Process for the preparation of ethylbenzene by alkylation and transalkylation

Abstract: A process for the production of ethylbenzene by the alkylation of benzene over a molecular sieve aromatic alkylation catalyst followed by transalkylation of polyalkylated aromatic components and then liquid phase alkylation. A feedstock containing benzene and ethylene is supplied to a first alkylation reaction zone containing a molecular sieve aromatic alkylation catalyst. The reaction zone is operated at temperature and pressure conditions to cause gas phase ethylation of the benzene with the production of an alkylation product comprising a mixture of ethylbenzene and polyalkylated aromatic components including diethylbenzene with xylene present in only small amounts. The output from the alkylation reaction zone is supplied to an intermediate recovery zone for the separation and recovery of ethylbenzene, a polyalkylated aromatic compound component including diethylbenzene. At least a portion of the polyalkylated component is supplied along with benzene to a transalkylation reaction zone operated in the liquid phase. Disproportionation product from the transalkylation reaction zone is supplied to a second alkylation zone containing a molecular sieve aromatic alkylation catalyst. Ethylene is also supplied to the second alkylation zone which is operated under temperature and pressure conditions to cause liquid phase ethylation of the benzene. The alkylation product from the second alkylation zone is supplied to the aforementioned intermediate recovery zone.

Description: This invention involves an aromatic alkylation/transalkylation process involving vapor phase ethylation of benzene over a silicalite aromatic alkylation catalyst with separate transalkylation followed by liquid phase alkylation of the transalkylation zone output. production. BACKGROUND OF THE INVENTION Aromatic conversion processes which are carried out over molecular sieve catalyst are well known in the chemical processing industry. Such aromatic conversion reactions include the alkylation of aromatic substrates such as benzene to produce alkyl aromatics such as ethylbenzene, ethyltoluene, cumene or higher aromatics and the transalkylation of polyalkyl benzenes to monoalkyl benzenes. Typically, an alkylation reactor which produces a mixture of mono- and polyalkyl benzenes may be coupled through various separation stages to a downstream transalkylation reactor. Such alkylation and transalkylation conversion processes can be carried out in the liquid phase, in the vapor phase or under conditions in which both liquid and vapor phases are present. Alkylation and transalkylation reactions may occur simultaneously within a single reactor. For example, where various series-connected catalyst beds are employed in an alkylation reactor as described below, it is a conventional practice to employ interstage injection of the aromatic substrate between the catalyst beds, which tends to enhance transalkylation reactions within the alkylation reactor. For example, in the ethylation of benzene with ethylene to produce ethylbenzene, the alkylation product within the reactor includes not only ethylbenzene but also polyethylbenzene, principally diethylbenzene with reduced amounts of triethylbenzene, as well as other alkylated aromatics such as cumene and butylbenzene. The interstage injection of the ethylene results not only further in alkylation reactions but also transalkylation reactions where, for example, benzene and diethylbenzene undergo transalkylation to produce ethylbenzene. Thus, even though a separate transalkylation reactor is connected downstream through a series of separation stages, it is the accepted practice to minimize polyalkylation within the alkylation reactor in order to facilitate the subsequent treatment and separation steps. An example of vapor phase alkylation is found in U.S. Patent No. 4,107,224 to Dwyer. Here, vapor phase ethylation of benzene over a zeolite catalyst is accomplished in a down flow reactor having four series connected catalyst beds. The output from the reactor is passed to a separation system in which ethylbenzene product is recovered, with the recycle of polyethylbenzenes to the alkylation reactor where they undergo transalkylation reactions with benzene. The Dwyer catalysts are characterized in terms of those having a constraint index within the approximate range of 1-12 and include, with the constraint index in parenthesis, ZSM-5 (8.3), ZSM-11 (8.7), ZSM-12 (2), ZSM-35 (4.5), ZSM-38 (2), and similar materials. The molecular sieve silicalite is a well-known alkylation catalyst. For example, U.S. Patent No. 4,520,220 to Watson et al discloses the use of silicalite catalysts having an average crystal size of less than 8 microns and a silica/alumina ratio of at least about 200 in the ethylation of an aromatic substrate such as benzene or toluene to produce ethylbenzene or ethyltoluene, respectively. As disclosed in Watson et al, the alkylation procedure can be carried out in a multi-bed alkylation reactor at temperatures ranging from about 350°-500°C. and, more desirably, about 400°-475°C, with or without a steam co-feed. The reactor conditions in Watson et al are such as to provide generally for vapor phase alkylation conditions. Another procedure employing silicalite and involving the ethylation of benzene under vapor phase reaction conditions coupled with the recycle of polyethylbenzene containing products back to the alkylation reactor is disclosed in U.S. Patent No. 4,922,053 to Wagnespack. Here, alkylation is carried out at temperatures generally in the range of 370°C. to about 470°C. and pressures ranging from atmospheric up to about 25 atmospheres over a catalyst such as silicalite or ZSM-5. The catalysts are described as being moisture sensitive and care is taken to prevent the presence of moisture in the reaction zone. The alkylation/transalkylation reactor comprises four series connected catalyst beds. Benzene and ethylene are introduced into the top of the reactor to the first catalyst bed coupled by recycle of a polyethylbenzene fraction to the top of the first catalyst bed as well as the interstage injection of polyethylbenzene and benzene at different points in the reactor. Another process involving the use of a silicalite as an alkylation catalyst involves the alkylation of an alkylbenzene substrate in order to produce dialkylbenzene of a suppressed ortho isomer content. Thus, as disclosed in U.S. Patent No. 4,489,214 to Butler et al, silicalite is employed as a catalyst in the alkylation of a monoalkylated substrate, toluene or ethylbenzene, in order to produce the corresponding dialkylbenzene, such as ethyl toluene or diethylbenzene. Specifically disclosed in Butler et al is the ethylation of toluene to produce ethyltoluene under vapor phase conditions at temperatures ranging from 350°-500°C. As disclosed in Butler, the presence of ortho ethyltoluene in the reaction product is substantially less than the thermodynamic equilibrium amount at the vapor phase reaction conditions employed. U.S. Patent No. 4,185,040 to Ward et al discloses an alkylation process employing a molecular sieve catalyst of low sodium content which is said to be especially useful in the production of ethylbenzene from benzene and ethylene and cumene from benzene and propylene. The Na 2 O content of the zeolite should be less than 0.5 wt.%. Examples of suitable zeolites include molecular sieves of the X, Y, L, B, ZSM-5, and omega crystal types, with steam stabilized hydrogen Y zeolite being preferred. Specifically disclosed is a steam stabilized ammonium Y zeolite containing about 0.2% Na 2 O. Various catalyst shapes are disclosed in the Ward et al patent. While cylindrical extrudates may be employed, a particularly preferred catalyst shape is a so-called "trilobal" shape which is configured as something in the nature of a three leaf clover. The surface area/volume ratio of the extrudate should be within the range of 85-160 in. -1 . The alkylation process may be carried out with either upward or downward flow, the latter being preferred, and preferably under temperature and pressure conditions so that at least some liquid phase is present, at least until substantially all of the olefin alkylating agent is consumed. Ward et al states that rapid catalyst deactivation occurs under most alkylating conditions when no liquid phase is present. U.S. Patent No. 4,169,111 to Wight discloses an alkylation/transalkylation process for the manufacture of ethylbenzene employing crystalline aluminosilicates in the alkylation and transalkylation reactors. The catalysts in the alkylation and transalkylation reactors may be the same or different and include low sodium zeolites having silica/alumina mole ratios between 2 and 80, preferably between 4-12. Exemplary zeolites include molecular sieves of the X, Y, L, B, ZSM-5 and omega crystal types with steam stabilized Y zeolite containing about 0.2% Na 2 O being preferred. The alkylation reactor is operated in a downflow mode and under temperature and pressure conditions in which some liquid phase is present. The output from the alkylating reactor is cooled in a heat exchanger and supplied to benzene separation columns from which benzene is recovered overhead and recycled to the alkylation reactor. The initial higher boiling bottoms fraction from the benzene columns comprising ethylbenzene and polyethylbenzene is supplied to an initial ethylbenzene column from which the ethylbenzene is recovered as the process product. The bottoms product from the ethylbenzene column is supplied to a third column which is operated to provide a substantially pure diethylbenzene overheads fraction which contains from 10 to 90%, preferably 20 to 60% of diethylbenzene. The diethylbenzene overheads fraction is recycled to the alkylation reactor while a side cut containing the remaining diethylbenzene and triethylbenzene and higher molecular weight compounds is supplied to the reactor along with benzene. The effluent from the reactor is recycled through the heat exchanger to the benzene column. U.S. Patent No. 4,774,377 to Barger et al discloses an alkylation/transalkylation process which, involves the use of separate alkylation and transalkylation reaction zones, with recycle of the transalkylated product to an intermediate separation zone. In the Barger process, the temperature and pressure conditions are adjusted so that the alkylation and transalkylation reactions take place in essentially the liquid phase. The transalkylation catalyst is an aluminosilicate molecular sieve including X-type, Y-type, ultrastable-Y, L-type, omega type and mordenite type zeolites with the latter being preferred. The catalyst employed in the alkylation reaction zone is a solid phosphoric acid containing material. Aluminosilicate alkylation catalysts may also be employed and water varying from 0.01 to 6 volume percent is supplied to the alkylation reaction zone. The output from the alkylation reaction zone is supplied to first and second separation zones. Water is recovered in the first separation zone. In the second separation zone, intermediate aromatic products and trialkylaromatic and heavier products are separated to provide an input to the transalkylation reaction zone having only