BRPI0721567A2 - Process for the production of para-xylene - Google Patents

Abstract

PROCESS FOR THE PRODUCTION OF PAR-XYLENE. An improved process for the selective disproportionation of toluene is disclosed. The process selectively uses a disproportionate catalyst comprising pentasyl-type zeolite such as MFI which is alloyed with aluminum phosphate. Conducting the process at a toluene conversion of greater than 30 wt% at a hydrogen-to-hydrocarbon ratio of less than 3.0, and especially a ratio of 0.1 to 1.0, optimizes the maximum yield for -xylene. Periodic rejuvenation by increasing the hydrogen-to-hydrocarbon ratio removes some carbon deposits and restores catalyst activity. An inert diluent gas also assists in selective catalyst pre-coking.

Description

BACKGROUND OF THE INVENTION This invention relates to an improved process for the conversion of aromatic hydrocarbons, such as the conversion of toluene to para-xylene. More specifically, the present invention concerns the selectivation and operation of a disproportionate process at low hydrogen levels to enable favorable coke formation and conversion.

of aromatics.

Xylene isomers are produced in large volumes from petroleum as raw materials for a variety of important industrial chemical compounds. The most important of the xylene isomers is para-xylene, the main polyester feedstock that continues to enjoy a high growth rate from extensive base demand. Ortho-xylene is used to produce italic anhydride, which has high volume markets nevertheless developed. Methylene is used in smaller but increasing volumes for such products as plasticizers, azo dyes and wood preservatives. Ethyl benzene is generally present in xylene mixtures and is occasionally recovered for styrene production, but is usually considered a less desirable component of C8 aromatics. Among aromatic hydrocarbons, the full importance of

xylenes rivals that of benzene as a feedstock for industrial chemical compounds. Neither xylenes nor benzene are produced from petroleum by naphta remediating in sufficient volume to meet demand, and conversion of other hydrocarbons is required to increase xylene and benzene yield. Often toluene is selectively disproportionate to give benzene and C8 aromatics of which the isomers of

Individual xylene are recovered.

A current goal of many petrochemical complexes and

aromatics is to increase the yield of xylenes and to de-emphasize the production of 15

benzene. Demand is growing faster for xdene derivatives than for benzene derivatives. Refinery modifications are being made to reduce the benzene content of gasoline in industrialized countries, which will increase the supply of benzene available to meet demand. A higher yield of xylenes at the expense of benzene is therefore a favorable goal, and processes for converting toluene have been>

marketed for high xylene yields.

US 4,016,219 discloses a process for disproportionate

of toluene using a catalyst comprising zeolite that has been

modified by the addition of phosphorus in an amount of at least 0,5%

in large scale. The zeolite crystals are contacted with a phosphorus compound.

to effect a reaction of the zeolite and the phosphorus compound. The zeohta

The modified material may then be inoculated into indicated matrix materials. US

No. 4,097,543 discloses disproportionate toluene for the selective production of

para-xylene using a zeolite that has undergone controlled pre-coking.

Zeolite can be ion exchanged with a variety of elements from

IB to VIII and composed of a variety of clays and other matenats of

20

matrix

25

US 4,182,923 discloses a process for toluene disproportionation with a high conversion of toluene to benzene and para-xylene by the use of a silica to alumina ratio aluminosilicate zeolite of 12 and which has been modified by hydrogen phosphate treatment. of ammonium to deposit phosphorus. US 4,629,717 releva, phosphorus-modified pupil hydrogel formed by the gelation of a homogeneous hydrosol. The composite has a relatively high surface area of 140 to 450 m2 / g and high selectivity and activity in 1–2 conversion tests.

heptene.

US 6 114 592 discloses an improved process combination for toluene disproportionation. The combination comprises selective hydrogenation of a toluene feedstock followed by a zeolitic catalyst. US 6,359,185 discloses an oil dripped zeolitic catalyst in an amorphous aluminum phosphate binder

which intensifies selectivity.

6,191,331 discloses a method of precoking

avoids a large temperature rise by using low pressure in the presence of nitrogen and in a low hydrogen-to-hydrocarbon ratio. US 6,429,347 discloses that conducting a process at a hydrogen-to-hydrocarbon ratio between 0.2 and 0.5 improves para-xylene selectivity and decreases benzene selectivity.

