KR20090100449A - Y-85 and modified lz-210 zeolites - Google Patents

Abstract

Catalysts for converting polyalkylaromatics to monoalkylaromatics, particularly cumene and ethyl benzene are disclosed which comprise aY-85 or a modified LZ-210 zeolite. For cumene and ethylbenzene production, a disclosed catalyst, made of 80 wt% zeolite and 20 wt% alumina binder on a volatile-free basis, has one or more of the following physical characteristics: (1) an absolute intensity of the Y-85 or modified LZ-210 zeolite as measured by X-ray diffraction (XRD) of preferably at least 50 and (2) a framework aluminum of the Y-85 or modified LZ-210 zeolite of preferably at least 60 % of the aluminum of the Y-85 or modified LZ-210 zeolite.

Description

Y-85 and modified L-210 zeolites {Y-85 AND MODIFIED LZ-210 ZEOLITES}

Y-85 and modified LZ-210 zeolites that can be used as catalysts in the transalkylation of polyalkylaromatics such as PIPB and PEB to cumene and ethyl benzene are disclosed with their preparation.

The following description specifically refers to the use of the catalysts disclosed herein in the transalkylation of polyisopropylbenzenes (PIPB) with benzene, which can yield cumene, but this is done only for the purpose of clarity and simplification. Should be considered. The broader scope of the present application will be often referred to herein for emphasis.

Cumene is a major commercial product, one of its main uses being a source of phenol and acetone through acid-catalyzed decomposition of intermediate peroxides after its air oxidation.

Since both phenols and acetones are important as commercial chemicals, great emphasis has been placed on the preparation of cumene, and the document mainly describes its manufacturing process. The most common and probably the most direct method of cumene production is the alkylation of benzene with propylene, in particular using acid catalysts.

Another common method for preparing cumene is the transalkylation of benzene with PIPB, in particular with di-isopropylbenzene (DIPB) and tri-isopropylbenzene (TIPB), in particular using acid catalysts. Any commercially suitable transalkylation process must satisfy the requirements of high conversion of polyalkylated aromatics and high selectivity of monoalkylated products.

The main direction of the reaction of benzene with PIPB yielding cumene corresponds to the Markownikoff addition of propyl groups. However, a small but very significant amount of reaction occurs through the addition of half-markicon to produce n-propylbenzene (NPB). It is meaningful that NPB formation interferes with the oxidation of cumene to phenol and acetone, and therefore cumene used for oxidation must be very pure in terms of NPB content.

Since cumene and NPB are difficult to separate by conventional means (eg, distillation), the production of cumene by transalkylation of benzene with PIPB should be carried out with a minimum of NPB production. One important factor to consider is that the use of acid catalysts for transalkylation increases NPB formation with increasing temperature. Therefore, in order to minimize NPB formation, the transalkylation should be carried out at the lowest possible temperature.

Since DIPB and TIPB are not only common feeds for the transalkylation of benzene by PIPB, but are also common by-products of the alkylation of benzene by propylene at cumene formation, transalkylation is usually carried out in conjunction with alkylation to produce less valuable byproducts. Minimize and generate additional cumene. In this combination process, cumene produced by both alkylation and transalkylation is typically recovered in a single product stream. Since NPB is also formed in alkylation and NPB formation in alkylation increases as the temperature rises, NPB production in both alkylation and transalkylation must be managed in relation to each other so that the NPB is relatively free from cumene product streams.

There is a need for an optimal transalkylation catalyst, for example for the production of cumene or ethylene benzene, having sufficient activity to affect transalkylation at an acceptable reaction rate at a temperature low enough to avoid unacceptable NPB formation. Since Y zeolites exhibit significantly higher activity than many other zeolites, they have been closely investigated as catalysts in aromatic transalkylation. However, there is a problem that Y zeolites result in unacceptably low rates of transalkylation at low temperatures desirable to minimize NPB formation.

Thus, in order for a commercial process based on Y zeolite to be realized, it is necessary to increase the catalytic activity, that is, increase the production rate of cumene or ethyl benzene at a predetermined low temperature.

Summary of the Invention

In order to meet the above requirements, a catalyst is disclosed which comprises a modified Y zeolite and which has a metal hydrogenation component of less than 0.2% by weight.

One modified Y zeolite is first prepared by ammonium ion exchange of sodium Y zeolite to produce a sodium-containing Y zeolite containing sodium cations, wherein the sodium content is based on the weight of the sodium-containing Y zeolite on anhydrous basis. NaO ₂ is 3% by weight and the Y zeolite has a first unit cell size. Next, the low sodium Y zeolite was hydrothermally steamed at a temperature in the range of 550 ° C. (1022 ° F.) to 850 ° C. (1562 ° F.) to contain sodium cations and to have a first bulk Si / Al ₂ molar ratio and a first unit. A steamed Y zeolite having a second unit cell size smaller than the cell size is produced. Finally, the steamed Y zeolite, wherein at least a portion of the sodium cation in the steamed Y zeolite is exchanged for ammonium ions and is greater than the first bulk Si / Al ₂ molar ratio, preferably in the range from 6.5 to 27 Contact is made with a sufficient amount of an aqueous solution of ammonium ion having a pH of less than 4, preferably 2-4, for a time sufficient to produce a modified Y zeolite having a 2 bulk Si / Al ₂ molar ratio. The unit cell size of the modified Y zeolite is in the range of 24.34 to 24.58 mm 3.

Another modified Y zeolite is prepared by treating a starting material such as Y-74 or Y-54 zeolite with an aqueous fluorosilicate solution to produce an LZ-210 zeolite having a first unit cell size. The fluorosilicate treated sample is then steamed at a temperature in the range of 550 ° C. (1022 ° F.) to 850 ° C. (1562 ° F.) to contain sodium cations and have a first bulk Si / Al ₂ molar ratio and a first unit cell size. To produce a steamed LZ-210 zeolite having a smaller second unit cell size. Finally, the steamed LZ-210 zeolite, wherein at least a portion of the sodium cations in the steamed LZ-210 zeolite is exchanged for ammonium ions and is greater than the first bulk Si / Al ₂ molar ratio and is in the range of 6.5-20 Contact with a sufficient amount of an aqueous solution of ammonium ion having a pH of less than 4 for a time sufficient to produce a modified LZ-210 zeolite having a 2 bulk Si / Al ₂ molar ratio. The unit cell size of the modified Y zeolite is in the range of 24.34 to 24.58 mm 3. Acid extraction can then be carried out to remove excess skeletal aluminum. Before or after treating Y zeolite with fluorosilicate salts, the catalyst is treated with ammonium ion exchange (s) to reduce the sodium content of the catalyst to less than or equal to 1% by weight of Na ₂ O at the same time as the first bulk Si The / Al ₂ molar ratio can be maintained. In another embodiment, the fluorosilicate treated Y zeolite (or LZ-210) may be ammonium exchanged without going through a steam treatment step to produce materials suitable for this disclosure by further lowering the Na ₂ O content.

The disclosed fabrication techniques affect the number and properties of extra skeletal aluminum (and Lewis acid moieties) as indicated by Si / Al ₂ changes and unit cell size changes to improve diffusion properties, increase catalytic activity and improve NPB formation. Lower.

One disclosed catalyst comprises a zeolite and a binder (1) at least 50 absolute intensities of modified Y zeolites measured by X-ray diffraction (XRD); And (2) skeletal aluminum in the modified Y zeolite, preferably at least 60%.

In one embodiment, the final catalyst for cumene production is a product of the absolute strength of the modified Y zeolite, measured by XRD, with the% skeletal aluminum in the aluminum of the modified zeolite, greater than 4200.

In another embodiment, the catalyst for ethyl benzene production has a product of the absolute strength of the modified Y zeolite, measured by XRD, with the% backbone aluminum in the aluminum of the modified Y zeolite, greater than 4500.

Another embodiment of the methods disclosed herein is described in the detailed description.

Brief description of the drawings

FIG. 1 graphically illustrates DIPB conversion (y-axis,%) versus temperature (x-axis, ° C.) for catalysts prepared according to Examples 2-4 and 7 of the disclosure for Control Examples 1 and 5.

2 shows the ratio of NPB to cumene in the product (y-axis, weight ppm) to DIPB conversion (x-axis,%) for the catalysts of Examples 2-4 and 7 of the present disclosure for Control Examples 1 and 5; Illustrate the graph.

FIG. 3 graphically plots DIPB conversion (y-axis,%) versus temperature (x-axis, ° C.) for the catalyst of Example 3 before and after regeneration (Example 7) and after regeneration (Example 9) for Control Example 1. FIG. To illustrate.

4 shows the ratio (y-axis, weight ppm) of NPB to cumene in the product for the catalyst of Example 3 before and after regeneration (Example 7) and after regeneration (Example 9) for Control Example 1 (x axis,%) is illustrated graphically.

FIG. 5 graphically illustrates DEB conversion (y-axis,%) vs. temperature (x-axis, ° C.) for the catalyst of Example 2 of the disclosure allowing the alkyl group and the catalyst disclosed above to work well in addition to propyl for Control Example 1. FIG. It was.

FIG. 6 graphically illustrates DIPB conversion (y-axis,%) versus temperature (x-axis, ° C.) for catalysts prepared according to Examples 14-16 of the present disclosure for Control Example 11. FIG.

FIG. 7 graphically illustrates NPB ratio (y-axis, weight ppm) versus DIPB conversion (x-axis,%) for cumene in the product for the catalyst of Examples 14-16 of the present disclosure for Control Example 11.

FIG. 8 graphically illustrates DIPB conversion (y-axis,%) versus temperature (x-axis, ° C.) for the catalyst of Example 14 before and after regeneration (Example 9) for Control Example 11. FIG.

9 shows the ratio of NPB to cumene (y-axis, ppm weight) to DIPB conversion (x-axis,%) in the product for the catalyst of Example 14 before and after regeneration (Example 19) for Control Example 11 Illustrated in the diagram.

