TWI586635B - Process for the production of xylenes - Google Patents

Description

Method for producing xylenes Cross-reference to related applications

The priority and benefit of U.S. Provisional Application Serial No. 62/018,724, filed on Jun. 30, 2014, which is incorporated herein by reference in its entirety, the entire entire entire entire entire entire entire entire entire entire entire entire entire entire entire content The related application is US Provisional Application No. 62/018,726 (with Agent Case No. 2014EM165) and U.S.S.N (with Agent Case No. 2015EM128) filed on June 30, 2014.

This invention relates to a process and apparatus for the manufacture of xylenes, and in particular to the manufacture of para-xylene.

The main source of xylenes is a catalytically reconstituted oil which is produced by contact of a petroleum brain with a hydrogenation/dehydrogenation catalyst on a support. The resulting recombinant Paraffin oil is a complex mixture of hydrocarbons heavier aromatic hydrocarbons and effective amount of a C ₆ to C ₈ aromatic addition. After removal of the light (C _5-) paraffinic hydrocarbon component, the remaining components of the oil is generally used recombinant multiple distillation steps to separate into _7-, C ₈ and C ₉₊ fraction containing the various C. The benzene can then be recovered from the C7 _- containing fraction to leave a toluene-rich fraction which is typically used to disproportionate and/or transalkylate toluene by utilizing a portion of the C9 ₊ aromatic-containing fraction. In addition, C ₈ aromatics. Containing the C ₈ fraction is fed to a circuit for producing xylenes, which is generally recovered by crystallization or adsorption on - xylene, and the resulting depletion of the through - flow catalytic conversion of p-xylene in the xylenes and the like Isomerization returns an equilibrium distribution. The resulting isomerized xylene stream can then be recycled to the para-xylene recovery unit.

Although benzene and toluene are important aromatic hydrocarbons, the demand for xylenes, especially for para-xylene, exceeds the demand for benzene and toluene and is now growing at a rate of 5-7% per year. Accordingly, there is a continuing need to develop aromatic manufacturing techniques that maximize the manufacture of para-xylene while minimizing associated capital and operating costs.

In accordance with the present invention, improved methods and apparatus for producing para-xylene and optionally benzene and/or o-xylene have been developed in which methylation is used instead of transalkylation using C9 ₊ aromatics. Toluene in the conversion of the recombinant oil or similar aromatic fraction and optionally benzene are added to the other xylenes. Methylation of toluene produces more p-xylene and less ethylbenzene than _{transalkylation} with C9 ₊ aromatics. As a result, the manufacturing and operating costs of the xylene-based separation zone can be reduced and a less expensive liquid phase process can be used for a portion of the xylene isomerization zone.

Accordingly, in one aspect, the present invention resides in a method of manufacturing of - of xylene, which comprises separating a feed stream of C ₆₊ hydrocarbons into aromatic containing at least one ilk toluene and C ₈ aromatic hydrocarbons containing a ilk and At least a portion of the toluene-containing stream is contacted with a methylating agent under conditions effective to convert toluene to xylenes and to produce a methylated effluent stream. Recovered from this ilk containing C ₈ aromatic hydrocarbons and the methylated on the effluent stream - xylene, to produce at least a pair of depletion by - xylene ilk. At least a portion of the depleted para-xylene stream is produced under conditions effective to isomerize the xylenes in the depleted para-xylene stream and to produce the first isomerized stream. Contacted with a xylene isomerization catalyst, and is effective in isomerizing xylenes and dealkylating or isomerizing ethylbenzene in the depleted para-xylene stream and making a second difference At least a portion of the depleted para-xylene stream is contacted with the xylene isomerization catalyst under vapor phase conditions of the structured stream. At least a portion of the first and second isomerized streams are then recycled to the para-xylene recovery step.

Another aspect, the present invention resides in a method of manufacturing of - of xylene, which comprises separating a feed stream of C ₆₊ hydrocarbons into aromatic containing at least one ilk toluene and C ₈ aromatic hydrocarbons containing a ilk, and At least a portion of the toluene-containing stream is contacted with the disproportionation catalyst under conditions effective to convert toluene to benzene and xylenes and to produce a disproportionate effluent stream. At least a portion of the benzene in the disproportionated effluent stream is contacted with a methylating agent under conditions effective to convert benzene to toluene and xylenes and to produce a methylated effluent stream. At least a portion of the toluene in the methylated effluent stream is recycled to the disproportionation step of the toluene.

Recovered from the effluent stream containing C ₈ aromatic hydrocarbons by disproportionation of the ilk and to - xylene, to produce at least a pair of depletion by - xylene ilk. At least a portion of the depleted para-xylene stream is reacted under conditions effective to isomerize the xylenes in the depleted xylene stream and to produce the first isomerized stream. Xylene isomerization catalyst contact, and in the isomerization of xylenes and dealkylation or isomerization of ethylbenzene in the depleted xylene stream and the manufacture of a second isomerized At least a portion of the depleted para-xylene stream is contacted with the xylene isomerization catalyst under vapor phase conditions of the stream. At least a portion of the first and second isomerized streams are then recycled to the para-xylene recovery step.

In another aspect, the invention resides in a device for producing para-xylene comprising a catalytic _{reformer for} producing a reconstituted oil stream comprising a C6 ₊ aromatic hydrocarbon, the separated oil stream being separated into a C7 _- containing aromatic hydrocarbon a first separation system comprising a stream of C8 ₊ aromatic hydrocarbons, methylating benzene and/or toluene in the C7 _- aromatic hydrocarbon-containing stream to produce a toluene of the methylated effluent stream a methylation unit, the second separation system for producing at least one depleted para-xylene stream from the C8 ₊ aromatic hydrocarbon-containing stream and the methylated effluent stream recovering para-xylene Forming a liquid phase xylene isomerization unit in the at least one depleted para-xylene stream to produce a first isomerized stream, isomerizing xylenes and dealkylating Or isomerizing ethylbenzene in the at least one depleted para-xylene stream to produce a gas phase xylene isomerization unit of the second isomerized stream, and at least a portion of the first The isomerized stream and the second isomerized stream are recycled to the recycle system of the second separation system.

11,13,14,16,18,19,22,23,24,32,34,36,38,39,41,43,44,46,48,49,51,52,55,56,111, 113,114,116,118,119,122,123,125,126,127,129,130,131,132,134,137,139,141,142,144,145,147,149,151,152, 153,155,156‧‧‧ pipeline

12,112‧‧‧ catalytic recombiner

15,115‧‧‧Reconstituted oil splitter

17,117‧‧‧ extraction area

21,121‧‧‧Benza

31,133‧‧‧methylation zone of toluene

32‧‧‧ethylbenzene removal unit

35,136‧‧‧xylene distillation tower

37,138‧‧‧Separation zone

42,143‧‧‧liquid (xylene) isomerization zone

45,146‧‧‧ gas phase isomerization zone

47,148‧‧‧Stabilizer Tower

54,154‧‧‧o-xylene tower

53‧‧‧p-xylene separation zone

124‧‧‧Selective toluene disproportionation zone

128‧‧‧BTX fractionation zone

BRIEF DESCRIPTION OF THE DRAWINGS Figure 1 is a flow chart showing a process for producing p-xylene from a catalytically reconstituted oil according to a first embodiment of the present invention.

