Process for isomerizing C8 aromatic hydrocarbons

Abstract

A liquid phase isomerization process comprising co-feeding molecular hydrogen at a feed rate of ≥ 100 ppm by weight, even for ≥ 5 hours ‑1 At a high WHSV of , the C8 aromatic hydrocarbon isomerization feed can be efficiently converted in the presence of an isomerization catalyst at a very low catalyst deactivation rate.

Description

[0001] CROSS-REFERENCE TO RELATED APPLICATIONS

[0002] This application claims priority to and the benefit of USSN 62 / 890,935 filed on August 23, 2019, and European Patent Application No. 19209114.8 filed on November 14, 2019, the disclosures of which are incorporated herein by reference. Technical Field

[0003] The present disclosure relates to the isomerization of C8 aromatic hydrocarbons. Specifically, the present disclosure relates to the isomerization of C8 aromatic hydrocarbons under conditions where most of the C8 aromatic hydrocarbons are present in the liquid phase. The present disclosure can be used, for example, in a method for producing para-xylene. Background Art

[0004] A high purity p-xylene product is typically produced by separating p-xylene from a p-xylene-rich aromatic hydrocarbon mixture comprising p-xylene, o-xylene, m-xylene, and optionally ethylbenzene ("EB") in a p-xylene separation / recovery system. The p-xylene recovery system may comprise, for example, a crystallizer and / or an adsorption chromatography separation system as known in the art. The p-xylene-depleted stream produced from the p-xylene recovery system (the "filtrate" from the crystallizer upon separation of p-xylene crystals, or the "residue" from the adsorption chromatography separation system, collectively referred to as "residue" in the present disclosure) is rich in m-xylene and o-xylene, and contains a p-xylene concentration that is typically lower than its concentration in an equilibrium mixture consisting of m-xylene, o-xylene, and p-xylene. In order to increase the yield of p-xylene, the residue stream may be fed to an isomerization unit, in which the xylenes undergo an isomerization reaction when contacting an isomerization catalyst system, thereby producing an isomerized effluent that is rich in p-xylene compared to the residue. At least a portion of the isomerized effluent may be recycled to the para-xylene recovery system, forming a "xylene loop", after optional separation and removal of lighter hydrocarbons that may be produced in the isomerization unit.

[0005] Xylene isomerization can be carried out under conditions where the C8 aromatics are substantially in the vapor phase in the presence of an isomerization catalyst (vapor phase isomerization or "VPI").

[0006] Newer production technologies have been developed to allow xylene isomerization at significantly lower temperatures in the presence of an isomerization catalyst, with the C8 aromatics being essentially in the liquid phase (liquid phase isomerization or "LPI"). Use of LPI can reduce the number of phase changes (liquid to vapor / vapor to liquid) required to process the C8 aromatic feed compared to conventional VPI. This provides the process with a sustainability advantage in the form of significant energy savings. For any paraxylene production facility, it would be highly advantageous to arrange an LPI unit as a supplement to or in place of a VPI unit. For existing paraxylene production facilities that lack an LPI, it would be highly advantageous to add an LPI unit to supplement the VPI unit or to replace the VPI unit.

[0007] Exemplary LPI processes and catalyst systems useful therefor are described in U.S. Patent Application Publication Nos. 2011 / 0263918 and 2011 / 0319688, 2017 / 0297977, 2016 / 0257631, and U.S. Patent No. 9,890,094, all of which are incorporated herein by reference in their entirety. In the LPI processes described in these references, MFI framework type zeolites (e.g., ZSM-5) are typically used as catalysts.

[0008] Due to the many advantages of the LPI process and the need to employ this technology, there is also a need for improvements in this technology. It has been found that LPI catalysts can experience aging over time at various deactivation rates, especially at high weight hourly space velocities ("WHSV"). It would be highly desirable to reduce catalyst deactivation in the LPI process to achieve long catalyst service life and therefore reduce catalyst regeneration and replacement costs.

[0009] The present disclosure satisfies this need and other needs.

[0010] Overview

[0011] It has been found that co-feeding molecular hydrogen with the liquid phase of the C8 aromatic hydrocarbon feed at a high feed rate of at least 100 ppm can significantly reduce the deactivation of the isomerization catalyst in the LPI process, even over a period of at least 5 hours. -1 And even at least 20 hours -1 Under high WHSV.

[0012] Thus, a first aspect of the present disclosure is directed to a method comprising: feeding molecular hydrogen and a liquid phase isomerized hydrocarbon feed comprising C8 aromatic hydrocarbons into an isomerization reactor having an isomerization catalyst disposed therein, wherein the molecular hydrogen is fed at a feed rate of 100 ppm to 5000 ppm by weight, based on the total weight of the isomerized hydrocarbon feed; and contacting the molecular hydrogen and the C8 aromatic hydrocarbons with the isomerization catalyst in the isomerization reactor under isomerization conditions to produce an isomerization effluent, wherein the isomerization conditions include a reaction pressure in the isomerization reactor of 1,700 kPa-gauge to 3,500 kPa-gauge and a reaction temperature such that the C8 aromatic hydrocarbons are substantially in a liquid phase in the isomerization reactor, and a weight hourly space velocity of 5 hours per hour. -1 -20 hours -1 .

[0013] A second aspect of the present disclosure relates to a method for producing paraxylene, the method comprising: feeding molecular hydrogen and a liquid phase isomerized hydrocarbon feed comprising C8 aromatic hydrocarbons into an isomerization reactor having an isomerization catalyst disposed therein, wherein the molecular hydrogen is fed at a feed rate of 100 ppm to 5000 ppm by weight, based on the total weight of the isomerized hydrocarbon feed; and contacting the molecular hydrogen and the C8 aromatic hydrocarbons with the isomerization catalyst in the isomerization reactor under isomerization conditions to produce an isomerization effluent, wherein the isomerization conditions include a reaction pressure of 1,700 kPa-gauge to 3,500 kPa-gauge and a reaction temperature of 200 to 300° C. such that the C8 aromatic hydrocarbons are substantially in a liquid phase in the isomerization reactor, and a weight hourly space velocity of 5 hours per hour. -1 -20 hours -1 ; the isomerization effluent comprises para-xylene at a higher concentration than the isomerization hydrocarbon feed; and recovering at least a portion of the para-xylene from the isomerization effluent.

[0014] BRIEF DESCRIPTION OF THE DRAWINGS

[0015] Figure 1 is a graph showing para-xylene / xylene as a function of reaction temperature at various WHSVs in a simulated LPI process of the present disclosure.

[0016] Figure 2 is a graph showing xylene loss as a function of reaction temperature at various WHSVs in a simulated LPI process of the present disclosure.

[0017] Figure 3 is a graph showing the relative activity loss of an isomerization catalyst as a function of reaction temperature at a WHSV of 2.5 hr-1 in a simulated LPI process of the present disclosure.

[0018] Figure 4is a graph showing the activity loss of the isomerization catalyst as a function of para-xylene / xylenes at various temperatures in a simulated LPI process of the present disclosure.

[0019] Figure 5 is a schematic illustration of the aromatics production complex from naphtha reforming to produce para-xylene product.

[0020] Figure 6 is a graph of the para-xylene concentration in the isomerization effluent as a function of the cumulative grams per gram of isomerization catalyst fed ("CGpGC") for an LPI process operating at various isomerization conditions, indicating various catalyst deactivation rates, particularly at various feed rates of molecular hydrogen and various WHSVs.

[0021] Figure 7 is a graph showing the para-xylene concentration in the isomerization effluent of a comparative LPI process as a function of WHSV, which is operated under isomerization conditions including bubbling molecular hydrogen into the isomerization hydrocarbon feed at 9 ppm and feeding a C8 aromatic hydrocarbon feed containing ethylbenzene at a low concentration of about 1 weight %, indicating rapid catalyst deactivation in the presence of low molecular hydrogen concentrations in the isomerization hydrocarbon feed.

[0022] Figure 8 is a graph showing the concentration of para-xylene (wt%) in the isomerization effluent of another comparative LPI process as a function of run time, which is -1 Operating at a WHSV of 1.33 WHSV but without co-feeding to the LPI reactor, where the isomerized hydrocarbon feed contained a high concentration of 9 wt% p-xylene, indicates rapid catalyst deactivation in the absence of co-fed molecular hydrogen, even though it is presumed that the isomerized hydrocarbon feed is unlikely to cause catalyst deactivation.

[0023] Details

[0024] Various specific embodiments, variations and examples of the present invention will now be described, including preferred embodiments and definitions adopted for the purpose of understanding the claimed invention. Although the following detailed description gives specific preferred embodiments, those skilled in the art will appreciate that these embodiments are exemplary only, and that the present invention may be implemented in other ways. For the purpose of determining infringement, the scope of the present invention will refer to any one or more of the appended claims, including their equivalents and elements or limitations equivalent to those recorded. Any reference to the "invention" may refer to one or more of the present inventions defined by the claims, but not necessarily all.

[0025] In the present disclosure, a method is described as comprising at least one "step". It should be understood that each step is an action or operation that can be performed one or more times in the method in a continuous or discontinuous manner. Unless otherwise specified or the context clearly indicates otherwise, multiple steps in the method can be performed sequentially in the order in which they are listed, overlapping or not overlapping with one or more other steps, or in any other order, as the case may be. In addition, one or more or even all steps can be performed simultaneously with respect to the same or different batches of materials. For example, in a continuous method, when the first step in the method can be performed with respect to the raw material just fed to the beginning of the method, the second step can be performed simultaneously with respect to the intermediate material, which is produced by processing the raw material fed to the method at an earlier time in the first step. Preferably, the steps are performed in the order described.

[0026] Unless otherwise indicated, all numbers indicating quantities in this disclosure are understood to be modified by the term "about" in all cases. It should also be understood that the precise numerical values ​​used in the specification and claims constitute specific embodiments. Efforts have been made to ensure that the data in the examples are accurate. However, it should be understood that due to the limitations of the technology and equipment used to make the measurements, any measured data inherently contain a certain level of error.

