product flow was allowed to either vent or was directed to a gas chromatography (Agilent 6890N Gas Chromatograph, using CHROMPERFECT ^® - Analysis, Version 6.1.10 for data evaluation) via a sampling loop. [3059] For the MRU testing, a pie-mixed feed gas was fed through the reactor. The pie-mixed feed gas entering the reactor was 36 moi.% ethane, 18 mol.% oxygen, and 46 moi.% nitrogen. Further, the pre-mixed feed gas flow was adjusted by a calibrated mass flow controller to obtain a gas hourly space velocity (GHSV) of about 3,000 h ¹ , based on the catalyst volume in the catalyst bed.

[3060] The flow rate of the feed gas was between about 70 standard cubic centimeters per minute (seem) to about 80 sccin (e.g., 76.1 scan). The catalyst bed placed in the reactor tube can include the catalyst or catal st material and a filler. With reference to the MRU’s catal st bed, a filler refers to a material that does not participate in the oxidative dehydrogenation of ethane or have other catalytic activity, such as non-selective oxidation under the MRU test conditions. The filler was quartz sand. The 35% conversion temperature was determined at a weight hourly space velocity (WHS V) of 2.90 h ¹ , with the WHSV based on the amount of catal st or the amount of catalyst used to prepare the catalyst material, and a gas hourly space velocity (GHSV) of about 3.000 h ¹ . Whereby WHSV was defined as mass flow of feed gas to the reactor divided by the weight of the catalyst in the catalyst bed, GHSV was defined as volumetric flow of the reactor feed gas divided by the volume of the catalyst bed.

[3061] Typically, the inlet pressure was in the range of about 1 pound per square inch gauge (psig) to about 2.5 psig and the outlet pressure was in the range of about 0 psig to about 0.5 psig. The gas feed exiting the catalyst bed was be analyzed by gas chromatography to determine the percent of various hydrocarbons (e.g.. ethane and ethylene) and, optionally other gases such as C¾, €(¾ and CO. [3062] Setup AgA was employed for all samples except Catalyst Material Ag4.1 and Catalyst Material Ag4.1.1.

[3063] Setup AgB

[3064] The MRU included a reactor tube made from SS316L stainless-steel SWAGELOK ^® Tubing, which had an outer diameter of 0.5 inches, an internal diameter of about 0.4 inches, and a length of about 15 inches. A 6- point WIKA Instruments Ltd. K-type thermocouple having an outer diameter of 0.125 inches was inserted axially through the center of the reactor, which was used to measure and control the temperature within the catalyst bed. A room temperature 316 SS sealed condenser was located after the reactor to collect water/acidic acid condensates.

The gas product flow was allowed to either vent or was directed to a gas chromatography (Agilent 6890N Gas Chromatograph, using CHROMPERFECT ^® - Analysis, Version 6.1.10 for data evaluation) via a sampling loop. [3065] For the MRU testing, a pre-mixed feed gas was fed through the reactor. The pre-mixed feed gas entering the reactor was 36 rnol.% ethane, 18 mol.% oxygen, and 46 mol.% nitrogen. Further, the pre-mixed feed gas flow was adjusted by a calibrated mass flow' controller to obtain a gas hourly space velocity (GHSV) of about 5,619 h ¹ , based on the catalyst volume in the catalyst bed.

[3066] The flow rate of the feed gas was about 150 seem. The catalyst bed placed in the reactor tube can include the catalyst or catalyst material and a filler. With reference to the MRU's catalyst bed, a filler refers to a material that does not participate in the oxidative dehydrogenation of ethane or have other catalytic activity such as non-selective oxidation under the MRU test conditions. The filler was quartz sand. The 35% conversion temperature was determined at a weight hourly space velocity (WHSV) of 9.16 h ¹ , with the WHSV based on the amount of catalyst or the amount of catalyst used to prepare the catalyst material, and a gas hourly space velocity (GHSV) of about 5,619 h ¹ . Whereby WHSV was defined as mass flow of feed gas to the reactor divided by the weight of the catalyst in the catalyst bed, GHSV was defined as volumetric flow' of the reactor feed gas divided by the volume of the catalyst bed.

[3067] Typically, the inlet pressure w¾s in the range of about 1 pound per square inch gauge (psig) to about 2.5 psig and the outlet pressure was in the range of about 0 psig to about 0.5 psig. The gas feed exiting the catalyst bed was analyzed by gas chromatography to determine the percent of various hydrocarbons (e.g., ethane and ethylene) and, optionally other gases such as (¼, C0 ₂ , and CO.

[3068] Setup AgB was employed for Catal st Material Ag4.1 and Catalyst Material Ag4.!.l.

[3069] Common to both Setup AgA and Setup AgB

[3070] To prepare catalyst and catalyst materials for testing in the MRU, the catalyst or catalyst material was loaded into a I -inch round die and pressed with 8 tons of compression force for 10 to 15 seconds of dwelling time. The pressed catalyst or catalyst material was then crashed into small pieces using a mortar and pestle. Note: for catal st materials which were pressed on the CPR-6 Automated press the 3 x 3 mm pellets were gently crushed with a mortar and pestle. The crushed catal st or catal st material was then sieved and a particle sizes between 425 mpi and 1 m n were collected to be loaded for testing on the MRU.

[3071] For MRU experiments, the catalyst bed was prepared by physically mixing 1.00-2.00 g of catalyst with quartz sand such that foe catalyst bed had a total volume of about 6-8 inL. The catalyst bed was loaded in the middle zone of the reactor — located between points 2 and 5 of the thermocouple — and the remaining volume of the reactor was packed with quartz sand (Figures 136 and 137). The load was then secured with glass wool on the top and the bottom of reactor.

[3072] The gas exiting the reactor was analyzed by gas chromatography (Agilent 6890N Gas Chromatograph, using CHROMPERFECT ^® - Analysis, Version 6.1.10 for data evaluation) to determine the percent of various hydrocarbons (e.g., ethane and ethylene) and, optionally other gases such as (.¾, €(¾, and CO and acetylene.

[3073] A catalyst or catalyst material’s 35% conversion temperature was determined as follows. Conversion of the feed gas was calculated as a mass flow' rate change of ethane in the product compared to feed ethane mass flow rate using the following formula:

[3074] In the above equation, C is the percent of feed gas that has been converted from ethane to another product (i.e., ethane conversion) and X is the molar concentration of the corresponding compound in the gaseous effluent exiting the reactor at corresponding temperature. The ethane conversion was then plotted as a function of temperature to acquire a linear algebraic equation. The linear equation for ethane conversion was solved to determine the temperature in which the ethane conversion was 35% (i.e. the 35% conversion temperature).

