KR20100040936A

KR20100040936A - Xylene production processes and apparatus with integrated feedstock treatment

Info

Publication number: KR20100040936A
Application number: KR1020107003498A
Authority: KR
Inventors: 안토니 네기즈; 에드윈 피 볼드잉흐; 제임스 이 레코스케; 에릭 제이 베이커
Original assignee: 유오피 엘엘씨
Priority date: 2007-07-30
Filing date: 2008-07-07
Publication date: 2010-04-21
Also published as: CN101808963A; JP2010535223A; PL391673A1; WO2009017937A3; MY149724A; TWI365183B; WO2009017937A2; TW200914415A

Abstract

A method and apparatus for the alkyl exchange of aromatics and olefin reduction in a feed stream are disclosed. The transalkylation conditions provide products with increased xylene concentrations and reduced olefin concentrations relative to the feed. Such methods and apparatus can be used in xylene production facilities to minimize or avoid the need for pretreatment of feedstocks such as hydration, hydrotreatment, or treatment with clay and / or molecular sieves.

Description

XYLENE PRODUCTION PROCESSES AND APPARATUS WITH INTEGRATED FEEDSTOCK TREATMENT}

The present invention relates to a process for transalkylation of aromatics and apparatus thereof. In particular, it is possible to reduce the olefin content of the feedstock in the process of the present invention, such as the transalkylation reaction of aromatics to produce xylene. The present invention is applicable to eliminating or reducing the need for pretreatment of feedstocks such as hydrotreating, hydrogenation, and treatment with clay and / or molecular sieves in xylene manufacturing facilities.

Xylenes, namely para-xylene, meta-xylene and ortho-xylene, are important intermediate products that find a wide variety of applications in chemical synthesis. Para-xylene produces terephthalic acid which, upon oxidation, is used for the production of synthetic fabric fibers and resins. Meta-xylene is used to make plasticizers, azo dyes, wood preservers, and the like. Ortho-xylene is a feedstock for phthalic anhydride. The distribution of xylene isomers from catalyst reforming or other sources generally does not match the distribution of the desired isomers for chemical intermediates, so manufacturers are converting feedstocks to produce more desired isomers.

The production of xylene is commercially practiced on large scale plants and is highly competitive. The concern is not only for the efficient conversion of feedstock to produce xylene through one or more of isomerization, transalkylation, and disproportionation, but also for such equipment, including capital and energy costs. There are other competitive aspects involved. Prior art aromatics complex flow schemes are described in Part 2 of Meyers' "Handbook of Petroleum Refining Processes" (2nd edition) published by McGraw-Hill.

Various sources have been proposed for monocyclic aromatics as feedstock for xylene production equipment. The most common sources are catalytic reforming of the naphtha fraction or hydration of the naphtha fraction after pyrolysis. Such processes typically produce a wide range of chemical compounds, including the desired monocyclic aromatics as well as polycyclic aromatics and olefins. Polycyclic aromatics and olefins are usually undesirable impurities in xylene production equipment. They can adversely affect product quality and process efficiency by requiring additional process steps, reducing the life of the catalyst, reducing the stability of the product, or causing undesired product color. Polycyclic aromatics are usually removed from the desired monocyclic aromatics by distillation. The polycyclic aromatics removed are then not worth discarding in any suitable manner, usually as fuel. It is also known that polycyclic aromatics can be converted to useful monocyclic aromatics such as toluene, xylene and C9 + monocyclic aromatics.

The quality of the feed stream for the various process units in the xylene production facility is also defined to ensure proper performance. For example, the olefin content of the streams supplied to some process units of xylene production equipment, including transalkylation units, is limited. Thus, olefins are recognized as contaminants in the transalkylation feed, so that in normal practice, the olefin content can be changed using various olefin removal processes such as hydration, hydrotreatment, treatment with clay and / or molecular sieves. Feed specification limit). Olefins are usually removed by clay treatment from the feedstock of the xylene production plant and / or intermediate product streams at various points within the plant. In clay treaters, olefins are converted to oligomers that can contaminate clay. The cost of operating a clay treater, including refilling the new clay into a clay treater or discarding organic contaminated used clay, can be a significant financial burden for commercial xylene producers. In addition, clay treaters can also alkylate olefins with aromatic rings. Thus, the effluent from the clay treater may contain aromatic rings with C 2+ substituents, such as ethylbenzene, propylbenzene, and methylethylbenzene. Thus, the value of the aromatic feedstock for the production of benzene, toluene and xylene is lowered.

The present invention makes it possible to remove olefins in an alkyl exchange process that reacts aromatics to produce xylene.

In one embodiment, the present invention provides a process for transalkylation and olefin removal from a feed stream comprising an aromatic compound having a Bromine Index of at least 50 and having at least 6 carbon atoms, wherein the concentration of xylenes is increased. Contacting the feed stream with an alkyl exchange catalyst comprising an acidic molecular sieve and at least one metal component under alkyl exchange reaction conditions to provide an alkyl exchange reaction product having a bromine index at least 60% lower than the bromine index of the feed stream. It relates to a method comprising the making.