[0027] As used herein, the indefinite article "a" or "an" should mean "at least one", unless specified to the contrary or the context clearly indicates otherwise. Therefore, unless specified to the contrary or the context clearly indicates that only one fractionation column is used, embodiments using "a fractionation column" include embodiments using one, two or more fractionation columns.

[0028] As used herein, "consisting essentially of" means that the composition, feed or effluent comprises a given component in a concentration of at least 60 weight %, preferably at least 70 weight %, more preferably at least 80 weight %, more preferably at least 90 weight %, and even more preferably at least 95 weight %, based on the total weight of the composition, feed or effluent in question.

[0029] "Substantially completely" means at least 95% by weight, preferably ≥98% by weight, preferably ≥99% by weight, preferably ≥99.5% by weight, preferably ≥99.9% by weight. Thus, when the C8 aromatic hydrocarbons are substantially completely present in the reactor in the liquid phase, at least 95% by weight, preferably ≥98% by weight, preferably ≥99% by weight, preferably ≥99.5% by weight, preferably ≥99.9% by weight of such C8 aromatic hydrocarbons are present in the reactor in the liquid phase. When molecular hydrogen is substantially completely dissolved in the isomerized hydrocarbon feed in the liquid phase, at least 95% by weight, preferably ≥98% by weight, preferably ≥99% by weight, preferably ≥99.5% by weight, preferably ≥99.9% by weight of such molecular hydrogen is dissolved in the isomerized hydrocarbon feed in the liquid phase.

[0030] The term "hydrocarbon" means (i) any compound consisting of hydrogen and carbon atoms or (ii) any mixture of two or more such compounds of (i). The term "Cn hydrocarbon", where n is a positive integer, means (i) any hydrocarbon compound containing carbon atoms (one or more) in a total number of n in its molecule, or (ii) any mixture of two or more such hydrocarbon compounds of (i). Thus, C2 hydrocarbons can be ethane, ethylene, acetylene, or any mixture of at least two of them in any proportion. "Cm to Cn hydrocarbons" or "Cm-Cn hydrocarbons", where m and n are positive integers and m < n, means any one of Cm, Cm+1, Cm+2, …, Cn-1, Cn hydrocarbons, or any mixture of two or more of them. Thus, "C2 to C3 hydrocarbons" or "C2-C3 hydrocarbons" can be any one of ethane, ethylene, acetylene, propane, propylene, propyne, allene, cyclopropane, and any mixture of two or more of them in any proportion between the components. "Saturated C2-C3 hydrocarbons" can be ethane, propane, cyclopropane, or any mixture of two or more of them in any proportion. "Cn+ hydrocarbons" means (i) any hydrocarbon compound containing carbon atoms (one or more) in a total number of at least n in its molecule, or (ii) any mixture of two or more such hydrocarbon compounds of (i). "Cn- hydrocarbons" means (i) any hydrocarbon compound containing carbon atoms in a total number of at most n in its molecule, or (ii) any mixture of two or more such hydrocarbon compounds of (i). "Cm hydrocarbon stream" means a hydrocarbon stream consisting essentially of Cm hydrocarbon(s). "Cm-Cn hydrocarbon stream" means a hydrocarbon stream consisting essentially of Cm-Cn hydrocarbon(s).

[0031] "Light hydrocarbon" means any C5- hydrocarbon in the present disclosure.

[0032] "Liquid-phase isomerization" and "LPI" are interchangeably meant to be a method for isomerizing C8 aromatic hydrocarbons in the presence of an isomerization catalyst in an isomerization zone, where xylene isomerizes under such isomerization conditions that the C8 aromatic hydrocarbons present in the isomerization zone are substantially in the liquid phase. "Substantially in the liquid phase" means ≥80 wt%, preferably ≥85 wt%, preferably ≥90 wt%, preferably ≥95 wt%, preferably ≥99 wt%, preferably all in the liquid phase. Such isomerization conditions are called liquid-phase isomerization conditions.

[0033] "Gas phase isomerization" and "VPI" interchangeably mean a C8 aromatics isomerization process in the presence of an isomerization catalyst in an isomerization zone, wherein the xylenes are isomerized under isomerization conditions such that the xylenes present in the isomerization zone are substantially in the vapor phase. "Substantially in the vapor phase" means ≥ 90 wt%, preferably ≥ 95 wt%, preferably ≥ 99 wt%, preferably all in the vapor phase. Such isomerization conditions are referred to as vapor phase isomerization conditions.

[0034] As used herein, "wt%" means percent by weight, "vol%" means percent by volume, "mol%" means percent by mole, "ppm" means parts per million and "ppm wt" and "wppm" are used interchangeably, meaning parts per million on a weight basis. All concentrations herein are expressed based on the total amount of the composition in question. Unless otherwise specified or indicated, all ranges expressed herein shall include both endpoints as two specific embodiments.

[0035] The nomenclature of elements and element groups used herein is based on the periodic table used by the International Union of Pure and Applied Chemistry after 1988. An example of a periodic table is shown on the inside cover of Advanced Inorganic Chemistry, 6th edition, by F. Albert Cotton et al. (John Wiley & Sons, Inc., 1999).

[0036] The method of the present disclosure can be used to prepare a p-xylene product containing a high concentration of p-xylene, and / or an o-xylene product containing a high concentration of o-xylene, and / or an m-xylene product containing a high concentration of m-xylene. Considering the higher economic value of p-xylene and o-xylene relative to m-xylene, the method of the present disclosure is preferably used to prepare p-xylene and / or o-xylene products, more preferably p-xylene products. In a conventional C8 aromatic hydrocarbon isomerization process, a C8 aromatic hydrocarbon feed containing a non-equilibrium concentration of xylene is contacted with an isomerization catalyst in an isomerization reactor under isomerization conditions to achieve a time period for the conversion of xylenes to produce an isomerization effluent leaving the isomerization reactor containing a concentration of xylene closer to their equilibrium concentration. Therefore, for the purpose of producing a para-xylene product, the C8 aromatic hydrocarbon feed added to the isomerization reactor generally contains (i) para-xylene, whose concentration (based on the total amount of xylenes in the isomerized hydrocarbon feed) is lower than its equilibrium concentration in the xylene mixture, and (ii) ortho-xylene and meta-xylene, whose total concentration (based on the total amount of xylenes in the isomerized hydrocarbon feed) is significantly higher than their total equilibrium concentration. When contacting the isomerization catalyst under isomerization conditions in the isomerization reactor, a portion of meta-xylene and / or ortho-xylene is converted into para-xylene to produce an isomerization effluent containing a higher concentration of para-xylene (based on the total amount of xylenes in the effluent). It is highly desirable that the para-xylene concentration (based on the total amount of xylenes in the isomerization effluent) is close to its equilibrium concentration in the xylene mixture. The para-xylene in the effluent can be recovered to produce a para-xylene product. "Xylene mixture" means a mixture consisting of para-xylene, meta-xylene and ortho-xylene.

[0037] The process of the present disclosure includes an LPI step in which a large amount of molecular hydrogen is co-fed into the isomerization reactor. It has been surprisingly found that the presence of such a large amount of molecular hydrogen can significantly reduce the deactivation rate of the LPI catalyst, even at ≥5.0, ≥7.5, ≥10, ≥12.5, ≥15, ≥17.5, and even ≥20 hours. -1 Under high WHSV.

[0038] Various embodiments of the process of the present disclosure include introducing a feed comprising C8 hydrocarbons and molecular hydrogen into a reactor having a xylene isomerization catalyst disposed therein, wherein the isomerized hydrocarbon feed comprises from 100 ppm to 5000 ppm by weight of molecular hydrogen, based on the total weight of the isomerized hydrocarbon feed, and reacting the isomerized hydrocarbon feed in a substantially liquid phase in the presence of the xylene isomerization catalyst to produce an isomerization effluent having an increased para-xylene concentration relative to the para-xylene concentration in the isomerized hydrocarbon feed, wherein the isomerized hydrocarbon feed is heated to 100 ppm at a gage pressure of from about 1,700 kPa to about 3,500 kPa at a weight hourly space velocity of about 5 hours. -1 About 20 hours -1Xylene isomerization catalyst in contact.

[0039] Isomerized hydrocarbon feed containing C8 aromatic hydrocarbons in the process of the present disclosure

[0040] The isomerization hydrocarbon feedstock containing C8 aromatic hydrocarbons fed to the isomerization reactor can contain one or more of xylenes at various concentrations. For example, the total concentration of xylenes that the isomerization hydrocarbon feedstock can contain is from c(xylene)1 to c(xylene)2 wt%, based on the total weight of the isomerization hydrocarbon feedstock, where c(xylene)1 and c(xylene)2 can independently be 30, 40, 50, 55, 60, 65, 70, 75, 80, 85, 90, 91, 92, 93, 94, 95, 96, 97, 98, 99, or even 100, provided that c(xylene)1 < c(xylene)2. Preferably, c(xylene)1 = 70. More preferably, c(xylene)1 = 80. Preferably, the isomerization hydrocarbon feedstock consists essentially of xylenes.

[0041] The isomerization hydrocarbon feedstock can contain ethylbenzene at various concentrations, such as from c(EB)1 to c(EB)2 wt%, based on the total weight of the isomerization hydrocarbon feedstock, where c(EB)1 and c(EB)2 can independently be 1, 2, 3, 4, 5, 6, 7, 8, 9, 10, 11, 12, 13, 14, 15, 16, 17, 18, 19, 20, 21, 22, 23, 24, 25, 26, 27, 28, 29, 30, provided that c(EB)1 < c(EB)2. Preferably, C(EB)2 = 20. Preferably, C(EB)2 = 10. Preferably, C(EB)2 = 5.

