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Title:
PRODUCTION OF AMMONIA FROM HYDROCARBONACEOUS FEEDSTOCK
Document Type and Number:
WIPO Patent Application WO/1990/006281
Kind Code:
A1
Abstract:
Ammonia is produced from hydrocarbonaceous feedstock by catalytic partial oxidation of the feedstock under temperature and steam conditions producing synthesis gas without free carbon; conversion of carbon monoxide in the synthesis gas to carbon dioxide by a water shift gas reaction; removal of carbon dioxide; adjustment of the hydrogen-to-nitrogen molar ratio to between 2:1 and 4:1; and reaction of hydrogen and nitrogen under ammonia-producing conditions.

Inventors:
Korchnak, Joseph D.
Dunster, Michael English Alan
Application Number:
PCT/US1989/005371
Publication Date:
June 14, 1990
Filing Date:
November 30, 1989
Export Citation:
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Assignee:
DAVY MCKEE CORPORATION.
International Classes:
B01J8/02; C01B3/02; C01B3/38; C01C1/04; (IPC1-7): C01B3/02; C01B3/38; C01C1/04
Domestic Patent References:
WO1986000286A11986-01-16
Foreign References:
EP0303438A21989-02-15
EP0130846A21985-01-09
EP0178833A21986-04-23
Other References:
CHEMICAL ABSTRACTS, vol. 104, no. 10, 10 March 1986 Columbus, Ohio, USA page 142; left-hand column; ref. no. 71253M &JP-A-85161303(Osaka Gas)23.08.1985 see abstract
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Claims:
CLAIMS
1. A process for producing ammonia from hydrocarbonaceous feedstock which comprises: thoroughly mixing the hydrocarbonaceous feedstock with an oxygen containing gas at an oxygentocarbon molar ratio in the range from 0.3:1 to 0.8:1 and with water vapor at a steamtocarbon molar ratio in the range from 0:1 to 3.0:1; catalytically partially oxidizing the mixture at a temperature equal to or greater than a minimum noncarbonizing temperature selected in the range from 870°C to 1030°C as a linear function of the steamtocarbon molar ratio being equal to or greater than a ratio in a corresponding range from 0.4:1 to 0:1 and at a space velocity in the range from 20,000 hour"1 to 50,000 hour"1, thereby producing a synthesis gas comprising hydrogen, carbon dioxide and carbon monoxide; converting carbon monoxide in the synthesis gas to carbon dioxide by subjecting the synthesis gas to a water gas shift reaction; removing carbon dioxide from the synthesis gas; adjusting the nitrogen content of the gas to obtain a hydrogento nitrogen molar ratio from about 2:1 to 4:1; and reacting the hydrogen and nitrogen in an ammonia synthesis process to produce ammonia. 2.
2. A process as claimed in claim 1, wherein the oxidant used for catalytic partial oxidation contains in excess of 70 mole percent of oxygen.
3. A process as claimed in claim 2, wherein the oxidant used for catalytic partial oxidation contains in excess of 90 mole percent of oxygen.
4. A process as claimed in claim 1, wherein the water gas shift reaction is carried out in a tubular reactor in which heat of reaction is recovered by generating steam. UBSTITUTESHEET .
5. A process as claimed in claim 1, wherein the carbon dioxide is removed from the synthesis gas by contacting the gas steam with a counter current liquid stream of carbon dioxide absorbing medium.
6. A process as claimed in claim 5, wherein pressure swing adsorption is used, following carbon dioxide removal, to remove components of the gas stream in order to produce a purified hydrogen steam; and nitrogen is added to the gas stream prior to ammonia synthesis to obtain the desired hydrogentonitrogen molar ratio.
7. A process as claimed in claim 6, wherein beds used in the pressure swing adsorption are desorbed with nitrogen.
8. A process as claimed in claim 6 whereby a portion of tail gas from the pressure swing adsorption is recycled to the catalytic partial oxidation step.
9. A process as claimed in claim 1, wherein carbon dioxide in removed together with other impurities by pressure swing adsorption to produce a purified hydrogen stream; and nitrogen is added to the gas stream prior to ammonia synthesis to obtain the desired hydrogentonitrogen molar ratio.
10. A process as claimed in claim 9, wherein beds used in the pressure swing adsorption are desorbed with nitrogen.
11. A process as claimed in claim 9, wherein tail gas produced by pressure swing adsorption is catalytically combusted, whereby carbon dioxide is produced as a combustion product; and the carbon dioxide is recovered.
12. A process as claimed in claim 6, wherein purge gas from the ammonia synthesis process is recycled to the process stream at a point upstream of the pressure swing adsorption step.
13. A process as claimed in claim 9, wherein purge gas from the ammonia synthesis process in recycled to the process stream at a point upstream of the pressure swing adsorption step.
14. A process for producing ammonia from hydrocarbonaceous feedstock which comprises: (a) introducing to a catalytic partial oxidation zone an essentially completely mixed gaseous mixture of a hydrocarbonaceous feedstock, oxygen or a gas containing at least 70 mole percent of oxygen and, optionally, steam in which the steamtocarbon molar ratio is from 0:1 to 3.0:1 and the oxygentocarbon molar ratio is from 0.3:1 to 0.8:1, said mixture being introduced to the catalytic partial oxidation zone at a temperature not lower than 200°F (93°C) below its catalytic autoignition temperature; (b) partially oxidizing the hydrocarbonaceous feedstock in the catalytic partial oxidation zone to produce a gas consisting essentially of methane, carbon oxides, hydrogen and steam by passing the mixture through a catalyst capable of catalyzing the oxidation of the hydrocarbons, said catalyst having a ratio of geometric surface area to volume of at least 5 cm2/cm3 and a total volume corresponding to a space velocity of between 20,000 hr"1 and 500,000 hr"1, thereby producing synthesis gas containing hydrogen, carbon monoxide and carbon dioxide; (c) contacting the synthesis gas with a shift reaction catalyst under water shift gas reaction conditions which cause carbon monoxide to be converted to carbon dioxide; (d) removing carbon dioxide from the gas stream; (e) adjusting the nitrogen content of the gas stream to obtain a hydrogentonitrogen molar ratio between about 2:1 and 4:1; and (f) reacting the hydrogen and nitrogen in the gas stream under ammoniaproducing conditions.
15. A process as claimed in claim 14, wherein the molar ratio of hydrogentonitrogen is adjusted to between about 2.5:1 and 3.5:1 prior to reacting the hydrogen and nitrogen under ammonia producing conditions.
16. A process for producing ammonia from a hydrocarbon gas containing principally methane, the process comprising: (a) mixing the hydrocarbon gas with steam and an oxygen containing gas at a steamtocarbon molar ratio in the range from 0:1 to 3.0:1 and at an oxygengas mole to carbon atom ratio in the range from 0.3:1 to 0.8:1 under conditions providing thorough even mixing without combustion; (b) partially oxidizing the hydrocarbonsteam oxygen gas mixture in a catalytic partial oxidation zone having a catalyst capable of promoting the oxidation of methane, said catalyst having a geometric surface area to volume of at least 5 cm2/cm3, said catalyst having a volume such as to produce a space velocity in the range from 20,000 hr"1 to 500,000 hr"1 to generate synthesis gas containing hydrogen, carbon monoxide and carbon dioxide; (c) contacting the synthesis gas with a shift reaction catalyst under water shift gas reaction conditions to convert carbon monoxide and water to carbon dioxide and hydrogen; (d) removing carbon dioxide from an output stream of the water shift reaction; and (e) reacting hydrogen in an output stream from the carbon dioxide removal step with nitrogen under ammoniaproducing conditions to produce ammonia.
17. A process as claimed in claim 16 wherein the partial oxidizing is performed at a temperature equal to or greater than a minimum temperature selected in the range from 870°C to 1030°C as a linear function of the steamtocarbon molar ratio being equal to or greater than a ratio in a corresponding range from 0.4:1 to 0:1.
Description:
PRODUCTION OF AMMONIA FROM HYDROCARBONACEOUS FEEDSTOCK

Field of the Invention The present invention relates to the production of ammonia from hydrocarbonaceous feedstocks by a process which includes the partial oxidation of a feedstock to produce a hydrogen-rich synthesis gas, which is further processed and fed into an ammonia-synthesis loop.

Description of the Prior Art

Ammonia has been produced by reacting hydrogen and nitrogen under ammonia-producing conditions in a so-called ammonia-synthesis loop according to the equation: N 2 + 3H 2 < > 2NH 3

(1)

Hydrocarbonaceous feedstocks, such as natural gases recovered from sites near petroleum deposits, are convenient sources of hydrogen for use in ammonia synthesis. Typically, natural gases contain, as their principle constituent, methane, with minor amounts of ethane, propane and butane. Also included in the conversion in some instances, may be low-boiling liquid hydrocarbons. In order to convert a hydrocarbonaceous feedstock into a hydrogen-containing stream suitable for introduction to an ammonia synthesis loop, the

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feedstock is first converted into a synthesis gas containing a major amount of hydrogen, together with minor amounts of carbon monoxide, carbon dioxide and methane. Steam reforming process exemplified for methane by the endothermic equation: CH 4 + H 2 0 > CO + 3H 2

(2) and partial oxidation process exemplified for methane by the exothermic equation CH 4 + 0 2 > CO + 2H 2

(3) or a combination of steam reforming and oxidation have been employed to produce synthesis gas. The steam reforming process equation shows production of an additional mole of hydrogen compared to the partial oxidation process equation. The synthesis gas produced by steam reforming and/or partial oxidation is treated, such as by a water shift reaction CO + H 2 0 > C0 2 + H 2 (4) and a carbon oxide removal process, to produce a hydrogen feed to the ammonia synthesis loop.