Fig. 2 is a flow chart showing a process for producing p-xylene from a catalytically reconstituted oil according to a modification of the first embodiment of the present invention.

Figure 3 is a flow chart showing a method for producing p-xylene from a catalytically reconstituted oil according to a second specific example of the present invention.

The present invention describes a process and apparatus for producing para-xylene and optionally benzene and/or o-xylene from a recombinant oil or similar aromatic fraction, wherein methylation is used instead of using C9 ₊ aromatics. Alkylation to convert toluene and/or benzene in the reconstituted oil fraction to additional xylene. By adding methylation to the aromatic composite unit, all of the aromatic rings can be converted to p-xylene if the benzene product is not desired. Further, when a benzene product is sometimes desired, an aromatic composite device having a methylation unit can produce benzene and p-xylene when benzene is preferentially produced, and when benzene is not preferentially produced, little or No benzene is produced.

Any method known in the art for adding a methyl group to a benzene ring can be used in the methylation step of the process. However, in certain preferred embodiments, the methylation step utilizes a highly para-selective methylation catalyst, such as those used in U.S. Patent Nos. 6,423,879 and 6,504,072, the entireties of each of which are incorporated by reference. The way is incorporated herein. The catalyst comprises a molecular sieve which, when measured at a temperature of 120 ° C and a pressure of 2 to 2 dimethylbutane of 60 torr (8 kPa), is about 0.1 to 15 sec ^-1 , such as 0.5 to 10 sec ^-1 . , 2-dimethylbutane diffusion parameters. As used herein, the diffusion parameter of a particular porous crystalline material is defined as D/r ² × 10 ⁶ , where D is the diffusion coefficient (cm ² /sec) and r is the radius (cm) of the crystal. The required diffusion parameters can be derived from the absorption measurements, assuming that the flat sheet model describes the diffusion process. Thus, for a particular sorbate loading Q, Q / Q _∞ value of (where Q _∞ is the equilibrium sorbate loading) mathematically related to (Dt / r ^{2) 1/2,} where t is the sorption reached The time (sec) required for the load Q. The illustration of this flat panel model is given by J. Crank, "The Mathematics of Diffusion", Oxford University Press, Ely House, Lodon, 1967, the entire contents of which is incorporated herein by reference.

The molecular sieve utilized in this para-selective methylation process is typically a medium pore size aluminum silicate zeolite. Medium pore zeolites are generally defined as having a pore size of from about 5 to about 7 angstroms such that the zeolite readily absorbs molecules such as n-hexane, 3-methylpentane, benzene, and para-xylene. Another general definition of medium pore zeolites includes the Constraint Index test described in U.S. Patent No. 4,016,218, which is incorporated herein by reference. In this case, the medium pore zeolite has a limiting index of about 1 to 12, as measured for the zeolite alone, without introducing an oxide modifier and before any steam treatment to adjust the diffusivity of the catalyst. . Specific examples of suitable medium pore zeolites include ZSM-5, ZSM-11, ZSM-12, ZSM-22, ZSM-23, ZSM-35, ZSM-48, and MCM-22, while ZSM-5 and ZSM-11 It is especially good.

The above-mentioned iso-porous zeolite is particularly effective for this methylation process because its pore size and shape facilitate the manufacture of para-xylene rather than other xylene isomers. These zeolites of a common form have diffusion parameter values in the range of 0.1-15 sec ^{-1 as} indicated above. However, the diffusivity required for the catalyst can be achieved by rigorously steam treating the catalyst to controllably reduce the micropore volume of the catalyst to not less than the vapor-free catalyst. 50% of the micropore volume, preferably 50-90%. The reduction in micropore volume was obtained by measuring the n-hexane adsorption capacity of the catalyst at 90 ° C and a pressure of 75 torr of n-hexane before and after the vapor treatment.

The vapor treatment of the zeolite is carried out at a temperature of at least about 950 ° C, preferably from about 950 to about 1075 ° C, and most preferably from about 1000 to about 1050 ° C for from about 10 minutes to about 10 hours, preferably from 30 minutes to 5 minutes. hour.

To effect a controlled reduction in the desired diffusivity and micropore volume, the zeolite is combined with at least one oxide modifier prior to vapor treatment (such as at least one element selected from Groups 2 to 4 and 13 to 16 of the Periodic Table). Oxide) bonding may be convenient. Most preferably, the at least one oxide modifier is selected from the group consisting of oxides of boron, magnesium, calcium, strontium, and oxides of the preferred phosphorus. In some cases, the zeolite may be combined with more than one oxide modifier, such as a combination of phosphorus and calcium and/or magnesium, as it may reduce the severity of the steam treatment required to achieve the target diffusion value in this manner. . In some embodiments, the total amount of oxide modifier present in the catalyst, as measured on an elemental basis, may be between about 0.05 and about 20% by weight based on the weight of the catalyst. Preferably, and preferably between about 0.1 and about 10% by weight.

In the case where the modifier includes phosphorus, the modifier is incorporated into the Catalysts are conveniently obtained by the methods described in U.S. Patent Nos. 4,356,338, 5,110, 776, 5, 231, 064, and 5, 348, 643, the entireties of each of Treatment with a phosphorus-containing compound can be readily accomplished by contacting the zeolite (alone or in combination with a binder or matrix material) with a solution of a suitable phosphorus compound, followed by drying and calcining to convert the phosphorus to its oxidized form. . Contact with the phosphorus-containing compound is typically carried out at a temperature of about 25 ° C and about 125 ° C for a period of between about 15 minutes and about 20 hours. The concentration of phosphorus in the contact mixture can be between about 0.01 and about 30% by weight. Suitable phosphorus compounds include, but are not limited to, phosphonic acids, phosphinic acids, phosphorous acid and phosphoric acid, salts and esters of such acids, and phosphorus halides.

After contact with the phosphorus-containing compound, the porous crystalline material can be dried and calcined to convert the phosphorus to an oxide form. The calcination may be carried out in an inert gas atmosphere or in the presence of oxygen, for example, in air at a temperature of from about 150 to 750 ° C, preferably from about 300 to 500 ° C for at least one hour, preferably from 3-5 hours. Similar techniques known in the art can be used to incorporate other modifier oxidizing agents into the catalyst utilized in the alkylation process.

In addition to the zeolite and the upgrade oxide, the catalyst utilized in the methylation process can also include one or more binders or matrix materials that are resistant to the temperatures and other conditions utilized in the process. Such materials include active and passive materials such as clay, vermiculite, and/or metal oxides such as alumina. The latter may be in the form of a gelatinous precipitate or gel which is naturally occurring or comprises a mixture of vermiculite and metal oxide. Active material The use of the catalyst tends to change the conversion rate and/or selectivity of the catalyst, and is therefore generally not preferred. Passive materials are suitable for use as diluents to control the amount of conversion in a particular process such that the product can be obtained economically and orderly without the use of other measures for controlling the rate of reaction. These materials can be incorporated into naturally occurring clays such as bentonite and kaolin to improve the crush strength of the catalyst under commercial operating conditions. These materials (i.e., clay, oxide, etc.) act as binders for the catalyst. It is convenient to provide a catalyst having good crush strength because it is convenient to prevent the catalyst from breaking into a powdery material during commercial use. These clay and/or oxide binders have been utilized for the sole purpose of improving the crush strength of the catalyst.