[0042] For the purpose of producing p-xylene product, it is highly desirable that the isomerization hydrocarbon feed containing C8 aromatic hydrocarbons fed to the isomerization zone contains a p-xylene concentration (based on the total amount of xylenes in the isomerization hydrocarbon feed) lower than its concentration in the equilibrium xylene mixture. Thus, the p-xylene in the isomerization hydrocarbon feed may have a concentration c(pX) in the range of c(pX)1 to c(pX)2 wt% based on the total amount of xylenes in the isomerization hydrocarbon feed, where c(pX)1 and c(pX)2 may independently be, for example, 0, 0.5, 1, 2, 3, 4, 5, 6, 7, 8, 9, 10, 11, 12, 13, 14, 15, 16, 17, 18, 19, 20, 21, 22, 23, 24, 25, 26, 27, 28, 29, 30, provided that c(pX)1 < c(pX)2. Preferably, c(pX)2 = 20. Preferably, c(pX)2 = 15. Preferably, c(pX)2 = 12. Preferably, c(pX)2 = 10. Preferably, c(pX)2 = 8. Preferably, c(pX)2 = 6. Preferably, c(pX)2 = 5. Preferably, c(pX)2 = 4. Preferably, c(pX)2 = 2. Preferably, c(pX)2 = 1. The m-xylene in the isomerization hydrocarbon feed may have a concentration c(mX) in the range of c(mX)1 to c(mX)2 wt% based on the total amount of xylenes in the isomerization hydrocarbon feed, where c(mX)1 and c(mX)2 may independently be, for example, 10, 15, 20, 25, 30, 35, 40, 45, 50, 55, 60, 65, 70, 75, 80, 85, 90, 95, provided that c(mX)1 < c(mX)2. Preferably, c(mX)1 = 30 and c(mX)2 = 80. Preferably, c(mX)1 = 40 and c(mX)2 = 80. The o-xylene in the isomerization hydrocarbon feed may have a concentration c(oX) in the range of c(oX)1 to c(oX)2 wt% based on the total amount of xylenes in the isomerization hydrocarbon feed, where c(oX)1 and c(oX)2 may independently be, for example, 10, 15, 20, 25, 30, 35, 40, 45, 50, 55, 60, 65, 70, 75, 80, 85, 90, 95, provided that c(oX)1 < c(oX)2. Preferably, c(oX)1 = 10 and c(oX)2 = 80. Preferably, c(oX)1 = 10 and c(oX)2 = 60. Preferably, c(oX)1 = 10 and c(oX)2 = 50. Preferably, c(oX)1 = 15 and c(oX)2 = 30. Such a feed may be the raffinate stream leaving the p-xylene recovery subsystem, for example, the raffinate stream using separation techniques based on adsorption chromatography.

[0043] The isomerized hydrocarbon feed may contain benzene, toluene, and C9+ hydrocarbons, but is desirably present in low amounts. The isomerized hydrocarbon feed may contain benzene and toluene in total in the range of c(BT)1 to c(BT)2 wt%, based on the total weight of the isomerized hydrocarbon feed, where c(BT)1 and c(BT)2 may independently be, for example, 0.01, 0.1, 1.0, 2.0, 3.0, 5.0, 8.0, 10.0, 15.0, 20.0, provided that c(BT)1 < c(BT)2. Preferably, c(BT)2 = 10.0. Preferably, c(BT)2 = 5.0. Preferably, c(BT)2 = 3.0. Toluene may be the major component between benzene and toluene. For example, the isomerized hydrocarbon feed may be substantially free of benzene. The isomerized hydrocarbon feed may contain C9+ hydrocarbons in total in the range of c(C9+)1 to c(C9+)2 wt%, based on the total weight of the isomerized hydrocarbon feed, where c(C9+)1 and c(C9+)2 may independently be, for example, 0.01, 0.1, 1.0, 5.0, 10.0, 20.0, provided that c(C9+)1 < c(C9+)2.

[0044] In various embodiments of the methods of the present disclosure, at least 80 wt%, preferably ≥85 wt%, preferably ≥90 wt%, preferably 95 wt%, preferably ≥98 wt%, preferably ≥99 wt%, preferably about 100 wt% of the isomerized hydrocarbon feed comprising C8 aromatic hydrocarbons is in the liquid phase at the inlet of the isomerization reactor. The isomerized hydrocarbon feed may have an inlet temperature in the range of T1 to T2 °C, where T1 and T2 may independently be, for example, 200, 210, 220, 230, 240, 250, 255, 260, 265, 270, 275, 280, 285, 290, 295, 300, provided that T1 < T2. The relatively low inlet temperature of the isomerized hydrocarbon feed, together with other isomerization conditions such as the pressure described below, enables the LPI process inside the isomerization reactor.

[0045] Molecular hydrogen co-feed

[0046] Some LPI processes in the prior art are carried out without co-feeding molecular hydrogen to the isomerization zone. We have found that in the absence of co-feeding any molecular hydrogen to the reactor, the isomerization catalyst can deactivate relatively rapidly over time, especially at ≥5 hours -1In the case where molecular hydrogen is supplied to the isomerization reactor at a low feed rate, for example at ≤ 10 ppm by weight (based on the total weight of the aromatic hydrocarbon feed entering the reactor), the isomerization catalyst may be deactivated at a rate that is lower than when no molecular hydrogen is co-fed at all, but may nevertheless be appreciable. Even in the case where the isomerization hydrocarbon feed is relatively easy to isomerize, for example, in the case where the isomerization feed contains a high concentration (e.g., ≥ 5 wt%, ≥ 8 wt%, ≥ 10 wt%, based on the total weight of the xylenes) of para-xylene and / or the isomerization feed contains a low concentration (e.g., ≤ 8 wt%, ≤ 6 wt%, ≤ 5 wt%, ≤ 4 wt%, ≤ 2 wt%) of ethylbenzene, at a low feed rate of molecular hydrogen or in the absence of co-fed molecular hydrogen, deactivation of the isomerization catalyst at an appreciable rate may be observed.

[0047] When molecular hydrogen is co-added at a high addition rate of ≥ 100 ppm by weight (based on the total weight of the isomerization hydrocarbon feed including the C8 aromatic hydrocarbon feed) in accordance with the present disclosure, the deactivation rate of the isomerization catalyst can be significantly reduced compared to both (i) no co-addition of molecular hydrogen at all and (ii) co-addition of molecular hydrogen at a low rate, e.g., ≤ 10 ppm by weight. In a completely unexpected manner, the deactivation rate of the isomerization catalyst can be significantly reduced even at ≥ 5 hours. -1 , ≥7.5 hours -1 , ≥10 hours -1 , ≥12.5 hours -1 , ≥15 hours -1 , ≥17.5 hours -1 , and even ≥20 hours -1 Low deactivation rates were observed with ≥100 ppm co-fed molecular hydrogen at high WHSVs.

[0048] In certain embodiments, molecular hydrogen or a portion thereof co-fed to the isomerization reactor may be injected into the isomerization reactor via an inlet as a pressurized gas. Additionally or alternatively, molecular hydrogen or a portion thereof may be fed into a feed line, vessel, or storage tank containing an isomerization hydrocarbon feed of C8 aromatics, where the molecular hydrogen forms a mixture with the isomerization hydrocarbon feed and is then supplied to the isomerization reactor. It is highly desirable that a portion, preferably a majority (e.g., ≥50%, ≥60%, ≥70%, ≥80%, ≥90%, ≥95%, ≥98%), more preferably substantially all (≥99%) of the co-fed molecular hydrogen is dissolved in the liquid phase in the isomerization reactor. To achieve a higher concentration of molecular hydrogen dissolved in the liquid phase of the isomerization hydrocarbon feed, a higher pressure may be applied. To maintain a majority of the molecular hydrogen dissolved in the liquid phase in the isomerization reactor, it is highly desirable that the feed rate of molecular hydrogen into the isomerization reactor is not higher than 5000 ppm by weight, based on the total weight of the isomerization hydrocarbon feed. The following Table I gives the approximate solubility of molecular hydrogen in liquid phase C8 aromatics in ppm by weight at various pressures at 25 °C, based on the total weight of C8 aromatics:

[0049] Table I

[0050]

[0051] Thus, molecular hydrogen may be fed to the isomerization reactor at a feed rate of r(H2)1 to r(H2)2 ppm by weight, based on the total weight of the isomerization hydrocarbon feed, where r(H2)1 and r(H2)2 may independently be, for example, 100, 150, 200, 250, 300, 350, 400, 450, 500, 550, 600, 650, 700, 750, 800, 850, 950, 1000, 1500, 2000, 2500, 3000, 3500, 4000, 4500, 5000, provided that r(H2)1 < r(H2)2. Preferably, r(H2)2 = 3000. Preferably, r(H2)2 = 2000. Preferably, r(H2)2 = 1000. Preferably, r(H2)2 = 800. Preferably, r(H2)2 = 600. Preferably, r(H2)2 = 500.

[0052] If molecular hydrogen is premixed with the isomerization hydrocarbon feed, it will be fed to the isomerization reactor at the inlet temperature as discussed above. If fed separately from the isomerization hydrocarbon feed, it may be fed to the isomerization reactor as a stream gas at a pressure sufficient to enable it to dissolve in the liquid phase in the isomerization reactor at a temperature preferably close to the inlet temperature of the isomerization hydrocarbon feed to achieve LPI in the isomerization reactor.

[0053] LPI conditions

[0054] The LPI step in the process of the present disclosure is carried out in an LPI reactor in the presence of an LPI catalyst under LPI conditions.

[0055] The LPI conditions can include a reaction gauge pressure range in the isomerization reactor from p1 to p2 kPa, where p1 and p2 can independently be, for example, 1700, 1800, 1900, 2000, 2100, 2200, 2300, 2400, 2500, 2600, 2700, 2800, 2900, 3000, 3100, 3200, 3300, 3400, 3500, provided that p1 < p2. Preferably, p2 = 3000. Preferably, p2 = 2500. The higher the reaction pressure, the greater the amount of molecular hydrogen that can dissolve in the liquid phase of the hydrocarbons present in the reactor, as indicated above.

