[0001] The present disclosure relates to processes and apparatuses for alkylation of isoparaffins and, in particular, to processes and apparatuses for alkylation of isoparaffins with olefins to produce high octane rated additive for fuels, such as gasoline.

BACKGROUND OF THE INVENTION

[0002] The alkylation of isoparaffins, such as isobutane, is an important refinery process for the production of high octane alkylate as a blend component for gasoline. Alkylation involves the addition of an alkyl group to an organic molecule. Thus, an isoparaffin can be reacted with an olefin to provide an isoparaffin of higher molecular weight. The product is a valuable blending component for gasoline due to its high octane rating, low sulfur, low olefin, and low aromatic content. Industrially, alkylation often involves the reaction of C2-C5 olefins with, for example, isobutane in the presence of an acidic catalyst to form alkylates. Alkylates are valuable blending components for the manufacture of premium gasolines due to their high octane ratings.

[0003] In the past, alkylation processes have included the use of liquid acids, such as hydrofluoric acid or sulfuric acid as catalysts. The use of liquid acids provides challenges in disposal of spent acid streams. Furthermore, consideration has been given by regulatory authorities to the restriction of the use of liquid acids in industrial alkylation reactions. An alternative to liquid acids are solid acids, such as zeolites. However, some solid acids, such as faujasite, typically have short catalyst lifetimes which lead to frequent catalyst regeneration and increased costs and may further require the use of precious metals such as platinum and palladium in catalyst regeneration. [0004] Recent efforts in further improving alkylation catalysts over liquid acid catalysts and previous solid acid catalysts have been focused on the development and use of solid acid catalysts, including zeolites, such as zeolites having the MWW framework type, e.g. MCM-22, MCM-36 and MCM-49 for the catalytic alkylation of an olefin with an isoparaffin. (US 4,992,615, US 5,254,792, US 5,304,698, US 5,354,718, US 5,516,962). The previous approaches in alkylation of isoparaffins focused on using a single stage reactor where the feed isobutane to olefin ratio (i:o ratio), a volume to volume ratio, was set by the composition of the gas entering the single stage reactor. For liquid acids the i:o ratio has typically been 4:1 to 10:1, and for solid catalysts the i:o ratio has typically been 40: 1 to 50:1, both based solely on the composition of the feedstock entering a single stage reactor.

[0005] The use of a single stage reactor may limit the ability to convert olefins and isoparaffins into higher octane rated fuel additives. For example, a single stage reactor does not permit splitting of the olefin feedstock creating a lower local concentration of olefin and a greater i:o ratio. [0006] There remains a need for improved isoparaffin-olefin alkylation processes that can be catalyzed by a solid acid catalyst with high conversion and high activity that maintains product quality of existing liquid phase processes.

SUMMARY OF THE INVENTION [0007] The present disclosure is related to processes for the alkylation of an isoparaffin. The processes may include introducing, in a multistage reactor, an isoparaffin feed, an olefin feed, and a hydrogen feed to a solid acid catalyst and a hydrogenation co-catalyst, where the solid acid catalyst includes a zeolite.

[0008] Processes for the alkylation of an isoparaffin may further include introducing, in a multistage reactor, an isoparaffin feed, an olefin feed, and a hydrogen feed to a solid acid catalyst and a hydrogenation co-catalyst, where the solid acid catalyst includes a crystalline microporous material of the MWW framework type and the hydrogenation co-catalyst includes a noble metal. [0009] Additionally, the present disclosure relates to catalyst beds for the alkylation of an isoparaffin with an olefin and hydrogenation of olefin, the catalyst bed having a volume. The catalyst beds may include an upstream zone including a zeolite and a downstream zone including a hydrogenation co-catalyst.

BRIEF DESCRIPTION OF THE DRAWING

[0010] FIG. 1A is a depiction of a reactor with one stage configured to receive an olefin feed and an isoparaffin feed. [0011] FIG. IB is a depiction of a reactor with two stages configured to receive an olefin feed and an isoparaffin feed, according to an embodiment.

[0012] FIG. 1C is a depiction of a reactor with four stages configured to receive an olefin feed and an isoparaffin feed, according to an embodiment.

[0013] FIG. ID is a depiction of a reactor with eight stages configured to receive an olefin feed and an isoparaffin feed, according to an embodiment.

DETAILED DESCRIPTION OF THE INVENTION

[0014] Previous alkylation processes and systems described in previous patent applications relied on a single stage reactor. Single stage alkylation reactors may provide lower conversion of isoparaffins and olefins into higher octane rated fuel additives, increased by-product formation, and can be limited in flow rate or i:o ratio. It was believed that the addition of multiple stages would improve conversion because olefin interactions with active catalyst site could increase with additional stages including catalyst. Additionally, the use of multiple stages may allow for use of similar feeds, but provide a larger i:o ratio at each stage than a single stage reactor could provide. A multiple stage reactor may offer improvements over single stage processes including splitting of olefin introduction into various stages which decreases the local concentration of olefin in a catalyst bed, which may provide improved i:o ratios, decreased issues with catalyst deactivation, decreased byproduct formation, and improved conversion of isoparaffins. Improved conversion may result from increased olefin interactions with active catalyst sites resulting from passing over catalyst beds within additional reactor stages.

[0015] However, it has been discovered that performing the alkylation in multiple stages shows a decrease in conversion and increased loss of catalyst activity at each subsequent stage. Without being limited by theory, it is believed that olefin oligomerization creates higher olefins that block catalyst active sites decreasing alkylation of isoparaffins and reducing catalyst activity. Furthermore, the feed entering a later stage does not have the same chemical composition as the feed entering previous stages because the feed after the first stage contains some quantity of higher olefins. Additionally, the problem of higher olefin content may be additive because a decrease in catalytic activity lowers isoparaffin conversion rates. At lower isoparaffin conversion rates more olefin oligomers are formed and, therefore, catalyst activity may decrease more quickly.

[0016] The benefits of a multistage reactor can be realized by removing the olefin oligomers produced via interstage hydrogenation.

[0017] It has been discovered that high conversion rates, maintenance of high catalyst activity, and decreased production of olefin oligomers can be achieved by hydrogenation of byproducts after reaction in one or more stages of a multistage reactor. Hydrogenating the product mixture may remove olefin oligomers thereby decreasing catalyst deactivation and increasing conversion rates.

Definitions

[0018] The term “Cn” compound (olefin or paraffin) where n is a positive integer, e.g., 1, 2, 3, 4, 5, etc., means a compound having n number of carbon atom(s) per molecule. The term “Cn+” compound where n is a positive integer, e.g., 1, 2, 3, 4, 5, etc., means a compound having at least n number of carbon atom(s) per molecule. The term “Cn-” compound where n is a positive integer, e.g., 1, 2, 3, 4, 5, etc., means a compound having no more than n number of carbon atom(s) per molecule.

[0019] The term “critical point” is the liquid-vapor end point of a phase equilibrium curve that designates conditions under which a liquid and vapor may coexist. At temperatures higher than the critical point (a “critical temperature”) a gas cannot be liquefied by pressure alone. At temperatures and pressures higher than the critical point the material is a supercritical fluid. For the purposes of this disclosure the critical point for isobutane is 134.6 °C and 3650 kPa, and the critical point for isopentane is 187.2 °C and 3378 kPa. [0020] As used herein, the term “molecular sieve” means a substance having pores of molecular dimensions that permit the passage of molecules below a certain size. Examples of molecular sieves include but are not limited to zeolites, silicoaluminophosphate molecular sieves, and the like.

