KR20100017555A - Process for selective aromatics disproportionation with increased conversion - Google Patents

Abstract

An improved process is disclosed for the selective disproportionation of toluene. The process preferably uses a disproportionation catalyst comprising a pentasil type zeolite such as MFI that is bound with aluminum-phosphate. Running the process at a toluene conversion greater than 30 wt-% and at a hydrogen-to-hydrocarbon ratio less than 3.0, and especially a ratio of 0.1 to 1.0, improves the maximum yield of para-xylene. Periodic rejuvenation by increasing the hydrogen-to-hydrocarbon ratio removes some carbon deposits and restores catalyst activity. An inert diluent gas assists in selective pre-coking of the catalyst as well.

Description

Selective disproportionation method of aromatics with increased conversion {PROCESS FOR SELECTIVE AROMATICS DISPROPORTIONATION WITH INCREASED CONVERSION}

The present invention relates to an improved process for converting aromatic hydrocarbons, such as an improved process for converting toluene to paraxylene. More specifically, the present invention relates to the operation and selectivation of the disproportionation process at low levels of hydrogen to enable preferred coke formation and aromatic conversion.

Xylene isomers are prepared in large quantities from petroleum as raw materials for various important industrial chemicals. The most important xylene isomer is paraxylene, which is the main feedstock for polyester and its production continues to increase at a high rate in response to high basic demands. Orzaylene is used to make phthalic anhydride, which has a huge market but is already saturated. Metaxylene is used in smaller amounts in products such as plasticizers, azo dyes and wood preservatives, but its use is increasing. Ethylbenzene is generally present in xylene mixtures and is often recovered for the production of styrene, but is often considered to be a less desirable component of C ₈ aromatics.

The overall importance of xylene as feedstock for industrial chemicals in aromatic hydrocarbons is comparable to that of benzene. Since neither xylene or benzene obtained from petroleum by reforming naphtha is produced in an amount sufficient to meet demand, conversion of other hydrocarbons is necessary to increase the yield of xylene and benzene. The most common method is to selectively disproportionate toluene to obtain benzene and C ₈ aromatics, from which the individual xylene isomers are recovered.

Recent objectives for many petrochemicals and aromatic complexes are to increase the yield of xylenes and to reduce the importance of the production of benzene. The demand for jaalene derivatives is growing faster than benzene derivatives. Industrialized countries are implementing modified refinery processes to reduce the benzene content of gasoline, resulting in increased supply of benzene to meet demand. Therefore, it is a preferred object to obtain a high yield of xylene even at the expense of benzene, and a method of converting toluene to obtain a high yield of xylene has been commercially used.

U.S. Patent 4,016,219 discloses a toluene disproportionation method using a catalyst comprising a zeolite modified by adding phosphorus in an amount of at least 0.5% by mass. Zeolite crystals are contacted with a phosphorus compound to react the zeolite with the phosphorus compound. The modified zeolite can then be incorporated into the indicated matrix material. US Pat. No. 4,097,543 discloses a toluene disproportionation method for the selective preparation of paraxylene using zeolites with controlled precoke treatment. Zeolites are ion exchanged with various elements selected from Groups IB to VIII and can be synthesized with various clays and other porous matrix materials.

U.S. Patent No. 4,182,923 discloses a toluene disproportionation process which has a high silica to alumina ratio of greater than 12 and is highly converted toluene to benzene and paraxylene using aluminosilicate zeolites modified with phosphorus deposition by treatment with ammonium hydrogen phosphate. Is disclosed. US Pat. No. 4,629,717 discloses phosphorus modified alumina hydrogels formed by gelling homogeneous hydrosols. This composite has a relatively high surface area of 140 to 450 m 2 / g and high activity and selectivity in the 1-heptene conversion test.

US Pat. No. 6,114,592 discloses an improved combination method for the selective disproportionation of toluene. This combination includes disproportionation using zeolite-based catalysts after selective hydrogenation of the toluene feedstock. U.S. Patent No. 6,359,185 discloses an oil drop of a zeolite-based catalyst in an amorphous aluminum phosphate binder to increase selectivity.