dialkyl aromatic components, or diethylbenzene in the case of an ethylbenzene manufacturing procedure or diisopropylbenzene in the case of cumene production. A benzene substrate is also supplied to the transalkylation zone for the transalkylation reaction and the output from the transalkylation zone is recycled to the first separation zone. The alkylation and transalkylation zones may be operated in downflow, upflow, or horizontal flow configurations. EPA publication 467,007 to Butler discloses other processes having separate alkylation and transalkylation zones employing various molecular sieve catalysts and with the output from the transalkylation reactor being recycled to an intermediate separation zone. Here, a benzene separation zone, from which an ethylbenzene/polyethylbenzene fraction is recovered from the bottom with recycling of the overhead benzene fraction to the alkylation reactor is preceded by a prefractionation zone. The prefractionation zone produces an overhead benzene fraction which is recycled along with the overheads from the benzene column and a bottom fraction which comprises benzene, ethylbenzene and polyethylbenzene. Two subsequent separation zones are interposed between the benzene separation zone and the transalkylation reactor to provide for recovery of ethylbenzene as the process product and a heavier residue fraction. The polyethylbenzene fraction from the last separation zone is applied to the transalkylation reactor and the output there is applied directly to the second benzene separation column or indirectly through a separator and then to the second benzene separation column. Butler discloses that the alkylation reactor may be operated in the liquid phase with a catalyst such as zeolite-β, zeolite-Y or zeolite-Ω or in the vapor phase employing a catalyst such as silicalite or ZSM-5. In the Butler process, where vapor phase alkylation is followed by liquid phase transalkylation, substantial quantities of water may be included in the feedstream to the alkylation reactor. In this case, the feed to the transalkylation reactor may be dehydrated to lower the water content. The transalkylation catalyst may take the form of a zeolite-Y or zeolite-Ω. WO 99/26904 discloses a process of producing alkylaromatics wherein the feedstock is fed to a transalkylation zone and the effluent from the transalkylation zone is then fed into an alkylation zone along with an olefin alkylating agent. EP 1002778 A1 discloses a process of gas phase alkylation comprising a multi-stage alkylation reaction zone having a plurality of series-connected catalyst beds. Monoclinic silicalite catalysts of differing silica/alumina ratios are utilized. SUMMARY OF THE INVENTION In accordance with the present invention there is provided a process for the production of ethylbenzene by the alkylation of benzene over a molecular sieve aromatic alkylation catalyst, preferably monoclinic silicalite having a silica/alumina ratio of at least 275, followed by transalkylation of polyalkylated aromatic components, followed by liquid phase alkylation. In carrying out the invention, a feedstock containing benzene and ethylene is supplied to a first alkylation reaction zone containing a molecular sieve aromatic alkylation catalyst. The reaction zone is operated at temperature and pressure conditions to maintain the benzene in the gaseous phase and to cause gas phase ethylation of the benzene with the production of an alkylation product comprising a mixture of ethylbenzene and polyalkylated aromatic components including diethylbenzene with xylene present in only small amounts. The output from the gas phase alkylation reaction zone is supplied to an intermediate recovery zone for the separation and recovery of ethylbenzene and the separation and recovery of a polyalkylated aromatic compound component including diethylbenzene. At least a portion of the polyalkylated component is supplied along with benzene to a transalkylation reaction zone. The transalkylation reaction zone is operated, preferably in the liquid phase, to cause disproportionation of the polyalkylated aromatic fraction to produce a disproportionation product containing unreacted benzene and having a reduced diethylbenzene content and an enhanced ethylbenzene content. Disproportionation product from the transalkylation reaction zone is supplied to a second alkylation zone containing a molecular sieve aromatic alkylation catalyst, preferably a molecular sieve having a pore size greater than the molecular sieve in the first alkylation reaction zone. Ethylene is also supplied to the second alkylation zone which is operated under temperature and pressure conditions in which the unreacted benzene is maintained in the liquid phase to cause liquid phase ethylation of the benzene. The alkylation product from the second alkylation zone is supplied to the aforementioned intermediate recovery zone. Preferably, the first alkylation reaction zone is operated under conditions in which the concentration of xylene, and the product recovered therefrom, is no more than about 0.1 wt.% based upon ethylbenzene in the product. The output from the second liquid phase alkylation zone has an even lower xylene content. In yet a further aspect of the invention, the intermediate separation zone includes a benzene recovery zone which is operated in first and second stages with the recycle of benzene to the first alkylation zone. In yet a further aspect of the invention, the molecular sieve aromatic alkylation catalyst in the second liquid phase alkylation zone is zeolite Y or zeolite beta. Preferably, the transalkylation zone contains zeolite Y transalkylation catalyst and is operated under liquid phase conditions, and the second alkylation zone contains zeolite beta or zeolite lanthanum-beta (La-Beta). BRIEF DESCRIPTION OF THE DRAWINGS Fig. 1 is an idealized schematic block diagram of an alkylation/transalkylation alkylation process embodying the present invention. Fig. 2 is a schematic illustration of a preferred embodiment of the invention incorporating separate parallel-connected alkylation and transalkylation reactors with an intermediate multi-stage recovery zone for the separation and recycling of components. Fig. 2A is a schematic illustration of an alkylation zone comprising a plurality of series-connected catalyst beds with the interstage injection of feed components. Fig. 3 is a graphical presentation showing the results of the experimental work and illustrating the xylene content relative to ethylbenzene for various silicalite catalysts at different space velocities as a function of time. Fig. 4 is a graphical presentation to indicate the amount of diethylbenzene relative to ethylbenzene for various catalysts and space velocities as a function of time. Fig. 5 is a graphical presentation showing the relative amount of propylbenzene relative to ethylbenzene as a function of time for the various catalysts and space velocities. Fig. 6 is a graphical presentation showing the relative amount of butylbenzene relative to ethylbenzene for the different catalysts and space velocities as a function of time. Fig. 7 is a graphical presentation of the amount of heavies content relative to ethylbenzene as a function of time for the various catalysts and space velocities. Fig. 8 is an illustration of data points illustrating the effect of space velocity on ethylbenzene production plotted as function of time for the various catalysts at different space velocities. Fig. 9 is a graphical presentation showing the results of experimental work illustrating the xylene content relative to ethylbenzene for a high silica/alumina ratio silicalite catalyst at different benzene/ethylene ratios embodying the present invention at various ethylene/benzene ratios and in comparison with a similar silicalite catalyst of lower silica/alumina ratio. Fig. 10 is a graphical presentation showing the results of experimental work illustrating a heavies content relative to ethylbenzene for a silicalite catalyst at different benzene/ethylene ratios embodying the present invention at various ethylene/benzene ratios and in comparison with a similar silicalite catalyst of lower silica/alumina ratio. Fig. 11 is a graphical presentation showing the results of experimental work illustrating a cumene content relative to ethylbenzene for a silicalite catalyst at different benzene/ethylene ratios embodying the present invention at various ethylene/benzene ratios and in comparison with a similar silicalite catalyst of lower silica/alumina ratio. Fig. 12 is a graphical presentation showing the results of experimental work illustrating a propylbenzene content relative to ethylbenzene for a silicalite catalyst at different benzene/ethylene ratios embodying the present invention at various ethylene/benzene ratios and in comparison with a similar silicalite catalyst of lower silica/alumina ratio. Fig. 13 is a graphical presentation showing the results of experimental work illustrating a butylbenzene content relative to ethylbenzene for a silicalite catalyst at different benzene/ethylene ratios embodying the present invention at various ethylene/benzene ratios and in comparison with a similar silicalite catalyst of lower silica/alumina ratio. Fig. 14 is a graphical presentation showing the results of experimental work illustrating a diethylbenzene content relative to ethylbenzene for a silicalite catalyst at different benzene/ethylene ratios embodying the present invention at various ethylene/benzene ratios and in comparison with a similar silicalite catalyst of lower silica/alumina ratio. Fig. 15 is a graphical presentation illustrating the impurities content of certain components relative to ethylbenzene as a function of ethylbenzene production as indicated by ethylene injection. DETAILED DESCRIPTION OF THE INVENTION The present invention involves the vapor-phase alkylation of benzene over a silicalite alkylation catalyst followed by recovery of ethylbenzene with transalkylation of a residual product with subsequent liquid phase alkylation of the transalkylated product. The preferred silicalite catalyst provides for extremely low activity for the production of undesirable polyalkylated aromatics, specifically xylene and high boiling point alkyl aromatic components boiling at about 185°C and above. In the production of ethylbenzene, the benzene substrate and the ethylene alkylating agent are injected in relative amounts so that the benzene is well in excess of the stoichiometric equivalent amount for the alkylation reaction. Under these circumstances, the ethylene constituent of the feedstream is the stoichiometric controlling reactant. That is, the ethylbenzene production rate can be increased by increasing the ethylene injection rate with an attendant reduction in the benzene/ethylene mole ratio or decreased by decreasing the ethylene injection rate with an attendant increase in the benzene/ethylene mole ratio. The foregoing assumes, of course, that the benzene feed rate remains constant. A difficulty encountered in increasing the ethylene flow rate, thus decreasing the benzene/ethylene ratio below a designated design parameter, is that the production of unwanted side products is often increased. A preferred embodiment of the present invention involves the use of an