WO2022060352A1 - Système et procédé d'alkylation d'isoparaffine - Google Patents

Système et procédé d'alkylation d'isoparaffine Download PDF

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WO2022060352A1
WO2022060352A1 PCT/US2020/050944 US2020050944W WO2022060352A1 WO 2022060352 A1 WO2022060352 A1 WO 2022060352A1 US 2020050944 W US2020050944 W US 2020050944W WO 2022060352 A1 WO2022060352 A1 WO 2022060352A1
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reactor
isoparaffin
alkylation
feed
stage
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PCT/US2020/050944
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English (en)
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Doron Levin
Ajit B. Dandekar
Christopher L. DEAN
Christopher L. GOEN
Rance N. FORD
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Exxonmobil Research And Engineering Company
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Priority to PCT/US2020/050944 priority Critical patent/WO2022060352A1/fr
Publication of WO2022060352A1 publication Critical patent/WO2022060352A1/fr

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    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G29/00—Refining of hydrocarbon oils, in the absence of hydrogen, with other chemicals
    • C10G29/20—Organic compounds not containing metal atoms
    • C10G29/205—Organic compounds not containing metal atoms by reaction with hydrocarbons added to the hydrocarbon oil

Definitions

  • This application relates to systems, methods, and apparatuses for alkylation of isoparaffins and, in particular, to systems, methods, and apparatuses for alkylation of isoparaffins with olefins to produce high octane alkylate, such as for use as a fuel additive.
  • alkylation of isoparaffins 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.
  • an isoparaffin can be reacted with an olefin to provide an isoparaffin of higher molecular weight.
  • alkylation often involves the reaction of C3-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, low sulfur, low olefin, low aromatic content, and low vapor pressure.
  • liquid acids such as hydrofluoric acid (HF) or sulfuric acid (H2SO4) as catalysts under relatively low temperature conditions.
  • HF hydrofluoric acid
  • H2SO4 sulfuric acid
  • An alternative to liquid acids are solid acids, such as zeolites.
  • zeolites such as zeolites.
  • the MWW framework solid acid catalysts may be used for catalytic alkylation of an olefin with an isoparaffin comprising contacting an olefin feed with an isoparaffincontaining feed under alkylation conversion conditions at a temperature at least equal to the critical temperature of the principal isoparaffin component of the feed.
  • Single-stage alkylation 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, which may, in turn, cause more rapid catalyst deactivation.
  • Multi-stage alkylation may be used with interphase olefin injection to enable increased local I/O ratios compared to single-stage alkylation.
  • interphase olefin injection to enable increased local I/O ratios compared to single-stage alkylation.
  • This application relates to systems, methods, and apparatuses for alkylation of isoparaffins and, in particular, to systems, methods, and apparatuses for alkylation of isoparaffins with olefins to produce high octane alkylate, such as for use as a fuel additive.
  • the present disclosure provides a system including a multistage reactor for alkylation of an isoparaffin using a solid acid catalyst and a liquid acid reactor for alkylation of an isoparaffin using a liquid acid catalyst.
  • the multistage reactor comprises a first inlet to receive an olefin feed, a second inlet to receive an isoparaffin feed, and a first outlet to output a first alkylation mixture comprising alkylate and isoparaffin.
  • the liquid acid reactor comprises a third inlet to receive an olefin feed, an isoparaffin feed, and at least a portion of the first alkylation mixture output from the multistage reactor, and a second outlet to output a second alkylation mixture comprising alkylate and isoparaffin.
  • the present disclosure provides a method including introducing, in a multistage reactor, a solid acid catalyst to an isoparaffin feed and an olefin feed to form a first alkylation mixture comprising alkylate and isoparaffin; and introducing, in a liquid acid reactor, a liquid acid catalyst to an isoparaffin feed, an olefin feed, and at least a portion of the first alkylation mixture to form a second alkylation mixture comprising alkylate and isoparaffin.
  • FIG. 1A is a depiction of a reactor with one stage configured to receive an olefin feed and an isoparaffin feed.
  • FIG. IB is a depiction of a reactor with two stages configured to receive an olefin feed and an isoparaffin feed, according to one or more aspects of the present disclosure.
  • FIG. 1C is a depiction of a reactor with four stages configured to receive an olefin feed and an isoparaffin feed, according to one or more aspects of the present disclosure.
  • FIG. 2 is a depiction of a generic flowsheet for a liquid acid alkylation system.
  • FIGS. 3 - 6 are depictions of flowsheets for acid alkylation systems, according to one or more aspects of the present disclosure.
  • This application relates to systems, methods, and apparatuses for alkylation of isoparaffins and, in particular, to systems, methods, and apparatuses for alkylation of isoparaffins with olefins to produce high octane alkylate, such as for use as a fuel additive.
  • the present disclosure leverages isobutane available from an existing alkylation process using liquid acid catalyst(s), such as HF and/or H2SO4, in combination with solid acid catalyst(s) to increase I/O ratio.
  • the solid acid catalyst(s) may be located in a multistage reactor (also referred to as a multi-bed reactor) upstream of the liquid acid reactor and the effluent of the solid acid reactor flowed directly to the liquid acid reactor, thereby allowing incremental production of alkylate without increasing the size of the liquid acid alkylation reactor or downstream separation systems.
  • Compositions of said solid acid catalyst(s) are also provided herein.
  • 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 has been discovered that performing alkylation using solid acid catalysts in multiple stages in combination with interstage olefin injection can be used to desirably increase I/O ratio.
  • the I/O ratio is a key variable impacting the activity and selectivity in an alkylation process. While multiple reactions are taking place in an alkylation reactor, there are two primary competing reactions for olefin: (1) alkylation with isoparaffin to form the desired paraffinic alkylate molecule (e.g. , 2-butene + isobutane -> C8 paraffin) and (2) oligomerization of olefin with other olefins to produce an undesired heavier olefin (e.g., 2-butene + 2-butene -> C8 olefin).
  • the undesired C8 olefin, once formed, may oligomerize further, leading to the formation of byproducts which diminish the activity and effectiveness of the acid catalyst.
  • byproducts which diminish the activity and effectiveness of the acid catalyst.
  • an acid soluble oil byproduct may be formed, and in solid acid alkylation heavier molecule byproducts that adsorb on the catalyst surface may be formed.
  • Selectivity of the desired alkylation pathway can be influenced by operating at high I/O ratio(s), thereby decreasing the probability of a feed olefin reacting with another olefin to form undesirable byproducts.
  • the I/O ratio can be increased by splitting the feed amongst multiple beds of a reactor.
  • the present disclosure provides use of a multistage alkylation reactor using solid acid catalyst in combination with a traditional single-stage alkylation reactor to achieve desirable I/O ratios and enable increased production capacity. That is, the present disclosure leverages existing liquid acid alkylation reactor flowsheets by combining a solid acid catalyst multistage reactor.
