KR101474891B1 - Y-85 and modified lz-210 zeolites - Google Patents

Abstract

The present invention discloses a catalyst comprising Y-85 or a modified LZ-210 zeolite for converting polyalkylaromatics to monoalkylaromatics, especially cumene and ethylbenzene. For the production of cumene and ethylbenzene, the disclosed catalyst consisting of 80% zeolite and 20% alumina binder on a volatiless basis has at least one of the following physical properties: (1) Y-85 or modified LZ-210 zeolite Is preferably 50 or more, as measured by X-ray diffraction (XRD); And (2) the skeletal aluminum of Y-85 or the modified LZ-210 zeolite is preferably at least 60% of the aluminum of Y-85 or modified LZ-210 zeolite.

Description

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

Y-85 and modified LZ-210 zeolites, which can be used as catalysts in the polyalkylaromatics, for example PIPB and the transalkylation of PEB to cumene and ethylbenzene, are disclosed, together with their preparation methods.

The following description makes specific reference to the use of the catalysts disclosed herein for the transalkylation of polyisopropylbenzene (PIPB) with benzene to yield cumene, which is done for the purpose of clarifying and simplifying the description Should be considered. In the present specification, the broader scope of the present application will be frequently mentioned for emphasis.

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

Since both phenol and acetone are important as commercial chemicals, a great deal of emphasis has been placed on the manufacture of cumene, and the literature describes primarily the process for its manufacture. The most common and perhaps most direct way of making cumene is the alkylation of benzene with propylene, especially using acid catalysts.

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

The main direction of the reaction of benzene with PIPB yielding cumene corresponds to the addition of Markownikoff of the propyl group. However, a very small but significant amount of reaction takes place via the addition of semi-marconycopy to produce n-propylbenzene (NPB). NPB formation is significant in that it inhibits the oxidation of cumene to phenol and acetone, and therefore the cumene used for oxidation should be very pure with respect to NPB content.

Since cumene and NPB are difficult to separate by conventional means (e.g., distillation), transalkylation of benzene by PIPB to produce cumene should be carried out with minimal NPB yield. One important factor to consider is that the use of acid catalysts for transalkylation increases NPB formation with increasing temperature. Thus, in order to minimize NPB formation, the transalkylation should be carried out at as low a temperature as possible.

Since DIPB and TIPB are not only a common feed for the transalkylation of benzene by PIPB but also are a common by-product of the alkylation of benzene by propylene in the formation of cumene, transalkylation is usually carried out with alkylation to produce less valuable by-products Minimize and generate additional cumene. In this combined process, cumene produced with both alkylation and transalkylation is typically recovered as a single product stream. Since NPB is also formed in the alkylation and the amount of NPB formation in the alkylation increases as the temperature rises, NPB generation in both the alkylation and transalkylation must be managed in relation to one another so that there is relatively no NPB in the cumene product stream.

There is a need for an optimal transalkylation catalyst for the production of, for example, cumene or ethylbenzene, which has 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 scrutinized as catalysts in aromatic transalkylation. However, there is a problem that Y zeolite results in unacceptably low rates of transalkylation at low temperatures desirable to minimize NPB formation.

Therefore, 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 ethylbenzene at a predetermined low temperature.

Summary of the Invention

In meeting the above-mentioned requirements, a catalyst comprising a modified Y zeolite and having a metal hydride content of less than 0.2 wt.% Is disclosed.

A modified Y zeolite is first prepared by ammonium ion-exchanging 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 less than 3% by weight, and the Y zeolite has a first unit cell size. Next, the sodium-containing Y zeolite is hydrothermally treated at a temperature in the range of from 550 ° C (1022 ° F) to 850 ° C (1562 ° F) to contain the sodium cations and having a first bulk Si / Al ₂ molar ratio, Producing a steamed Y zeolite having a second unit cell size smaller than the cell size. Finally, the steamed Y zeolite is contacted with a solution of at least a portion of the sodium cations in the steamed Y zeolite exchanged with ammonium ions and having a first bulk Si / Al ₂ molar ratio, preferably in the range of 6.5 to 27 Is contacted with a sufficient amount of an aqueous solution of ammonium ions having a pH of less than 4, preferably 2 to 4, for a period of time sufficient to produce a modified Y zeolite having a bulk Si / Al ₂ molar ratio. The unit cell size of the modified Y zeolite ranges from 24.34 to 24.58 Angstroms.

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 a LZ-210 zeolite having a first unit cell size. The fluorosilicate treated sample is then steamed at a temperature in the range of about 550 ° C (1022 ° F) to about 850 ° C (1562 ° F) to contain the sodium cations and having a first bulk Si / Al ₂ molar ratio, To produce a steamed LZ-210 zeolite having a smaller second unit cell size. Finally, the steam-treated LZ-210 zeolite is contacted with a solution of at least a portion of the sodium cations in the steamed LZ-210 zeolite exchanged with ammonium ions and having a first bulk Si / Al ₂ molar ratio greater than 6.5: Is contacted with a sufficient amount of an aqueous solution of ammonium ions having a pH of less than 4 for a time sufficient to produce a modified LZ-210 zeolite having a bulk Si / Al ₂ molar ratio. The unit cell size of the modified Y zeolite ranges from 24.34 to 24.58 Angstroms. Subsequently, acid extraction may be performed to remove excess skeletal aluminum. Y zeolite is treated with ammonium ion exchange (s) to reduce the sodium content of the catalyst to less than 1 wt.% Of Na ₂ O, prior to or after treatment of the zeolite with the fluorosilicate salt, / Al ₂ molar ratio can be maintained. In another embodiment, the fluorosilicate treated Y zeolite (or LZ-210) may be ammonium exchanged without further steam treatment to further reduce the Na ₂ O content to produce a material suitable for the above disclosure.

The disclosed fabrication technique affects the number and properties of the extra skeletal aluminum (and Lewis acid part) as shown by the Si / Al ₂ ratio change and the unit cell size change, which improves the diffusion characteristics, increases the catalytic activity, .

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

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

In another embodiment, the catalyst for ethylbenzene production has a product of the absolute strength of the modified Y zeolite measured by XRD and the product of the% skeletal aluminum in the aluminum of the modified Y zeolite of more than 4500.

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

Brief Description of Drawings

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

2 shows the ratio (y-axis, ppm by weight) of NPB to cumene versus DIPB conversion (x-axis,%) for the catalysts of Examples 2 to 4 and 7 of the present disclosure for Comparative Examples 1 and 5 Graphs.

3 is a graph showing the DIPB conversion (y-axis,%) versus temperature (x-axis, ° C) of the catalyst of Example 3 before regeneration (Example 7) and after regeneration For example.

4 shows the ratio (y-axis, weight ppm) of NPB to cumene in the product to the catalyst of Example 3 before regeneration (Example 7) and after regeneration (Example 9) for Control Example 1 versus DIPB conversion (x-axis,%) as a graph.

5 graphically illustrates the DEB conversion (y-axis,%) versus temperature (x-axis, ° C) for the catalyst of Example 2, above, in which the alkyl group and the catalyst described above work well in addition to the profile for Control 1 Respectively.

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

Figure 7 graphically illustrates the NPB ratio (y-axis, weight ppm) versus DIPB conversion (x-axis,%) for cumene among the products for the catalysts of Examples 14-16 of this disclosure for Control Example 11.

FIG. 8 graphically shows the 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 against Control Example 11.

9 shows the ratio (y-axis, ppm weight) of NPB to cumene versus DIPB conversion (x-axis,%) of the products before and after regeneration (Example 19) Illustrated as a diagram.