The most commonly employed method for converting hydrocarbonaceous feedstocks to synthesis gas includes catalytic steam reforming, for example the above equation (2). In this process, the hydrocarbonaceous feedstock is reacted with steam in the presence of a catalyst, usually a nickel-containing catalyst, at a temperature between about 1200°F (650°C) and 1900°F (1040°C). This reaction has been performed in catalyst filled tubes within a furnace. The hydrocarbons react with steam under these conditions to produce carbon monoxide and hydrogen. Catalytic steam reforming is an expensive process to carry out. Not

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only is the nickel-containing catalyst very expensive, but also, the reactions are highly endothermic. Consequently, a great deal of energy must be provided to drive the reaction. Frequently, air is provided to the reforming reaction in order to provide energy through partial oxidation of hydrocarbons and thus reduce the external heat requirements. Air reforming can also be performed as a secondary reforming step to reduce unreacted methane (methane slippage) to less than one percent on a volumetric basis. Upon exiting the primary steam reformer, the unreacted methane is converted in the secondary steam reformer by the injection of air, whereby the heat of reaction is supplied by the combustion of methane, hydrogen and carbon monoxide. The quantity of internal combustion air injection is such that it will supply the nitrogen requirements of the ammonia synthesis loop. Carbon oxides, which are poisonous to the ammonia synthesis catalyst must be removed from the exit gas of the secondary reformer prior to entry into the ammonia synthesis loop. Synthesis gas production for ammonia synthesis may also be carried out autothermally in an autothermal reactor by adding an oxidant such as air to the steam and hydrocarbon mixture. The endothermic heat of reaction is supplied by the exothermic combustion reactions:

CH 4 + 20 2 > C0 2 + 2H 2 0

(5) H 2 + h 0 2 > H 2 0

(6) co + h o 2 —> co 2

(7)

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The autothermal reactor typically consists of two catalyst beds, the first bed providing a high outlet temperature sufficient for steam reforming in the second bed. Alternatively, the reactants can be partially reformed in a steam reforming furnace and enter the autothermal reactor at a temperature sufficiently high to ignite spontaneously with the entering oxygen, thus producing a higher temperature sufficient for reforming in the downstream catalyst bed. Following autothermal reforming, carbon oxides are removed so that the synthesis gas can be provided to the ammonia synthesis loop. Autothermal reforming generally takes place at relatively low throughputs. The process is carried out at space velocities of the order of 8,000 hr "1 to 12,000 hr "1 . "Space velocity" can be defined as the volumetric hourly rate of throughput per volume of catalyst. All figures quoted herein refer to the volumetric hourly rate at standard conditions of temperature and pressure. The foregoing procedures for producing synthesis gas suitable for use in ammonia production have the drawbacks of requiring expensive catalysts; large volumes of catalyst; relatively low rates of throughput; equipment that is expensive and, in some cases, takes up excessive amounts of space; and, in some cases, requires unacceptably large amounts of energy to drive the process.

Partial oxidation of hydrocarbonaceous feedstocks represents one alternative to steam reforming in the production of synthesis gas. Most of the partial oxidation processes that have been employed commercially are non-catalytic processes. Non-catalytic partial oxidation reactions, however, are relatively inefficent. They operate at high

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temperatures, i.e., in the range of 2,200°F (1200°C) to 2,800°F (1500°C) and require large amounts of oxygen. Typically, the oxygen-to-carbon ratio required in non-catalytic partial oxidation is greater than 0.8:1 and often greater than 1:1. Furthermore partial oxidation produces free carbon.

U.S. Patent 4,390,347 issued to Dille at al. describes a process for the production of synthesis gas by the non-catalytic partial oxidation of a liquid hydrocarbonaceous fuel. The hydrocarbonaceous feedstock is reacted with a free oxygen-containing gas in the presence of steam at an autogenously maintained temperature within the range of 1700°F (930°C) to 3000°F (1650°C) at a pressure in the range of about 1 to 23 atmospheres absolute (1 to 23 bar). The oxygen-to-carbon molar ratio is said to be from 0.7:1 to 1.5:1. Steam is mixed with the hydrocarbon stream to moderate the temperature. Generated carbon soot prevents damage to the refractory lining of the generator. A water quench and scrub removes the free carbon.

U.S. Patent 3,890,113 issued to Child et al, describes the production of a methane-rich steam in which non-catalytic partial oxidation of a hydrocarbonaceous feedstock is carried ou in the presence of steam and oxygen. The ratio of free oxygen in the oxidant to carbon in the feedstock is in the range of 0.8:1 to 1.5:1. Particulate carbon is removed from the effluent gas stream in a gas cleaning zone. The product synthesis gas is subjected to a water gas shift reaction to increase the amount of hydrogen in the gas.

U.S. Patent 3,927,998 issued to Child et al. , relates to the production of a methane rich stream by

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the partial oxidation of a hydrocarbonaceous fuel employing a steam to fuel weight ratio of 2.2:1 to 2.9:1 and an oxygen-to-carbon molar ratio of 0.8:1 to 0.84:1. The partial oxidation is carried out in the absence of catalysts. The synthesis gas is cooled and water, carbon dioxide, particulate carbon and other impurities are removed. The hydrogen and carbon monoxide in the gas are reacted in a catalytic methanation zone to produce a methane-rich stream. Conversion efficency of oxidation processes can generally be improved by the use of catalysts; but where the oxidation process in only partial, i.e. with insufficient oxygen to completely oxidize the hydrocarbon, then the catalyst is subject to carbon deposit and blockage.

However, U.S. Patent 4,087,259, issued to Fujitani et al., describes employment of a rhodium catalyst in a process wherein liquid hydrocarbonaceous feedstock is vaporized and then partially oxidized in contact with the rhodium catalyst at a temperature in the range of

690° to 900°C with optional steam added as a coolant at rate not more than 0.5 by volume relative to the volume of the liquid hydrocarbon in terms of the equivalent amount of water. The rhodium catalyst enables partial oxidation without causing deposition of carbon, but at temperatures greater than 900°C, thermal decomposition occurs producing ethylene or acetylene impurities. When steam is added, the quantity of hydrogen produced is increased while the yield of carbon monoxide remains constant due to catalytic decomposition of the steam to hydrogen gas and oxygen. A "LHSV" (Liquid Hourly Space Velocity) from 0.5 to 25 1/hour is disclosed; particularly, a high yield from partial oxidation of gasoline vapor, without steam, is produced at a

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temperature of 725°C and at a LHSV of 20, and with steam, is produced at temperatures of 700°C and 800°C and at a LHSV of 2.

In order to obtain acceptable levels of conversion using catalytic reforming processes of the prior art it has generally been necessary to use space velocities below about 12,000 hr "1 . For example, U.S. Patent No. 4,522,894, issued to Hwang et al., describes the production of a hydrogen-rich gas to be used as fuel for a fuel cell. The process reacts hydrocarbon feed with steam and an oxidant in an autothermal reformer using two catalyst zones. The total hourly space velocity is between 1,960 hr -1 and 18,000 hr "1 . Because the prior art processes must be carried out at low space velocity, catalytic reactors of the prior art have had to have large catalyst beds in order to achieve the throughputs desired in commercial operation. This increases the size and cost of the reactor. It is an object of the present invention to provide a process for the production of ammonia from hydrocarbonaceous feedstock which is energy efficient, is capable of using low cost catalysts and employs relatively small equipment volume to achieve commercially acceptable throughput.

It is a further object of the invention to provide a process for the production of ammonia from hydrocarbonaceous feedstock with a relatively low oxygen demand, thereby increasing throughput of hydrocarbonaceous feed.

These and other objects of the invention are achieved by a process which is described below.

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Summary of the Invention This invention provides a process for the production of ammonia in which a hydrogen-rich synthesis gas is generated by the catalytic partial oxidation of a hydrocarbonaceous feedstock, such as natural gas, with an oxidant stream under temperature and steam conditions producing essential no free carbon at a space velocity in the range from 20,000 hour "1 to 500,000 hour "1 ; treatment of the resultant synthesis gas to remove components other than hydrogen and nitrogen; adjustment of the nitrogen content of the hydrogen-containing stream; and reaction of the hydrogen and nitrogen to produce ammonia.

In one embodiment, the invention provides a process for producing ammonia from hydrocarbonaceous feedstock which comprises:

(a) introducing to a catalytic partial oxidation zone an essentially completely mixed gaseous mixture of a hydrocarbonaceous feedstock, oxygen or an oxygen-containing gas and, optionally, steam in which the steam-to-carbon molar ratio is from 0:1 to 3.0:1 and the oxygen-to-carbon molar ratio is from 0.3:1 to 0.8:1, said mixture being introduced to the catalytic partial oxidation zone at a temperature not lower than 200°F (93°C) below its catalytic autoignition temperature;

(b) partially oxidizing the hydrocarbonaceous feedstock in the catalytic partial oxidation zone at a temperature equal to or greater than a minimum non-carbonizing temperature selected in the range from 1600°F (870°C) to 1900°F (1030°C) as a linear function of the steam-to-

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carbon molar ratio being equal to or greater than a ratio in a corresponding range from 0.4:1 to 0:1 to produce a synthesis gas contain hydrogen, carbon monoxide and carbon dioxide by passing the mixture through a catalyst capable of catalyzing the partial oxidation of the hydrocarbons at a space velocity in a range from 20,000 hour "1 to 500,000 hour "1 , said catalyst having a ratio of geometric surface area to volume of at least 5 cmVcm 3 ;

(c) contacting the synthesis gas with a shift catalyst under water gas shift reaction conditions which cause carbon monoxide to be converted to carbon dioxide and hydrogen; (d) removing carbon dioxide from the gas stream; (e) adjusting the nitrogen content of the carbon dioxide free gas stream to obtain a hydrogen-to-nitrogen molar ratio between about 2:1 and 4:1; and (f) reacting the hydrogen and nitrogen in the adjusted gas stream under ammonia-producing conditions to produce ammonia.