Naturally occurring clays that can be combined with the porous crystalline material include smectite and kaolinites, including the class of bentonites, and the kaolinites commonly known as Dixie, McNamee, Georgia, and Florida clays, or others. The main mineral component is kaolinite, kaolinite, dickite, pearlite, or eucalyptus kaolinite. Such clays can be used as raw materials that were originally mined or initially subjected to calcination, acid treatment, or chemical upgrading.

In addition to the foregoing materials, the porous crystalline material may be combined with a porous matrix material (such as vermiculite-alumina, vermiculite-magnesia, vermiculite-zirconia, vermiculite-yttria, vermiculite-yttria, vermiculite- Titanium oxide and ternary compositions such as vermiculite-alumina-yttria, vermiculite-alumina-zirconia, vermiculite-alumina-magnesia, and vermiculite-magnesia-zirconia are combined.

The relative proportions of the porous crystalline material and the inorganic oxide matrix vary widely, and the content of the former ranges from about 1 to about 90% by weight and when When the composite is prepared in the form of beads, it is more often in the range of from about 2 to about 80% by weight of the composite. Preferably, the matrix material comprises vermiculite or kaolinite clay.

The methylation catalyst used in the process can be optionally co-coked. This pre-coking step can be carried out by initially loading a non-coking catalyst into the methylation reactor. Then, as the reaction proceeds, coke is deposited on the surface of the catalyst and thereafter periodically regenerated by exposure to an oxygen-containing gas atmosphere at elevated temperatures, typically in the desired range, typically about 1 to About 20% by weight and preferably from about 1 to about 5% by weight.

Methylation of toluene according to the present process can be carried out using any known methylating agent, but preferred methylating agents include methanol and/or a mixture of carbon monoxide and hydrogen.

Suitable conditions for the methylation reaction include a temperature of 350 to 700 ° C (such as 500 to 600 ° C), an absolute pressure of 100 to 7000 kPa, an hourly weight space velocity of 0.5 to 1000 hr ^-1 , and at least about 0.2 (for example, A molar ratio of toluene to methanol of from about 0.2 to about 20). The method can suitably be carried out in a fixed, mobile or fluid catalytic bed. A moving or fluidized bed configuration is preferred if it is desired to continuously control the degree of coke loading. With a mobile or fluid bed configuration, the degree of coke loading can be controlled by varying the severity and/or frequency of continuous oxidative regeneration in the catalyst regenerator. An example of a fluidized bed process suitable for methylated toluene includes injection of the methylating agent at one or more locations downstream of the toluene feed location. This method is described in U.S. Patent No. 6,642,426, the disclosure of which is incorporated herein in its entirety by reference.

Using the present method, toluene can be alkylated with methanol, by weight of at least about 75 Selectivity (based on total C ₈ aromatic product basis) percent, at least about 15% by weight per pass conversion of the aromatic and less than 1 wt% The level of trimethylbenzene is manufactured to produce para-xylene. Unreacted toluene and methylating agent and a portion of the water by-product can be recycled to the methylation reactor and the heavy by-products can be discharged to the fuel disposal point. The discharged to the fraction of C ₈ - p-xylene separation zone, generally by fractional crystallization or by selective adsorption or both to be operated by the alkylation effluent stream to recover - xylene yield and leaves the main stream containing a C ₇ and C ₈ hydrocarbons on the depletion - xylene ilk. Since the methylation unit of toluene strengthens the para-xylene content of the C ₈ fraction of the reconstituted oil, the size of the para-xylene separation zone can be reduced. This is a significant advantage because the para-xylene separation zone is one of the most expensive processes in aromatic composite plants in view of capital cost and operating costs.

After the para-xylene is recovered in the para-xylene separation zone, the residual depleted para-xylene stream is isomerized back to equilibrium and recycled back to the para-xylene separation zone. In this process, the isomerization of the depleted para-xylene stream is carried out by liquid phase and gas phase isomerization units operating in parallel and simultaneously or alternately. Thus, methylation of toluene producing almost no aromatic C ₈ stream of ethylbenzene, and C ₈ aromatic stream producing recombinant obvious ethylbenzene content. Parallel implementation of liquid phase isomerization (low operating cost but limited ethylbenzene removal) and gas phase isomerization (higher operating cost but very efficient ethylbenzene removal) gives different C ₈ aromatics The stream energy is isomerized to p-xylene at reduced operating costs.

Any liquid phase catalytic isomerization process known to those skilled in the art can be used in the liquid phase xylene isomerization unit, but a preferred catalytic system is disclosed in U.S. Patent Application Publication Nos. 2011/0263918 and 2011/0319688. The entire contents of each of them are incorporated herein by reference. Suitable conditions for the liquid phase isomerization process as used herein include temperatures from about 230 ° C to about 300 ° C and pressures from about 1300 to about 3500 kPa, which are selected such that the depleted para-xylene stream is substantially Maintain the liquid phase. In some embodiments, the hourly weight space velocity (WHSV) can be from about 0.5 to about 10 hr ^-1 .

The gas phase isomerization unit can also utilize any known isomerization catalyst system, but preferably utilizes an ethylbenzene effective to convert some or all of the spent para-xylene stream and The xylenes return to a catalytic system of equilibrium concentration. The removal of ethylbenzene can be carried out by dealkylation to benzene or by isomerization to xylenes. A preferred gas phase isomerization process is described in U.S. Patent No. 5,516,956, the disclosure of which is incorporated herein by reference. Suitable conditions for the gas phase isomerization process include a temperature of from about 660 °F to about 900 °F (350 °C to about 480 °C) and a pressure of from about 50 to about 400 psig (446 to 2860 kPa) at about 3 and about 50 hr ^- A WHSV between ¹ and a hydrogen to hydrocarbon molar ratio of from about 0.7 to about 5.

In some embodiments, it may be desirable to add a disproportionation unit of toluene upstream of the methylation unit of the toluene. The toluene disproportionation unit converts toluene in the reformed oil to benzene and xylenes (especially para-xylene) so that the size and operating cost of the toluene methylation unit can be reduced. Although any disproportionation method of toluene can be utilized, a method of selectively converting toluene into p-xylene is preferably used. This method can be utilized A catalyst comprising an aluminosilicate zeolite of intermediate pore size (such as ZSM-5) that has been selected by vermiculite and/or coke. In a preferred embodiment, the method is operable to include an initial conditioning period wherein the toluene is contacted with the meteorite-selective catalyst under conditions sufficient to increase the para-selectivity of the catalyst, followed by steady state The toluene is contacted with the catalyst under conditions sufficient to achieve a substantially fixed level of toluene conversion and para-xylene selectivity. This method is described in U.S. Patent No. 5,625,103, the disclosure of which is incorporated herein in its entirety by reference.

The invention will now be described more particularly with reference to the drawings.