alkylation catalyst comprising a predominantly monoclinic silicalite of a high silica/alumina ratio, preferably one having a small crystal size, which enables the ethylene content of the feedstream to be materially increased with only a relatively small or even no increase in xylenes content and high boiling point polyalkylated aromatic components. Specifically, the catalyst can be characterized in terms of a xylene response to a 25% increase in ethylene rate (a decrease in the benzene/ethylene ratio of 20%) such that the increase in xylene production is no more than 10% relative to the production of ethylbenzene. Similarly, for the heavy polyalkylated aromatic components boiling at 185°C or above, the corresponding increase under these conditions is no more than 5% relative to the production of ethylbenzene. The low xylene content in the effluent from the alkylation reactor is particularly significant in terms of the downstream separation of ethylbenzene since the xylenes, particularly the meta- and para-isomers have boiling points very close to the boiling point of ethylbenzene. Similarly, a low heavies content is desirable in terms of downstream separation procedures carried out prior to the supply of diethylbenzene, which has a maximum boiling point of about 185°C, to a downstream transalkylation reactor. The practice of the present invention will normally follow the accepted practice of vapor-phase alkylation of an aromatic substrate in a first alkylation reaction zone followed by a separate transalkylation reaction. However, rather than recycling of the transalkylation output to an intermediate separation zone for the recovery of ethylbenzene, this product is supplied to a second alkylation reactor which is operated in the liquid phase. A preferred embodiment of the present invention involves the use of a high silica/alumina ratio silicalite alkylation catalyst in a mode that actually increases the diethylbenzene output from the alkylation reaction during a portion of a cycle of operation in which one of a plurality of parallel reactors is placed in a regeneration mode. The alkylation reactor can be operated under a relatively high space velocity such that the alkylation over the silicalite catalyst is carried out to provide a diethylbenzene content substantially above what is achieved under normal operating conditions. Specifically, the diethylbenzene content is increased by incremental value of about 0.2 or more of the diethylbenzene content produced at a space velocity of one-half of the enhanced space velocity. This enhanced space velocity occurs during a relatively short period of time after which a reduced space velocity is encountered in which the diethylbenzene content is reduced during normal operating conditions to a value near the thermodynamic equilibrium value. The enhanced diethylbenzene production is offset by an accompanying selectivity toward the production of ethylbenzene relative to the production of xylenes as a by-product. Stated otherwise, the xylene content in the product is diminished preferably to a value of less than 0.06 wt.% based upon the ethylbenzene in the product. Further, the ortho xylene content is contained at a relatively low level, less than the thermodynamic equilibrium level of ortho xylene at temperature and pressure conditions of the alkylation reactor zone. Specifically, the ortho xylene content can be diminished to a value of about one-half or less than the equilibrium value. In this respect the equilibrium ratio of the three isomers of xylene at a desired alkylation temperature of about 400°C. is 24.3% ortho xylene, 52.3% meta xylene, and 23.4% para xylene. The practice of the present invention can result in a ortho xylene content in the reaction product of no more than about 10 wt.% of the total xylene content of the reaction product. The preferred silicalite employed in the present invention, in addition to having a relative high silica aluminum ratio, has a smaller crystal size than the silicalite traditionally employed in aromatic alkylation procedures. Silicalite, as is well known in the art, is a molecular sieve catalyst which is similar to the ZSM-5 zeolites but is typically characterized by a higher silica/alumina ratio providing an aluminum/unit cell ratio of less than 1, and, in addition, is normally characterized as having a somewhat larger than average crystal size than is commonly associated with the ZSM zeolites. As is well known in the art, silicalite, which in the as synthesized form is characterized by orthorhombic symmetry, can be converted to monoclinic symmetry by a calcination procedure as disclosed, for example, in U.S. Patent No. 4,599,473 to DeBras et al. As described in detail in DeBras et al, "Physico-chemical characterization of pentasil type materials, I. Precursors and calcined zeolites, and II. Thermal analysis of the precursors," Zeolites , 1985, Vol. 5, pp. 369-383, the silicalite typically has a relatively large crystal size. Thus, at an average of less than one aluminum atom per unit cell (a silica/alumina ratio of about 200) silicalite typically has an average crystal size of perhaps 5-10 microns or more. The aforementioned Patent No. 4,489,214 to Butler et al discloses experimental work involving the ethylation of toluene over silicalite of a crystal size greater than one micron, ranging from 1-2 microns up to 8 microns. The silicalite is further characterized in terms of a variable aluminum gradient such that the aluminum gradient is positive when going from the interior to the surface of the molecular sieve crystal. That is, the silicalite can be characterized by a core portion which is relatively aluminum deficient with an outer shell portion which is relatively aluminum rich. It is to be understood that the term "aluminum rich" is a relative term and that for silicalite even the outer shell portion of the crystallite has a low aluminum content. The preferred embodiment of the present invention involves vapor phase ethylation of benzene in a multistage reaction zone containing high silica/alumina ratio silicalite followed by liquid phase transalkylation followed in turn by liquid phase alkylation of the disproportion product recovered from the transalkylation reactor in which the alkylation and transalkylation reactors are integrated with an intermediate recovery zone. This zone preferably involves a plurality of separation zones operated in a manner to effectively provide feed streams to the reactors with recycle of the output from the second liquid phase alkylation reactor to a benzene recovery zone downstream of the first vapor phase alkylation reactor. In this integrated mode of operation, the transalkylation/liquid phase alkylation product is applied to one or both stages of a multistage benzene recovery zone. Subsequent separation steps can be carried out in a manner to apply a split feed to the transalkylation reactor. The alkylation reactor is a multistage reaction zone containing at least three series connected catalyst beds which contain the silicalite alkylation catalyst. More preferably, four or more beds are employed. As described in greater detail below, the silicalite alkylation catalyst preferably is silicalite characterized as having a high monoclinicity and a small sodium content, both in terms of sodium in the crystalline molecular sieve structure and in the binder component. The preferred catalyst used in the transalkylation reactor is a molecular sieve having a pore size greater than the pore size of the silicalite catalyst. Preferably, the transalkylation catalyst is zeolite Y. The preferred catalyst in the second liquid phase alkylation reactor is preferably zeolite Y or zeolite beta, with zeolite beta or zeolite La-Beta being particularly preferred. The aforementioned zeolites are in themselves well known in the art. For example, zeolite Y is disclosed in the aforementioned Patent No. 4,185,040 to Ward, and zeolite beta is disclosed in U.S. Patent Nos. 3,308,069 to Wadlinger and 4,642,226 to Calvert et al. Zeolite lanthanum beta (LA-BETA) is disclosed in EPO Publication No. 507,761 B1 to Shamshoum et al and may be prepared from zeolite beta by ion exchange to incorporate lanthanum into a precursor zeolite beta molecular sieve structure. For example, La-Beta can be prepared by ion exchange of anhydrous ammonium beta zeolite with an aqueous solution of lanthanum nitrate followed by mulling with nitrite acid-treated alumina, extrusion, and calcination. For a further description of zeolite Y, zeolite beta, and zeolite La-Beta, reference is made to the aforementioned Patent Nos. 4,185,040 to Ward, 3308,069 to Wadlinger, and 4,642,226 to Calvert et al and to EPO 507,761 B1 to Shamshoum et al, the entire disclosures of which are incorporated herein by reference. A preferred application of the invention is in a system involving a multistage alkylation reactor with the output coupled to a four-stage separation system which in turn supplies a polyethylbenzene feed to a transalkylation reactor. In the embodiment of the invention described herein, parallel alkylation and transalkylation reactors are employed. This results in a preferred mode of operation in which the parallel alkylation reactors are simultaneously operated in an alkylation mode while periodically one reactor can be taken off-stream with the feedstream completely supplied to the on-stream reactor. In the embodiment illustrated and described below, two parallel reactors are employed although it is to be recognized that three or more reactors can likewise be employed in parallel. A similar configuration is employed for the transalkylation reactors and for the liquid phase alkylation reactors. The result is that simultaneous catalyst regeneration can occur in one reactor during operation of the remaining alkylation and/or transalkylation reactors. Assuming that two parallel reactors are employed, it can be seen that this mode of operation will, for the same flow rate of feedstream, result in the operation of the reactors at two different space velocities, with the space velocity during regeneration of a reactor being about twice that with both parallel reactors in operation. Preferably, the first alkylation reactor comprises at least four catalyst beds as described above. More beds can be provided, and it will sometimes be advantageous to provide at least five catalyst beds in the alkylation reactor. The reactor is operated so as to provide vapor phase alkylation (both the aromatic substrate and the alkylating agent are in the vapor phase) at temperatures ranging from about 332°C - 427°C ( 630°F.- 800°F ) . at the inlet to about 332°C - 454°C ( 700 °F.