U. S. Patent No. 6,191, 331 discloses a preliminary coke treatment method which avoids a significant temperature rise by using a low ratio of hydrogen to hydrocarbon and low pressure in the presence of nitrogen. U. S. Patent No. 6,429, 347 discloses implementing the process at a hydrogen to hydrocarbon ratio between 0.2 and 0.5 to improve the selectivity of paraxylene and reduce the selectivity of benzene.

Researchers in the field of aromatic disproportionation continue to study catalysts and methods for the particularly high conversion of toluene to paraxylene, with desirable selectivity and stability.

Summary of the Invention

It is an object of the present invention to provide an improved process for the disproportionation of aromatic hydrocarbons. The specific purpose is to obtain the highest yield of xylene by the selective disproportionation of toluene.

The present invention is based on the unexpected finding that the process at low levels of hydrogen to hydrocarbons promotes an increase in the yield of paraxylene by allowing higher conversion of toluene than the process has already been used. In addition, low levels of hydrogen also improve the condition and selective precoking of the zeolite-based catalyst when combined with the nitrogen diluent prior to use in high conversion processes.

The present invention relates to a process for the production of xylene comprising a selective disproportionation zone with conditions comprising greater than 30% by weight toluene conversion levels and a hydrogen to hydrocarbon ratio of 0.1 to 1.0. In the disproportionation zone, the stream is contacted with the disproportionation catalyst under disproportionation conditions. If the disproportionation inlet temperature is increased by 20 ° C. relative to the initial disproportionation inlet temperature, the disproportionation catalyst is regenerated by increasing the hydrogen to hydrocarbon ratio by at least 0.5. The disproportionation catalyst preferably comprises pentasil zeolite-based aluminosilicate, most preferably MFI. This catalyst is subjected to a preliminary coke stage before being used in the disproportionation zone to increase its selectivity to paraxylene in the product above its equilibrium concentration.

The foregoing description, other objects, and embodiments will be apparent from the detailed description of the invention.

Description of Preferred Embodiments

A broad embodiment of the present invention is a selective toluene disproportionation process that operates at low hydrogen to hydrocarbon ratios to increase selectivity to paraxylene. Thus, one essential component of the process is a zeolite-based catalyst which, prior to being used for disproportionation, deposits a controlled concentration of carbon on the catalyst and is subjected to a preliminary coke step to increase paraxylene selectivity. The paraxylene content of the paraxylene rich product from the disproportionation reaction of the present invention is an amount exceeding its equilibrium concentration at disproportionation conditions.

Selective disproportionation process areas of the present invention include molecular sieves and refractory inorganic oxides. Preferred molecular sieves are zeolite-based aluminosilicates, or zeolites, any of which may have a SiO ₂ ZAl ₂ O ₃ ratio of greater than 10, preferably greater than 20, and a pore diameter of 5 to 8 angstroms. have. Specific examples of zeolites that can be used are zeolites of the MFI, MEL, EUO, FER, MFS, MTT, MTW, TON, MOR and FAU types. Pentasil zeolites MFI, MEL, MTW and TON are preferred, and MFI-type zeolites, commonly designated ZSM-5, are particularly preferred.

Preferred methods for preparing MFI-type zeolites are well known in the art. Zeolites are preferably prepared by crystallizing a mixture containing alumina source, silica source, alkali metal source, water and alkyl ammonium compound or precursor thereof.

Refractory binders or matrices are used to facilitate the production of disproportionation catalysts, provide strength, and reduce manufacturing costs. The binder should be uniform in composition and fire resistant relative to the conditions used in the process. Suitable binders include at least one of inorganic oxides such as alumina, magnesia, zirconia, chromia, titania, boria, toria, phosphorus oxide, zinc oxide and silica. Preferred binders are alumina and / or silica. The amount of zeolite present in the bound catalyst can vary considerably, but is usually present in an amount of 30 to 90 mass%, preferably 50 to 80 mass% of the catalyst.

Preferred binder or matrix components are phosphorus containing alumina (hereinafter referred to as aluminum phosphate) components. Phosphorus can be synthesized with alumina in any acceptable manner known in the art. The zeolite and aluminum phosphate binders are mixed and then shaped into particles using methods well known in the art, such as gelling, peeling, granulation, sizing, spray drying, extrusion or any combination of these techniques. Preferred methods of preparing the zeolite / aluminum phosphate support include adding zeolite to an alumina sol or phosphorus compound, shaping the mixture of alumina sol / zeolite / phosphorus compound into particles using the oil dropping method described below, and spherical Calcination of the particles.