Aromatic disproportionate workers continue to look for processes and catalysts having exceptionally high conversion to para-xylene from toluene combined with

favorable selectivity and stability.

SUMMARY OF THE INVENTION An object of the present invention is to provide an improved process for the disproportionation of aromatic hydrocarbons. A specific objective is to obtain a maximally high yield of xylenes per

selective disproportionation of toluene.

This invention is based on the surprising discovery that operation at low hydrogen-to-hydrocarbon levels promotes increased yields of para-xylene by permitting the higher toluene conversion operation than previously used. Low hydrogen levels also improve conditioning and selective pre-coking of zeolitic catalysts when combined with a diluent.

nitrogen before use in the high conversion process.

The present invention is directed to a process for the production of xylene comprising a selective disproportionate zone under conditions comprising a toluene conversion level greater than 30 wt% and hydrogen to hydrocarbon ratio of 0.1 to 1, 0 In the disproportionate zone the current is contacted with a disproportionate catalyst under disproportionate conditions. When the disproportionate inlet temperature has increased by at least 20 ° C above the disproportionate inlet temperature, the disproportionate catalyst is rejuvenated by increasing the hydrogen-to-hydrocarbon ratio by at least 0.5. The disproportionate catalyst preferably comprises a pentasyl zeolitic aluminosilicate, more preferably MFI. This catalyst undergoes a pre-coking step prior to its use in the disproportionate zone for the purpose of increasing its para-xylene selectivity in the product beyond its equilibrium concentration.

These as well as other objectives and modalities will become

evident from the detailed description of the invention. FIG. 1 shows para-xylene yields in various ratios

hydrogen-to-hydrocarbon emissions as toluene conversion increases

on selectively precoking catalyst.

FIG. 2 shows benzene yields in various ratios of

hydrogen-to-hydrocarbon as toluene conversion increases

on selectively preheated catalyst.

DESCRIPTION OF PREFERRED EMBODIMENTS

A broad embodiment of the present invention is a selective toluene disproportionation process operating at low hydrogen-to-hydrocarbon ratio for increased para-xylene selectivity. Accordingly, a necessary element of the process is a zeolitic catalyst which has been subjected to a pre-coking step prior to its use for disproportionation with the purpose of depositing a controlled concentration of carbon on the catalyst and increasing para-selectivity. xylene. The para-xylene content of the disproportionate para-xylene rich product of the present invention is above its equilibrium concentration under disproportionate conditions.

The selective disproportionate process zone of the present invention comprises a molecular sieve and a refractory inorganic oxide. Molecular sieves are zeolitic aluminosilicates, or zeolites, which may be any of those having a SiCVAl2O3 ratio greater than 10, preferably greater than 20, and a pore diameter of 0.5 to, 8 nanometers (nm). Specific examples of zeolites that may be used are the zeolite types MFI, MEL, EUO, FER, MFS, MTT, MTW, OR, MOR and FAU. Menta, MEL, MTW and TON pentasila zeolites are preferred, and MFI type zeolites, often referred to as ZSM-5, are

especially preferred.

Preparation of the preferred MFI type zeolite is well

Known in the art. The zeolite is preferably prepared by crystallization of a mixture containing an alumina source, a silica source, an alkali metal source, water and an alkyl ammonium compound or precursor thereof.

A refractory matrix or binder (a) is used to facilitate fabrication of the disproportionate catalyst, provide strength and reduce manufacturing costs. The binder should be uniform in composition and relatively refractory to the conditions used in the process. Suitable binders include inorganic oxides such as one or more of alumina, magnesia, zirconia, chromia, titania, boria, thorium, phosphorus oxide, zinc oxide and silica. Alumina and / or silica are preferred binders. The amount of zeolite present in the bonded catalyst can vary considerably but is usually present in an amount of from 90 to mass% and preferably from 50 to 80 mass% of catalyst.