Detailed description of the invention

Improved catalysts comprising crystalline zeolite molecular sieves are disclosed. Molecular sieves for use in the catalysts disclosed herein include Y zeolites such as Y-85 and modified LZ-210 zeolites.

Y-85 zeolite

Regarding the Y zeolites of the above disclosure, US type 3,130,007, which is incorporated herein by reference in its entirety, describes Y-type zeolites. Modified Y zeolites suitable for use in preparing the catalysts disclosed herein significantly modify the Y zeolite framework structure and composition from Y zeolites, and generally increase the bulk Si / Al ₂ molar ratio to typically greater than 6.5, And / or by treatment to reduce the unit cell size. However, in converting the Y zeolite starting material to modified Y zeolites useful in the methods disclosed herein, the resulting modified Y zeolites have an X-ray powder diffraction pattern for Y zeolites as described in the '007 patent. It will be understood that may not be exactly the same as. The modified Y zeolite may have an X-ray powder diffraction pattern similar to the '007 patent, but as the person skilled in the art will understand, the d interval is somewhat shifted due to cation exchange, firing, etc., which makes the Y zeolite catalytically active and stable. It is usually necessary to convert to form.

The modified Y zeolites disclosed herein have a unit cell size of 24.34 to 24.58 mm 3, preferably 24.36 to 24.55 mm 3. The modified Y zeolite has a bulk Si / Al ₂ molar ratio of 6.5 to 23.

In preparing the modified Y zeolite component in the catalyst disclosed above, the starting material may be Y zeolite in the form of an alkali metal (eg sodium) as described in the '007 patent. Y zeolites in the form of alkali metals are ion exchanged with ammonium ions or ammonium ion precursors, such as quaternary ammonium or other nitrogen-containing organic cations, so as to dry the alkali metal content to alkali metal oxides (e.g., Na ₂ O) on a dry basis. To less than 4% by weight, preferably less than 3% by weight, more preferably less than 2.5% by weight. As used herein, the weight of an anhydrous or dry basis zeolite means the weight after maintaining the zeolite at a temperature of 900 ° C. (1652 ° F.) for approximately 2 hours.

Optionally, the starting zeolite also contains a rare earth cation as RE ₂ O _{3 such} that the rare earth content constitutes 0.1 to 12.5% by weight, preferably 8.5 to 12% by weight of the zeolite (anhydrous basis) or at some stage of the modification procedure. It may be ion exchanged to contain cations. Those skilled in the art will appreciate that the ion exchange capacity of zeolites for introducing rare earth cations decreases during the disclosed treatment process. Thus, if rare earth cation exchange occurs, for example, as a final step in the manufacturing process, it may be impossible to introduce only the desired amount of rare earth cations. The skeletal Si / Al ₂ ratio of the starting Y zeolite may be in the range of 3 to 6, but is advantageously greater than 4.8.

The manner in which this first ammonium ion exchange is carried out is not an important factor and can be performed in a manner known to those skilled in the art. For example, such conventional ammonium ion exchange is performed at a pH of 4 or more. It is advantageous to use a three step procedure with 15% by weight aqueous ammonium nitrate solution at a rate such that the initial weight ratio of ammonium salt to zeolite in each step is equal to one. The contact time between the zeolite and the exchange medium is 1 hour in each stage and the temperature is 85 ° C. (185 ° F.). The zeolite is washed with 7.5 liters (2 gal) of water per 0.45 kg (1 lb) of zeolite between steps. The exchanged zeolite was then dried to 100 ° C. (212 ° F.) resulting in 20% by weight loss of ignition (LOI) at 1000 ° C. FIG. When using a rare earth cation, it is preferred to contact the already ammonium exchanged form of the zeolite with an aqueous solution of the rare earth salt in a known manner. The mixed rare earth chloride salt is added to an aqueous slurry of ammonium-exchanged Y zeolite (0.31 g of RECl ₃ per g of zeolite) in the temperature range of 85 to 95 ° C. so that the rare earth content is generally 8.5 to 12 as rare earth as RE ₂ O ₃ . Calculate the weight percent zeolite product.

After the ammonium ion exchange is complete, the steam treatment of the ammonium exchanged, optionally rare earth exchanged Y zeolite may be carried out at 550-850 ° C. (1022-98 ° C.) for a time sufficient to reduce the unit cell size to a range of 24.60 kPa, preferably 24.34-24.58 kPa. By contacting a steam environment containing at least 2 psia, preferably 100% steam, at a temperature of 1562 ° F.). Steam at a concentration of 100% can be used for 1 hour at temperatures in the range of 600-725 ° C. (1112-1337 ° F.). Note that this steaming step is not necessary for starting zeolites with Si / Al ₂ ratios of 6.5 or higher, as exemplified by the fluorosilicate treated material, with high Si / Al ₂ ratios being followed by acid extraction treatments and catalysts. Sufficient stability for the preparation and hydrocarbon conversion process.

Low pH, ammonium ion exchange is an important aspect of making modified Y zeolites consisting of the catalysts used in the methods disclosed herein. This exchange can be carried out in the same manner as in the case of the initial ammonium exchange, except that the pH of the exchange medium is reduced to at least 4, preferably at most 3, during at least part of the ion exchange procedure. The reduction in pH is readily accomplished by adding the appropriate inorganic or organic acid to the ammonium ion solution. Nitric acid is particularly suitable for this purpose. It is desirable to avoid acids that form insoluble ammonium salts. In performing low pH ammonium ion exchange, the pH of the exchange medium, the amount of exchange medium to the zeolite, and the contact time of the exchange medium with the zeolite are all important factors. As long as the exchange medium has a pH of 4 or less, it is confirmed that the sodium cation is exchanged with the hydrogen cation in the zeolite, and at least some ammonium is extracted mainly with non-skeleton ammonium and some skeletal ammonium. However, the effectiveness of the process is improved by acidifying the ion exchange medium with more acid than necessary to lower the pH to only 4 or less. As will be apparent from the data described below, the more acidic the exchange medium, the greater the tendency to extract not only skeletal aluminum but also non-skeletal aluminum from the zeolite. The extraction procedure is carried out to a degree sufficient to produce a zeolite product having a bulk Si / Al ₂ ratio of 6.5 to 35. In other embodiments, the bulk Si / Al ₂ ratio is 6.5-23, more preferably 6.5-20.

Typical Y zeolites having total silica to alumina Y modified Y zeolites used in the catalysts of the methods disclosed herein contain a Y zeolite represented by Y-85. US 5,013,699 and 5,207,892, which are incorporated herein by reference, describe Y-85 zeolites and their preparation and, therefore, need not be described in detail herein.

As illustrated in FIGS. 1-5 and the examples below, the disclosed catalysts provide for increased catalytic activity and low NPB formation in the case of cumene production. In the case of producing ethylbenzene from poly-ethylbenzene (FIG. 5), the diffusion properties of the disclosed catalysts appear to be important even if the internal isomerization of the ethyl groups is rarely concerned and the ethyl group is less than propyl.

Although the catalysts disclosed above may contain metal hydrogenation catalyst components, such components are not a requirement. Based on the weight of the catalyst, such metal hydrogenation catalyst components may be present in less than 0.2% by weight or less than 0.1% by weight, calculated as individual monoxide of the metal component, or the catalyst may lack any metal hydrogenation catalyst component. . When present, the metal hydrogenation catalyst component may be present in the final catalyst composition as a compound such as an oxide, sulfide, halide, or the like, or in the elemental metal state. As used herein, the term 'metal hydrogenation catalyst component' includes these various compound forms of the metal. The catalytically active metal is contained within the zeolite constituents on the outer surface of the zeolite crystals, i.e., contained in the internal adsorption zone, i. The metal can be imparted to the entire composition by any method to obtain a high dispersion state. Among the suitable methods are implantation, adsorption, ion exchange and intensive mixing. The metal may be copper, silver, gold, titanium, chromium, molybdenum, tungsten, rhenium, manganese, zinc, vanadium or any element of IUPAC Groups 8-10, in particular platinum, palladium, rhodium, cobalt and nickel. Mixtures of metals can be used.

The final catalyst composition may contain 10 to 95% by weight, preferably 15 to 50% by weight, of the general binder component. The binder may generally be an inorganic oxide or a mixture thereof. Both amorphous and crystalline can be applied. Examples of suitable binders are silica, alumina, silica-alumina, clay, zirconia, silica-zirconia and silica-boria. Alonia is the preferred binder material.

In the cumene preparation, the final catalyst prepared with 80% by weight zeolite and 20% by weight alumina binder on a volatile free basis has preferably one, more preferably both of the following properties: (1) X-ray diffraction The absolute strength of the modified Y zeolite measured by (XRD) is preferably at least 50, more preferably at least 60; And (2) the framework aluminum of the modified Y zeolite is preferably at least 60%, more preferably at least 70% of the aluminum of the modified Y zeolite. In one embodiment, the final catalyst for cumene production is a product of the absolute strength of the modified Y zeolite, measured by XRD, with the% skeletal aluminum in the aluminum of the modified Y zeolite, greater than 4200. In ethylbenzene production, the final catalyst preferably has one, more preferably both of the following characteristics: (1) The absolute strength of the modified Y zeolite as determined by X-ray diffraction (XRD) is preferably Is at least 65, more preferably at least 75; And (2) the framework aluminum of the modified Y zeolite is preferably at least 50%, more preferably at least 60% of the aluminum of the modified Y zeolite. In one embodiment, the final catalyst for the cumene product has a product of the absolute strength of the modified Y zeolite and the% skeletal aluminum in the aluminum of the modified Y zeolite, measured by XRD, greater than 4500.

In one embodiment, the methods disclosed herein use a substantially dry catalyst. After low pH ammonium ion exchange, a firing step is removed which removes substantially all of the water present. It has been found that the catalytic performance in the process described herein is improved by removing water. In order to maintain high activity and low NPB formation, it was found that the water content of the zeolite must be relatively low before it can be used in the transalkyl process.