[3075] Further, the gas exiting the reactor was analyzed by gas chromatography to determine catalyst selectivity to ethylene (i.e., the percentage on a molar basis of ethane that forms ethylene). Selectivity to ethylene was determined using the following equation:

[3076] In the above equation, S _Ethyiene is the selectivity to ethylene and X is the molar concentration of the corresponding compound in the gaseous effluent exiting the reactor at corresponding temperature. The selectivity to ethylene w¾s determined at the 35% conversion temperature, unless otherwise indicated. As such, after the 35% conversion temperature was determined, the above equation for selectivity was solved using the corresponding values for X^i _ene» Xco ₂ , and Xco at the 35% conversion temperature.

[3077] When reported, acetic acid production w'as determined by running MRU testing long enough to collect an aqueous condensate in the condenser (e.g., 1-3 days). After collecting a sample of the condensate, the sample was submitted for liquid GC analysis (Agilent 6890N Gas Chromatograph, Using CHROMPERFECT ^® - Analysis, Version 6.1.10 for data evaluation). To perform the liquid GC analysis, 300-450 mg of liquid sample was transferred to a scintillation vial Next, 25 mg of isopropanoi (IP A) was added as an internal standard. Further, 18-20 rnL of distilled ¾0 was added to dilute the sample. Prepared samples were then transferred to GC vials and set in sequence to tested using an auto sampler. The GC analysis w ^; as a split injection method with a temperature program and FID detector. Further, a set of 3 calibration standards were run in duplicate for the relative response factor used for calculating acetic acid content in sample. [3078] For MRU experiments, the catalyst bed was prepared by Method AgA or Method AgB.

[3079] Method AgA: Any added beryllium oxide was considered part of the catalyst weight loading. Catalyst weight loadings and gas flow's are kept the same (2.00 g and a WHSV of 2,9 h ¹ ). Catalyst bed volumes are kept constant at 6 mL. This method was used for the loading and running of Catalyst Agl.l as well as MoVBeC\ catalyst materials.

[3080] Method AgB: Any added aluminum oxide was not considered part of the catalyst weight loading. For example, if the catalyst material includes aluminum oxide in a 60 to 40 weight percentage ratio (MoVBeO _* wt. % to AIOc wt. %), tiie typical loading of 2.00 was divided by 0.6 to calculate the target Method AgB loading of: 3.33 g. Catalyst bed volumes were in the range of 5 mL to 8 mL, depending on the density. Catalyst weight loadings and gas flows are kept the same (2.00 g and a WHSV of 2.9 h ¹ ). This method w'as used for the MoVBeA10 _x catalyst materials, whereby no diluent was required as the mixture had 80% alumina in the mixture. Both sand and aluminum oxide are considered to be inert for the purposes of this MRU method.

[3081] All catalyst materials with added aluminum oxide in tins study were loaded in Method AgB. When Method AgB was used for MoVBeA10 _x catalyst material, the beiyllia was considered for catalyst weight loading and any additional aluminum oxide was not considered for catal st weight loading. Therefore, for a catal st material was prepared from 40 wt. % alumina (e.g. boehmite), 57.9 wt. % MoVO _x and 2.1 wt. % betyllia, then 60% of the resulting catalyst material was considered for cataly st weight loading. Again catalyst weight loadings and gas flows are kept the same (2.00 g and a WHSV of 2.9 h ¹ ).

[3082] Acetic acid conversions were not measured on the GC because this product condenses out in the water product, where water was a product of the ODH process. Additionally, not enough aqueous acetic acid condensate was produced during an MRU screening run (single day run) to be accurate quantified by GC. Longer collection times (roughly x time) were required to quantify the amount of acetic acid produced as byproduct

[3083] ICP-MS

[3084] Samples were prepared according to one of the foilowing two digestion methods:

[3085] Sodium Hydroxide Preparation

[3086] Digestion of sample was conducted to bring the sample into solution prior to dilution in nitric acid. Sample (10 mg) was placed into a scintillation vial with 3 mL sodium hydroxide solution (6.25 mol/L). The sample solution was stirred via stir-bar in a 90 °C oil bath. Once the sample was digested, the solution was transferred into an ICP-MS containment vessel with the scintillation vial being rinsed three times with a total of 15 mL ICP grade water. The rinses are added to the ICP-MS containment vessel. The solution was then brought up to 25 mL with ICP grade water. The solution was analyzed via ICP-MS. Weights are recorded throughout the preparation process to be entered into the ICP-MS software for result calculations.

[3087] Lithium Metaborate Fusion Preparation

[3088] Fusion of sample was conducted for amalgamation prior to dissoivingxdilution in nitric acid. Sample (10 mg) w'as placed into a platinum crucible with 0.1 g Lithium metaborate (98.5%) / Lithium Bromide (1.5%) covering the sample. The crucible was placed into a muffle furnace at room temperature and brought up to 1000 °C over 2 hours. Once at 1000 °C, the sample remains in the muffle furnace for 20 minutes before the temperature program was turned off. The sample cools down in the muffle fiimaee until the muffle furnace reaches 500 °C, at which point the crucible was removed and placed at a cooling station to continue to cool to room temperature. The crucible with amalgamated sample was placed on a stir plate and slowly stirred via stir-bar with 5 mL 5% nitric acid for 2 hours to dissolve the sample into solution. The solution was transferred into an ICP-MS containment vessel with the crucible being rinsed three times with a total of 15 mL 5% nitric acid. The rinses are added to the ICP-MS containment vessel. An additional lOOx dilution with 5% nitric acid was conducted prior to analysis via ICP-MS. Weights are recorded throughout the preparation process to be entered into the ICP-MS software for result calculations.

[30891 XRD analysis

[3090] Powder X-Ray Diffractometiy (PXRD) data was collected using a PANaly tical Aeris X-ray diffractometer by SEMx Incorporated. This diffractometer instrument consisted of three basic elements: X-ray tube, sample holder, and X-ray detector. X-rays were generated in a cathode ray tube (Cu source with Ka radiation = 1.5418 A) with the resulting X-ra s being directed onto the sample. As the sample and detector are rotated, die intensity of the reflected X-rays was recorded to produce characteristic X-ray spectra. When the incident X-rays reflecting off the sample satisfies the Bragg Equation (hl=2ά sin Q), constructive interference occurs and a peak in intensity occurs (y-axis). X-ray diffractometers were setup such that the sample rotates in the path of the X-ray beams at an angle 0, while the X-ray detector was mounted on an arm to collect the diffracted X-rays and rotates at an angle of 20 from ~5° to 70° (x-axis).

[3091] Qualitative XRD analysis and Rietveid Refinement was performed using HighScore Plus XRD analysis software. The samples were finely ground to reduce particle size and to obtain a uniform mixture. They were then loaded onto the XRD sample holder and the XRD spectrum was acquired. The Rietveid Refinement results were combined with Highscore Plus and EDS results to perform qualitative and quantitative analysis.