1 illustrates a method of producing para-xylene according to a first embodiment of the present invention, in which a petroleum brain raw material is supplied to a catalytic reformer via a line 11 (for example, a semi-regeneration recombiner, a recycle recombiner, or a continuous catalytic recombiner). ) 12. The effluent from the catalytic reforming unit 12 is a complex mixture of aliphatic and aromatic hydrocarbons, and the pentane distilled off (not shown) after removing the C _5- fraction in the residue of the C ₆₊ fraction by Line 13 is fed to a reconstituted oil splitter 15. Hydrogen is also produced in the catalytic reformer 12 and removed via line 14 for use in the gas phase isomerization zone described below, or in a different unit of the refinery, or in a cyclohexane unit or any other petrochemical In the method, if the aromatic composite device is not located beside the refinery. Alternatively, the hydrogen can be sold as an output, or used in a fuel or burned.

In one embodiment, the reconstituted oil separator 15 (which may optionally be a dividing wall distillation column) divides the C ₆₊ fraction in line 13 into a top stream containing C _6- , a middle stream containing C ₇ , And the bottom stream containing C ₈₊ . According to a particular economic targets of the recombinant splitter overhead oil may also contain some or all of the line 13 in the presence of toluene and / or C ₈ aromatic, as well as of their non-aromatic azeotrope. In another specific embodiment (not shown), the oil recombinant splitter 15 in the line 13 C ₆₊ fraction into an overhead stream containing the C _7- C ₈₊ and a bottom stream containing the recovered omit the intermediate portion of stream . Again, depending on the particular economic goals, the C-containing overhead stream may also contain some of the _7- or all of the aromatic _{C. 8,} as well as of their non-aromatic azeotrope present in the line 13.

Returning to Fig. 1, the C _6- containing overhead stream, or in the alternative embodiment, the C _7- containing overhead stream, is sent from the recombined oil splitter 15 via line 16 to the extraction zone 17, It may be a liquid-liquid extraction method, an extractive distillation type method or a composite method thereof. The non-aromatic raffinate is removed from the extraction zone 17 via line 18 and can be used in an olefin oligomerization or reconstitution oil alkylation unit, or as a feed to a steam cracker or to a gasoline tank of the refinery. Material, or as a fuel. The raffinate can also be used as a feed to the aromatization unit to produce additional aromatic molecules while consuming hydrogen. Optionally, after pretreatment with a clay or molecular sieve catalyst to remove traces of olefins or other low concentrations of impurities, the aromatic product from extraction zone 17 is removed via line 19 and supplied to benzene column 21. The entrained water is removed from the aromatic extraction product in benzene column 21 and the benzene product is collected via line 22, typically as a side stream from the benzene column 21. The bottom product of the benzene column is rich in toluene, although it may also contain some traces of xylenes and heavier alkyl aromatics and is sent via line 23 to the methylation zone 31 of toluene. The benzene in line 22 can be recovered for sale or hydrogenation to produce cyclohexane or can be fed to the methylation zone 31 of the toluene for additional xylenes.

The toluene methylation zone 31 also receives the C7 _- intermediate stream from the reconstituted oil splitter 15 via line 32 and supplies a methylating agent, typically methanol (not shown in Figure 1) and two Methyl ether. It should be noted that a split between line 16 (the top stream containing C _6- from the reconstituted oil splitter 15) and line 32 (the intermediate stream containing C ₇ from the reconstituted oil splitter 15) can be used. The substance is effectively controlled to the level of non-aromatics sent to the methylation zone 31 of the toluene because the non-aromatic species exiting the reconstituted oil splitter 15 via line 16 to the extraction zone 17 will be shifted via line 18. except. Thus, additional flow through line 16 will reduce the overall non-aromatic content of the feed to the methylation zone 31 of the toluene.

In the toluene methylation zone 31, toluene from lines 23 and 32 is optionally reacted with benzene and methanol in line 22 from column 21 to produce xylenes and water. In some examples, also via lines 32 and 23, the C ₈ aromatic feed to zone 31 toluene methylation, ethylbenzene for removal of toluene in the methylated regions 31 become alkylated benzene In this zone 31, there may be subsequent methylation of benzene to toluene or xylenes.

The toluene may be routed through a toluene furnace and/or a heat exchange unit (not shown) prior to entering the methylation zone 31 of the toluene to evaporate and heat the toluene to the temperature required to maintain the methylation reaction. This depends on the type of catalyst used in the methylation process. Some catalysts require that the toluene be preheated to 400 ° C, while other catalysts require that the toluene be preheated to 600 ° C. The toluene can be heated to these temperatures in the heat exchanger unit and/or furnace of the process, depending on the heat sink available in the process. Toluene heated to high temperature, For example, in a furnace, the temperature at which the toluene is decomposed into coke or heavier hydrocarbons (which may affect the rate of heat transfer) can be achieved. This rate of decomposition is reduced by feeding a diluent (such as nitrogen, hydrogen, fuel gas, vapor, or a combination thereof) with the toluene upstream of the heat transfer device. The molar ratio of these diluents to toluene may be from 0.01 to more than 10. Metallurgy suitable for tube management can also be used in any of the convection or radiation zones to manage toluene decomposition, as will be appreciated by those skilled in the art. Examples include carbon steel, stainless steel, titanium, or other alloys. Special coating or application can also be used to minimize the effects of toluene decomposition and minimize coking. Additionally, additives can be used to minimize toluene coking.

The efficiency of this methylation reaction is improved as the methylating agent (typically methanol) is widely and generally dispersed throughout the reactor. The methylating agent can be introduced into the fixed bed or fluid bed reactor in a number of different ways, such as via a single injection point, several injection points, or even via a sprayer. The methylating agent can be dispersed in the reactor through a plurality of nozzles embedded in the reaction vessel or through an internal dispersion network. The number of nozzles embedded in the reactor can be one, several or many. Alternatively, the methylating agent can be introduced into the fixed bed or fluid bed via an internal diffuser. The internal diffuser can be a single injection point, some injection points, or many injection points. In the case of some or many injection points, the spreader may contain a main branch branched by one or more common manifold heads, and additional branch branches are branched by each main line to form a network of supervisors. The mains can be designed to have a uniform diameter that is the same or different from the common manifold head, or that is tapered to different diameters and lengths. One or several or many nozzles along each common manifold head or main pipe To introduce the methylating agent. The size and length of the nozzles may be similar or different depending on the desired methylation agent dispersed in the reactor. The internal spreader, main pipe, and nozzle can be insulated if used in a fluid bed or fixed bed reactor. The decision to insulate or not can change the metallurgical requirements, which can range from carbon steel or stainless steel to titanium or other types of alloys commonly used. The overall temperature of the methylation fluid and the surface temperature within the distribution network are preferably below the decomposition temperature of the methylating agent, as is known to those skilled in the art. The rate of decomposition of the methylating agent can be reduced by co-feeding a diluent such as nitrogen, hydrogen, a fuel gas, or a combination thereof. The molar ratio of these diluents to the methylating agent may range from 0.01 to greater than 10. A preferred dispensing system for the methylating agent is a fractal diffuser that contains an order of magnitude of radiation and axially distributed throughout the reaction zone. A fractal dispensing system can be designed to introduce the methanation agent axially into the interior of the reactor at the same or different rates. The axial dispersion can also be controlled to have two or more fractal dispersers, and the rate of methylating agent is controlled externally by the reactor via conventional engineering methods (i.e., valves, pumps, conditioning orifices, etc.).