DETAILED DESCRIPTION OF THE INVENTION

An improved catalyst comprising a crystalline zeolite molecular sieve is 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

With regard to the Y zeolites of the disclosure, the Y type zeolite is described in US 3,130,007, which is incorporated herein by reference in its entirety. Modified Y zeolites suitable for use in making the catalysts disclosed herein exhibit substantial modification of the Y zeolite framework structure and composition from the Y zeolite and generally increase the bulk Si / Al ₂ molar ratio to a value typically above 6.5, And / or by decreasing the unit cell size. However, in converting the Y zeolite starting material to a modified Y zeolite useful in the presently disclosed process, the resulting modified Y zeolite has an X-ray powder diffraction pattern for the Y zeolite as described in the '007 patent ≪ / RTI > The modified Y zeolite may have an X-ray powder diffraction pattern similar to that of the '007 patent, but as will be appreciated by those skilled in the art, the d-spacing shifts somewhat due to cation exchange, calcination, etc., It is generally necessary to convert it into a form.

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

In preparing the modified Y zeolite component in the disclosed catalyst, the starting material may be a Y zeolite in the alkali metal (e.g., sodium) form as described in the '007 patent. The Y zeolite in the alkali metal form is ion exchanged with an ammonium ion or an ammonium ion precursor such as quaternary ammonium or other nitrogen containing organic cation to convert the alkali metal content to an alkali metal oxide (e.g., Na ₂ O) 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 a zeolite on a dry or dry basis refers to the weight after maintaining the zeolite at a temperature of 900 ° C (1652 ° F) for approximately 2 hours.

Optionally, the starting zeolite may also contain rare earth cations in such a degree that the rare earth content as RE ₂ O ₃ constitutes from 0.1 to 12.5% by weight, preferably from 8.5 to 12% by weight of the zeolite (on anhydrous basis) And may be ion-exchanged to contain cations. Those skilled in the art will understand that the ion exchange capacity of zeolites for introducing rare earth cations is reduced during the disclosed process steps. Thus, when rare-earth cation exchange occurs, for example, as a final step in the manufacturing process, it may not be possible to introduce only the desired amount of rare earth cations. The skeletal Si / Al ₂ ratio of the starting Y zeolite may range from 3 to 6, but it is advantageous that it is greater than 4.8.

The manner of carrying out this primary ammonium ion exchange is not an important factor and can be carried out in a manner known to those skilled in the art. For example, such conventional ammonium ion exchange is carried out at a pH of 4 or above. It is advantageous to use a three step procedure with a 15 wt% aqueous ammonium nitrate solution in such a ratio that the initial weight ratio of the ammonium salt to the zeolite is 1 at each step. The contact time between the zeolite and the exchange medium is 1 hour at each step and the temperature is 85 ° C (185 ° F). The zeolite is washed with 7.5 l (2 gal) of water per 0.45 kg (1 lb) of zeolite between the stages. The exchanged zeolite is then dried at 100 ° C (212 ° F) to reach a loss on ignition (LOI) of 20% by weight at 1000 ° C. When using rare earth cations, it is preferred to contact the already-ammonium-exchanged form of the zeolite with an aqueous solution of a rare earth salt in a known manner. Adding the mixed rare earth chloride salt to an aqueous slurry of ammonium Y zeolite (the zeolite 1 RECl ₃ 0.386 g per g) exchange in the temperature range of 85~95 ℃ by rare earth content is generally a rare earth as RE ₂ O ₃ 8.5~12 % ≪ / RTI > by weight of the zeolite product.

After the ammonium ion exchange is complete, the steam treatment of the ammonium-exchanged, optionally rare-earth exchanged Y zeolite is carried out at 550 to 850 ° C (1022 ° C) for a time sufficient to reduce the unit cell size to less than 24.60 Å, preferably in the range of 24.34 to 24.58 Å. To 1562 ° F) or 600 to 725 ° C (1112 to 1337 ° F) with a steam environment containing at least 2 psia of steam, preferably 100% steam. 100% steam can be used for 1 hour at a temperature in the range of 600 to 725 ° C (1112 to 1337 ° F). The steam treatment step is to be noted that it is not necessary for the starting Y zeolite or higher Si / Al ₂ ratio of 6.5, as illustrated by the silicate treatment material fluoro, higher Si / Al ₂ ratio of extracting acid further processing And provides sufficient stability to undergo catalyst preparation and hydrocarbon conversion processes.

Low pH, ammonium ion exchange is an important aspect of making a modified Y zeolite composed of the catalyst used in the process disclosed herein. This exchange can be carried out in the same way as in the case of the initial ammonium exchanges, except that the pH of the exchange medium is reduced to at most 4, preferably less than 3, at least during the part of the ion exchange procedure. The reduction in pH is easily accomplished by adding an appropriate inorganic or organic acid to the ammonium ion solution. Nitric acid is particularly suitable for this purpose. It is preferred to avoid acids which form insoluble ammonium salts. In conducting low pH ammonium ion exchange, both the pH of the exchange medium, the amount of exchange medium for the zeolite, and the contact time of the exchange medium and the zeolite are important factors. As long as the pH of the exchange medium is below 4, it is confirmed that the sodium cations are exchanged for the hydrogen cations in the zeolite, and at least some aluminum, mainly non-framework aluminum and some framework aluminum are extracted. However, the effect of the method is improved by acidifying the ion exchange medium with an acid more than necessary to lower the pH to just under 4. As will be apparent from the data to be described later, the more acid exchanged the medium, the greater the tendency to extract the skeletal aluminum as well as the non-skeletal aluminum from the zeolite. The extraction procedure is carried out to a sufficient degree to produce a zeolite product having a bulk Si / Al ₂ ratio of 6.5 to 35. In another embodiment, the bulk Si / Al ₂ ratio is from 6.5 to 23, more preferably from 6.5 to 20.

A typical Y zeolite having a total silica to alumina Y modified Y zeolite used in the catalyst of the process disclosed herein contains Y zeolite represented by Y-85. U.S. Pat. Nos. 5,013,699 and 5,207,892, which are incorporated herein by reference, describe Y-85 zeolite and its preparation, and need not be described here in detail.

As illustrated in Figures 1-5 and in the Examples below, the catalysts disclosed above provide increased catalyst activity and low NPB formation in the case of cumene production. In the case of producing ethylbenzene from poly-ethylbenzene (FIG. 5), the diffusion characteristics of the disclosed catalysts appear to be important, even though the internal isomerization of the ethyl group is scarcely concerned and the ethyl group is less than the propyl group.

While the catalysts disclosed above may contain metal hydrogenation catalyst components, such components are not a requirement. Based on the weight of the catalyst, such a metal hydrogenation catalyst component may be present in an amount less than 0.2 wt.% Or less than 0.1 wt.%, Calculated as the individual monohydrate 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, a sulfide, a halide or the like, or in an elemental metal state. As used herein, the term " metal hydrogenation catalyst component " includes these various compound forms of the metal. The catalytically active metal may be attached to or supported by the internal adsorption region of the zeolite component on the external surface of the zeolite crystal, i. E., Contained within the pore system, or if a binder, diluent or other component is used. The metal may be imparted to the overall composition by any method to obtain a highly dispersed state. Suitable methods include injection, adsorption, cation exchange and intensive mixing. The metal may be copper, silver, gold, titanium, chromium, molybdenum, tungsten, rhenium, manganese, zinc, vanadium or any of the elements of the IUPAC groups 8-10, particularly platinum, palladium, rhodium, cobalt and nickel. Mixtures of metals may be used.

The final catalyst composition may contain conventional binder components in an amount of 10 to 95 wt%, preferably 15 to 50 wt%. The binder may generally be an inorganic oxide or a mixture thereof. Both amorphous and crystalline can be applied. Examples of suitable binders include silica, alumina, silica-alumina, clay, zirconia, silica-zirconia and silica-boria. Alunia is the preferred binder material.