Preferred oil dropping methods for preparing aluminum phosphate are described in US Pat. No. 4,629,717, the contents of which are incorporated by reference. The technique described in this' 717 patent involves gelling a hydrosol of alumina containing a phosphorus compound using known oil dropping methods. In general, such techniques involve hydrolysis by the decomposition of aluminum in aqueous hydrochloric acid at reflux temperatures of about 80 ° C to 105 ° C. The ratio of aluminum to chlorine in the sol is in the range of 0.7: 1 to 1.5: 1 mass ratio. The phosphorus compound is then added to the sol. Preferred phosphorus compounds are phosphoric acid, phosphorous acid and ammonium phosphate. The relative amounts of phosphorus and aluminum expressed in molar ratios range from 10: 1 to 1: 100, respectively, based on the element. Zeolite is added to aluminum phosphate hydrosol and the mixture is gelled. One method of gelling this mixture involves combining the mixture with a gelling agent and then dispersing the resulting formulation into an oil or tower heated to a high temperature so that gelling occurs to form elliptical particles. Gelling agents which can be used in this process are hexamethylene tetraamine, urea or mixtures thereof. The gelling agent releases ammonia at high temperatures to cure or convert the hydrosol spheres into hydrogel spheres. The blended mixture is preferably dispersed in the oil bath in the form of droplets from a nozzle, an orifice or a rotating disk. The spheres are then recovered continuously from the oil bath and it is common to undergo specific drying and aging treatments in ammonia solution and in oil to further improve the physical properties. The resulting aged and gelled particles are washed, dried at relatively low temperatures of 100 ° C. to 150 ° C., and calcined at temperatures of 450 ° C. to 700 ° C. for 1 to 20 hours.

Alternatively, the particles can be formed by spray drying the mixture. In any case, the conditions and equipment should be selected to obtain small spherical particles and the particles should preferably have an average diameter of less than 1.0 mm, more preferably 0.2 to 0.8 mm and optimally 0.3 to 0.8 mm.

The amount of phosphorus-containing alumina component (oxide form) present in the catalyst may range from 10 to 70 mass%, preferably from 20 to 50 mass%. Aluminum phosphate binder / matrix can be added to the hydrosol prior to dropping, magnesia, beryllia, boria, silica, germania, tin oxide, zinc oxide, titania, zirconia, vanadia, iron oxide, chromia, cobalt oxide Other inorganic oxides, including the like, may optionally be contained in small amounts, but are not limited thereto.

Aluminum-phosphate binders are generally amorphous, ie, the binder material is essentially amorphous in nature. Less than 10 mass% per binder pore volume is a micropore volume unique to the crystalline material, and the micropore volume is more preferably less than 5% of the pore volume, optimally less than 2% of the pore volume. Crystalline aluminophosphate is generally a binder material that is not suitable for making strong fracture resistant catalysts. Since the material that is not in the amorphous phase is generally present as gamma-alumina, the proportion of crystalline material increases as the phosphorus content of amorphous aluminum phosphate decreases. The average bulk density of the spheres depends on the phosphorus content; the higher the proportion of phosphorus, the lower the average bulk density. In addition, the surface area is controlled by the phosphorus content and oil-dropped gamma-alumina spherical particles typically have a surface area of up to 250 m 2 / g, but the surface area of the elliptical particles of aluminum phosphate may be up to 450 m 2 / g. The Al / P atomic ratio of the binder / matrix is generally in the range of 1/10 to 100/1, more typically in the range of 1/5 to 20/1, and typically between 1: 1 and 5: 1.

The catalyst may contain a metal component and is preferably selected from components of the group consisting of gallium, rhenium and bismuth. However, the catalyst is preferably composed mainly of a zeolite-based aluminosilicate having a pore diameter of 5 to 8 mm 3 and an aluminum phosphate binder.

