Patent Application
20190359542
This disclosure relates to transalkylation of heavy (C9+) aromatic hydrocarbon feedstocks to produce xylenes, particularly para-xylene.
Xylenes are important aromatic hydrocarbons, for which the worldwide demand is steadily increasing. The demand for xylenes, particularly para-xylene, has increased in proportion to the increase in demand for polyester fibers and film and typically grows at a rate of 5-7% per year. An important source of xylenes and other aromatic hydrocarbons is catalytic reformate, which is produced by contacting a mixture of petroleum naphtha and hydrogen with a strong hydrogenation/dehydrogenation catalyst, such as platinum, on a moderately acidic support, such as a halogen-treated alumina. The resulting reformate is a complex mixture of paraffins and the C6 to C8 aromatics, in addition to a significant quantity of heavier aromatic hydrocarbons. After removing the light (C5−) paraffinic components, the remainder of the reformate is normally separated into C7−, C8 and C9+-containing fractions using a plurality of distillation steps. The C8-containing fraction is then fed to a xylene production loop where para-xylene is recovered, generally by adsorption or crystallization, and the resultant para-xylene depleted stream is subjected to catalytic conversion to isomerize the xylenes back towards equilibrium distribution and to reduce the level of ethylbenzene that would otherwise build up in the xylene production loop.
However, the quantity of xylene available from reformate C8 fractions is limited and so recently refineries have also focused on the production of xylene by transalkylation of heavy (C9+) aromatic hydrocarbons (both from reformate and other sources) with benzene and/or toluene over noble metal-containing zeolite catalysts. For example, U.S. Pat. No. 5,942,651 discloses a process for the transalkylation of heavy aromatics comprising contacting a feed comprising C9+ aromatic hydrocarbons and toluene with a first catalyst composition comprising a molecular sieve having a constraint index ranging from 0.5 to 3, such as ZSM-12, and a hydrogenation component under transalkylation reaction conditions to produce a transalkylation reaction product comprising benzene and xylene. The transalkylation reaction product is then contacted with a second catalyst composition which comprises a molecular sieve having a constraint index ranging from 3 to 12, such as ZSM-5, and which may be in a separate bed or a separate reactor from the first catalyst composition, under conditions to remove benzene co-boilers in the product.
One problem associated with heavy aromatics transalkylation processes is catalyst aging since, as the catalyst loses activity with increasing time on stream, higher temperatures tend to be required to maintain constant conversion. When the maximum reactor temperature is reached, the catalyst needs to be replaced or regenerated, normally by oxidation. In particular, it has been found that the aging rate of existing transalkylation catalysts is strongly dependent on the presence in the feed of aromatic compounds having alkyl substituents with two or more carbon atoms, such as ethyl and propyl groups. Thus these compounds tend to undergo reactions such as disproportionation and dealkylation/realkylation to produce C10+ coke precursors.
To address the problem of C9+ feeds containing high levels of ethyl and propyl substituents, US Publication No. 2009/0112034 discloses a catalyst system adapted for transalkylation of a C9+ aromatic feedstock with a C6-C7 aromatic feedstock comprising: (a) a first catalyst comprising a first molecular sieve having a Constraint Index in the range of 3-12 and 0.01 to 5 wt. % of at least one source of a first metal element of Groups 6-10; and (b) a second catalyst comprising a second molecular sieve having a Constraint Index less than 3 and 0 to 5 wt. % of at least one source of a second metal element of Groups 6-10, wherein the weight ratio of said first catalyst to said second catalyst is in the range of 5:95 to 75:25. The first catalyst, which is optimized for dealkylation of the ethyl and propyl groups in the feed, is located in front of the second catalyst, which is optimized for transalkylation, when they are brought into contact with a C9+ aromatic feedstock and a C6-C7 aromatic feedstock in the presence of hydrogen.
However, despite these and other advances, there remains a number of unresolved problems with existing C9+ aromatic conversion processes. One such problem is the high costs associated with handling the large recycle streams associated with existing processes. Thus, even under the vapor phase conditions employed in existing processes, the level of conversion of the C9+ aromatic hydrocarbons in the transalkylation step is relatively low (typically of the order of 40 wt. % to 70 wt. %). As a result, the effluent from the transalkylation reactor contains high levels of residual C9+ aromatic hydrocarbons which must be separated and revaporized before being recycled to the transalkylation step. In addition, supplying fresh and in some cases recycle benzene and toluene to the combined dealkylation/transalkylation process not only increases vaporization loads and hence energy costs, but increases ring loss and hydrogen usage over the entire process.
According to the present disclosure, it has now been found that the capital and operating costs associated with C9+ aromatic hydrocarbon conversion processes can be reduced by separating the dealkylation and transalkylation steps and conducting the transalkylation step in the liquid phase. In this way, the need for vaporization of recycled C9+ aromatic hydrocarbons can be eliminated or at least significantly reduced. In addition, by supplying the fresh and any recycled benzene and toluene to the liquid phase transalkylation step, energy usage required for feed vaporization can be further reduced along with ring loss and hydrogen usage associated with the dealkylation step. Also, since it only processes fresh C9+ aromatic hydrocarbons, with minimal recycle, it is expected that the size of the dealkylation reactor can be reduced.