Brief Description of the Drawings Fig. 1 is an elevated cross-section view of a partial oxidation reactor having at its input a mixer and distributor suitable for introducing the reactants to the catalyst bed for use in the process of the invention.

Fig. 2 is an enlarged elevational cross-section view of a broken-away portion of the mixer and distributor of Fig. 1.

Fig. 3 is a top view of a broken-away quarter section of the mixer and distributor of Fig. 1.

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Fig. 4 is a bottom view of a broken-away quarter section of the mixer and distributor of Fig. 1.

Fig. 5 is a diagrammatic elevational cross-sectional illustration of a broken-away portion of the mixer and feeder of Figs. 1 and 2 showing critical dimensions.

Fig. 6 is a block flow diagram of one embodiment in accordance with the process of the invention for ammonia production using air as an oxidant in catalytic partial oxidation, and using cryogenic removal of nitrogen prior to the ammonia synthesis loop.

Fig. 7 is a block flow diagram of a modified embodiment similar to Fig. 6, but employing pressure swing adsorption to adjust nitrogen content prior to the ammonia synthesis loop.

Fig. 8 is block flow diagram of another modified embodiment similar to Fig. 6, but using oxygen-enriched air in the catalytic partial oxidation step without the adjustment of nitrogen content prior to the ammonia synthesis loop.

Fig. 9 is a block flow diagram of a still further modified embodiment of the process of the invention for producing ammonia which employs oxygen or oxygen-rich gas as the oxidant in the catalytic partial oxidation steps and is designed for low-capital cost.

Fig. 10 is a block flow diagram of yet another modified embodiment similar to Fig. 9 but is designed for low energy consumption. Fig. 11 is a graph plotting oxygen-to-carbon molar ratio vs. steam-to-carbon molar ratio for three different operating temperatures at an operating pressure of 400 psig. (2760 KPa).

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Fig. 12 is a graph plotting the hydrogen-to-carbon monoxide molar ratio in the catalytic partial oxidation reaction product vs. the steam-to-carbon molar ratio for three different operating temperatures at an operating pressure of 400 psig. (2760 KPa) .

Fig. 13 is a graph plotting the volume % methane in the catalytic partial oxidation product vs. the steam-to-carbon molar ratio for three different operating temperatures at an operating pressure of 400 psig. (2760 KPa) .

Fig. 14 is a graph plotting the volume % carbon dioxide in the catalytic partial oxidation product vs. steam-to-carbon molar ratio for three different operating temperatures at an operating pressure of 400 psig. (2760 KPa).

Fig. 15 is a graph plotting the molar ratio of total hydrogen and carbon monoxide in the product to total hydrogen and carbon in the feedstock vs. steam- to-carbon molar ratio for three different operating temperatures at an operating pressure of 400 psig. (2760 KPa) .

Fig. 16 is a detailed process flow diagram of a first portion of a process in accordance with the invention. Fig. 17 is a detailed process flow diagram of a second portion of the process of Fig. 16.

Detailed Description of the Preferred Embodiments The process of the present invention can be used to produce ammonia from any gaseous or low-boiling hydrocarbonaceous feedstock. Typically, a gaseous hydrocarbonaceous feedstock used to produce synthesis gas is a gas containing principally methane such as

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natural gas having the following approximate composition: methane, 93%; ethane, 5%; propane 1.5%; butane and higher hydrocarbons, 0.5%.

In general, the process of the invention involves the steps of catalytic partial oxidation of hydrocarbonaceous feedstock under temperature and water content conditions to produce synthesis gas without free carbon; treatment of the resultant synthesis gas to remove components other than hydrogen and nitrogen (e.g., carbon oxides) and to recover a carbon dioxide stream; adjustment of the nitrogen content of the hydrogen-containing stream; and reaction of the hydrogen and nitrogen to produce ammonia.

Catalytic Partial Oxidation Of Hydrocarbonaceous Feedstock

One particular aspect of the invention is the substantial capital cost savings and/or advantageous operating economy resulting from the employment of catalytic partial oxidation to produce the raw synthesis gas employed in the ammonia producing process. This is made possible by the discovery that catalytic partial oxidation performed at a temperature equal to or greater than a minimum non-carbon-forming temperature selected in the range from 1600°F (870°C) to 1900°F (1040°C) as a linear function of the steam- to-carbon molar ratio being equal to or greater than a ratio in a corresponding range from 0.4:1 to 0:1 and at a space velocity in the range from 20,000 hour "1 to 500,000 hour "1 produces essentially no free carbon deposits on the catalyst. Further, it is found that products of the partial catalytic oxidation in the process of the invention consist essentially of hydrogen, carbon monoxide and carbon dioxide at oxidation temperatures equal to or greater than the

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minimum temperature, rhodium catalysts are not required to prevent carbon formation. For example in Fig. 11, dotted line 25 represents a generally linear function which, at a steam/carbon ratio of 0, corresponds to a minimum partial oxidation temperature of about 1900°F (1040°C), and at a steam/carbon ratio of 0.4 corresponds to a minimum partial oxidation temperature of about 1600°F (870°C); favorable catalytic partial oxidation without producing free carbon occurs at temperatures and steam/carbon ratios equal to or greater than points on the line. Further, lower minimum temperatures at corresponding steam/carbon ratios greater than 0.4 can be extrapolated from the linear function represented by line 25. In the catalytic partial oxidation step of the process of the invention, reactant gases are introduced to the catalytic partial oxidation reaction zone, i.e. the catalyst bed, at an inlet temperature not lower than 200°F (93°C) below the catalytic autoignition temperature of the feed mixture.

Preferably the reactant gases are introduced at a temperature at or above the catalytic autoignition temperature of the mixture. The reactants should be completely mixed prior to the reaction. Introducing the thoroughly mixed reactant gases at the proper temperature ensures that the partial oxidation reactions will be mass transfer controlled. Consequently, the reaction rate is relatively independent of catalyst activity, but dependent on the surface-area-to-volume ratio of the catalyst. It is possible to use any of a wide variety of materials as a catalyst, provided that the catalyst has the desired surface-area-to-volume ratio. It is not necessary that the catalyst have specific catalytic activity for steam

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reforming. Even materials normally considered to be non-catalytic can promote the production of synthesis gas herein when used as a catalyst in the proper configuration. The term "catalyst", as used herein, is intended to encompass such materials.

The catalytic partial oxidation step can be understood with reference to the figures. The catalytic partial oxidation zone is typically the catalyst bed of a reactor such as that indicated generally in Fig. 1 at 28. The reactor 28 includes an input mixing and distributor section indicated generally at 30 which mixes the feedstock with an oxidant and distributes the mixture to the entrance of a catalytic reactor section indicated generally at 32. In the catalytic reactor section 32, the feedstock is partially oxidized to produce a product which is then passed through the exit section indicated generally at 34.

The reactor includes an outer shell 40 of structural metal such as carbon steel with a top 42 secured thereon by bolts (not shown) or the like. A layer 44 of insulation, such as 2300°F (1260°C) BPCF ceramic fiber insulation, is secured to the inside of the upper portion of the shell 40 including the top 42. In the lower portion of the mixing section 30, in the reactor section 32 and in the outlet section 34, there are secured insulating layers 46, 48 and 50 on the inside of the shell. The layer 46 is a castable or ecpiivalent insulation such as 2000°F (1090°C) ceramic insulation. The layer 48 is also a castable or equivalent layer of insulation but containing 60% alumina for withstanding 3000°F (1650° C). The internal layer 50 is a refractory or equivalent layer such as 97% alumina with ceramic anchors or 97% alumina

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brick for withstanding the interior environment of the reactor section.

The catalytic reactor section 32 contains one or more catalyst discs 54. As shown, the reactor contains a sequence of discs 54 separated by high alumina rings 58 between each adjacent pair of discs. The stack is supported by a grill with high alumina bars 56. A sample port 60 is formed in the lower end of the reaction section and has a tube, such as type 309 stainless steel tube 62, extending below the bottom refractory disc 54 for withdrawing samples of the product.

The outlet section 34 is suitably formed for being connected to a downstream heat recovery boiler (not shown) and/or other processing equipment.

The catalyst comprises a high surface area material capable of catalyzing the partial oxidation of the hydrocarbonaceous feedstock. The catalyst is in a configuration that provides a surface area to volume ratio of at least 5 cm 2 /cm 3 . Preferably, the catalyst has a geometric surface area to volume ratio of at least 20 cm 2 /cm 3 . While there is no strict upper limit of surface area to volume ratio, it normally does not exceed about 40 cm 2 /cm 3 . A wide variety of materials can be used in the construction of the catalyst including materials not normally considered to have catalytic activity, provided that the catalyst configuration has the desired surface area to volume ratio. The catalyst disc 54 can be, for example, a monolithic structure having a honeycomb type cross-sectional configuration. Suitable monolithic structures of this type are produced commercially, in sizes smaller than those used in the process of the

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invention, as structural substrates for use in the catalytic conversion of automobile exhausts and as catalytic combustion chambers of gas turbines or for catalytic oxidation of waste streams. Typically, the monolithic structure is an extruded material containing a plurality of closely packed channels running through the length of the structure to form a honeycomb structure. The channels are typically square and may be packed in a density as high as 1,200 per square inch of cross section. The monolithic structure can be constructed of any of a variety of materials, including cordierite (MgO/Al 2 0 3 /Si0 2 ) , Mn/MgO cordierite (Mn-MgO/Al 2 0 3 Si0 2 ), mullite (Al 2 0 3 /Si0 2 ), mullite aluminum titanate (Al 2 0 3 /Si0 2 -(Al,Fe) 2 0 3 /Ti0 2 ), zirconia spinel (Zr0 2 /MgO/Al 2 0 3 ), spinel (MgO/Al 2 0 3 ), alumina (Al 2 0 3 ) and high nickel alloys. The monolithic catalyst may consist solely of any of these structural materials, even though these materials are not normally considered to have catalytic activity by themselves. Using honeycombed substrates, surface area to volume ratios up to 40 cm 2 /cm 3 or higher can be obtained. Alternatively, the monolithic substrate can be coated with any of the metals or metal oxides known to have activity as oxidation catalysts. These include, for example, palladium, platinum, rhodium, irridium, osmium, ruthenium, nickel, chromium, cobalt, cerium, lanthanum and mixtures thereof. Other metals which can be used to coat the catalyst disc 54 include noble metals and metals of groups IA, IIA, III, IV, VB, VIB, or VIIB of the periodic table of elements.