-850°F ) . at the outlet. The pressure may be within the range of about 17 to 31 bar ( 250 to 450 psia ) with the pressure decreasing from one bed to the next as the temperature increases. By way of example, the benzene and ethylene supplied to the top of the reactor may enter the reactor at a temperature of about 393°C ( 740°F ) and a pressure of [about] 29.6 bar ( 430 psia ) . The alkylation reaction is exothermic so that the temperature progressively increases from the first to the last catalyst bed by a way of example. The interstage temperatures may increase from 399°C ( 750°F ) for the first catalyst bed to 407°C ( 765°F ) after the second catalyst bed to 820°F after the third catalyst bed to a temperature of about 449°C ( 840°F ) after the last catalyst bed. in tlle Normally in the operation of multi-stage reaction zone of the type involved in the present invention, a benzene-ethylene mixture is introduced to the first catalyst bed at the top of the reaction zone and also in between the several successive stages of catalyst beds. In the present invention, ethylene is supplied along with benzene to the top of the first catalyst bed top at the upper end of the reactor. In addition, interstage injection of ethylene and benzene is provided for between the subsequent catalyst beds. The benzene to ethylene mole ratio is about 18 as injected into the top of the alkylation reactor and progressively decreases because of the interstage injection of ethylene and coupled with the alkylation of the benzene to ethylbenzene and polyethylbenzenes. The silicalite alkylation catalyst employed in the present invention does not require the presence of water to stabilize the catalyst, so a water or steam co-feed, as is sometimes used in connection with silicalite, is not called for in this invention. As noted above, interstage injection of ethylene is normally employed, and the interstage injection of benzene can also be provided for. The mole ratio of the benzene to the ethylene at the interstage injection points can vary from zero (no benzene injection) up to about five. The benzene in many cases will be employed in an amount less than the amount of ethylene on a mole basis. Stated otherwise, benzene can either not be injected between the catalyst beds or, if injected, can be employed in a relatively minor amount, i . e ., a mole ratio of benzene to ethylene of less than one. On the other hand, the benzene/ethylene mole ratio can be as high as five. This is coupled with a somewhat lower operating temperature than would normally be the case for vapor phase alkylation. In the preferred embodiment of the invention, the temperature of the benzene stream into the top of the alkylation reactor will be in the order of 382°C ( 720°F ) . or lower. The alkylation reaction is, of course, an exothermic reaction so that the temperature will be increased progressively throughout the alkylation column as noted previously. The silicalite alkylation catalyst employed in the present invention is a molecular sieve from the pentasil family of high silica molecular sieves. Such pentasil molecular sieves are described, for example, in Kokotailo et al, "Pentasil Family of High Silica Crystalline Materials," Chem. Soc. Special Publ. 33, 133-139 (1980). The silicalite molecular sieve alkylation catalyst has a somewhat smaller pore size than the preferred zeolite-Y employed in the transalkylation reactor and the preferred zeolite beta used in the second alkylation reactor. The silicalite catalyst has an effective pore size or window within the range of 5-6 angstroms. Zeolite Y has a pore size of about 7 angstroms. Zeolite beta has an elliptical pore having a major axis of about 7.5 angstroms and a minor axis of about 5.5 angstroms. Zeolite La-Beta has a pore size geometry similar to zeolite beta. The preferred silicalite catalyst has a somewhat smaller crystal size, less than one micron, than is usually the case. Preferably, the crystal size is even somewhat smaller, providing an average crystal size of 0.5µ or less (as contrasted with a crystal size of perhaps 1-2µ up to about 8 microns for similar catalysts such as disclosed in the aforementioned Patent No. 4,489,214 to Butler et al.) and is preferably formulated with an alumina binder to provide catalyst particules having a surface area/volume ratio of at least 152cm -1 (60 in -1 ). A preferred silicalite for use in the present invention is extruded with an alumina binder in a "trilobe" shape having a nominal diameter of about 0.159 cm ( 1/16" ) and a length of the extrudate of about 0.317 - 0.635 cm ( 1/8-1/4." ) As discussed below, the silicalite catalyst has a low sodium content and this is complemented in the preferred embodiment of the invention by the use of an alumina binder which is of unusually high purity and unusually large pore size as described in greater detail below. The "trilobe" cross sectional shape is something on the order of a three leaf clover. The purpose of this shape is to increase the geometric surface area of the extruded catalyst beyond what one would expect with a normal cylindrical extrudate. The silicalite catalyst is characterized as monoclinic silicalite. Monoclinic silicalite may be prepared as disclosed in U.S. Patent Nos. 4,781,906 to Cahen et al and 4,772,456 to DeClippeleir et al. Preferably the catalysts will have near 100% monoclinicity) although silicalite catalysts that are 70-80% monoclinic and about 20-30% orthorhombic symmetry may be used in the preferred embodiment of the invention. The silicalite preferably is present in an amount of 75-80 wt.% with the alumina binder being present in an amount of 20-25 wt.%. The silica/alumina ratio of the silicalite is at least 300. An especially preferred silica/alumina ratio is 300-350, and silicalite within this range was used in experimental work respecting the invention as described hereafter. The silicalite may have an alpha value of about 20-30. The "alpha value" is characterized in terms of the activity of a catalyst for cracking hexane as disclosed in U.S. Patent Nos. 4,284,529 to Shihabi and 4,559,314 to Shihabi. The silicalite catalyst typically contains small amounts of sodium and iron. As noted previously, the silicalite alkylation catalyst has a crystal structure characterized by an aluminum rich outer shell and an aluminum deficient interior portion when compared with the outer shell. The silicalite catalyst is dry and has no appreciable or intended water content. Specifically, the silicalite catalyst contains no more than about 200 ppm sodium, preferably no more than 100 ppm sodium and no more than about 500 ppm iron, preferably no more than 300 ppm iron. The alumna binder is a high purity alumina such as "catapal alumina." Preferably, the alumina binder is characterized in terms of an unusually high pore size and unusually low sodium content. As noted previously, the silicalite itself has a low sodium content in its crystalline structure. By maintaining a low sodium content in the alumina, a high portion of the catalyst sites in the silicalite structure are maintained in the active hydrogen form -- that is, the low sodium content of the binder tends to minimize neutralization of the crystalline catalyst sites due to ion exchange between sodium in the binder and the acid sites in the catalyst. The alumina binder is further characterized in terms of a relatively large pore size after the catalyst is extruded and divided into particles. Specifically, the pore size of the binder, which can be termed the "maximum" pore size to avoid confusion with the pore size of the silicalite itself, is about 1,000 angstroms or more. A preferred pore size range is within the range of about 1,000 to about 1,800 angstroms. This relatively large pore size binder can enhance the efficiency of the catalyst by avoiding, or at least minimizing, an alumina-diffusing mechanism as applied to the catalyst particles themselves, thus enhancing access to the silicalite molecular sieve within the catalyst particles. The pore size of the molecular sieve structure itself normally can be expected to be on the order of about 5-6 angstroms. The silicalite catalyst preferably contains only a small amount of sodium, about 70-200 ppm sodium and contains only a small amount of iron, about 200-500 ppm. The catalyst need not contain any additional "promoter" metals incorporated during the synthesis of the catalyst. Turning now to the drawings and referring first to Fig. 1, there is illustrated a schematic block diagram of an alkylation/transalkylation process carried out in accordance with the present invention. As shown in Fig. 1, a product stream comprising a mixture of ethylene and benzene in a mole ratio of benzene to ethylene about 10 to 20 is supplied via line 1 through a heat exchanger 2 to an alkylation zone 4. Alkylation zone 4 comprises one or more multi-stage reactors having a plurality of series-connected catalyst beds containing the preferred high silica/alumina ratio silicalite as described in greater detail below. The alkylation zone 4 is operated at temperature and pressure conditions to maintain the alkylation reaction in the vapor phase, i . e . the aromatic substrate is in the vapor phase, and at a feed rate to provide a space velocity enhancing diethylbenzene production while retarding xylene production. The output from the first alkylation reactor 4 is supplied via line 5 to a intermediate benzene separation zone 6 which may take the form of one or more distillation columns. Benzene is recovered through line 8 and recycled to line 1 to the input of the first alkylation reactor. The bottoms fraction from the benzene separation zone 6, which includes ethylbenzene and polyalkylated benzenes including polyethylbenzene and xylene, is supplied via line 9 to an ethylbenzene separation zone 10. The ethylbenzene separation zone may likewise comprise one or more sequentially-connected distillation columns. The ethylbenzene is recovered through line 12 and applied for any suitable purpose, such as in the production of vinylbenzene. The bottoms fraction from the ethylbenzene separation zone 10 which comprises polyethylbenzene, principally diethylbenzene, is supplied via line 14 to a transalkylation reactor 16. Benzene is supplied to the transalkylation reaction zone through line 18. The transalkylation reactor, which preferably is operated under liquid phase conditions, contains a molecular sieve catalyst, preferably zeolite-Y, which has a somewhat larger pore size than the silicalite catalyst used in the first alkylation reaction zone. The output from the transalkylation reaction zone is supplied via line 20 to a second alkylation reaction zone along with ethylene supplied through line 23. The second alkylation zone 22 is operated under temperature and pressure conditions in which the unreacted benzene is maintained in the liquid phase to cause liquid-phase ethylation of the benzene. The output from the second liquid-phase alkylation reactor is recycled via line 24 to the benzene separation zone 6. In this mode of operation, which is discussed in greater detail below, the concentration of ethylbenzene, which is recovered from the bottom of the transalkylation reaction, can be enhanced from a value of about 18-20% ethylbenzene to about 30-40% ethylbenzene in the liquid-phase alkylation reactor. The liquid-phase alkylation reactor, because of the moderate liquid-phase conditions, will produce a lower xylene content than the xylene content produced at the output of the first stage gas-phase alkylation reactor. Thus, by recycling the output from the liquid-phase alkylation reactor, the overall xylene content in the product stream is reduced. Referring now to Fig. 2, there is illustrated in greater detail a suitable system incorporating a multi-stage intermediate recovery zone for the separation and recycling of components involved in the alkylation/transalkylation/alkylation process. As shown