Excess water can reduce the number of active sites and limit the diffusion to those sites, which does not effectively promote transalkylation. In order to solve this problem, dehydration of the catalyst particles to contain a predetermined amount of water can be carried out prior to initiation with a desiccant which can be introduced into the transalkyl reaction zone, wherein the temperature in the reaction zone is either aromatic or It may be increased slowly before introducing the transalkylatable aromatic. During this initial heating period, the water content of the zeolite is determined by the equilibrium between the amount of zeolite, catalyst, desiccant and water in the reaction zone, if any, at the temperature of the reaction zone. The zeolite portion of the catalyst is very hydrophilic and the hydration level is controlled by controlling the rate at which the desiccant crosses the catalyst and the temperature during the dehydration step. The desiccant can be any agent which removes water and does not adversely affect the catalyst, such as molecular nitrogen, air or benzene. The temperature during this dehydration step is maintained between 25-500 ° C. (77-932 ° F.). The water content of the catalyst is calculated by measurement of the loss on ignition (LOI), which is generally calculated after the weight of heating at 900 ° C. (1652 ° F.) for 2 hours, and then the weight loss due to the decomposition of ammonium ions into ammonia. Measure by measuring the amount. Catalysts containing water in excess of a predetermined amount, i.e., in an amount greater than the equilibrium amount of water that the catalyst will contain at any time during the start of the process, will lose water immediately after equilibrium during initiation, The dehydration step for providing the amount of water to the catalyst may, although desirable, not necessarily be carried out.

Any desirable properties of the catalyst, such as crush strength and ammonium ion concentration, are obtained by controlling the time and temperature conditions at which the extruded catalyst particles are calcined. In some cases, firing at high temperatures leaves the required amount of water in the catalyst, making it unnecessary to carry out the individual dehydration step. Thus, as used herein, 'dehydrating' and 'dehydrating' refer not only to the separation step in which the water is removed with the catalyst after the firing step but also encompasses the firing step carried out under conditions in which a certain amount of water remains in the catalyst particles. do.

The dehydration procedure described above is part of the actual process for producing the catalysts disclosed above in the production plant. However, procedures other than those described above can be used to dewater the catalyst at the time the catalyst is produced in the production plant, or at any other time in the production plant or elsewhere. For example, the extruded catalyst particles may contain a water deficiency containing gas such as dry molecular nitrogen or air, or a dry reactant such as a dry aromatic substrate (eg benzene) or a dry transalkylable aromatic (eg, in a transalkylation reactor). For example, DIPB or TIPB) may be dehydrated in the system at relatively high temperatures on the catalyst until the catalyst contains a predetermined amount of water. In the in-system dehydration step, the water deficient gas or reactant typically contains less than 30 ppm by weight of water and the contacting step is carried out at a temperature between 25 ° C. (77 ° F.) and 500 ° C. (932 ° F.). In one embodiment, the catalyst is contacted with gaseous flow dry nitrogen at 250 ° C. (482 ° F.). The catalyst is, for example, 130 ° C (266 ° F) to 260 ° C (500 ° F), 160 ° C (320 ° F) to 210 ° C (410 ° F), 180 ° C (356 ° F) to 200 ° C (392 ° F) or 150 Contact with liquid flow dry benzene from < RTI ID = 0.0 > The catalyst particles may also be stored in a manufacturing plant or elsewhere in contact with the surrounding gas until a predetermined amount of water is desorbed.

Typically, the LOI of the catalyst charged to the transalkylation reactor is in the range of 2-4% by weight. After charging into the reactor, and preferably prior to promoting the transalkylation reaction with the catalyst, the catalyst can be subjected to a dehydration step to reduce the water content in the catalyst. The nitrogen content of the catalyst is also preferably minimized.

The catalysts disclosed above are useful for transalkylating transalkylable aromatics. The transalkylation process disclosed herein preferably permits transalkylationable hydrocarbons as feed with the aromatic substrate. Transalkylable hydrocarbons useful in the transalkylation process consist of aromatic compounds characterized by constituting molecules based on the aromatic substrate with one or more alkylating agents instead of one or more hydrogen atoms around the aromatic substrate ring structure.

The chelating agent compound is a group of various substances including monoolefins, diolefins, polyolefins, acetylene hydrocarbons, and alkyl halides, alcohols, ether esters (the latter being various esters of alkylsulfates, alkylphosphates and carboxylic acids). Can be selected from. Preferred olefin functional compounds are olefinic hydrocarbons containing monoolefins containing one double bond per molecule. Monoolefins that can be used as the olefin functional compound in the processes disclosed above are generally gaseous or generally liquid, and include ethylene, propylene, 1-butene, 2-butene, isobutylene, high molecular weight general liquid olefins, Such as various pentenes, hexenes, heptenes, octenes and mixtures thereof, and high molecular weight liquid olefins, the latter having various olefin oligomers having 9 to 18 carbon atoms per molecule, such as propylene trimers, propylene tetramers, propylene Pentameric and the like. Although equilibrium results are not essential, C ₉ -C ₁₈ normal olefins can also be used as cyclo olefins such as cyclopentene, methylcyclopentene, cyclohexene, methylcyclohexene and the like can be used. It is preferable that the said monoolefin contains 2 or more and less than 14 carbon atoms. More particularly, the monoolefin is preferably propylene. The alkylating agent compound is preferably a C ₂ -C ₁₄ aliphatic hydrocarbon, more preferably propylene.

Aromatic substrates useful as part of the feed to the transalkylation process can be selected from the group of aromatic compounds comprising individually and in admixture of benzene and monocyclic alkylsubstituted benzenes having the structure:

In said formula, R is a hydrocarbon containing 1-14 carbon atoms, n is an integer of 1-5. On the other hand, the aromatic base portion of the feedstock may be benzene, benzene containing 1 to 5 alkyl group substituents and mixtures thereof. Non-limiting examples of the feedstock compound include benzene, toluene, xylene, ethylbenzene, mesitylene (1,3,5-trimethylbenzene), cumene, n-propylbenzene, butylbenzene, dodecylbenzene, tetradecylbenzene and their And mixtures. It is particularly preferable that the aromatic substrate is benzene.

The disclosed transalkylation process can have multiple processes. In one process, the catalyst of the transalkylation reaction zone is used to remove excess alkylating agent compound from the ring structure of the polyalkylated aromatic compound, and to transfer the alkylating agent compound to the previously unalkylated aromatic base molecule prepared by the process. Increase the amount of the desired aromatic compound. For related purposes, the reactions carried out in the transalkylation reaction zone include removing all alkylating agents from substituted aromatic compounds, and converting the aromatic substrate to benzene in the meantime.

The feed mixture is preferably 20 ppm by weight, more preferably 10 ppm by weight, based on the weight of the transalkylationable aromatic and aromatic substrates through which the concentrations of water and oxygen containing compounds in the blended feed pass into the reaction zone, Even more preferably 2 ppm by weight. How very low concentrations in the feed mixture are achieved is not critical to the method disclosed herein. Generally, one stream containing transalkylable aromatics and another stream containing aromatic substrates are provided, each stream having water and water so that the feed mixture formed by the combination of said individual streams has a predetermined concentration. It has a concentration of oxygen-containing compound precursor. Water and oxygen containing compounds can be removed from the individual streams or feed mixtures by conventional methods such as drying, adsorption or stripping. The oxygen containing compound may be any alcohol, aldehyde, epoxide, ketone, phenol or ether having a molecular weight or boiling point within the molecular weight or boiling range of the hydrocarbon in the feed mixture.

In order to transalkylate the polyalkylaromatic with the aromatic substrate, a feed mixture containing the aromatic substrate and the polyalkylated aromatic compound in a molar ratio of 1: 1 to 50: 1, preferably 1: 1 to 10: 1, was prepared. Continuous or intermittent injection into the transalkylation reaction zone containing the catalyst disclosed above at transalkylation conditions including temperatures of 390 [deg.] C. (140-734 [deg.] F.), in particular 70-200 [deg.] F. (158-392 [deg.] F.). A pressure suitable for use herein is preferably 1 atm (101.3 kPa (a)) but should not exceed 130 atm (13169 kPa (a)). A particularly preferred pressure range is 10-40 atm (1013-4052 kPa (a)). Preference is given to weight space time velocities (WHSV) of from 0.1 to 50 h ^-1 , in particular from 0.5 to 5 h ^-1 , based on the polyalkylaromatic feed rate and the dry weight of the catalyst. Although the process disclosed herein can be carried out in the vapor phase, it should be noted that the temperature and pressure combinations applied in the transalkylation reaction zone preferably ensure that the transalkylation reaction is carried out in the liquid phase. In the liquid phase transalkylation process to produce monoalkylaromatics, the catalyst is washed successively with the reactants to prevent coke precursor formation on the catalyst. This reduces the amount of carbon formation on the catalyst, in which case the catalyst cycle life is extended compared to the gas phase transalkylation process where coke formation and catalyst deactivation are major issues. In addition, the selectivity to monoalkylaromatic production, in particular cumene production, is higher in the catalytic liquid phase transalkylation reaction as compared to the catalytic gas phase transalkylation reaction.

Transalkylation conditions for the processes disclosed herein generally include a molar ratio of aromatic ring groups per alkyl group of from 1: 1 to 25: 1. It is contemplated that the molar ratio may be less than 1: 1 and the molar ratio may be 0.75: 1 or less. It is preferable that the molar ratio of the aromatic ring group per alkyl propyl group (propyl group in cumene production) is 6: 1 or less.

In the transalkylation conditions, the catalyst particles are typically measured in water by Karl Fischer titration and preferably at most 4% by weight, more preferably at most 3% by weight, even more preferably at most 2% by weight. It is contained in an amount, and nitrogen is contained in an amount of preferably 0.05% by weight or less as determined by micro (CHN) (carbon-hydrogen-nitrogen) analysis.