The catalyst discs 54 may also consist of structural packing materials, such as that used in packing absorption columns. These packing materials generally comprise thin sheets of corrugated metal

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tightly packed together to form elongate channels running therethrough. The structural packing materials may consist of corrugated sheets of metals such as high temperature alloys, stainless steels, chromium, manganese, molybdenum and refractory materials. These materials can, if desired, be coated with metals or metal oxides known to have catalytic activity for the oxidation reaction, such as palladium, platinum, rhodium, irridium, osmium, ruthenium, nickel, chromium, cobalt, cerium, lanthanum and mixtures thereof.

The catalyst discs 54 can also consist of dense wire mesh, such as high temperature alloys or platinum mesh. If desired, the wire mesh can also be coated with a metal or metal oxide having catalytic activity for the oxidation reaction, including palladium, platinum, rhodium, irridium, osmium, ruthenium, nickel, chromium cobalt, cesium, lanthanum and mixtures thereof.

The surface area to volume ratio of any of the aforementioned catalyst configurations can be increased by coating the surfaces thereof with an aqueous slurry containing about 1% or less by weight of particulate metal or metal oxide such as alumina, or metals of groups IA, IIA, III, IV, VB, VIB and VIIB and firing the coated surface at high temperature to adhere the particulate metal to the surface, but not so high as to cause sintering of the surface. The particles employed should have a BET (Brunnauer-Emmett-Teller) surface area greater than about 10 m 2 /gram, preferably greater than about 200 m 2 /gram.

In the practice of the invention, a gaseous mixture of hydrocarbonaceous feedstock, oxygen or an oxygen-containing gas, which can be air, oxygen-enriched air, or other oxygen-rich gas, and,

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optionally, steam is introduced into the catalytic partial oxidation zone at a temperature not lower than 200°F (93°C) below its catalytic autoignition temperature. Preferably, the gaseous mixture enters the catalytic partial oxidation zone at a temperature equal to or greater than its catalytic autoignition temperature. It is possible to operate the reactor in a mass transfer controlled mode with the reactants entering the reaction zone at a temperature somewhat below the autoignition temperature since the heat of reaction will provide the necessary energy to raise the reactant temperature within the reaction zone. In such a case, however, it will generally be necessary to provide heat input at the entrance to the reaction zone, for example by a sparking device, or by preheating the contents of the reactor, including the catalyst, to a temperature in excess of the autoignition temperature prior to the introduction of the reactants to initiate the reaction. If the reactant temperature at the input to the reaction zone is lower than the autoiginition temperature by more than about 200° F (93°C), the reaction becomes unstable.

When the reactant mixture enters the catalytic partial oxidation zone at a temperature exceeding its autoignition temperature, it is necessary to introduce the mixture to the catalyst bed immediately after mixing; that is, the mixture of hydrocarbonaceous feedstock and oxidant should preferably be introduced to the catalyst bed before the autoignition delay time elapses. It is also essential that the gaseous reactants be thoroughly mixed. Failure to mix the reactants thoroughly reduces the quality of the product and can lead to overheating. A

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suitable apparatus for mixing and distributing the hydrocarbonaceous feedstock and oxygen or oxygen-containing gas so as to provide thorough mixing and to introduce the heated reactants into the reaction zone in a sufficiently short time is illustrated in Figs. 1-5 and described in more detail in copending commonly assigned patent Application Serial No. 085,159, filed August 14, 1987 in the names of J.D. Korchnak, M. Dunster and J.H. Marten. Referring to Fig. 1, one of the feed gases, i.e. hydrocarbonaceous gas or oxygen-containing gas, is introduced into the input section 30 through a first inlet port 66 through the top 42 which communicates to an upper feed cone 68 which forms a first chamber. The cone 68 is fastened by supports 69 in the top 42. The other feed gas is introduced into the input section 30 through second inlets 70 extending through side ports of the shell 40 and communicating to a second chamber

72 which is interposed between the upper chamber 68 and the inlet of the catalyst reaction section 32. A ring

73 mounted on the central portion of an upper wall 75 of the chamber 72 sealingly engages the lower edge of the cone 68 so that the wall 75 forms a common wall between the upper chamber 68 and lower chamber 72. The chamber 72 has an upper outer annular portion 74, see also Figs. 2 and 3, which is supported on the top surface of the refractory layer 50. A lower portion of the chamber 72 has a tubular wall 76 which extends downward in the refractory sleeve 50. The bottom of the chamber 76 is formed by a cast member 78.

Optionally, steam can be introduced into either or both of the hydrocarbonaceous feedstock and oxygen or oxygen-containing gas. The gases are fed to the reactor in relative proportions such that the

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steam-to-carbon molar ratio is from 0:1 to 3.0:1, preferably from 0.3:1 to 2.0:1. The oxygen-to-carbon ratio is from 0.4:1 to 0.8:1, preferably from 0.45:1 to 0.65:1. The reactant mixture preferably enters the catalytic reactor section 32 at a temperature at or above its autoignition temperature. Depending on the particular proportions of reactant gases, the reactor operating pressure and the catalyst used, this will generally be between about 550°F (290°C) and 1,100°F (590°C) Preferably, hydrocarbonaceous feedstock and steam are admixed and heated to a temperature from 650°F (340°C) to 1,200°F (650°C) prior to passage through inlet port(s) 70 or 66. Oxygen or oxygen-containing gas, such as air, is heated to a temperature from 150°F (65°C) to 1200°F (650°C) and passes through the other inlet port(s) 66 or 70.

Referring to Figs. 2, 3 and 4, the mixing and distributing means comprises a plurality of elongated tubes 80 having upper ends mounted in the upper wall 75 of the chamber 72. The lumens of the tubes at the upper end communicate with the upper chamber 68. The bottom ends of the tubes 80 are secured to the member 78 with the lumens of the tubes communicating with the upper ends of passageways 84 formed vertically through the member 78. Orifices 86 are formed in the walls of the tubes 80 for directing streams of gas from the chamber 72 into the lumens of the tubes 80. The inlets 66 and 70, the cone 68, the supports 69 are formed from a conventional corrosion and heat resistant metal while the chamber 72, tubes 80 and member 78 are formed from a conventional high temperature alloy or refractory type material.

SUBS TIT UTESHEET

The number of tubes 80, the internal diameter 90 (see Fig. 5) of the tubes 80, the size and number of the orifices 86 in each tube are selected relative to

*-. the gas input velocities and pressures through inlets 66 and 70 so as to produce turbulent flow within the tubes 80 at a velocity exceeding the flashback velocity of the mixture. The minimum distance 92 of the orifices 86 from the bottom end of the tube 80 at the opening into the diverging passageways 84 is selected to be equal to or greater than that required for providing substantially complete mixing of the gas streams from chambers 68 and 72 under the conditions of turbulence therein. The size of the internal diameter 90 of the tubes 80 as well as the length 94 of the tubes is designed to produce a sufficient pressure drop in the gas passing from the chamber 68 to the reaction chamber so as to provide for substantially uniform gas flow through the tubes 80 from the chamber 68. Likewise the size of the orifices 86 is selected to provide sufficient pressure drop between the chamber 72 and the interior of the tubes 80 relative to the velocity and pressures of the gas entering through inlets 70 so as to provide substantially uniform volumes of gas flows through the orifices 86 into the tubes 80.

The diverging passageways 84 in the member 78 are formed in a manner to provide for reduction of the velocity of the gas and to produce uniform gas distribution over the inlet of the catalyst. The rate of increase of the cross-section of the passageway 84 as it proceeds downward, i.e., the angle 98 that the wall of the passageway 84 makes with the straight wall of the tubes 80, must generally be equal to or less than about 15° and preferably equal to or less than 7°

in order to minimize or avoid creating vortices within the passageways 84. This assures that the essentially completely mixed gases, at a temperature near to or exceeding the autoignition temperature, will pass into the catalyst bed in a time preferably less than autoignition delay time. The configuration of the bottom end of the passageways, as shown in Fig. 4, is circular, but other configuration such as hexagonal, square, etc. are possible. The catalytic partial oxidation reaction is pre erably carried out in the catalytic reaction section 32 at a pressure greater than 100 psig (690 KPa), more preferably at a pressure greater than 250 psig (1720 KPa). The catalytic partial oxidation reaction is carried out at a temperature between about 1400°F (760°C) and 2000°F (1090°C).

The product gas exiting the outlet section 34 consists essentially of hydrogen, carbon oxides, i.e. carbon monoxide and carbon dioxide, methane, water vapor and any inert components (e.g. nitrogen or argon) introduced with the feedstock. Trace amounts of C 2 - and higher hydrocarbons may be present in the product gas. As used herein "trace amounts" means less than about 0.1% by weight.

Removal of Carbon Oxides

Since carbon oxides, i.e. carbon dioxide and carbon monoxide, are present in the synthesis gas and they are poisonous to the ammonia synthesis catalyst, they are removed from the synthesis gas prior to the ammonia synthesis loop.