In the manufacture of cumene, the final catalyst prepared from 80% by weight of zeolite and 20% by weight of alumina binder on a volatile-free basis preferably has one or more of the following properties: (1) X-ray diffraction The absolute strength of the modified Y zeolite measured by differential scanning calorimetry (XRD) is preferably not less than 50, more preferably not less than 60; And (2) the skeleton 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 the manufacture of cumene has an absolute intensity of the modified Y zeolite measured by XRD and a product of% framework aluminum of aluminum of the modified Y zeolite above 4200. In ethylbenzene production, the final catalyst preferably has one or more of the following characteristics: (1) The absolute strength of the modified Y zeolite, as measured by X-ray diffraction (XRD), is preferably Is at least 65, more preferably at least 75; And (2) the skeleton 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 product of the% skeletal aluminum in the aluminum of the modified Y zeolite is greater than 4500.

In one embodiment, the process disclosed herein employs a substantially dry catalyst. After the low pH ammonium ion exchange, a calcination step to remove substantially all of the water present is not necessary. It has been confirmed 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 has been confirmed that the water content of the zeolite must be relatively low before being used in the transalkyl process.

Excessive water can reduce the number of active sites and limit diffusion to that site, thereby not effectively promoting transalkylation. To solve this problem, dehydration of the catalyst particles and incorporation of a predetermined amount of water can be carried out prior to initiation by a desiccant which can be introduced into the transalkylation reaction zone, The transalkylatable aromatics can be slowly increased prior to introduction. During this initial heating period, the water content of the zeolite is determined by the equilibrium between the amounts 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 highly hydrophilic and the hydration level is controlled by controlling the rate at which the desiccant goes beyond the catalyst and the temperature during the dehydration step. The desiccant may be any agent that 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 and 500 ° C (77-932 ° F). The water content of the catalyst is calculated by measuring the ignition loss (LOI), which is calculated by calculating the weight after heating for 2 hours at 900 ° C (1652 ° F), and then calculating the weight loss due to ammonium ion decomposition Measure by subtracting the amount. A catalyst containing water in an amount larger than a predetermined amount, that is, an amount larger than the equilibrium amount of water contained in the catalyst at any time during the start of the process, will lose water as soon as equilibrium is established during initiation, The dehydration step to provide the amount of water to the catalyst may be desirable but need not be done.

Any desirable properties of the catalyst, such as crushing strength and ammonium ion concentration, are obtained by controlling the time and temperature conditions under which the extruded catalyst particles are fired. In some cases, calcination at high temperatures leaves a required amount of water in the catalyst, making it unnecessary to carry out the individual dehydration step. Thus, 'dewatering' and 'dewatering' as used herein means not only a separation step in which water is removed by the catalyst after the calcination step but also a calcination step carried out under the condition that a predetermined amount of water remains in the catalyst particles do.

The dewatering procedure described above is part of the actual process of making the catalysts described above in the production plant. However, procedures other than those described above can be used to dehydrate 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 be treated in a transalkylation reactor with a water-deficient gas, such as dry molecular nitrogen or air, or a dry reactant such as a dry aromatic substrate (e.g., benzene) or a dry transalkylatable aromatic For example, DIPB or TIPB) can be dehydrated in the system at a relatively high temperature on the catalyst until the catalyst contains a predetermined amount of water. In the dehydration step in the system, 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 [deg.] C (77 [deg.] F) and 500 [deg.] C (932 [deg.] F). In one embodiment, the catalyst is contacted with flow-drying nitrogen on a gaseous phase at 250 캜 (482 ℉). The catalyst may be heated to a temperature of, for example, from about 266 DEG F to about 500 DEG F, from about 320 DEG F to about 210 DEG C, from about 180 DEG C to about 356 DEG F to about 392 DEG F, (302 ° F) to 180 ° C (356 ° F) liquid stream of dry benzene. The catalyst particles may also be stored in the manufacturing plant or elsewhere to contact the surrounding gas until a predetermined amount of water is desorbed.

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

The disclosed catalysts are useful for transalkylating transalkylatable aromatics. The transalkylation process disclosed herein preferably allows transalkylatable hydrocarbons as a feed with an aromatic base. The transalkylatable hydrocarbons useful in the transalkylation process are comprised of aromatic compounds characterized in that they constitute molecules consisting essentially of an aromatic substrate replacing at least one hydrogen atom around the aromatic-based ring structure with at least one alkylating agent.

The alkylating agent compounds may be selected from the group of various materials including monoolefins, diolefins, polyolefins, acetylenic hydrocarbons, and alkyl halides, alcohols, ether esters, the latter being the various esters of alkyl sulfates, alkyl phosphates and carboxylic acids. Can be selected. Preferred olefinic functional compounds are olefinic hydrocarbons containing monoolefins containing one double bond per molecule. The monoolefins that can be used as the olefin functional compounds in the process described above are generally gaseous or generally liquid and include ethylene, propylene, 1-butene, 2-butene, isobutylene, Such as various pentenes, hexenes, heptenes, octenes, and mixtures thereof, and high molecular weight liquid olefins, the latter including various olefin oligomers with 9 to 18 carbon atoms per molecule, such as propylene trimers, propylene tetramers, And the like. C ₉ -C ₁₈ normal olefins may also be used, as the cycloolefin, such as cyclopentene, methylcyclopentene, cyclohexene, methylcyclohexene, etc., may be used, although the equilibrium results are not essential. The monoolefin preferably contains 2 to 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 may be selected individually from the group of aromatic compounds comprising benzene and a mixture of benzene and a single ring alkyl substituted benzene having the structure:

Wherein R is a hydrocarbon containing 1 to 14 carbon atoms and n is an integer of 1 to 5. On the other hand, the aromatic-based portion of the feedstock may be benzene, benzene containing from 1 to 5 alkyl 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 mixtures thereof. It is particularly preferred that the aromatic base is benzene.

The disclosed transalkylation process may have multiple purposes. In one object, a catalyst in the transalkylation reaction zone is used to remove excess alkylating agent compounds from the ring structure of the polyalkylated aromatic compound, transferring the alkylating agent compound to a previously un-alkylated aromatic based molecule, Thereby increasing the amount of the predetermined aromatic compound. For related purposes, the reaction carried out in the transalkylation reaction zone involves removing all alkylating agents from the substituted aromatic compounds, and converting the aromatic base to benzene during this.
The feed mixture is preferably present at a concentration of 20 ppm by weight, more preferably 10 ppm by weight, based on the weight of the transalkylatable aromatic and aromatic substrates through which the concentration of water and oxygen-containing compounds in the compounded feed passes into the reaction zone, Even more preferably 2 ppm by weight. The way in which very low concentrations of the feed mixture are achieved is not important in the process disclosed herein. In general, there is provided one stream containing a transalkylatable aromatic, and another stream containing an aromatic substrate, each stream comprising water and / or a mixture of feeds formed from the combination of the individual streams, Containing compound precursor. The water and oxygen containing compounds may be removed from the individual stream or feed mixture 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.

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In order to transalkylate the polyalkylaromatics by aromatic substrates, a feed mixture containing aromatic substrates and polyalkylated aromatic compounds in a molar ratio of 1: 1 to 50: 1, preferably 1: 1 to 10: 1, Continuously or intermittently into a transalkylation reaction zone containing the catalysts described above under transalkylation conditions comprising a temperature of from about < RTI ID = 0.0 > 190 C to < / RTI > The pressure suitable for use in this application is preferably 1 atm (101.3 kPa (a)), but should not exceed 130 atm (13169 kPa (a)). A particularly preferred pressure range is 10 to 40 atm (1013 to 4052 kPa (a)). The weight hourly space velocity (WHSV) of 0.1 to 50 hr ^-1 , especially 0.5 to 5 hr ^-1 , based on the catalyst total weight of the polyalkylaromatic feed rate and dry weight is preferred. It should be noted that although the processes disclosed herein can be carried out in a vapor phase, the temperature and pressure combinations applied in the transalkylation reaction zone are preferably such that the transalkylation reaction is carried out in liquid phase. In a liquid transalkylation process for preparing monoalkylaromatics, the catalyst is continuously washed with the reactants to prevent coke precursor formation on the catalyst. This reduces the amount of carbon formation on the catalyst, and in this case, the catalyst circulation life is prolonged as compared to the gas phase transalkylation process, where coke formation and catalyst inertness are the major problems. In addition, the selectivity to monoalkyl aromatic production, particularly cumene production, is higher than that of the catalytic gas phase transalkylation reaction.