Reactor Design and Conditions [0021] The processes described can be conducted in any suitable multistage reactor, such as one including fixed-beds, moving beds, swing beds, fluidized beds (including turbulent beds), and/or one or more combinations thereof. A reactor stage begins at the point in which olefin is introduced and ends at either an interstage space or where additional olefin is introduced. A multistage reactor may have one or more interstage spaces between stages. An interstage space may be an open space, a filled space, a separation barrier, a distribution plate or system, or an injection point. Multistage reactors of the present disclosure may be configured to receive an olefin feed at multiple sites or inlets, and the introduction of olefin marks a new reactor stage. In addition, the reactor may include multiple catalyst beds located in the same or different housing. A multistage reactor or a stage within the reactor may include a bed of catalyst particles where the particles have insignificant motion in relation to the bed (a fixed bed). In addition, injection of the olefin feed can be effected at a single point in the reactor or at multiple points spaced along the reactor. The isoparaffin feed and the olefin feed may be premixed before entering the reactor. [0022] In certain embodiments of the present disclosure, the multistage reactor includes a plurality of fixed beds, continuous flow-type reactor stages in either a down flow or up flow mode, where the reactor stages may be arranged in series or parallel. The multistage reactor may include multiple reactor stages in series and/or in parallel, for example, a multistage reactor may include 2 stages, 4 stages, 8 stages, 10 stages, 12 stages, or any other plurality of stages. A reactor stage includes a catalyst bed. The reactor stage may have various configurations such as: multiple horizontal beds, multiple parallel packed tubes, multiple beds each in its own reactor shell, or multiple beds within a single reactor shell. In certain embodiments, a reactor stage includes a fixed bed which provides uniform flow distribution over the entire width and length of the bed to utilize substantially all of the catalyst. In at least one embodiment, the multistage reactor can provide heat transfer from reactor stages or catalyst beds in order to provide effective methods for controlling temperature.

[0023] A catalyst bed includes catalyst and hydrogenation co-catalyst. The catalyst and co catalyst might not be equally distributed throughout the catalyst bed, and are such that the catalyst and co-catalyst may form a gradient of catalyst to co-catalyst in the direction of flow. Therefore, the co-catalyst may be partially or completely downstream of the catalyst in a catalyst bed. The gradient within a catalyst bed may be stepwise (e.g. a number of zones) or gradual, e.g. such a large number of zones that a single zone is no longer distinguishable from adjacent zones, but the zones together form a gradient of catalyst to co-catalyst. In some embodiments, the gradient is stepwise with two zones, where the catalyst bed has an upstream zone including catalyst and a downstream zone including co-catalyst. In at least one embodiment, the upstream zone is substantially free of co-catalyst or substantially free of noble metals. Additionally, the catalyst and co-catalyst may be unequal in volume or weight and therefore a zone may be greater or smaller in volume than another zone within a catalyst bed. For example, an upstream zone including catalyst may be about 60% or more of the volume of a catalyst bed and the downstream zone including co catalyst occupies the remaining volume, such as an upstream zone including catalyst that is about 65% or more, about 70% or more, about 75% or more, about 80% or more, about 85% or more, about 90% or more, about 95% or more, about 98% or more, about 99% or more of the volume of the catalyst bed, with the remaining volume occupied by a downstream zone including co-catalyst. [0024] The efficiency of a multistage reactor containing fixed beds of catalyst may be affected by the pressure drop across a fixed bed. The pressure drop depends on various factors such as the path length, the catalyst particle size, and pore size. A pressure drop that is too large may cause channeling through the catalyst bed, poor efficiency. In some embodiments, the reactor has a cylindrical geometry with axial flows through the catalyst beds.

[0025] The various designs of the multistage reactor may accommodate control of specific process conditions, e.g. pressure, temperature, and LHSV. The LHSV determines volume and residence time that may provide the desired conversion.

[0026] Operating pressures may be controlled to reduce or eliminate oligomerization reactions and/or favor alkylation reactions. Additionally, increased reactor pressures may improve conversion rates for the olefin feed and improve selectivity towards the alkylated paraffin over olefin oligomers. Operating pressure may be from about 300 to about 1500 psig (about 2068 to about 10342 kPag), such as from about 400 to about 1200 psig (about 2758 to about 8274 kPag), from about 450 psig to about 1000 psig (about 3102 to about 6895 kPag), from about 550 psig to about 950 psig (about 500 to about 6550 kPag), from about 650 psig to about 950 psig (about 4482 to about 6550 kPag), from about 750 psig to about 950 psig (about 5171 to about 6550 kPag), or from about 800 psig to about 950 psig (about 5516 to about 6550 kPag). In some embodiments, the operating temperature and pressure remain above the critical point for the isoparaffin feed during the reactor run.

[0027] Additionally, operating temperatures may be controlled to reduce or eliminate olefin oligomerization reactions and/or favor alkylation of isoparaffins. Operating temperature may be from about 100 °C or greater, such as about 130 °C or greater, about 140 °C or greater, about 150 °C or greater, or about 160 °C or greater, such as from about 100 °C to about 200 °C, from about 130 °C to about 170 °C, or from about 140 °C to about 160 °C. Operating temperatures may exceed the critical temperature of the isoparaffin feed, or the principal component in the isoparaffin feed. The term “principal component” is defined as the component of highest concentration in the feedstock. For example, isobutane is the principal component in a feedstock consisting of isobutane and 2-methylbutane in isobutane:2-methylbutane weight ratio of about 50:1.

[0028] The temperature of the multistage reactor or an individual stage within the reactor may affect by-product formation and a temperature higher than 130 °C may decrease heavier olefin concentrations. Furthermore, an increase in temperature may improve conversion of the olefin feed. However, for certain olefins, a higher temperature increases olefin isomerization, and olefin isomerization may lead to the formation of alkylation products that are lower in value. For example, in the alkylation of isobutane with 2-butene, a main component of the alkylation product mixture is trimethylpentane which has an octane rating of 100, but if 2-butene is isomerized to 1 -butene the alkylation shifts to higher production of dimethylhexane which has an octane rating of 70, providing less value as a fuel additive. Therefore, temperature may be used to reduce or eliminate heavier olefin concentrations, especially in cases where the olefin is not affected by isomerization, such as propene or isobutene. In some embodiments, the alkylation product mixture contains about 10 wt% or less, such as about 5 wt% or less, about 2 wt% or less, about 1 wt% or less, or is substantially free of products of olefin oligomerization.

[0029] Hydrocarbon flow through a reactor stage containing the catalyst is typically controlled to provide an olefin liquid hourly space velocity (OLHSV) sufficient to convert about 99 percent, or more, by weight of the fresh olefin to alkylation product. In some embodiments, OLHSV values are from about 0.01 hr'to about 10 hr ¹ , such as about 0.02 hr'to about 1 hr ¹ , or such as about 0.03 hr ¹ to about 0.1 hr ¹ . The liquid hourly space velocity of the isoparaffin is controlled to meet a target i:o ratio. Because the i:o ratio is vokvol, the isoparaffin liquid hourly space velocity is directly correlated to the OLHSV.

[0030] FIG. 1A depicts an alkylation reactor 100 A with a single reactor stage 101. Reactor stages(s) may individually or collectively be termed an alkylation zone and include catalyst, such as a solid acid catalyst including zeolite of the MWW framework type. The olefin feed is introduced to reactor stage 101 via line 103 and the isoparaffin feed through line 105. An alkylation product mixture exits the reactor through line 107. In alkylation reactor 100A the i:o ratio is controlled solely by the composition of the olefin feed and the isoparaffin feed entering reactor bed 101 [0031] FIG. IB depicts a multistage alkylation reactor 100B with two reactor stages: first stage 101 A and second stage 10 IB. The olefin feed is introduced to the reactor beds via lines 103 A and 103B and OLHSV values are from about 0.01 hr'to about 10 hr ¹ , such as about 0.02 hr 'to about 1 hr ¹ , or such as about 0.03 hr'to about 0.1 hr ¹ . The split introduction of the olefin feed allows a lower concentration (half) of the olefin feed to be introduced locally to each of the first stage 101 A and the second stage 101B. The isoparaffin feed is introduced to alkylation reactor 100B through line 105. Alkylation reactor 100B has an interstage space 109 between first stage 101 A and the second stage 101B to allow for introduction of additional olefin feed through line 103B. Reactor stages 101 A and 101B may individually include an upstream zone and a downstream zone. For example, in reactor stage 101 A, a downstream zone is proximate interstage space 109 and may include hydrogenation co-catalyst. Furthermore, the downstream zone may simply be a thin layer of hydrogenation catalyst packed downstream in a fixed bed. Upstream and downstream zones within a reactor stage are not pictured because there may not be a distinct division between and upstream zone and a downstream zone, for example, in at least one embodiment, where hydrogenation co-catalyst is supported or bound with catalyst.