  • critical point refers to 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.
  • critical point for isobutane is about 134.6 °C and about 3650 kPa
  • critical point for isopentane is about 187.2 °C and about 3378 kPa.
  • isobutane is the principal component in a feedstock consisting of 2-butene and isobutane in an isobutane/2 -butene volume ratio of 50/1.
  • the term “heavy olefin,” and grammatical variants thereof, refers to a C8+ hydrocarbon containing at least one carbon-carbon double bond.
  • the term “light olefin,” and grammatical variants thereof, refers to a C2-C7 hydrocarbon containing at least one carbon-carbon double bond.
  • the term “inert gas,” and grammatical variants thereof, refers to a gas that does not undergo reaction in the presence of a catalyst, when there is no olefin present.
  • the term “MWW framework type,” and grammatical variants thereof, 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, Henk van Koningsveld (Elsevier, Amsterdam, 2007), incorporated herein by reference in its entirety.
  • molecular sieve As defined herein, the term “molecular sieve,” and grammatical variants thereof refers to 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.
  • crystalline microporous material of the MWW framework type refers to 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.
  • 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 or more aspects, one c-unit cell thickness;
  • 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
  • accessible volume refers to the unit cell volume remaining after the van der Waals atomic sphere volumes are subtracted.
  • the accessible volume is reported as a percentage of accessible volume out of the total volume.
  • the accessible volume is determined using the water absorption test of ASTM C830.
  • BET surface area refers to the Brunauer-Emmett-Teller method of measuring surface area of a solid via adsorption of gas molecules.
  • the BET surface area is calculated using the ISO 9277 standard.
  • deisoparaffinizer refers to equipment used to separate an isoparaffin of interest from a mixture of other hydrocarbons and/or other chemical compounds.
  • a deisobutanizer is a deisoparaffinizer that separates isobutane from a mixture of butane and heavier hydrocarbons.
  • Various aspects of the present disclosure can be conducted in any suitable single or 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.
  • the various aspects of the present disclosure conduct alkylation using a liquid acid catalyst in either a single-stage or a multistage reactor, and alkylation using a solid acid catalyst in a multistage reactor.
  • 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.
  • the reactor may include multiple catalyst beds located in the same or different housing.
  • a reactor or a stage within a multistage reactor may include a bed of catalyst particles where the particles have insignificant motion in relation to the bed (a fixed bed).
  • injection of the olefin feed can be effected at a single point in the reactor or at multiple points spaced along the reactor.
  • An isoparaffin feed is also introduced into the one or more stages of the reactor, and may be a separate feed from the olefin feed or, alternatively, the isoparaffin feed and the olefin feed may be premixed before entering the reactor or fed simultaneously.
  • 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.
  • a multistage reactor may include two (2), four (4) stages, ten (10) stages, twelve (12) stages, or any other plurality of stages, without departing from the scope of the present disclosure.
  • a reactor stage includes a catalyst bed.
  • Reactor stage(s) 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.
  • 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 acid catalyst therein.
  • the multistage reactor can provide heat transfer from reactor stages or catalyst beds in order to provide effective methods for controlling temperature.
  • the efficiency of a single or 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, but not limited to, the path length, the catalyst particle size, pore size, and the like, and any combination thereof.
  • a pressure drop that is too large may cause channeling through the catalyst bed, poor efficiency, and increased catalyst deactivation, which increases the frequency of catalyst rejuvenation.
  • the reactor has a cylindrical geometry with axial flows through the catalyst beds.
  • Various designs of the multistage reactor may accommodate control of specific process conditions, such as, but not limited to, pressure, temperature, liquid hourly space velocity (LHSV), and olefin liquid hourly space velocity (OLHSV), and any combination thereof.
  • LHSV liquid hourly space velocity
  • OLHSV olefin liquid hourly space velocity
  • Operating pressures may be controlled to reduce or eliminate oligomerization reactions and/or favor alkylation reactions, as described above. 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 psig to about 1500 psig (about 2068 to 10342 kPag), such as from about 400 psig to about 1200 psig (about 2758 to about 8274 kPag), from about 450 psig to about 1000 psig (about 3102 kPag to 6895 about kPag), from about 550 psig to about 950 psig (about 3792 kPag to about 6550 kPag), from about 650 psig to about 950 psig (about 4481 kPag to about 6550 kPag), from about 750 psig to about 950 psig (about 5171 kPag to about 6550 kPag), or from about 800 psig to about 950 psig (about 5516 to about 6550 kPag), encompassing any value and subset therebetween.
  • the operating temperature and pressure remain above the critical point for the isoparaffin feed during the
  • operating temperatures may be controlled to reduce or eliminate olefin oligomerization reactions and/or favor alkylation of isoparaffins, which may reduce byproduct formation and decrease the frequency of catalyst rejuvenations.
  • Operating temperature may be 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, encompassing any value and subset therebetween.
  • Operating temperatures may exceed the critical temperature of the isoparaffin feed, or the principal component in the isoparaffin feed.
  • the temperature of the multistage reactor or an individual stage within the reactor may affect byproduct 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 by reducing byproduct formation and decreasing the frequency of catalyst rejuvenations. 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.
  • 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.
  • the alkylation product mixture contains about 10 wt% or less, about 5 wt% or less, about 2 wt% or less, about 1 wt% or less, including about 10 wt% to about 0.1 wt%, or substantially free of products of olefin oligomerization, encompassing any value and subset therebetween.
  • Hydrocarbon flow through a reactor stage containing a catalyst is typically controlled to provide an OLHSV sufficient to convert about 99 wt% or more, by weight of fresh olefin to alkylation product.
  • OLHSV values are from about 0.01 hr 1 to about 10 hr 1 , such as about 0.02 hr 1 to about 1 hr 1 , or 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.
  • FIG. 1A depicts an alkylation reactor 100 A with a single reactor stage 101.
  • Reactor stage(s) may individually or collectively be termed an alkylation zone and include catalyst, such as a solid acid catalyst including a 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.
  • the I/O ratio is controlled solely by the composition of the olefin feed and the isoparaffin feed entering reactor bed 101. For purposes of explanation, it will be assumed that the alkylation reactor 100A is operating at an I/O ratio of 4/1.
  • FIG. IB depicts a multistage alkylation reactor 100B with two reactor stages: first stage 101 A and second stage 101B.
  • the olefin feed is introduced to the reactor beds via lines 103 A and 103B.
  • 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 101A 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 the first stage 101 A and the second stage 10 IB to allow for introduction of additional olefin feed through line 103B.
  • Reactor stages 101A and 101B may individually include an upstream zone and a downstream zone.
  • a downstream zone is proximate to interstage space 109 and may include hydrogenation co- catalyst.
  • 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 an upstream zone and a downstream zone, for example, in at least one aspect, where hydrogenation co-catalyst is supported or bound with catalyst.