The venting of the process from the methylation zone 31 of the toluene is collected by line 33 and can be used in an olefin oligomerization unit or an alkylation unit of a reconstituted oil, or can be sent to a steam cracker or The refiner is used for the recovery of olefins or as a fuel gas. The xylene-containing product from the toluene methylation zone 31 is fed via line 34 to a xylene distillation column 35, which may also receive a C8 ₊ bottoms stream from the recombination oil splitter 15 via line 24. Since the C8 ₊ bottoms stream in line 24 is heavier than the methylated product stream of toluene in line 34, the line 24 can be connected to a lower position of the xylene distillation column 35 than the line 34. on.

Prior to the xylene distillation column 35, the product stream from the toluene methylation zone 31 can be fed through a toluene distillation column (not shown) to recover unconverted toluene from the xylenes and heavier components. Fresh toluene can also be fed through the toluene distillation column. The feed point of the product stream and fresh toluene to the distillation column may be the same or different, as will be determined by those skilled in the art. In addition, there may be other streams that may be fed to the distillation column, such as from petroleum brain recombiners, xylene isomerization units, disproportionation units, transalkylation units or any other toluene and heavier aromatics. The unit of xylenes and heavier mass flows. After condensation using conventional cooling methods such as air heat sinks, water coolers or process coolers or combinations of them in parallel or series configuration, the toluene from the distillation unit is typically recovered as a liquid overhead product. The toluene is also recovered as a vapor product at the top of the distillation column, upstream of any cooling unit, or recovered from the distillation column as a side stream. Likewise, the toluene can be recovered as a liquid product from one of the trays in the distillation column (e.g., 3-5 trays below the top of the distillation column). This is particularly effective if the distillation column contains a lighter component than toluene, such as water or a light hydrocarbon, which reduces the concentration of toluene by dilution. The distillation column which separates toluene from heavier aromatics and impurities may also be a wall column having one or more separators. The recovered toluene can be recycled back to the methylation zone 31 of the toluene and the heavier components are sent downstream for further processing.

The operation of the xylene distillation column 35 to produce at least one of enriched - C ₈ aromatic xylene overhead stream, which is sent to a separation zone 37 via line 36, in which - based paraxylene product is recovered via line 38. The separation unit 37 can be based on an adsorption method or a crystallization method or any combination of the two, but in any case, it can be optimized to manage para-xylene and two different streams (ie having ~20%) Separation of the p-xylene content (C ₈ portion of the reconstituted oil), and preferably the p-xylene content of ≧ 75% (the effluent of the methylation process of toluene). This optimization will result in a substantial reduction in the size of the entire separation zone 37 and a significant savings in utility consumption. Such optimization may include feeding the para-concentrated xylene stream unrelated to the equilibrium para-xylene stream, as in U.S. Patent Nos. 8,168,845; 8,529,757; 8,481,798; 8,569,564; 8,580,120; U.S. Patent Application Serial No. 2012/ U.S. Provisional Patent Application Serial No. 61/946,052, the entire contents of each of which is incorporated herein by reference. Invariably, a small amount of toluene will be present in the xylene feed fed to the para-xylene separation zone 37. If a simulated moving bed (SMB) adsorption unit is used to recover para-xylene, a portion of the toluene present in the xylene feed will be fractionated as a "crude" toluene product, which may contain trace amounts of Xylene or water. In the absence of any treatment to remove traces of xylenes or water, this stream can be sent directly to the methylation zone 31 of the toluene because the feed to the methylation zone 31 of toluene typically contains water. Co-feed to improve the utilization of methanol and to inhibit preheating coking of the feed. In the separation zone 37, the combination of both the adsorption method and the crystallization method may include a small SMB unit (not shown) and a small crystallization unit (not shown) operated in series or in parallel, and the SMB is originally dedicated to the balance by two. The toluene stream separates p-xylene and the crystallization unit is initially dedicated to the separation of p-xylene from the concentrated p-xylene stream.

After recovery of the para-xylene, the remaining depleted para-xylene liquid phase effluent from separation zone 37 is collected via line 39 and can be fed in liquid phase via line 41 to liquid phase xylene isomerization. Zone 42 wherein the xylenes are isomerized to equilibrium. The effluent from the liquid phase isomerization zone 42 and collected in line 43 contains near equilibrium para-xylene (~24%) and is recycled to the xylene column 35. In other specific examples (not shown), the effluent from the liquid phase isomerization zone 42 can be sent directly to the separation zone 37 as long as the concentration of the heavy aromatic product over the liquid phase isomerization zone 42 is Within the specifications of the separation method used in separation zone 37. Aromatic C ₉₊ maximum allowable concentration of U.S. Patent No. 7,989,672 (the entire contents of which are incorporated herein by reference) teaches the crystallization unit, within a certain range which is also applicable to a moving bed adsorption process or a crystallization of analog and analog A composite type of moving bed adsorption method.

Alternatively, the dexylene-depleted xylenes in line 39 can be vaporized by a heater (not shown) and fed to the gas phase xylene isomerization zone 45 in the gas phase via line 44. The vapor phase effluent from the isomerization zone contains 45 to approach equilibrium of - xylene (~ 24%) and collected in line 46, and then fed to a stabilizer column 47, wherein the overhead stream containing the C ₇ lines This is removed via line 48 and the C8 ₊ bottoms stream is collected and fed to the xylene distillation column 35 via line 49. When the gas phase isomerization process used in isomerization zone 45 is of the ethylbenzene dealkylation type, the overhead in line 48 contains benzene and some toluene byproducts. When the gas phase isomerization process used in isomerization zone 45 is of the ethylbenzene isomerization type, the overhead in line 48 contains less benzene and toluene by-products. In either case, benzene can be fed to the extraction zone 17 and sold as a product or sent to a cyclohexane unit; benzene can also be treated in the methylation zone 31 of the toluene for additional xylenes. . The toluene effluent from isomerization zone 45 will be treated in the methylation zone 31 of the toluene for additional xylene production. The combined benzene/toluene stream in line 48 can be sent directly to the toluene methylation unit 31, thereby reducing fractionation costs and maximizing capital utilization.

The xylene distillation column 35 also produces a bottoms stream containing C9 ₊ hydrocarbons primarily produced in the catalytic _reformer 12 and which is collected via line 51 and conveyed via line 52 for sale, sent to solvent, and sent to The gasoline tank is sent to the fuel oil tank and/or sent to the olefin refining process. Additional fractionation facilities (not shown) may be required to optimize the configuration of the C9 ₊ bottom stream component. However, since the amount of C ₉ aromatic compound is small, the distillation column in the bottom circuit (i.e. reboiler loop) residence time of the C ₉ aromatic may be extremely long. Then, the C ₉ aromatic polymerizable upon prolonged exposure to high temperatures or high carbon condensation of the hydrocarbon components, and can block the bottom of the loop or heat exchanger. Additives can be used to control the rate of heavy polymerization or condensation. Alternatively, other sources may be added an aromatic C ₉ to the distillation column of C ₉ aromatic diluted from the methylation of toluene in the method. This C ₉ aromatic another source of which may be continuously or in a batch mode or in semi-batch mode introduced, and the accompanying methylation of toluene with aromatic C ₉ by the system continuously or in divided Flush out in batch or semi-batch mode. Any position of the additional source of C ₉ aromatic art may decide those technologies so on of the distillation column is introduced into the distillation column.