[0032] If lines 103 and 105 have the same composition as in FIG. 1A, then the local i:o ratio is doubled in the configuration of FIG. IB because the olefin feed is divided into two lines 103 A and 103B and the olefin introduced via line 103A to first stage 101A can be converted, such as about 90 wt% or greater, 95 wt% or greater, 98 wt% or greater, or 99 wt% or greater is converted in the reaction within first stage 101A, based on the total weight of olefin in the olefin feed introduced via line 103 A. Additionally, only a small portion of the isoparaffin feed is converted by the reaction in first stage 101A, such as about 10 wt% or less, 5 wt% or less, 2 wt% or less, lwt% or less, 0.5 wt% or less, or 0.1 wt% or less of the isoparaffin feed is converted based on the total weight of isoparaffin. Therefore, the amount of isoparaffin introduced to interstage 109 and, therefore, introduced to second stage 101B is similar to that introduced to first stage 101A. For example, if an i:o ratio of 100:1 is introduced to first stage 101A and there is an olefin conversion of 100% then the i:o ratio in second stage 101B would be -99:1, if no additional isoparaffin was added. Furthermore, the olefin introduced to interstage 109 (either via line 103B or from the effluent of first stage 101 A) and, therefore, introduced to second stage 101B is similar in quantity to that introduced to first stage 101A. Additionally, because the selected isoparaffin may be consumed in each stage, such as in amounts of about 10 wt% or less, about 5 wt% or less, about 2 wt% or less, about 1 wt% or less, about 0.5 wt% or less, or about 0.1 wt% or less, additional isoparaffin may be added in an interstage space so as to maintain a consistent i:o ratio throughout the multistage reactor. Similarly to FIG. 1 A, an alkylation product mixture exits the reactor through line 107.

[0033] FIG. 1 C depicts a multistage alkylation reactor 100C with four reactor stages: first stage 101 A, second stage 101B, third stage 101C, and fourth stage 10 ID. The olefin feed is introduced to the reactor beds via lines 103A, 103B, 103C and 103D and OLHSV values are from about 0.01 hr Ho about 10 hr ¹ , such as about 0.02 hr ho about 1 hr ¹ , or such as about 0.03 hr ho about 0.1 hr f The split introduction of the olefin feed allows a lower concentration (one quarter) of the olefin feed to be introduced locally to each of the first stage 101 A, second stage 101B, third stage 101C, and fourth stage 101D. The isoparaffin feed is introduced to alkylation reactor lOOC through line 105. Alkylation reactor lOOC has multiple interstage spaces: first interstage space 109A, second interstage space 109B, and third interstage space 109C between reactor stages 101A, 101B, 101C, and 101D to allow for introduction of additional olefin feed through lines 103B, 103C, and 103D. Reactor stages 101 A, 101 B, 101 C, and 10 ID may individually include an upstream zone and a downstream zone. For example, in reactor stage 101A a downstream zone is close to interstage space 109A and may include hydrogenation co-catalyst. Furthermore, the downstream zone may simply be a thin layer of hydrogenation catalyst packed downstream in a fixed bed. Upstream and downstream zones within a reactor stage are not pictured because there may not be a distinct division between and upstream zone and a downstream zone, for example, in at least one embodiment, where hydrogenation co-catalyst is supported or bound with catalyst.

[0034] If lines 103 and 105 have the same composition as in FIG. 1A, then the local i:o ratio is 4 times that found in FIG. 1A, because the olefin feed is divided into four lines 103 A, 103B, 103C, and 103D. The i:o ratio in a single stage is only slightly affected by prior stage(s) because the olefin introduced to a prior stage can be largely converted within that stage, but the isoparaffin is introduced at such a ratio that the amount converted may have little effect on the ratio in later stages. For example, the olefin introduced via line 103A to first stage 101A is converted, such as about 90 wt% or greater, 95 wt% or greater, 98 wt% or greater, or 99 wt% or greater is converted in first stage 101A, based on the total weight of olefin in the olefin feed introduced via line 103A. Additionally, only a small portion of the isoparaffin feed may be converted by the reaction in first stage 101A (such as about 10 wt% or less), 5 wt% or less, 2 wt% or less, lwt% or less, 0.5 wt% or less, or 0.1 wt% or less of the isoparaffin feed is converted based on the total weight of isoparaffin introduced to first stage 101 A. Therefore, the amount of isoparaffin introduced to interstage 109A and, therefore, introduced to second stage 101B is similar to that introduced to first stage 101A and the olefin introduced to interstage space 109A via line 103B and from the effluent of first stage 101A serves to bring the olefin level back up to a desired i:o ratio. For example, if an i:o ratio of 100: 1 is introduced to first stage 101 A and there is an olefin conversion of 100%, then the i:o ratio in second stage 101B would be -99:1, if no additional isoparaffin was added. The combination of isoparaffin and olefin is then introduced to second stage 101B, where the olefin may be converted in second stage 101B, such as about 90 wt% or greater, 95 wt% or greater, 98 wt% or greater, or 99 wt% or greater is converted in second stage 101B, based on the total weight of olefin introduced to interstage space 109A. Similarly, only a small portion of the isoparaffin feed may be converted by the reaction in second stage 101B, such as about 10 wt% or less, 5 wt% or less, 2 wt% or less, lwt% or less, 0.5 wt% or less, or 0.1 wt% or less of the isoparaffin feed is converted based on the total weight of isoparaffin introduced to interstage space 109 A. As the isoparaffin feed enters additional interstage spaces (such as second interstage space 109B, and third interstage space 109C) more olefin may be introduced (via lines 103Cand 103D) to adjust the i:o ratio as the combined feeds are introduced to additional stages (such as third stage 101C and fourth stage 101D). Additionally, because the selected isoparaffin may be consumed in each stage (such as in amounts of about 10 wt% or less, about 5 wt% or less, about 2 wt% or less, about 1 wt% or less, about 0.5 wt% or less, or about 0.1 wt% or less), additional isoparaffin may be added in an interstage space so as to maintain a consistent i:o ratio throughout the multistage reactor. Similarly to FIG. 1A, an alkylation product mixture exits the reactor through line 107.

[0035] FIG. ID depicts a multistage alkylation reactor 100D with eight reactor stages: first stage 101A, second stage 101B, third stage 101C, fourth stage 101D. The olefin feed is introduced to the reactor beds via lines 103A, 103B, 103C and 103D and OLHSV values are from about 0.01 hr'to about 10 hr ¹ , such as about 0.02 hr Ho about 1 hr ¹ , or such as about 0.03 hr Ho about 0.1 hr f The split introduction of the olefin feed allows a lower concentration (one quarter) of the olefin feed to be introduced locally to each of the first stage 101 A, second stage 101B, third stage 101C, fourth stage 101D, fifth stage 101E, sixth stage 101F, seventh stage 101G, and eighth stage 101H. The isoparaffin feed is introduced to alkylation reactor 100D through line 105. Alkylation reactor 100D has multiple interstage spaces: first interstage space 109A, second interstage space 109B, third interstage space 109C, fourth interstage space 109D, fifth interstage space 109E, sixth interstage space 109F, and seventh interstage space 109G between reactor stages 101A, 101B, 101 C, 101 D, 101 E, 101 F, 101 G, and 101H to allow for introduction of additional olefin feed through lines 103B, 103C, 103D, 103E, 103F, 103G, and 103H. Reactor stages 101 A, 101B, 101C, 101 D, 101 E, 101 F, 101 G, and 101H may individually include an upstream zone and a downstream zone. For example, in reactor stage 101 A a downstream zone is close to interstage space 109A and may include hydrogenation co-catalyst. Furthermore, the downstream zone may simply be a thin layer of hydrogenation catalyst packed downstream in a fixed bed. Upstream and downstream zones within a reactor stage are not pictured because there may not be a distinct division between and upstream zone and a downstream zone, for example, in at least one embodiment, where hydrogenation co-catalyst is supported or bound with catalyst.