in Fig. 2, an input feed stream is supplied by fresh ethylene through line 31 and fresh benzene through line 32. Line 32 is provided with a preheater 34 to heat the benzene stream consisting of fresh and recycled benzene to the desired temperature for the alkylation reaction. The feedstream is applied through a two-way, three-position valve 36 and inlet line 30 to the top of one or both parallel alkylation reaction zones 38 and 40 comprising a plurality of series connected catalyst beds each of which contains a silicalite alkylation catalyst. The reactors are operated at an average temperature, preferably within the range of 332°C-427°C (700°F-800°F) and at pressure conditions of about 14-24 bar (200 to 350 psia), to maintain the benzene in the gaseous phase. In normal operation of the system depicted in Fig. 2, both reaction zones 38 and 40 will, during most of a cycle of operation, be operated in a parallel mode of operation in which they are both in service at the same time. In this case, valve 36 is configured so that the input stream in line 30 is roughly split in two to provide flow to both reactors in approximately equal amounts. Periodically, one reactor can be taken off-stream for regeneration of the catalyst. Valve 36 is configured so that all of the feedstream from line 30 can be supplied to reactor 38 while the catalyst beds in reactor 40 are regenerated and visa versa. The regeneration procedure will be described in detail below but normally will take place over a relatively short period of time relative to the operation of the reactor in parallel alkylation mode. When regeneration of the catalyst beds in reactor 40 is completed, this catalyst can then be returned on-stream, and at an appropriate point, the reactor 38 can be taken off-stream for regeneration. This mode of operation in operation of the individual catalyst beds at relatively lower space velocities for prolonged periods of time with periodic relatively short periods of operation at enhanced, relatively higher space velocities when one reactor is taken off-stream. By way of example, during normal operation of the system with both reactors 38 and 40 on-stream, the feedstream is supplied to each reactor to provide a space velocity of about 35 hr. -1 LHSV. When reactor 40 is taken off-stream and the feed rate continues unabated, the space velocity for reactor 38 will approximately double to 70 hr. -1 LHSV. When the regeneration of reactor 40 is completed, it is placed back on-stream, and again the feed stream rate space velocity for each reactor will decrease to 35 hr. -1 until such point as reactor 38 is taken off-stream, in which the case the flow rate to reactor 20 will, of course, increase, resulting again in a transient space velocity in reactor 20 of 70 hr -1 LHSV. A preferred reactor configuration is shown in detail in Fig. 2A. As illustrated there, the reactor 38 comprises four series connected catalyst beds designated as beds A, B, C and D. An ethylene feed stream is supplied via line 39 and proportionating valves 39a, 39b and 39c to provide for the appropriate interstage injection of ethylene. Benzene can also be introduced between the catalyst stages by means of secondary benzene supply lines 41a, 41b and 41c, respectively. As will be recognized, the parallel reactor 40 will be configured with similar manifolding as shown in Fig. 2A with respect to reactor 38. Returning to Fig. 2, the effluent stream from one or both of the alkylation reactors 38 and 40 is supplied through a two-way, three-position outlet valve 44 and outlet line 45 to a two-stage benzene recovery zone which comprises as the first stage a prefractionation column 47. Column 47 is operated to provide a light overhead fraction including benzene which is supplied via line 48 to the input side of heater 34 where it is mixed with benzene in line 32 and then to the alkylation reactor input line 30. A heavier liquid fraction containing benzene, ethylbenzene and polyethylbenzene is supplied via line 50 to the second stage 52 of the benzene separation zone. Stages 47 and 52 may take the form of distillation columns of any suitable type, typically, columns having from about 20-60 trays. The overheads fraction from column 52 contains the remaining benzene which is recycled via line 54 to the alkylation reactor input. Thus, line 54 corresponds to the output line 8 of Fig. 1. The heavier bottoms fraction from column 52 is supplied via line 56 to a secondary separation zone 58 for the recovery of ethylbenzene. The overheads fraction from column 58 comprises relatively pure ethylbenzene which is supplied to storage or to any suitable product destination by way of line 60, corresponding generally to output line 4a of Fig. 1. By way of example, the ethylbenzene may be used as a feedstream to a styrene plant in which styrene is produced by the dehydrogenation of ethylbenzene. The bottoms fraction containing polyethylbenzenes, heavier aromatics such as cumene and butylbenzene, and normally only a small amount of ethylbenzene is supplied through line 61 to a tertiary polyethylbenzene separation zone 62. As described below, line 61 is provided with a proportioning valve 63 which can be used to divert a portion of the bottoms fraction directly to the transalkylation reactor. The bottoms fraction of column 62 comprises a residue which can be withdrawn from the process via line 64 for further use in any suitable manner. The overhead fraction from column 62 comprises a polyalkylated aromatic component containing diethylbenzene and triethylbenzene (usually in relatively small quantities) and a minor amount of ethylbenzene is supplied to an on stream transalkylation reaction zone. Similarly as described above with respect to the alkylation reactors, parallel transalkylation reactors 65 and 66 are provided through inlet and outlet manifolding involving valves 67 and 68. Both of reactors 65 and 66 can be placed on stream at the same time so that both are in service in a parallel mode of operation. Alternatively, only one transalkylation reactor can be on-stream with the other undergoing regeneration operation in order to burn coke off the catalyst beds by minimizing the amount of ethylbenzene recovered from the bottom of column 58, the ethylbenzene content of the transalkylation feedstream can be kept small in order to drive the transalkylation reaction in the direction of ethylbenzene production. The polyethylbenzene fraction withdrawn overhead from column 62 is supplied through line 69 and mixed with benzene supplied via line 70. This mixture is then supplied to the on-line transalkylation reactor 65 via line 71. Preferably, the benzene feed supplied via line 70 is of relatively low water content, about 0.05 wt.% or less. Preferably, the water content is reduced to a level of about 0.02 wt.% or less and more preferably to no more than 0.01 wt.%. The transalkylation reactor is operated as described before in order to maintain the benzene and alkylated benzenes within the transalkylation reactor in the liquid phase. Typically, the alkylation reactor and the transalkylation reactor may be operated to provide an average temperature within the transalkylation reactor of about 65.5°C - 288°C ( 150°F.-550°F). and an average pressure of about 41 bar ( 600 psi ) . The preferred catalyst employed in the transalkylation reactor is zeolite Y having the characteristics described previously. The weight ratio of benzene to polyethylbenzene should be at least 1:1 and preferably is within the range of 1:1 to 4:1. The output from the transalkylation reactor containing benzene, ethylbenzene, and diminished amounts of polyethylbenzene is supplied through line 72 to the input manifolding stage of a liquid phase alkylation zone. Similarly, as described previously with respect to the first stage alkylation reactor and transalkylation reactor, the liquid-phase transalkylation zone preferably comprises parallel alkylation reactors 76 and 77. The product stream 72 is combined with fresh benzene supplied through line 78 and supplied through valve 80 to one or both of the parallel liquid-phase transalkylation reactors 76 and 79. The liquid-phase alkylation reactors are typically operated under temperature conditions of about 204 - 260°C ( 400°-500°F ) with an average pressure of about 34 - 52 bar ( 500-750 psi ) . The liquid-phase alkylation reactor preferably contains a molecular sieve aromatic alkylation catalyst which has an average pore size greater than the average pore size than the silicalite molecular sieve employed in the first alkylation reactor. The zeolite in the liquid-phase alkylation reactor preferably is zeolite Y and/or zeolite beta. Preferably, the catalyst used in the transalkylation reactor is zeolite-Y as described previously, and the catalyst employed in the second liquid-phase alkylation reactor is zeolite beta or zeolite La-Beta. The output from the liquid-phase alkylation reactors is applied through output manifolding including a two-way, three-position outlet valve 82 to line 84 where it is recycled to the benzene separation zone. Outlet line 84 is connected through valves 85 and 86 to the inlet lines of prefractionation column 47 and the second stage column 52. The effluent from the liquid-phase transalkylation reactor may be supplied as indicated to either or both of distillation columns 47 and 52. Preferably, the system will be operated in which all or most of the liquid phase alkylation product is supplied to column 47. Returning to the operation of the separation system, in one mode of operation the entire bottoms fraction from the ethylbenzene separation column 58 is applied to the tertiary separation column 52 with overhead fractions from this zone then applied to the transalkylation reactor. This mode of operation offers the advantage of relatively long cycle lengths of the catalyst in the transalkylation reactor between regeneration of the catalyst to increase the catalyst activity. Another mode of operation of the invention achieves this advantage by supplying a portion of the output from the ethylbenzene separation column 58 through valve 63 directly to the transalkylation reactor. By employing vapor phase alkylation in combination with liquid phase transalkylation and subsequent liquid phase alkylation in accordance with the present invention, a significant quantity of the bottoms fraction from the ethylbenzene column can be sent directly to the transalkylation reactor, thus decreasing the amount of residue which is lost from the process. This mode of operation is consistent with and particularly advantageous in combination with the operation of the alkylation reactor to retard transalkylation and enhance ethylbenzene production. While applicants' invention is not to be limited by theory, it is believed that direct application of a substantial portion of the output from the ethylbenzene separation zone to the transalkylation reactor is made possible, at least in part, by the low water content in the process stream resulting from low water content introduced initially into the transalkylation reactor. As shown in Fig. 2, a portion of the bottoms fraction from the secondary separation zone 58 bypasses column 62 and is supplied directly to the transalkylation reactor 65 via valve 63 and line 88. A second