  • 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 compared to FIG. 1A because the olefin feed is divided into two lines 103A and 103B and the olefin introduced via line 103A to first stage 101A can be converted. For example, greater than about 90 wt% or more, including up to 100 wt% of olefin may be converted within the first stage 101 A, based on the total weight of olefin in the olefin feed line 103 A.
  • the amount of isoparaffin introduced to interstage 109 and, thus, to second stage 101B is similar to that introduced to first stage 101A. Accordingly, the I/O ratio of the second stage 101B is similar but somewhat reduced compared to the I/O ratio of the first stage 101 A, if no additional isoparaffin is added.
  • Additional isoparaffin may be added at interstage space 109, for example, so as to maintain a consistent I/O ratio throughout the multistage reactor 100B.
  • the olefin introduced to interstage space 109 (either via line 103B or from the effluent of first stage 101A) and, therefore, introduced to second stage 101B is similar or identical in quantity to that introduced to first stage 101A.
  • the alkylation product mixture exits the reactor 100B through line 107. Therefore, the I/O ratio of 4/1 using the single-stage reactor of FIG. 1 A can be increased to 8/1 at bed inlet based on the multistage reactor configuration shown in FIG. IB.
  • FIG. 1C depicted is a multistage alkylation reactor 100C with four reactor stages: first stage 101A, second stage 101B, third stage 101C, and fourth stage 101D.
  • the olefin feed is introduced to the reactor beds via lines 103A, 103B, 103C, and 103D.
  • 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 101A, second stage 101B, third stage 101C, and fourth stage 101D.
  • the isoparaffin feed is introduced to alkylation reactor 100C through line 105.
  • alkylation reactor 100C has multiple interstage spaces: first interstage space 109 A, 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 101A, 101B, 101C, and 101D may individually include an upstream zone and a downstream zone.
  • a downstream zone is proximate to interstage space 109 A and may include hydrogenation co-catalyst.
  • 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 an upstream zone and a downstream zone, for example, in at least one aspect, where hydrogenation co-catalyst is supported or bound with catalyst.
  • lines 103 and 105 have the same composition as in FIG. 1A, then the local I/O ratio is 4 times in the configuration of FIG. 1C compared to FIG. 1A because the olefin feed is divided into four lines 103A, 103B, 103C, and 103D.
  • the I/O ratio in each of the single-stages 101A, 101B, 101C, and 101D is only slightly affected by any prior stage(s) because the olefin introduced to any single-stage can be largely converted therein.
  • greater than about 90 wt% or more, including up to 100 wt% of olefin may be converted within each stage 101A, 101B, 101C, and 101D, based on the total weight of olefin in the olefin feed line 103A, 103B, 103C, and 103D, respectively.
  • Only a small portion of the isoparaffin feed 105 is converted by the reaction in each of stages 101A, 101B, 101C, and 101D, such as less than about 10 wt%, or in the range of about 10 wt% to about 0.1 wt%, of the isoparaffin feed is converted based on the total weight of isoparaffin.
  • the amount of isoparaffin introduced to interstages 109A, 109B, and 109C and, thus, to second stage 101B, third stage 101C, and fourth stage 101D, respectively, is similar to that introduced to first stage 101A. Accordingly, the I/O ratio of each of the four stages 101A, 101B, 101C, and 101D is similar, although each of stages 101B, 101C, and 101D will have increasingly reduced I/O ratios compared to first stage 101 A, if no additional isoparaffin is added. Additional isoparaffin may be added at any of the interstage spaces, for example, so as to maintain a consistent I/O ratio throughout the multistage reactor 100C.
  • the olefin introduced to interstage spaces 109A, 109B, and 109C (either via line 103B, 103C, or 103D or from the effluent of a prior stage) and, therefore, introduced to downstream stages is similar or identical in quantity to that introduced to all of the stages.
  • the alkylation product mixture exits the reactor 100C through line 107. Therefore, the I/O ratio of 4/1 using the single-stage reactor of FIG. 1A can be increased to 16/1 at bed inlet based on the multistage reactor configuration shown in FIG. 1C.
  • the particular location of olefin feed and the isoparaffin feed inlets shown in FIGS. 1A-1C are non-limiting and may be introduced at any location provided that the feeds are supplied to the one or more reactor stages.
  • the specific configuration, shape, interspace volume, stage volume, outlet location, and the like for the multistage reactors described in the present disclosure is non-limiting and may depend on a number of factors including, but not limited to, refining facility design, desired amount or volume of alkylation product mixture, and the like.
  • the amount of stages within a multistage reactor may also be three (3) or greater than four (4), without departing from the scope of the present disclosure depending on the desired I/O ratio.
  • the I/O ratio at bed inlet could be increased to 40/1 by increasing the number of stages to ten (10).
  • Feedstocks useful in the present alkylation process include at least one isoparaffin feed and at least one olefin feed.
  • the isoparaffin feed used in alkylation processes of the present disclosure may have from about 4 to about 7 carbon atoms (about C4 to about C7), encompassing any value and subset therebetween.
  • Representative examples of such isoparaffins include, but are not limited to, isobutane, isopentane, 3 -methylhexane, 2-methyIhexane, 2,3-dimethyIbutane, and mixture(s) thereof, typically isobutane.
  • the olefin component of the feedstock may include at least one olefin having from about 2 to about 12 carbon atoms (about C2 to about C12), encompassing any value and subset therebetween.
  • Representative examples of such olefins include, but are not limited to, 2-butene, isobutylene, 1 -butene, propylene, ethylene, pentene, hexene, octene, heptene, or mixture(s) thereof.
  • the olefin component of the feedstock is selected from the group consisting of propylene, butene, pentene, and any combination or mixture(s) thereof.
  • 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 150:1, such as from about 0.1:1 to about 1:1, encompassing any value and subset therebetween.
  • 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 150:1, such as from about 0.1: 1 to about 1:1, encompassing any value and subset therebetween.
  • the concentration of olefin feed can be adjusted by, such as, for example, by staged additions thereof. Using staged additions, isoparaffin/olefin feed concentrations (and therefore the I/O ratio) can be maintained at levels to improve conversion and reduce catalyst deactivation.
  • the ratio of isoparaffin to olefin ratio by volume 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, or about 120/1 to about 500/1, or about 160/1 to about 480/1, or about 200/1 to about 450/1, or about 220/1 to about 450/1, or about 240/1 to about 420/1, or about 240/1 to about 400/1, encompassing any value and subset therebetween.
  • the production of olefin oligomers increases with lower I/O ratios.
  • an I/O ratio of about 100/1 or greater may be used.
  • the efficiency of the alkylation process can be reduced at higher I/O ratios, due to a 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 isobutane and lighter alkanes from the nC4 and C5+ hydrocarbons produced in the reactor (e.g., a deisobutanizer).