[3092] The weight percentage of amosphous content was determined by external standard. With an external standard phase, the instrument intensity constant K-factor, was determined. Corundum was used as the external standard and was measured with the same instrument configuration shortly after the unknown sample was measured. The K-factor approach is described by O’Connor and Raven: 1988, Powder Diffraction, 3 (1), 2-6. For each sample, the weight percentage of the crystalline MoVO _x orthorhombic phase had to be determined in order to assign weight percentages to the amorphous content. The Degree of Crystallinity (DOC) Method, based on the estimation that the total intensity of area contributed to the overall diffraction pattern by each component in the analysis, was used to determine the amount of amorphous phase.The degree of crystallinity was calculated from the total areas under the defined crystalline and amorphous components from:

DOC = Crystalline Area Crystalline Area + Amorphous Area where the weight fraction of the amorphous material was calculated from:

W amorphous = 1 - DOC.

[3093] Comparative raw data analysis

[3094] The PXRD raw data was also analyzed using a Python code through the program Spyder. The code generated overlaid plots. It also analyzed the data by comparing the peak prominence of all the local maxima and generated a plot with peaks meeting an established threshold. Relevant catalyst peaks are highlighted in the plot with vertical lines and the range of the relative peak intensities were provided by the code.

[3095] SEM

[3096] Scanning electron microscope (SEM) images were collected using a JSM-IT300LV

INTOU CHSCOPE™. Samples were prepared on an aluminum stud with double sided carbon tape.

[3097] SEM-EDS

[3098] Energy-dispersive X-ray spectroscopy (EDS) was conducted using a JEOL JED-2300 DRY SDD EDS detector. Samples were sent to SEMx Incorporated for EDS analysis. The samples were finely ground to reduce particle size and obtain (he uniform mixture. They were then loaded onto EDS stub for analysis by SEM. EDS was used for elemental analysis and surface examination. EDS is a micro-analytical technique that provides a semi- quantitative elemental analysis of the surface of a sample (e.g., the top 1 to 3 microns). The SEM was used to examine the surface morphology at magnifications ranging from 20 to 100,000 times. The EDS instrument delects elements with an atomic number equal or greater than sodium, but also lias light element capability, which means that it can also detect carbon, nitrogen, oxygen, and fluorine. The estimated lower detectable limit for any given element generally is between about 0.2 and 0.5 wt. %

[3099] Particle size distribution by SEM

[3100] Samples were sent to SEMx Incorporated for particle size analysis using scanning electron microscopy (SEM), model JEOL - JSM300 LV. SEM was used to observe and count the particles in the sample to obtain the Particle Size Distribution (PSD). For the PSD measurements, the SEM instrument took pictures at different magnifications. Measurements were done for 400-800 particles at different magnifications to cover the size range (statistical population). Size was measured by length in micrometers and was measured on the longest dimension of the particles. SEM based PSD is a method for analyzing samples where particles are agglomerated (stuck together) because the analyst can visually see this through the microscope and make the judicious decision to measure the distinct particles rather than the agglomerates. Statistics and analysis were based on total counts measured by SEM. [3101] Yield calculations

[3102] Theoretical yield calculations were based on the weight of each reagent used.

The weight of each reagent used in grains was divided by the molecular weight in grams per mol. For example: Weight (NFL-O _ft MO _/ CEi · 4¾0 (g) / Molecular weight (NH ₄ ) _& q ₇ q ₂₄ · 4¾0 (g/inol) = (NH ₄ ) ₆ Mq ₇ q ₂₄ · 4H ₂ 0 mol [3103] This calculation was performed for the vanadium as well.

[3104] The theoretical moles of the final product were calculated by assuming that both the molybdenum and the vanadium have attained the highest oxidation states in the final product. Thus, molybdenum and vanadium formed Mo(¾ and V2O5 respectively. For example:

[3105] The moles of the starting material were used and multiplied by die respective molar equivalents of each of the total oxidized species. The moles were then multiplied by the predicted theoretical weight of the fully oxidized final product in order to get the final theoretical weight of the catalyst. For example: [3106] Theoretical weight of molybdenum in the fully oxidized state:

(NH ₄ )QMO ₇ 0 ₂₄ · 4H ₂ 0 mol x 7 - 7(MoO _; . mol)

((MoO inol)) x M0O3 g/rnol) = Theoretical weight of fully oxidized molybdenum in grams [3107] Theoretical weight of vanadium in the fully oxidized state:

(VOSO4 · 3.46H2O mol x ½ , ½ V2O5 mol)

(V2O5 mol x V2O5 g/mol) = Theoretical weight of fully oxidized vanadium in grams

Total theore tical weight in g = (Theoretical weight of fully oxidized molybdenum in grams) + (Theore tical weight of fully oxidized vanadium in grains)

Percent yield = (Actual measured yield (g) / Theoretical yield (g)) x 100

[3108] The percent yield was determined by diving the actual measured yield by the theoretical yield and multiplying by 100.

[3109] Crush strength testing

[3110] The crush strength testing was done with the use of the standard ASTM method D4179-11, Standard Test Method for Single Pellet Crush Strength of Formed Catalyst and Catalyst Carriers using a Torbal force gauge FB500. The maximum force capacity of the gauge was 500 N with a resolution of 0. 1 N.

[3111] Bulk Density measurement

[3112] To a 10 nil graduated cylinder. The graduated cylinder was tarred and filled to the 10 ml mark with pelletized catalyst. The graduated cylinder was tapped such that the pelletized catalyst settled in the cylinder. The weight of the catalyst that fits in the cylindrical 10 ml portion was recorded. This weight was divided by 10 mL to get the bulk density measurement.

[3113] MRU Results

[3114] The MRU 35% ethane conversion results for select samples, as well as beryllium oxide powder and calcium carbonate powder are presented in Table Agl, Table Ag2, Table Ag3, and Table Ag4. The overlaid MRU performance for Catalyst Material Ag2.4 and Catalyst Material Ag2.6 is presented in Figure 128. The 24h + collection of data in Figure 128 is experimentally equivalent given experimental error. The overlaid MRU performance for Catalyst Material Ag4 1 and Catalyst Material Ag4. 1 .1 is presented in Figure 129. The data presented in Figure 129 show's that the addition of 5% calcium carbonate to Catalyst Material Ag4.1 via the “dry mixing method” (Catalyst Material Ag4.1.1) provided a slight increase in selectivity' comparatively to Catalyst Material Ag4.1. As used herein, a diy mixing method mixes all dry ingredients for a catalyst together, prior to pressing the dry ingrediaents into a catalyst pellet. This improvement in selectivity comes with a slight decrease in activity' . The addition of calcium carbonate is beneficial for pressing the Catalyst Material (Ag4.1 vs Ag4.1.1), without damaging the die sets. [3115] Table Agl

† Temperature of 380°C provided 0 29 mo3.% ethane conversion. No higher conversion could be obtained through direct measurement.