Selective precoking of the zeolite-based catalyst increases the proportion of paraxylene in the paraxylene rich product to exceed the equilibrium level at disproportionation conditions. The proportion of paraxylene in the product above the equilibrium level under disproportionation conditions is generally at least 80 mass%, preferably at least 90 mass% of the C ₈ aromatics. Precoking is carried out on fresh or regenerated catalyst for a time of 0.5 hours to 10 days before being used for the disproportionation reaction. The catalyst may be precoked in-situ or ex-situ to increase the proportion of paraxylene in the C ₈ aromatic product.

Preliminary coking is carried out at conditions comprising at least one of higher temperature, lower pressure and higher space velocity as compared to the conditions of the subsequent disproportionation step. Such preliminary coking conditions include an absolute pressure of 100 kPa to 4 MPa and a liquid space velocity per hour of 0.2 to 20 hr ⁻¹ . The conditions include at least one of the inlet temperature higher than at least 50 ° C., the pressure lower than at least 100 kPa, or preferably no more than half the pressure used in the subsequent disproportionation step. Lower pressures and / or less hydrogen / hydrocarbon ratios will reduce the rate of exothermic aromatic-saturation reactions, thus limiting the temperature rise and therefore should be a relatively flat temperature profile. Thus, a typical temperature range would be 300 ° C to 700 ° C and a typical hydrogen to coke forming feed range would be 0.01 to 5.

The use of nitrogen or other similar inert diluent gases such as methane, ethane, or propane is considered to be very advantageous when included with hydrogen during the preliminary coking step. These thermally inert diluent gases assist in adjusting the temperature profile and are present in a molar ratio of 0.01 to 10, preferably greater than 1, relative to the coke forming feed. The temperature profile is believed to affect the rate of cocosification in various parts of the catalyst bed. Therefore, steep temperature gradients will result in non-uniform coke deposition, and thus the selectivity of the catalyst layer in different portions will change, resulting in inferior performance in subsequent disproportionation reactions. Thus, a typical temperature difference across the bed of catalyst during selective precoking will be between 10 ° C. increase or decrease, preferably between 3 ° C. increase and 4 ° C. decrease.

Preliminary coking is carried out with 5 to 40 mass% carbon content or catalytic coke, with 10 to 30 mass% carbon being preferred. The coke forming feed for preliminary coking may comprise a feedstock for the disproportionation step as described below, such as toluene, or preferably a mixture or other specific known in the art comprising aromatics. Hydrocarbons may be used. Further details on preliminary coking are disclosed in US Pat. No. 4,097,543 and US Pat. No. 6,191,331, which is incorporated herein by reference.

The feedstock for this process may be formulated in any combination to obtain more valuable alkylaromatics of the general formula C ₆ H ( _6-n ) R _n (wherein n is from 0 to 5, R is CH ₃ , C ₂ H ₅ , C ₃ H ₇ , or C ₄ H ₉ ). Suitable alkylaromatic hydrocarbons are, for example, benzene, toluene, xylene, ethylbenzene, trimethylbenzene, ethyltoluene, propylbenzene, tetramethylbenzene, ethyldimethylbenzene, diethylbenzene, methylpropylbenzene, ethylpropylbenzene, triethyl Benzene, diisopropylbenzene and mixtures thereof, including but not limited to.

The feedstock comprises toluene, optionally in combination with a C ₉ aromatic, suitably derived from one or a variety of sources. The feedstock can be prepared synthetically, for example, according to the manner in which naphtha is subjected to catalytic reforming or pyrolysis to water treatment to obtain an aromatic rich product. The feedstock may be derived from a product having an appropriate purity by extracting an aromatic hydrocarbon from a mixture of aromatic and non-aromatic hydrocarbons and fractionating the extract. For example, aromatics can be recovered from reformate. The reformate may be prepared by any method known in the art. The aromatics can then be recovered from the reformate by using one of the optional solvents such as sulfolane types in the liquid-liquid extraction zone. The recovered aromatics can then be separated by fractional classification into a stream having a desired carbon number range. The feedstock should contain up to 10 mass% non-aromatics, and the content of benzene and C ₈ aromatics is typically an economical determination of the dilution of toluene from these aromatics. If the severity of the reforming or pyrolysis is high enough, no extraction is necessary and fractional distillation may be sufficient to produce the feedstock.