  • a higher I/O ratio can provide greater quantities of isobutane and lighter alkanes separated from the alkylation product mixture that can be recycled to the reactor.
  • a hydrogen feed may be fed to the reactor including hydrogen and, in some instances, inert gases to decrease hydrogen concentration within the feed.
  • concentration of hydrogen in the multistage reactor can be adjusted by, e.g., staged additions thereof.
  • hydrogen/olefin feed concentrations can be maintained at levels sufficient to reduce or eliminate olefin oligomers formed in a stage of the multistage reactor.
  • the molar ratio of hydrogen to olefin is from about 1:1000 to about 1:1, about 1:500 to about 1:2, or about 1:100 to about 1:5, encompassing any value and subset therebetween.
  • the isoparaffin feed, the olefin feed, and/or the hydrogen feed may be treated to remove catalyst poisons.
  • 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.
  • 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. 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.
  • 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.
  • the crystalline microporous material is of the MWW framework type, such as a zeolite.
  • crystalline microporous materials of the MWW framework type include, but are not limited to, MCM-22 (U.S. Patent No. 4,954,325, incorporated herein by reference in its entirety), PSH-3 (U.S. Patent No. 4,439,409, incorporated herein by reference in its entirety), SSZ-25 (U.S. Patent No.
  • ERB-1 European Patent No. 0293032
  • ITQ-1 U.S. Patent No. 6,077,498, incorporated herein by reference in its entirety
  • ITQ-2 International Publication No. WO97/17290, incorporated herein by reference in its entirety
  • MCM-36 U.S. Patent No. 5,250,277, incorporated herein by reference in its entirety
  • MCM-49 U.S. Patent No. 5,236,575, incorporated herein by reference in its entirety
  • MCM-56 U.S. Patent No. 5,362,697, incorporated herein by reference in its entirety
  • UZM-8 U.S. Patent No.
  • UZM-8HS U.S. Patent No. 7,713,513, incorporated herein by reference in its entirety
  • UZM-37 U.S. Patent No. 7,982,084, incorporated herein by reference in its entirety
  • EMM-10 U.S. Patent No. 7,842,277, incorporated herein by reference in its entirety
  • EMM-12 U.S. Patent No. 8,704,025, incorporated herein by reference in its entirety
  • EMM-13 U.S. Patent No. 8,704,023, incorporated herein by reference in its entirety
  • UCB-3 U.S. Patent No.
  • the crystalline microporous material of the MWW framework type may be contaminated with other crystalline materials, such as mordenite, ferrierite or quartz. These contaminants may be present in quantities of about 10 wt% or less, such as about 5 wt% or less, including 0% (no detectable impurities).
  • 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.
  • 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; and any combination thereof, such as the combination of two noble metals, two non-noble metals, or a combination of noble and non-noble metals.
  • noble metals such as Pd, Pt, Rh, Ru, Ir, Os, Ag, Au
  • non-noble metals such as Mo, Co, Ni, Fe
  • any combination thereof such as the combination of two noble metals, two non-noble metals, or a combination of noble and non-noble metals.
  • the co-catalyst may be supported, suitable support materials may include clay, alumina, silica, titania, zirconia, aluminosilicates, zeolites, carbon, and combination thereof.
  • the silica support can be an amorphous silica support.
  • 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, but are not limited to, zeolites, such as MCM-41, other M41S structures, SBA-15, and the like, and any combination thereof.
  • a silica support may be modified with alumina and the amount of alumina added to modify a silica support may vary.
  • 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%, encompassing any value and subset therebetween.
  • 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%, encompassing any value and subset therebetween.
  • 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, encompassing any value and subset therebetween.
  • the amount of 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.
  • 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), encompassing any value and subset therebetween.
  • 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), encompassing any value and subset therebetween.
  • 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), encompassing any value and subset therebetween.
  • 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.
  • the combined amount of metal(s) can be about 0.05 wt% to about 5.0 wt%, or about 0.05 wt% to about
  • 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) , encompassing any value and subset therebetween.
  • 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, encompassing any value and subset therebetween.
  • 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).
  • 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, encompassing any value and subset therebetween.
  • Catalysts and co-catalysts suitable for use in the systems and processes described include an optional binder.
  • 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 60 wt% or less, for example about 50 wt% or less, such as about 40 wt% or less, for example about 30 wt% or less, such as about 20 wt% or less, including 0%. Where a non-alumina binder is present, the amount employed may be as little as about 1 wt%, or about 5 wt% or more, for example about 10 wt% or more, or in the range of about 1% to about 20%, encompassing any value and subset therebetween.
  • a silica binder is employed such as disclosed in U.S. Pat. No. 5,053,374, incorporated herein by reference in its entirety.
  • a zirconia or titania binder is used.
  • 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 herein by reference in its entirety.
  • the binder material may contain alumina, including amorphous alumina.
  • a catalyst composition may be prepared according to various aspects of the present disclosure by introducing a zeolite to an optional binder and water to form a first mixture which is extruded to form an extrudate.
  • the extrudate is dried before addition of an exchange fluid.
  • the mixture after having been introduced to an exchange fluid, may be calcined.
  • a catalyst composition may include zeolite and optional binder.
  • a catalyst composition may include zeolite in about 30 wt% or greater, such as about 40 wt% or greater, about 50 wt% or greater, about 60 wt% or greater, about 70 wt% or greater, about 80 wt% or greater, about 82 wt% or greater, about 84 wt% or greater, about 86 wt% or greater, about 88 wt% or greater, about 90 wt% or greater, about 92 wt% or greater, about 94 wt% or greater, about 95 wt% or greater, about 96 wt% or greater, about 97 wt% or greater, about 98 wt% or greater, about 99 wt% or greater, or about 99.5 wt% or greater, such as from about 40 wt% to about 99.99 wt%, from about 50 wt% to about 99.95 wt%, from about 60wt% to about 99.9 wt%, from about 70 wt% to about 99
  • a catalyst composition of the present disclosure may be prepared by adding an exchange fluid to a mixture of the zeolite and optional binder.
  • Such exchange fluids contain cations that exchange with cations associated with the zeolite framework, such as sodium cations.
  • the catalyst composition of the present disclosure may be prepared by treating the mixture of zeolite and optional binder with a swelling agent which may cause the zeolite layers to swell or separate and are removable by calcination.
  • Suitable exchange fluids include sources of cations, such as quaternary ammonium cations, such as organoammonium cations, or inorganic ammonium cations, and any combination thereof.
  • Suitable swelling agents may include, but are not limited to, a source of organic cations such as quaternary organoammonium cations or organophosphonium cations, in order to affect an exchange of interspathic cations.