[0056] The LPI conditions can include a reaction temperature range in the isomerization reactor from T1 to T2 °C, where T1 and T2 can independently be, for example, 200, 210, 220, 230, 240, 250, 255, 260, 265, 270, 275, 280, 285, 290, 295, 300, provided that T1 < T2. The relatively low reaction temperature of the LPI process can be particularly energy-efficient because it requires less energy to heat the isomerized hydrocarbon feed and it does not require condensing a large amount of high-temperature gaseous isomerization effluent into a liquid for downstream processing.

[0057] The LPI conditions in the process of the present disclosure can particularly advantageously include a high WHSV in the range of w1 to w2 hours -1 where w1 and w2 can be, for example, 5.0, 5.5, 6.0, 6.5, 7.0, 7.5, 8.0, 8.5, 9.0, 9.5, 10, 11, 12, 12.5, 13, 14, 15, 16, 17, 17.5, 18, 19, 20, provided that w1 < w2. As indicated above and demonstrated in the following comparative examples, an LPI process without co-fed molecular hydrogen or with co-fed molecular hydrogen at a low feed rate can experience a high rate of catalyst deactivation, even at a WHSV of 4 hours -1 Such a high WHSV of the process of the present disclosure enables a compact LPI reactor design for a new LPI with a given capacity, an increased capacity for an existing LPI reactor with a given catalyst loading, and a reduced catalyst consumption for producing a given amount of p-xylene product.

[0058] In certain embodiments of the processes of the present disclosure, the reaction temperature and / or weight hourly space velocity may be increased without substantially reducing the p-xylene / xylene weight ratio in the isomerization effluent. "Substantially reducing the p-xylene / xylene weight ratio" means reducing the p-xylene / xylene weight ratio in the isomerization effluent by ≥ 10% based on the p-xylene / xylene weight ratio before increasing the reaction temperature and / or weight hourly space velocity. The "p-xylene / xylene" weight ratio means the weight ratio of p-xylene to all xylenes present in a mixture or stream.

[0059] In certain other embodiments of the process of the present disclosure, the reaction temperature and / or weight hourly space velocity may be increased without substantially increasing xylene losses in the process. "Substantially increasing xylene losses" means increasing xylene losses by ≥ 10% based on xylene losses prior to increasing the reaction temperature and / or weight hourly space velocity. "Xylene losses" means the concentration of all xylenes present in the effluent stream minus the concentration of all xylenes present in the feed, expressed as a weight percent.

[0060] In some other embodiments of the method of the present disclosure, molecular hydrogen can be fed at a first rate in the initial stage of the catalyst cycle, and then subsequently during the normal operation of the isomerization reactor at a second rate, wherein the first rate is lower than the second rate. The initial stage of the catalyst cycle means the catalyst cycle period when the catalyst exhibits extremely high activity, which can lead to excessive production of side reactions and by-products. For example, in the initial stage, molecular hydrogen can be fed to the reactor at a rate of 0-500, or 0-300, or 0-200, or 0-150ppm. Under the low first feed rate of molecular hydrogen, the isomerization catalyst is deactivated ("de-edge") at a relatively high rate, thereby reducing the generation of side reactions and by-products. At the end of the initial stage, when the de-edge of the isomerization catalyst is completed, the molecular hydrogen feed rate can be increased to a higher level to reduce the deactivation of the isomerization catalyst to a desired level.

[0061] Isomerization catalyst

[0062] U.S. Patent Nos. 6,180,550, 6,448,459, 6,872,866, 7,244,409, 7,371,913, 7,495,137, 7,592,499, 8,221,707, 8,273,934, and 8,697,929 describe LPI processes and / or catalysts suitable for aromatic isomerization processes, the relevant portions of which are incorporated herein by reference in their entireties. Any suitable LPI catalyst known in the art may be used in the process of the present disclosure.

[0063] The isomerization catalyst that can be used in the LPI step of the method of the present disclosure may include a molecular sieve such as a zeolite. Such a zeolite can be selected from, but is not limited to, a zeolite having the following framework structure, and a combination thereof: MFI, MEL, MWW, MOR, etc. Preferably, the isomerization catalyst includes an MFI framework zeolite such as ZSM-5. Preferably, the isomerization catalyst includes a zeolite having a 10-member or 12-member ring structure such as MWW or MOR framework zeolite, or a mixture thereof. Preferably, the isomerization catalyst includes a first zeolite such as ZSM-5 having an MFI framework structure, and a second zeolite having a 10-member or 12-member ring structure different from the first zeolite. Non-limiting examples of MWW zeolites that can be used for the isomerization catalyst used in the method of the present disclosure include: MWW-22, MWW-49, MWW-54, and combinations thereof. The zeolite present in the isomerization catalyst can advantageously be in hydrogen form. The following combinations of zeolites are particularly advantageous for isomerization catalysts: ZSM-5 and MWW-22, ZSM-5 and MWW-49, and ZSM-5 and MWW-56.

[0064] In certain embodiments, the isomerization catalyst may include a first metal element selected from Fe, Co, Ni, Ru, Rh, Pd, Re, Os, Ir, Pt, and combinations thereof, and an optional second metal selected from Sn, Zn, Ag, and combinations thereof. The first metal element may catalyze the hydrogenation of olefins, which may be produced in an isomerization reaction by, for example, dealkylation of ethylbenzene. The second metal element may promote the catalytic effect of the first metal element. In other embodiments, the isomerization catalyst does not contain noble metals (i.e., Ru, Rh, Pd, Os, Ir, and Pt). In other embodiments, the isomerization catalyst may not contain any Group 7-10 metals. In other embodiments, the isomerization catalyst may not contain any Group 7-15 metals, except aluminum.

[0065] The ZSM-5 zeolite useful in the isomerization catalyst may have one or more of the following characteristics: being in the hydrogen form (HZSM-5), having a crystal size of ≤ 0.1 micrometer, having a mesopore surface area (MSA) ≥ 45 m 2 / g, a total surface area to mesopore surface area ratio of ≤9, and a silica to alumina molar ratio in the range of 20 to 50.

[0066] The isomerization catalyst can be a self - binding zeolite catalyst substantially free of binder. Alternatively, the isomerization catalyst can contain a binder material in addition to a molecular sieve such as zeolite. Examples of suitable binder materials include, but are not limited to: alumina, silica, aluminosilicate, zirconia, zircon, titanium dioxide, clays (such as montmorillonite, bentonite, hectorite), kaolins (such as Dixie, McNamee, Georgia and Florida kaolins or other kaolins whose main mineral components are halloysite, kaolinite, nacrite or dickite), combinations thereof, their compounds, etc. Any suitable amount of binder can be present in the isomerization catalyst. For example, the binder material can be included in the isomerization catalyst at a concentration of c1 to c2 wt%, based on the total weight of the catalyst, where c1 and c2 can independently be 1, 5, 10, 15, 20, 25, 30, 35, 40, 45, 50, 55, 60, 65, 70, 75, 80, 85, 90, 95, 98, 99, provided that c1 < c2. Inclusion of a binder in the isomerization catalyst can enhance its mechanical strength, etc.

[0067] The isomerization catalyst can be prepared by methods known in the art. For example, the components of the catalyst such as zeolite and an optional binder can be mixed to form a dense mixture, which is then extruded into a desired shape, dried, calcined and optionally selectivated to produce the isomerization catalyst. The raw catalyst can be activated outside the reactor or in the isomerization reactor prior to the C8 aromatic hydrocarbon isomerization operation. Activation can be carried out, for example, by exposing the catalyst to a stream of gas containing molecular hydrogen.

[0068] The isomerization catalyst can be a newly prepared catalyst or a regenerated catalyst, or a mixture thereof. Catalyst regeneration can be carried out in the isomerization reactor, for example, by exposing the catalyst to a stream of gas containing molecular hydrogen after the catalyst activity has decreased to a threshold level at the end of the catalyst cycle. Alternatively, non - in - situ regeneration of the catalyst can be achieved, where the spent catalyst is removed from the isomerization reactor and heated and / or exposed to a stream of gas containing molecular hydrogen in an oxygen - rich environment to reduce the coke on its surface.

[0069] In some embodiments, the reactor can be a fixed - bed reactor, a fluidized - bed reactor, or a moving - bed reactor. The hydrocarbon liquid in the reactor can flow upward, downward, or in a radial manner.

[0070] Isomerization effluent

[0071] As a result of the isomerization reaction in the reactor, the isomerization effluent desirably contains para-xylene at a concentration higher than in the isomerization hydrocarbon feed, and meta-xylene and ortho-xylene at a total concentration lower than in the isomerization hydrocarbon feed. The isomerization effluent can contain para-xylene at a concentration of c(pX)1 to c(pX)2 wt%, based on the total weight of xylenes in the isomerization effluent, where c(pX)1 and c(pX)2 can independently be, for example, 15, 16, 17, 18, 19, 20, 21, 22, 23, provided that c(pX)1 < c(pX)2. Preferably, c(pX)1 = 18. Preferably, c(pX)1 = 20. Preferably, c(pX)1 = 21. Preferably, c(pX)1 = 22. Preferably, c(pX) is close to its concentration in the equilibrium xylene mixture at the isomerization temperature. In the process of the present disclosure, by co-feeding molecular hydrogen at a feeding rate of ≥ 100 ppm by weight (based on the total weight of the isomerization hydrocarbon feed), even at high WHSVs of ≥ 5 hours -1 , ≥ 7.5 hours -1 , ≥ 10 hours -1 , ≥ 12.5 hours -1 , ≥ 15 hours -1 , ≥ 17.5 hours -1 and even ≥ 20 hours -1 , high isomerization catalyst activity can be achieved, and it is possible to achieve and maintain c(pX)1 ≥ 20 wt% for a long period of time, which is very surprising.