In this application, the initial initial labeling of all elements of the Periodic Table (CRC Handbook of Chemistry and Physics, ISBN 0-8493-0480-6, CRC Press, Boca Raton, Florida, USA, 80th Edition, 1999-2000) See IUPAC 'New Symbol Method' on the Periodic Table of Elements.

As used herein, the molar ratio of aromatic ring groups per alkyl group is defined as follows. The ratio of molecules is the number of molecules of the aromatic ring group passing through the reaction zone for a certain period of time. The number of molecules of the aromatic ring group is the sum of all aromatic ring groups, regardless of the compound from which the aromatic ring group is to be generated. For example, in cumene production, 1 mol of benzene, cumene, 1 mol, 1 mol of DIPB and 1 mol of TIPB each contribute 1 mol of the aromatic ring group to the sum of the aromatic ring groups. In ethylbenzene (EB) preparation, 1 mol of benzene, 1 mol of EB and 1 mol of di-ethylbenzene (DEB) each contribute 1 mol of the aromatic ring group to the sum of the aromatic ring groups. The denominator of this ratio is the number of moles of alkyl groups that have the same number of carbon atoms as the number of alkyl groups on a given monoalkylated aromatic and pass through the reaction zone for a certain period of time. The number of moles of alkyl groups is the sum of all alkyl and alkenyl groups having the same number of carbon atoms as the number of alkyl groups on a given monoalkylated aromatic, regardless of the compound from which the alkyl or alkenyl groups occur, except that no paraffins are included. Thus, the number of moles of propyl groups causes iso-propyl, n-propyl or propenyl groups to occur except that paraffins such as propane, n-butane, isobutane, pentane and higher paraffins are excluded from the calculation of the number of moles of propyl groups. It is the sum of all iso-propyl, n-propyl and propenyl groups, irrespective of the compound being. For example, 1 mol of propyl, 1 mol of cumene, and 1 mol of NPB each contribute 1 mol of the propyl group to the sum of the propyl groups, while correlating with the distribution of three groups between iso-propyl and n-propyl groups Without, 1 mol of DIPB contributes 2 mol of propyl group and 1 mol of tri-propylbenzene contributes 3 mol of propyl group. 1 mol of ethylene and 1 mol of EB each contribute 1 mol of ethyl group to the sum of the ethyl groups, while 1 mol of DEB contributes 2 mol of ethyl groups and 1 mol of tri-ethylbenzene contributes 3 mol of ethyl groups. Ethane does not contribute to the moles of ethyl groups.

As used herein, WHSV means weight spacetime velocity, which is defined as the hourly weight flow rate divided by the catalyst weight, where the weight flow rate and catalyst weight are in the same weight units.

As used herein, DIPB conversion is defined as the difference between the moles of DIPB in the feed and the moles of DIPB in the product divided by the moles of DIPB in the feed and multiplied by 100.

The surface area in this application is ASTM D4365-95 [Standard Test Method for Determining Micropore Volume and Zeolite Area of a Catalyst, and in the article by S. Brunauer et al., J. Am. Chem. Soc, 60 (2), 309-319 (1938)], using a Brunauer-Emmett-Teller (BET) model method using nitrogen adsorption techniques to obtain nitrogen partial pressure p / po data points ranging from 0.03 to 0.30. Calculated using.

As referenced herein, the absolute intensity by X-ray powder diffraction (XRD) of the Y zeolitic material calculates the standard sum of the intensities of several selected XRD peaks of the Y zeolitic material and adds the sum to the first standard and the US Department of Commerce. It was measured by dividing by the standard sum of several XRD peak intensities of the alpha-alumina NBS 674a intensity standard, certified by the National Institute of Standards and Technology (NIST), an agent of. The absolute strength of Y zeolite is the ratio of the sum multiplied by 100:

Absolute Intensity = (Standard Intensity of Y Zeolite Material Feet) / (Standard Intensity of Alpha-Alumina Standard Peak) x 100

The scan parameters of the Y zeolite material and alpha-alumina standard are shown in Table 1.

matter Y zeolite Alpha-alumina standard 2T range 4-56 24.6-26.6, 34.2-36.2, 42.4-44.4 Step time 1 second / step or more depending on zeolite content 1 second / step Step width 0.02 0.01 peak (511,333), (440), (533), (642), (751,555) + (660,822), (664) (012), (104), (113)

For the purposes of this disclosure, the absolute strength of Y zeolites mixed with a non-zeolitic binder to yield a mixture of Y zeolite Z parts by weight and non-zeolitic binder (100-Z) parts by weight on a dry basis is such that the mixture is derived from absolute strength. It can be calculated using the formula A = C · (100 / Z), where A is the absolute strength of the Y zeolite and C is the absolute strength of the mixture. For example, when Y zeolite is mixed with HNO ₃ peptized Pural SB alumina to yield a mixture of 80 parts by weight of zeolite and 20 parts by weight of Al ₂ O ₃ binder on a dry basis and the measured absolute strength of the mixture is 60, The absolute strength of Y zeolite is calculated to be (60). (100/80) or 75.

As used herein, the unit cell size, sometimes referred to as the lattice parameter, is the (642), (822), (555), (840) and (664) peaks of the sporesite and the silicon (111) peak for correction. The unit cell size calculated using the method of using the profile fitting to find the XRD peak position of.

As used herein, the bulk Si / Al ₂ molar ratio of zeolite is based on silica (Al ₂ O ₃₎ to alumina as measured based on the total or total amount of aluminum and silicon (skeleton and non-skeleton) present in the zeolite. Molar ratio of SiO ₂ ) to, and sometimes herein, the total molar ratio of silica to Alumina (SiO ₂ to Al ₂ O ₃ ). The bulk Si / Al ₂ molar ratio is generally obtained by conventional chemical analysis involving all forms of aluminum and silicon present.

As used herein, the aluminum fraction of zeolite, which is skeletal aluminum, can be found in bulk compositions and in paper [G.T. Kerr, A.W. Chester, and D.H. Olson, Acta. Phys. Chem., 1978, 24, 169 and the paper [G.T. Kerr, Zeolites, 1989, 9, 350], based on the Kerr-Dempsey equation for skeletal aluminum.

As used herein, dry matter means based on weight after drying time for 1 hour in flow air at a temperature of 900 ° C. (1652 ° F.).

The following examples are presented for illustrative purposes only and are not intended to limit the scope of the present disclosure.

Example 1-Control

A sample of Y-74 zeolite was slurried in 15 wt% NH ₄ NO ₃ aqueous solution and the solution temperature was raised to 75 ° C. (167 ° F.). Y-74 zeolites were stabilized sodium Y zeolites having a bulk Si / Al ₂ ratio of approximately 5.2, a unit cell size of approximately 24.53, and a sodium content of approximately 2.7 wt%, calculated as dry Na ₂ O. Y-74 zeolites have a bulk Si / Al ₂ ratio of approximately 4.9, a unit cell size of approximately 24.67, and a sodium content of dry Na ₂ O (column 4, line 47 to column 5, line 2 of US 5,324,877). Later ammonium exchange by steps (1) and (2) to remove approximately 75% of Na and steamed at approximately 600 ° C. (1112 ° F.) to dealuminate) from the sodium Y zeolite of approximately 9.4% by weight. Prepared. Y-74 zeolites were prepared and obtained from UOP LLC, Des Planes, Illinois, USA. After 1 hour of contact at 75 ° C. (167 ° F.), the slurry was filtered and the filter cake was washed with excess warm deionized water. The NH ₄ ⁺ ion exchange, filtration and water washing steps were repeated two or more times, and the resulting filter cake had a bulk Si / Al ₂ ratio of 5.2, a sodium content of 0.13% by weight, calculated as dry Na ₂ O, The unit cell size was 24.572 Hz and the absolute intensity measured by X-ray diffraction was 96. The resulting filter cake was dried to an appropriate wet level and mixed with HNO ₃ peptized Pural SB alumina to yield a mixture of 80 parts by weight zeolite and 20 parts by weight Al ₂ O ₃ binder, followed by 1.59 mm (1/16 in) ) Extruded into a cylindrical extrudate of diameter. The extrudate was dried and calcined for 1 hour in flowing air at approximately 600 ° C. (1112 ° F.). The catalyst represented the prior art. The catalyst had a unit cell size of 24.494 kPa, an XRD absolute strength of 61.1, and 57.2% as a percentage of aluminum of the Y zeolite with modified skeletal aluminum.

Example 2

Another sample of Y-74 zeolite used in Example 1 was slurried in 15 wt% NH ₄ NO ₃ aqueous solution. The pH of the slurry was lowered from 4 to 2 by addition of a sufficient amount of 17 wt% HNO ₃ solution. The slurry temperature was then heated to 75 ° C. (167 ° F.) and maintained for 1 hour. After 1 hour of contact at 75 ° C. (167 ° F.), the slurry was filtered and the filter cage was washed with excess warm deionized water. The acid extraction, filtration, and water washing steps in the presence of NH ₄ ⁺ ion exchange were repeated once, and the resulting filter cake had a bulk Si / Al ₂ ratio of 11.5 and a sodium content of 0.01 Na as measured by dry Na ₂ O. Less than% by weight and unit cell size was 24.47 mm 3. The resulting filter cake was dried to an appropriate wet level and mixed with HNO ₃ peptized Pural SB alumina to yield a mixture of 80 parts by weight zeolite and 20 parts by weight Al ₂ O ₃ binder, followed by 1.59 mm (1/16 in) ) Extruded into a cylindrical extrudate of diameter. The extrudate was dried and calcined for 1 hour in flowing air at approximately 600 ° C. (1112 ° F.). The properties of the catalyst were 68.2% by weight SiO ₂ in bulk and dry weight, 30.5% in weight in Al ₂ O ₃ dry weight, 0.04% in weight sodium Na ₂ O, 0.03% in dry weight (NH ₄ ) ₂ O 0.03 Weight%, unit cell size 24.456 mm 3, absolute XRD intensity 66.5, skeletal aluminum as a percentage of ammonium in modified Y zeolite, 92.2%, BET surface area 708 m ² / g.