  • Suitable exchange fluids may include, but are not limited to, aqueous or non-aqueous solutions. Additionally, exchange fluids may include ammonium cations. Also, suitable exchange fluids may have a normality of from about 0.1 N to about 5 N, such as from about 0.2 N to about 4 N, from about 0.4 N to about 3 N, or from about 0.5 N to about 2 N, encompassing any value and subset therebetween.
  • Suitable sources of ammonium cations may include, but are not limited to, ammonium nitrate, ammonium hydroxide, ammonium acetate, ammonium chloride, ammonium carbonate, tetramethylammonium nitrate, tetramethylammonium hydroxide, n-octylammonium nitrate, n- octylammonium hydroxide, cetyltrimethylammonium nitrate, cetyltrimethylammonium hydroxide, and any combination thereof.
  • a pH range of about 4 to about 14, such as about 4.5 to about 13.5, encompassing any value and subset therebetween, is typically employed during treatment with the exchange fluid.
  • the catalyst composition is dried prior to the addition of an exchange fluid. Drying the catalyst composition may include thermal treatment at about 300 °C or less (below calcination temperatures). Drying may take place at a temperature of from about 100 °C to about 300 °C, such as from about 105 °C to about 250 °C, from about 110 °C to about 220 °C, from about 115 °C to about 200 °C, or from about 120 °C to about 180 °C, encompassing any value and subset therebetween.
  • the catalyst composition is not dried prior to the addition of an exchange fluid and is dried and calcined thereafter.
  • the catalyst composition before drying or calcining may have a solids content of about 60 wt% or less, such as about 50 wt% or less, about 45 wt% or less, about 40 wt% or less, such as in the range of about 60 wt% to about 35 wt%, encompassing any value and subset therebetween.
  • Calcining can be performed by heating the catalyst composition at temperature of about 350 °C or greater, about 375 °C or greater, about 400 °C or greater, about 425 °C or greater, about 450 °C or greater, about 475 °C or greater, about 500 °C or greater, about 525 °C or greater, or about 550 °C or greater, such as from about 250 °C to about 1000 °C, from about 300 °C to about 900 °C, from about 350 °C to about 800 °C, from about 400 °C to about 700 °C, or from about 450 °C to about 600 °C, encompassing any value and subset therebetween.
  • Calcination may occur in a time frame of from about 1 minute to about 72 hours, such as from about 5 minutes to about 48 hours, from about 10 minutes to about 36 hours, from about 15 minutes to about 24 hours, from about 20 minutes to about 20 hours, from about 25 minutes to about 18 hours, or from about 30 minutes to about 16 hours, encompassing any value and subset therebetween.
  • Calcination may be performed in the presence of inert gas such as nitrogen or argon, or in the presence of non-inert gases such as oxygen or air, or in mixtures thereof, for example mixtures of air and nitrogen.
  • inert gas such as nitrogen or argon
  • non-inert gases such as oxygen or air
  • the exchange fluid is decomposed or oxidized by the presence of oxygen or air during calcination. While subatmospheric pressure can be employed for the calcination, atmospheric pressure is typical used simply for reasons of convenience.
  • the addition of exchange fluid may result in the formation of a layered oxide of enhanced interlayer separation, compared to the layered oxide before introduction to the exchange fluid.
  • the interlayer separation may be dependent upon the steric volume of the cation introduced.
  • a series of cation exchanges can be carried out. For example, a cation may be exchanged with a cation of greater size, thus increasing the interlayer separation in a step-wise fashion, as compared to cation exchange performed using a cation of smaller size.
  • water may be trapped between the layers of the zeolite.
  • the calcined catalyst composition may include layers, which can exhibit high BET surface area (e.g., greater than 400 m 2 /g), making them highly useful as catalysts or catalytic supports, for hydrocarbon conversion processes, such as alkylation.
  • a calcined catalyst composition of the present disclosure may exhibit a BET surface area of about 400 m 2 /g or greater, such as about 450 m 2 /g or greater, about 500 m 2 /g or greater, about 550 m 2 /g or greater, or about 600 m 2 /g or greater, such as from about 400 m 2 /g to about 2000 m 2 /g, from about 450 m 2 /g to about 1500 m 2 /g, from about 500 m 2 /g to about 1000 m 2 /g, from about 550 m 2 /g to about 900 m 2 /g, or from about 600 m 2 /g to about 800 m 2 /g, encompassing any value and subset therebetween.
  • the calcined catalyst composition may further exhibit an accessible volume of about 10% or greater, such as 12% or greater, 15% or greater, or 17% or greater, such as from about 10% to about 40%, from about 12% to about 35%, from about 15% to about 30%, or from about 17% to about 25%, encompassing any value and subset therebetween.
  • 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, and 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.
  • the product may include about 15-35 wt% of C5-C7 hydrocarbons, 50-85 wt% of C8 hydrocarbons and 1-10 wt% of C9+ hydrocarbons.
  • a process can be selective to desirable high octane components so that, in the case of alkylation of isobutane with C3-C4 olefins, the C6 fraction typically includes at least about 40 wt%, such as at least about 70 wt%, of 2,3-dimethylbutane, the C7 fraction typically includes at least about 40 wt%, such as at least about 80 wt%, of 2,3-dimethylpentane and the C8 fraction typically includes at least about 50 wt%, such as at least about 70 wt%, of 2,3,4-trimethylpentane, 2,3,3-trimethylpentane, and 2,2,4- trimethylpentane.
  • the product may include about 20-40 wt% of C5 hydrocarbons, about 15- 35 wt% of C9 hydrocarbons, about 20-35 wt% of C8 hydrocarbons, and about 2-10 wt% of C10+ hydrocarbons.
  • 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.
  • the molar ratio of trimethylpentane to dimethylhexane can be about 3 or more, such as about 4 to about 5, or about 3 to about 6.
  • the molar ratio of trimethylhexane to dimethylheptane can be about 1 or more, such as about 1.5 or more, or from about 1 to about 3, encompassing any value and subset therebetween.
  • 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 isoparaffin to be recycled as part or all of the isoparaffin feed (e.g., a deisobutanizer). Furthermore, depending on alkylate demand, part or all of a C9+ fraction can be recovered for use as a distillate blending stock.
  • a separation system such as a distillation train
  • Various portions of the multistage reactors, single-stage reactors, and/or separation systems described herein may be in fluid communication, such as by feed lines, hoses, pipes, troughs, or other conduits, and may be equipped with suitable pumps and valving to facilitate and/or control flow therebetween.
  • a multistage reactor may be in fluid communication with one or both of a single-stage reactor and a deisoparaffinizer, typically in single directional flow.
  • other conduits may be in fluid communication with the multistage reactors, single-stage reactors, and/or separation systems directly or by way of other fluid communication conduits, without departing from the scope of the present disclosure.