In the disproportionation process, the feed is usually first heated by indirect heat exchange with the discharge of the reaction zone and then further heated in the calcined heater. The resulting gaseous stream is passed through a reaction zone which may comprise one or more individual reactors. While the use of a single reaction vessel with a fixed cylinder bed of catalyst is preferred, other reaction schemes or radial flow reactors using catalysts of moving bed may be used if desired. Passing the combined feed through the reaction zone produces a gaseous effluent stream comprising hydrogen and product and unconverted feed hydrocarbons. This discharge is typically cooled through indirect heat exchange with the stream entering the reaction zone and is further cooled using air or cooling water. The temperature of the effluent stream is generally lowered by heat exchange to effect sufficient condensation of the hydrocarbon product having substantially all feeds and six or more carbon atoms per molecule. The resulting mixed phase stream is passed through a vapor-liquid separator where these two phases are separated and the hydrogen rich vapor from this separator is recycled to the reaction zone. The condensate from the separator is passed to a stripping column where substantially all of the C ₅ and light hydrocarbons present in the exhaust are concentrated into an overhead stream and removed from the process. The aromatic rich stream, referred to herein as the disproportionate effluent stream, is recovered as pure strip bottom product.

Conditions used in the disproportionation process zone usually include temperatures of 200 ° C to 600 ° C, preferably 350 ° C to 575 ° C. As the catalyst slowly loses activity during the process, the temperature required to maintain the desired conversion will rise. Thus, the normal end of operation temperature may be at least 65 ° C. above the temperature at the time of operation.

The disproportionation zone generally operates at a hydrogen to hydrocarbon ratio of 0.1 to 1.0, with a hydrogen to hydrocarbon ratio of 0.2 to 0.5 being preferred. The ratio of hydrogen to hydrocarbon is calculated based on the molar ratio of free hydrogen to feedstock hydrocarbons. The periodic increase in hydrogen to hydrocarbon of at least 0.5 allows for catalyst regeneration by hydrogenation of soft coke. The hydrogen to hydrocarbon ratio during regeneration is preferably in the range of 1-5.

The disproportionation zone is operated at moderately high pressures ranging from 100 kPa to 6 MPa absolute. The preferred pressure range is 2 to 3.5 MPa. The disproportionation reaction can be carried out over a wide range of space velocities, and at higher space velocities a higher proportion of paraxylene can be obtained at the expense of conversion. The liquid space velocity per hour is generally in the range of 0.2 to 20 hr ⁻¹ .

The disproportionate effluent stream is separated into a light recycle stream, a paraxylene rich mixed C ₈ aromatic product and a heavy aromatic stream. The para xylene rich product can be sent to the xylene separation zone for recovery of pure para xylene and optionally other xylenes and ethylbenzene can also be recovered as pure product. The paraxylene rich stream preferably contains paraxylene in an amount exceeding the equilibrium concentration relative to total xylene under disproportionation conditions, more preferably at least 80 mass% of paraxylene, and 85 mass% It is more preferable to contain the above para xylene. The light recycle stream may be diverted to other uses such as benzene and toluene recovery, but optionally some are recycled to the disproportionation zone because they contain not only benzene but also large amounts of non-aromatics that remain with this benzene to reduce commercial value. . The heavy recycle stream contains substantially all C ₉ and heavier aromatics and can be discharged as a product of the process.

The xylene separation zone may recover substantially pure paraxylene from the paraxylene rich stream in the xylene separation zone using one or more different separation techniques such as fractional fractionation, crystallization or selective adsorption. Common crystallizations are disclosed in US 3,177,255, US 3,467,724 and US 3,662,013. Various other crystallization alternatives are discussed in US Pat. No. 5,329,061, which is incorporated by reference. In embodiments wherein the paraxylene rich product comprises paraxylene in an amount substantially above the equilibrium concentration, the recovery of paraxylene uses a single step of crystallization corresponding to a higher temperature than the purification step of conventional crystallization. Can be carried out.