The synthesis gas exiting the catalytic partial oxidation zone is cooled to a temperature from about 350°F (175°C) to about 750°F (400°C) using

conventional heat exchange methods, either by heating the hydrocarbon and steam feedstock, heating the oxidant stream, superheating steam, raising steam in a boiler, preheating boiler feedwater or a combination thereof.

The first step in the removal of carbon oxides is the conversion of carbon monoxide to carbon dioxide by the water gas shift reaction in which carbon monoxide is reacted with water to produce carbon dioxide and hydrogen. The water gas shift reaction is known, and suitable equipment for carrying out the reaction is commercially available. The water gas shift reaction can be carried out in two stages, i.e. a high temperature shift and a low temperature shift. In this procedure, the synthesis gas is first reacted with water vapor at a temperature from about 580°F (300°C) to 750°F (400°C) and a pressure from about 15 atm. (1520 KPa) to 40 atm. (4050 KPa), followed by reaction at a temperature from about 350°F (175°C) to 500°F (260°C) and a pressure from about 15 atm. (1520 KPa) to 40 atm. (4050 KPa). Alternatively the water gas shift reaction can be carried out in a single stage, low temperature tubular, steam-raising reactor shift vessel. In this procedure, the water vapor and synthesis gas are reacted at a temperature from about 350°F (175°C) to 500°F (260°C) and a pressure from about 15 atm. (1520 KPa) to 40 atm. (4050 KPa). The exit stream from the water gas shift reaction zones has a carbon monoxide content less than about 0.5 percent on a volumetric basis.

Essentially all of the remaining carbon monoxide can be converted to carbon dioxide by catalytic selective oxidation. In this procedure, the exit stream from the water gas shift reaction zone.

after heat removal to reduce its temperature to about 100°F (35°C) to 250°F (120°C), is reacted with air in the presence of a catalyst that is highly selective for the oxidation of carbon monoxide under conditions in which little or no hydrogen is oxidized. The catalytic selective oxidation procedure is known in the art and described by U.S. Patents No. 3,216,782, No. 3,216,783 and No. 3,631,073. Suitable process equipment for carrying out is commercially available for example, under the trademark Selectoxo.

Alternatively to the catalytic selective oxidation procedure, remaining carbon monoxide can be methanated using known procedures. However, since methanation reacts each mole of carbon monoxide with three moles of hydrogen, this procedure consumes hydrogen that would otherwise be useful in the ammonia synthesis loop. Furthermore, although methane is inert in the ammonia synthesis loop, increasing the amount of methane in the feed to the ammonia synthesis loop increases the loop purge requirements.

Any of the other procedures known in the art for removing carbon monoxide can be employed to remove traces of carbon monoxide from the gas stream.

After conversion of carbon monoxide to carbon dioxide, carbon dioxide is removed from the gas stream and recovered using known procedures such as, for example, passing the gas through a countercurrent stream of a liquid absorbent medium, such as potassium carbonate, which absorbs the carbon dioxide. Commercial processing units for carbon dioxide removal are available for example, under the trademarks Selexol, Amine Guard, and Benfield. These processes absorb the carbon dioxide into a chemical or physical absorption medium at relatively high pressure and low

SUBSTITUTE SHEET

temperature, allowing other gases to pass through essentially unchanged. The chemical or physical absorbent is then regenerated by pressure let down into

* -> a lower pressure vessel and, if a chemical absorbent is used, stripped of the carbon dioxide by a countercurrent stream of steam. The carbon dioxide gas is discharged from the top of the regenerator and the absorbent returned to the absorber to recover more carbon dioxide.

Adjustment of Nitrogen Content

Following removal of carbon dioxide from the gas stream, the nitrogen content of the gas stream is adjusted to provide a hydrogen to nitrogen ratio suitable for ammonia synthesis. Generally, the molar ratio of hydrogen-to-nitrogen is adjusted to between about 2:1 to 4:1 and preferably between from about 2.5:1 to 3.5:1 for ammonia synthesis. Any suitable means of adjusting the nitrogen content can be employed. When air or oxygen-enriched air is used as the oxidant in the catalytic partial oxidation step, the amount of nitrogen present in the synthesis gas exiting the catalytic partial oxidation zone is normally in molar excess to the amount required for ammonia synthesis and therefore, nitrogen must be removed from the gas stream. When oxygen or an oxygen-rich (>70 mole.%) gas is employed as the oxidant in the catalytic partial oxidation step, the synthesis gas exiting the catalytic partial oxidation zone normally requires the addition of nitrogen for ammonia synthesis.

One method for adjusting the nitrogen content of the gas stream is by cryogenic separation and

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removal of nitrogen. Cryogenic separation of gases is a known procedure whereby gases are fractionated according to their liquefaction temperatures. Commercially available cryogenic separators can be employed to remove nitrogen from the gas stream.

Alternatively, the nitrogen content of the gas stream can be adjusted by pressure swing adsorption. Pressure swing adsorption involves the adsorption of components to be removed at high pressure followed by their desorption at low pressure. The process operates on a repeated cycle having two basis steps, adsorption and regeneration. Not all the hydrogen is recovered as some is lost in the waste gas during the regeneration stage. By careful selection of the frequency and sequence of steps within the cycle, however, the recovery of hydrogen can be maximized and the ratio of hydrogen to nitrogen in the product effluent gas can be strictly controlled to give the desired ratio.

Regeneration of the adsorbent is carried out in three basic steps: (a) The adsorber is depressurized to the low pressure. Some of the waste components are desorbed during this step. (b) The adsorbent is purged at low pressure, with the product hydrogen removing the remaining waste components, (c) The adsorber is repressurized to adsorption pressure ready for service. The waste gases evolved during regeneration are collected in a waste gas surge drum and then used as fuel.

Pressure swing adsorption can also be used to remove carbon dioxide, methane, water vapor and other trace contaminants such as H 2 S. Accordingly, the pressure swing adsorption unit can serve both to remove carbon dioxide and to adjust the nitrogen content of the gas stream. Cryogenic separation can be employed

to remove methane, however, water vapor and final traces of carbon dioxide must still be removed by a separate procedure. A suitable method for removing water vapor and carbon dioxide prior to ammonia synthesis is by passing the gas stream over any of the commercially available molecular sieve materials.

Ammonia Synthesis As previously indicated, methane removal is optional because methane is inert in the ammonia synthesis loop. It is, however, preferred to remove methane from the gas stream. After components other than hydrogen and nitrogen have been removed and the hydrogen-to-nitrogen ratio has been adjusted, the gas stream is ready to enter the ammonia-synthesis loop. Any suitable procedure for reacting the hydrogen and nitrogen to obtain ammonia can be employed. Advantageously, the basic ammonia synthesis procedure employed is derived from the so-called "Haber-Bosch" process. In this process, the gas stream circulates under pressure in a loop wherein it is passed into a heated reaction chamber where it is reacted in contact with an ammonia synthesis catalyst. The gases containing the product ammonia then leave the reaction chamber and ammonia is recovered by condensation. The unreacted gases are recirculated by means of a compressor and are admixed with feed gas prior to re-entering the reaction chamber. The ammonia synthesis reaction is carried out at a temperature from about 650°F (340°C) to 770°F (410°C) and a pressure from about 80 atm (8100 KPa) to 150 atm (15200 KPa).

The nitrogen and hydrogen are reacted in the ammonia synthesis loop in contact with a conventional ammonia synthesis catalyst. Suitable catalysts

include, by way of example, singly or doubly promoted iron catalysts. Catalyst promoters include Al 2 0 3 , alone or in combination with K 2 0; Zr0 2 , alone or in combination with K 2 0; or Si0 2 , alone or in combination with K0 2 .

The various embodiments of the process of the invention can be further understood with reference to Figs. 6-10.

Figs. 6, 7 and 8 illustrate schematically three embodiments of the invention which employ air or oxygen enriched air to convert hydrocarbonaceous feedstock to ammonia.

According to the process of Fig. 6, hydrocar¬ bonaceous feedstock is first optionally treated in desulfurization step 100 to remove sulfur from the feedstock. Sulfur removal can be effected by any suitable means, such as by absorption on zinc oxide. Sulfur removal prior to catalytic partial oxidation is optional in each embodiment inasmuch as the catalytic partial oxidation process is sulfur tolerant, downstream steps, such as acid gas removal or pressure swing adsorption, can be used to remove gaseous sulfur compounds.

After desulfurization, the feedstock enters the catalytic partial oxidation zone together with air and steam. Catalytic partial oxidation takes place at step 102. The effluent gas, containing hydrogen, methane, carbon oxides and nitrogen, exits the catalytic partial oxidation zone at a temperature of about 1650°F (900°C) and is passed through heat exchanger(s) at step 104 to reduce its temperature to between 350°F (175°C) and 750°F (400°C) .

The gas is then passed to a water gas shift reactor where carbon monoxide is converted to carbon

SUBSTSTUTESHEET

dioxide at step 106 using the shift reaction previously described. The exit gas from the shift reaction zone is passed to a heat exchanger 108, where the gas temperature is reduced to between 100°F (38°C) and 250°F (120°C). The gas is then passed to a selective oxidation zone, where remaining carbon monoxide is converted to carbon dioxide at step 110 by the previously described selective oxidation process. The gas then undergoes removal of carbon dioxide by contacting the gas stream with a countercurrent flow of liquid which absorbs the carbon dioxide from the gas at step 112. The carbon dioxide can be recovered and sold as a valuable article of commerce. The exit gas from the carbon dioxide absorber contains trace amounts of water and carbon dioxide, which are removed at step 116 by passing the gas stream through a molecular sieve.