Thus, in one aspect, the present disclosure relates to a process for producing xylenes from C9+ aromatic hydrocarbons, the process comprising:
(a) contacting a first feedstock comprising C9+ aromatic hydrocarbons with a first catalyst in the presence of 0 wt. % or more of hydrogen under effective vapor phase dealkylation conditions to dealkylate part of the C9+ aromatic hydrocarbons and produce a first product comprising benzene, toluene and residual C9+ aromatic hydrocarbons; and
(b) contacting a second feedstock comprising C9+ aromatic hydrocarbons and benzene and/or toluene with a second catalyst under effective liquid phase C9+ transalkylation conditions to transalkylate at least part of the C9+ aromatic hydrocarbons and produce a second product comprising xylenes.
In a further aspect, the present disclosure relates to a process for producing xylenes from C9+ aromatic hydrocarbons, the process comprising:
(a) contacting a first feedstock comprising C9+ aromatic hydrocarbons with a first catalyst in the presence of 0 wt. % or more of hydrogen under effective vapor phase dealkylation conditions to dealkylate part of the C9+ aromatic hydrocarbons and produce a first product comprising benzene, toluene and residual C9+ aromatic hydrocarbons;
(b) contacting at least part of the C9+ aromatic hydrocarbons from the first product together with benzene and/or toluene with a second catalyst under effective liquid phase C9+ transalkylation conditions to transalkylate at least part of the C9+ aromatic hydrocarbons and produce a second product comprising benzene, toluene, xylenes and residual C9+ aromatic hydrocarbons; and
(c) separating xylenes from the second product.
As used herein, the numbering scheme for the Periodic Table Groups is as described in Chemical and Engineering News, 63(5), 27 (1985).
As used in this specification, the term “framework type” is used in the sense described in the “Atlas of Zeolite Framework Types,” 2001.
The term “aromatic” is used herein in accordance with its art-recognized scope which includes alkyl substituted and unsubstituted mono- and polynuclear compounds.
The term “catalyst” is used interchangeably with the term “catalyst composition”.
The term “ethyl-aromatic compounds” means aromatic compounds having an ethyl group attached to the aromatic ring. The term “propyl-aromatic compounds” means aromatic compounds having a propyl group attached to the aromatic ring.
The term “Cn” hydrocarbon wherein n is an positive integer, e.g., 1, 2, 3, 4, 5, 6, 7, 8, 9, 10, 11, 12, as used herein means a hydrocarbon having n number of carbon atom(s) per molecular. For example, Cn aromatics means an aromatic hydrocarbon having n number of carbon atom(s) per molecule. The term “Cn+” hydrocarbon wherein n is an positive integer, e.g., 1, 2, 3, 4, 5, 6, 7, 8, 9, 10, 11, 12, as used herein means a hydrocarbon having at least n number of carbon atom(s) per molecule. The term “Cn-” hydrocarbon wherein n is an positive integer, e.g., 1, 2, 3, 4, 5, 6, 7, 8, 9, 10, 11, 12, as used herein means a hydrocarbon having no more than n number of carbon atom(s) per molecule.
The term “effective vapor phase dealkylation conditions” means that the relevant reaction is conducted under conditions of temperature and pressure such that at least part of the aromatic components of the reaction mixture are in the vapor phase. In some embodiments, the mole fraction of the aromatic components in the vapor phase, relative to the total aromatics in the reaction mixture, can be at least 0.75, such as at least 0.85, up to 1 (all the aromatic components in the vapor phase).
The term “effective liquid phase C9+ transalkylation conditions” means that the transalkylation reaction is conducted under conditions of temperature and pressure such that at least part of the aromatic components of the transalkylation reaction mixture are in the vapor phase. In some embodiments, the mole fraction of aromatic compounds in the liquid phase, relative to the total aromatics, can be at least 0.01, or at least 0.05, or at least 0.08, or at least 0.1, or at least 0.15, or at least 0.2, or at least 0.3, or at least 0.4, or at least 0.5, and optionally up to having substantially all aromatic compounds in the liquid phase.
The term “mordenite” as used herein includes, but is not limited to, a mordenite zeolite having a very small crystal size and having a high mesopore surface area made by the particular selection of the synthesis mixture composition, as disclosed in WO 2016/126431.
The term “xylenes” as used herein is intended to include a mixture of the isomers of xylene of ortho-xylene, meta-xylene and para-xylene.