A preferred matrix or binder component is a phosphorus-containing alumina component (hereinafter called aluminum phosphate). Phosphorus may be composed of alumina in any acceptable manner known in the art. The zeolite and aluminum phosphate binder are mixed and particulate by means well known in the art such as gelation, coating, nodulating, marumerizing, spray drying, extrusion or any combination of these techniques. A preferred method of preparing the aluminum phosphate / zeolite support involves adding the zeolite to either an alumina sol or a phosphorus compound, transforming a particulate alumina / zeolite / phosphorus sol mix by the use of a oil drip method as described here below and particle calcination

spherical.

The preferred oil drip method for preparing aluminum phosphate is described in US 4,629,717 which is incorporated by reference. The technique described in the '717 patent involves the gelling of an alumina hydrosol containing a phosphorus compound using the well-known droplet drip method. Generally this technique involves preparing a hydrosol by digestion of aluminum in aqueous hydrochloric acid at reflux temperatures of 80 ° to 105 ° C. The mass ratio of aluminum to chloride in the sun ranges from 0.7: 1 to 1.5: 1. A phosphorus compound is then added to the sun. Preferred phosphorus compounds are phosphoric acid, phosphorous acid and ammonium phosphate. The relative amount of phosphorus and aluminum expressed in molar ratios ranges from 10: 1 to 1: 100, respectively, on an elemental basis. The zeolite is added to the aluminum phosphate hydrosol and the mixture is gelled. One method of gelling this mixture involves combining a gelling agent with the mixture and then dispersing the resulting combination in an oil bath or oil tower that has been heated to elevated temperatures such that gelling occurs with particulate formation. spheroidal. The gelling agents that may be used in this process are hexamethylene tetramine, urea or mixtures thereof. Gelling agents release ammonia at elevated temperatures which hardens or converts the hydrosol spheres to hydrogel spheres. The combined mixture is preferably dispersed into the oil bath as droplets from a nozzle, orifice or rotating disc. The spheres are then continuously removed from the oil bath and typically subjected to specific aging and drying treatments in oil and ammonia solution to further improve their physical characteristics. The resulting gelled and aged particles are then washed and dried at a relatively low temperature of 100 ° to 150 ° C and subjected to a calcination procedure at a temperature of 100 ° C.

450 ° to 700 ° C for a period of 1 to 20 hours.

Alternatively, the particles may be formed by spray drying the mixture. In any case, conditions and equipment must be selected to obtain small spherical particles; the particles preferably should 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 (such as oxide) in the catalyst may range from 10 to 70 mass% and preferably from 20 to 50 mass%. The aluminum phosphate matrix / binder may optimally contain smaller proportions of other inorganic oxides including, but not limited to, magnesia, beryllium, boron, silica, germanium, tin oxide, zinc oxide, titania, zirconia, vanadia, iron, chromia, cobalt oxide and the like that can be added in the

hydrosol before dripping. The aluminum-containing binder is generally amorphous, i.e. the

Binder material is essentially amorphous. Preferably

less than 10% by weight of the binder pore volume is the volume of

characteristic of crystalline material, and the highest micropore volume

preferably it is less than 5% and optimally less than 2% of pore volume. Crystalline aluminophosphate is generally an inappropriate binder material to prepare a strong, crush-resistant catalyst. Material that is not an amorphous phase is usually present as gamma-alumina; As the phosphorus content of amorphous alumina phosphate is decreased, therefore, the proportion of crystalline material is increased. The average apparent density of the spheres also varies with the phosphorus content, because a higher proportion of phosphorus decreases the average apparent density. Surface area is also controlled by phosphorus content: Oil-dripped gamma-alumina spherical particles typically have surface areas of up to 250 m2 / g, whereas aluminum phosphate spheroidal particles may have surface areas of up to 450 m2 / g. Atomic ratios of binder / matrix Al / P generally range from 1/10 to 100/1, more typically from 1/5 to 20/1, and are often between 1: 1 and 5: 1.

The catalyst may contain a metal component, preferably selected from the group consisting of gallium, rhenium and bismuth. Preferably, however, the catalyst consists essentially of a zeolitic aluminosilicate having a pore diameter of 0.5 to 0.8 nm and an aluminum phosphate binder.