.j. Temperature of 380°C provided 0.21 mol.% ethane conversion. No higher conversion could be obtained through direct measurement.

JDelta T (°C) is defined as the difference between the 35% conversion temperature relative to Catalyst Agl.1.

[3116] Table Ag2

JDelta T (°C) is defined as the difference between the 35% conversion temperature relative to Catalyst Agl.2. *Only two MRU data points used to establish data set. More performance data is presented in Figure 128. Note: 35% ethane conversion temperatures reported above 380°C are extrapolated, not interpolated.

[3117] Table Ag3

† value obtained from linear algebraic expression was below zero.

$ Delta T (°C) is defined as the difference between the 35% conversion temperature relative to Catalyst Agl.3. 35% ethane conversion temperatures reported above 380°C are extrapolated, not interpolated.

[3118] Table Ag4 Delta T (°C) is defined as the difference between the 35% conversion temperature relative to Catalyst Agl.4. Note: 35% ethane conversion temperatures reported above 380°C are extrapolated, not interpolated.

[3119] Elemental Analysis

[3120] The ICP-MS analysis and EDS analysis for Catalyst Agl .1 as well as Catalyst Materials Agl 1 and Agl .2 are presented in Table 5. EDS is not well suited for identifying elements lighter than Na. As such, the contents of Be cannot be identified by this technique. Ranges were established by assuming all of the alumina was either AiOOH or AI2O3. Catalyst base material ranges were established from ICP-MS measurements of various catalyst active phase batches

[3121] Table Ag5

[3122] SEM and PSD Analysis

[3123] Table Ag6 presents the particle size analysis from SEM for Catalyst Agl . 1 as well as Catalyst Materials Ag2.i and Ag2.2. The SEM image of Catalyst Agi .1 at a IO,OOOc magnification is presented in Figure 130. The SEM image of Catalyst Material Agl. 1 at a IO,OOOc magnification is presented in Figure 131. The SEM image of Catalyst Material Agl.2 at a IO,OOOc magnification is presented in Figure 132, The SEM image of Catalyst Material Ag2.2.1 at a IO,OOOc magnification is presented in Figure 133. The SEM image of Catalyst Material Ag2.6 at a 5,000x magnification is presented in Figure 134. The SEM image of Catalyst Material Ag2.4 at a IO,OOOc magnification is presented in Figure 135. The SEM image of BeO at a IO,OOOc magnification is presented in Figure 136.

[3124] Table Ag6

[3125] XRD Analysis

[3126] PXRD diffractograms for Catalyst Agl.l, Catalyst Material Agl.l, Catalyst Material Agl.2, Catalyst Material Ag2.2.1, Catalyst Material Ag2.4, and Catalyst Material Ag2.6 are presented in Figure 141.

[3127] Table Ag7 presents the crystallite size for Catalyst Agl.l, Catalyst Material Agl.l, and Catalyst Material Agl .2, which was calculated by the Scherrer equation, using the main peak at 22, 1 Rad 2Q, and the raw data from the PXRD diffractograms. Since this peak in indicative of the M1/M2 phase (and possibly other mixed metal oxide phases), the reported crystallite sizes only reflect the crystalline phases. Addition of BeO to Catalyst Agl.l resulted in a decrease in crystallite size. However, addition of alumina to Catalyst Material Agl.l resulted in an increase of the crystallite size, even comparatively to Catalyst Agi.1. This could suggest that the promoter and support interact with the crystalline phases.

[3129] The XRD phase fitting and amorphous content analysis for Catalyst Agl.l, Catalyst Materials Agl.l, Catalyst Material Agl.2, Catalyst Material Ag2.2.1 , Catalyst Material Ag2.4, and Catalyst Material Ag2.6, as well as beryllium oxide and VERSA!.™ 250 Alumina are presented in Table Ag8. The theoretical amorphous content was calculated using the amorphous content of Catalyst Agl .1 , BeO, and VERSAL™ alumina, and is meant to represent what amorphous content should have been observed if no phase changes occurred after the addition of promoter and support. The theoretical amorphous content could not be calculated for Catalyst Material Agl.l.1 because the amorphous content of calcium carbonate was not identified. In all examples, the amorphous content increased comparatively to the theoretical amount.

[3130] Table Ag8

[3131] Ml orthorhombic phase is identified as a phase fitted with either 04-022-1665 or 04-022-1664, or a combination of both

[3132] Analysis of the final formulated catalyst material by XRD and SEM, indicated that promoting and supporting the MoVO _x catalyst with beryllium and aluminum oxides respectively is accompanied by a change in crystallinity, as shown in Table Ag8. The crystallinity of the overall mixture significantly decreases, which implies that a recrystallization or phase transition has occurred during the wet mixing process. The resulting active phase crystal size and overall particle size also increased as a result (Table Ag6, Table Agl). The phase transition is surprising as it would not be expected that a simple wet mixing of metal oxides would be accompanied by a change in crystallinity. This is likely due to the alumina and baseline catalyst interaction causing a recrystallization of the catalyst mixture. This in turn increases the crystal size which increases the overall pastiele size of the catalyst, beryllium and aluminum oxide mixture.

[3133] A Python code (scipy .singal.find peaks) was used to identity peaks in the PXRD raw data for Catalyst Agl.l, Catalyst Materials Agl.l, Catalyst Material Agl.2, Catalyst Material Ag2.2.1, Catalyst Material Ag2.4, and Catalyst Material Ag2,6 (Table Ag9, Figure 141). This code identifies peaks by analyzing the prominence of maxima (prominence^ 50, wlen=100). Due to the setting used, very broad peaks overlapping with sharp peaks were not always identified by the code. Peak 9, located at approximately 22.2° 2Q, is the reference peak for relative intensities. Versa!™ 250 Alumina lias broad peaks at 13.91, 28.21, 38.46, 49.11, 55.66, 64.99° 2Q.

[3134J Table Ag9

[3135] Catalyst Material Ag2.4 and Catalyst Material Ag2.6 have more VERSAL™ 250 Alumina in their formulations. The boehmite peaks from VERSAL™ 250 Alumina obscure keys peaks which the Python code could not resolve, as can be seen in Table Ag9. Therefore, the PXRD diffractograms for Catalyst Material Ag2.4 and Catalyst Material Ag2.6 were not included in the peak range and relative intensity analysis presented in Table Agio and Figure 142 (peak range analysis for Catalyst Agl.l, Catalyst Material Agl.l, Catalyst Material Agi.2, Catalyst Material Ag2.2 1). Peak 9, located at approximately 22.2° 20, is the reference peak for relative intensities. Since Catalyst Material Agl.l and Catalyst Material Ag2 2.1 contained less VERSAL™ 250 Alumina, almost all peaks could be resolved, with the exception of peak 20 for Catalyst Material Agl.l and peak 34 for Catalyst Material Ag2.2.1. These two specific data points are therefore omitted from the data ranges presented in Table Agio. [3136] Table Agio

*Catalyst Material Agl. 1 omited. Python code could not resolve overlapped peak.