[0036] If lines 103 and 105 have the same composition as in FIG. 1A, then the local i:o ratio is 8 times that found in FIG. 1A, because the olefin feed is divided into eight lines 103 A, 103B, 103C, 103D, 103E, 103F, 103G, and 103H. The i:o ratio in a single stage is only slightly affected by prior stage(s) because the olefin introduced to a prior stage can be largely converted within that stage, but the isoparaffin is introduced at such a ratio that the amount converted may have little effect on the ratio in later stages. For example, the olefin introduced via line 103A to first stage 101 A is converted, such as about 90 wt% or greater, about 95 wt% or greater, about 98 wt% or greater, or about 99 wt% or greater is converted in first stage 101 A, based on the total weight of olefin in the olefin feed introduced via line 103 A. Additionally, only a small portion of the isoparaffin feed may be converted by the reaction in first stage 101 A, such as about 10 wt% or less of the isoparaffin feed is converted based on the total weight of isoparaffin introduced to first stage

IOIA, such as about 5 wt% or less, about 2 wt% or less, about 1 wt% or less, about 0.5 wt% or less, or about 0.1 wt% or less. Therefore, the amount of isoparaffin introduced to interstage 109A and, therefore, introduced to second stage 101B is similar or slightly less than that introduced to first stage 101A and the olefin introduced to interstage space 109A via line 103B and from the effluent of first stage 101A serves to bring the olefin level back up to a desired i:o ratio. Therefore, for example, if an i:o ratio of 100:1 is introduced to first stage 101 A and there is an olefin conversion of 100% then the i:o ratio in second stage 101B would be -99:1, if no additional isoparaffin was added. The combination of isoparaffin and olefin is then introduced to second stage

IOIB, where the olefin may be converted in second stage 101B, such as about 90 wt% or greater, about 95 wt% or greater, about 98 wt% or greater, or about 99 wt% or greater is converted in second stage 101B, based on the total weight of olefin introduced to interstage space 109A. Similarly, only a small portion of the isoparaffin feed may be converted by the reaction in second stage 101B, such as about 10 wt% or less of the isoparaffin feed is converted based on the total weight of isoparaffin introduced to interstage space 109A, such as about 5 wt% or less, about 2 wt% or less, about 1 wt% or less, about 0.5 wt% or less, or about 0.1 wt% or less. As the isoparaffin feed enters additional interstage spaces (such as second interstage space 109B, and third interstage space 109C) more olefin may be introduced (via lines 103C, 103D, 103E, 103F, 103G, and 103H) to adjust the i:o ratio as the combined feeds are introduced to additional stages (such as third stage

IOIC, fourth interstage space 109D, fifth interstage space 109E, sixth interstage space 109F, and seventh interstage space 109G). Additionally, because the selected isoparaffin may be consumed in each stage, such as in amounts of about 10 wt% or less, about 5 wt% or less, about 2 wt% or less, about 1 wt% or less, about 0.5 wt% or less, or about 0.1 wt% or less, additional isoparaffin may be added in an interstage space so as to maintain a consistent i:o ratio throughout the multistage reactor. Similarly to FIG. 1A, an alkylation product mixture exits the reactor through line 107.

Feedstocks

[0037] Feedstocks useful in the present alkylation process include at least one isoparaffin feed, at least one olefin feed, and a hydrogen feed. The isoparaffin feed used in alkylation processes of the present disclosure may have from about 4 to about 7 carbon atoms. Representative examples of such isoparaffins include isobutane, isopentane, 3-methylhexane, 2-methylhexane, 2,3- dimethylbutane, and mixture(s) thereof, typically isobutane.

[0038] The olefin component of the feedstock may include at least one olefin having from 2 to 12 carbon atoms. In some embodiments, the olefin feed comprises one or more C2 to C5 olefins, such as one or more C3 to C5 olefins. Representative examples of such olefins include ethylene, propylene, 1 -butene, 2-butene, isobutylene, 1-pentene, 2-pentene, 3-pentene, 2-methy 1-1 -butene, 3-methyl- 1 -butene, 2-methyl-2-butene, hexene, octene, heptene, or mixture(s) thereof. In some embodiments, the olefin component of the feedstock is selected from the group consisting of propene, 1 -butene, 2-butene, isobutylene, 1-pentene, 2-pentene, 3-pentene, 2-methy 1-1 -butene, 3- methyl-1 -butene, 2-methy 1-2-butene, and mixture(s) thereof. For example, in one embodiment, the olefin component of the feedstock may include a mixture of propylene and at least one butene, such as 2-butene, where the weight ratio of propylene to butene is from about 0.01:1 to about 1.5:1, such as from about 0.1:1 to about 1:1. In another embodiment, the olefin component of the feedstock may include a mixture of propylene and at least one pentene, where the weight ratio of propylene to pentene is from about 0.01:1 to about 1.5:1, such as from about 0.1:1 to about 1:1. [0039] The concentration of olefin feed can be adjusted by, e.g., staged additions thereof. By staged additions, isoparaffin/olefin feed concentrations (and therefore the i:o ratio) can be maintained at levels to improve conversion and reduce catalyst deactivation. In at least one embodiment, the ratio of isoparaffin to olefin ratio by volume, referred to as the i:o ratio is: about 100:1 or greater, about 120:1 or greater, about 140:1 or greater, about 160:1 or greater, about 180:1 or greater, about 200:1 or greater, about 220:1 or greater, about 240:1 or greater, about 260:1 or greater, about 280: 1 or greater, or about 300: 1 or greater, such as from about 100: 1 to about 500: 1, about 120:1 to about 500:1, about 160:1 to about 480:1, about 200:1 to about 450:1, about 220:1 to about 450: 1, about 240: 1 to about 420: 1, or about 240: 1 to about 400: 1. [0040] The production of olefin oligomers increases with lower i:o ratios. To reduce or eliminate the production of olefin oligomers an i:o ratio of about 100: 1 or greater may be used. On the other hand, the efficiency of the alkylation process can be reduced at higher i:o ratios, due to large quantity of isoparaffin present in the alkylation product mixture, which is then separated and recycled to the reactor. The separation and recycling of isoparaffin may occur in a distillation apparatus that allows for distillation of low C5- alkane from C6+ alkanes and alkenes produced in the reactor. A higher i:o ratio can provide greater quantities of C5- alkane separated from the alkylation product mixture that can be recycled to the reactor.

[0041] The hydrogen feed includes hydrogen and may include inert gases to decrease hydrogen concentration within the feed. The concentration of hydrogen in the multistage reactor can be adjusted by, e.g., staged additions thereof. By staged additions, hydrogen/isoparaffin feed concentrations can be maintained at levels sufficient to reduce or eliminate olefin oligomers formed in a stage of the multistage reactor. In at least one embodiment, the molar ratio of hydrogen to isoparaffin is from about 1:100 to about 1:1, about 1:50 to about 1:2, or about 1:10 to about 1:5. [0042] Before being sent to the reactor, the isoparaffin feed, the olefin feed, and/or the hydrogen feed may be treated to remove catalyst poisons. For example, catalyst poisons may be removed using guard beds with specific absorbents for reducing the level of S, N, and/or oxygenates to values which do not affect catalyst stability, activity, and selectivity.