The transalkylation conditions for the processes disclosed herein generally include a molar ratio of aromatic ring groups per alkyl group of 1: 1 to 25: 1. The molar ratio may be less than 1: 1, and the molar ratio may be 0.75: 1 or less. The molar ratio of the aromatic ring group per alkyl propyl group (propyl group in the production of cumene) is preferably 6: 1 or less.

Under transalkylation conditions, the catalyst particles typically have a water content of less than or equal to 4 wt%, more preferably less than or equal to 3 wt%, and even more preferably less than or equal to 2 wt%, measured by Karl Fischer titration, , And contains nitrogen in an amount of preferably 0.05% by weight or less, as measured by micro (CHN) (carbon-hydrogen-nitrogen) analysis.

In the present work, all of the elements of the periodic table are firstly marked on the inside of 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 the 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 specific period of time. The molecular number of the aromatic ring group is the sum of all the aromatic ring groups, regardless of the compound in which the aromatic ring group is generated. For example, in the manufacture of cumene, 1 mole of benzene, 1 mole of cumene, 1 mole of DIPB and 1 mole of TIPB each contribute 1 mole of aromatic ring group to the sum of aromatic ring groups. In the production of ethylbenzene (EB), 1 mol of benzene, 1 mol of EB and 1 mol of di-ethylbenzene (DEB) each contribute 1 mol of aromatic ring group to the sum of the aromatic ring groups. The denominator in 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 phase and pass through the reaction zone for a specified 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 group, regardless of the compound on which the alkyl or alkenyl group is to occur, except that paraffin is not included. Thus, the number of moles of the propyl group is such that an iso-propyl, n-propyl or propenyl group occurs except that paraffins, e.g., propane, n-butane, isobutane, pentane and higher paraffins are excluded from the calculation of the number of moles of propyl groups N-propyl and propenyl groups, irrespective of the compound being used. For example, 1 mole of propylene, 1 mole of cumene, and 1 mole of NPB each contribute 1 mole of the propyl group to the sum of the propyl groups, while correlating with the distribution of the three groups between iso-propyl and n- 1 mol of DIPB contributes 2 mol of propyl groups and 1 mol of tri-propylbenzene contributes 3 mol of propyl groups. 1 mole of ethylene and 1 mole of EB each contribute 1 mole of ethyl group to the total of ethyl groups while 1 mole of DEB contributes 2 mole of ethyl group and 1 mole of triethylbenzene contributes 3 mole of ethyl group. Ethane does not contribute to the moles of the ethyl group.

As used herein, WHSV refers to the weight hourly space velocity, defined as the weight flow rate per hour divided by the weight of the catalyst, wherein the weight flow rate and catalyst weight are the same weight units.

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

The surface area in this section is determined according to 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. The nitrogen partial pressure p / po data points in the range 0.03-0.30 were determined using the BET (Brunauer-Emmett-Teller) model method using the nitrogen adsorption technique as described in Soc, 60 (2), 309-319 Respectively.

As referred to herein, the absolute intensity by X-ray powder diffraction (XRD) of the Y zeolite material is calculated by calculating the standard sum of the intensities of several selected XRD peaks of the Y zeolite material and adding the sum to the first standard, By the standard sum of several XRD peak intensities of the alpha-alumina NBS 674a strength standard certified by the American Institute of Standards and Technology (NIST). The absolute strength of the Y zeolite is the ratio of the sum multiplied by 100:

Absolute intensity = (standard intensity of Y zeolite material peak) / (standard intensity of alpha-alumina standard peak) x 100

Y < / RTI > zeolite material and the alpha-alumina standard are given 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 Depending on the zeolite content, 1 second / step or more 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 the Y zeolite, which is blended with the non-zeolitic binder to yield a mixture of Z parts by weight of Y zeolite and 100 parts by weight of non-zeolitic binder (100-Z) on a dry weight basis, Can be calculated using the equation A = C. (100 / Z), where A is the absolute intensity of the Y zeolite and C is the absolute intensity of the mixture. For example, if the Y zeolite is mixed with HNO ₃ demineralized 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 the Y zeolite is calculated as (60) · (100/80) or 75.

As used herein, a unit cell size, sometimes referred to as a lattice parameter, is defined as the (111) peak for the 642, (822), (555), (840) Quot; refers to a unit cell size calculated using a method of using a profile fitting for finding an XRD peak position of the unit cell.

As used herein, the bulk Si / Al ₂ molar ratio of the zeolite is determined by the addition of silica to Al ₂ O ₃ (Al ₂ O ₃ ), based on the total or total amount of aluminum and silicon (skeletal and non-skeletal) present in the zeolite To SiO ₂ ), and sometimes represents the total molar ratio of silica (SiO ₂ to Al ₂ O ₃ ) to alumina in this round. 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 the zeolite, which is the skeletal aluminum, Kerr, A.W. Chester, and D.H. Olson, Acta. Phys. Chem., 1978, 24, 169] and the paper [G.T. Kerr, Zeolites, 1989, 9, 350].

As used herein, dry weight means that the weight is based on weight after drying time of 1 hour at flow 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 disclosure.

Example 1 - Control Example

A sample of Y-74 zeolite was slurried in a 15 wt% aqueous NH ₄ NO ₃ solution and the solution temperature was raised to 75 ° C (167 ° F). The Y-74 zeolite was a stabilized sodium Y zeolite having a bulk Si / Al ₂ ratio of about 5.2, a unit cell size of about 24.53, and a sodium content of about 2.7% by weight, calculated as dry Na ₂ O. The Y-74 zeolite has a bulk Si / Al ₂ ratio of approximately 4.9, a unit cell size of approximately 24.67, and a sodium content of greater than or equal to the normal of the process described in column 4 of the dry Na ₂ O (US 5,324,877, line 47 to column 5, From about 9.4% by weight of sodium Y zeolite, calculated as ammonium exchanged to remove about 75% of Na by subsequent steps (1) and (2) and de-aluminumed by steaming at about 1100 F . Y-74 zeolite was prepared and obtained from UOP LLC, Des Plaines, 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 more times, and the resulting filter cake had a bulk Si / Al ₂ ratio of 5.2 and a sodium content of 0.13 wt% calculated as a dry Na ₂ O, The unit cell size was 24.572 ANGSTROM, and the absolute intensity measured by X-ray diffraction was 96. The resulting filter cake was dried to the appropriate wetting level and mixed with HNO ₃ degassed Pural SB alumina to yield a mixture of 80 parts by weight zeolite and 20 parts by weight Al ₂ O ₃ binder in dry weight, ) Diameter cylindrical extrudate. The extrudate was dried and calcined in flowing air at approximately 11O < 0 > F (600 < 0 > C) for 1 hour. The catalyst represented the prior art. The catalyst had a unit cell size of 24.494 Å, an XRD absolute strength of 61.1, and a skeletal aluminum of 57.2% as a percentage of aluminum in modified Y zeolite.