portion of the bottoms fraction from the ethylbenzene column is applied to the tertiary separation column 62 through valve 63 and line 90. The overhead fraction from column 62 is commingled with the bypass effluent in line 88 and the resulting mixture is fed to the transalkylation reactor via line 67. By bypassing the column 62 with a substantial portion of the bottoms product from column 58, the residue which is lost from the system can be reduced. Preferably in this mode of operation a substantial amount of the bottoms product from column 58 is sent directly to the transalkylation reactor, bypassing the polyethylbenzene column 62. Normally, the weight ratio of the first portion supplied via line 88 directly to the transalkylation reactor to the second portion supplied initially via line 90 to the polyethylbenzene would be within the range of about 1:2 to about 2:1. However, the relative amounts may vary more widely to be within the range of a weight ratio of the first portion to the second portion in a ratio of about 1:3 to 3:1. As stated previously, the molecular sieve catalyst employed in the vapor phase alkylation reactor(s) preferably is monoclinic silicalite. A preferred silicalite is a high silica/alumina ratio silicalite of small crystal size of about 0.5 or less as described previously. In experimental work, vapor phase alkylation was carried out over the preferred form of silicalite catalyst at a first relatively low space velocity favoring transalkylation in the alkylation reactor and at a second relatively high space velocity exemplifying operation in accordance with the present invention to retarding transalkylation and enhancing diethylbenzene production. All of the experimental work was carried out in a single pass reactor operated at an inlet temperature of 400°C. (752°F.) and a pressure of 21 bar ( 300 psig ) . In all of the experimental work, the benzene feed had a purity in excess of 99%, an ethylbenzene content varying from about 0.2-0.7%, a non-aromatic content of about 0.1% or less, or other aromatic, principally toluene and C 3 -C 4 alkylbenzene of about 0.1 wt.% or less. The benzene feed was supplied to the reactor at two flow rates, one providing a liquid hourly space velocity (LHSV) of 70 hr. -1 and the other at 35 LHSV hr. -1 . The molar ratio of benzene to ethylene was initially maintained constant at 10. The one silicalite catalyst employed in this experimental work was a predominantly monoclinic silicalite having a silica/alumina ratio of about 320. This catalyst (designated herein as Catalyst B) was extruded with a high purity alumina binder (about 20 wt.%) in a trilobe configuration as described above. A second silicalite catalyst (Catalyst A) used in the experimental work was very similar to Catalyst B except this silicalite had a silica/alumina ratio of about 225. Catalyst A was likewise used at LHSV of 35 and 70 hr. -1 . Both the high and low silica/alumina ratio silicalite catalyst employed in the experimental work were monoclinic silicalite extruded with the high purity alumina binder in a trilobe shape as described previously. The results of this experimental work are set forth in Figures 3-8, which illustrate the experimental results in terms of various effluent component characteristics plotted on the ordinate versus the time in days plotted on the abscissa. In each of Figures 3-7 the time of the run "D," and thus the age of the catalyst in days since the inception of the run, is plotted on the abscissa. In Figures 3-7, the results over the lower silica/alumina ratio catalyst, Catalyst A, are shown in broken lines. Runs carried out with Catalyst B are indicated by solid lines with space velocities of 35 hr. -1 and 70 hr. -1 . In Figures 3-8, data points associated with the experimental work carried out over the lower silica/alumina ratio Catalyst A at 35 hr. -1 LHSV are indicated by the symbol ▪. Data points for this same catalyst at 70 hr -1 LHSV are indicated by the symbol □. The experimental work carried out over the high silica/alumina Catalyst B of the present invention at 35 hr. -1 LHSV is indicated by the symbol Δ. Data points for Catalyst B at 70 hr -1 are designated by the symbol σ. In Figures 3-7 curves illustrative of data for Catalyst A incorporate "A," e.g. curve Al as in Fig. 3 and curve 4A1 as in Fig. 4, and for Catalyst B the legend "B," e.g. curve 4B1 in Fig. 4. Two test runs were carried out over fresh catalysts having the high silica/alumina ratio at 70 hr. -1 LHSV and the data points for these runs are indicated by the symbols ▴ (Test 2) and ⋄ (Test 3). The catalyst employed in Test 3 was regenerated by burning coke off of the catalyst particles and was used again at a space velocity of 70 hr. -1 LHSV. The regeneration procedure involved the initial injection of an air/nitrogen mixture having a total oxygen content initially of about 4 vol.%. at a temperature of about 380°-420°C. Regeneration lasted for about six hours during which time the amount of nitrogen in the air/nitrogen mixture was progressively decreased until pure air (about 20% oxygen) was injected during the latter stages of the regeneration run. The regeneration step was carried out at a temperature of 450-500°C. The results of the work with this regenerated catalyst are indicated by the symbol ○. Referring first to Fig. 3, the xylene content (Xl) of the effluent relative to the ethylbenzene content measured in parts per million of xylene is plotted on the ordinate. In Fig. 3, the test run carried out with the lower silica/alumina ratio Catalyst A at the lower space velocity of 35 hr -1 LHSV is designated by the broken line curve A1. The corresponding test on the higher silica/alumina ratio Catalyst B at 35 hr -1 LHSV is indicated by the corresponding solid line curve B1. By increasing the space velocity over Catalyst A to 70 hr. -1 indicated by Curve A-2, the xylene content was decreased substantially but still remained above the xylene content achieved with Catalyst B at the lower space velocity. The data given for xylene content in the product, as well as for the various other by-products, are presented in terms of the amount of ethylbenzene in the product. The ethylbenzene content in the reactor effluent typically will be about 12 wt.%. Thus, data presented in terms of 600 ppm (or 0.06 wt.%) relative to ethylbenzene would be equivalent to about 70 ppm xylene in the total reactor effluent. Thus, the high silica/alumina ratio Catalyst B consistently produced lower xylenes content in the effluent than Catalyst A. By increasing the space velocity for Catalyst B from 35 to 70 hr. -1 LHSV (Curves B2 and B3), even lower xylenes contents were achieved. When the catalyst used in Run 3 was regenerated and again used at an LHSV of 70 hr. -1 , (Curve R3), the xylene content was substantially the same as that achieved with the fresh catalyst. Fig. 4 illustrates the diethylbenzene content (DEB) in weight percent relative to the ethylbenzene content plotted on the ordinate versus the time in days (D) plotted on the abscissa. As shown in Fig. 4, by curves 4A1 and 4B1 for Catalysts A and B respectively at 35 hr. -1 LHSV, both catalysts produced substantially the same diethylbenzene content, indicating that transalkylation occurred during the test runs at the lower space velocity. When the space velocity was increased for Catalyst B to 70 hr. -1 , substantially higher diethylbenzene contents were observed. This was observed in both test runs of the fresh Catalyst B and also in the Catalyst B used in the second test run in which Catalyst B was regenerated, as indicated by curves 4B2, 4B3, and 4R3. The incremental increase in diethylbenzene content when going from a space velocity of 35 to 70 hr. -1 is, as shown in Fig. 4, about 0.2 of the space velocity at 35 hr. -1 . Stated otherwise, the ratio of the diethylbenzene content produced at a space velocity of 70 hr. -1 to the diethylbenzene content produced at one-half of this space velocity level (35 hr. -1 ) is shown initially to be about 1.2. With time and aging during the use of the high silica/alumina ratio catalyst, this ratio appears to increase somewhat, for example, to a value of about 1.3 at a catalyst age of 10 days. On the other hand, when the space velocity for the lower silica/alumina ratio Catalyst A was doubled from 35 to 70 hr. -1 LHSV, as indicated by curve 4A2, only a modest increase in diethylbenzene content was observed. This indicated significant transalkylation activity over the lower silica/alumina ratio catalyst even at the higher space velocity. For the higher silica/alumina ratio Catalyst B, on the other hand, the transalkylation activity diminished substantially resulting in a substantially higher diethylbenzene content. Fig. 5 illustrates the results of the experimental work in terms of propylbenzene content (PB) in parts per million relative to ethylbenzene versus the catalyst age in days (D). Corresponding values of the butylbenzene content (BB) relative to ethylbenzene are shown in weight parts per million on the ordinate of Fig. 6. As indicated by Fig. 5, Catalyst B (curve 5B1) showed somewhat lower propylbenzene content in the effluent than Catalyst A (curve 5A1) at the lower space velocity of 35 hr. -1 LHSV. At the higher space velocity of 70 hr. -1 , Catalyst B in both tests showed generally somewhat lower propylbenzene content than at the higher space velocities, as indicated by curves 5B2 and 5B3. This was true for the regenerated catalyst (curve RB3) at the higher space velocity. Referring to Fig. 6, at space velocity of 35 hr. -1 the higher silica/alumina ratio Catalyst B (curve 6B1) showed slightly higher butylbenzene content than was observed for Catalyst A (curve 6A1). Space velocity appeared to show very little effect on the butylbenzene made for the Catalyst B. In Fig. 7 the heavies content (HC) in weight percent relative to ethylbenzene is charted on the ordinate versus the catalyst age (D) in days on the abscissa. As shown in Fig. 7, the higher silica/alumina ratio Catalyst B (curve 7B1) showed substantially lower heavies content relative to ethylbenzene at an LHSV of 35 hr. -1 than for Catalyst A (curve 7A1). The heavies content was generally further reduced by operating with Catalyst B at the higher space velocity. The heavies content was generally considered to include triethylbenzene and higher boiling point components in about the 185°C and above range. This can be seen from an examination of Fig. 7. The heavies content at the lower space velocity of 35 hr. -1 was reduced from a value of 0.25 wt.% relative to ethylbenzene for the Catalyst A (curve 7A1) to a value of about 0.15 wt.% relative to ethylbenzene for the higher silica/aluminum ratio Catalyst B (curve 7B 1). In terms of product concentration, and assuming an ethylbenzene content of about 12% in the reactor effluent, this indicates that operating in accordance with the present invention can substantially reduce the heavies content of the reactor effluent to a value of about 0.02 wt.