Described herein are various processes for producing xylene from C9+ aromatic hydrocarbons. In these processes, a first feedstock comprising C9+ aromatic hydrocarbons is contacted with a first catalyst in the presence of hydrogen under effective vapor phase dealkylation conditions to dealkylate part of the C9+ aromatic hydrocarbons and produce a first product comprising benzene, toluene and residual C9+ aromatic hydrocarbons. A second feedstock comprising C9+ aromatic hydrocarbons, such as from the first product, together with benzene and/or toluene is then contacted with a second catalyst in the presence of hydrogen under effective liquid phase C9+ transalkylation conditions to transalkylate at least part of the C9+ aromatic hydrocarbons and produce a second product comprising xylenes. Para-xylene is then recovered from the second product.
The aromatic feed used in the present process comprises one or more aromatic hydrocarbons containing at least 9 carbon atoms. Specific C9+ aromatic compounds found in a typical feed include mesitylene (1,3,5-trimethylbenzene), durene (1,2,4,5-tetramethylbenzene), hemimellitene (1,2,4-trimethylbenzene), pseudocumene (1,2,4-trimethylbenzene), 1,2-methylethylbenzene, 1,3-methylethylbenzene, 1,4-methylethylbenzene, propyl-substituted benzenes, butyl-substituted benzenes, and dimethylethylbenzenes. Suitable sources of the C9+ aromatics are any C9+ fraction from any refinery process that is rich in aromatics. This aromatics fraction may contain a substantial proportion of C9+ aromatics, e.g., at least 50 wt. %, such as at least 80 wt. % C9+ aromatics, wherein preferably at least 80 wt. %, and more preferably more than 90 wt. %, of the hydrocarbons will range from C9 to C12. Typical refinery fractions which may be useful include catalytic reformate, FCC naphtha or TCC naphtha.
The first stage of the present process comprises contacting the C9+ aromatic hydrocarbon feedstock in a first reaction zone with a first catalyst effective to dealkylate C2+ alkyl-containing compounds, particularly ethyl-aromatic compounds and propyl-aromatic compounds, to produce mainly benzene and toluene and the corresponding alkenes. The total feed to the first reaction zone therefore normally includes 0 wt. % or more of hydrogen to convert the alkenes to the corresponding alkanes Thus, for example, when hydrogen is present, the hydrogen/hydrocarbon molar ratio in the total feed to the first reaction zone may be from 0.05 to 10, for example from 0.1 to 5.
Any known dealkylation catalyst can be used in the first reaction zone. However, in embodiments, the first catalyst comprises a first molecular sieve having a Constraint Index in the range of about 3 to about 12, optionally together with at least one hydrogenation component. In this respect, Constraint Index is a convenient measure of the extent to which an aluminosilicate or other molecular sieve provides controlled access to molecules of varying sizes to its internal structure. For example, molecular sieves which provide a highly restricted access to and egress from its internal structure have a high value for the Constraint Index. Molecular sieves of this kind usually have pores of small diameter, e.g., less than 5 Angstroms. On the other hand, molecular sieves which provide relatively free access to their internal pore structure have a low value for the constraint index, and usually pores of large size. The method by which constraint index is determined is described fully in U.S. Pat. No. 4,016,218, which is incorporated herein by reference for the details of the method.
Suitable molecular sieves for use in the first catalyst composition comprise at least one of ZSM-5, ZSM-11, ZSM-22, ZSM-23, ZSM-35, ZSM-48, ZSM-57 and ZSM-58. ZSM-5 is described in detail in U.S. Pat. No. 3,702,886 and Re. 29,948. ZSM-11 is described in detail in U.S. Pat. No. 3,709,979. ZSM-22 is described in U.S. Pat. Nos. 4,556,477 and 5,336,478. ZSM-23 is described in U.S. Pat. No. 4,076,842. ZSM-35 is described in U.S. Pat. No. 4,016,245. ZSM-48 is more particularly described in U.S. Pat. Nos. 4,234,231 and 4,375,573. ZSM-57 is described in U.S. Pat. No. 4,873,067. ZSM-58 is described in U.S. Pat. No. 4,698,217.
In one preferred embodiment, the first molecular sieve comprises ZSM-5 and especially ZSM-5 having an average crystal size of less than 0. 1 micron, for example such that the ZSM-5 crystals have an external surface area in excess of 100 m2/g as determined by the t-plot method for nitrogen physisorption. Suitable ZSM-5 compositions are disclosed in US Publication No. 2015/0298981, the entire contents of which are incorporated herein by reference.
Conveniently, the first molecular sieve has an alpha value in the range of about 100 to about 1500, such as about 150 to about 1000, for example about 150 to about 600. Alpha value is a measure of the cracking activity of a catalyst and is described in U.S. Pat. No. 3,354,078 and in the Journal of Catalysis, Vol. 4, p. 527 (1965); Vol. 6, p. 278 (1966); and Vol. 61, p. 395 (1980), each incorporated herein by reference as to that description. The experimental conditions of the test used herein include a constant temperature of 538° C. and a variable flow rate as described in detail in the Journal of Catalysis, Vol. 61, page 395.
Generally, the first molecular sieve is an aluminosilicate having a silica to alumina molar ratio of less than 1000, typically from about 10 to about 100.