Optionally, in the case where it is desired to produce ortho-xylene, the operation of the xylene distillation column 35 is adjusted to allow a portion of the ortho-xylene and C9 ₊ hydrocarbons to be collected via line 51, and some or all of The bottom stream 51 of the xylene column can be fed to the o-xylene column 54 via line 53. The o-xylene product collected in the overhead line 56 of the o-xylene column 54 will always contain the oxygenate produced in the methylation process of the toluene. These oxygenates will generally be removed in an oxygenate removal process (not shown) prior to obtaining the o-xylene product. Any oxygenate removal method can be used, but a preferred method is described in U.S. Patent Application Serial No. 2013/0324780, the entire disclosure of which is incorporated herein by . Alternatively, the phenolic compound can be effectively removed using a caustic lotion application, as described in U.S. Patent Application Serial No. 2012/0316, the entire disclosure of which is incorporated herein by reference. The bottom heavy from the o-xylene column 54 is sent via line 55 to the gasoline tank and/or fuel sump. If excess ortho-xylene is produced in excess of the manufacturing requirements, some or all of the ortho-xylene can be treated to make more pairs when it passes over the liquid phase isomerization zone 42 or the gas phase isomerization zone 45. Xylene.

In a modification (not shown) of the method depicted in Figure 1, extracts from zone 17 containing benzene, toluene and possibly even xylenes and saturated water, traces of olefins not injected into stream 18 and / or other non-aromatic materials can be fed directly to the methylation zone 31 of toluene without prior separation in the benzene column. Therefore, the flow in line 19 is fed to the methyl group of toluene. It is not necessary to remove the water or the olefins or other non-aromatics prior to the zone 31. The benzene in stream 19 will be methylated to toluene and the toluene will be methylated to xylenes. The non-aromatics and olefins will be cleaved into light gases in zone 31, with some coke formation, and if zone 31 is a fluidized bed type reactor, the limit on the amount of non-aromatics present is economical. The decision is not limited by this method. Thus, zone 17 can be operated in a very energy efficient mode to not exclude nearly 100% of the non-aromatic materials in stream 18. The general energy savings across zone 17 can exceed 10% of normal energy costs. Additionally, olefin removal techniques are not required and avoid expensive fractionation of benzene from toluene and heavies.

Another modification of the method illustrated in Figure 1 is illustrated in Figure 2, in which like reference numerals are used to identify components that are similar to those shown in Figure 1. In particular, in the method shown in Fig. 2, it is not provided for the recovery of non-aromatic or benzene and thus the extraction zone 17 and the benzene column 21 of Fig. 1 are omitted. Thus, in this modification, after the C _-5 fraction of the recombiner effluent is removed in a pentane distillator (not shown), the effluent is fed via line 13 to a reconstituted oil splitter zone 15 The overhead stream containing C ₆ /C ₇ is separated from the bottom stream containing C ₈₊ . The top stream containing C ₆ /C ₇ is fed via line 16 to the methylation zone 31 of the toluene without a benzene extraction step, and as in the specific example of Figure 1, the bottom containing C ₈₊ The stream is fed to the xylene distillation column 35 via line 24. Another important effect of changing the stabilizer-containing overhead stream of column 47 C ₆ / C ₇ of which is recycled to the inlet line of the toluene methylation zone 31 via line 48. All of the benzene is finally converted to xylenes and the aromatic complexing process described does not produce a benzene product.

In another modification (not shown) of the process illustrated in Figure 1, the product from the methylation zone 31 of the toluene is fed to a separation drum to produce three distinct phases, including hydrocarbon streams, water, and A methanol stream and an olefin-containing exhaust stream. The separation can be carried out in one or more drums with cooling between the drums consisting of air, cooling water or some suitable coolant stream, including freezing. The drums can be horizontal or vertical or a combination thereof. The horizontal drum can contain an internal baffle. The horizontal drum can contain a water tank to collect the water phase. An internal dehumidification pad can be utilized to minimize the liquid carried by the exhaust. The vertical drum may also contain the same features as the horizontal drum, as will be appreciated by those skilled in the art. A combination of coolers can also be used to cool the flow between the drums. The coolant exchanger can also be located inside the separation drum. The separation drum can also be combined with a quench tower to save capital.

The hydrocarbon stream recovered by the separation drum can be further processed through a distillation zone (such as a distillation column of toluene and/or a distillation column 35 of xylenes) to further separate the hydrocarbons. The water/methanol stream is fed to a methanol stripper to remove hydrocarbons, methanol, and other oxygenated compounds from the water. The resulting stream containing methanol, hydrocarbons, and other oxygen containing compounds can be recycled back to the methylation zone 31 of the toluene. The water stream may contain acids such as formic acid, acetic acid or the like which lower the pH of the stream. The water stream can be neutralized by treatment with caustic, ammonia, sodium carbonate, or a neutralizing agent known to anyone skilled in the art. The wastewater stream can be treated at various locations; such as in the reactor effluent at the bottom of the methanol stripper, or at any location therebetween. The olefin-containing stream is sent to another before the final recovery of the valuable olefin component External treatment to remove contaminants.

In one embodiment, the toluene methylation zone 31 comprises a reactor, a regenerator for the catalyst, a catalyst cooler, a heat exchange device, and a gas/solids separation device. The reactor effluent may contain catalyst particles, which may be combined using gas/solids separation devices such as cyclones, centrifuges, filters, filtrates, scrubbers or even columns, tanks, precipitators, or combinations thereof. The catalyst particles are separated from the effluent stream of the reactor. The apparatus may be located inside the reaction tank, such as a cyclone or multiple cyclones, but is preferably located outside of the reactor. The gas/solids unit can be located upstream of or downstream of any heat exchange unit used to recover heat from the reactor effluent stream. The heat exchange unit includes a steam generator for producing steam having a pressure in the range of 10 psig to 1200 psig, or a heat exchange unit using heat from the reactor effluent stream to heat the treatment fluid, or a combination thereof. The flue gas from the regenerator may also contain catalyst fines which must be reduced for discharge into the atmosphere. These fines can be recovered from the flue gas using a number of different techniques including cyclones or multiple cyclones, electrostatic precipitators, washing columns, centrifuges, or combinations thereof. The flue gas solids recovery unit can be upstream or downstream of any treated heat exchange unit, such as a CO reboiler or any other heat exchange unit commonly used in the operation of flue gas. Catalyst particles recovered from the reactor effluent or the regenerator may be returned to the reaction zone, the regeneration zone, or both, either directly or indirectly (e.g., via an intermediate storage tank), or may be discharged by the system.