'Catalyst Material Ag2.2.1 omitted. Python code could not resolve overlapped peak from boehmite alumina

[3137] Weight Percent Analysis

[3138J The bulk chemical weight percent, comparing the XRD phase fitting and the EDS measurements for Catalyst Agl.l, Catalyst Materials Agl.l, Catalyst Material Agl.2, Catalyst Material Ag2,2.1, Catalyst Material Ag2.4, and Catalyst Material Ag2.6 are presented in Table Agl 1.

[3139] Table Agl 1

* EDS is not well suited for identifying elements lighter than Na. As such, the contents of Be cannot be identified by this technique.

[3140] FTIR Analysis

[3141] The FTIR analysis for Catalyst Agl .1 and Catal st Material Ag2.2 are presented in Table Agl2 as well as Figure 140.

[3142] Table Agl2

[3143] Cnish Strength and Bulk Density

[3144] The crash strength of Catalyst Material Agl.2 is presented in Table Agl3. [3145] Table Agl3

[3146] Press speeds ranged from 1750 to 2500, even when not recorded. The radius of all pellets was measured to be 3 min.

[3147] The bulk density for pelletized Catalyst Material Agl.2 was measured to be 0.51 g/niL.

[3148] Synthesis of Samples

[3149] Sample Summary

[3150] Table Agl4

[3151] Synthesis of Catalyst Agl.l

[3152] A solution of (NH ₄ ) ₆ Mq ₇ q ₂₄ ·4H ₂ 0 (44.20 g, 35.77 mmol, white solid) in 600 mJL of dH ₂ G was prepared in a 2-L RBF equipped with magnetic stir bar. A solution of VOSCV3.461¼0 (14.07 g, 62,95 rnrnol, bright blue solid) in 600 niL of dH ₂ 0 w'as prepared in a 1-L beaker equipped with magnetic stir bar. Both solutions w'ere stirred in a 60 °C water bath until homogeneous. The blue vanadium solution was then added to the clear colorless molybdenum solution. This resulted in a dark purple solution with a fine suspension. Sodium dodecyl sulfate (SDS) (13.57 g, 47.06 mmol, white solid) was added to the reaction mixture. The purple slurry was left to stir at 60 °C for 30 minutes.

[3153] The reaction mixture was transferred to a glass liner, with a total volume of about 1375 niL measured after rinsing. The liner was loaded into a 2-L PARR reactor mid the gap filled witli dl-LO. The reactor was sealed and the head space evacuated and backfilled witli N ₂ gas lOx times. The headspace was left under 15 psig N ₂ gas and sealed. A heating mantel and insulation was used to heat the reaction for 24 hours at 230 °C (heating mantel controller set to 240 °C). Once cooled to room temperature, the reactor was vented, and the contents filtered using a Buchner funnel and 4 quantitative filter papers. The filter cake was rinsed with 1180 nxL of room temperature dH ₂ 0 and 900 mL of 80 °C d¾0. The filtrate was a dark blue color and the product was a silvery /grey purple color.

[3154] The filter cake was dried in the oven at 90 °C overnight with 42.42 g of product being recovered. The dry powder product was roughly crushed using a spatula, and then loaded into the muffle furnace for an air treatment (26 hours at 280 °C). There w'as 20.92 g of product recovered after the air treatment. The dry powder product w'as loaded into two quartz boats and centered in the quartz tube of the QRU. The quartz tube w'as purged with purified nitrogen overnight. The furnace was then ramped up to 400 °C at a rate of 1.6 °C/min. The catal st was calcined at 400 °C for 2 horns and then cooled to ambient temperature naturally. An SEM image of Catal st Agl.l is show'll in Figure 137.

[3155] Synthesis of Catalyst Material Agl.1

[3156] To a 400 mL beaker was charged 18.3035 g of Catalyst Agl.1 and 1.5949 g of beryllium oxide with 75 mL of distilled w'ater forming a purple slurry. The beaker was clamped in an oil bath and an overhead agitator assembly was set up using a one-inch Teflon stir blade and a glass stir shaft. The overhead agitator assembly was set to 100 rpm and the oil bath w'as set to 100 °C. The purple slurry was stirred and heated for 1.16 hours until the slurry became a black / purple paste. The beaker containing the paste was removed from the oil bath. The oil on the outside of the beaker was removed using heptane and the beaker was placed in a 90 °C oven to dry for 18 hours. After such time the beak r containing the dried black/purple paste was transferred to a muffle furnace where it was calcined for 2 hours with a 30-minute ramp time. Subsequently, the beaker containing the black/purple powder was removed from the muffle furnace yielding 19.0480 g of dark purple powder. An SEM image of Catalyst Ag2.1 is shown in Figure 138.

[3157] Synthesis of Catalyst Material Agl .2

[3158] To a 400 mL beaker was charged 16.0480 g of Catalyst Material Agl .1 and 24 0858 g of VERSAL™ 250 Alumina along with 75 mL of distilled water forming a dark purple slurry. The beaker was clamped in an oil bath and an overhead agitator was assembled using a one-inch Teflon stir blade and a glass stir shaft. The overhead agitator was set to 100 rpm and the oil bath was set to 100 °C. The dark purple slum' was heated for 55 minutes until the slum' became a dark purple paste. The oil on the outside of the beaker containing the dark purple paste was removed using heptane and the beaker and paste were dried in a 90 °C oven for 18 hours. Following the drying step at 90 °C the dried paste in the beaker was removed, the powder was ground using a mortar and pestle and muffle furnace calcined at 350 °C for 2 horns with a rarnp time of 30 minutes. An SEM image of Catalyst Ag2.2 is shown in Figure 139.

[3159] The catalyst w'as pelletized using the auto press and both radial and axial crush crash strength measurements were taken at three different speeds of the auto press. Crash strength was determined using ASTM D4179.

[3160] The built density was also measured for this catalyst using a standard bulk density measurement procedure. [3161] Synthesis of Catalyst Agl .2

[3162] Three separate vessels were filled with distilled water and heated to the desired temperature. The starting materials were dissolved in each vessel while stirring. Table Agl 5 shows the conditions of the starting chemicals preparation.