Zeolitic catalyst is pre-coking

selective to increase the proportion of para-xylene in the para-xylene-rich product above equilibrium levels under disproportionate conditions. The proportion of para-xylene in the product above equilibrium levels under disproportionate conditions is generally at least 80 mass% and preferably 90 mass% or more of the C8 aromatics. Pre-coking is performed on a new or regenerated catalyst prior to use for disproportionation for a time ranging from 0.5 hours to 10 days. The catalyst may be subjected to pre-coking either in situ or ex situ in order to increase the proportion of para-xylene in the product.

of aromatics Cg. Pre-coking is performed under conditions relating to the subsequent disproportionation step comprising one or more of a higher temperature, lower pressure, and higher spatial velocity. Such pre-coking conditions comprise a pressure of 100 kPa at 4 MPa absolute, and a liquid hourly space velocity of 0.2 to 20 hr. 1. The conditions comprise one or more of an inlet temperature of at least 50 ° C. at least 100 kPa lower, or preferably not greater than half the pressure used in the subsequent disproportionation step. Lower pressure and / or lower hydrogen / hydrocarbon ratio will decrease the proportion of exothermic saturation reactions. thus restrict the temperature increase, the result should be a relatively flatter temperature profile, so a typical temperature range would be 300 ° to 700 ° C and a typical hydrogen range for coke feed would be 0.01 to 5. It is believed that the use of nitrogen or other diluent gas

Similarly inert such as methane, ethane, or propane is highly beneficial when included with hydrogen during the pre-coking phase. Such thermally inert diluent gas assists in controlling the temperature profile and is present at a molar ratio to the coke forming feed of 0.01 to 10, preferably within a ratio greater than 1. It is believed that the coke profile is Temperature affects the coking rate in various parts of the catalyst bed. An abrupt temperature gradient will therefore cause non-uniform coke deposition, and parts of the catalyst bed will thus be made selective to a different extent causing more unsatisfactory performance in subsequent disproportionate reactions. Thus a typical temperature differential across the catalyst bed during selective pre-coking would be between an increase or decrease of 10 ° C, and preferably between an increase of 3 ° C and a decrease of 4 ° C. Pre-coking results in a carbon or coke content

in a catalyst of between 5 and 40 mass% of carbon, and preferably between

and 30 wt% carbon. A coke formation feed for

Pre-coking can comprise the feed load for the

disproportionate as described herein below, such as toluene, or other

specific hydrocarbons or mixtures known in the art preferably

comprising aromatics may be used. Other details regarding pre-

coking plants are disclosed in US 4,097,543 and US 6,191,331 herein

incorporated as references. The feed charge for the present process comprises

alkyl aromatic hydrocarbons of the general formula C6H (6.n) Rn, where η ranges from 0 to 5 and R is CH3, C2H5, C3H7, or C4H9, in any combination to obtain more valuable alkyl aromatics. Suitable alkyl aromatic hydrocarbons include, for example without limitation the invention, benzene, toluene, xylenes, ethylbenzene, trimethylbenzenes, ethyl toluenes, propylbenzenes, tetramethylbenzenes, ethyl dimethylbenzenes, diethylbenzenes, methylpropylbenzenes, ethylpropylbenzenes, triethylbenzenes, diisopropylbenzenes, and mixtures thereof.

The feedstock preferably comprises toluene, optionally in combination with C9 aromatics, and suitably is derived from one or a variety of sources. Raw materials may be synthetically produced, for example, from naphtha by catalytic reforming or by pyrolysis followed by hydrogenation to give a product rich in aromatics. The feedstock may be derived from such product of adequate purity by extracting aromatic hydrocarbons from a mixture of aromatic and non-aromatic hydrocarbons and extract fractionation. For example, aromatics can be recovered from a pensioner. The reformed may be produced by any of the processes known in the art. Aromatics can then be recovered from the reformed by the use of a selective solvent, such as a sulfolane type, in a liquid-liquid extraction zone. The recovered aromatics can then be separated into streams having the desired carbon number range per fractionation. The feedstock must contain no more than 10% by weight of non-aromatics; The content of benzene and aromatics C8 is mainly an economic decision related to the toluene dilution of these aromatics. When the severity of reform or pyrolysis is high enough, extraction may be unnecessary and fractionation may be sufficient to

prepare the feed load. Within the process of disproportionate feeding

is usually first heated by indirect heat exchange against the

effluent from the reaction zone and is then further heated in a

ignited heater. The resulting vaporized current is then passed through

of a reaction zone that may comprise one or more reactors

individual. The use of a single reaction vessel having a fixed cylindrical bed

catalyst is preferred, but other reaction configurations using

moving catalyst beds or radial flow reactors may be used if

desired. Passing combined feed through reaction zone

performs the production of a vaporized effluent stream comprising

hydrogen and both product and feed hydrocarbons not

converted. This effluent is normally cooled by indirect heat exchange.