Example 3

Another sample of Y-74 zeolite used in Example 1 was slurried in 15 wt% NH ₄ NO ₃ aqueous solution. A sufficient amount of 17 wt% HNO ₃ solution was added over a period of 30 minutes to remove excess skeletal aluminum. The slurry temperature was then heated to 175 ° F. and maintained for 90 minutes. After 90 minutes of contact at 79 ° C. (175 ° F.), the slurry was filtered and the filter cake was washed with 22% ammonium nitrate solution and then washed with excess warm deionized water. Unlike Example 2, the acid extraction was not repeated twice in the presence of ammonium nitrate. The resulting filter cake had a bulk Si / Al ₂ ratio of 8.52 and a sodium content of 0.18% by weight as measured as a dry Na ₂ O. The resulting filter cake was dried by the method described in Example 2, mixed with HNO ₃ peptized Pural SB alumina, extruded, dried and calcined. The catalyst had a unit cell size of 24.486 kPa, an absolute XRD intensity of 65.8, a percentage of ammonium in the modified Y zeolite as 81.1% skeletal aluminum, and a BET surface area of 698 m ² / g.

Example 4

Example 4 followed the same procedure as described in Example 3 except that an increase of 33% HNO ₃ was used as compared to Example 3. The same stable Y-74 used in Example 1 was slurried in 15 wt% NH ₄ NO ₃ aqueous solution. A sufficient amount of 17 wt% HNO ₃ was added over a period of 30 minutes to remove excess skeletal aluminum. The slurry temperature was then heated to 175 ° F. and maintained for 90 minutes. After 90 minutes of contact at 79 ° C. (175 ° F.), the slurry was filtered and the filter cake was washed with excess warm deionized water. Unlike Example 2, the NH ₄ ⁺ ion exchange, filtration and water washing steps were not repeated. The resulting filter cake had a bulk Si / Al ₂ ratio of 10.10 and a sodium content of 0.16% by weight as measured as a dry Na ₂ O. The resulting filter cake was dried by the method described in Example 2, mixed with HNO ₃ peptized Pural SB alumina, extruded, dried and calcined. The catalyst had a unit cell size of 24.434 mm 3, absolute XRD strength of 53.6, percentage of ammonium in the modified Y zeolite as 74.9% skeletal aluminum, and a BET surface area of 732 m ² / g.

Example 5-Control

Example 5 followed the same procedure as described in Example 3 except that an increase of 52% HNO ₃ was used as compared to Example 3. The same stable Y-74 used in Example 1 was slurried in 15 wt% NH ₄ NO ₃ aqueous solution. A sufficient amount of 17 wt% HNO ₃ was added over a 30 minute period of time to increase the bulk Si / Al ₂ ratio. The slurry temperature was then heated to 175 ° F. and maintained for 90 minutes. After 90 minutes of contact at 79 ° C. (175 ° F.), the slurry was filtered and the filter cake was washed with excess warm deionized water. Unlike Example 2, the NH ₄ ⁺ ion exchange, filtration and water washing steps were not repeated. The resulting filter cake had a bulk Si / Al ₂ ratio of 11.15 and a sodium content of 0.08 wt%, measured as a dry Na ₂ O. The resulting filter cake was dried to an appropriate wet level and mixed with HNO ₃ peptized Pural SB alumina to yield a dry mixture of 80 parts by weight of zeolite and 20 parts by weight of Al ₂ O ₃ binder, followed by 1.59 mm (1 / Extruded into a 16 in) diameter cylindrical extrudate. The extrudate was dried and calcined for 1 hour in flowing air at approximately 600 ° C. (1112 ° F.). The catalyst had a unit cell size of 24.418 kPa, an absolute XRD intensity of 44.8, 75.2% skeletal aluminum as a percentage of ammonium in the modified Y zeolite, and a BET surface area of 756 m ² / g.

Example 6

The same stabilized Y-74 used in Example 1 was slurried in a 15 wt% NH ₄ NO ₃ aqueous solution. The total amount of HNO ₃ used in this example is the same as that in Example 5. However, instead of performing acid extraction in a single step as described in Example 5, the acid extraction was used with 85% of the total HNO ₃ acid in the first step and the remaining 15% of the total acid in the second step. It was carried out in two steps to use. The acid extraction procedure / conditions in each of the two separate steps were the same as described in Example 5. A solution of 17 wt% HNO ₃ was added to the slurry consisting of Y-74 and NH ₄ NO ₃ solution. The slurry temperature was then heated to 175 ° F. and maintained for 90 minutes. After 90 minutes of contact at 79 ° C. (175 ° F.), the slurry was filtered and the filter cake was washed with excess warm deionized water. Acid extraction in the presence of NH ₄ ⁺ (to the remaining 15% of the total HNO ₃ used), filtration and water washing steps were repeated, and the resulting filter cake had a bulk Si / Al ₂ ratio of 11.14 and a sodium content of dry measured as Na ₂ O was 0.09% by weight. The resulting filter cake was dried to an appropriate wet level and mixed with HNO ₃ peptized Pural SB alumina to yield a mixture of 80 parts by weight zeolite and 20 parts by weight Al ₂ O ₃ binder, followed by 1.59 mm (1/16 in) ) Extruded into a cylindrical extrudate of diameter. The extrudate was dried and calcined for 1 hour in flowing air at approximately 600 ° C. (1112 ° F.). The catalyst had a unit cell size of 24.411 kPa, an absolute XRD intensity of 56.1, a percentage of ammonium in the modified Y zeolite as 72.5% skeletal aluminum, and a BET surface area of 763 m ² / g.

Example 7

The same stabilized Y-74 used in Example 3 was slurried in 18 wt% ammonium nitrate solution. To this solution, 17% sulfuric acid solution was added over 30 minutes. The batch was then heated to 175 ° F (79 ° C) and held for 90 minutes. Heat was removed and the batch was then quenched with treated water to reduce the temperature to 62 ° C. (143 ° F.) and filter. The Y zeolitic material was then reslurried in a 6.4 wt% ammonium sulfate solution and held at 79 ° C. (175 ° F.) for 1 hour. The material was then filtered and water washed. The resulting filter cake had a bulk Si / Al ₂ ratio of 7.71 and a sodium content of 0.16% by weight as measured as a dry Na ₂ O. The filter cake produced in the manner described in Example 2 was dried, mixed with HNO ₃ peptized Pural SB alumina, extruded, dried and calcined. The catalyst had a unit cell size of 24.489 mm 3, absolute XRD strength of 65.3, and 75.7% skeletal aluminum as a percentage of ammonium in the modified Y zeolite.

Table 2 summarizes the properties of the catalysts prepared in Examples 1-7.

Example One 2 3 4 5 6 7 Example Type Comparative example Example Example Example Comparative example Example Example Drawing by run data 1-5 1-2, 5 1-2 1-2 1-2 none 1-4 Y zeolite bulk Si / Al ₂ ratio, mol 5.20 11.50 8.52 10.10 11.15 11.14 7.71 Y zeolite unit cell size, Å 24.494 24.456 24.486 24.434 24.418 24.411 24.489 Catalytic XRD Absolute Strength 61.1 66.5 65.8 53.6 44.8 56.1 65.3 Y zeolite XRD absolute strength 76.4 83.1 82.3 67 56 70.1 81.6 Y zeolite skeletal aluminum, atomic% of total aluminum 57.2 92.2 81.1 74.9 75.2 72.5 75.7 Catalytic BET Surface Area, m ² / g - 708 698 732 756 763 -

Example 8

The catalysts prepared in Examples 1-5 and 7 were tested for transalkylation performance with feeds containing benzene and polyalkylated benzenes. The feed was prepared in combination with polyalkylated benzenes obtained from commercial transalkylation units with benzene. The feed blend prepared above means a typical transalkylated feed composition having a molar ratio of aromatic ring groups to propyl groups of approximately 2.3. Catalysts prepared by the processes disclosed herein have been found to provide the same advantages when processing feeds at substantially low or high feed molar ratios. The feed composition measured by gas chromatography is summarized in Table 3. The test was conducted once under conditions of DIPB WHSV over a reactor pressure of 3447 kPa (g) (500 psi (g)), a molar ratio of aromatic ring groups to a propyl group of 2.3 and a reaction temperature range of 0.8 h- ¹ . through a fixed bed reactor in mode. The reactor was allowed to achieve a substantially steady state at each reaction temperature and the product was sampled for analysis. Substantially no catalyst deactivation occurred during the test. Prior to feeding the feed, each catalyst was contacted with a flowing nitrogen stream containing less than 10 ppm by weight of water at 250 ° C. (482 ° F.) for 6 hours to undergo a drying procedure.

ingredient Concentration, wt% benzene 63.832 Non-aromatic 0.038 toluene 0.002 Ethylbenzene 0.000 Cumen 0.880 NPB 0.002 Butylbenzene 0.071 Pentylbenzene 0.021 m-DIPB 20.776 o-DIPB 0.520 p-DIPB 13.472 Hexylbenzene 0.308 1,3,5-TIPB 0.029 1,2,4-TIPB 0.012 Tetra-isopropylbenzene 0.003 Nonylbenzene 0.004 Unknown 0.030 Sum 100.000

This example demonstrates the advantages of high activity and product purity in the transalkylation of poly-alkylates to cumene, which contributes to the catalysts produced by the methods disclosed herein.

Example 9-Regeneration

Samples of the catalyst prepared in Example 7 were tested in the manner described in Example 8 as described above. After the test, the catalyst used was placed in a ceramic dish and placed in a muffle furnace. While passing the flow air through the muffle, the furnace temperature is raised from 70 ° C. (158 ° F.) to 550 ° C. (1022 ° F.) at a rate of 1 ° C. (1.8 ° F.) per minute, and at 6 ° C. (1022 ° F.). After holding for hours, it was cooled to 110 ° C. (230 ° F.). After regeneration, the catalyst was again tested in the manner described in Example 8.