% or less. When the space velocity is increased to 70 hr. -1 , an even further reduction, down to the level of about 0.01 wt. % or below, can be achieved. Fig. 8 shows the weight percent of ethylbenzene (EB) in the reactor effluent as a function of catalyst age for Catalysts A and B at space velocities of 35 and 70 hr. -1 . Given the scatter of the data and the overlapping data points for the various catalysts at the various space velocities, no curves are presented in Fig. 8. However, it can be seen from the examination of the data presented in Fig. 8 that there was no appreciable difference in the ethylbenzene yields for the catalysts at either the lower or higher space velocity. As can be seen from the foregoing experimental work, use of the higher silica/alumina ratio catalyst consistently lowers the xylene and heavies content from that associated with use of the comparable Catalyst A of a lower silica/alumina ratio. This is true for both the lower and higher space velocities, thus the higher silica/alumina ratio Catalyst B is particularly well-suited for the operation of a system employing parallel reactors and separate alkylation and transalkylation reactors with subsequent liquid phase alkylation as described above with reference to Fig. 2. When operating the alkylation reactors in a parallel mode (during normal operations without regeneration), the xylene yield relative to ethylbenzene can be maintained relatively low at a value of about 600 ppm xylene or lower. This contrasts with 900 or more ppm xylene using the lower silica/alumina ratio Catalyst A. In either case, the xylene content is kept to a level where it is no more than 0.1 wt.% of the ethylbenzene content of the product. When operating a transalkylation reactor at a higher space velocity, as will be involved during a regeneration stage of one of the parallel reactors, the higher silica/alumina ratio Catalyst B results in yet a further dramatic reduction in the xylene yield. While this is accompanied by a significant increase in diethylbenzene yield, as shown in Fig. 4, this can readily be accommodated by the use of the separate transalkylation zone. This is particularly so given that the effect is relatively transitory since the operation of a reactor in the regeneration mode will usually occupy no more than about 5 to 10 percent of the time during which the reactor is operated in the direct alkylation mode. Moreover, the substantially-reduced xylene yield for the high silica/alumina ratio Catalyst B is observed at the higher space velocities associated with operation of one of the alkylation reactors in a regeneration mode. In the experimental work thus far described, the benzene/ethylene ratio was maintained constant at 10:1 molar. In further experimental work respecting the invention, the ethylene feed rate was increased to provide enhanced ethylbenzene content in the product stream while maintaining the benzene feed rate constant. In the vapor phase ethylation procedure described previously, the controlling parameter limiting ethylbenzene production is the ethylene injection rate. Stated otherwise, the rate of ethylbenzene production can be increased by increasing the ethylene injection rate with an attendant reduction, with a benzene injection rate maintained constant, in the benzene/ethylene molar ratio. A difficulty normally encountered when attempting to increase the ethylene injection rate above that normally associated with a more or less "optimum" benzene/ethylene ratio is a substantial increase in the attendant production of impurities and undesirable side products, principally xylene, diethylbenzene, cumene, and "heavies," i . e ., the post-185°C fraction. The preferred monoclinic silicalite alkylation catalysts produce substantially enhanced ethylbenzene production with no significant increase in xylenes and heavies content, accompanied by the increase in ethylene injection rate. This is demonstrated by further experimental work carried out employing the catalyst identified above as "A" and "B" in which the benzene/ethylene ratio was varied by a factor of up to 50% for runs conducted with Catalyst B. The results of this experimental work, in terms of the effect of an increase of ethylene injection and hence a decrease in benzene/ethylation ratio and an increase in ethylbenzene production, is shown in Figures 9-14 which are similar to Figure 8 in that only data points are presented without curves drawn through like data points.. In each of Figures 9-14 the age of the catalyst since the inception of the run is plotted on the abscissa, and the selectivity in terms of effluent component characteristics relative to ethylbenzene is plotted on the ordinate. In each of Figures 9-14, the data corresponding to Catalyst B at a given benzene/ethylene ratio is denoted by the symbols ○ (or=in Fig. 14), ▴ or ⋄ for the following benzene/ethylene ratios: 10 = ○, 8 = ◆, and 6.6 = ▴. Hence, for example in Fig. 9, data points ▪ show the ppm of xylene relative to ethylbenzene observed for Catalyst A at a benzene/ethylene ratio of 10. Similarly, the data points ○ show the results obtained for Catalyst B at a benzene/ethylene ratio of 10, whereas the data points ◆ indicate the results obtained for Catalyst B at a benzene/ethylene ratio of 8 -- that is, a 25% increase in the ethylene injection rate. Similarly, data points ▴ show the results obtained over Catalyst B with an increase in an ethylene injection rate of 50%, corresponding to a benzene/ethylene molar ratio of 6.6. In each of the runs reported in Figures 9-14, the results are for benzene injection at a liquid hourly space velocity of 70 hr. -1 . Referring first to Fig. 9, the xylene content XL relative to the ethylbenzene content measured in ppm is plotted on the ordinate versus the time of the run D in days on the abscissa. As will be recognized by those skilled in the art, a low xylene content is all important because of the close proximity of the distillation point, particularly for meta and para xylene to the distillation point of ethylbenzene. As shown in Fig. 9, the preferred catalyst shows lower xylene make than the already low xylene make associate with the somewhat lower silica/alumina ratio silicalite, Catalyst A. Perhaps, more importantly, the data of Fig. 9 confirms that an increase in ethylene rate to enhance ethylbenzene production showed only a modest increase in xylene production. In fact, as depicted by data points ◆, a 25% increase in ethylene injection resulted in only an initial, modest decrease over the standard injection rate depicted by ○, and as the catalyst aged during the run, even this increase appeared to substantially disappear. The increase in ethylene injection by 50%, as indicated by data points ▴ resulted in a consistently higher xylene make, but even here, the xylene produced was less than the already low xylene level achieved through the use of Catalyst B at a benzene/ethylene ratio of 10. Fig. 10 shows the heavies content HC in a weight percent relative to ethylbenzene plotted on the ordinate versus the catalyst age HD plotted on the abscissa. For Catalyst B the increase in ethylene rate by 25% above the base rate showed a response in heavies content which is virtually the same as the xylene response, that is, no increase in heavies content. When the ethylene rate was increased by 50%, as shown by data points ▴, there was a modest increase in heavies production relative to ethyl benzene, but it was still well below the heavies content associated with Catalyst A at the substantially lower ethylene rate. As a practical matter, the data presented in Figures 9 and 10 shows that the higher silica/alumina ratio of Catalyst B can be used to achieve produce high ethylbenzene production rates without any serious increase in either xylene or heavies content. Figures 11-13 show plots of cumene (CU), normal propylbenzene (PB), and butylbenzene (BB), respectively, plotted on the ordinate versus the time D in days on the abscissa. In each case the values given are parts per million relative to ethylbenzene. As shown in Fig. 11, the selectivity of cumene for Catalyst B was slightly higher than for Catalyst A over the life of the run. However, as shown in Fig. 12, n-propyl benzene production for Catalyst B was less than for Catalyst A, and the response to the two catalysts at a benzene/ethylene ratio of 10 was about the same for butylbenzene as shown in Fig. 13. The increase in cumene and normal propylbenzene with an increase in ethylene rate was observed, as ethylene injection rate was increased from a 125% rate to a 150% rate. The increased impurities content was still low given the increase in ethylbenzene production, assuming a direct correlation between ethylene injection rate and ethylbenzene production. The increase in butylbenzene with an increasing injection rate, as shown in Fig. 13, was substantially more pronounced than the cumene and N-propylbenzene, but even then, at a 25% percent increase in ethylene injection, it was still less than the corresponding increase in ethylbenzene production. Fig. 14, which shows the diethylbenzene content, DEB, in weight percent relative to ethylbenzene, plotted on the ordinate as a function of the time of the run D in days plotted on the abscissa. As shown in Fig. 14, the higher silica/alumina ratio, Catalyst B, shows a material increase in diethylbenzene content over the diethylbenzene content observed for the lower silica/alumina ratio silicalite, Catalyst A. Further increases for Catalyst B are observed when going from a benzene/ethylene ratio of 10, data points=, to progressively increasing ethylene rates as indicated by data points Curves ⋄and ▴, respectively. The higher diethylbenzene rate for the Catalyst B of the present invention is indicative of the lower transalkylation activity of Catalyst B than for the lower silica/alumina ratio silicalite, Catalyst A. This is readily accommodated by the use of a separate downstream transalkylation unit in accordance with the present invention, as described previously. Fig. 15 shows the percent of impurities relative to ethylbenzene, I, plotted on the ordinate versus the increase in production rate R plotted on the abscissa for xylene, Curve 20, heavies, Curve 21, normal propylbenzene, Curve 22, cumene, Curve 23, and butylbenzene, Curve 24. The impurities content reflect values taken for each of the five impurities at the tenth day of the run, at 100%, 125%, and 150% ethylene rates. As indicated in Fig. 15, and consistent with the data presented in Figures 9 and 10, the xylene content and the heavies content relative to ethylbenzene remain constant, even though the ethylbenzene production is increased by 25%, as indicated by the increased ethylene injection rate. As indicated by the aforementioned experimental work, use of the small crystal size monoclinic silicalite results in a low xylene yield in the initial gas phase alkylation reaction. The lowest xylene yield is achieved through the use of the higher silica/alumina ratio silicalite, Catalyst (B). The provision of the second liquid phase alkylation reaction zone, in accordance with the present invention, can be expected to produce an even lower xylene mate than that produced in the initial gas phase alkylation reaction. By recycling the output from the liquid phase alkylation reactors via line 84 to the benzene separation zone, the overall xylene make of the process can be even further reduced. As indicated previously, the output from the liquid phase alkylation reaction zone is preferably supplied to the first stage benzene recovery zone because.