Alternative separation regions include layers of molecular sieves that operate according to the teachings of US 3,201,491 that facilitate the use of molecular sieves of continuous moving layers. Subsequent improvements to the method are described in US 3,696,107 and US 3,626,020. Details of the operation of the xylene separation zone can be obtained from US 4,039,599 and US 4,184,943. The xylene separation zone also shifts the isomers of orthoxylene and metaxylene to paraxylene, as well as catalytic alkyl- in the separation loop to isomerize ethyl benzene to xylene or dealkylate to benzene. It can be incorporated into the aromatic isomerization zone. The benzene produced here can also be sent to the transalkylation zone. The xylene separation zone may also use the simulated simultaneous adsorptive separation method of US 4,402,832. Extracts and raffinate streams can be handled as described in these references or as described in US 4,381,410 and US 5,495,061.

Various combinations of the methods described above are within the scope of the present invention. For example, toluene as well as benzene can be charged to the disproportionation zone as a supplemental feedstock. The xylene separation zone may use one or more of a number of known separation techniques such as adsorption, crystallization and fractionation. Orzaylene and / or metaxylene may be recovered as pure product from the xylene separation zone using one or more such techniques.

The disproportionation process can be carried out until the conversion of toluene is no longer economically preferred by catalyst reduction, destruction, or inactivation. When the initial conversion is reduced as measured by an increase in inlet temperature of 20 ° C. or more, a typical economic target arises, at which point the catalyst is regenerated by an increase of at least 0.5 molar ratio of free hydrogen to feedstock hydrocarbons. Thus, preferred regeneration conditions include a molar ratio of free hydrogen present to feedstock hydrocarbons 1 to 5, inlet temperature 200 ° C. to 600 ° C., absolute pressure 100 kPa to 6 MPa and liquid space velocity 0.2 to 20 hr ⁻¹ per hour. .

1 shows the yield of paraxylene at various hydrogen to hydrocarbon ratios as the toluene conversion is increased on the optionally precoated catalyst.

FIG. 2 shows the yield of benzene at various hydrogen to hydrocarbon ratios as the toluene conversion is increased on the optionally precoated catalyst.

The following examples are present to demonstrate the invention and to illustrate certain embodiments thereof. These examples do not limit the scope of the invention as described in the claims. There are many possible variations that would be appreciated by one of ordinary skill in the art to be within the spirit of the present invention.

Example 1

Alumina-phosphate coupled MFI catalysts were prepared to evaluate the present invention. The first solution was prepared by adding phosphoric acid to an aqueous solution of hexamethylenetetraamine (HMT) in an amount such that the ratio of aluminum to phosphorus in the binder is 1: 1 so that the phosphorus content of the final catalyst was 3.8 mass%. The second solution was prepared by adding MFI-type zeolite having a Si / Al ₂ ratio of 39 to a sufficient amount of alumina sol prepared by digesting metallic aluminum in hydrochloric acid so that the zeolite content in the final catalyst was 70% by mass. These two solutions were mixed to obtain a homogeneous mixture of HMT, phosphorus, alumina sol and zeolite. This mixture was dispersed in droplet form in an oil bath maintained at 93 ° C. The droplets were held in the oil tank until the droplets cured to form hydrogel spheres with a diameter of 1.6 mm. The spheres were then taken out of the oil bath, washed with water, air dried and then calcined at a temperature of 650 ° C. This disproportionation catalyst was used in the disproportionation test and preliminary coking described below.

Example 2

The catalyst was subjected to approximately 90 mol% of total xylene in xylene at conditions including a temperature of 560 ° C., a pressure of 0.72 MPa and a 4 weight hourly space velocity (WHSV) in the presence of a hydrogen to hydrocarbon molar ratio of 0.5. Precoking for a sufficient time. The disproportionation of pure toluene was then carried out at 2.45 MPa and 4 WHSV in the presence of pure hydrogen at various temperatures required to achieve a range of toluene conversion levels.

To illustrate the invention, test runs were performed at hydrogen to hydrocarbon ratios 3.0, 2.0, 1.0, 0.5 and 0.2. FIG. 1 shows the yield of paraxylene at these hydrogen to hydrocarbon ratios as the toluene conversion is increased on the optionally precoated catalyst. 2 shows the yield of benzene at these hydrogen to hydrocarbon ratios. Surprisingly, the maximum critical yield of paraxylene was found around 30% by weight conversion, and these maximum values shift to even higher conversions as the hydrogen to hydrocarbon ratio is reduced below 3.0.