The catalyst can be withdrawn from the regenerator and fed to a heat exchange device also known as a catalyst cooler for removal in the regenerator by the catalyst The heat generated by the combustion of coke and other hydrocarbons. The withdrawal of the catalyst by the regenerator can be based on continuous or batch mode and with different rates. The cooled catalyst is then fed back to the regenerator. The temperature of the catalyst bed in the regenerator is controlled by controlling the flow of the catalyst through the catalyst cooler and/or the removed heat. Depending on the amount of coke to be burned in the regenerator, the catalytic converter can operate between a maximum rate and a complete shutdown. The flow rate of the catalyst withdrawn from the regenerator is used to control solids (including solids fluidized by a suitable vapor stream (venting medium) which is injected into and out of the catalyst cooler) The flow of the sliding valve or other suitable valve is controlled. The venting medium can be air, steam, nitrogen, hydrocarbons, and/or other suitable gases, which can also be injected into the catalyst cooler to ensure fluidization of the solids within the catalyst cooler, and control The heat transfer coefficient of the fluidized catalyst thus ensures proper heat transfer of the thermal catalyst to the cooling medium. The catalyst cooler can also be used to preheat the boiler feed water, produce steam of different pressures, preheat and/or vaporize the process steam, or heat the air. The catalyst cooler is typically coupled to the regenerator, to a separate structure for support, or to be fully or partially enclosed (stuck) inside the regenerator tank.

Figure 3 illustrates a kind of a second manufacturing embodiment of the present invention in particular - of xylene, added selective toluene disproportionation (the STDP) region, to C ₇ containing recombinant oil fractions of toluene conversion is at least partially into a two Toluene, thereby allowing the size of the methylation zone of the toluene to be reduced. Thus, with reference to Figure 3, the petroleum brain feedstock is supplied via line 111 to a catalytic reformer (e.g., a semi-regenerative recombiner, a recycle recombiner, or a continuous catalytic reformer) 112. After removing the co-produced hydrogen via line 113 and after removing the C _5- fraction in the pentane distillator (not shown), the remaining C ₆₊ effluent from the recombiner 112 is Line 114 is fed to a reconstituted oil splitter 115. In a particular embodiment as those in the FIG. 1, the oil recombinant splitter 115 (which may be optionally wall of the distillation column) in the C ₆₊ fraction of the line 114 into an overhead stream containing the C _6-, C ₇ containing The middle stream and the bottom stream containing C ₈₊ .

The C ₆ -containing overhead stream from the reconstituted oil splitter 115 is sent via line 116 to extraction zone 117, which may be a liquid-liquid extraction process, an extractive distillation process or a composite process thereof. The non-aromatic raffinate is removed from extraction zone 117 via line 118, leaving an aromatic product stream that is optionally pretreated with clay or molecular sieve catalyst to remove traces of olefins or other low concentration impurities. Thereafter, it is supplied to the benzene column 121 via a line 119. The trapped water is removed from the aromatic extract product in benzene column 121 and the benzene-rich stream is removed via line 122 as a side stream from benzene column 121. The bottom product of the benzene column is rich in toluene, although it also contains some traces of xylenes and heavier alkyl aromatics, and is sent via line 123 to the disproportionation zone 124 of the selective toluene.

The selective disproportionation of toluene (the STDP) region 124 can be received by the catalytic reformer 112 via line 125 containing an intermediate portion of the stream C _7, and as discussed above, is operated to selectively converted into benzene and toluene Para-xylene. Typically, the effluent from the selective disproportionation of toluene in the region of 124 - C ₈ component comprising p-xylene concentration in the effluent is greater than 90% by weight of. The exhaust from the selective toluene disproportionation zone 124 is collected via line 126 and can be used as a fuel gas, or as a feed to a steam cracker or can be in the methylation zone of toluene described below deal with.

The effluent from the selective toluene disproportionation zone 124 is supplied via line 127 to the BTX fractionation zone 128, wherein at least a portion of the unreacted toluene is separated from the effluent and recycled to the STDP zone 124 via line 129. The BTX is also removed from the fractionation zone 128 in line 131 STDP effluent containing C in the overhead stream ₆ leaving the bottom of the C ₈₊ cut, it is collected and discharged from the BTX fractionation zone 128 via line 132. As shown in Figure 3, the BTX fractionation zone 128 can be a dividing wall distillation column.

In line 131 the ₆ C-containing overhead stream and the benzene-rich stream in line 122 is fed to the toluene methylation zone 133, along with methanol supply (not shown). In the toluene methylation zone 133, benzene from lines 122 and 131 is reacted with methanol to produce toluene, xylenes and water, while any toluene present in these streams will be converted to additional xylenes. . The methylated effluent is removed from the methylation zone 133 of toluene via line 134 and recycled to the BTX fractionation zone 128, wherein the toluene is separated from the xylenes for recycle to the STDP zone 124, At the same time the xylenes are collected as part of the C8 ₊ bottoms fraction in line 132. Thus, in this particular example, it will be seen that the toluene methylation zone 133 is primarily used for the methylation of benzene and to produce toluene and high para-xylene content of xylenes. Thus, the top stream 122 of the benzene column is treated in the methylation zone 133 of the toluene while the bottom stream 123 of the benzene column is treated in the STDP zone 124. Alternatively, when biased towards benzene production, some or all of the benzene-rich stream in line 122 or the benzene-rich stream in line 131 can be recovered for sale or hydrogenation to produce cyclohexane.

The process venting from the methylation zone 133 of toluene is collected by line 135 and, as in the previous specific examples, may be used in an olefin oligomerization unit or a reconstituted oil alkylation unit, or may be sent to a vapor A cracker or refiner for olefin recovery or as a fuel gas.

As noted above, the exhaust from the STDP zone 124 collected by line 126 can be treated in the methylation zone 133 of the toluene if, for example, the methylation zone 133 of toluene utilizes a fluid bed or moving bed reactor. For the fluid bed unit, the exhaust gas can be used as a diluent for benzene/toluene and/or methanol instead of steam. Additionally, exhaust from zone 124 can be used as flushing gas in zone 133 instead of steam or nitrogen or hydrogen typically used to improve fluid bed transportability. If the off-gas in line 126 contains heavier (C4 ₊ ) hydrocarbons, these will be cleaved in zone 133 into lighter olefins and paraffinic products. The olefin can then be recovered in a downstream facility such as a steam cracker.

The C8 ₊ bottoms fraction from the BTX fractionation zone 128 is fed via line 132 to a xylene distillation column 136, which is also received by the recombination oil splitter 115 via line 130 for the C8 ₊ bottoms stream. Because the C8 ₊ bottoms stream in line 130 is heavier than the C8 ₊ bottoms fraction in line 132, the line 130 can be connected to the lower portion of the xylene distillation column 136 than the line 132. The xylene distillation column 136 is enriched in at least one manufacturing operation - C ₈ aromatic overhead xylene stream, sent to a separation zone which line 138 via line 137, in which - based paraxylene product is recovered via line 139. The separation zone 138 can be based on an adsorption process or a crystallization process or any combination of the two, and can operate in the same manner as the separation zone 37 of FIG.