Following removal of carbon oxides, the nitrogen content of the gas is adjusted to a hydrogen-to-nitrogen molar ratio between about 2:1 to 4:1, preferably from about 2.5:1 to 3.5:1. Since the embodiment illustrated in Fig. 6 employs air as an oxidant in the catalytic partial oxidation step, nitrogen is removed from the gas stream to obtain the desired hydrogen-to-nitrogen ratio. The removal of nitrogen is achieved in a cryogenic separator at step 128. Essentially, all of the methane in the gas stream is also removed in this step.

The gas stream is then compressed (not shown) to between about 80 atm (8100 KPa) and 150 atm (15200 KPa) and enters the ammonia synthesis loop 126 where the hydrogen and nitrogen are reacted under ammonia-producing conditions in the presence of a catalyst to produce ammonia.

Fig. 7 schematically illustrates a variation of the process of Fig. 6 wherein, following the conversion of carbon monoxide to carbon dioxide at steps 106 and 110, the gas stream is subjected to pressure swing adsorption at step 118 to remove carbon dioxide, methane, water vapor and a portion of the nitrogen. The exit stream from the pressure swing adsorption unit is fed directly to the ammonia synthesis loop 126. Because pressure swing adsorption is capable of removing essentially all of the carbon dioxide and water vapor from the gas stream, it is unnecessary to employ a carbon dioxide removal step 112 or a molecular sieving step 116 in this embodiment of the invention. Fig. 8 schematically illustrates an embodiment of the invention which employs oxygen-enriched air as the oxidant in the catalytic partial oxidation step 102. By properly adjusting the ratio of oxygen to air in the oxidant feed, and thereby the amount of nitrogen entering the process, it is possible to eliminate the need for downstream adjustment of nitrogen. In order to obtain the desired hydrogen-to-nitrogen ratio for ammonia production, the air-to-oxygen molar ratio employed in the catalytic partial oxidation step of the process embodiment illustrated in Fig. 8 is between about 0.2 and 0.3. Since downstream nitrogen adjustment is not necessary in the embodiment of Fig. 8, it is preferred not to employ pressure swing adsorption, but rather, to remove carbon dioxide and water vapor by the same sequence of steps employed in the embodiment of Fig. 6, that is, by absorption of carbon dioxide in a countercurrent absorber at step 112, followed by molecular sieving at step 116 to remove traces of carbon dioxide and water vapor.

Figs. 9 and 10 schematically illustrate the two preferred embodiments of the invention that produce ammonia in processes that employ oxygen or an oxygen-rich gas as the oxidant in the catalytic partial oxidation of hydrocarbonaceous feedstock. These embodiments are economically advantageous because they reduce capital equipment costs and/or energy requirements. As used herein, the term "oxygen-rich gas" refers to a gas having an oxygen content of at least 70%, preferably at least 90%. Since the use of oxygen or oxygen-rich gas in the catalytic partial oxidation process eliminates or greatly reduces the nitrogen load at the front end of the process, the size of synthesis gas generating equipment and downstream conditioning equipment required is greatly reduced.

Furthermore, energy requirements to bring nitrogen to the catalytic partial oxidation reactor conditions is reduced or eliminated. Oxygen or oxygen-rich gas for use in the catalytic partial oxidation process can be generated using known techniques such as cryogenic fractionation of air.

Fig. 9 illustrates an embodiment of the invention that incurs relatively low capital equipment costs. In this embodiment, sulfur is optionally removed from the hydrocarbonaceous feedstock at step

100, after which the feedstock, together with steam and oxygen or oxygen-rich gas, is fed to the catalytic partial oxidation zone as previously described. Catalytic partial oxidation takes place at step 102, thereby producing a synthesis gas containing hydrogen, carbon monoxide, carbon dioxide, methane and little or no nitrogen, i.e. less' than 30% and preferably less than 10% nitrogen.

After heat exchange 104, shift gas reaction 106 and heat exchange 108 as previously described, the nitrogen content of the gas stream is adjusted to achieve a hydrogen-to nitrogen molar ratio from about 2:1 to 4:1, preferably from about 2.5:1 to 3.5:1 by the addition of nitrogen.

Impurities, including methane, carbon dioxide, carbon monoxide and H 2 S are removed in a pressure swing adsorption unit 118. Recovery of hydrogen in the pressure swing adsorption step 118 can be marginally improved by depressurizing the adsorption beds and using nitrogen to desorb the beds. A portion of the stripping nitrogen is carried forward with the product hydrogen. The tailgas from the pressure swing adsorption step 118, containing carbon monoxide, carbon dioxide, hydrogen, nitrogen, methane and water vapor, is fed to a catalytic combustion unit at step 122. The temperature of the tailgas entering the catalytic combustion unit is from about 570°F (300°C) to 1100°F (590°C) . Catalytic combustion is effected at a combustion temperature from about 600°F (316°C) to 1800°F (980°C) and a pressure from 1 atmosphere (100 KPa) to 2 atm. (200 KPa). Space velocities from about 8000 hr "1 to 500,000 hr "1 can be achieved. The exit gas from the catalytic combustion step 122 contains primarily carbon dioxide and water, together with minor amounts of nitrogen, methane and any other inert materials, e.g. argon. Water is condensed from the exit gas by cooling the gas through a heat exchanger at step 124. Carbon dioxide is thus recovered from the gas stream and sold as a valuable item of commerce.

The product exit gas from the pressure swing adsorption step 118 is admixed with a sufficient amount

of nitrogen to bring the hydrogen-to-nitrogen molar ratio to a value from about 2:1 to 4:1, preferably from about 2.5:1 to 3.5:1. The gas, consisting essentially of hydrogen and nitrogen, is compressed to a pressure between about 80 atm (8100 KPa) and 150 atm (15200 KPa) and fed to the ammonia synthesis loop 126, where the hydrogen and nitrogen are reacted under ammonia-producing conditions, as previously described, to produce ammonia. Fig. 10 illustrates an embodiment of the invention in which hydrocarbonaceous feedstock is converted to ammonia by an embodiment of the invention that employs relatively small amounts of energy.

Following optional desulfurization 100, the hydrocarbonacoeus feedstock, together with steam and oxygen or oxygen-rich gas, is fed to the catalytic partial oxidation zone 102 where they undergo catalytic partial oxidation, as previously described, to produce synthesis gas containing hydrogen, carbon dioxide, carbon monoxide, methane and little or no nitrogen. After heat exchange 104, shift gas reaction 106 and heat exchange 108, as previously described, carbon dioxide is removed at step 112 by means such as absorption of carbon dioxide by contact with a countercurrent stream of liquid absorbent, as previously described in connection with Fig. 6. Carbon dioxide from the carbon dioxide removal step 112 can be recovered and sold as a valuable item of commerce. The product gas stream exiting the carbon dioxide removal step 112 is fed to a pressure swing absorber where impurities including methane, carbon dioxide, carbon monoxide and H 2 S are removed. An improved recovery of hydrogen in the pressure swing adsorption step 118 can be achieved by depressurizing

the adsorption beds and using nitrogen to desorb the beds. A portion of the stripping nitrogen is carried forward with the product hydrogen.

In the embodiment of Fig. 10, a portion of the waste (off-gas) from the pressure swing adsorption unit is recycled to the catalytic partial oxidation zone where it provides heat to the reactants. Preferably, between about 50% and 80% of the waste gas is recycled from the pressure swing adsorption unit and the remainder is used as fuel. The product hydrogen exiting the pressure swing adsorption step 118 is admixed with sufficient nitrogen to bring the nitrogen-to-hydrogen molar ratio to between about 2:1 and 4:1, preferably between about 2.5:1 and 3.5:1. The gas is then compressed and fed to the ammonia synthesis loop, where hydrogen and nitrogen are reacted at step 126 under ammonia-producing conditions, as previously described, to produce ammonia.

An ammonia plant for producing ammonia from natural gas is illustrated in Figs. 16 and 17. This system employs a process similar to that of Fig. 9 and has the catalytic oxidation step 102, heat removal step 104, carbon monoxide shift step 106, heat removal step 108, pressure swing absorbtion step 118, combustion step 122, and ammonia loop process 126 indicated generally therein. The natural gas feedstock is received on line 200 and is passed through saturator 202 countercurrent to a heated water flow from line 204 fed into the top of the saturator. The saturator 202 is designed to saturate the natural gas with water so as to reduce the steam requirements for the process. The saturated feed gas stream 206 from the saturator 202 is then further mixed with steam from branch 208 of the output of a high pressure steam drum 210 to produce

the desired steam and natural gas mixture. This mixture on line 212 is then fed through heating coils 213 in a fired heater 214 to produce the desired input temperature for the catalytic partial oxidation. The fired heater is operated, at least partially, by waste fuel in line 216 produced by the process. The fired heater 214 also heats input water in line 218 which is fed to the steam drum 210, and heats steam in coils 221 from branch 220 from the output of drum 210 to produce superheated steam in line 222 which is utilized in the process, for example to drive a turbine compressor.

A heated hydrocarbon process stream 226 from the output of the fired heater 214 is then fed to the catalytic partial oxidation reactor 28, shown in Fig. 1, where it is mixed with oxygen fed through line 228 and fed to the catalytic reaction zone of the reactor 28 to catalytically partially oxidize the natural gas and produce synthesis gas. The synthesis gas from the reactor 28 in line 230 is passed through a heat exchanger 232 which is cooled by a water stream from the high pressure steam drum 210 to partially cool the process stream. From the heat exchanger 232, the process stream is passed through line 234 which is mixed with additional steam applied through line 236 to form an input 238 to a high temperature shift reactor 240. From high temperature shift reactor, the process stream is fed line 242 to heat exchanger 244, where it is further cooled by a water stream from the steam drum 210. The process stream output 246 from heat exchanger 244 is further cooled by water spray from line 248 to form a process stream 250 which is fed to low temperature shift reactor 252. The high temperature shift reactor 240 and low temperature shift reactor 242 perform the water shift reaction step 106 to convert

carbon monoxide in the process stream into hydrogen and carbon dioxide. Process gas stream 254 from the reactor 252 is then passed through heat exchanger 256, line 258, heat exchanger 262, line 263, heat exchanger 264 and line 266 to a knock-out drum 268 where water is removed from the process stream. Purged gas stream 270 from the ammonia loop is combined with the process stream in the drum 268. The resulting process stream 272 from the drum 268 is then applied to a pressure swing adsorption unit 274 where carbon oxides and other impurities are removed from the process stream. Nitrogen on line 276 is applied to the pressure swing adsorption unit 274 to aid in desorption. Also branch 278 of the nitrogen feed stream is combined with output stream 280 of the pressure stream adsorption unit to form the ammonia makeup gas stream 282 which is fed to the ammonia loop 126 of Fig. 17.