Example 2

Another sample of the Y-74 zeolite used in Example 1 was slurried in an aqueous 15 wt% NH ₄ NO ₃ solution. The pH of the slurry was lowered from 4 to 2 by adding a sufficient amount of a 17 wt% HNO ₃ solution. The slurry temperature was then heated to 75 ° C (167 ° F) and held for 1 hour. After 1 hour of contact at 75 ° C (167 ° F), the slurry was filtered and the filter cake washed with excess warm deionized water. NH ₄ ⁺ was an ion present under the acid extraction of the exchange, filtration and water washing step was repeated once, the resulting filter Kate 0.01 measured as Na ₂ O in the bulk Si / Al ₂ ratio of 11.5, and a sodium content of dry weight %, And the unit cell size was 24.47 ANGSTROM. The resulting filter cake was dried to the appropriate wetting level and mixed with HNO ₃ degassed Pural SB alumina to yield a mixture of 80 parts by weight zeolite and 20 parts by weight Al ₂ O ₃ binder in dry weight, ) Diameter cylindrical extrudate. The extrudate was dried and calcined in flowing air at approximately 11O < 0 > F (600 < 0 > C) for 1 hour. The characteristics of the catalyst were as follows: 68.2 wt% of SiO ₂ in bulk and dry weight, 30.5 wt% in Al ₂ O ₃ dry weight, 0.04 wt% in sodium dry weight Na ₂ O, 0.03 wt% in (NH ₄ ) ₂ O dry weight The unit cell size was 24.456 ANGSTROM, the absolute XRD intensity was 66.5, the skeletal aluminum as a percentage of ammonium in the modified Y zeolite was 92.2%, and the BET surface area was 708 m < ² > / g.

Example 3

Another sample of the Y-74 zeolite used in Example 1 was slurried in an aqueous 15 wt% NH ₄ NO ₃ 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 < RTI ID = 0.0 > 79 C < / RTI > (175 F) and held for 90 minutes. After 90 minutes of contact at 175 DEG F (79 DEG C), the slurry was filtered, 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 is a bulk Si / Al ₂ ratio of 8.52, was measured as the sodium content of the dry weight of Na ₂ O 0.18% by weight. The resulting filter cake was dried in the manner described in Example 2, mixed with HNO ₃ demineralized Pural SB alumina, extruded, dried, and calcined. The characteristics of the catalyst were as follows: the unit cell size was 24.486 Å, the absolute XRD intensity was 65.8, the percentage of ammonium in the modified Y zeolite was 81.1%, and the BET surface area was 698 m ² / g.

Example 4

Example 4 follows the same procedure 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 an aqueous 15 wt% NH ₄ NO ₃ solution. It was added over a 17 wt.% HNO ₃ in sufficient amount for a 30 minute time to remove the extra framework aluminum. The slurry temperature was then heated to < RTI ID = 0.0 > 79 C < / RTI > (175 F) and held for 90 minutes. After 90 minutes of contact at 79 DEG C (175 DEG 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, measured as a dry Na ₂ O. The resulting filter cake was dried in the manner described in Example 2, mixed with HNO ₃ demineralized Pural SB alumina, extruded, dried, and calcined. The characteristics of the catalyst were as follows: the unit cell size was 24.434 Å, the absolute XRD intensity was 53.6, the percentage of ammonium in the modified Y zeolite was 74.9%, and the BET surface area was 732 m ² / g.

Example 5 - Control Example

Example 5 follows the same procedure 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 an aqueous 15 wt% NH ₄ NO ₃ solution. A sufficient amount of 17 wt% HNO ₃ was added over a period of 30 minutes to increase the bulk Si / Al ₂ ratio. The slurry temperature was then heated to < RTI ID = 0.0 > 79 C < / RTI > (175 F) and held for 90 minutes. After 90 minutes of contact at 79 DEG C (175 DEG 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 is a bulk Si / Al ₂ ratio of 11.15, a sodium content was measured as the dry weight of Na ₂ O 0.08% by weight. The resulting filter cake was dried to the appropriate wetting level and mixed with HNO ₃ degassed Pural SB alumina to yield a mixture of 80 parts by weight Zeolite and 20 parts by weight Al ₂ O ₃ binder in dry weight, 16 in.) Diameter cylindrical extrudate. The extrudate was dried and calcined in flowing air at approximately 11O < 0 > F (600 < 0 > C) for 1 hour. The characteristics of the catalyst were as follows: the unit cell size was 24.418 Å, the absolute XRD strength was 44.8, the percentage of ammonium in the modified Y zeolite was 75.2%, and the BET surface area was 756 m ² / g.

Example 6

The same stabilized Y-74 used in Example 1 was slurried in an aqueous 15 wt% NH ₄ NO ₃ solution. The total amount of HNO ₃ used in this example is the same as that in Example 5. However, instead of carrying out the acid extraction in a single step as described in Example 5, the acid extraction used 85% of the total HNO ₃ acid in the first stage and the remaining 15% of the total acid in the second stage Using the two-step method. 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 < RTI ID = 0.0 > 79 C < / RTI > (175 F) and held for 90 minutes. After 90 minutes of contact at 79 DEG C (175 DEG F), the slurry was filtered and the filter cake was washed with excess warm deionized water. The filtration and water washing steps were repeated with acid extraction in the presence of NH ₄ ⁺ (with the remaining 15% of the total HNO ₃ used) and the resulting filter cake had a bulk Si / Al ₂ ratio of 11.14 and a sodium content of Measured as Na ₂ O, was 0.09% by weight. The resulting filter cake was dried to the appropriate wetting level and mixed with HNO ₃ degassed Pural SB alumina to yield a mixture of 80 parts by weight zeolite and 20 parts by weight Al ₂ O ₃ binder in dry weight, ) Diameter cylindrical extrudate. The extrudate was dried and calcined in flowing air at approximately 11O < 0 > F (600 < 0 > C) for 1 hour. The characteristics of the catalyst were as follows: the unit cell size was 24.411 Å, the absolute XRD intensity was 56.1, the percentage of ammonium in the modified Y zeolite was 72.5%, and the BET surface area was 763 m ² / g.

Example 7

The same stabilized Y-74 used in Example 3 was slurried in an 18 wt% ammonium sulfate solution. To the solution, 17% sulfuric acid solution was added over 30 minutes. The batch was then heated to < RTI ID = 0.0 > 79 C < / RTI > (175 F) and held for 90 minutes. The heat was removed and the batch was then quenched with processing water to reduce the temperature to 62 ° C (143 ° F) and filter. The Y zeolite material was then reslurried in a 6.4 wt% ammonium sulfate solution and maintained at 79 DEG C (175 DEG F) for 1 hour. The material was then filtered and washed with water. The resulting filter cake had a bulk Si / Al ₂ ratio of 7.71 and a sodium content of 0.16% by weight, measured as a dry Na ₂ O. The filter cake produced in the manner described in Example 2 was dried, mixed with HNO ₃ degassed Pural SB alumina, extruded, dried and calcined. The characteristics of the catalyst were a unit cell size of 24.489 Å, an absolute XRD strength of 65.3, and a skeletal aluminum content of 75.7% as a percentage of ammonium in the modified Y zeolite.