Typically, the first catalyst composition comprises at least 1 wt. %, preferably at least 10 wt. %, more preferably at least 25 wt. %, and most preferably at least 50 wt. %, of the first molecular sieve. In one embodiment, the first catalyst composition comprises from 55 to 80 wt. % of the first molecular sieve.
Additionally or alternatively, the first catalyst composition may comprise a combination of the first molecular sieve having a Constraint Index in the range of about 3 to about 12 and an additional molecular sieve having a Constraint Index less than 3, such zeolite beta, mordenite or faujasite. For example, the first catalyst composition may comprise a combination of ZSM-5 and mordenite.
In addition to a molecular sieve having a Constraint Index in the range of about 3 to about 12, the first catalyst composition comprises at least one hydrogenation component, such as at least one metal or compound thereof of Groups 6 to 12 of the Periodic Table of the Elements. Suitable hydrogenation components include platinum, palladium, iridium, rhenium and mixtures and compounds thereof, preferably platinum, rhenium and compounds thereof. In some embodiments, the first catalyst composition comprises two or more hydrogenation components including a first metal or compound thereof selected from platinum, palladium, iridium, rhenium and mixtures thereof and a second metal or compound chosen so as to lower the benzene saturation activity of the first metal. Examples of suitable second metals include at least one of copper, silver, gold, ruthenium, iron, tungsten, molybdenum, cobalt, nickel, tin and zinc. Conveniently, the first metal is present in the first catalyst in an amount from 0.001 to 1 wt %, such as from 0.01 to 0.1 wt %, of the first catalyst and the second metal is present in the first catalyst in amount from 0.001 to 10 wt %, 0.1 to 1 wt %, of the first catalyst.
In some embodiments, the first metal comprises platinum and/or rhenium and the second metal comprises copper and/or tin. In one preferred embodiment, the first metal comprises platinum and the second metal comprises tin, desirably at a molar ratio of platinum to tin from 0.1:1 to 1:1, such as from 0.2:1 to 0.4:1.
In some embodiments, the first catalyst composition may comprise one or more of the hydrogenation components described above on a refractory oxide, with or without the presence of a molecular sieve. Suitable refractory oxides comprise silica, alumina, silica-alumina and titania.
One or more of the hydrogenation components can be incorporated into the first catalyst composition by any known method, including co-crystallization, ion exchange into the composition to the extent a Group 13 element, e.g., aluminum, is in the molecular sieve structure, impregnated therein, or mixed with the molecular sieve and binder. In some embodiments, ion exchange may be preferred. After incorporation of the hydrogenation component(s), the catalyst composition is usually dried by heating at a temperature of 65° C. to 160° C., typically 110° C. to 143° C., for at least 1 minute and generally not longer than 24 hours, at pressures ranging from 100 to 200 kPa-a. Thereafter, the catalyst composition may be calcined in a stream of dry gas, such as air or nitrogen, at temperatures of from 260° C. to 650° C. for 1 to 20 hours. Calcination is typically conducted at pressures ranging from 100 to 300 kPa-a.
The first catalyst composition may be self-bound (that is without a separate binder) or may also comprise a binder or matrix material that is resistant to the temperatures and other conditions employed in the present process. Where such a binder or matrix material is present, it is substantially free of amorphous alumina, since it is found that the exclusion of a binder containing amorphous alumina reduces external catalytic sites for coke production and hence increases catalyst cycle length. One preferred binder material for the first catalyst composition comprises silica since extrusion with silica ensures that the catalyst has high mesoporosity and hence high activity. Alternatively, the binder or matrix material may be a crystalline molecular sieve material, which may be isostructural with, or have a different structure than, the first molecular sieve.
Where the first catalyst composition contains a binder or matrix material, the latter may be present in an amount ranging from 5 to 95 wt. %, and typically from 10 to 60 wt. %, of the total catalyst composition.
Examples of specific catalyst compositions useful in the dealkylation stage of the present process include Pt supported on ZSM-5, Pt-supported on silica or alumina, Pt/Sn on a combination of mordenite and ZSM-5, Re on a combination of mordenite and ZSM-5, and Mo on a combination of mordenite and ZSM-5.
The first catalyst composition may be extruded into particles of any desired shape before being loaded into the first catalyst bed. In some embodiments, it may be desirable to control the shape and size of the catalyst particles so as to maximize the external surface area of the catalyst. For example, it may be desirable to control the catalyst particle configuration such that the particles have a surface to volume ratio of about 80 to <200 inch-1, preferably about 100 to 150 inch-1. Suitable particle configurations for achieving such a surface to volume ratio include grooved cylindrical extrudates and hollow or solid polylobal extrudates, such as quadrulobal extrudates.
In operation, the first reaction zone is maintained under vapor phase conditions effective to dealkylate aromatic hydrocarbons containing C2+ alkyl groups in the heavy aromatic feedstock and to saturate the resulting C2+ olefins. Suitable conditions for operation of the first catalyst bed comprise a temperature in the range of about 100 to about 800° C., preferably about 300 to about 500° C., a pressure in the range of about 790 to about 7000 kPa-a, preferably about 2170 to 3000 kPa-a, a H2:HC molar ratio in the range of about 0.01 to about 20, preferably about 1 to about 10, and a WHSV in the range of about 0.01 to about 100 hr-1, preferably about 2 to about 20 hr-1.