against the current entering the reaction zone and then it is additionally

cooled through the use of cooling water or air. The temperature of

effluent stream is usually lowered by sufficiently heat exchange

to condense substantially all of the hydrocarbons in the

food or product having six or more carbon atoms per

molecule. The resulting mixed phase current is passed through a

liquid vapor into which the two phases are separated and from which vapor rich in

hydrogen is recycled to the reaction zone. Separator condensate is passed through an extraction column in which substantially all C5 or lighter hydrocarbons present in the effluent are concentrated in a top stream and removed from the process. A stream rich in aromatics which is referred to herein as the disproportionate effluent stream is recovered as the extractor bottom product.

Conditions used in the disproportionate process zone typically include a temperature of from 200 ° to 600 ° C, and preferably from 350 ° to 575 ° C. The temperature required to maintain the desired degree of conversion will increase as the catalyst gradually loses activity during processing. Normal end-of-process temperatures may therefore exceed start-up temperatures.

process at 65 ° C or higher.

The disproportionate zone is generally operated in

hydrogen to hydrocarbon ratios of from 0.1 to 1.0, preferably from 0.2 to 0.5. The hydrogen-to-hydrocarbon ratio is calculated based on the free hydrogen molar ratio compared to the feedstock hydrocarbon. Periodic hydrogen-to-hydrocarbon ratio increases of at least 0.5 allow catalyst rejuvenation by soft coke hydrogenation. Preferably the hydrogen to hydrocarbon ratio during rejuvenation is within the range of 1 to 5.

The disproportionate zone is operated at moderately high pressures widely ranging from 100 kPa to 6 absolute MPa. A preferred pressure range is 2 to 3.5 MPa. The disproportionate reaction can be performed over a wide range of space velocities, with higher space velocities resulting in a larger proportion of para-xylene at the expense of conversion. Hourly liquid space velocity is usually within the range of 0.2 to 20 h-.

The disproportionate effluent stream is separated into a lightweight recycle stream, a para-xylene rich mixed aromatics product stream and a heavy aromatics stream. The para-xylene rich product may be sent to a xylene separation zone for recovery of pure para-xylene; optionally other xylenes and ethyl benzene may also be recovered as pure products. The para-xylene rich stream preferably contains para-xylene in proportion to total xylenes above their equilibrium concentration under disproportionate conditions, more preferably at least 80% by weight of para-xylene, and most preferably at least 85% by weight. paraxylene mass. The lightweight recycle stream may be diverted to other uses such as benzene and toluene recovery, but optionally a portion is recycled because it contains not only benzene and toluene but also non-aromatic amounts that would remain with benzene and reduce its value. commercial. The heavy recycle stream contains substantially all C9 and heavier aromatics and can be removed.

as a process product.

The xylene separation zone may use one or more

different separation techniques such as fractionation, crystallization or selective adsorption to recover substantially pure para-xylene from the para-xylene rich stream in the xylene separation zone. Conventional crystallization is disclosed in US 3,177,255, US 3,467,724 and US 3,662,013. Several other crystallization alternatives are discussed in US 5,329,061, incorporated by reference. In an embodiment in which the para-xylene rich product has a para-xylene content substantially above equilibrium concentration, para-xylene recovery may be conducted using only a single crystallization stage corresponding to the highest temperature purification stage. conventional crystallization

An alternative separation zone comprises a molecular sieve bed operated in accordance with the teachings of US 3,201,491 to encourage the use of a continuously mobile molecular sieve bed. Subsequent improvements are described in US 3,696,107 and US 3,626,020. Details on the operation of the xylene separation zone can be obtained from US 4,039,599 and US 4,184,943. The xylene separation zone may also incorporate an alkyl aromatic catalytic isomerization zone within the separation circuit for the purpose of displacing the ortho- and meta-xylene isomers towards para-xylene as well as to isomerize ethyl benzene to xylene if not to alkylate it to benzene. The benzene produced herein can also be sent to the transalkylation zone. The xylene separation zone may also utilize a simulated concurrent adsorptive separation process of US 4,402,832. Extract and raffinate streams may be handled as described in these references or as described in US 4,381,410 and US 5,495,061.