[0072] The total concentration of meta-xylene and ortho-xylene in the isomerization effluent is lower than that in the isomerization feed. Desirably, the concentration of meta-xylene in the isomerization effluent, based on the total weight of xylenes in the isomerization effluent, is close to its concentration in the equilibrium xylene mixture. Thus, the isomerization effluent may contain meta-xylene at a concentration of c(mX)1 to c(mX)2 wt%, based on the total weight of xylenes in the isomerization effluent, where c(mX)1 and c(mX)2 may independently be, for example, 40, 42, 44, 45, 46, 48, 49, 50, 51, 52, 54, 55, 56, 58, 60, 62, 64, 65, 66, 68, 70, provided that c(mX)1 < c(mX)2. Preferably, c(mX)2 = 60. Preferably, c(mX)2 = 58. Preferably, c(mX)2 = 55. Preferably, c(mX)2 = 53. Preferably, c(mX)2 = 50. Desirably, the concentration of ortho-xylene in the isomerization effluent, based on the total weight of xylenes in the isomerization effluent, is close to its concentration in the equilibrium xylene mixture. Thus, the isomerization effluent may contain ortho-xylene at a concentration of c(oX)1 to c(oX)2 wt%, based on the total weight of xylenes in the isomerization effluent, where c(oX)1 and c(oX)2 may independently be, for example, 15, 16, 18, 20, 22, 24, 25, 26, 28, 30, 32, 34, 35, provided that c(oX)1 < c(oX)2. Preferably, c(oX)1 = 20 and c(oX)2 = 30. Preferably, c(oX)1 = 20 and c(oX)2 = 26.

[0073] The isomerization effluent may contain ethylbenzene at various concentrations, depending on the concentration of ethylbenzene present in the isomerization hydrocarbon feed.

[0074] Recovery of paraxylene

[0075] The isomerization effluent, which is rich in p-xylene and poor in the total of meta-xylene and ortho-xylene compared to the isomerization hydrocarbon feed, can be fed to a p-xylene recovery subsystem from which a high-purity p-xylene product can be produced.

[0076] The p-xylene recovery system can use adsorption chromatography methods to separate p-xylene from the meta-xylene, ortho-xylene, and ethylbenzene present in the isomerization effluent. Exemplary adsorption chromatography methods and systems are described, for example, in U.S. Patent Nos. 3,040,777, 3,201,491, 3,422,848, 9,302,201, 3,761,533, 4,029,717, 6,149,874, and 9,302,201, the relevant portions of which are incorporated herein by reference in their entireties.

[0077] The para-xylene recovery system can use a crystallization method to separate para-xylene from other C8 aromatic hydrocarbons present in the isomerization effluent. Exemplary crystallization separation methods and systems are described in, for example, 3,662,013, 5,329,061, 5,498,822, 6,147,272, and 6,600,083, the relevant portions of which are incorporated herein in their entirety.

[0078] The para-xylene recovery system may use a combination of the adsorption chromatography method and the crystallization method as described above.

[0079] In addition to the isomerization effluent produced from the isomerization reactor in the process of the present disclosure, the paraxylene recovery system can also receive other paraxylene-containing feed streams, such as one or more paraxylene-rich feed streams produced by a C7- / C9+ aromatic hydrocarbon transalkylation process, a toluene disproportionation process, a process for methylating benzene / toluene with methanol, and the like.

[0080] The method of the present disclosure uses an LPI reactor. In certain embodiments, a single LPI reactor may be sufficient to process the entire isomerized hydrocarbon feed. In other embodiments, it may be desirable to use two or more LPI reactors connected in series and / or in parallel. When multiple LPI reactors are connected in parallel, they may be configured to have overlapping but different operating cycles, so that when one reactor is shut down for catalyst replacement or repair, the others may continue to operate to ensure the continuity of the operation of other components in the overall process of aromatic compound production. Additionally or alternatively, when multiple LPI reactors are connected in series, they may be configured to have overlapping but different operating cycles, so that when one reactor is in the time for catalyst replacement or repair, the others may continue to operate to ensure the continuity of the operation of other components in the overall process of aromatic compound production. When two LPI reactors are connected in series to include an upstream reactor (i.e., a pre-reactor) and a downstream reactor (i.e., a post-reactor), it is contemplated that the two reactors may be online at the same time or at different times, and the two reactors may switch positions as pre- or post-reactors. It is also contemplated that multiple LPI reactors may be operated in series for a certain period of time, and then operated in parallel for another period of time.

[0081] The method of the present disclosure may also include an LPI reactor and a VPI reactor. The LPI reactor is operated under isomerization conditions as described above. The VPI reactor is operated under VPI conditions to implement a VPI process carried out in the gas phase in the presence of a VPI catalyst. VPI methods and reactors known in the art can be used. References describing VPI methods and / or reactors and / or catalysts include, but are not limited to, U.S. Patent Nos. 3,651,162, 3,856,872, 3,919,339, 4,098,836, the relevant parts of which are incorporated herein by reference in their entirety. The LPI reactor and the VPI reactor may be operated in parallel or in series. Preferably, the LPI reactor and the VPI reactor are operated in series. Preferably, the LPI is a pre-reactor and the VPI reactor is a post-reactor. The VPI reactor may be more effective than the LPI reactor in converting ethylbenzene into usable products such as benzene and / or xylene. Thus, a combination of an LPI reactor and a VPI reactor may be advantageously used to process an isomerized hydrocarbon feed comprising elevated concentrations of ethylbenzene, e.g., ≥8 wt%, ≥10 wt%, ≥12 wt%, ≥15 wt%, 20 wt%, based on the total weight of the isomerized hydrocarbon feed.

[0082] In the following cases (case numbers 1-7), various specific embodiments of the process of the present disclosure are described. The simulations in cases 1-5 apply to an LPI process in which molecular hydrogen is co-fed into the LPI reactor in accordance with an embodiment of the present disclosure, and similarly apply to an LPI process in which molecular hydrogen is not co-fed into the isomerization reactor.

[0083] Case 1: Controlling temperature to minimize catalyst loading size while optimizing xylene isomerization is relatively undesirable The extent of the xylene loss reaction

[0084] Reactor temperature can be optimized to reduce catalyst loading size (increase WHSV) while maintaining low / acceptable xylene losses. Figure 1 and 2 The ratio of the weight of p-xylene to the total weight of all xylenes (p-xylene / xylene ratio) and the xylene losses as a function of WHSV and temperature in certain simulated LPI processes are shown, respectively. Figure 1 It is shown that in this case the design WHSV can be doubled by increasing the operating temperature by approximately 20-30°C while maintaining a comparable paraxylene / xylene ratio in the LPI product mixture effluent ("isomerization effluent"). As an example, the design WHSV for a 2.5 hour -1 The LPI reactor operating at 240°C is expected to achieve a paraxylene / xylene ratio of 0.2347 in the isomerization effluent from the model feed composition. Designed for a WHSV of 5.0 hours -1 (This only needs to be designed for WHSV of 2.5 hours-1 An LPI reactor operating at 260°C (half the catalyst loading in the reactor) is expected to achieve a paraxylene / xylene ratio of 0.2339 in the isomerization effluent. In a second example, a reactor designed for a WHSV of 10 hours -1 The reactor operating at 290°C is expected to achieve a paraxylene / xylene ratio of 0.2355 in the isomerization effluent from the model feed composition. By increasing the operating temperature by an additional 30°C from 260°C, with a design for a WHSV of 5.0 hours -1 The reactor catalyst loading size was again halved compared to the reactor of .These examples clearly illustrate the ability to optimize the design of catalyst loading size versus operating temperature to achieve similar paraxylene / xylene ratio targets in the isomerization effluent.

[0085] Continuing with the example provided in the above section, Figure 2 The trade-off of xylene loss when optimizing reactor temperature to reduce catalyst loading size is illustrated. In this disclosure, "xylene loss" is calculated as the total concentration of all xylenes in the feed composition (expressed as xylene weight percent based on the total weight of the feed composition) minus the total concentration of xylenes in the product mixture effluent (expressed as xylene weight percent based on the total weight of the product mixture effluent). In the case of the first example (from 240°C for 2.5 hours) -1 WHSV shift to 260°C for 5 hours -1 The estimated impact of xylene loss increases from 0.37% per pass to 0.53% WHSV per hour. This per pass xylene loss is still low after the increase and is still comparable or better than conventional gas phase isomerization processes. Therefore, the reduction in catalyst loading size is beneficial to reduce capital investment at the expense of only a minor yield debit (i.e., a minor increase in xylene loss). In a second example, the WHSV is further increased from 5 to 10 hours. -1 , and the operating temperature was increased from 260 to 290°C. For this example, the xylene loss increased from 0.53% (in 5 hours -1 WHSV and 260°C) increased to 0.71% (at 10 hours -1 and 290°C). Again, xylene losses are still low even after the increase, comparable to or better than conventional gas phase isomerization processes. Thus, a reduction in catalyst loading size is beneficial in reducing capital investment at the expense of only a minor yield loss (i.e., a minor increase in xylene losses).

[0086] The above examples illustrate the ability to optimize the loading size and operating temperature to meet the target para-xylene / xylene ratio in the isomerization effluent while reducing capital investment. Those skilled in the art will recognize that reverse optimization is also an option. A higher catalyst loading can be selected and the single pass xylene loss reaction can be reduced by reducing the temperature. However, if Figure 2 As shown in Figure 1, there will be a point of diminishing returns in reducing xylene losses as the temperature is further reduced. A single catalyst system was assumed in the simulations in Case 1. Although alternative catalyst systems (including multiple catalyst bed systems) can be used to change the temperature range and effect size, an optimization approach can generally be applied.

[0087] Case 2: Using temperature to manage catalyst activity loss / aging

[0088] Reactor temperature can also be used as a means to maintain / optimize reactor yield in response to catalyst activity loss (ie, aging). Figure 3 Illustrate the relative activity loss of the catalyst as a function of temperature in order to maintain a WHSV of 2.5 hours for a particular isomerization feed composition. -1 A constant para-xylene / xylene ratio in the product mixture effluent of a simulated LPI process operated under Figure 3 It is seen that as catalyst activity decreases (i.e., as relative activity loss increases), the reaction temperature can be increased to maintain a constant p-xylene / xylene ratio in the product mixture effluent. It will be appreciated by those skilled in the art that when catalyst activity loss occurs, relative xylene losses will also decrease at various rates (at constant operating temperature, WHSV, feed concentration). Temperature increases can be limited / set to offset activity losses by optimizing between the p-xylene / xylene ratio in the product and the single pass xylene losses in the reactor.