The dealkylation stage of the present process can be conducted in any known reactor system including, but not limited to, a fixed bed reactor, a moving bed reactor, a fluidized bed reactor and a reactive distillation unit, with a fixed bed reactor being preferred.
The product of the dealkylation stage mainly comprises residual (unreacted) C9+ aromatic hydrocarbons together with smaller amounts of benzene, toluene, xylenes, and lower alkanes and residual hydrogen. The dealkylation product is then fed to a separator where the lower alkanes and any residual hydrogen are removed before the remainder of the alkylation product is supplied, in some embodiments without an intermediate separation step, to a second transalkylation reaction zone. Any separated hydrogen can then be recycled to the first reaction zone, or supplied to a another reaction zone, such as, for example, the liquid phase transalkylation reaction zone, discussed in detail below. The separated lower alkanes can be recovered for use as fuel.
In the second, liquid phase transalkylation stage of the present process, at least part of the unreacted C9+ aromatic hydrocarbons from the dealkylation product, benzene and/or toluene are contacted in a second liquid phase transalkylation reaction zone with a second catalyst composition comprising a second molecular sieve and optionally one or more hydrogenation components. Fresh benzene and/or toluene, preferably fresh toluene, can also be supplied to the transalkylation reaction zone.
In one embodiment, a suitable molecular sieve for the second catalyst composition includes a molecular sieve with a framework structure having a 3-dimensional network of 12-member ring pore channels. Examples of framework structures having a 3-dimensional 12-member ring are the framework structures corresponding to faujasite (such as zeolite X or Y, including USY), *BEA (such as zeolite Beta), BEC (polymorph C of Beta), CIT-1 (CON), MCM-68 (MSE), hexagonal faujasite (EMT), ITQ-7 (ISV), ITQ-24 (IWR), and ITQ-27 (IWV), preferably faujasite, hexagonal faujasite, and Beta (including all polymorphs of Beta). It is noted that the materials having a framework structure including a 3-dimensional network of 12-member ring pore channels can correspond to zeolites, silicoaluminophosphates, aluminophosphates, and/or any other convenient combination of framework atoms.
Additionally or alternately, a suitable transalkylation catalyst includes a molecular sieve with a framework structure having a 1-dimensional network of 12-member ring pore channels, where the pore channel has a pore channel size of at least 6.0 Angstroms, or at least 6.3 Angstroms. The pore channel size of a pore channel is defined herein to refer to the maximum size sphere that can diffuse along a channel Examples of framework structures having a 1-dimensional 12-member ring pore channel can include, but are not limited to, mordenite (MOR), zeolite L (LTL), and ZSM-18 (MEI). It is noted that the materials having a framework structure including a 1-dimensional network of 12-member ring pore channels can correspond to zeolites, silicoaluminophosphates, aluminophosphates, and/or any other convenient combination of framework atoms.
Additionally or alternately, a suitable transalkylation catalyst includes a molecular sieve having the MWW framework structure. Although the MWW framework structure does not have 12-member ring pore channels, the MWW framework structure does include surface sites that have features similar to a 12-member ring opening. Examples of molecular sieves having MWW framework structure include MCM-22, MCM-49, MCM-56, MCM-36, EMM-10, EMM-10-P, EMM-13, PSH-3, SSZ-25, ERB-1, ITQ-1, ITQ-2, UZM-8, UZM-8HS, UZM-37, MIT-1, and interlayer expanded zeolites. It is noted that the materials having an MWW framework structure can correspond to zeolites, silicoaluminophosphates, aluminophosphates, and/or any other convenient combination of framework atoms.
Additionally or alternately, a suitable transalkylation catalyst includes an acidic microporous material that has a largest pore channel corresponding to a 12-member ring or larger, and/or that has a pore channel size of at least 6.0 Angstroms, or at least 6.3 Angstroms and/or that has another active surface having a size of at least 6.0 Angstroms. It is noted that such microporous materials can correspond to zeolites, silicoaluminophosphates, aluminophosphates, and/or materials that are different from molecular sieve type materials.