[3163] Table Agl5

[3164] The solutions were each stirred for about 10 minutes until homogeneous solution were obtained. The total water used for the reaction, including rinsing the vessels, was 57 L. The solution of ammonium molybdate tetrahydrate from Vessel- Agl was pumped into the 100-L reactor vessel at the pump rate of 3.2 L/min. Once transferred, the reactor was stirred and the solution of vanadyl sulfate hydrate from Vessel~Ag2 was added to the 100-L reactor vessel at the pump rate of 3.2 L/min. Lastly, the sodium dodecyl sulfate solution in Vessel-Ag3 was pumped into the 100-L reactor vessel at the rate of 3.2 L/min. The reaction mixture was allowed to stir in the reactor for 30 minutes. While the reaction w ^; as stirring, the headspace of the reactor was purged with N2 to displace all the air present. Upon hitting the 30-minute mark, the reactor w'as sealed, and die reactor heaters were set to 230 °C.

After having reached an internal temperature of 220 °C and pressure of 390 psig, the hydrothermal reaction was allowed to proceed for 24 hours. The reactor was then cooled to 50 °C and vented. The contents of the reactor were filtered, and the filtrate was rinsed with 140 L of distilled water. The wet catalyst cake was dried in mi oven at 90 °C for 48 horns.

[3165] The catalyst was loaded in a tube furnace and heated at 285 °C for 26 hours under a low' flow of air (500 seem). After the air treatment, the catalyst was calcined at 400 °C for 3 hours under a flow of N ₂ (800-1000 seem) in the same tube furnace. After the N ₂ calcination, the catalyst was treated a second time at 350 °C for 3 hours in air.

[3166] Synthesis of Catalyst Material Ag2.1

[3167] To a 100 inL beaker was charged 2.3684 g of Catalyst Agl.2, 12.0920 g of VERS AL™ 250 Alumina, 0.6472 g of beryllium oxide and 33 ml, of distilled wnter. The beaker was clamped into an oil bath and an overhead agitator w'as set up with a glass stir rod and a 0.5-inch Teflon stir blade. The aqueous mixture was stirred for 3 hours and 15 minutes in a 100°C oil bath with the overhead agitator stirring at 100 rpm. The resulting paste was dried at 90°C for 18 hours. Subsequently, the dried powder was then muffle furnace treated at 350°C for 2 hours with a ramp time of 30 minutes yielding 18.2475 g of final catalyst material.

[3168J Synthesis of Catalyst Material Ag2.2

[3169] To a 400-mL beaker was charged 25.7860 g of Catalyst Agl.2, 42.0149 g of VERSAL™ -250 Alumina, 2.2420 g of Beryllium Oxide and 150 mL of distilled water. These additions formed a light purple aqueous mixture. The beaker was clamped into an oil bath and an overhead agitator was assembled using a ½” Teflon stir blade and a glass stir shaft. The oil bath was heated to 100°C and the overhead agitator was set to 99 rprn. The aqueous mixture was heated and stirred for 5 hours and 30 minutes forming a light purple paste. The paste was dried in an oven at 90°C for 18 hours forming a light purple powder. The light purple powder was transferred into two smaller beakers. The catal st powder in these beakers were muffle furnace calcined at 350°C for 2 hours with a ramp time of 30 minutes yielding 59.36 g. A portion of Catalyst Material Ag2.2 was pressed.

[3170] Synthesis of Catalyst Material Ag2.2.1

[3171] To a 250 mL beaker was charged 20.0837 g of Catalyst Material Ag2.2, 0.4016 g of calcium carbonate and 40 ml of distilled water forming a black/purple aqueous mixture. The beaker was clamped into an oil bath, the oil bath was heated to 100°C. An overhead agitator assembly was assembled using a glass stir rod and a 0.5-inch Teflon stir blade. The slurry was stirred and heated at 100°C for 1 hour. The resulting paste was dried in an oven for 18 hours. The resulting powder was further dried at 200°C for 2 hours. The resulting powder was pelletized on the auto-press forming 9.61 g of pellets and 8.51 g of powder.

[3172] Synthesis of Catalyst Material Ag2.3

[3173] Catalyst Agl .2 in the amount of 27.6320 g was charged a 600 ml beaker, followed by addition of 92.5221 g of VERSAL™ 250 .Alumina, 2.4056 g ofberyllium oxide and 250 ml of distilled water. The beaker was clamped into an oil bath and an overhead agitator was assembled using a glass stir rod and a 0.5-inch Teflon stir blade. The oil bath was heated to 1 Q0°C and the overhead agitator was set to 100 rpm. The mixture was left to stir for 2 hours and 20 minutes, after which it became a purple-black paste. The beaker containing the paste was heated in a 90°C oven for about 18 hours (overnight drying). Subsequently, the dried paste was calcined in a muffle furnace at 350°C for two hours with a 30-minute ramp time in air atmosphere with convective air exchange.

[3174] Synthesis of Catalyst Material Ag2.3.1

[3175] To a 250 mL beaker was charged 1.6984 g of calcium carbonate and 18.374 g of Catalyst Material Ag2.3. To the beaker was charged 57 mL of distilled water. The Catalyst Material Ag3.1 and calcium carbonate mixture was bubbled and the solution turned a yellow/green color. The beaker was clamped into an oil bath and an overhead agitator was assembled using a glass stir rod and a 0.5 -inch Teflon stir blade. The oil bath was healed to 100°C and the overhead agitator was set to 100 rpm. The mixture was left to stir for 2 hours, after winch it became a grey purple paste. The beaker containing the paste was heated in a 90 °C oven for 18 hours. Subsequently, the dried paste was ground, and muffle furnace calcined at 350°C for two hours with a 30-minute ramp time. The resulting grey powder yielding 15.2 g was ground and sieved to 500 mhi and pelletized using an auto-press. [3176] Synthesis of Catalyst Material Ag2.4

[3177] To a 250-mL beaker was charged 10.0263 g of Catalyst Agl.2, 0.8056 g of beryllium oxide, 43.2188 g of VERSAL™ 250 Alumina, and 150 mL of distilled water. The beaker was clamped in a 100°C oil bath and an overhead agitator assembly was setup with a 0.5-inch Teflon stir blade and a glass stir rod. The agitator speed was set to 100 ipni and the suspension was allowed to stir for 3 hours to form a thick paste. The mixture was removed from the oil bath and dried in an oven at 90°C over the -weekend. The beaker was then placed in a muffle furnace and calcined at 350°C for 2 hours, with a 30-minute ramp time.