Catalysts

[0043] One class of catalysts suitable for use in a process of this disclosure is a molecular sieve or zeolite. The molecular sieve may have a Constraint Index of about 5 or less, and may be a crystalline microporous material of the MWW framework type. MWW framework type refers to a type of crystalline microporous material that includes at least two independent sets of 10- membered ring channels and has composite building units of d6r (t-hpr) and mel as defined and discussed in Compendium of Zeolite Framework Types. Building Schemes and Type Characteristics Van Koningsveld, Henk, (Elsevier, Amsterdam, 2007), incorporated by reference. Crystalline microporous materials of the MWW framework type can include those molecular sieves having an X-ray diffraction pattern comprising d-spacing maxima at 12.4±0.25, 6.9±0.15, 3.57±0.07 and 3.42±0.07 Angstrom. The X-ray diffraction data used to characterize the material are obtained by standard techniques using the K-alpha doublet of copper as incident radiation and a diffractometer equipped with a scintillation counter and associated computer as the collection system. Crystalline microporous materials of the MWW framework type include molecular sieves having natural tiling units of t-dac-1, t-euo, t-hpr, t-kah, t-kzd, t-mel, t-mww-1, t-mww-2, and t-srs as defined and discussed in Three -periodic Nets and Tilings: Natural Tilings for Nets, V. A. Blatov, O. Delgado-Friedrichs, M. O'Keeffe and D. M. Proserpio, Acta Crystallogr. A 63, 418-425 (2007), incorporated by reference.

[0044] In at least one embodiment, the crystalline microporous material is a zeolite. As used herein, the term “crystalline microporous material of the MWW framework type” includes one or more of:

(a) molecular sieves made from a common first degree crystalline building block unit cell, which unit cell has the MWW framework topology. (A unit cell is a spatial arrangement of atoms which if tiled in three-dimensional space describes the crystal structure. Such crystal structures are discussed in the “Atlas of Zeolite Framework Types”, Fifth edition, 2001, incorporated herein by reference);

(b) molecular sieves made from a second degree building block, being a 2-dimensional tiling of such MWW framework topology unit cells, forming a monolayer of one unit cell thickness, in one embodiment, one c-unit cell thickness;

(c) molecular sieves made from common second degree building blocks, being layers of one or more than one unit cell thickness, where the layer of more than one unit cell thickness is made from stacking, packing, or binding at least two monolayers of MWW framework topology unit cells. The stacking of such second degree building blocks can be in a regular fashion, an irregular fashion, a random fashion, or any combination thereof; and

(d) molecular sieves made by any regular or random 2-dimensional or 3-dimensional combination of unit cells having the MWW framework topology.

[0045] Examples of crystalline microporous materials of the MWW framework type include MCM-22 (U.S. Patent No. 4,954,325), PSH-3 (U.S. Patent No. 4,439,409), SSZ-25 (U.S. Patent No. 4,826,667), ERB-1 (European Patent No. 0293032), ITQ-1 (U.S. Patent No. 6,077,498), ITQ- 2 (International Publication No. WO97/17290), MCM-36 (U.S. Patent No. 5,250,277), MCM-49 (U.S. Patent No. 5,236,575), MCM-56 (U.S. Patent No. 5,362,697), UZM-8 (U.S. Patent No. 6,756,030), UZM-8HS (U.S. Patent No. 7,713,513), UZM-37 (U.S. Patent No. 7,982,084), EMM- 10 (U.S. Patent No. 7,842,277), EMM-12 (U.S. Patent No. 8,704,025), EMM-13 (U.S. Patent No. 8,704,023), UCB-3 (U. S. Patent No. 9,790, 143B2), MIT-1 (Luo, et. al, Chem Sci. 2015 November 1; 6(11): 6320-6324), and mixtures thereof.

[0046] In some embodiments, the crystalline microporous material of the MWW framework type may be contaminated with other crystalline materials, such as ferrierite or quartz. These contaminants may be present in quantities of about 10 wt% or less, such as about 5 wt% or less. In some embodiments, the crystalline microporous material of the MWW framework type employed may be an aluminosilicate material having a silica to alumina molar ratio of about 10 or more, such as from about 10 to about 50.

Hydrogenation Co-Catalyst [0047] Any suitable hydrogenation catalyst may be used as a co-catalyst, including noble metals, such as Pd, Pt, Rh, Ru, Ir, Os, Ag, Au; or non-noble metals, such as Mo, Co, Ni, Fe; or combination(s) thereof, such as the combination of two noble metals, two non-noble metals, or a combination of noble and non-noble metals.

[0048] The co-catalyst may be supported, suitable support materials may include clay, alumina, silica, titania, zirconia, aluminosilicates, zeolites, carbon, or combination(s) thereof. In some embodiments, the silica support can be an amorphous silica support. In other embodiments, the support can be a mesoporous crystalline or semi-crystalline support material. Examples of mesoporous silica materials suitable for use as a support can include zeolites, such as MCM-41, other M41S structures, or SBA-15. Additionally, a silica support may be modified with alumina and the amount of alumina added to modify a silica support may vary. For example, the amount of alumina added to a silica support can be about 0.3 wt% to about 3.0 wt%, about 0.5 wt% to about 2.5 wt%, about 0.5 wt% to about 1.8 wt%, about 0.75 wt% to about 1.6 wt%, about 1.0 wt% to about 1.5 wt%, about 1.1 wt% to about 1.5 wt%, or about 1.25 wt% to about 1.5 wt%. In some embodiments, the amount of alumina added to a silica support can be about 0.3 wt% to about 2.5 wt%, or about 1.0 wt% to about 2.5 wt%, or about 1.1 wt% to about 2.2 wt%.

[0049] When one metal is present, the amount of the one metal can be about 0.05 wt% or more based on the total weight of the co-catalyst and support material (if any), for example about 0.1 wt% or more, or about 0.2 wt% or more, or about 0.5 wt% or more. Additionally, when one metal is present, the amount of that hydrogenation metal can be about 5.0 wt% or less based on the total weight of the co-catalyst and support material (if any), for example about 3.5 wt% or less, about 2.5 wt% or less, about 2.0 wt% or less, about 1.5 wt% or less, about 1.0 wt% or less, about 0.9 wt% or less, about 0.75 wt% or less, or about 0.6 wt% or less. For example, the amount of hydrogenation metal can be about 0.05 wt% to about 5.0 wt%, or about 0.1 wt% to about 2.5 wt%, or about 0.1 wt% to about 2.0 wt%, or about 0.1 wt% to about 1.5 wt%, or about 0.2 wt% to about 5.0 wt%, or about 0.2 wt% to about 2.5 wt%, or about 0.2 wt% to about 1.5 wt%, or about 0.5 wt% to about 5.0 wt%, or about 0.5 wt% to about 2.5 wt%, or about 0.5 wt% to about 1.5 wt% based on the total weight of the co-catalyst and support material (if any). In some embodiments, the amount of hydrogenation metal can be about 0.05 wt% to about 5.0 wt%, or about 0.1 wt% to about 2.0 wt%, or about 0.2 wt% to about 1.0 wt% based on the total weight of the co-catalyst and support material (if any). [0050] When more than one hydrogenation metal is present, the collective amount of hydrogenation metals can be about 0.05 wt% or more based on the total weight of the co-catalyst, such as about 0.1 wt% or more, about 0.2 wt% or more, about 0.3 wt% or more, about 0.4 wt% or more, or about 0.5 wt% or more based on the total weight of the co-catalyst and support material (if any). Additionally or alternately, when more than one hydrogenation metal is present, the collective amount of hydrogenation metals can be about 5.0 wt% or less based on the total weight of the co-catalyst and support material (if any), for example about 3.5 wt% or less, about 2.5 wt% or less, about 1.5 wt% or less, about 1.0 wt% or less, about 0.9 wt% or less, about 0.75 wt% or less, or about 0.6 wt% or less. For example, the combined amount of metal(s) can be about 0.05 wt% to about 5.0 wt%, or about 0.05 wt% to about 2.5 wt%, or about 0.05 wt% to about 2.0 wt%, or about 0.05 wt% to about 1.5 wt%, or about 0.1 wt% to about 5.0 wt%, or about 0.1 wt% to about 2.5 wt%, or about 0.1 wt% to about 1.5 wt%, or about 0.2 wt% to about 5.0 wt%, or about 0.2 wt% to about 2.5 wt%, or about 0.2 wt% to about 1.5 wt% based on the total weight of the co-catalyst and support material (if any). In some embodiments, the amount of metal(s) can be about 0.05 wt% to about 5.0 wt%, or about 0.1 wt% to about 2.5 wt%, or about 0.2 wt% to about 5.0 wt% based on the total weight of the co-catalyst and support material (if any). For embodiments where both Pt and Pd are present as the hydrogenation metals, the ratio of Pt to Pd can be from about 1:3 to about 4:1, or from about 1 :4 to about 3: 1, or from about 1 :2 to about 4: 1, or from about 1 :2 to about 3:1. The amounts of metal(s) may be measured by methods specified by ASTM for individual metals, including but not limited to atomic absorption spectroscopy (AAS).