  • Clause 1 A system comprising: a multistage reactor for alkylation of an isoparaffin using a solid acid catalyst, the multistage reactor comprising: a first inlet to receive an olefin feed; a second inlet to receive an isoparaffin feed; and a first outlet to output a first alkylation mixture comprising alkylate and isoparaffin; a liquid acid reactor for alkylation of an isoparaffin using a liquid acid catalyst, the liquid acid reactor comprising: a third inlet to receive an olefin feed, an isoparaffin feed, and at least a portion of the first alkylation mixture output from the multistage reactor; and a second outlet to output a second alkylation mixture comprising alkylate and isoparaffin.
  • Clause 2 The system of Clause 1, wherein the first outlet is in fluid communication with the third inlet to receive the at least a portion of the first alkylation mixture output from the multistage reactor.
  • Clause 3 The system of Clause 1 or Clause 2, further comprising a deisoparaffinizer having: a fourth inlet to receive the second alkylation mixture from the liquid acid reactor, the deisoparaffinizer to separate at least a portion of the isoparaffin from the second alkylation mixture, thereby resulting in separated isoparaffin; a third outlet to output the separated isoparaffin; and a fourth outlet to output the second alkylation mixture having the separated isoparaffin separated therefrom.
  • a deisoparaffinizer having: a fourth inlet to receive the second alkylation mixture from the liquid acid reactor, the deisoparaffinizer to separate at least a portion of the isoparaffin from the second alkylation mixture, thereby resulting in separated isoparaffin; a third outlet to output the separated isoparaffin; and a fourth outlet to output the second alkylation mixture having the separated isoparaffin separated therefrom.
  • Clause 4 The system of Clause 3, wherein the third outlet is in fluid communication with the third inlet to receive the at least a portion of the separated isoparaffin output from the isoparaffin.
  • Clause 5 The system of Clause 3 or Clause 4, wherein at least a portion of the isoparaffin feed received by the second inlet is formed of the separated isoparaffin output from the isoparaffin.
  • Clause 6 The system of any of the preceding Clauses, wherein the third outlet is in fluid communication with the second inlet to receive the at least a portion of the separated isoparaffin output from the isoparaffin.
  • Clause 7 The system of any of the preceding Clauses, wherein the multistage reactor is a fixed bed multistage reactor.
  • Clause 8 The system of any of the preceding Clauses, wherein the multistage reactor comprises at least a first stage and a second stage, and a first interstage space connecting the first stage and the second stage.
  • a method comprising: introducing, in a multistage reactor, a solid acid catalyst to an isoparaffin feed and an olefin feed to form a first alkylation mixture comprising alkylate and isoparaffin; and introducing, in a liquid acid reactor, a liquid acid catalyst to an isoparaffin feed, an olefin feed, and at least a portion of the first alkylation mixture to form a second alkylation mixture comprising alkylate and isoparaffin.
  • Clause 10 The method of Clause 9, wherein the multistage reactor is in fluid communication with the liquid acid reactor to introduce the at least a portion of the first alkylation mixture from the multistage reactor and to the liquid acid reactor.
  • Clause 11 The method of Clause 9 or Clause 10, further comprising separating, in a deisoparaffinizer, at least a portion of the isoparaffin from the second alkylation mixture, thereby resulting in separated isoparaffin.
  • Clause 12 The method of Clause 11, wherein at least a portion of the isoparaffin feed introduced to the multistage reactor is formed of the separated isoparaffin.
  • Clause 13 The method of Clause 12, wherein the deisoparaffinizer is in fluid communication with the multistage reactor to introduce the separated isoparaffin from the deisoparaffinizer to the multistage reactor.
  • Clause 14 The method of Clause 11 to Clause 13, wherein the isoparaffin feed introduced to the multistage reactor and the isoparaffin feed introduced to the single-stage reactor each comprise isobutane, and the deisoparaffinizer is a deisobutanizer.
  • Clause 15 The method of Clause 9 to Clause 14, wherein the isoparaffin feed introduced to the multistage reactor and the isoparaffin feed introduced to the liquid acid reactor each comprise isobutane.
  • Clause 16 The method of Clause 9 to Clause 15, wherein the solid acid catalyst comprises a zeolite.
  • Clause 17 The method of Clause 16, wherein the zeolite is a crystalline microporous material of MWW framework type.
  • Clause 18 The method of Clause 9 to Clause 17, wherein the second alkylation mixture comprises a greater concentration of alkylate per unit compared to the first alkylation mixture.
  • Clause 19 The method of Clause 9 to Clause 18, wherein the multistage reactor is a fixed bed multistage reactor.
  • Clause 20 The method of Clause 9 to Clause 19, wherein the multistage reactor comprises at least a first stage and a second stage, and a first interstage space connecting the first stage and the second stage.
  • any I/O ratio may be achieved with any increase in production capacity up to the limit of the amount of isobutane available, either in the recycle loop or in the makeup stream, as illustrated below.
  • high I/O ratio is preferable, and the I/O ratio using of solid acid catalysts can be maximized by increasing the number of beds in a reactor and/or increasing the amount of the isoparaffin (e.g. , isobutane) fed to the fixed bed reactor.
  • Examples 1 through 4 below describe synthesis of catalysts used in accordance with one or more aspects of the present disclosure.
  • Examples 5 and 6 below describe use of the synthesized catalysts for alkylation of isobutane with 2-butene to demonstrate the impact of I/O ratio on catalyst performance.
  • Examples 7 through 10 describe leveraging existing liquid acid alkylation reactor flowsheets in combination with a solid catalyst multistage reactor to achieve desirable high I/O ratios and increased production, in accordance with various aspects of the present disclosure.
  • Isobutane was obtained from a commercial source and used as received.
  • the isobutane purity was 99.6 % with the balance n-butane.
  • Catalysts used for isobutane alkylation with light olefins are dried in the reactor under nitrogen flow at 250°C for at least 4 hours prior to use.
  • Catalyst 1 (“Cl”) 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 about 10 to 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).
  • the dried extrudate was heated to 1,000 °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).
  • Catalyst 2 (“C2”) was prepared by combining 95 parts MCM-49 zeolite crystals 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 about 10 to 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).
  • 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).
  • Catalyst 3 (“C3”) was prepared by combining 95 parts MCM-49 zeolite crystals 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 about 10 to 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).
  • the dried extrudate was ion exchanged with 0.75 N ammonium nitrate solution followed by washing with deionized water and drying.
  • the dried extrudate was then was heated to 1000 °F (538 °C) under flowing nitrogen and finally calcined in a nitrogen/air mixture to a temperature of 1000 °F (538 °C).
  • Catalyst 4 (“C4”) was prepared by combining 95 parts MCM-49 zeolite crystals with 2.5 parts precipitated silica and 2.5 parts colloidal silica, on a calcined dry weight basis.
  • the MCM- 49 and precipitated silica dry powders were placed in a muller or a mixer and mixed for about 5 to 20 minutes.