The molecular sieve can optionally be characterized based on having a composition with a molar ratio YO2 over X2O3 equal to n, wherein X is a trivalent element, such as aluminum, boron, iron, indium and/or gallium, preferably aluminum and/or gallium, and Y is a tetravalent element, such as silicon, tin and/or germanium, preferably silicon. For example, when Y is silicon and X is aluminum, the molar ratio of YO2 over X203 is the silica-to-alumina molar ratio. For a MWW framework molecular sieve, n can be less than about 50, e.g., from about 2 to less than about 50, usually from about 10 to less than about 50, more usually from about 15 to about 40. For a molecular sieve having the framework structure of Beta and/or its polymorphs, n can be about 10 to about 60, or about 10 to about 50, or about 10 to about 40, or about 20 to about 60, or about 20 to about 50, or about 20 to about 40, or about 60 to about 250, or about 80 to about 250, or about 80 to about 220, or about 10 to about 400, or about 10 to about 250, or about 60 to about 400, or about 80 to about 400. For a molecular sieve having the framework structure FAU, n can be about 2 to about 400, or about 2 to about 100, or about 2 to about 80, or about 5 to about 400, or about 5 to about 100, or about 5 to about 80, or about 10 to about 400, or about 10 to about 100, or about 10 to about 80. Optionally, the above n values can correspond to n values for a ratio of silica to alumina in the MWW, *BEA, and/or FAU framework molecular sieve. In such optional aspects, the molecular sieve can optionally correspond to an aluminosilicate and/or a zeolite.
Optionally, the catalyst comprises 0.01 wt. % to 5.0 wt. %, or 0.01 wt. % to 2.0 wt. %, or 0.01 wt. % to 1.0 wt. %, or 0.05 wt. % to 5.0 wt. %, or 0.05 wt. % to 2.0 wt. %, or 0.05 wt. % to 1.0 wt. %, or 0.1 wt. % to 5.0 wt. %, or 0.1 to 2.0 wt. %, or 0.1 wt. % to 1.0 wt. %, of a metal element of Groups 5-11 (according to the IUPAC Periodic Table). The metal element may be at least one hydrogenation component, such as one or more metals selected from Group 5-11 and 14 of the Periodic Table of the Elements, or a mixture of such metals, such as a bimetallic (or other multimetallic) hydrogenation component. Optionally, the metal can be selected from Groups 8-10, such as a Group 8-10 noble metal. Specific examples of useful metals are iron, tungsten, vanadium, molybdenum, rhenium, chromium, manganese, ruthenium, osmium, nickel, cobalt, rhodium, iridium, copper, tin, noble metals such as platinum or palladium, and combinations thereof. Specific examples of useful bimetallic combinations (or multimetallic combinations) are those where Pt is one of the metals, such as Pt/Sn, Pt/Pd, Pt/Cu, and Pt/Rh. In some aspects, the hydrogenation component is palladium, platinum, rhodium, copper, tin, or a combination thereof. The amount of the hydrogenation component can be selected according to a balance between hydrogenation activity and catalytic functionality. For a hydrogenation component including two or more metals (such as a bimetallic hydrogenation component), the ratio of a first metal to a second metal can range from 1:1 to about 1:100 or more, preferably 1:1 to 1:10.
Optionally, a suitable transalkylation catalyst can be a molecular sieve that has a constraint index of 1-12, optionally but preferably less than 3. The constraint index can be determined by the method described in U.S. Pat. No. 4,016,218, which is incorporated herein by reference with regard to the details of determining a constraint index.
Additionally or alternately, a transalkylation catalyst (such as a transalkylation catalyst system) can be used that has a reduced or minimized activity for dealkylation. The Alpha value of a catalyst can provide an indication of the activity of a catalyst for dealkylation. In various aspects, the transalkylation catalyst can have an Alpha value of about 100 or less, or about 50 or less, or about 20 or less, or about 10 or less, or about 1 or less. The alpha value test is a measure of the cracking activity of a catalyst and is described in U.S. Pat. No. 3,354,078 and in the Journal of Catalysis, Vol. 4, p. 527 (1965); Vol. 6, p. 278 (1966); and Vol. 61, p. 395 (1980), each incorporated herein by reference as to that description. The experimental conditions of the test used herein include a constant temperature of 538° C. and a variable flow rate as described in detail in the Journal of Catalysis, Vol. 61, p. 395.
In addition to the second molecular sieve, it may be desirable to incorporate in the second catalyst composition another material that is resistant to the temperatures and other conditions employed in the transalkylation process of the disclosure. Such materials include active and inactive materials and synthetic or naturally occurring zeolites, as well as inorganic materials such as clays, silica, hydrotalcites, perovskites, spinels, inverse spinels, mixed metal oxides, and/or metal oxides such as alumina, lanthanum oxide, cerium oxide, zirconium oxide, and titania. The inorganic material may be either naturally occurring, or in the form of gelatinous precipitates or gels including mixtures of silica and metal oxides.
Use of a material in conjunction with each molecular sieve, i.e., combined therewith or present during its synthesis, which itself is catalytically active, may change the conversion and/or selectivity of the catalyst composition. Inactive materials suitably serve as diluents to control the amount of conversion so that transalkylated products can be obtained in an economical and orderly manner without employing other means for controlling the rate of reaction. These catalytically active or inactive materials may be incorporated into, for example, alumina, to improve the crush strength of the catalyst composition under commercial operating conditions. It is desirable to provide a catalyst composition having good crush strength because in commercial use, it is desirable to prevent the catalyst composition from breaking down into powder-like materials.