[0051] Quantities of catalyst and co-catalyst might not be equal in weight, for example, the weight ratio of co-catalyst to catalyst may be from about 1 : 1000 to about 1:1, about 1 : 500 to about 1:2, about 1:100 to about 1:5, about 1:100 to about 1:10, or about 1:50 to about 1:10.

Binder

[0052] Catalysts and co-catalysts suitable for use in the systems and processes described include a binder.

[0053] Binder materials may include inorganic oxides, such as alumina, silica, titania, zirconia and mixtures and compounds thereof, may be present in the catalyst in amounts about 90 wt% or less, for example about 80 wt% or less, such as about 70 wt% or less, for example about 60 wt% or less, such as about 50 wt% or less. Where a non-alumina binder is present, the amount employed may be as little as about 1 wt%, such as about 5 wt% or more, for example about 10 wt% or more. In at least one embodiment, a silica binder is employed such as disclosed in U.S. Pat. No. 5,053,374, incorporated by reference. In other embodiments, a zirconia or titania binder is used. [0054] In other embodiments, the binder may be a crystalline oxide material such as the zeolite- bound-zeolites described in U.S. Pat. Nos. 5,665,325 and 5,993,642, incorporated by reference. In the case of crystalline binders, the binder material may contain alumina, including amorphous alumina.

Product

[0055] The product of the alkylation reaction (also referred to as the alkylation product mixture) can include: alkanes resulting from the alkylation of isoparaffin with olefin, unreacted isoparaffin, unreacted olefin, olefin oligomers, other byproducts, including other alkanes and alkenes. The product composition of the isoparaffin-olefin alkylation reaction described is dependent on the reaction conditions and the composition of the olefin feed and isoparaffin feed. The product is a complex mixture of hydrocarbons, since alkylation of the feed isoparaffin by the feed olefin is accompanied by a variety of competing reactions including cracking, olefin oligomerization, and/or further alkylation of the alkylate product by the feed olefin. For example, in the case of alkylation of isobutane with C3-C5 olefins, such as 2-butene, the product may include about 20-30 wt% of C5-C7 hydrocarbons, 50-75 wt% of C8 hydrocarbons and 2.5-20 wt% of C9+ hydrocarbons. Moreover, using an MWW type molecular sieve as the catalyst, it has been discovered that processes can be selective to desirable high octane components so that, in the case of alkylation of isobutane with C3-C5 olefins, the C6 fraction typically includes at least 40 wt%, such as at least 70 wt%, of 2,3-dimethylbutane, the C7 fraction typically includes at least 40 wt%, such as at least 80 wt%, of 2,3-dimethylpentane and the C8 fraction typically includes at least 50 wt%, such as at least 70 wt%, of 2,3,4-trimethylpentane, 2,3,3-trimethylpentane, and 2,2,4- trimethylpentane.

[0056] Additionally, in the case of alkylation of isobutane with C5 olefins, such as n-pentene and 2-methyl-2-butene, the product may include about 30-40 wt% of C5 hydrocarbons, 15-25 wt% of C9 hydrocarbons, 25-35 wt% of C8 hydrocarbons, and 2.5-10 wt% of C10+ hydrocarbons. Moreover, using an MWW type molecular sieve as the catalyst, it has been found that a process can be selective to desirable high octane components so that, in the case of alkylation of isobutane with C5 olefins, the C8 and C9 fractions typically include a higher molar ratio of trimethyl isomers to dimethyl isomers, which is beneficial for increasing octane. For the C8 fraction, the molar ratio of trimethylpentane to dimethylhexane can be about 3 or more, e.g. about 4 to about 5, or about 3 to about 6. For the C9 fraction, the molar ratio of trimethylhexane to dimethylheptane can be about 1 or more, e.g. about 1.5 or more, or from about 1 to about 3.

[0057] The product of the isoparaffin-olefin alkylation reaction may be fed to a separation system, such as a distillation train, to recover a C5+ fraction for use as a gasoline octane enhancer. Additionally, the separation system may separate the C4-C6 isoparaffin to be recycled as part or all of the isoparaffin feed. Furthermore, depending on alkylate demand, part or all of a C9+ fraction can be recovered for use as a distillate blending stock.

Other Embodiments of the Present Disclosure [0058] Clause 1. A process for the alkylation of an isoparaffin, the process including: introducing, in a multistage reactor, an isoparaffin feed, an olefin feed, and a hydrogen feed to a solid acid catalyst and a hydrogenation co-catalyst, where the solid acid catalyst includes a zeolite.

[0059] Clause 2. The process of clause 1, where the zeolite is a crystalline microporous material of the MWW framework type.

[0060] Clause 3. The process of clause 2, wherein the crystalline microporous material of the MWW framework type is selected from the group consisting of MCM-22, PSH-3, SSZ-25, ERB- 1, ITQ-1, ITQ-2, MCM-36, MCM-49, MCM-56, EMM-10, EMM-12, EMM-13, UZM-8, UZM- 8HS, UZM-37, UCB-3, or mixture(s) thereof. [0061] Clause 4. The process of clause 3, wherein the crystalline microporous material of the

MWW framework type is selected from the group consisting of MCM-22, MCM-49, MCM-56, EMM- 10, or mixture(s) thereof.

[0062] Clause 5. The process of clause 1, where a stage of the multistage reactor includes an upstream zone and a downstream zone. [0063] Clause 6. The process of clause 5, where the upstream zone includes the solid acid catalyst.

[0064] Clause 7. The process of clause 5, where the downstream zone includes the hydrogenation co-catalyst.

[0065] Clause 8. The process of clause 1, where the molar ratio of the hydrogen feed to isoparaffin feed is about 1 : 100 to about 1:1.

[0066] Clause 9. The process of clause 8, where the molar ratio of the hydrogen feed to isoparaffin feed is about 1:10 to about 1:5.

[0067] Clause 10. The process of clause 1, where the hydrogenation co-catalyst includes Pd, Pt, Rh, Ru, Ir, Co, Ni, Fe, or a combination thereof. [0068] Clause 11. The process of clause 10, where the hydrogenation co-catalyst includes Pd,

Pt, or a combination thereof.

[0069] Clause 12. The process of clause 1, where the hydrogenation co-catalyst includes a support.

[0070] Clause 13. The process of clause 12, where the support includes a zeolite. [0071] Clause 14. The process of clause 12, where the support includes an alumina modified silica.

[0072] Clause 15. The process of clause 1, where the weight ratio of the hydrogenation co catalyst to the solid acid catalyst is from about 1 : 1000 to about 1:1. [0073] Clause 16. A process for the alkylation of an isoparaffin, the process including: introducing, in a multistage reactor, an isoparaffin feed, an olefin feed, and a hydrogen feed to a solid acid catalyst and a hydrogenation co-catalyst, where the solid acid catalyst includes a crystalline microporous material of the MWW framework type and the hydrogenation co-catalyst includes a noble metal. [0074] Clause 17. The process of clause 16, where the hydrogenation co catalyst further includes a support.

[0075] Clause 18. The process of clause 16, where the crystalline microporous material of the MWW framework type is selected from the group consisting of MCM-22, PSH-3, SSZ-25, ERB- 1, ITQ-1, ITQ-2, MCM-36, MCM-49, MCM-56, EMM-10, EMM-12, EMM-13, UZM-8, UZM- 8HS, UZM-37, UCB-3, or mixture(s) thereof.

[0076] Clause 19. The process of clause 18, wherein the crystalline microporous material of the MWW framework type is selected from the group consisting of MCM-22, MCM-49, MCM- 56, EMM- 10, or mixture(s) thereof.