3 and 4 show test results before play (indicated by 'Example 7') and before play (indicated by 'Example 9'). The results indicate that the catalyst before and after regeneration was similar to the activity and product purity better than the curve for the Example 1 catalyst, thus showing good catalyst reproducibility.

Example 10

Samples of the catalysts prepared in Examples 1 and 2 were evaluated for transalkylation of poly-ethylbenzene. Each catalyst was tested using a feed consisting of a combination of 63.6 wt% benzene and 36.4 wt% para-diethylbenzene (p-DEB). After introducing the catalyst into the reactor, the catalyst was dried by contacting a flowing nitrogen stream containing less than 10% by weight of water at 250 ° C. (482 ° F.) for 6 hours. Each of the catalysts was carried out to 2:00 over the reaction temperature range of ^-1 of p-DEB (446 ℉) 170 ℃ (338 ℉) ~230 ℃ in WHSV. The reactor was allowed to achieve a substantially steady state at each reaction temperature and the product was sampled for analysis. Substantially no catalyst deactivation occurred during the test. 5 shows the results for both catalysts. The results indicate that the catalyst prepared in Example 2 possesses similar or superior activity and stability to the curves for the catalyst prepared in Example 1 and can be used in commercial poly-ethylbenzene transalkyl operation.

A summary of this data is provided in Figures 1-5. In FIG. 1, the DIPB conversions for Examples 2-4 and 7 are substantially higher than for Examples 1 and 5, where Example 1 is indicated by line 101. In Figure 2, the NPB / cumene ratio is lower for Examples 2-4 and 7 compared to Example 1 (indicated by line 201). In FIG. 3, the DIPB conversion is higher for the non-regenerated catalyst of Example 7 and the regenerated catalyst of Example 9 compared to Example 1 (indicated by line 101 of FIG. 1). In FIG. 4, the NPB / cumene ratio is lower for the non-regenerated and regenerated catalysts of Examples 7 and 9, respectively, compared to Example 1 (indicated by line 201 in FIG. 2). In FIG. 5, Example 2 shows a better DEB conversion than Example 1 (indicated by line 501). The low activity and inferior product purity for the catalyst prepared in Comparative Example 5 is believed to be due to too much acid extraction conditions. Thus, excessive acid extraction conditions can reduce the crystallinity of the Y zeolite.

LZ-210

Y zeolites can be used in the processes disclosed herein and can be prepared by dealuminating Y zeolites having a total molar ratio of silica to alumina of 5 or less, and are described herein by US 4,503,023, 4,597,956, 4,735,928 And 5,275,720. The '023 patent discloses another procedure for dealumination of Y zeolites comprising contacting the Y zeolite with an aqueous solution of fluorosilicate salts using controlled rates, temperature and pH conditions to avoid aluminum extraction without silicon substitution. do. In the '023 patent, fluorosilicate salts are used as aluminum extracts and also as exotic silicon feeds that are inserted into the Y zeolite structure instead of the extracted aluminum. The salt has the formula:

(A) _{2 / b} SiF ₆

In the above formula, A is a metallic or nonmetallic cation other than H ⁺ having a 'b' value. The cations represented by 'A' include alkylammonium, NH 4 ⁺ , Mg ⁺⁺ , Li ⁺ , Na ⁺ , K ⁺ , Ba ⁺⁺ , Cd ⁺⁺ , Cu ⁺⁺ , H ⁺ , Ca ⁺⁺ , Cs ⁺ , Fe ⁺⁺ , Co ⁺⁺ , Pb ⁺⁺ , Mn ⁺⁺ , Rb ⁺ , Ag ⁺ , Sr ⁺⁺ , Ti ⁺ and Zn ⁺⁺ .

A preferred member of this group of Y zeolites is known as LZ-210, a zeolite-based aluminosilicate molecular sieve described in the '023 patent. LZ-210 zeolites and other zeolites in this group are readily prepared from Y zeolite starting materials. In one embodiment, the total molar ratio of silica to LZ-210 zeolite alumina is 5.0-11.0. The unit cell size is in the range of 24.38 to 24.50 ms, preferably 24.40 to 24.44 ms. The LZ-210 class of zeolites used in this process and the compositions disclosed herein have compositions represented in terms of molar ratios of oxides, as in the formula:

(0.85-1.1) M _{2 / n} O: Al ₂ O ₃ : XSiO ₂

In the above formula, 'M' is a cation having an 'n' value, and 'x' has a value of 5.0 to 11.0.

In general, LZ-210 zeolites can be prepared by dealuminating Y type zeolites using an aqueous solution of fluorosilicate salts, preferably a solution of ammonium hexafluorosilicate. The dealumination can be accomplished by exchanging Y zeolites, generally non-essentially ammonium exchanged Y zeolites, with a solution of an aqueous reaction medium, such as ammonium acetate, and slowly adding an aqueous solution of ammonium fluorosilicate. After allowing the reaction to proceed, a zeolite is produced with an increased total molar ratio of silica to alumina. The extent of the increase depends at least in part on the amount of fluorosilicate solution in contact with the zeolite and on the reaction time allowed. In general, a reaction time of 10 to 24 hours is sufficient to achieve equilibrium. The resulting solid product, which can be separated from the aqueous reaction medium by conventional filtration techniques, is in the form of LZ-210 zeolite. In some cases, the product may be steam calcined by methods known in the art. For example, the product may be brought into contact with water at a partial pressure of at least 1.4 kPa (a) (0.2 psi (a)) for 1/4 to 3 hours at temperatures between 482 ° C. (900 ° F.) and 816 ° C. (1500 ° F.). It can provide high crystalline stability. In some cases, the product of steam firing may be ammonium exchange treated by methods known in the art. For example, the product may be slurried with water after adding ammonium salts to the slurry. The resulting mixture is typically heated for a period of time, filtered and washed with water. Steam treatment and ammonium exchange methods for LZ-210 zeolites are described in US Pat. Nos. 4,503,023, 4,735,928 and 5,275,720.

In one embodiment, treatment with an aqueous solution of the fluorosilicate salt after ammonium exchange increases the Si / Al ₂ ratio, increases hydrothermal stability, and lowers the tendency to form excess skeletal aluminum.

The final low pH ammonium ion exchange of the preferred LZ-210 zeolite is Y zeolite (and / or the LZ discussed above) except lowering the pH of the exchange medium to 4 or less, preferably 3 or less during at least part of the ion exchange procedure. -210 zeolite) in the same manner as in the case of the initial ammonium exchange. Reduction of pH is easily carried out by adding an appropriate inorganic acid or organic acid to the ammonium ion solution. Nitric acid is particularly suitable for this purpose. It is preferable to avoid acids forming insoluble ammonium salts. In performing low pH ammonium ion exchange, the pH of both the exchange medium, the amount of exchange medium to the zeolite and the contact time of the exchange medium and the zeolite are important factors. As long as the exchange medium has a pH of 4 or less, the sodium cation is exchanged with hydrogen cations in the zeolite, and at least some ammonium, mainly non-skeleton ammonium and some skeletal ammonium, are extracted. However, the effect of the process is enhanced by acidifying the ion exchange medium with more acid than necessary to lower the pH to only 4 or less. As will be apparent from the data mentioned below, the more acidic the exchange medium, the greater the tendency to extract not only the skeleton but also the non-skeletal aluminum from the zeolite. The extraction procedure is carried out to a degree sufficient to produce a zeolite product having a bulk Si / Al ₂ molar ratio of 6.5 to 27. In another embodiment, the bulk Si / Al ₂ molar ratio is in the range of 6.5-23, or more preferably 6.5-20.

The following LZ-210 examples are for illustrative purposes only and are not intended to limit the scope of the present disclosure.

Example 11-Control

A sample of Y-74 zeolite was slurried in 15 wt% NH ₄ NO ₃ aqueous solution and the solution temperature was raised to 75 ° C. (167 ° F.). Y-74 zeolites were stabilized sodium Y zeolites having a bulk Si / Al ₂ ratio of approximately 5.2, a unit cell size of approximately 24.53, and a sodium content of approximately 2.7 wt%, calculated as dry Na ₂ O. Y-74 zeolites have a bulk Si / Al ₂ ratio of approximately 4.9, a unit cell size of approximately 24.67, and a sodium content of dry Na ₂ O (column 4, line 47 to column 5, line 2 of US 5,324,877). Ammonium exchanged by steps (1) and (2) to remove approximately 75% of Na and steamed at approximately 600 ° C. (1112 ° F.) to dealuminate) to prepare from sodium Y zeolite of approximately 9.4% by weight. It was. Y-74 zeolites were prepared and obtained from UOP LLC, Des Planes, Illinois, USA. After 1 hour of contact at 75 ° C. (167 ° F.), the slurry was filtered and the filter cake was washed with excess warm deionized water. The NH ₄ ⁺ ion exchange, filtration and water washing steps were repeated two or more times, and the resulting filter cake had a bulk Si / Al ₂ ratio of 5.2, a sodium content of 0.13% by weight, calculated as dry Na ₂ O, The unit cell size was 24.572 Hz and the absolute intensity measured by X-ray diffraction was 96. The resulting filter cake was dried to an appropriate wet level and mixed with HNO ₃ peptized Pural SB alumina to yield a mixture of 80 parts by weight zeolite and 20 parts by weight Al ₂ O ₃ binder, followed by 1.59 mm (1/16 in) ) Extruded into a cylindrical extrudate of diameter. The extrudate was dried and calcined for 1 hour in flowing air at approximately 600 ° C. (1112 ° F.). The catalyst represented the prior art. The catalyst had a unit cell size of 24.494 kPa, an XRD absolute strength of 61.1, and 57.2% as a percentage of aluminum of the Y zeolite with modified skeletal aluminum.