Table 2 summarizes the characteristics 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 Drawings 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 Catalyst 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 skeleton Aluminum, total atomic% 57.2 92.2 81.1 74.9 75.2 72.5 75.7 Catalyst BET surface area, m ² / g - 708 698 732 756 763 -

Example 8

The catalysts prepared in Examples 1 to 5 and 7 were tested for transalkylation performance with feeds containing benzene and polyalkylated benzene. The feed was prepared by combining with polyalkylated benzene obtained from a commercial transalkylation unit having benzene. The prepared feed formulation means a typical transalkylation feed composition in which the molar ratio of aromatic ring groups to propyl groups is approximately 2.3. It has been found that catalysts produced by the processes disclosed herein provide the same advantages when processing feeds at substantially lower or higher feed molar ratios. The feed compositions as determined by gas chromatography are summarized in Table 3. The test is 3447 kPa (g) (500 psi (g)) reactor pressure, at a molar ratio of 0.8 and a ring group for groups of 2.3 profile once-through under the conditions of DIPB WHSV over the reaction temperature range from ^-1 (the once- through mode in a fixed bed reactor. The reactor was allowed to achieve a substantially steady state at each reaction temperature and the product was sampled for analysis. During the test, substantially no catalyst deactivation occurred. Prior to charging the feed, each catalyst was subjected to a drying procedure by contacting it with a stream of nitrogen stream containing less than 10 ppm by weight of water at 250 ° C (482 ° F) for 6 hours.

ingredient Concentration, weight% 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 transalkylation of the poly-alkylate to cumene contributing to the catalyst produced by the process 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 used catalyst was placed in a ceramic dish and it was introduced into a muffle furnace. The temperature of the furnace is raised from 70 ° C (158 ° F.) to 1022 ° F. at a rate of 1 ° C per minute (1.8 ° F) while passing stream air through the muffle furnace, Lt; 0 > C (230 < 0 > F). After regeneration, the catalyst was tested again in the manner described in Example 8. < RTI ID = 0.0 >

FIGS. 3 and 4 show the test results before regeneration (indicated by 'Example 7') and before regeneration (indicated by 'Example 9'). The results show that the catalyst before and after the regeneration is similar to the activity and product purity better than the curve for the catalyst of Example 1, thus exhibiting excellent catalyst regeneration.

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 the catalyst was fed into the reactor, the catalyst was contacted with a stream of nitrogen stream containing less than 10% by weight of water at 250 < 0 > C (482 [deg.] F) for 6 hours to dry. Each catalyst was run over a reaction temperature range of 170 ° C (338 ° F) to 230 ° C (446 ° F) at 2 ° ^-1 p-DEB WHSV. The reactor was allowed to achieve a substantially steady state at each reaction temperature and the product was sampled for analysis. During the test, substantially no catalyst deactivation occurred. Figure 5 shows the results for both catalysts. The results show that the catalyst prepared in Example 2 has similar or superior activity and stability to the curve for the catalyst prepared in Example 1 and can be used in commercial poly-ethylbenzene transalkyl operation.

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

LZ-210

Y zeolite may be used in the process disclosed herein and may be prepared by dealumination of Y zeolites having a total molar ratio of silica to alumina of 5 or less and may be prepared by methods described in US 4,503,023, 4,597,956, 4,735,928 And 5,275,720. The '023 patent also discloses another procedure for dealumination of Y zeolites, including contacting an aqueous solution of a fluorosilicate salt with a Y zeolite using controlled rate, temperature and pH conditions avoiding aluminum extraction without silicon substitution do. The '023 patent uses a fluorosilicate salt as an aluminum extract and also as an exogenous silicon feed that is inserted into the Y zeolite structure instead of the extracted aluminum. The salt has the formula:

(A) _{2 / b} SiF ₆

Wherein A is a metallic or non-metallic cation other than H ^& lt ^{; +} & gt ^; having the value ' b & apos ^; . A cation represented by 'A' are ^{alkylammonium, NH4 +, 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 the zeolite-based aluminosilicate molecule LZ-210 described in the '023 patent. The LZ-210 zeolite and other zeolites of this group are readily prepared from the Y zeolite starting material. In one embodiment, the total molar ratio of silica to LZ-210 zeolite alumina is 5.0 to 11.0. The unit cell size is in the range of 24.38 to 24.50 ANGSTROM, preferably 24.40 to 24.44 ANGSTROM. The LZ-210 family of zeolites used in the present process and compositions disclosed herein retain a composition expressed in terms of the molar ratio of the oxides as in the following formula:

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

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

Generally, the LZ-210 zeolite can be prepared by dealumination of a Y-type zeolite using a solution of a fluorosilicate salt, preferably ammonium hexafluorosilicate. The dealumination can be accomplished by exchanging the Y zeolite, typically an essentially non-essential ammonium exchanged Y zeolite, 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 with an increased overall molar ratio of silica to alumina is produced. The extent of the increase depends at least in part on the amount of fluorosilicate solution in contact with the zeolite, and on the acceptable reaction time. 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 treated by methods known in the art. For example, the product may be contacted with steam at a partial pressure of 1.4 kPa (a) (0.2 psi (a)) or more at a temperature of 482 ° C (900 ° F) to 816 ° C (1500 ° F) for a period of 1/4 to 3 hours High crystalline stability can be provided. In some cases, the product of the steam calcination may be subjected to ammonium exchange treatment by methods known in the art. For example, the product can be slurried with water after adding the ammonium salt to the slurry. The resulting mixture is typically heated for a period of time, filtered, and washed with water. Steam treatment and ammonium exchange processes for LZ-210 zeolites are described in US 4,503,023, 4,735,928 and 5,275,720.

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

The final low pH ammonium ion exchange of the preferred LZ-210 zeolite is carried out in the presence of Y zeolite (and / or the LZ discussed above), except that the pH of the exchange medium is lowered to at most 4, -210 zeolite) in the same manner as in the case of the initial ammonium exchange. Reduction of the pH is easily carried out by adding an appropriate inorganic or organic acid to the ammonium ion solution. Nitric acid is particularly suitable for this purpose. It is preferable to avoid the acid forming the insoluble ammonium salt. In carrying out low pH ammonium ion exchange, the pH of both of the exchange medium, the amount of exchange medium for the zeolite and the contact time of the exchange medium and the zeolite are important factors. As long as the pH of the exchange medium is below 4, the sodium cations are exchanged for hydrogen cations in the zeolite, and at least some aluminum, mainly non-skeletal aluminum and some skeletal aluminum are extracted. However, the effect of this process is enhanced by acidifying the ion exchange medium with more acid than is necessary to lower the pH to just below 4. As is evident from the data mentioned below, the more acid exchanged the acid, the greater the tendency to extract the skeleton as well as the non-framework aluminum from the zeolite. The extraction procedure is carried out to a sufficient degree to produce a zeolite product having a bulk Si / Al ₂ molar ratio of 6.5-27. In another embodiment, the bulk Si / Al ₂ molar ratio is in the range of 6.5 to 23, or more preferably in the range of 6.5 to 20.

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

Example 11 - Control Example

A sample of Y-74 zeolite was slurried in a 15 wt% aqueous NH ₄ NO ₃ solution and the solution temperature was raised to 75 ° C (167 ° F). The Y-74 zeolite was a stabilized sodium Y zeolite having a bulk Si / Al ₂ ratio of about 5.2, a unit cell size of about 24.53, and a sodium content of about 2.7% by weight, calculated as dry Na ₂ O. The Y-74 zeolite has a bulk Si / Al ₂ ratio of approximately 4.9, a unit cell size of approximately 24.67, and a sodium content of greater than or equal to the normal of the process described in column 4 of the dry Na ₂ O (US 5,324,877, line 47 to column 5, From sodium Y zeolite, approximately 9.4% by weight, calculated as ammonium exchanged by subsequent steps (1) and (2) to remove approximately 75% of Na and dealuminated at approximately 11O < 0 > Respectively. Y-74 zeolite was prepared and obtained from UOP LLC, Des Plaines, 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 more times, and the resulting filter cake had a bulk Si / Al ₂ ratio of 5.2 and a sodium content of 0.13 wt% calculated as a dry Na ₂ O, The unit cell size was 24.572 ANGSTROM, and the absolute intensity measured by X-ray diffraction was 96. The resulting filter cake was dried to the appropriate wetting level and mixed with HNO ₃ degassed Pural SB alumina to yield a mixture of 80 parts by weight zeolite and 20 parts by weight Al ₂ O ₃ binder in dry weight, ) Diameter cylindrical extrudate. The extrudate was dried and calcined in flowing air at approximately 11O < 0 > F (600 < 0 > C) for 1 hour. The catalyst represented the prior art. The catalyst had a unit cell size of 24.494 Å, an XRD absolute strength of 61.1, and a skeletal aluminum of 57.2% as a percentage of aluminum in modified Y zeolite.