Variations in the process combination described above are within the scope of the invention. For example, benzene as well as toluene may be charged to the disproportionate zone as an additional feed charge. The xylene separation zone may use one or more of several known separation techniques such as adsorption, crystallization and fractionation. Orthoxylene and / or para-xylene may be recovered by one or more of such techniques as pure zone products.

xylene separation.

The disproportionate process may be carried out until

Conversion of toluene is no longer economically favorable due to catalyst decline, deterioration, or deactivation. A typical economic target occurs when the initial conversion has decreased, as measured by an inlet temperature rise of 20 ° C or greater, at which point the catalyst is rejuvenated by the increase in the free hydrogen molar ratio to feedstock hydrocarbons at at least 0.5. Accordingly, preferred rejuvenation conditions include free hydrogen present at a molar ratio to hydrocarbon feedstock of 1 to 5, an inlet temperature of 200 ° to 600 ° C, a pressure of 100 kPa to 6 MPa absolute, and a speed hourly space

0.2 to 20 hours "1.

EXAMPLES

The following examples are presented to demonstrate the present invention and to illustrate certain specific embodiments thereof. These examples are not to be construed as limiting the scope of the invention shown in the claims. There are many other possible variations, as those ordinarily skilled in the art will recognize, that are within the spirit of the invention.

EXAMPLE I

An alumina phosphate-bonded MFI catalyst was prepared to evaluate the invention. A first solution was prepared by the addition of phosphoric acid in an aqueous hexamethylene tetraamine (HMT) solution in an amount to give a final catalyst phosphorus content of 3.8 mass% and an aluminum: phosphorus atomic ratio.

1: 1 binder.

A second solution was prepared by the addition of a zeolite.

MFI type having an SVAl2 ratio of 39 to sufficient alumina sol prepared by digestion of metallic aluminum in hydrochloric acid to give a finished catalyst zeolite content of 70% by weight. These two solutions were mixed to give a homogeneous mixture of HMT, phosphorus, alumina sol, and zeolite. The mixture was dispersed as droplets in an oil bath maintained at 93 ° C. The droplets remained in the oil bath until they hardened and formed hydrogel beads having a diameter of 1.6 mm. The beads were removed from the oil bath, washed with water, air dried, and calcined at a temperature of 650 ° C. This disproportionation catalyst was used in the pre-coking and disproportionation tests described herein below. EXAMPLE II

The catalyst was then precooked under conditions comprising a temperature of 560 ° C, a pressure of 0.72 MPa and a weighted hourly space velocity of 4 (WHSV) in the presence of a hydrogen to hydrocarbon ratio of 0.5 per cent. sufficient time to make approximately 90 mol% para-xylene in total xylenes. Pure toluene disproportionation was then conducted at 2.45 MPa and WHSV 4 in the presence of pure hydrogen at varying temperatures as required reaching a range of toluene conversion levels.

Test runs were conducted in hydrogen-to-hydrocarbon ratios of 3.0, 2.0, 1.0, 0.5 and 0.2 for the purpose of illustrating the invention. FIG. 1 shows the yields of para-xylene in these hydrogen-to-hydrocarbon ratios as toluene conversion increases over the selectively precoking catalyst. FIG. 2 shows benzene yields in these hydrogen-to-hydrocarbon ratios. Surprisingly, a critical maximum para-xylene yield was found close to a 30% by weight conversion level, and this maximum shifts to even higher conversion levels as the hydrogen-to-hydrocarbon ratio drops below 3.0

Following FIG. 1, the hydrogen to hydrocarbon ratio of 1.0 appears to provide a maximum yield of para-xylene within the range of 12.5 wt% over a conversion level of 30 to 33 wt%. Also, when the hydrogen-to-hydrocarbon ratio dropped below 1.0, the maximum shifts to an even higher conversion level. This conversion shift allows even higher yields of para-xylene, which are not available at higher ratios of conversion.

hydrogen-to-hydrocarbon.