Example 12

According to the procedure described in US 4,503,023, the synthesized Y-54 zeolite was treated with ammonium exchange followed by ammonium fluorosilicate. Y-54 zeolites are sodium Y zeolites having a bulk Si / Al ₂ ratio of approximately 4.9, a unit cell size of 24.67, and a sodium content of 9.4% by weight, calculated as dry Na ₂ O. Y-54 zeolite was prepared and obtained from UOP LLC, Des Planis, Illinois, USA. The resulting Y zeolite with a bulk Si / Al ₂ molar ratio of 6.5 was steamed at 100 ° C. for 1 hour at 600 ° C. (1112 ° F.) and then ammonium exchanged. The resulting filter cake was dried to an appropriate wet level and mixed with HNO ₃ peptized Pural SB alumina to yield a mixture of 80 parts by weight zeolite and 20 parts by weight Al ₂ O ₃ binder, followed by 1.59 mm (1/16 in) ) Extruded into a cylindrical extrudate of diameter. The extrudate was dried and calcined for 1 hour in flowing air at approximately 600 ° C. (1112 ° F.). The resulting catalyst had a unit cell size of 24.426 kPa, an absolute XRD strength of 81.6, and 63.2% of skeletal aluminum as a percentage of aluminum in the modified Y zeolite.

Example 13

According to the procedure described in US Pat. No. 4,503,023, the synthesized Y-54 zeolite was ammonium exchanged and then treated with ammonium fluorosilicate. The resulting Y zeolite, referred to as LZ-210 (9) with a bulk Si / Al ₂ ratio of 9.0, was steamed at 100 ° C. for 1 hour at 600 ° C. (1112 ° F.). A slurry consisting of 228 g of steamed LZ-210 (9) and 672 g of H ₂ O was first prepared. A NH ₄ NO ₃ solution consisting of 212 g of H ₂ O and 667 g of 50 wt% (NH ₄ ) NO ₃ was then added to the steamed LZ-210 (9) slurry. The resulting mixture was then raised to 85 ° C. (185 ° F.) and mixed for 15 minutes. To the mixture was added 5.7 g of 66 wt% HNO ₃ and the resulting mixture was maintained at 85 ° C. (185 ° F.) with continuous stirring for 60 minutes. At the end of the acid extraction, the mixture was filtered and the cake was washed with 1000 ml of H ₂ O and then dried at 100 ° C. (212 ° F.) overnight. In the second part, 200 g of dry cake was added to a solution consisting of 667 g of 50% by weight (NH ₄ ) NO ₃ and 650 g of H ₂ O, to which 20 g of 66% by weight HNO ₃ was added. The resulting slurry was mixed for 60 minutes. The mixture was then filtered, washed with 1000 ml of H ₂ O and the filter cake dried at 100 ° C. (212 ° F.) overnight. The resulting zeolite had a bulk Si / Al ₂ ratio of 10.82 and Na ₂ O of 0.026 wt%. The zeolite powder was mixed with HNO ₃ peptized Pural SB alumina to yield a mixture of 80 parts by weight of zeolite and 20 parts by weight of Al ₂ O ₃ binder in dry weight, and the moisture was adjusted to yield the appropriate dough fabric, followed by 1.59 mm (1 / Extruded into a 16 in) diameter cylindrical extrudate. The extrudate was dried and calcined for 1 hour in flowing air at approximately 600 ° C. (1112 ° F.). The resulting catalyst had a unit cell size of 24.430 mm 3, absolute XRD strength of 78.4, skeletal aluminum of 77.8%, and a BET surface area of 661 m ² / g.

Example 14

According to the procedure described in US Pat. No. 4,503,023, the synthesized Y-54 zeolite was ammonium exchanged and then treated with ammonium fluorosilicate. The resulting Y zeolite, referred to as LZ-210 (9) with a bulk Si / Al ₂ ratio of 9.0, was steamed at 100 ° C. for 1 hour at 600 ° C. (1112 ° F.). An amount of 256 g of steamed LZ-210 (9) was added to 1140 g of 22% by weight NH ₄ NO ₃ . To the zeolite slurry, 368 g of 17 wt% HNO ₃ was slowly added over 30 minutes. The slurry was then heated to 80 ° C. (176 ° F.) and held at 80 ° C. (176 ° F.) for 90 minutes. At the end of the acid extraction, the slurry was quenched with 1246 g of H ₂ O, filtered, washed with 1140 g of 22 wt% NH ₄ NO ₃ , washed with 1000 ml of H ₂ O, at 100 ° C. (212 ° F.) Oven dried overnight. The resulting zeolite had a bulk Si / Al ₂ ratio of 14.38 and Na ₂ O of 0.047% by weight. The resulting zeolite powder was mixed with HNO ₃ peptized Pural SB alumina to yield a mixture of 80 parts by weight of zeolite and 20 parts by weight of Al ₂ O ₃ binder in a dry amount, and adjusted moisture to yield a suitable dough fabric, followed by 1.59 mm ( Extruded into a 1/16 in) diameter cylindrical extrudate. The extrudate was dried and calcined for 1 hour in flowing air at approximately 600 ° C. (1112 ° F.). The resulting catalyst had a unit cell size of 24.393 kPa, absolute XRD strength of 79.6, skeletal aluminum of 81.8%, and a BET surface area of 749 m ² / g.

Example 15

According to the procedure described in US Pat. No. 4,503,023, the synthesized Y-54 zeolite was ammonium exchanged and then treated with ammonium fluorosilicate. The resulting Y zeolite, with a bulk Si / Al ₂ ratio of 12 and referred to as LZ-210 (12), was steamed at 100 ° C. for 1 hour at 600 ° C. (1112 ° F.). A slurry consisting of 231 g of steamed LZ-210 (12) and 668 g of H ₂ O was first prepared. A NH ₄ NO ₃ solution consisting of 212 g H ₂ O and 667 g of 50 wt% (NH ₄ ) NO ₃ was then added to the steamed LZ-210 (12) slurry. The resulting mixture was then raised to 85 ° C. (185 ° F.) and mixed for 15 minutes. 33.4 g of 66 wt% HNO ₃ was added to the mixture and the resulting mixture was maintained at 85 ° C. (185 ° F.) with continuous stirring for 60 minutes. At the end of the acid extraction, the mixture was filtered and the cake was washed with 1000 ml of H ₂ O and then dried at 100 ° C. (212 ° F.) overnight. In the second part, 200 g of dry cake was added to a solution consisting of 667 g of 50% by weight (NH ₄ ) NO ₃ and 650 g of H ₂ O, to which 20 g of 66% by weight HNO ₃ was added. The resulting slurry was stirred for 60 minutes. The mixture was then filtered, washed with 1000 ml of H ₂ O and the filter cake dried at 100 ° C. (212 ° F.) overnight. The resulting zeolite had a bulk Si / Al ₂ ratio of 17.24 and Na ₂ O of 0.01 wt%. The resulting zeolite powder was mixed with HNO ₃ peptized Pural SB alumina to yield a mixture of 80 parts by weight of zeolite and 20 parts by weight of Al ₂ O ₃ binder in a dry amount, and adjusted moisture to yield a suitable dough fabric, followed by 1.59 mm ( Extruded into a 1/16 in) diameter cylindrical extrudate. The extrudate was dried and calcined for 1 hour in flowing air at approximately 600 ° C. (1112 ° F.). The resulting catalyst had a unit cell size of 24.391 kPa, absolute XRD strength of 81.2, skeletal aluminum of 94.9%, and a BET surface area of 677 m ² / g.

Example 16

250 g of LZ-210 (120) of Example 15 (prior to steam treatment) were added to a NH ₄ NO ₃ solution consisting of 500 g of 50% NH ₄ NO ₃ and 625 g of H ₂ O. The slurry was heated to 95 ° C. (203 ° F.) and kept at temperature for 2 hours. The slurry was then filtered and water washed. The cake was then exchanged with NH ₄ NO ₃ following the same procedure and washed twice with water. The filter cake was oven dried overnight at 100 ° C. (212 ° F.). The resulting zeolite had a bulk Si / Al ₂ ratio of 12.62 and Na ₂ O of 0.05% by weight. The dried zeolite was mixed with HNO ₃ peptized Pural SB alumina to yield a mixture of 80 parts by weight of zeolite and 20 parts by weight of Al ₂ O ₃ binder in dry weight, and adjusted moisture to yield an appropriate dough fabric, followed by 1.59 mm (1 / 16 in) diameter cylindrical extrudate. The extrudate was dried and calcined for 1 hour in flowing air at approximately 600 ° C. (1112 ° F.). The resulting catalyst had a unit cell size of 24.431 mm 3, absolute XRD strength of 77.3, skeletal aluminum of 89.2%, and a BET surface area of 660 m ² / g.

Table 4 summarizes the properties of the catalysts prepared in Examples 11-16.

Example 11 12 13 14 15 16 18 Example Type Comparative example Example Example Example Example Example Example Drawing by run data 6-9 none none 6-9 6-7 6-7 6-7 Y zeolite bulk Si / Al ₂ ratio, mol 5.20 8.61 10.82 14.38 17.24 12.62 12.62 Y zeolite unit cell size, Å 24.494 24.426 24.430 24.393 24.391 24.431 24.439 Catalytic XRD Absolute Strength 61.1 81.6 78.4 79.6 81.2 77.3 72.5 Y zeolite XRD absolute strength 76.4 102 98 99.5 101.5 96.6 90.6 Y zeolite skeletal aluminum, atomic% of total aluminum 57.2 63.2 77.8 81.8 94.9 89.2 92.6 Catalytic BET Surface Area, m ² / g - - 661 749 677 660 660

Example 17

The catalysts prepared in Examples 11 and 14-16 were tested for transalkylation performance with feeds containing benzene and polyalkylated benzenes. The feed was prepared in combination with polyalkylated benzenes obtained from commercial transalkylation units with benzene. The feed composition measured by gas chromatography is summarized in Table 2 above. The test was a perfusion fixed bed reactor under conditions of DIPB WHSV over a reactor pressure of 3447 kPa (g) (500 psi (g)), a molar ratio of aromatic ring groups to a propyl group of 2.3, and a reaction temperature range of 0.8 hour ^-1 . It was carried out in. The reactor was allowed to achieve a substantially steady state at each reaction temperature and the product was sampled for analysis. Substantially no catalyst deactivation occurred during the test. Prior to feeding the feed, each catalyst was contacted with a flowing nitrogen stream containing less than 10 ppm by weight of water at 250 ° C. (482 ° F.) for 6 hours to undergo a drying procedure.