Condensate on line 286 from the knock-out drum 268 is fed through branch 288 for use in other processes, and through pump 290 to form portions 292 and 294. The portion 292 is combined with a water recycle stream from the water output 296 of saturator 202 through pump 295 to form the cooling stream 298 through the heat exchanger 256. This stream 298, heated by the heat exchanger 256 forms the heated water input stream for the saturator 202. The remaining portion of the output 296 of the saturator 202 is passed through line 304 for offsite blowdown. The condensate portion 294 is combined with heated water stream 302 to form the water spray stream 248 which is used to chill the input 250 of the low temperature shift reactor.

In the ammonia process 126 shown in Fig. 17, the loop makeup gas input stream 284 passes through

compressor 310, flash drum 312, water cooled heat exchanger 314, compressor 316, flash drum 318, water cooled heat exchanger 320 and line 322 which is combined with ammonia loop recycle stream 324 to form process stream 326. The process stream 326 is circulated by compressor 324 through line 330 and heat exchanger 332 to line 334 connected to an internal heat exchanger input of ammonia converter 336. After heating in the converter 336, the process stream from is passed through line 338 and the ammonia converter 336 wherein hydrogen and nitrogen are reacted in the presence of a catalyst to form a portion of the product stream into ammonia. The output of the ammonia converter on line 340 passes through heat exchanger 342 where it is cooled by heat exchange with water flow from drum 344. From heat exchanger 342 the process stream passes through line 346 to the heat exchanger 332 where it heats the incoming stream 330. The process stream then passes through line 348, water cooled heat exchanger 350, line 352 and heat exchanger 354 which is cooled by the recycle stream 356 from an ammonia condensation apparatus. From heat exchanger 354, the incoming strean passes through line 358 and successive chiller sections 360 and 362 to a catch pot 364 where liquid ammonia is gathered. From the catch pot 364, the non-condensed overhead forms the recycle stream 356. The purge gas stream 270 is taken from the stream 356.

Liquid ammonia from the catch pot 364 is transfered through line 366 to an ammonia receiver section 368, and from there is withdrawn by pump 370 to ammonia product line 372. Sections 374 and 376 of the ammonia receiver provide ammonia streams for the respective chiller sections 360 and 362 to condense

ammonia in the product stream. Gaseous product from the ammonia receiver sections 368, 376 and 374 are compressed by compressors 380, 382 and 384, and passed through line 386 to a heat exchanger 388 and a refrigeration loop receiver 390 from which liquid refrigerant over line 392 is fed back to ammonia receiver section 374. Flash gas condenser 394 receives a portion of the stream from line 392 and returns the further cooled portion through line 396 to the receiver section 368. A turbine 385 drives the compressors 380, 382 and 386.

When compared to present day commercial processes, the catalytic partial oxidation process using, as oxidant, a stream containing in excess of 70 mole percent of oxygen, as described herein, the process of the invention offers the following advantages.

(1) The high cost steam reforming furnace and secondary reformer are eliminated when compared to the conventional commercial process.

(2) Low oxygen consumption when compared to conventional partial oxidation.

(3) Low water consumption when compared to steam reforming. (4) Low cost when compared to catalytic partial oxidation processes which use air or enriched air to produce a nitrogen rich or stoichiometric synthesis gas at the exit of the catalytic partial oxidation reactor. (5) Reduced area requirement when compared to the steam reforming route (particularly suitable for offshore application). (6) High efficiency when compared to conventional commercial ammonia production

processes and when compared to catalytic partial oxidation using air or enriched air containing less than 70 mole percent of oxygen.

*--

(7) Lower in capital cost then all present commercial processes.

The following examples are intended to illustrate further the invention described herein and are not intended to limit the scope of the invention in any way.

EXAMPLE I

Natural gas is converted to synthesis gas in a catalytic partial oxidation reactor of the construction shown in Fig. 1. There are included nine catalyst discs 54, each having a diameter of 30 inches (0.76m) and a thickness of 10 inches (0.25m). The discs are formed from a honeycomb monolith of cordierite material with a geometric surface area of approximately 25 cm 2 /cm 3 . A high surface area alumina layer is deposited on the cordierite to serve as a support upon which finely dispersed catalytic metal components are distended. The catalytic metal components are approximately 50% by weight platinum and 50% by weight palladium. Space velocity of the catalyst is 97,000 hr. "1 Natural gas (>95% methane) is mixed with steam at various steam-to-carbon molar ratios, heated and supplied through 10-inch diameter inlet 66 at a pressure of 400 psig (2760 KPa). Air is heated and supplied through two 8-inch inlets 70 at a pressure of 410 psig (2830 KPa). The diameter of the lower portion 76 of the chamber 72 is 27 inches (0.68m) with the diameter of the upper portion 74 being 36 inches

(0.91m). There are 261 tubes 80 having 0.5 inch (12.7mm) internal diameters and having lengths of 20 inches (0.51m). Six orifices 86 of 0.123-inch (3.2m) diameter are formed in each tube with four of the orifices evenly spaced around each tube at a distance of 4 inches (0.102m) above the lower end of the tube and with the remaining two orifices formed opposite each other at a distance 6 inches (0.152m) above the lower end of the tube. The bottom member 78 has a thickness of 5 inches (0.127m) and the passageway sections 84 are conical with upper diameters of 0.5 inches (12.7 mm) and lower diameters of 1.75 inches (44.5 mm). Pressures within the chambers 68 and 72 are maintained at essentially the inlet pressures. The temperature of the mixed reactant gases is 1,100°F (590°C). Fig. 11 shows oxygen consumption for the catalytic partial oxidation process as a function of steam-to-carbon molar ratio, for reaction temperatures of 1,600°F (870°C), 1,750°F (950°C) and 1,900°F (1040°C) and an operating pressure of 400 psig (2700 KPa). It can be seen from the graph that oxygen consumption, expressed as oxygen-to-carbon molar ratio, is relatively low for the process of the invention as compared with present commercial partial oxidation processes. The dashed line 25 in Fig. 11 represents the linear function of minimum temperatures and steam/carbon ratios required to prevent carbon deposits.

Fig. 12 shows the molar ratio of hydrogen, as H 2 , to carbon monoxide in the product as a function of the steam-to-carbon ratio for reaction temperatures of 1,600° F (870°C), 1,750°F (950°C) and 1,900°F (1040°C).

Figs 13 and 14, respectively, show the amounts of methane and carbon dioxide, as volume %, in

the product as a function of the steam-to-carbon ratio for reaction temperatures of 1,600°F (870°C), 1,750°F (950°C) and 1,900°F (1040°C).

Fig. 15 shows the effective H 2 production of the process, expressed as total moles of H 2 and carbon monoxide in the product divided by total moles of H 2 and carbon in the feedstock.

Example II The following example describes the production of ammonia from a gaseous hydrocarbonceous feedstock using the process of the invention represented by Figs. 9 and 10.

Hydrocarbonceous feedstock is desulfurized using conventional methods depending on the quantity and type of sulfur compounds, contained in the feedstock. Desulfurization may, for example, be conveniently carried out by preheating the hydrocarbonceous feedstock at a temperature between 250°F (120°C) and 750°F (400°C) and absorbing the sulfur compounds, into zinc oxide contained in one or more desulfurization vessels.

Steam is added to the desulfurized feedstock to give a steam-to-carbon ratio of between 1.0 and 1.7 to 1.0. The steam may be added either directly or by feed gas saturation. The hot water for feed gas saturation is conveniently provided by recovering heat from the synthesis gas at a point downstream of the shift reactors.

The mixed feedstock is preheated to approximately 1100°F (590°C) in a fired heater and passed to the catalytic partial oxidation reactor where it is admixed with preheated oxygen as oxygen containing gas containing at least 70 mole percent of

oxygen in a ratio of between 0.5 and 0.55 moles of oxygen per atom of carbon contained in the feedstock, before passing to the partial oxidation catalyst where the above reactions (3) and (4) occur. The exit temperature from the partial oxidation reactor is about 1700°F (930°C). Heat is recovered from the effluent gas from the reactor, by raising steam in a boiler, before additional steam is added to the synthesis gas passed to the high temperature shift reactor at a temperature of approximately 700°F (370°C) where more of the carbon monoxide is reacted according to the above equation (4) . The synthesis gas emerges from the high temperature shift reactor at about 850°F (450°C) and is cooled to about 650°F (340°C) in a second boiler which also generates high pressure steam. A water quench is used to reduce the synthesis gas temperature to 425°F (220°C) at which point it undergoes low temperature shift according to the above equation (4) to reduce the carbon monoxide content and increase the hydrogen content further. Heat is recovered at the exit of the low temperature shift reactor by preheating water, which is used to saturate the hydrocarbonceous feedstock with steam. Further heat is recovered by preheating demineralized water which is passed to a deaerator for use as boiler feedwater. Then the synthesis gas is cooled to approximately 100°F (38°C) with cooling water and condensed water is separated in a knock-out drum.