[0089] Case 3: Determining relative activity loss for a given reactor or series of reactors

[0090] The relative activity loss of a single catalyst system or multiple catalyst systems can be estimated by monitoring various performance parameters as a function of temperature, feed rate, WHSV and / or feed concentration, and the performance parameters may include, but are not limited to, paraxylene / xylene ratio in the product, paraxylene / xylene ratio delta across the reactor, xylene losses, toluene in the product, benzene in the product, trimethylbenzenes in the product, methylethylbenzene in the product, total C9 aromatics in the product ("A9"), total C10 aromatics in the product ("A10"), total C9+ aromatics in the product ("A9+"), toluene variation across reactors, benzene variation across reactors, trimethylbenzene variation across reactors, methylethylbenzene variation across reactors, total A9 variation across reactors, total A10 variation across reactors, total A9+ variation across reactors, ethylbenzene conversion across reactors, non-aromatic conversion across reactors (single species, a subset of species, or all non-aromatics), C5- products (single species, a subset of species, or all C5-), C5- increment across reactors (single species, a subset of species, or all C5-), or any combination of these parameters.

[0091] As an example, Figure 4 The relative activity loss is estimated using the p-xylene / xylene in the product mixture effluent. For this example, a constant 2.5 hour -1 WHSV operation. In this example, in 2.5 hours -1 WHSV establishes a temperature curve that relates the para-xylene / xylene ratio in the product to the relative activity loss. For illustrative purposes, let us assume that at 2.5 hours -1 The catalyst system operated at WHSV and 240°C achieved a para-xylene / xylene ratio in the product of 0.2234. This corresponds to a relative activity loss of ~50% from the fresh catalyst activity. In a second example, let us assume that at 2.5 hours -1 The catalyst system operating at WHSV and 280° C. achieved a para-xylene / xylene ratio in the product of 0.2321. This performance represents a relative activity loss of approximately 75% from the fresh catalyst activity.

[0092] In the case of a pre-post or multiple reactor system, the estimated relative activity loss or any parameter used in determining the estimated activity loss can be used to indicate / indicate when to take a reactor / catalyst bed out of service. Figure 4In the example provided in , the paraxylene / xylene ratio curve illustrates that as the activity loss extends beyond about 75%, a sudden drop occurs. Therefore, a plausible alternative strategy would be to use temperature increases to maintain yield until the approximately 75% activity loss condition is met, and then trigger the decommissioning of the catalyst bed once the 75% activity loss condition is met / exceeded.

[0093] Case 4: Catalyst activity demarcation at the start of operation; catalyst activity management during operating rate decay

[0094] In contrast to Case 2, it is common for a fresh catalyst bed to have a higher than expected initial catalyst activity at the beginning of the cycle, which results in higher xylene losses or undesirable levels of byproducts such as benzene, toluene or A9+ molecules. This can also happen in the middle or at the end of a run if the unit rate is significantly turned down to less than the design WHSV. In order to minimize the production of undesirable byproducts at the beginning of a cycle or during a significantly decayed operation, the temperature can be lowered to reduce the yield of undesirable byproducts. In the case of "de-marginalizing" the initial new catalyst activity, usually during the first 1-2 months, but in extreme cases it may be as much as 1 year or several years, as the byproduct production rate decreases, the temperature can be gradually increased. In the case of unit rate decay, the temperature can be adjusted as needed to manage the yield. The absolute magnitude of the temperature change will depend on the percentage of decay from the design rate and the relative activity loss of the catalyst in the operation that caused the decay event.

[0095] Temperature control to manage above the desired initial catalyst activity can also be used for multi-bed / multi-reactor systems, such as pre-post configurations. In the example of a pre-post configuration, a fresh catalyst bed (whether pre- or post-) can be started at a lower temperature relative to the reactor already in operation. Independent temperature control of a single reactor can be managed by any number of configurations, which may include, but are not limited to: independent heaters / coolers for the feed to each reactor, independent bypass lines around the heater / cooler of each reactor to allow temperature control of each reactor, injection of a cooled or heated fresh stream into a post reactor / post reactors, which is mixed with the effluent of the previous reactor / reactors to control the temperature of a post reactor / post reactors, and utilizing natural heat loss between a pre-post reactor / pre-post reactors to operate a pre reactor / pre reactors at a lower temperature than a post reactor / post reactors. Any combination of these temperature control methods can be used.

[0096] Case 5: Managing a front-end configuration where the reactor / catalyst bed size varies between the front and the back

[0097] As shown in Case 1, the reactor operating temperature can be controlled to allow similar yields for reactors with different catalyst loadings or WHSVs. In a pre-post or multi-reactor configuration, the reactors can be optimized / designed to have different amounts of catalyst. See Table II for an illustration of Cases 5a and 5b. In the tables of the present disclosure, for purposes of brevity, EB means ethylbenzene, L-paraffins means linear paraffins, BCNA means branched or cyclic non-aromatic hydrocarbons, MEBZ means methylethylbenzene, DEBZ means diethylbenzene, TriMBZ means trimethylbenzene, and DMEBZ means dimethylethylbenzene.

[0098] Case 5a estimates the yield for a pre-post configuration, where the pre-catalyst bed is -1 WHSV and 240 °C operation and post-catalyst bed for 10 hours -1 WHSV and 260 ° C. In this case, the bulk of the isomerization reaction takes place in the larger front catalyst bed, which allows the use of a 10-hour -1 The smaller post-bed operated at 260°C WHSV and slightly increased temperature serves as the finishing reactor for xylene isomerization. Note that this combination of pre-post systems with different WHSVs and temperatures produces the same 2.5 hour saturation as presented in the data in Tables 1 and 2 above. -1 Similar results were obtained for the larger single bed design case at WHSV and 240 °C.

[0099] Case 5b estimates the yield for a pre-post configuration, where the pre- and post-catalyst beds are each heated for an equivalent of 10 hours. -1 WHSV operation, but controlling the operating temperature between the lead and tail reactors (240°C for the lead reactor and 270°C for the tail reactor). This configuration results in a similar paraxylene / xylene ratio in the product as Case 5a, but does have a slightly higher estimated xylene loss of 0.53% relative to the 0.46% xylene loss of Case 5a.

[0100] Table II

[0101]

[0102] Case 6: Managing Hydrogen Co-feed in Pre-Post or Multi-reactor Systems

[0103] It has been surprisingly found that low concentration hydrogen injection mitigates / mitigates activity loss in liquid phase isomerization systems. It should be noted that the hydrogen co-feed level of a pre-post or multi-reactor system can be controlled only to the inlet of the first reactor, or each individual reactor or any subset of reactors / catalyst beds can be controlled separately. The hydrogen level in the feed to each reactor can be controlled by any of the following means, but not limited to the subsets listed herein: independent hydrogen injection to any given reactor / reactors, adding fresh feed to the subsequent reactor / reactors, which is combined with at least a portion of the effluent from the previous reactor / reactors, flashing the effluent of the previous reactor / reactors to a lower pressure to remove hydrogen from the liquid phase, and then feeding at least a portion of the remaining liquid stream to the subsequent reactor / reactors.

[0104] The hydrogen co-feed of a single reactor or any subset of reactors can be optimized to minimize aging at different WHSV, temperature or feed concentrations. It can also be controlled / optimized over the cycle time of the catalyst in response to a drop in activity. For example, at the beginning of a cycle, if the fresh catalyst activity for a given bed is undesirably high and results in high generation of undesirable byproducts, the hydrogen co-feed can be reduced. Reducing the hydrogen co-feed during this "de-marginalization" period will help promote / accelerate the initial activity loss, thereby reducing the high undesirable byproduct generation time at the beginning of the cycle. As the generation of undesirable byproducts decreases, the hydrogen co-feed can be increased to reduce long-term activity losses and ultimately extend the run time of the catalyst bed.

[0105] If an increased rate of activity loss is observed under various operating conditions, the hydrogen co-feed can be optimized throughout the cycle. For example, during times when the observed rate of activity decline is high, the hydrogen co-feed can be increased. During times when the observed rate of activity decline is low / negligible, the hydrogen co-feed can be decreased to reduce operating costs.

[0106] Case 7: Optimization of WHSV, temperature based on feed composition for a single bed or for a pre-post or multi-reactor system Degree, hydrogen co-feed

[0107] Feed concentration affects both the p-xylene / xylene ratio and the degree of undesirable by-product formation in the isomerization effluent. The concepts outlined in the above cases 1-6 can also be applied to optimize capital and yields, where significant feed composition differences or changes are expected. Please refer to Table III for illustrating the control of a single bed reactor temperature to achieve a similar yield for a standard feed composition relative to a high o-xylene feed. In this example, the standard feed contains a low p-xylene (<1 wt%) content, a low to medium EB content (~5.3 wt%), and a medium o-xylene content (~37 wt%). The feed rich in o-xylene (Ox) also contains low p-xylene (<1%), but also has a low EB and m-xylene content (both <1 wt%). In this example, when processing an o-xylene-rich feed, when compared to the 2.5 hour -1 The single catalyst bed produces a lower para-xylene / xylene ratio and lower xylene losses in the product when compared to processing a standard feed at the same conditions at WHSV and 240° C. However, in this example, for the o-xylene-rich feed case, the para-xylene / xylene ratio in the product can be increased to match the standard feed product by increasing the reactor temperature from 240° C. to 250° C. As discussed in the above section, optimizing WHSV, temperature, or a combination of both can be used to achieve a product comparable to the standard feed case.