After recovery of the para-xylene, the remaining depleted para-xylene liquid phase stream from the separation zone 138 is collected in line 141 and fed to the liquid phase xylenes via liquid line 142 via line 142. Isomerization zone 143 in which the xylenes are isomerized to equilibrium. The effluent collected from the liquid phase isomerization zone 143 contains near equilibrium para-xylene (~24%) and is recycled via line 144 to the xylene distillation column 136.

Alternatively, the depleted paraxylene in line 141 can be vaporized by a heater (not shown) and fed to the vapor phase xylene isomerization zone 146 via gas line 145. Again, the vapor phase effluent from the isomerization zone 146 containing the balance of the proximity of - xylene (~ 24%) and are collected in line 147, is then fed to a stabilizer column 148, wherein the C-containing overhead stream ₇ This is removed via line 149 and the C8 ₊ bottoms stream is collected and fed to the xylene distillation column 136 via line 151. The flow in line 149 can be treated in the same manner as the equivalent flow in line 48 of the specific example of FIG.

The xylene distillation column 136 also produces a bottoms stream containing C9 ₊ hydrocarbons primarily produced in the catalytic reformer 112 and which are collected via line 152 and sent to a point of sale, to a solvent, to a gasoline tank. And/or fuel oil tanks. Optionally, in the case where it is desired to produce ortho-xylene, the operation of the xylene distillation column 136 is adjusted to allow a portion of the ortho-xylene and C9 ₊ hydrocarbons to be collected via line 152, and partially The or all of the xylene bottoms stream can be fed via line 153 to the o-xylene column 154 where the o-xylene product is collected in overhead line 155. The bottom heavy material from the o-xylene column 154 is sent via line 156 to the gasoline tank and/or fuel sump. If more than ortho-xylene is produced in excess of the manufacturing requirements, some or all of the ortho-xylene may be treated to cross the liquid phase isomerization zone 143 or the vapor phase isomerization zone 146 to produce more pairs of Toluene.

The invention will now be more particularly described with reference to the following non-limiting examples.

[Example]

This simulation example illustrates the alkylation of toluene using a methanol unit rather than a transalkylation unit, based on a general aromatic-based composite unit in which the xylenes are produced in the recombination and transalkylation zone. The aromatic composite unit of the same raw material has a minimal effect on the total output of para-xylene. In this example, it is assumed that all of the xylenes are converted to p-xylene (manufactured without o-xylene). The results are shown in Table 1 below.

In Table 1, each of the aromatic composite devices utilizes the same raw materials (1265.3 kTa of petroleum brain) qualitatively and quantitatively. In addition, the recombination zone provides the same product composition in all cases, and the product composition is listed in column #1 entitled "CCR Recombinant Oil". Column #2, titled "(only) xylene recovery", shows the production of para-xylene, if only the recombiner xylenes are recovered (no transalkylation unit). Column #3, entitled "Xylene Recovery and Transalkylation", shows the manufacture of para-xylene in a general aromatic composite unit in which a transalkylation unit is added to produce additional xylenes. Column #4, entitled "Recovery of Dylene by TAM (No Transalkylation)", shows the production of para-xylene from an aromatic composite unit in which the alkylation of toluene using a methanol unit has been added but the transalkylation is omitted. Base unit.

As seen and based on the same raw material and recombination zone output, the p-xylene production for a general aromatic composite unit is 560.9 kTa, and is also used for an alkyl group having toluene using a methanol unit without a transalkylation unit. The p-xylene produced by the composite unit was 526.3 kTa. This means that the para-xylene production is essentially the same, but the manufacturing cost is substantially reduced because (1) the high para-xylene content in the methylated effluent of the toluene significantly reduces the size of the separation zone and (2) The methylation of toluene using very little ethylbenzene in the effluent allows the use of liquid phase isomerization techniques which significantly reduce the overall cost of the xylene isomerization process. In addition, the production of para-xylene is often advantageous over benzene production due to the higher marginal benefit. Benzene can be fed to the transalkylation zone for further manufacture of xylenes, but this is limited by the ratio of methyl to ring. However, when the methylation zone of toluene is available, all of the benzene can be converted to xylenes and further turned as needed. It turns into p-xylene. Thus, in the case of column #4, an additional 83.6 kTa of benzene can be used for further para-xylene production.

While the invention has been described and illustrated with reference to the specific embodiments thereof, However, for this reason, the scope of the appended claims should be limited only to determine the true scope of the invention.

18,19,22,23,24,32,33,34,36,38,39,41,44,51,52,55,56‧‧‧ pipeline

21‧‧ Benzene

31‧‧‧methylation zone of toluene

35‧‧‧ xylene distillation tower

37‧‧‧Separation zone

42‧‧‧liquid (xylene) isomerization zone

53‧‧‧p-xylene separation zone

54‧‧‧o-xylene tower

Claims (9)

A method of manufacturing of - of xylene, the method comprising: (a1) a feed stream comprising an aromatic C ₆₊ hydrocarbons into at least one of ilk containing toluene and C ₈ aromatic hydrocarbons containing ilk; (B1) Contacting at least a portion of the toluene-containing stream in the extraction zone to remove non-aromatic species present in at least a portion of the stream and to produce an extracted toluene-containing stream; (c1) efficiently converting toluene to xylenes and producing via effluent stream under conditions of methylation of at least some portion of toluene with a methylating agent in contact ilk containing; (D1) from the effluent containing C ₈ aromatic hydrocarbons and the ilk stream recovered by methylation of Para-xylene to produce at least one depleted para-xylene stream, (e1) in the isomerization of the xylenes in the depleted para-xylene stream and to produce the first isomerization At least a portion of the at least one depleted para-xylene stream is contacted with a xylene isomerization catalyst; (f1) in an isomerized xylene group and dealkylated or Isomerization of the ethylbenzene in the depleted para-xylene stream and the production of a second isomerized stream under gas phase conditions At least a portion of the at least one depleted para-xylene stream is contacted with the xylene isomerization catalyst; and (g1) recycling at least a portion of the first and second isomerized streams to (c1) . The method according to Claim 1 patentable scope, wherein the feed stream (a1) contained in the mixture by the aromatic and aliphatic C ₆₊ hydrocarbons produced of C _5- hydrocarbons removed from the recombinant oil flow. The method of claim 1, wherein the separating (a1) also produces a stream comprising benzene. The method of claim 1, wherein the separating (a1) is carried out by a dividing wall distillation column. The method of claim 3, wherein at least a portion of the benzene-containing stream is supplied to the contact (b1) and then to the contact (c1). The method of claim 1, wherein the methylating agent comprises methanol. The method of claim 1, wherein the contacting (c1) is carried out in the presence of a catalyst comprising a temperature of 120 ° C and a pressure of 2 to 2 dimethylbutane of 60 torr (8 kPa). When measured, there was a porous crystalline material having a diffusion parameter of 2,2-dimethylbutane of about 0.1 to 15 sec ^-1 . The method of claim 7, wherein the porous crystalline material comprises ZSM-5 which has been pretreated with steam at a temperature of at least 950 °C. The method according to Claim 1 patentable scope, and further comprising: (h1) recovering at least one of the ortho-containing C ₈ aromatic hydrocarbons and the ilk by methylation of the effluent stream - xylene.