Following FIG. 2, an increase in benzene yield was observed as the toluene conversion level increased in all cases. Also, at each conversion level benzene yield decreases as the hydrogen-to-hydrocarbon ratio decreases. When the conversion level was less than 33 wt% with the hydrogen-to-hydrocarbon ratio less than 1.0, benzene yield was less than 15 wt% in all cases.

EXAMPLE III

Nitrogen addition during the selection phase was investigated by conducting a first nitrogen-free test and a test with a nitrogen-to-hydrogen ratio of 2.5 while maintaining a hydrogen-to-hydrocarbon ratio of 0. , 5 in both tests. Temperatures were maintained at 560 ° C, pressures at 0.72 MPa, and WHSV at 3 h "1. Disproportionation was subsequently performed with pure toluene at 2.45 MPa, a 4 h'1 WHSY and a hydrogen-to-hydrogen ratio. 3.0 hydrocarbon to achieve a 30% toluene conversion to

Weight.

Data obtained from the disproportionation test showed that a para-xylene to total xylene ratio of 90% by weight, the pure hydrogen selection process achieved a benzene to total xylene ratio of 1.6. However, the nitrogen selectivation procedure achieved a ratio of benzene to total xylenes of 1.3. Accordingly, it has been confirmed that the presence of an inert gas such as nitrogen during the selection procedure has a beneficial effect on reducing benzene production.

Claims (7)

Process for the production of para-xylene, characterized in that it comprises: a) disproportionate a toluene-containing feedstock by contacting the feedstock with a selectively pre-cached catalyst in the presence of an inert gas, said contact occurring at disproportionate conditions comprising free hydrogen present at a molar ratio to the feedstock hydrocarbons of 0.1 to 1.0, an inlet temperature of 200 ° to 600 ° C, a pressure of 100 kPa to 6 MPa absolute, and a liquid hourly spatial velocity of 0.2 to 20 h "1 to obtain a para-xylene-rich product containing para-xylene above its equilibrium concentration by converting more than 30% by weight of the toluene present in the b) recover para-xylene from the para-xylene rich product by one or both of adsorption and crystallization, and c) perform step (a) for a period of time until the inlet temperature begins has increased by 20 ° C or more, at which point the catalyst is rejuvenated by increasing the free hydrogen to hydrocarbon molar ratio of the feedstock by at least 0.5. Process according to Claim 1, characterized in that the conversion of toluene is 33 wt.% Or greater. Process according to Claim 1, characterized in that the conversion of toluene is 30 to 33% by weight. Process according to Claim 3, characterized in that the para-xylene rich product of step (b) further comprises benzene present in an amount not greater than 15% by weight calculated on a toluene feed base. . Process according to claim 1, characterized in that the rejuvenation conditions of step (c) further comprise free hydrogen present at a molar ratio to feed charge hydrocarbons of 1 to 5, an inlet temperature of 200 ° C. ° at 600 ° C, a pressure of 100 kPa at 6 MPa absolute, and a net hourly space velocity of 0,2 to 20 h "1. Process according to Claim 1, characterized in that the catalyst is selectively pre-coking by contacting a pentasilic zeolite selected from the group consisting of MFI, MEL, MTW, and ΤΟΝ with a coking formation feed at the feedstock. presence of a gas comprising hydrogen and an inert pre-coking diluent gas comprising an inlet temperature of 300 ° to 700 ° C, a pressure of 100 kPa to 4 MPa absolute, a free hydrogen molar ratio for forming feed 0.1 to 5, a molar ratio of inert diluent gas to coke-forming feed from 0.01 to 10, and a liquid hourly space velocity of 0.2 to 20 h "1 to deposit between 5 and 40% by weight of carbon on the catalyst and obtain a selectively precoking catalyst. Process according to Claim 6, characterized in that the inert diluent gas is selected from the group consisting of nitrogen, methane, ethane, propane, and mixtures thereof.