Naturally occurring clays that can be composited with each molecular sieve as a binder for the catalyst composition include the montmorillonite and kaolin family, which families include the subbentonites, and the kaolins commonly known as Dixie, McNamee, Georgia and Florida clays or others in which the main mineral constituent is halloysite, kaolinite, dickite, nacrite or anauxite. Such clays can be used in the raw state as originally mined or initially subjected to calcination, acid treatment or chemical modification.
In addition to the foregoing materials, each molecular sieve (and/or other microporous material) can be composited with a binder or matrix material, such as an inorganic oxide selected from the group consisting of silica, alumina, zirconia, titania, thoria, beryllia, magnesia, lanthanum oxide, cerium oxide, manganese oxide, yttrium oxide, calcium oxide, hydrotalcites, perovskites, spinels, inverse spinels, and combinations thereof, such as silica-alumina, silica-magnesia, silica-zirconia, silica-thoria, silica-beryllia, silica-titania, as well as ternary compositions such as silica-alumina-thoria, silica-alumina-zirconia, silica-alumina-magnesia and silica-magnesia-zirconia. It may also be advantageous to provide at least a part of the foregoing porous matrix binder material in colloidal form so as to facilitate extrusion of the catalyst composition.
In some aspects, a molecular sieve (and/or other microporous material) can be used without an additional matrix or binder. In other aspects, a molecular sieve/microporous material can be admixed with a binder or matrix material so that the final catalyst composition contains the binder or matrix material in an amount ranging from 5 to 95 wt. %, and typically from 10 to 60 wt. %.
Prior to use, steam treatment of the catalyst composition may be employed to minimize the aromatic hydrogenation activity of the catalyst composition. In the steaming process, the catalyst composition is usually contacted with from 5 to 100% steam, at a temperature of at least 260° C. to 650° C. for at least one hour, specifically 1 to 20 hours, at a pressure of 100 to 2590 kPa-a.
A hydrogenation component can be incorporated into the catalyst composition by any convenient method. Such incorporation methods can include co-crystallization, exchange into the catalyst composition, liquid phase and/or vapor phase impregnation, or mixing with the molecular sieve and binder, and combinations thereof. For example, in the case of platinum, a platinum hydrogenation component can be incorporated into the catalyst by treating the molecular sieve with a solution containing a platinum metal-containing ion. Suitable platinum compounds for impregnating the catalyst with platinum include chloroplatinic acid, platinous chloride and various compounds containing the platinum ammine complex, such as Pt(NH3)4Cl2.H2O or (NH3)4Pt(NO3)2.H2O. Palladium can be impregnated on a catalyst in a similar manner
Alternatively, a compound of the hydrogenation component may be added to the molecular sieve when it is being composited with a binder, or after the molecular sieve and binder have been formed into particles by extrusion or pelletizing. Still another option can be to use a binder that is a hydrogenation component and/or that includes a hydrogenation component.
After treatment with the hydrogenation component, the catalyst is usually dried by heating at a temperature of 65° C. to 160° C., typically 110° C. to 143° C., for at least 1 minute and generally not longer than 24 hours, at pressures ranging from 100 to 200 kPa-a. Thereafter, the molecular sieve may be calcined in a stream of dry gas, such as air or nitrogen, at temperatures of from 260° C. to 650° C. for 1 to 20 hours. Calcination is typically conducted at pressures ranging from 100 to 300 kPa-a.
In addition, prior to contacting the catalyst composition with the hydrocarbon feed, the hydrogenation component can optionally be sulfided. This is conveniently accomplished by contacting the catalyst with a source of sulfur, such as hydrogen sulfide, at a temperature ranging from about 320° C. to 480° C. The source of sulfur can be contacted with the catalyst via a carrier gas, such as hydrogen or nitrogen. Sulfiding per se is known and sulfiding of the hydrogenation component can be accomplished without more than routine experimentation by one of ordinary skill in the art in possession of the present disclosure.
Generally, the conditions employed in a liquid phase transalkylation process can include a temperature of about 400° C. or less, or about 360° C. or less, or about 320° C. or less, and/or at least about 100° C., or at least about 200° C., such as between 100° C. to 400° C., or 100° C. to 340° C., or 230° C. to 300° C.; a pressure of 2.0 MPa-g to 10.0 MPa-g, or 3.0 MPa-g to 8.0 MPa-g, or 3.5 MPa-g to 6.0 MPa-g; an H2:hydrocarbon molar ratio of 0 to 20, or 0.01 to 20, or 0.1 to 10; and a weight hourly space velocity (“WHSV”) for total hydrocarbon feed to the reactor(s) of 0.1 to 100 hr-1, or 1 to 20 hr-1. Optionally, the pressure during transalkylation can be at least 4.0 MPa-g. It is noted that H2 is not necessarily required during the transalkylation reaction, so optionally the transalkylation can be performed without introduction of H2. The feed can be exposed to the transalkylation catalyst under fixed bed conditions, fluidized bed conditions, or other conditions that are suitable when a substantial liquid phase is present in the reaction environment.