[0108] Table III

[0109]

[0110] Figure 5 : Illustrative Aromatics Production Overall Process for Producing Para-Xylene Product

[0111] Figure 5A method 101 for preparing xylene, particularly para-xylene product, from a reformate stream is schematically illustrated. In this figure, a heavy naphtha stream 103 produced by a crude oil refining process is supplied to a reforming region 105. The reforming region 105 may include any conventional naphtha catalytic reforming reactor (s) such as one or more fixed bed reactors (s) and fluidized bed reactors (s) known in the art. The reforming catalyst is disposed in the reforming region. When the reforming catalyst is contacted under reforming conditions such as those known in the art, the hydrocarbons in the heavy naphtha stream 103 undergo a series of chemical reactions, including but not limited to isomerization, aromatization, dehydrogenation and cyclization, etc., whereby at least a portion of the paraffins and cycloalkanes are converted into aromatic hydrocarbons. A reforming effluent 107 containing C6+ aromatic hydrocarbons (including benzene, toluene, xylene, ethylbenzene and C9+ aromatic hydrocarbons) can be obtained from the reforming region. The reformate effluent 107 or a portion thereof is then supplied to a reformate separator 109 (e.g., a single distillation column or a series of distillation columns), thereby producing a stream 111 rich in C6-C7 hydrocarbons and a stream 113 rich in C8+ aromatic hydrocarbons. The stream 111 rich in C6-C7 hydrocarbons comprises benzene, toluene, and their azeotropic paraffins and cycloalkanes, etc. The stream 113 rich in C8+ aromatic hydrocarbons may comprise C8 aromatic hydrocarbons and C9+ aromatic hydrocarbons. The stream 113 rich in C8+ aromatic hydrocarbons, optionally together with other streams rich in C8+ aromatic compounds (one or more) such as stream 145 (described below) as a combined stream 114, is supplied to a xylene separator 115, thereby producing a stream 117 rich in xylenes and a stream 129 rich in C9+ aromatic hydrocarbons. The combined stream 114 is rich in C8+ aromatic hydrocarbons and depleted in benzene, toluene, and their azeotropes compared to stream 107. The xylene-rich stream 117 is rich in xylene and ethylbenzene. The concentration of ethylbenzene in the stream 117 can vary in a wide range. The stream 117 can contain para-xylene in various concentrations, depending on the composition (one or more) of the stream (one or more) rich in C8+ aromatic hydrocarbons supplied to the xylene separator 115. For the purpose of producing a para-xylene product, the xylene-rich stream 117 is usually supplied to a para-xylene recovery subsystem 119, thereby producing a para-xylene-rich para-xylene product stream 121 and a para-xylene-poor stream 123. The para-xylene recovery subsystem 119 can be any para-xylene separation system based on crystallization and / or based on adsorption chromatography known in the art. The para-xylene-poor stream 123, which is usually rich in meta-xylene, ortho-xylene and ethylbenzene compared to the stream 117, is at least partially supplied to an isomerization zone 125 containing an isomerization catalyst disposed therein and operating under isomerization conditions. A portion of the meta-xylene and ortho-xylene in the feed stream 125 supplied to the isomerization zone 125 is converted to para-xylene upon contacting the isomerization catalyst under isomerization conditions. The isomerization effluent 127 exiting the isomerization zone 125 contains a higher concentration of para-xylene than the para-xylene-depleted feed stream 123.The isomerization effluent 127 or a portion thereof is then supplied to the xylene splitter 115. The xylene splitter 115, the paraxylene recovery subsystem 119 and the isomerization zone 125 form a xylene loop.

[0112] Available in Figure 5 The isomerization methods of the present disclosure are advantageously used in the methods described.

[0113] like Figure 5 , a C9+ aromatics-rich stream 129, which is typically C9, C10, and C11+ aromatics, produced by a xylene splitter 115 is then separated in a distillation column 131 to obtain a C9-C10 aromatics-rich stream 133 and a C11+ aromatics-rich stream 135. Stream 135 is typically directed away and used, for example, as a motor gasoline blending stock, fuel oil, etc. Stream 133 is then supplied to a transalkylation zone 147 having a transalkylation catalyst disposed therein, along with a benzene / toluene-rich stream 146. In the presence of a transalkylation catalyst and under transalkylation conditions, C9-C10 aromatics react with benzene / toluene to produce xylenes. Stream 111, which is typically C6-C7 hydrocarbon-rich, is supplied to an extractive distillation zone 137, where a C6-C7 aromatics-rich stream 139 and an aromatics-depleted residual liquid stream 138 are produced. Stream 139 is then supplied to a benzene column 141, thereby producing a benzene product stream 143, a toluene-rich stream 146, and a C8+ aromatics-rich stream 145. The toluene-rich stream 146, or a portion thereof, is supplied to a transalkylation 147 along with the C9-C10 aromatics-rich stream 133 as described above. The C8+ aromatics-rich stream 145 is then supplied to a xylene splitter 115 along with stream 113, as described above.

[0114] The present disclosure is further illustrated by the following non-limiting examples. Example

[0115] In the following examples, a C8 aromatic hydrocarbon feed stream consisting essentially of ethylbenzene, para-xylene, ortho-xylene and ethylbenzene is fed to an LPI reactor preloaded with a given amount of an isomerization catalyst based on ZSM-5. A molecular hydrogen bottle is connected to the LPI reactor or a C8 aromatic hydrocarbon feed stream source so that molecular hydrogen can be co-fed to the reactor at various feed rates. The C8 aromatic hydrocarbon feed stream is heated to achieve a reactor inlet temperature in the range of 200-300°C. The pressure in the reactor is in the range of 689-3447 kPa (gauge pressure). Under such conditions, the C8 aromatic hydrocarbons present in the reactor are substantially completely in the liquid phase, and the molecular hydrogen (if present) co-fed to the reactor is substantially completely dissolved in the liquid phase of the hydrocarbon. The isomerization conditions are changed to test the deactivation of the isomerization catalyst. The isomerization effluent leaving the isomerization reactor is then analyzed by using gas chromatography to determine the concentrations of p-xylene, o-xylene and m-xylene (wt %), based on the total weight of the isomerization effluent. The measured p-xylene concentration (C(pX), wt %), o-xylene concentration (C(oX), wt %) and m-xylene concentration (C(mX), wt %) are then used to calculate the p-xylene concentration (C(pX / X), %) based on the total amount of xylenes according to the following formula:

[0116]

[0117] In the description of these embodiments, "CGpGC" means the cumulative weight (grams) of C8 aromatic hydrocarbon feed that has been processed through a batch of isomerization catalyst / gram of isomerization catalyst. Under the assumption that the feed rate of the C8 aromatic hydrocarbon feed is maintained constant, CGpGC will correspond to the product of the feed rate and the operating time of the catalyst.

[0118] Example 1

[0119] In this example, the C8 aromatic hydrocarbon feed stream contains 3.4 wt% para-xylene, 64.1 wt% meta-xylene, 18.3 wt% ortho-xylene, 12.3 wt% ethylbenzene and 1.3 wt% other hydrocarbons. In the isomerization process, isomerization conditions such as the feed rate and WHSV of molecular hydrogen are changed. The reactor pressure is changed so as to maintain the molecular hydrogen co-fed to the isomerization reactor at various feed rates to be substantially completely dissolved in the liquid phase C8 aromatic hydrocarbons. The reaction temperature in the isomerization reactor is maintained at about 280°C. Figure 6 The concentration of para-xylene in the isomerization effluent as a function of CGpGC in the process is shown in FIG. Figure 6 The feed rate of molecular hydrogen (C(H2), ppm by weight of molecular hydrogen, based on the weight of the C8 aromatic hydrocarbon feed stream) and WHSV (hours) during the various CGpGC ranges shown in -1 ).exist Figure 6 , the CGpGC range 201 is defined by two vertical lines L-1 and L-2, the range 203 is defined by vertical lines L-2 and L-3, the range 205 is defined by vertical lines L-3 and L-4, the range 207 is defined by vertical lines L-5 and L-6, the range 209 is defined by vertical lines L-7 and L-8, and the range 211 is defined by vertical lines L-8 and L-9.

[0120] Table II

[0121] CGpGC range 201 203 205 207 209 211 <![CDATA[WHSV (hours -1 )]]> 10 10 10 20 20 20 <![CDATA[H2 co-feeding rate (ppm)]]> 7 0 255 255 0 0

[0122] If available Figure 6 As can be seen, in the CGpGC range 201, at WHSV of 10 hours -1 The concentration of para-xylene in the isomerization effluent (C(pX), wt%) steadily decreased at a molecular hydrogen feed rate of 7 ppm, indicating a steady decrease in the activity of the catalyst. -1 In the absence of co-feed molecular hydrogen to the isomerization reactor, C(pX) continued its decline similar to range 201, indicating a similar deactivation rate of the isomerization catalyst in the absence of co-feed molecular hydrogen. The data in ranges 201 and 203 illustrate that at 10 hours -1 At the high WHSV of 1.3 ppm, the isomerization catalyst deactivated relatively quickly when no molecular hydrogen was co-fed and when molecular hydrogen was co-fed at a low rate of 7 ppm.

[0123] At the end of CGpGC range 203 and the beginning of range 205, the molecular hydrogen feed rate was increased to 255 ppm while the WHSV was maintained at 10 hours. -1 , which is then maintained throughout range 205. At the beginning of range 205, C(pX) is significantly increased compared to any CGpGC in ranges 203 and 201, indicating that the activity of the isomerization catalyst is significantly increased in the presence of significantly increased molecular hydrogen concentration in the reactor. Throughout range 205, C(pX) is substantially stable, indicating that after 10 hours -1 The isomerization catalyst showed little deactivation in the presence of 255 ppm molecular hydrogen at a high WHSV.

[0124] At the end of CGpGC range 205, WHSV was ramped up while the molecular hydrogen feed rate was maintained at 255 ppm. During range 207, the molecular hydrogen feed rate and WHSV were maintained at 255 ppm and 20 hours, respectively. -1. The C(pX) in range 207 is lower than that in range 205, which is understandable due to the twice higher WHSV in range 207. Nevertheless, the C(pX) in the entire range 207 is significantly higher than at the end of ranges 201 and 203, even though the WHSV in range 207 is twice as high as that in ranges 201 and 203, clearly illustrating the significantly higher isomerization catalyst activity in range 207 than in ranges 201 and 203, and the effect of the high molecular hydrogen concentration in the isomerization reactor on the catalyst activity. In addition, during range 207, once C(pX) stabilizes, it decreases very little, showing that even at 20 hours -1 The very low deactivation rate of the isomerization catalyst at the very high WHSV indicates the substantial effect imparted by the high concentration of molecular hydrogen in the isomerization reactor.