Claims: In the production of ethylbenzene and the separate transalkylation of polyethylbenzene, the method comprising: (a) providing a first alkylation reaction zone containing a molecular sieve aromatic alkylation catalyst; (b) supplying a feedstock containing benzene and ethylene to said first alkylation reaction zone; (c) operating said alkylation reaction zone at temperature and pressure conditions in which benzene is in the gaseous phase to cause gas-phase ethylation of said benzene in the presence of said molecular sieve catalyst to produce an alkylation product comprising a mixture of xylene, ethylbenzene and polyalkylated aromatic; (d) recovering said alkylation product from said reaction zone and supplying said product from said reaction zone to an intermediate recovery zone for the separation and recovery of ethylbenzene from the alkylation product and the separation and recovery of a polyalkylated aromatic component including diethylbenzene; (e) supplying at least a portion of said polyalkylated aromatic component including diethylbenzene in said polyalkylated component to a transalkylation zone containing a molecular sieve transalkylation catalyst; (f) supplying benzene to said transalkylation reaction zone; (g) operating said transalkylation reaction zone under temperature and pressure conditions to cause disproportionation of said polyalkylated aromatic fraction to produce a disproportionation product containing unreacted benzene and having a reduced diethylbenzene content and an enhanced ethylbenzene content; (h) supplying disproportionation product from said transalkylation reaction zone to a second alkylation reaction zone containing a molecular sieve aromatic alkylation catalyst; (i) supplying ethylene to said second alkylation zone and operating said second alkylation reaction zone at temperature and pressure conditions in which said unreacted benzene is maintained in the liquid phase to cause liquid phase ethylation of said benzene to produce an alkylation product containing xylene in a lower xylene concentration than the xylene concentration in the alkylation product from said first recited alkylation reaction zone; and (j) supplying the alkylation product from said second alkylation zone to said intermediate recovery zone. The process of claim 1 wherein the molecular sieve aromatic alkylation catalyst in said first alkylation reaction zone comprises monoclinic silicalite and the molecular sieve aromatic alkylation catalyst in said second alkylation reaction zone has an effective pore size greater than the effective pore size of said monoclinic silicalite catalyst in said first recited reaction zone. The method of claim 2 wherein the molecular sieve aromatic alkylation catalyst in said second alkylation reaction zone is selected from the group consisting of zeolite Y, zeolite beta, and zeolite La-Beta. The method of claim 3 wherein said catalyst in said second alkylation reaction zone is zeolite beta. The method of claim 3 wherein said catalyst in said second alkylation reaction zone is zeolite La-Beta. The method of claim 1 wherein said catalyst in said first alkylation reaction zone is silicalite having a silica/alumina ratio of at least 275. The method of claim 6 wherein said silicalite catalyst is monoclinic silicalite having an average crystal size of less than one micron. The method of claim 7 wherein said monoclinic silicalite alkylation catalyst has an average crystal size of 0.5 microns or less and is formulated with an alumina binder to provide catalyst particles having a surface area/volume ratio of at least 152 cm -1 (60 in. -1 ). The method of claim 1 wherein said transalkylation reaction zone contains a zeolite Y transalkylation catalyst and is operated under temperature and pressure conditions effective to maintain the feedstock in said transalkylation zone in the liquid phase. The method of claim 1 wherein the first alkylation reaction zone comprises a plurality of pentasil molecular sieve aromatic alkylation catalyst comprising predominantly monoclinic silicalite; and wherein the concentration of xylene in said product from the first recited alkylation reaction zone is no more than 0.1 percent based upon ethylbenzene in the product. The method of claim 10 wherein said monoclinic silicalite alkylation catalyst has an average crystal size of 0.5 microns or less and is formulated with an alumina binder to provide catalyst particles having a surface area/volume ratio of at least 152 cm -1 ( 60 in. -1 ). In the production of ethylbenzene and the separation transalkylation of polyethylbenzene, the method comprising: (a) providing a first alkylation reaction zone molecular sieve aromatic alkylation catalyst; (b) supplying a feedstock containing benzene and ethylene to said first alkylation reaction zone; (c) operating said alkylation reaction zone at temperature and pressure conditions in which benzene is in the gaseous phase to cause gas-phase ethylation of said benzene in the presence of said molecular sieve catalyst to produce an alkylation product comprising a mixture of xylene, ethylbenzene and polyalkylated aromatic components including diethylbenzene; (d) recovering said alkylated product from said first reaction zone and supplying said product from said first reaction zone to a benzene recovery zone for the separation of benzene substrate from alkylation product; (e) operating said benzene recovery zone to produce a lower boiling benzene-containing fraction and a higher boiling fraction comprising a mixture of monoalkylated aromatic and polyalkylated aromatic component; (f) recycling benzene from said benzene recovery zone to said reaction zone; (g) supplying said higher boiling point fraction from said benzene recovery zone to a secondary separation zone; (h) operating said secondary separation zone to produce a secondary lower boiling fraction comprising a monoalkylated aromatic component and a higher boiling fraction comprising a heavier polyalkylated aromatic component; (i) supplying at least a portion of said polyalkylated aromatic component including diethylbenzene in said polyalkylated component to a transalkylation reaction zone containing a molecular sieve transalkylation catalyst to produce an alkylation product containing xylene in a lower xylene concentration than the xylene concentration in the alkylation product from said first recited alkylation reaction zone; (j) supplying benzene to said transalkylation reaction zone; (k) operating said transalkylation reaction zone under temperature and pressure conditions to cause disproportionation of said polyalkylated aromatic fraction to produce a disproportionation product containing unreacted benzene and having a reduced diethylbenzene content and an enhanced ethylbenzene content; (1) supplying disproportionation product from said transalkylation reaction zone to a second alkylation reaction zone containing a molecular sieve aromatic alkylation catalyst; (m) supplying ethylene to said second alkylation zone and operating said second alkylation reaction zone at temperature and pressure conditions in which said unreacted benzene is maintained in the liquid phase to cause liquid phase ethylation of said benzene; and (n) supplying the alkylation product from said second alkylation zone to said benzene recovery zone. The method of claim 12 wherein said benzene recovery zone is operated in a first stage in which a portion of benzene is recovered overhead from said alkylated product and a second stage in which additional benzene is recovered overhead from said alkylation product and benzene is recycled from both said first and second stage of said benzene recovery zone to said first alkylation reaction zone. The method of claim 13 wherein set at least portion of said alkylation product from said second alkylation zone is supplied to said first stage of said benzene recovery zone. The method of claim 14 wherein the molecular sieve aromatic alkylation catalyst in said first alkylation reaction zone comprises monoclinic silicalite and the molecular sieve aromatic alkylation catalyst in said second alkylation reaction zone has an effective pore size greater than the effective pore size of said monoclinic silicalite catalyst in said first recited reaction zone. The method of claim 15 wherein the molecular sieve aromatic alkylation catalyst in said second alkylation reaction zone is selected from the group consisting of zeolite Y, zeolite beta, and zeolite La-Beta. The process of claim 16 wherein said transalkylation reaction zone is operated under temperature and pressure conditions effective to maintain the feedstock in said transalkylation zone in the liquid phase.

Inventors: BUTLER JAMES R, MERILL JAMES, BUTLER, JAMES R., MERILL, JAMES
Applicants: FINA TECHNOLOGY, FINA TECHNOLOGY, INC.
Application year: 2000
Main classification: C07C15
IPC codes: C07C15