[3178] Synthesis of Catalyst Material Ag2.5

[3179] To a 250-mL beaker was charged 10.0807 g of Catalyst Agl.2, 0.8097 g of beryllium oxide, 42.2380 g of VERSAL™ 250 Alumina and 150 mL of distilled water. The beaker was clamped into an oil bath and an overhead agitator was assembled using a 0.5-inch Teflon stir blade and a glass stir shaft. The oil bath was heated to 100°C and Site overhead agitator w'as set to 100 rpm. The purple / grey aqueous mixture w'as heated and stirred for 3 hours. The resulting grey /purple paste was dried in a 90°C oven for 18 hours. Subsequently, the dried light purple powder was calcined in a muffle furnace at 350°C for 2 hours with a 30-minute ramp time yielding 48.6073 g of light purple pow'der.

[3180] Synthesis of Catalyst Material Ag2.5.1

[3181] To Catalyst Material Ag2.5 w'as charged 0.9920 g of CaCCh and 153 mL of distilled water. The beaker was clamped into an oil bath and an overhead agitator was assembled using a 0.5-inch Teflon stir blade and glass stir shaft. The oil bath was heated to 100°C and the overhead agitator was set to 98 rpm. The aqueous mixture was heated and stirred for 5 hours and 30 minutes forming a paste. The light purple paste was dried in a 90°C oven for 18 hours yielding 45.5219 g of light purple powder. This material was then treated at 380°C.

[3182] Synthesis of Catalyst Material Ag2.6

[3183] To a 50-mL beaker was charged 20543 g of Catalyst Agl 2, 0 1602 g of beryllium oxide, 8.6317 g of VERSAL™ 250 Alumina and 20 mL of distilled water. The contents were manually mixed to form a uniform paste. The beaker was placed in an oven at 90°C overnight. The beaker was then transferred to a muffle furnace and calcined at 350°C for 2 hours with a 30 minute ramp time. The resulting purple catalyst material powder was ground and yielded 9.0129 g.

[3184] Synthesis of Catal st Agl .3

[3185] Three separate vessels were filled with distilled water and heated to tire desired temperature. The starting materials were dissolved in each vessel while stirring. Table Agl6 shows the conditions of the starting chemicals preparation: [3186] Table Agl6

[3187] The solutions were each stirred for about iO minutes until homogeneous solution were obtained. The total water used for the reaction, including rinsing the vessels, was 57 L. The solution of ammonium molybdate tetrahydrate from Vessel- Agl was pumped into the 100-L reactor vessel at the pump rate of 3.2 L/min. Once transferred, the reactor was stirred and the solution of vanadyl sulfate hydrate from Vesse3~Ag2 was added to the 100-L reactor vessel at the pump rate of 3.2 L/min. Lastly, the sodium dodeeyi sulfate solution in Vessei-Ag3 was pumped into the 100-L reactor vessel at the rate of 3.2 L/min. The reaction mixture was allowed to stir in the reactor for 30 minutes. While the reaction was stirring, the headspace of the reactor was purged with N2 to displace all the air present. Upon hitting the 30-minute mark, the reactor was sealed, and the reactor heaters were set to 230 °C.

After having reached an internal temperature of 220 °C and pressure of 390 psig, the hydrothermal reaction was allowed to proceed for 24 hours. The reactor was then cooled to 50 °C and vented. The contents of the reactor were filtered, and the filtrate was rinsed with 140 L of distilled water. The wet catalyst cake was dried in an oven at 90 °C for 48 horns.

[3188] The catalyst was loaded in a tube furnace and heated at 285 °C for 26 hours under a low flow of air (500 seem). After the air treatment, the catalyst was calcined at 400 °C for 3 hours under a flow of N ₂ (800-1000 seem) in the same tube furnace. After the N ₂ calcination die catal st was treated a second time at 350 °C for 3 hours in air.

[3189] Synthesis of Catalyst Material Ag3.1

[3190] To a 400-rnL beaker was loaded 49.97 g of Catalyst Agl.3 and 4.35 g of beryllium oxide. The mixture was stirred manually with a stir stick and then about 120 ml, of dH ₂ 0 was added. The beaker was placed in an oil bath at 100 °C and stirred at 100 rpm with an overhead stirrer. After about 2 hours, the mixture had formed a thick paste. The beaker was transferred to an oven at 90 °C and left overnight, yielding 56.99 g of catalyst material powder. The beaker was then placed in a muffle furnace at 350 °C for 2 hours (in addition to a 30-minute ramp to 350 °C) and left to cool overnight. The calcined catalyst was dark purplish-grey powder color and 54.53 g of material was recovered. The powder was submitted to MRU for testing. [3191] Synthesis of Catalyst Material Ag3.1.1

[3192] To a lOO-mL beaker was loaded 4.00 g of Catalyst Material Ag3.1 (dark purplish-grey powder) and 6.00 g ofVERSAL™ 250 (white powder). The mixture was stirred manually with a stir stick and then about 34 mL of dH ₂ 0 was added. The beaker was placed in an oil ba th at 100 °C and stirred at 100 rprn with an overhead stirrer. After about 1 hour, the mixture had formed a thick paste. The beaker was transferred to an oven at 90 °C and left overnight, yielding 10.16 g of catalyst material powder. The beaker was then placed in a muffle furnace at 350 °C for 2 hours with a 30-minute rarnp time and left to cool overnight, yielding 9.85 g of light grey catalyst material powder. The calcined material was ground using a mortar and pestle. The powder was submitted to MRU for testing. [31931 Synthesis of Catalyst Material Ag3.2

[3194] To a 100-mL beaker was loaded 8.00 g of baseline material Catalyst Material Ag3.1 (dark purplish- grey powder) and 6.00 g ofVERSAL™ 250 alumina (white powder). The mixture was stirred manually with a stir stick and then about 31 mL of dtftO was added. The beaker was placed in an oil bath at 100 °C and stirred at 100 rpm with an overhead stirrer. After about 1 hour, the mixture had formed a thick paste. The beaker was transferred to an oven at 90 °C and left overnight, yielding 13.32 g of light grey powder. The powder was placed in a muffle furnace at 350 °C for 2 hours (in addition to a 30-minute ramp to 350 °C) and left to cool overnight. This yielded 13.08 g of light grey material which was roughly crushed with a spatula.

[3195] Synthesis of Catalyst Material Ag3.2.1

[3196] To a 100-mL beaker containing Catalyst Material Ag3.2 was loaded 5.61 g of calcium carbonate and 32.3 ml, of distilled water and manually stirred with a stir stick. The beaker was placed in an oil bath at 100 °C and stirred at 100 rpm with an overhead stirrer. After about 1 hour, the mixture had formed a thick paste. The beaker was transferred to an oven at 90 °C and left overnight, yielding 16.53 g of grey/beige catalyst material. The material was ground using a mortar and pestle. The powder was submitted to MRU for testing.