[0077] Clause 20. A catalyst bed for the alkylation of an isoparaffin with an olefin and hydrogenation of olefin, the catalyst bed having a volume and including: an upstream zone including a zeolite, and a downstream zone including a hydrogenation co-catalyst.

[0078] Clause 21. The catalyst bed of clause 20, where the upstream zone is 70% or more of the volume of the catalyst bed.

[0079] Clause 22. The catalyst bed of clause 20, where the upstream zone is substantially free of noble metals.

Examples Feed Pretreatment

[0080] Isobutane was obtained from a commercial source and used as received. The isobutene purity was 99.6% with the balance n-butane. [0081] Propylene and 2-butene were obtained from a commercial specialty gases source and used as received. The 2-butene was a mixture of trans-2-butene and cis-2-butene.

Catalyst Preparation and Loading

[0082] Catalysts used for isobutene alkylation with light olefins are dried in the reactor under nitrogen flow at 250 °C for at least 4 hours prior to use. Example 1

[0083] The catalyst was prepared by combining 80 parts MCM-49 zeolite crystals with 20 parts pseudoboehmite alumina, on a calcined dry weight basis. The MCM-49 and pseudoboehmite alumina dry powder were placed in a muller or a mixer and mixed for 30 minutes. Sufficient water was added to the MCM-49 and alumina during the mixing process to produce an extrudable paste. The extrudable paste was formed into a 1/20 inch quadralobe extrudate using an extruder. After extrusion, the extrudate was dried at a temperature ranging from 250 °F (121 °C) to 325 °F (168 °C). After drying, the dried extrudate was heated to 1000 °F (538 °C) under flowing nitrogen. The extrudate was then cooled to ambient temperature, humidified with saturated air or steam and then ion exchanged with 0.75 N ammonium nitrate solution followed by washing with deionized water and drying. The extrudate was then calcined in a nitrogen/air mixture to a temperature of 1000 °F (538 °C).

Example 2

[0084] 95 parts MCM-49 zeolite crystals were combined with 5 parts pseudoboehmite alumina, on a calcined dry weight basis. The MCM-49 and pseudoboehmite alumina dry powder were placed in a muller or a mixer and mixed for 30 minutes. Sufficient water was added to the MCM-49 and alumina during the mixing process to produce an extrudable paste. The extrudable paste was formed into a 1/20 inch quadralobe extrudate using an extruder. After extrusion, the extrudate was dried at a temperature ranging from 250 °F (121 °C) to 325 °F (168 °C). After drying, the dried extrudate was heated to 1000 °F (538 °C) under flowing nitrogen. The extrudate was then cooled to ambient temperature, humidified with saturated air or steam and then ion exchanged with 0.75 N ammonium nitrate solution followed by washing with deionized water and drying. The extrudate was then calcined in a nitrogen/air mixture to a temperature of 1000 °F (538 °C). Example 3

[0085] The catalyst of Example 2 was impregnated with dilute palladium tetraamine nitrate solution at incipient wetness to achieve a target Pd level of 0.15%. After drying the impregnated extrudate at 250 °F (121 °C), the catalyst is calcined in air at 662 °F (350 °C) for 1 hour.

Example 4 [0086] The catalyst of Example 2 was impregnated with dilute platinum tetraamine nitrate solution at incipient wetness to achieve a target Pt level of 0.1%. After drying the impregnated extrudate at 250 °F (121 °C), the catalyst is calcined in air at 662 °F (350 °C) for 1 hour Example 5 [0087] The catalyst of Example 1 was loaded into a pilot plant and operated as a single stage reactor, as shown in FIG. 1A. The reactor was 60” long and made from ¾” O.D. Schedule 40 pipe. The reactor was loaded with 50 g of catalyst. The reactor was located in an isothermal sand bath maintained at 302 °F (150 °C). Reactor pressure was 750 psig (5171 kPag). Isobutane (99.6% purity) and 2- butene were independently fed to the top of the single stage reactor at a relative rate such that the isobutane to 2-butene ratio at the top of the catalyst bed was 40: 1. The reactor effluent was measured using a FID GC equipped with a 150 m Petrocol column. The 2-butene flow to the reactor was set to achieve an Olefin Liquid Hourly Space Velocity (OLHSV) of 0.06 h ¹ and subsequently 0.03 h ¹ . Isobutane flowrates were adjusted as olefin flowrates were adjusted to maintain a constant i:o of 40: 1 to the inlet to the catalyst bed. Average 2-butene conversion at 0.06 h ¹ was 80.7% and at 0.03 h ¹ the average 2-butene conversion was 93.7%.

[0088] The catalyst of Example 1 was loaded into a pilot plant and operated as a 2 stage reactor, as shown in FIG. IB. Each stage was 60” long and made from ¾” O.D. Schedule 40 pipe. Each stage was loaded with 50 g of catalyst. The two-stage reactor was located in an isothermal sand bath maintained at 302 °F (150 °C). Reactor pressure was 750 psig (5171 kPag). Isobutane (99.6% purity) was fed to the first stage of the reactor and the 2-butene flow was split evenly into 2 using Coriolis meters and independently fed to each stage. The relative rates of isobutane and 2-butene were set such that the isobutane to 2-butene ratio at the top of the first stage was 40:1. The alkylation product mixture exiting the reactor was measured using a FID GC equipped with a 150 m Petrocol column. The total 2-butene flow to the reactor was set to achieve an Olefin Liquid Hourly Space Velocity (OLHSV) of 0.06 h ¹ . Average 2-butene conversion at 0.06 h ¹ was 70.4%. As can be seen when comparing the results to Example 5, the operation of the 2 stage system resulted in significantly lower olefin conversion.

[0089] The catalyst of Example 1 was loaded into a pilot plant and operated as a 4-stage reactor, as shown in FIG. 1C. Each stage was 60” long and made from ¾” O.D. Schedule 40 pipe. Each stage was loaded with 50 g of catalyst. The four-stage reactor was located in an isothermal sand bath maintained at 302 °F (150 °C). Reactor pressure was 750 psig (5171 kPag). Isobutane (99.6% purity) was fed to the first stage and the 2-butene flow was split evenly into 4 using Coriolis meters and independently fed to each stage of the four-stage reactor. The relative rates of isobutane and 2-butene were set such that the isobutane to 2-butene ratio at the top of the first stage was 40: 1. The alkylation product mixture exiting the reactor was measured using a FID GC equipped with a 150 m Petrocol column. The total 2-butene flow to the reactor was set to achieve an Olefin Liquid Hourly Space Velocity (OLHSV) of 0.06 h ¹ and subsequently 0.03 h ¹ . Isobutane flowrates were adjusted as olefin flowrates were adjusted to maintain a constant i:o of 40: 1 to the inlet to the first stage. Average 2-butene conversion at 0.06 h ¹ was 61.4% and at 0.03 h ¹ the average 2-butene conversion was 77.5%. As can be seen when comparing the results to Example 5, the operation of the 4 bed system resulted in significantly lower olefin conversion.

[0090] The catalyst of Example 2 was loaded into a pilot plant and operated as a 4-stage reactor, as shown in FIG. 1C. Each stage was 60” long and made from ¾” O.D. Schedule 40 pipe. Each stage was loaded with 150 g of catalyst. The reactor was located in an isothermal sand bath maintained at 302 °F (150 °C). Reactor pressure was 750 psig (5171 kPag). Isobutane (99.6% purity) was fed to the first stage and the 2-butene flow was split evenly into 4 using Coriolis meters and independently fed to each reactor bed. The relative rates of isobutane and 2-butene were set such that the isobutane to 2-butene ratio at the top of the first stage was 40: 1. The alkylation product mixture exiting the multistage reactor was measured using a FID GC equipped with a 150 m Petrocol column. The total 2-butene flow to the reactor was set to achieve an Olefin Liquid Hourly Space Velocity (OLHSV) of 0.03 h ¹ . Average 2-butene conversion at 0.03 h ¹ was about 82%. The alkylation product mixture exiting the multistage reactor was sent to a distillation column for separation of C4 and lighter components from the reaction product. The alkylation product mixture was analyzed via offline GC and shown to have about 13.9% C5+ olefins.