Two options of the process route are included in the processes of Figs. 9 and 10 depending on the relative requirements for minimizing capital cost and maximizing feedstock conversion efficiency. Fig. 9 shows the minimum capital cost option. In Fig. 9 the remaining carbon monoxide, methane, water vapor and

carbon dioxide are removed from the synthesis gas by pressure swing adsorption, as described earlier, to yield a high purity hydrogen stream. Nitrogen, from the air separation plant is then added to produce a gas containing hydrogen and nitrogen in a mole ratio of between 2.5 and 3.5 moles of hydrogen per mole of nitrogen. The waste gas from the pressure swing adsorption unit (off-gas) is used either as fuel in the fired heater or, if carbon dioxide is required, for example for the downstream production of urea, it may be burned catalytically, with a portion of the oxidant stream, using an oxidation catalyst. The methane, carbon monoxide and hydrogen are converted to carbon dioxide and water and the effluent stream from the catalyst combustion therefore contains essentially only carbon dioxide, nitrogen and water vapor. Water is condensed from this stream by reducing its temperature and at the same time useful heat is recovered. The water is finally separated in a knock-out drum to yield the product carbon dioxide stream.

Fig. 10 shows the maximum efficiency option. In Fig. 10 carbon dioxide is recovered from the synthesis gas, in a conventional chemical or physical adsorption process 112, as previously described. The carbon dioxide may be sold as a commercial product.

Pressure swing adsorption 118 is then used to produce a pure hydrogen stream. The pressure swing adsorption off-gas contains essentially only hydrogen, carbon monoxide, nitrogen, methane and any residual water vapor and carbon dioxide. Part of this gas is used as fuel to the fired heater and the remainder is recycled to be used as feedstock in the catalytic partial oxidation reactor. Alternatively, it may be recycled to a point upstream of either the high temperature

shift reactor or the low temperature shift reactor. Nitrogen from the air separation plant is added to the purified hydrogen stream in a molar ratio of between 2.5 and 3.5 moles of hydrogen per mole of nitrogen to produce a gas suitable for ammonia synthesis.

Before it can be used for ammonia synthesis the synthesis gas must be compressed to approximately 100 atm (10130 KPa).

The make up gas is mixed with the circulating gas in the synthesis loop at the suction of the circulator. The bulk of the gas leaving the circulator is preheated in the loop interchanger after which the gas splits into two streams. One stream is used as quench gas for moderating the synthesis reaction temperature and is injected in between the first and the second beds of the ammonia converter. The other stream is the converter feed gas and this is preheated to reaction temperature, by heat exchange with the effluent gas leaving the second bed, in a heat exchanger located inside the ammonia converter.

Effluent gas leaving the converter at approximately 766°F (408°C) is cooled to 7°F (-13°C) in a boiler which raises high pressure steam, loop gas interchanger, loop cooler, recycle gas interchanger, and the loop chiller, and then passed into the catchpot. The uncondensed gas is taken from the top of the catchpot, preheated in the recycle gas interchanger and then recycled to the suction of the circulator. A small purge is taken from the recycle gas in order to avoid buildup of the inerts level in the ammonia synthesis loop. This purge is recycled to a point upstream of the pressure swing adsorption unit or alternatively upstream of the catalytic partial oxidation reactor.

Liquid ammonia leaves the catchpot under level control and is reduced to atmospheric pressure in the ammonia receiver. Ammonia and dissolved gases flash off and are separated from the liquid ammonia in the ammonia receiver, and recompressed in the three stage refrigeration compressor, to 240 psig (1650 KPa). From the discharge of the refrigeration compressor, the flash gases are cooled with cooling water to condense the bulk of the ammonia, which is separated in the refrigeration loop receiver.

The ammonia content of the flash gases is further reduced by cooling the flash gases to -14°F (- 26°C) with ammonia let down from the refrigeration loop receiver and returning to the atmospheric pressure section of the ammonia receiver. The majority of the condensed ammonia from the refrigeration loop receiver is let down to the high pressure section of the ammonia receiver, operating at 49 psig (338 KPa) and 33°F (1°C). The high pressure section of the ammonia receiver also acts as a flash vessel for the primary section of the chiller. Flash gases from this section enter the third step of compression of the refrigeration compressor while the liquid is let down to the medium pressure section of the ammonia receiver operating at 16 psig (110 KPa) and 0°F (-17°C). This section acts as a flash vessel for the secondary section of the chiller. The flash gases from the medium pressure section of the ammonia receiver enter the second stage of the refrigeration compression whilst the liquid phase is let down to the low pressure section of the ammonia receiver operating at atmospheric pressure and -28°F (-33°C). From here, the product ammonia is pumped to atmospheric storage.

Example III The following TABLES I, II and III contain moles/hour, mole percent, and parameters of pressure, temperature, water/steam flow, and heat transfer for an ammonia plant according to Figs. 16 and 17. The moles/hour are lb moles/hour (0.4536 kg moles/hour) and the plant produces 600 short ton NH 3 per day (544 X 10 3 Kg/day) .

DESCRIPTION CO Ar 02 H20 NH3 TOTALS

Natural Gas 1818.04

(line 200) S. Nat. Gas 1.09 2728.09 4546.09

(line 226) 0i7 en 4.64 923.85 928.49

(line 228) C?0 out 1137.57 537.99 3708.98 144.25 1.82 4.64 2364.43 7899.68

(line 232) HT Shift 255.56 1410.00 4580.98 144.25 1.82 4.64 3556.20 9963.45

(line 242) LT Shift 19.24 1656.32 4827.31 4.25 1.82 4.64 4263.99 10917.57

(line 254) Purge Gas 95.41 125.61 11.07 232.09

(line 270) PSA input 19.24 1654.81 4922.70 144.25 127.43 4.64 24.52 11.07 6908.76

(line 272) Condensate 1.51 .02 4239.37 .01 4240.91

(line 286) Nitrogen 1616.96 1616.96

(line 278) PSA out 4527.52 4527.52

(line 230) Loop Make-Up 4527.52 1616.96 6144.48

(line 282) PSA Vasts 19.24 1654.81 395.79 144.25 127.43 4.64 24.52 11.07 2381.85

(line 216) Combustion A..ir 0.90 2331.77 23.09 627.15 2987.92

(line 217) Stack Gas 1819.21 2477.23 32.66 10<f.27 734.20 13,43 5181.05 CNC2)

TABLE I (CONTINUED)

Moles / Hour

DESCRIPTION CO C02 H2 CH4 C2H6 C3H8 C4H10 N2 Ar 02 H20

Converter in 15949.18 16656.51

(line 330) Converter out 11521.92 15181.77

(line 340) Catchpot Gas 11517.07 15165.16

(line 356) Recycle Gas 11421.66 15039.55

(line 324) Catchpot Liquid 4.85 16.61

(line 365) Flash Gas 4.85 15.60

Aπ-monia Product (line 372)

TABLE I I

Mole Percent

CO C02 H2 CH4 C2H6 C3H8 C4H10 02 H20 NH3

0.06 99.69 0.11 0.03 0.01

0.02 39.87 0.04 0.01 0.00 50.01

99.50

14.40

2.6.

0.18

0.23

100.00

100.00

73.68

0.31 69.98 16.52 6.06

0.02 35.12

TABLE I I (CCNTINUED) Mole Percent

DESCRIPTION CO C02 H2 CH4 C2H6 C3H8 C4H10 N2 Ar 02 H20 NHj

Converter in ".00 . 49.08 3.92

(line 330)

Converter out 37.19 49.00 13.81

(line 340) _ β

Catchpot Gas 41.10 54.12 4.78

(line 356) _ a

Recycle Gas 41.10 54.12 4.73

(Line 324) βn __

Catchoot Liquid 0.16 0.56 9*5-28

(line 366) . ..

Flash Gas 20.43 69.92 9.o5

Airaioni .a _ Prod .uc .t 100.00

(line 372)

DESCRIPTION Net Heat Transfer MM3TU/HR 5CCAL 10 6 /HR

Natural Gas 200 Saturation Water 204 Saturated Gas 206 Steam 208 Steam Drum 210 Procss Gas/Steam 212 213 36.226 9.129 output 21

25.024 6.306 ea3i 222 225 232 58.518 14.772 4 787 419

37179 16864 233 t 242 244 17.056 4.293 6 17188 7796 LT Shift input 250 LT Shift output 254 Ξeat Exchanger 256 Process Gas 258 Heat Exchanger 262 Process Gas 26S Eeat Exchanger 26*» Process Gas 266 Purge Gas 170 PSA inaut 272

TABLE I I I (CCNT I NUED )

Parameters

DESCRIPTION Pressure Temperature Water/Steam Net Heat Transfer PSIG KPa DEG F DEG C LBS /HR KG/HR MMBTU/ HR KCAL 10 5 / HR

Nitrogen 276 240 1655 105 41 PSA Output 280 240 1655 105 41 Loop Make-Up 282 239 1648 Condensate 288 Pump 290 Water 292 Pump 295 Vater 294 Vater 302 267 131 "water 304 Process Gas 322 Process Gas 326 Process Gas 330 1500 103-42 Heat Exchanger 332 100.399 25 .300 u Process Gas 334 Converter Output 340 1470 10135 Heat Exchanger 342 51.120 12.881 Process Gas 346 Process Gas 348 Heat Exchanger 350 3.931 2.251 Process Gas 352 Heat Exchanger 354 15.977 <f .026 Process Gas 358 Chiller Section 360 15 .903 H.oal Chiller Section 362 15. 312 4MU Catch Pot ~ 6M Receiver Section 368 Receiver Section 376 Receiver Section 374 Ammonia Product 372

Since many variations, modifications and changes in detail can be made to the above described embodiments, it is intended that the subject described above and shown in the accompanying drawings be interpreted as illustrative and not in a limiting sense.