[3197] Synthesis of Catalyst Material Ag3.2.2

[3198] To a 100 mL beaker was charged 8.04 g of Catalyst Material Ag3.2.1. This beaker was then placed in a muffle furnace and calcined at 350 °C for 2 hours with a 30 minute ramp time and left to cool overnight. This yielded beige powder. The powder was submitted to MRU for testing.

[3199] S nthesis of Catalyst Material Ag3.3.1

[3200] To a 250-mL beaker was loaded 8.00 g of Catalyst Material Ag3.1, dark purplish-grey powder), 12.00 g of calcium carbonate (white powder). The mixture was stirred manually with a stir stick and then about 75 mL of dH ₂ 0 was added. The beaker was placed in an oil bath at 100°C and stirred at 100 rpm with an overhead stirrer. After about 2 hours, the mixture had formed a thick paste. The beaker was transferred to an oven at 90°C and left overnight, yielding 18.25 g of catalyst material powder. The powder (grey /beige) was ground using a mortar and pestle. The powder was submitted to MRU for testing.

[3201] Synthesis of Catalyst Material Ag3.3.2

[3202] To another 100 mL beaker was charged 9.22 g of Catal st Material Ag3.3.1, was then placed in a Lindberg Blue M programmable muffle furnace at 350°C for 2 hours (in addition to a 30-minute ramp to 350°C) and left to cool overnight. This yielded 8.92 g of beige pow'der. The powder was submitted to MRU for testing. [3203] Synthesis of Catalyst Material Ag3.4.1

[3204] To a 100-mL beaker was loaded 8.00 g of Catalyst Material Ag3.1 (dark purplish-grey powder) and 20.01 g of calcium carbonate (white powder). The mixture was stirred manually with a stir stick and then about 71 rnL of dlCO was added. The beaker was placed in an oil bath at 100 °C and stirred at 100 rpm with an overhead stirrer. After about 1 hour, the mixture had formed a thick paste. The beaker was transferred to an oven at 90 °C and left overnight, yielding 25.72 g of grey / beige catalyst material. The powder was ground using a mortar and pestle. The powder was submitted to MRU for testing.

[3205] Synthesis of Catalyst Material Ag3.4.2

[3206] The remaining powder was placed in a muffle furnace at 350 °C for 2 hours with a 30-minute ramp time and left to cool overnight. This yielded 13.14 g of beige material.

[3207] Synthesis of Catalyst Ag 1.4

[3208] Catalyst Agl.2 and Catalyst Agl.3 were combined.

[3209] Synthesis of Catalyst Material Ag4.1

[3210] In a Pyrex dish was mixed 325.00 g of Catalyst Agl.2, 28.2641 g of beryllium oxide, 1413.0 g of VERSAL™ 250 alumina and 3600 ml of dlTO. The mixture was manually mixed with large stainless-steel sewing utensils. In a second large glass Pyrex: dish was mixed 325.00 g of catalyst active phase, 28.2600 g of BeO, 1413.0 g of VERSAL™ 250 alumina and 3600 mL of dHiO. The mixture w'as manually mixed, and then transferred into the larger Pyrex dish. The combined mixture was further manually mixed, until it appeared to be homogeneous (approximately 20 minutes of mixing). It was noted that the paste was a little wet, and that 250 mL less could be used for the next batch. The large Pyrex dish was loaded into the oven at 90°C overnight. The resulting po wder was transferred into 3-L beakers (two beakers at time) and loaded into a Lindberg Blue M programmable muffle furnace and calcined at 350°C for 2 hours (in addition to a 30-minute ramp to 350°C), before being left to cool overnight. The powder was then ground using a Retsch ^® BB50 jaw crasher.

[3211] In a Pyrex dish was mixed 324.872 g of catalyst active phase Catalyst Agl.4, 28.262 g of BeO, 1413.0 g of VERSAL™ 250 alumina and 3500 mL of dH> _; 0. The mixture was manually mixed with large stainless-steel sewing utensils. In a second large glass Pyrex dish was mixed 325.075 g of catalyst active phase, 28 262 g of BeO, 1408.0 g of VERSAL™ 250 alumina and 3500 rnL of dPLO. The mixture was manually mixed, and then transferred into the larger Pyrex dish. The combined mixture was further manually mixed, until it appeared to be homogeneous (approximately 20 minutes of mixing). The large Pyrex dish was loaded into the oven at 90°C overnight. The resulting pow'der was transferred into 3-L beakers (two beakers at time) and loaded into a Lindberg Blue M programmable muffle furnace and calcined at 350°C for 2 hours (in addition to a 30-minute ramp to 350°C), before being left to cool overnight. The powder was then ground using a Retsch ^® BB50 jaw crusher.

[3212] In a Pyrex dish was mixed 324.998 g of catal st active phase Catalyst Agl.3, 28.2638 g of BeO, VERSAL™ 250 alumina (mass was not recorded, but assumed to be 1413.0) and 3500 mL of dH ₂ 0. The mixture was manually mixed with large stainless-steel serving utensils. In a second large glass Pyrex dish was mixed 298.117 g of catalyst active phase, 25.9232 g of BeO, 1296.0 g of Versa!™ 250 alumina and 3500 mL of dH ₂ 0. The mixture was manually mixed, and then transferred into the larger Pyrex dish. The combined mixture was further manually mixed until it appeared to be homogeneous (approximately 20 minutes of mixing). It was noted that the paste was a little wet, and that 250 mL less could be used for the next batch. The large Pyrex dish was loaded into the oven at 90°C overnight. The resulting powder was transferred into 3-L beakers and loaded into a Lindberg Blue M programmable muffle furnace and calcined at 350°C for 2 hours (in addition to a 30-minute ramp to 350°C), before being left to cool overnight. The powder was then ground using a Retsch ^® BB50 jaw crusher.

[3213J All catalyst materials from these three batches were combined for Catalyst Material Ag4.1.

[3214] Synthesis of Catalyst Material Ag4.1.1 i3215| FeMo Cat Ltd. received 10607 g of Catalyst Material Ag4.1. To this was added 530 g of calcium carbonate via dry mixing. A total 11137 g of catalyst material was fed into the RTF 41 press with 9898 g of pellets being recovered. Of the 1239 g lost, there was 185 g of recoverable run material and 550 g of dust; oilier losses totaled 504 g. The tableting yield was therefore 88.9%, though the process potential yield was 95.4%. Calcination in an air atmosphere of the resulting catalyst material pellets led to a 14.6 wt.% mass loss, and thus the final pellet mass was 8456 g. The calcination furnace was ramped to 200°C over 2 hours and dwelled at 200°C for 2 hours. It was then further ramped to 350°C over 2 hours and dwelled at 350°C for 2 hours, before cooling back down to 50°C over 2 hours.

[3216] A number of impleme tations have been described. Nevertheless, it will be understood that various modifications may be made without departing from the spirit and scope of the disclosure.