Example 12

The synthesized Y-54 zeolite was subjected to ammonium exchange treatment followed by ammonium fluorosilicate according to the procedure described in US 4,503,023. The Y-54 zeolite is a sodium Y zeolite having a bulk Si / Al ₂ ratio of about 4.9, a unit cell size of 24.67, and a sodium content of 9.4% by weight calculated as a dry Na ₂ O. Y-54 zeolite was prepared and obtained from UOP LLC, Des Plaines, Ill., USA. The resultant Y zeolite having a bulk Si / Al ₂ molar ratio of 6.5 was steamed at 100 ° C for 1 hour at 1120 ° F and then exchanged with ammonium. The resulting filter cake was dried to the appropriate wetting level and mixed with HNO ₃ degassed Pural SB alumina to yield a mixture of 80 parts by weight zeolite and 20 parts by weight Al ₂ O ₃ binder in dry weight, ) Diameter cylindrical extrudate. The extrudate was dried and calcined in flowing air at approximately 11O < 0 > F (600 < 0 > C) for 1 hour. The resulting catalyst had a unit cell size of 24.426 ANGSTROM, an absolute XRD strength of 81.6 and a framework aluminum content of 63.2% as a percentage of aluminum in the modified Y zeolite.

Example 13

According to the procedure described in US 4,503,023, the synthesized Y-54 zeolite was ammonium exchanged and then treated with ammonium fluorosilicate. The resulting Y zeolite having a bulk Si / Al ₂ ratio of 9.0 and referred to as LZ-210 (9) was steamed at 100 ° C (1112 ° F) for 1 hour with 100% steam. A slurry composed 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 캜 (185 ℉) 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, the cake washed with 1000 ml H ₂ O and dried overnight at 100 ° C (212 ° F). In the second part, 200 g of dry cake was added to a solution consisting of 667 g of 50 wt% (NH ₄ ) NO ₃ and 650 g of H ₂ O, and 20 g of 66 wt% HNO ₃ was added thereto. The resulting slurry was mixed for 60 minutes. The mixture was then filtered, washed with 1000 ml H ₂ O, and the filter cake dried overnight at 100 ° C (212 ° F). The resulting zeolite had a bulk Si / Al ₂ ratio of 10.82 and a Na ₂ O content of 0.026 wt%. The zeolite powder was mixed with HNO ₃ demineralized 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 weight and controlling the moisture to yield a suitable dough, 16 in.) Diameter cylindrical extrudate. The extrudate was dried and calcined in flowing air at approximately 11O < 0 > F (600 < 0 > C) for 1 hour. The resulting catalyst had a unit cell size of 24.430 Å, an absolute XRD strength of 78.4, a skeletal aluminum content of 77.8% and a BET surface area of 661 m ² / g.

Example 14

According to the procedure described in US 4,503,023, the synthesized Y-54 zeolite was ammonium exchanged and then treated with ammonium fluorosilicate. The resulting Y zeolite having a bulk Si / Al ₂ ratio of 9.0 and referred to as LZ-210 (9) was steamed at 100 ° C (1112 ° F) for 1 hour with 100% steam. An amount of 256 g of steamed LZ-210 (9) was added to 1140 g of 22 wt% 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 176 ° F (80 ° C) 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 and dried at 100 ° C. (212 ° F.) Oven-dried overnight. The resulting zeolite had a bulk Si / Al ₂ ratio of 14.38 and a Na ₂ O content of 0.047 wt%. The resulting zeolite powder was mixed with HNO ₃ demineralized 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 weight and controlling the moisture to yield a suitable dough fabric, 1/16 in.) Diameter cylindrical extrudate. The extrudate was dried and calcined in flowing air at approximately 11O < 0 > F (600 < 0 > C) for 1 hour. The resulting catalyst had a unit cell size of 24.393 Å, an absolute XRD strength of 79.6, a skeletal aluminum content of 81.8%, and a BET surface area of 749 m ² / g.

Example 15

According to the procedure described in US 4,503,023, the synthesized Y-54 zeolite was ammonium exchanged and then treated with ammonium fluorosilicate. The resulting Y zeolite, having a bulk Si / Al ₂ ratio of 12 and referred to as LZ-210 (12), was steamed at 100 ° C (1112 ° F) for 1 hour with 100% steam. A slurry composed 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 of 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 캜 (185 ℉) and mixed for 15 minutes. To the mixture was added 33.4 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, the cake washed with 1000 ml H ₂ O and dried overnight at 100 ° C (212 ° F). In the second part, 200 g of dry cake was added to a solution consisting of 667 g of 50 wt% (NH ₄ ) NO ₃ and 650 g of H ₂ O, to which was added 10 g of 66 wt% HNO ₃ . The resulting slurry was stirred for 60 minutes. The mixture was then filtered, washed with 1000 ml H ₂ O, and the filter cake dried overnight at 100 ° C (212 ° F). The resulting zeolite had a bulk Si / Al ₂ ratio of 17.24 and a Na ₂ O content of 0.01 wt%. The resulting zeolite powder was mixed with HNO ₃ demineralized 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 weight and controlling the moisture to yield a suitable dough fabric, 1/16 in.) Diameter cylindrical extrudate. The extrudate was dried and calcined in flowing air at approximately 11O < 0 > F (600 < 0 > C) for 1 hour. The resulting catalyst had a unit cell size of 24.391 Å, an absolute XRD strength of 81.2, a skeletal aluminum content of 94.9% and a BET surface area of 677 m ² / g.

Example 16

250 g of LZ-210 (120) of Example 15 (before steaming) was added to 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 < RTI ID = 0.0 > 95 C (203 F) < / RTI > The slurry was then filtered and washed with water. The cake was then NH ₄ NO ₃ exchanged 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 wt%. The dried zeolite was mixed with HNO ₃ demineralized 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 controlling the moisture to yield a suitable dough, / 16 in.) Diameter cylindrical extrudate. The extrudate was dried and calcined in flowing air at approximately 11O < 0 > F (600 < 0 > C) for 1 hour. The resulting catalyst had a unit cell size of 24.431 Å, an absolute XRD strength of 77.3, a skeletal aluminum content of 89.2% and a BET surface area of 660 m ² / g.

Table 4 summarizes the characteristics 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 Drawings 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 Catalyst 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 skeleton Aluminum, total atomic% 57.2 63.2 77.8 81.8 94.9 89.2 92.6 Catalyst 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 benzene. The feed was prepared by combining with polyalkylated benzene obtained from a commercial transalkylation unit having benzene. The feed compositions measured by gas chromatography are summarized in Table 2 above. The test is 3447 kPa (g) (500 psi (g)) of the reactor pressure, the reaction temperature over the range of molar ratio of 0.8 and ^-1 of the aromatic ring group for groups of 2.3 profile of once-through fixed bed reactor under the conditions of DIPB WHSV . The reactor was allowed to achieve a substantially steady state at each reaction temperature and the product was sampled for analysis. During the test, substantially no catalyst deactivation occurred. Prior to charging the feed, each catalyst was subjected to a drying procedure by contacting it with a stream of nitrogen stream containing less than 10 ppm by weight of water at 250 ° C (482 ° F) for 6 hours.