  • Colloidal silica available as LUDOX® HS-40 from W.R. Grace (Columbia, MD), was then added and mixed for about 5 to 10 minutes. Sufficient water and a 5% NaOH solution (2.5% NaOH by weight) were then added during the mixing process to produce an extrudable paste.
  • the extrudable paste was formed into a 1/20 inch cylindrical 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 ion exchanged with 0.75 N ammonium nitrate solution followed by washing with deionized water and drying. The dried extrudate was then is heated to 1000 °F (538 °C) under flowing nitrogen and finally calcined in a nitrogen/air mixture to a temperature of 1000 °F (538 °C).
  • Catalyst C2 (of Example 2) was loaded into a pilot plant and operated as a single bed, as shown in FIG. 1A.
  • the reactor was 14 inches long and made from 3/8 inch outside diameter (O.D.) stainless steel tubing.
  • the reactor was loaded with 4 grams (g) of catalyst.
  • the reactor was located in an electrically heated furnace and maintained at 302 °F (150 °C).
  • Reactor pressure was 750 pounds per square inch gauge (psig).
  • the reactor effluent was measured using gas chromatography with flameionization detection (FID GC) equipped with a 150 m PETROCOL® column.
  • FID GC flameionization detection
  • the flow to the reactor was set to achieve an Olefin Liquid Hourly Space Velocity (OLHSV) of 0.03 h 1 .
  • Performance data as a function of I/O ratio is shown in Table 1; selectivity data is in grams per gram of C5+.
  • Catalyst C4 (of Example 4) was loaded into a pilot plant and operated as a four (4) reactor bed system, as shown in FIG. 1C. Each reactor was 60 inches long and made from 3/4 inch O.D. schedule 40 pipe. Each reactor was loaded with -148 g of catalyst. The reactors were located in an isothermal sandbath maintained at 302 °F (150 °C). Reactor pressure was 850 psig. Isobutane (99.6% purity) was fed to the first reactor bed and 2-butene flow was split evenly into four (4) parts using Coriolis meters and independently fed to each of the four (4) reactor beds.
  • the relative rates of isobutane and 2-butene were set such that the isobutane to 2-butene ratio at the top of the first reactor bed was -110/1, -150/1, or -190/1, as provided in Table 2 below.
  • the reactor effluent exiting the last bed 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 OLHSV of 0.032 h 1 .
  • Performance data as a function of I/O ratio is shown in Table 2; selectivity data is in grams per gram of C5+.
  • Example 7 A generic flowsheet for a liquid acid alkylation system 200 is shown in FIG 2.
  • the liquid acid reactor system 200 operates at an I/O ratio of 4/1.
  • Flows have been normalized such that the pure volumetric olefin feed rate is set at 1 thousand barrels per day (“kbd”).
  • kbd the pure volumetric olefin feed rate
  • all non-olefin components of the olefin feed, such as n-butane have been omitted from FIG. 2.
  • the n-butane generally present in the make-up isobutane (also referenced herein as “iC4”) and recycled isobutane streams have been omitted.
  • a 1 kbd olefin feed 202 will react with -1.17 kbd of isobutane feed 204 in liquid acid reactor 206 to produce -1.78 kbd of alkylate product 208 recovered from the deisobutanizer tower (“DiB”) 212.
  • the 2.83 kbd of unreacted isobutane feed 210 recovered overhead from the DiB tower 212 is recycled to the front of the reactor where it mixes with the 1.17 kbd of isobutane makeup 204 to give the 4 kbd of isobutane (labeled “A”) that feeds the reactor 206 along with the 1 kbd of olefin feed 202.
  • FIG. 3 illustrates a combined liquid and solid alkylation system 300, according to one or more aspects of the present disclosure, combining an existing alkylation process using liquid acid catalyst(s) (e.g., as shown in FIG. 2) with a multistage solid acid catalyst alkylation process.
  • liquid acid catalyst(s) e.g., as shown in FIG. 2
  • FIG. 3 illustrates a combined liquid and solid alkylation system 300, according to one or more aspects of the present disclosure, combining an existing alkylation process using liquid acid catalyst(s) (e.g., as shown in FIG. 2) with a multistage solid acid catalyst alkylation process.
  • This example demonstrates that a refiner operating the liquid acid reactor of FIG. 2 (Example 7) can increase the production of alkylate by 10%, even though unable to feed additional olefin to the liquid acid reactor due to a variety of constraints (e.g., refrigeration limit, acid consumption/hydraulic limit, feed hydraulic limit, and the like).
  • An additional 0.117 kbd of isobutane is required (compared to Example 7) to react the additional 0.1 kbd of olefin 303, increasing the isobutane makeup feed 304 rate to 1.287 kbd.
  • the multistage reactor 301 comprises ten (10) separate beds containing solid acid catalyst, making the olefin flow to each bed 0.01 kbd, assuming equal flow to each bed.
  • sufficient isobutane is drawn from the isobutane recycle loop 310 (2.83 kbd).
  • a 2.117 kbd isobutane feed 305 from the recycle loop 310 is fed to the multistage reactor 301, making the local I/O ratio at the bed inlet equal to 212.
  • 0.178 kbd of alkylate is produced while consuming 0.117 kbd of isobutane.
  • the multistage reactor 301 effluent 307 (2 kbd of isobutane + 0.178 kbd of alkylate) is fed to the front end of the liquid acid reactor 306 together with the remaining 2 kbd of isobutane in the recycle loop 309 and the remaining 1 kbd of olefin feed 302b such that the I/O ratio going into the liquid acid reactor 306 remains at 4/1.
  • a total of 1.958 kbd of alkylate 308 is produced as a bottoms product in the DiB tower 312 at constant DiB overhead recycle rate when compared against Example 7.
  • FIG. 4 illustrates a combined liquid and solid alkylation system 400, according to one or more aspects of the present disclosure, combining an existing alkylation process using liquid acid catalyst(s) (e.g., as shown in FIG. 2) with a multistage solid acid catalyst alkylation process.
  • liquid acid catalyst(s) e.g., as shown in FIG. 2
  • FIG. 4 illustrates a refiner operating the liquid acid reactor of FIG. 2 (Example 7) can increase the production of alkylate by 10%, even though unable to feed additional olefin to the liquid acid reactor due to a variety of constraints (e.g., refrigeration limit, acid consumption/hydraulic limit, feed hydraulic limit, and the like).
  • An additional 10% (olefin feed 403 having 0.1 kbd) of olefin feed 402a (having a total of 1.1 kbd) is directed to the multistage reactor 401.
  • An additional 0.117 kbd of isobutane is required (compared to Example 7) to react the additional 0.1 kbd of olefin 403, increasing the isobutane makeup feed 404 rate to 1.287 kbd.
  • the multistage reactor 401 comprises ten (10) separate beds containing solid acid catalyst, making the olefin flow to each bed 0.01 kbd, assuming equal flow to each bed.