6 and 7 show the test results for the catalysts prepared in Examples 11 and 14-16. In Figure 6, the catalysts prepared in Examples 14-16 show higher activity (i.e. higher DIPB conversion at certain temperatures) compared to the curve 601 for Example 11. In Figure 7, the catalysts prepared in Examples 14-16 also exhibit better product purity (i.e., lower NPB / cumene at a given DIPB conversion) than the curve 701 for the catalyst prepared in Example 1. In Figures 6 and 7, the data for Example 16 indicate that the steam treatment and acid extraction steps are not needed in catalyst preparation because good performance can be achieved even if both are omitted. In addition, in FIGS. 6 and 7, the data for Example 14 show superior activity and compatibility with acid extraction after one-step steam treatment instead of the two-step acid extraction of Example 15 even if the acid extraction conditions are more excessive. Indicates that purity can be achieved.

Example 18

Samples of the catalyst prepared in Example 16 were tested in the manner described in Example 17 as described above. After the test, the catalyst used was placed in a ceramic dish and placed in a muffle. While passing the flow air through the muffle, the furnace temperature is raised from 70 ° C. (158 ° F.) to 550 ° C. (1022 ° F.) at a rate of 1 ° C. (1.8 ° F.) per minute, and at 6 ° C. (1022 ° F.). After holding for hours, it was cooled to 110 ° C. (230 ° F.). The regenerated catalyst has a unit cell size of 24.439 kPa, absolute XRD strength of 72.5, skeletal aluminum of 92.6%, and a BET surface area of 660 m ² / g. Table 4 summarizes the properties of the regenerated catalyst. After regeneration, the catalyst was tested again in the manner described in Example 17. The catalyst before and after regeneration had the same activity (ie DIPB conversion at a given temperature) and product purity (ie NPB / cumene at a given DIPB conversion), thus exhibiting good catalyst regeneration.

Example 19

Samples of the catalyst prepared in Example 14 were tested in the manner described in Example 17 as described above. After the test, the catalyst used was regenerated in the manner described in Example 18. After regeneration, the catalyst was tested again in the manner described in Example 17.

8 and 9 graphically show test results for catalysts prior to regeneration (labeled 'Example 14') and after regeneration (labeled 'Example 19'). The results indicate that the catalyst before and after regeneration has similar activity (ie, DIPB conversion at a given temperature) and product purity (ie, NPB / cumene at a given DIPB conversion), both of which are similar to those of the Example 11 catalyst. Curves 601 and 701 of 8 and 9, respectively, and thus excellent catalyst regeneration ability.

This example shows the advantage of high activity and product purity in the transalkylation of cumene and poly-alkylates such as DIPB and TIPB to cumene and DEB to EB contributing to the catalysts produced by the processes disclosed herein.

The disclosed catalysts may contain metal hydrogenation catalyst components, but such components are not a requirement. Based on the weight of the catalyst, the metal hydrogenation catalyst component may be present in less than 0.2% or less than 0.1% by weight, calculated as individual monoxide of the metal component, or the catalyst may be free of any metal hydrogenation catalyst component. When present, the metal hydrogenation catalyst component may be present in the final catalyst composite as a compound such as oxides, sulfides, halides, or in the elemental metal state. As used herein, the term 'metal hydrogenation catalyst component' includes these various compound forms of the metals. The catalytically active metal is contained within the zeolite constituents on the outer surface of the zeolite crystals, i.e., contained in the internal adsorption zone, i. The metal can be imparted to the entire composition by any method to obtain a high dispersion state. Among the suitable methods are implantation, adsorption, ion exchange and intensive mixing. The metal may be copper, silver, gold, titanium, chromium, molybdenum, tungsten, rhenium, manganese, zinc, vanadium or any element of IUPAC Groups 8-10, in particular platinum, palladium, rhodium, cobalt and nickel. Mixtures of metals can be used.

The final catalyst composition may contain 10 to 95% by weight, preferably 15 to 50% by weight, of the general binder component. The binder may generally be an inorganic oxide or a mixture thereof. Both amorphous and crystalline can be applied. Examples of suitable binders are silica, alumina, silica-alumina, clay, zirconia, silica-zirconia and silica-boria. Alonia is the preferred binder material.

In the cumene preparation, the final catalyst prepared with 80% by weight zeolite and 20% by weight alumina binder on a volatile free basis has preferably one, more preferably both of the following properties: (1) X-ray diffraction The absolute strength of the modified Y zeolite measured by (XRD) is preferably at least 50, more preferably at least 60; And (2) the framework aluminum of the modified Y zeolite is preferably at least 60%, more preferably at least 70% of the aluminum of the modified Y zeolite. In one embodiment, the final catalyst for cumene production is a product of the absolute strength of the modified Y zeolite, measured by XRD, and the percent backbone aluminum of aluminum in the modified Y zeolite, greater than 4200. In ethylbenzene production, the final catalyst preferably has one, more preferably both of the following characteristics: (1) The absolute strength of the modified Y zeolite as determined by X-ray diffraction (XRD) is preferably Is at least 65, more preferably at least 75; And (2) the framework aluminum of the modified Y zeolite is preferably at least 50%, more preferably at least 60% of the aluminum of the modified Y zeolite. In one embodiment, the final catalyst for the cumene product has a product of the absolute strength of the modified Y zeolite, as measured by XRD, and the percent backbone aluminum of aluminum in the modified Y zeolite, greater than 4500.

Although only specific embodiments have been mentioned, alternatives and modifications will become apparent to those skilled in the art from the foregoing description. Such and other alternatives are considered equivalent and are within the spirit and scope of the present disclosure and the appended claims.

Claims (17)

A catalyst comprising Y-85 or a modified LZ-210 zeolite, Zeolites having a skeletal aluminum content of at least 60% of the bulk aluminum content, including 60 to 90% by weight on the basis of no volatiles, the remainder containing alumina as a binder, having an absolute strength of 50 or more measured by X-ray diffraction (XRD) catalyst. The catalyst of claim 1 wherein the product of skeletal aluminum content expressed as an integral percentage and absolute strength is greater than 4200. 3. The catalyst according to claim 1 or 2, wherein the zeolite has a Na ₂ O content of less than 3% by weight based on the zeolite weight on anhydrous basis. 4. The catalyst of claim 1 wherein the zeolite has a bulk Si / Al ₂ molar ratio in the range of 6.5 to 27. 5. 5. The catalyst of claim 1, wherein the zeolite has an absolute strength of at least 60. 6. 5. The catalyst of claim 1 wherein the zeolite has at least 70% skeletal aluminum. The catalyst according to claim 1, wherein the loss on ignition (LOI) at 900 ° C. is in the range of 2 to 4 wt%. 8. The catalyst of claim 1 wherein the modified zeolite has a unit cell size of 24.58 mm 3 or less. 9. (a) treating the first zeolite with an aqueous ionic solution to yield a second zeolite having a first unit cell size; (b) hydrothermally treating the second zeolite at a temperature in the range of 550-850 ° C. to produce a third zeolite having a first bulk Si / Al ₂ molar ratio and having a second unit cell size smaller than the first unit cell size. ; And (c) Y-85 or modified having at least a portion of the sodium cation of the third zeolite exchanged with ammonium ions and having a second bulk Si / Al ₂ molar ratio in the range of 6.5 to 27 greater than the first bulk Sl / Al ₂ molar ratio; Contacting the third zeolite with an aqueous solution of a sufficient amount of ammonium ions having a pH of less than 4 for a time sufficient to produce LZ-210 zeolite, wherein the Y-85 or modified LZ-210 zeolite also has a Y- Up to 60% of all aluminum in 85 or modified LZ-210 zeolites with an absolute intensity of 50 or more measured by X-ray diffraction) Method for producing a Y-85 or modified LZ-210 zeolite catalyst comprising a. 10. The process of claim 9, wherein the ionic aqueous solution of step (a) comprises a fluorosilicate salt. The method according to claim 9 or 10, wherein step (a) comprises ammonium ion exchange of sodium in the first zeolite so that its sodium content is less than Na ₂ O 3% by weight based on the weight of the first zeolite on anhydrous basis. It further comprises reducing to. The process according to any one of claims 9 to 11, wherein the aqueous solution of step (c) comprises ions selected from the group consisting of nitrate ions and sulfate ions. The method of claim 9, wherein the contacting in step (c) comprises contacting the third zeolite with a first aqueous solution of ammonium ions having a pH of less than 4 to form a first mixture, wherein the first mixture Filtering to recover the filter cake, and further comprising contacting the filter cake with a second solution of ammonium ions having a pH of less than four. The process according to any one of claims 11 to 13, wherein the hydrothermal treatment in step (b) comprises a steam treatment. The method according to any one of claims 11 to 14, wherein after step (c) the Y-85 or modified LZ-210 zeolite is contacted with a dehydrating agent having a water concentration of less than 30 ppm by weight at a temperature in the range from 25 to 500 ° C. Phosphorus manufacturing method. A method of transalkylation of an aromatic comprising contacting the aromatic and the aromatic substrate with the catalyst of claim 1. 16. A process for transalkylation of an aromatic comprising contacting the aromatic and aromatic substrate with a catalyst prepared according to the method of any one of claims 9-15.

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