In addition to staying within the general conditions above, the transalkylation conditions can be selected so that a desired amount of the hydrocarbons (reactants and products) in the reactor are in the liquid phase. Referring now to
As shown in
The resulting effluent from the liquid phase transalkylation process can have a xylene yield, relative to the total weight of the hydrocarbons in the effluent, of at least about 4 wt. %, or at least about 6 wt. %, or at least about 8 wt. %, or at least about 10 wt. %. Other major components of the transalkylation effluent include benzene, toluene and residual C9+ aromatic hydrocarbons. Separation of these components can be achieved using any conventional separation system, such as a distillation train. In particular, the xylene can be recovered and supplied to a para-xylene recovery loop, while the toluene and residual C9+ aromatic hydrocarbons can be recycled to the liquid phase transalkylation reactor. In some cases, part of the residual C9+ aromatic hydrocarbons may need to be recycled to the dealkylation stage to further reduce the level of ethyl and propyl-substituted aromatics. The benzene can be recovered for other uses or can be recycled to the liquid phase transalkylation stage.
Embodiments of the vapor phase dealkylation/liquid phase transalkylation process disclosed herein have a number of advantages. For example, since vaporization of recycle C9+ aromatics is reduced or eliminated and the need to vaporize fresh and recycle benzene and/or toluene is avoided, energy requirements are reduced. In addition, minimizing feed to the vapor phase (dealkylation) reactor will result in lower ring loss and lower hydrogen consumption. A smaller hydrogen compressor will therefore be required in a grassroots application. Additionally, a once through hydrogen option may be considered, thereby eliminating the need for a compressor. Further, processing only C9+ aromatics in the vapor phase dealkylation reactor to dealkylate ethyl and propyl groups, without transalkylation, will allow optimization of both the catalyst and reactor. Compared to the traditional scheme, the vapor phase dealkylation reactor is expected to be smaller as it only processes fresh C9+ aromatics, with minimal recycle. Although the addition of the liquid phase transalkylation reactor will add capital cost, its impact is expected to me minimal.
One embodiment of the present process for producing xylenes, and particularly para-xylenes, from C9+ aromatic hydrocarbons is shown in
The effluent from the dealkylation reaction zone 12 is collected in line 14 and fed to a separator 15 where the hydrogen and lower alkanes are removed and the hydrogen is recycled via line 16 to the dealkylation reaction zone 12. The remainder of the dealkyation effluent is collected in line 17 and fed to a transalkylation reaction zone 18, which also receives a supply of fresh toluene via line 19, a supply of recycled toluene via line 21 and a supply of recycled C9+ aromatic hydrocarbons via line 22. The transalkylation reaction zone 18 houses a second catalyst composition comprising a molecular sieve, typically having a Constraint Index less than 3, such as MCM-49, and a hydrogenation component, such as platinum or palladium. The reaction zone 18 is maintained under liquid phase conditions effective for transalkylation of at least part of the C9+ aromatic hydrocarbons supplied by lines 17 and 22 with at least part of the toluene supplied by lines 17, 19 and 21 to produce benzene and an equilibrium mixture of xylene isomers.
The effluent from the transalkylation reaction zone 18 is collected in line 23 and fed to a fractionation system 24, where unreacted C9+ aromatic hydrocarbons and toluene are removed via lines 22 and 21, respectively, for recycle to the reaction zone 18, benzene is recovered via line 25 and the desired xylene product is removed via line 26. A heavies stream may be purged through line 27.
The equilibrium xylene stream collected in line 26 typically has from above equilibrium amounts of para-xylene (above 24 wt. %) up to about 60 wt. %. This para-xylene rich C8 stream is initially supplied to a para-xylene recovery unit, such as, for example, a para-xylene extraction unit or a simulated moving bed column (SMB) 28, where the para-xylene is selectively adsorbed and, after treatment with a suitable desorbant, such as for example paradiethylbenzene, paradifluorobenzene, diethylbenzene or toluene or mixtures thereof, is recovered via line 29 for further purification. After separation of the para-xylene, the remaining para-xylene depleted steam is fed by line 31 to a xylene isomerization section (not shown) which can be operated in the gas phase or the liquid phase to isomerize ortho-xylene and meta-xylene in para-xylene depleted steam to form additional para-xylene. The isomerized stream can then be recycled back to the SMB 28 to recover the additional para-xylene.
While the present disclosure has included description and illustration of particular embodiments, those of ordinary skill in the art will appreciate that this disclosure lends itself to variations not necessarily illustrated herein. For this reason, then, reference should be made solely to the appended claims for purposes of determining the scope of the invention.
This application claims priority to and the benefit of U.S. Provisional Application Ser. No. 62/403,757, filed Oct. 4, 2016, the disclosure of which is incorporated herein by reference in its entirety.
| Filing Document | Filing Date | Country | Kind |
|---|---|---|---|
| PCT/US2017/017309 | 2/10/2017 | WO | 00 |
| Number | Date | Country | |
|---|---|---|---|
| 62403757 | Oct 2016 | US |