  • sufficient isobutane is drawn from the isobutane recycle loop 410 (2.83 kbd).
  • a 3.117 kbd isobutane feed 405 from the recycle loop 410 is fed to the multistage reactor 401, making the local I/O ratio at the bed inlet equal to 312.
  • 0.178 kbd of alkylate is produced while consuming 0.117 kbd of isobutane.
  • the multistage reactor 401 effluent 407 (3 kbd of isobutane + 0.178 kbd of alkylate) is fed to the front end of the liquid acid reactor 406 together with the remaining 1 kbd of isobutane in the recycle loop 409 and the remaining 1 kbd of olefin feed 402b such that the I/O ratio going into the liquid acid reactor 406 remains at 4/1.
  • a total of 1.958 kbd of alkylate 408 is produced as a bottoms product in the DiB tower 412 at constant DiB overhead recycle rate when compared against Example 7.
  • FIG. 5 illustrates a combined liquid and solid alkylation system 500, according to one or more aspects of the present disclosure, combining an existing alkylation process using liquid acid catalyst(s) (e.g., as shown in FIG. 2) with a multistage solid acid catalyst alkylation process.
  • liquid acid catalyst(s) e.g., as shown in FIG. 2
  • FIG. 5 illustrates a combined liquid and solid alkylation system 500, according to one or more aspects of the present disclosure, combining an existing alkylation process using liquid acid catalyst(s) (e.g., as shown in FIG. 2) with a multistage solid acid catalyst alkylation process.
  • This example demonstrates that a refiner operating the liquid acid reactor of FIG. 2 (Example 7) can increase the production of alkylate by 15%, even though unable to feed additional olefin to the liquid acid reactor due to a variety of constraints (e.g., refrigeration limit, acid consumption/hydraulic limit, feed hydraulic limit, and the like).
  • the multistage reactor 501 comprises ten (10) separate beds containing solid acid catalyst, making the olefin flow to each bed 0.015 kbd, assuming equal flow to each bed.
  • sufficient isobutane is drawn from the isobutane recycle loop 510 (2.83 kbd).
  • a 3.176 kbd isobutane feed 505 from the recycle loop 510 is fed to the multistage reactor 501, making the local I/O ratio at the bed inlet equal to 212.
  • 0.267 kbd of alkylate is produced while consuming 0.176 kbd of isobutane.
  • the multistage reactor 501 effluent 507 (3 kbd of isobutane + 0.267kbd of alkylate) is fed to the front end of the liquid acid reactor 506 together with the remaining 1 kbd of isobutane in the recycle loop 509 and the remaining 1 kbd of olefin feed 502b such that the I/O ratio going into the liquid acid reactor 506 remains at 4/1.
  • a total of 2.047 kbd of alkylate 508 is produced as a bottoms product in the DiB tower 512 at constant DiB overhead recycle rate when compared against Example 7.
  • FIG. 6 illustrates a combined liquid and solid alkylation system 600, according to one or more aspects of the present disclosure, combining an existing alkylation process using liquid acid catalyst(s) (e.g., as shown in FIG. 2) with a multistage solid acid catalyst alkylation process.
  • liquid acid catalyst(s) e.g., as shown in FIG. 2
  • FIG. 6 illustrates that a refiner operating the liquid acid reactor of FIG. 2 (Example 7) can increase the production of alkylate by 10%, even though unable to feed additional olefin to the liquid acid reactor due to a variety of constraints e.g., refrigeration limit, acid consumption/hydraulic limit, feed hydraulic limit, and the like).
  • An additional 0.117 kbd of isobutane is required (compared to Example 7) to react the additional 0.1 kbd of olefin 603, increasing the isobutane makeup rate to 1.287 kbd.
  • the multistage reactor 601 comprises ten (10) separate beds containing solid acid catalyst, making the olefin flow to each bed 0.01 kbd, assuming equal flow to each bed.
  • the purity of the isobutane recycle stream 610 (2.83 kbd) is not sufficient to allow it to be used in the multistage reactor 601 (e.g., it may be contaminated with aqueous and/or acidic species).
  • all of the isobutane makeup stream 604 is directed to the multistage reactor 601.
  • 1.287 kbd isobutane makeup feed 604 is fed to the multistage reactor 601, making the local I/O ratio at the bed inlet equal to 129.
  • the I/O ratio may be further increased by increasing the number of olefin injection points or via other means.
  • 0.178 kbd of alkylate is produced while consuming 0.117 kbd of isobutane.
  • the multistage reactor 601 effluent 607 (1.17 kbd of isobutane + 0.178 kbd of alkylate) is fed to the front end of the liquid acid reactor 606 together with all of the 2.83 kbd of isobutane in the recycle loop 610 and the remaining 1 kbd of olefin feed 602b such that the I/O ratio going into the liquid acid reactor 606 remains at 4/1.
  • a total of 1.958 kbd of alkylate 608 is produced as a bottoms product in the DiB tower 612 at constant DiB overhead recycle rate when compared against Example 7.
  • compositions and methods are described herein in terms of “comprising” various components or steps, the compositions and methods can also “consist essentially of’ or “consist of’ the various components and steps.
  • compositions and methods are described in terms of “comprising,” “containing,” or “including” various components or steps, the compositions and methods can also “consist essentially of’ or “consist of’ the various components and steps. All numbers and ranges disclosed above may vary by some amount. Whenever a numerical range with a lower limit and an upper limit is disclosed, any number and any included range falling within the range is specifically disclosed. In particular, every range of values (of the form, “from about a to about b,” or, equivalently, “from approximately a to b,” or, equivalently, “from approximately a-b”) disclosed herein is to be understood to set forth every number and range encompassed within the broader range of values.

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Abstract

La présente divulgation concerne des systèmes, des procédés et des dispositifs destinés à l'alkylation d'une isoparaffine. Les systèmes comprennent un réacteur à étages multiples pour l'alkylation d'une isoparaffine à l'aide d'un catalyseur acide solide en communication fluidique avec un réacteur à une seule étage pour l'alkylation d'une isoparaffine à l'aide d'un catalyseur acide liquide. Le réacteur à une seule étage est conçu pour recevoir au moins une partie d'un mélange d'alkylation produit à l'intérieur du réacteur à étages multiples. Les procédés comprennent l'introduction, dans un réacteur à étages multiples, d'un catalyseur acide solide à une charge d'isoparaffine et une charge d'oléfine pour former un premier mélange d'alkylation comprenant un alkylate et une isoparaffine, et l'introduction, dans un réacteur à une seule étage ou à étages multiples, d'un catalyseur acide liquide à une charge d'isoparaffine, une charge d'oléfine, et au moins une partie du premier mélange d'alkylation pour former un second mélange d'alkylation comprenant un alkylate et une isoparaffine.
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