Figures 6 and 7 show test results for the catalysts prepared in Examples 11 and 14-16. In FIG. 6, the catalysts prepared in Examples 14-16 exhibit higher activity (i.e., higher DIPB conversion at a given temperature) than curve 601 for Example 11. In FIG. 7, the catalysts prepared in Examples 14-16 also exhibit product purity (i.e., lower NPB / cumene at a given DIPB conversion) than curve 701 for the catalyst prepared in Example 1. In Figures 6 and 7, the data for Example 16 shows that the steaming and acid extraction steps are not required in catalyst preparation, because excellent performance can be achieved even if both are omitted. Also, in Figures 6 and 7, the data for Example 14 demonstrates that acid extraction after one-step steam treatment instead of the two-step acid extraction of Example 15, even with more acid extraction conditions, 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, which was placed in a muffle furnace. The temperature of the furnace is raised from 70 ° C (158 ° F.) to 1022 ° F. at a rate of 1 ° C per minute (1.8 ° F) while passing stream air through the muffle furnace, Lt; 0 > C (230 < 0 > F). The regenerated catalyst had a unit cell size of 24.439 Å, an absolute XRD strength of 72.5, a skeletal aluminum content 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 catalysts before and after the regeneration were the same (i.e., DIPB conversion at a given temperature) and product purity (i.e., NPB / cumene at a given DIPB conversion) were the same, thus exhibiting excellent catalyst regeneration capability.

Example 19

A sample of the catalyst prepared in Example 14 was tested in the manner described in Example 17, as previously described. 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. [

Figures 8 and 9 graphically show the test results for the catalyst before regeneration (indicated as 'Example 14') and after regeneration (indicated as 'Example 19'). The results are similar for catalysts both before and after regeneration (i.e., DIPB conversion at a given temperature) and product purity (i.e., NPB / cumene at a given DIPB conversion) 8 and 9, respectively, which are superior to the respective curves 601 and 701), and thus exhibits excellent catalyst regeneration ability.

This example demonstrates the advantages of high activity and product purity in the transalkylation of cumene and DE-to-EB poly-alkylates, such as DIPB and TIPB, contributing to the catalyst produced by the process 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 an amount less than 0.2 wt.% Or less than 0.1 wt.%, Calculated as the 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 an oxide, a sulfide, a halide, or in an 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 in the internal adsorption region of the zeolite component on the outer surface of the zeolite crystal, i.e., in the pore system, or, if a binder, diluent or other component is used, it is adhered to or supported on these components. The metal may be imparted to the overall composition by any method to obtain a highly dispersed state. Suitable methods include injection, adsorption, cation exchange and intensive mixing. The metal may be copper, silver, gold, titanium, chromium, molybdenum, tungsten, rhenium, manganese, zinc, vanadium or any of the elements of the IUPAC groups 8-10, particularly platinum, palladium, rhodium, cobalt and nickel. Mixtures of metals may be used.

The final catalyst composition may contain conventional binder components in an amount of 10 to 95 wt%, preferably 15 to 50 wt%. The binder may generally be an inorganic oxide or a mixture thereof. Both amorphous and crystalline can be applied. Examples of suitable binders include silica, alumina, silica-alumina, clay, zirconia, silica-zirconia and silica-boria. Alunia is the preferred binder material.

In the manufacture of cumene, the final catalyst prepared from 80% by weight of zeolite and 20% by weight of alumina binder on a volatile-free basis preferably has one or more of the following properties: (1) X-ray diffraction The absolute strength of the modified Y zeolite measured by differential scanning calorimetry (XRD) is preferably not less than 50, more preferably not less than 60; And (2) the skeleton 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 the manufacture of cumene has a product of the absolute strength of the modified Y zeolite measured by XRD and the product of the% skeletal aluminum of aluminum in the modified Y zeolite above 4200. In ethylbenzene production, the final catalyst preferably has one or more of the following characteristics: (1) The absolute strength of the modified Y zeolite, as measured by X-ray diffraction (XRD), is preferably Is at least 65, more preferably at least 75; And (2) the skeleton 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 measured by XRD and the product of the% skeletal aluminum of the aluminum in the modified Y zeolite is greater than 4500.

While only certain embodiments have been described, alternatives and modifications will become apparent to those skilled in the art from the foregoing description. These and other alternatives are considered to be equivalent and are within the spirit and scope of the present disclosure and appended claims.

Claims (17)

As a catalyst comprising a modified LZ-210 zeolite, Said catalyst comprising 60 to 90 wt% zeolite based on volatiles free and the remainder comprising alumina as a binder and having an absolute intensity as measured by X-ray diffraction (XRD) of 50 or greater, said zeolite also having a bulk of the zeolite Wherein the catalyst has a skeletal aluminum content of at least 60% of the aluminum content. The catalyst of claim 1, wherein the product of the absolute strength and the skeletal aluminum content expressed as integral percentages is greater than 4,200. According to claim 1 or 2 wherein the zeolite is one of catalyst, based on the weight of zeolite in the anhydrous basis is less than 3% by weight Na ₂ O content. The catalyst according to claim 1 or 2, wherein the zeolite has a bulk Si / Al ₂ molar ratio ranging from 6.5 to 27. 3. The catalyst according to claim 1 or 2, wherein the zeolite has an absolute strength of 60 or more. The catalyst according to claim 1 or 2, wherein the zeolite has a skeletal aluminum of 70% or more. 3. The catalyst according to claim 1 or 2, wherein the ignition loss (LOI) at 900 DEG C is in the range of 2 to 4 wt%. 3. The catalyst according to claim 1 or 2, wherein the modified zeolite has a unit cell size of 24.58 ANGSTROM or less. (a) treating a first zeolite with an aqueous ionic solution to produce a second zeolite having a first unit cell size; (b) hydrothermally treating the second zeolite at a temperature in the range of 550 to 850 ° C to produce a third zeolite having a first bulk Si / Al ₂ molar ratio and a second unit cell size smaller than the first unit cell size ; And (c) The third zeolite, at least a part of the sodium cations was replaced with ammonium ion, the first bulk Sl / Al-modified LZ-210 zeolite having a large range from 6.5 to 27 second bulk Si / Al ₂ molar ratio of greater than ₂ molar ratio of Contacting the third zeolite with an aqueous solution of a sufficient amount of ammonium ions having a pH of less than 4, wherein the modified LZ-210 zeolite also comprises up to 60% of the total aluminum in the modified LZ-210 zeolite Having a skeletal aluminum content, and having an absolute intensity measured by X-ray diffraction of 50 or more Lt; RTI ID = 0.0 > LZ-210 < / RTI > zeolite catalyst. 10. The process of claim 9, wherein the ionic aqueous solution of step (a) comprises a fluorosilicate salt. 10. The method of claim 9, wherein (a) is that of the first to the sodium of the zeolite treatment, ammonium ion exchange of its sodium content by weight of the first zeolite in an anhydrous basis reduced to less than 3% by weight of Na ₂ O ≪ / RTI > 10. The method according to claim 9, wherein the aqueous solution of step (c) comprises ions selected from the group consisting of nitrate ions and sulfate ions. 10. 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, filtering the first mixture to recover the filter cake And contacting said filter cake with a second solution of ammonium ions having a pH of less than 4. < Desc / Clms Page number 13 > 14. The method according to any one of claims 11 to 13, wherein the hydrothermal treatment in step (b) comprises steam treatment. 14. The process according to any one of claims 11 to 13, wherein the modified LZ-210 zeolite after step (c) is contacted with a dehydrating agent having a water concentration of less than 30 ppm by weight at a temperature ranging from 25 to 500 < 0 > C. An aromatic transalkylation process comprising contacting an aromatic and an aromatic substrate with a catalyst of claim 1 or 2. A method of transalkylation of an aromatic comprising contacting an aromatic and aromatic substrate with a catalyst prepared according to the method of any of claims 9 to 13.

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