[0125] At the end of range 207, the feed rate of molecular hydrogen was reduced. During range 209, no molecular hydrogen was fed to the reactor and the WHSV was maintained at 20 hours. -1 As a result of the absence of molecular hydrogen in the isomerization reactor, C(pX) is greatly reduced in range 209 compared to range 207, again illustrating the effect of the presence of molecular hydrogen in the isomerization reactor on the activity of the isomerization catalyst. C(pX) in range 209 is similar to the end of range 203, where no molecular hydrogen is also fed to the reactor.

[0126] At the end of range 209, a catalyst deactivator is injected into the reactor to accelerate the deactivation of the isomerization catalyst. Thereafter, during range 211, the WHSV is maintained at 20 hours. -1 No molecular hydrogen was co-fed into the reactor. Due to the deactivated catalyst, 20 hours -1 Due to the high WHSV of and the absence of molecular hydrogen co-fed to the reactor, the C(pX) in range 211 is low.

[0127] This Example 1 clearly demonstrates that in a liquid phase LPI process, by co-feeding molecular hydrogen at a feed rate of ≥100 ppm (based on the total weight of the C8 aromatic hydrocarbon isomerization feed), the activity of the isomerization catalyst is increased and the deactivation of the catalyst is greatly reduced compared to feeding molecular hydrogen at a lower feed rate or without co-feeding molecular hydrogen. Considering the high activity and extremely low deactivation rate of the isomerization catalyst shown at a high molecular hydrogen concentration, by co-feeding molecular hydrogen to the isomerization reactor at ≥100 ppm, it is possible to achieve a high deactivation rate in, for example, ≥10 hours. -1 , ≥15 hours -1 , ≥17.5 hours -1 and even ≥20 hours -1 LPI method at very high WHSV.

[0128] Example 2 (Comparative)

[0129] In this example, the C8 aromatic hydrocarbon feed stream contains 1.1 wt% para-xylene, 67.7 wt% meta-xylene, 29.8 wt% ortho-xylene, 1 wt% ethylbenzene and 0.5 wt% other hydrocarbons. In the isomerization process, the WHSV is changed and the feed rate of molecular hydrogen is maintained at 9 ppm by weight, based on the weight of the C8 aromatic hydrocarbon feed stream. The reaction temperature in the isomerization reactor is maintained at about 239°C. The reactor pressure is sufficient to maintain the molecular hydrogen co-fed to the isomerization reactor substantially completely dissolved in the liquid phase C8 aromatic hydrocarbons. Figure 7 In the display method as WHSV (hours -1 ) is a function of the concentration of p-xylene in the isomerization effluent based on the total amount of xylene (C(pX / X), wt %).

[0130] The feed stream used in this example is believed to be relatively easy to isomerize due to the low concentration of ethylbenzene therein. Therefore, similar performance of the catalyst at a high WHSV will be appreciated. Figure 7 It is clearly shown that as WHSV increases from 2.5 to 6 hours -1 , C(pX / X) decreased significantly, indicating that the catalyst deactivated as WHSV increased, even at 2.5-6 h. -1 This example shows that co-feeding molecular hydrogen at a low feed rate of 9 ppm does not provide adequate benefit for isomerization catalyst deactivation, even with a mild feed containing a low concentration of ethylbenzene.

[0131] Example 3 (Comparative)

[0132] In this example, the C8 aromatic hydrocarbon feed stream contains 8.1 wt% para-xylene, 58.1 wt% meta-xylene, 24.4 wt% o-xylene, 7.5 wt% ethylbenzene and 1.9 wt% other hydrocarbons. In the isomerization process, the WHSV is maintained at 4 hours. -1 No molecular hydrogen was fed to the isomerization reactor. The reaction temperature in the isomerization reactor was maintained at about 239°C. Figure 8 The p-xylene concentration (C(pX / X), wt %) in the isomerization effluent based on the total amount of xylenes is shown in FIG. 5 as a function of the time on stream (TOS, days) in the process.

[0133] The feed stream used in this example is believed to be relatively easy to isomerize due to its high concentration of paraxylene. Therefore, it would be desirable to isomerize the feed stream within 4 hours. -1 The high performance of the catalyst under WHSV of Figure 8 Clearly, C(pX / X) decreased significantly over a period of only 25 days, indicating that-1 At relatively low WHSV, the catalyst activity deactivates significantly in a relatively short time.

[0134] Considering this example, the positive effect of co-feeding molecular hydrogen at the high feed rate of > 100 ppm in Example 1 in reducing the deactivation of the isomerization catalyst is further illustrated.

Claims

1. A method for isomerizing C8 aromatic hydrocarbons, comprising: co-feeding molecular hydrogen and a liquid phase isomerized hydrocarbon feed comprising C8 aromatic hydrocarbons into an isomerization reactor having an isomerization catalyst disposed therein, wherein the co-feeding comprises feeding the molecular hydrogen at a feed rate of 200 ppm to less than or equal to 1000 ppm by weight, based on the total weight of the isomerized hydrocarbon feed; and contacting molecular hydrogen and C8 aromatic hydrocarbons with an isomerization catalyst in an isomerization reactor under isomerization conditions to produce an isomerization effluent, wherein the isomerization conditions include a reaction pressure in the isomerization reactor of 1,700 kPa-gauge to 3,500 kPa-gauge and a reaction temperature in the range of 240 to 300° C. such that ≥80% by weight of the C8 aromatic hydrocarbons are in a liquid phase in the isomerization reactor, and a weight hourly space velocity of 10 to 20 hours -1 .

2. The process according to claim 1, wherein molecular hydrogen is co-fed into the isomerization reactor at a feed rate of 200 ppm to 500 ppm by weight, based on the total weight of the isomerization hydrocarbon feed.

3. A process according to claim 1 or claim 2, wherein at least 95 wt% of the C8 aromatic hydrocarbons are in the liquid phase in the isomerization reactor.

4. The process of claim 1 wherein at least a portion of the molecular hydrogen is dissolved in the liquid phase of the isomerized hydrocarbon feed prior to being fed to the isomerization reactor.

5. The process of claim 1, wherein the isomerization conditions include a reaction temperature in the isomerization reactor of 260-300°C.

6. The process of claim 1, wherein the isomerized hydrocarbon feed comprises ethylbenzene in a concentration of no greater than 20 weight percent, based on the total weight of the isomerized hydrocarbon feed.

7. The process of claim 1, wherein the isomerized hydrocarbon feed comprises paraxylene in a concentration of no greater than 10 wt%.

8. The process of claim 1, wherein the isomerization catalyst contains no precious metals.

9. The method of claim 1, wherein the isomerization catalyst comprises a first metal element selected from the group consisting of Fe, Co, Ni, Ru, Rh, Pd, Re, Os, Ir, Pt, and combinations thereof, and optionally a second metal element selected from the group consisting of Sn, Zn, Ag, and combinations thereof.

10. The method according to claim 1, further comprising: The reaction temperature and / or weight hourly space velocity are increased without substantially reducing the paraxylene / xylene weight ratio in the isomerization effluent.

11. The method according to claim 1, further comprising: The reaction temperature and / or weight hourly space velocity are increased without significantly increasing xylene losses.

12. The method of claim 1 further comprising increasing the molecular hydrogen feed rate after an initial stage of the catalyst cycle.

13. A method for producing paraxylene, the method comprising: co-feeding molecular hydrogen and a liquid phase isomerized hydrocarbon feed comprising C8 aromatic hydrocarbons into an isomerization reactor having an isomerization catalyst disposed therein, wherein the co-feeding comprises feeding the molecular hydrogen at a feed rate of 200 ppm to less than or equal to 1000 ppm by weight, based on the total weight of the isomerized hydrocarbon feed; and contacting molecular hydrogen and C8 aromatic hydrocarbons with an isomerization catalyst in an isomerization reactor under isomerization conditions to produce an isomerization effluent, wherein the isomerization conditions include a reaction pressure of 1,700 kPa-gauge to 3,500 kPa-gauge and a reaction temperature of 240-300° C. such that ≥80 wt % of the C8 aromatic hydrocarbons are in a liquid phase in the isomerization reactor, and a weight hourly space velocity of 10-20 hours -1 ; the isomerization effluent comprises para-xylene at a higher concentration than the isomerization hydrocarbon feed; and At least a portion of the para-xylene is recovered from the isomerization effluent.

14. The process according to claim 13, wherein molecular hydrogen is fed to the isomerization reactor at a feed rate of 200 ppm to 500 ppm by weight, based on the total weight of the isomerization hydrocarbon feed.

15. A process according to claim 13 or claim 14 wherein at least 95 wt% of the C8 aromatic hydrocarbons are in the liquid phase in the isomerisation reactor.

16. The process of claim 15, wherein at least 98% of the C8 aromatic hydrocarbons are in the liquid phase in the isomerization reactor.

17. The process of claim 15, wherein at least 98% of the molecular hydrogen is dissolved in the liquid phase of the C8 aromatic hydrocarbons in the isomerization reactor.

18. The process of claim 13, wherein the isomerization catalyst contains no precious metals.

19. The method of claim 13, wherein the isomerization catalyst comprises a first metal element selected from the group consisting of Fe, Co, Ni, Ru, Rh, Pd, Re, Os, Ir, Pt, and combinations thereof, and optionally a second metal element selected from the group consisting of Sn, Zn, Ag, and combinations thereof.

20. The method of claim 13, further comprising: The reaction temperature and / or weight hourly space velocity are increased without substantially reducing the paraxylene / xylene weight ratio in the isomerization effluent.

21. The method of claim 13, further comprising increasing the molecular hydrogen feed rate after an initial phase of the catalyst cycle.

Images