According to FIG. 1, it was shown that a hydrogen to hydrocarbon ratio of 1.0 provides a maximum yield of paraxylene in the range of 12.5 wt% for a conversion level of 30 to 33 wt%. In addition, when the hydrogen to hydrocarbon ratio is reduced below 1.0, the maximum value moves to a much higher conversion level. This conversion shift increases the yield of paraxylene to be achieved even more, which is impossible at higher ratios of hydrogen to hydrocarbon.

According to FIG. 2, the benzene yield was observed to increase with increasing conversion level of toluene in all cases. However, at all conversion levels, the yield of benzene decreases as the hydrogen to hydrocarbon ratio decreases. If the conversion level was less than 33% by weight and the hydrogen to hydrocarbon ratio was less than 1.0, the benzene yield was lower than 15% by weight in all cases.

Example 3

The addition of nitrogen during the selectivity change step was investigated by performing a first test with no nitrogen and a test with a nitrogen to hydrocarbon ratio of 2.5, both of which maintained a hydrogen to hydrocarbon ratio of 0.5. The temperature was maintained at 56 ° C., the pressure was 0.72 MPa, and the WHSV was maintained at 3 hr ⁻¹ . Subsequently, a disproportionation reaction was carried out using pure toluene at 2.45 MPa, WHSV of 4 hr ⁻¹ and hydrogen to hydrocarbon ratio of 3.0 to achieve 30% by weight of toluene conversion.

The data obtained in the disproportionation test demonstrated that, at a ratio of paraxylene to total xylene of 90% by weight, the pure hydrogen selectivity change process achieved a benzene to total xylene ratio of 1.6. However, the selectivity change process using nitrogen achieved a benzene to total xylene ratio of 1.3. Thus, it has been found that the presence of an inert gas such as nitrogen during the selectivity change process has the beneficial effect of reducing benzene production.

Claims (7)

a) disproportionation of the toluene-containing feedstock, wherein the feedstock is contacted with a catalyst which is optionally precoked in the presence of an inert gas, the contacting being in a molar ratio of 0.1 to 1.0 free hydrogen to feedstock hydrocarbon Greater than 30% by weight of toluene present in the feedstock, carried out under disproportionation conditions, including inlet temperatures of 200 ° C. to 600 ° C., absolute pressures of 100 kPa to 6 MPa and liquid hourly space velocity of 0.2 to 20 hr ⁻¹ . Is converted to obtain a paraxylene rich product containing paraxylene above its equilibrium concentration; b) recovering paraxylene from paraxylene rich product by one or both of adsorption and crystallization; And c) performing step (a) for a time until the initial inlet temperature increases by at least 20 ° C., at which point the catalyst is regenerated by increasing the molar ratio of free hydrogen to feedstock hydrocarbon by at least 0.5. Method for producing para xylene comprising a. The process of claim 1 wherein the conversion of toluene is at least 33% by weight. The process of claim 1 wherein the conversion of toluene is 30 to 33 weight percent. 4. The process of claim 3, wherein the paraxylene rich product of step (b) further comprises benzene in an amount of up to 15% by weight when calculated on the basis of toluene feed. The process of claim 1 wherein the regeneration conditions of step (c) are free hydrogen to feedstock hydrocarbons present in a molar ratio of 1 to 5, inlet temperature of 200 ° C. to 600 ° C., absolute pressure of 100 kPa to 6 MPa and 0.2 to 20. and an hourly liquid space velocity of hr ⁻¹ . The process of claim 1, wherein an inlet temperature of 300 ° C. to 700 ° C., an absolute pressure of 100 kPa to 4 MPa, a molar ratio of free hydrogen to coke-forming feed of 0.1 to 5, an inert diluent gas to coke forming feed of 0.01 to 10 Group consisting of coke forming feed and MFI, MEL, MTW and TON in the presence of a gas comprising hydrogen and an inert diluent gas at preliminary coking conditions comprising a molar ratio of water and a liquid hourly space velocity of 0.2 to 20 hr ⁻¹ Selectively precoking the catalyst by contacting a pentasil zeolite selected from to deposit 5-40% by mass of carbon on the catalyst and optionally obtaining a precoked catalyst. The method of claim 6, wherein the inert diluent gas is selected from the group consisting of nitrogen, methane, ethane, propane and mixtures thereof.

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