[0091] A sample of the alkylation product mixture produced in Example 7 was hydrogenated using a commercial MaxSat™ hydrogenation catalyst available from ExxonMobil Catalyst & Licensing. Hydrogenation took place in a batch reactor at 200 °C and 800 psig (5171 kPag) for 8 hours. The hydrogenated alkylate was analyzed by GC and shown to have <1% C5+ olefins.

[0092] The catalyst of Example 2 was loaded into a pilot plant single stage reactor, as shown in FIG. 1A. The reactor was 14” long and made from 3/8” O.D. stainless steel tubing. The reactor was loaded with 4 g of catalyst. The reactor was located in an electrically heated furnace and maintained at 302 °F (150 °C). Reactor pressure was 750 psig (5171 kPag). A pre-mixed gas blend with isobutane and 2-butene at a 40: 1 ratio was fed to the top of the reactor. The alkylation product mixture was analyzed using a FID GC equipped with a 150 m Petrocol column. The flow to the reactor was set to achieve an Olefin Liquid Hourly Space Velocity (OLHSV) of 0.057 h ¹ . Average 2-butene conversion at 0.057 h ¹ was 99.7%

Example 11 [0093] The catalyst of Example 2 was loaded into a pilot plant single stage reactor, as shown in FIG. 1A. The reactor was 14” long and made from 3/8” O.D. stainless steel tubing. The reactor was loaded with 4 g of catalyst. The reactor was located in an electrically heated furnace and maintained at 302 °F (150 °C). Reactor pressure was 750 psig (5171 kPag). A pre-mixed gas blend with isobutane and 2-butene at a 40: 1 ratio was fed to the top of the reactor. The alkylation product mixture was analyzed using a FID GC equipped with a 150 m Petrocol column. The flow to the reactor was set to achieve an Olefin Liquid Hourly Space Velocity (OLHSV) of 0.057 h ¹ . To simulate the operation of a 5 -stage multistage reactor configuration, alkylate produced in the 4- stage reactor from Example 8 was co-fed at a rate of 3.25 cc/hr. Average 2-butene conversion at 0.057 h ¹ was 58.1% at 10 days of co-feeding and continued to drop with days on stream. As demonstrated by this example, the presence of about 2% C5+ olefins in the feed caused the 2- butene conversion to decrease by about 41.6%.

[0094] The catalyst of Example 2 was loaded into a pilot plant single stage reactor , as shown in FIG. 1A. The reactor was 14” long and made from 3/8” O.D. stainless steel tubing. The reactor was loaded with 4 g of catalyst. The reactor was located in an electrically heated furnace and maintained at 302 °F (150 °C). Reactor pressure was 750 psig (5171 kPag). A pre-mixed gas blend with isobutane and 2-butene at a 40:1 ratio was fed to the top of the reactor bed. The alkylation product mixture was measured using a FID GC equipped with a 150 m Petrocol column. The flow to the reactor was set to achieve an Olefin Liquid Hourly Space Velocity (OLHSV) of 0.057 h ¹ . To simulate the operation of a 5-stage multistage reactor configuration where the heavier olefin content has been reduced by hydrogenation, the hydrogenated alkylate prepared in Example 9 was co-fed at a rate of 3.25 cc/hr. Average 2-butene conversion at 0.057 h ¹ was 98.8%. As demonstrated by this example, the presence of about 0.15% C5+ olefins in the feed caused the 2- butene conversion to decrease by less than 1% vs. the reference case in Example 10.

[0095] The catalysts of Examples 2, 3 and 4 were loaded into parallel microunits as dual catalyst beds as indicated in the Table 1 below. The catalysts were loaded such the Bed 1 was the first bed to see the feed. The catalysts were diluted with quartz at a level such that the volume of quartz diluent was 0.85 times the catalyst volume. The dual bed catalyst system was operated as a single reactor as shown in Figure 1 (a). The reactors were located in electrically heated furnaces and maintained at 302°F (150 °C). Reactor pressure was 700 psig (4826 kPag). A pre-mixed gas blend with isobutane and 2-butene at a 40:1 ratio was fed to the top of each reactor bed. Each reactor effluent was measured using a FID GC equipped with a 60 m DB-1 column. The flow to the reactor was set to achieve an Olefin Liquid Hourly Space Velocity (OLHSV) of 0.12 h ¹ .

Table 1 [0096] After 244 hours after the start of the feed to the system, a hydrogen co-feed was started at 0.11 moles of hydrogen per mole of C4 hydrocarbon fed to the system. Analysis of the reactor effluent for C8 olefins with and without hydrogen is shown in the Table 2 below.

Table 2 [0097] As the data in the table above shows, the presence of the hydrogenating metal on the alkylation catalyst leads to the lowering of the C8 olefin concentration, likely due to saturation of the C8 olefin, when ¾ is flowing. Increasing the amount of catalyst at the bottom of the reactor bed containing a hydrogenation functionality leads to lower amounts of C8 olefins exiting the reactor. As shown in examples above, the production of undesired olefins in the alkylation reactor leads to a nonlinear increase of heavier fraction in the alkylate products in the subsequent stages and leads to catalyst deactivation. The examples in this disclosure show the reduction in undesired heavies as well as an increase in the desired products can be achieved by including hydrogenation co-catalyst downstream of the catalyst in a catalyst bed.

[0098] Overall, it has been discovered that certain byproducts, including olefin oligomers, produced during the alkylation of isoparaffins with olefins in a multistage reactor may decrease catalyst activity and reduce conversion of olefins. It has also been discovered that the produced olefin oligomers may be reduced or eliminated by hydrogenation in the downstream portion of a catalyst bed within a reactor stage. A multistage reactor may provide greater conversion and production rates and decrease overall costs of production. The combination of using a reactor having two or more stages, one or more of which may include an upstream and a downstream zone, and hydrogenation at the downstream zone within a reactor stage may provide reduced olefin oligomers, increased olefin conversion, increased production, decreased catalyst deactivation, and/or improved product selectivity, as compared to conventional alkylation processes or multistage alkylation processes in the absence of a hydrogenation catalyst.

[0099] The phrases, unless otherwise specified, "consists essentially of and "consisting essentially of do not exclude the presence of other steps, elements, or materials, whether or not, specifically mentioned in this specification, so long as such steps, elements, or materials, do not affect the basic and novel characteristics of this disclosure, additionally, they do not exclude impurities and variances normally associated with the elements and materials used.

[0100] For the sake of brevity, only certain ranges are explicitly disclosed herein. However, ranges from any lower limit may be combined with any upper limit to recite a range not explicitly recited, as well as, ranges from any lower limit may be combined with any other lower limit to recite a range not explicitly recited, in the same way, ranges from any upper limit may be combined with any other upper limit to recite a range not explicitly recited. Additionally, within a range includes every point or individual value between its end points even though not explicitly recited. Thus, every point or individual value may serve as its own lower or upper limit combined with any other point or individual value or any other lower or upper limit, to recite a range not explicitly recited.

[0101] All documents described herein are incorporated by reference herein, including any priority documents and/or testing procedures to the extent they are not inconsistent with this text. As is apparent from the foregoing general description and the specific embodiments, while forms of this disclosure have been illustrated and described, various modifications can be made without departing from the spirit and scope of this disclosure. Accordingly, it is not intended that this disclosure be limited thereby. Likewise whenever a composition, an element or a group of elements is preceded with the transitional phrase “comprising,” it is understood that we also contemplate the same composition or group of elements with transitional phrases “consisting essentially of,” “consisting of,” “selected from the group of consisting of,” or “is” preceding the recitation of the composition, element, or elements and vice versa.

[0102] While the present disclosure has been described with respect to a number of embodiments and examples, those skilled in the art, having benefit of this disclosure, will appreciate that other embodiments can be devised which do not depart from the scope and spirit of the present disclosure.