In another embodiment, there is provided a method that can reduce capital and operating costs while generating xylene isomers from a feedstock comprising olefins and polycyclic aromatics in a xylene loop process. The xylene loop process separates the xylene loop feed by fractional distillation to provide a stream comprising xylene, subjecting the xylene stream to selective xylene isomer separation, and isomerizing the effluent stream from this selective xylene isomer separation By rebalancing the concentration of xylene isomers and circulating the isomerization effluent back to fractional distillation. The apparatus for this fractional distillation is also referred to herein as a xylene column.

In an embodiment of the present invention, at least a portion of the feedstock comprising C8 aromatics and olefins is mixed with a stream comprising C9 + aromatics, such as a high boiling fraction from the xylene column, and this mixed stream not only increases the concentration of xylenes. However, the olefin is removed and exposed to sufficient transalkylation conditions to convert the polycyclic aromatics to monocyclic aromatics.

Thus, the process of the present invention enables the use of aromatic feedstock containing olefins, thereby reducing or eliminating the need for pretreatment such as hydration, hydrotreatment, clay and / or molecular sieve treatment, thereby reducing capital and operating costs. .

1 is a schematic diagram illustrating an apparatus for carrying out a process according to the invention.
2 is a schematic diagram illustrating a reactor system having two separate reactor vessels for the transalkylation and isomerization reactions.
3 is a schematic representation of a reactor system with a single reactor vessel having two zones for the transalkylation and isomerization reactions.
4 is a schematic diagram illustrating an apparatus for carrying out a process according to the invention.

Processes and apparatus for the production of xylene isomers are disclosed, for example, in "Handbook of Petroleum Refining Processes" (2nd edition, Part 2, 1997) by Robert A. Meyers, published by McGraw-Hill. In conventional processes, the xylene production equipment feedstock comprising C8 aromatics and olefins is sorted to remove benzene and toluene and then fractionally distilled in a xylene column to provide a C8 aromatic stream from which one or more of the Recover the xylene isomer. The most desired isomer is para-xylene, but ortho-xylene and meta-xylene can also find commercial applications. After separation of the desired xylene isomer, the residual stream is isomerized and recycled to the xylene column, which also provides a high boiling fraction containing C9 + aromatics. Feedstock containing C8 may be classified in alternative ways, such as by first separating the C9 + aromatics and then by removing the C7- aromatics in the xylene column.

C8 containing aromatic feedstock for xylene production equipment is usually obtained from petroleum treatment, for example from catalytic reforming of naphtha fraction or from hydration treatment after pyrolysis treatment of naphtha fraction. This process typically produces a wide range of chemical compounds, including aliphatic (saturated and unsaturated) compounds and aromatic (monocyclic and polycyclic) compounds. The feedstock is usually characterized for the Initial Boiling Point and the End Boiling Point. The final boiling point is the temperature at which 99.5% by mass of the sample will boil, as determined by the simulated distillation GC method of ASTM method D2887. Typically, the final boiling point of the feedstock of the xylene production equipment is at least 210 ° C, and in other embodiments the final boiling point of the feedstock is 220 ° C. In one embodiment, the final boiling point is in the range of 240 ° C. to 280 ° C., and in another embodiment in the range of 340 ° C. to 360 ° C. Thus, the feedstock may contain highly substituted aromatic compounds and polycyclic aromatic compounds. Typically, the feedstock contains aliphatic compounds and low boiling point compounds, including aromatics such as benzene and toluene. Often, the feedstock has an initial boiling point of less than 80 ° C., and in some embodiments has an initial boiling point of less than 70 ° C. Thus, the feedstock may contain C5 and possibly lighter aliphatic according to any pretreatment such as distillation to remove C4 and lower hydrocarbons. Feedstocks can be custom made, such as by prefractionation, to include more selected groups of ingredients. In addition, multiple feedstocks may be provided simultaneously to the xylene production plant.

The feedstock of the xylene production plant may have the composition shown in Table 1.

ingredient Range (mass%) C6 and below aliphatic 2-15 C7 and C8 aliphatic 3-10 C8 or higher aliphatic 0-2 benzene 1-10 toluene 15-30 C8 aromatic 25-40 C9 aromatic 10-15 C10 + aromatic 1-7

Of the C 9+ aromatics in the feedstock, 1 to 50 mass%, often 3 to 30 mass%, are polycyclic aromatics. The olefin content of the feedstock and other streams is usually represented by the bromine index. Typically, the feedstock has a bromine index of at least 100, in some embodiments a bromine index of at least 300, and in other embodiments a bromine index of at least 600.

The bromine index is an indicator of olefin content. The bromine index is determined according to the procedure described in UOP Method 304-90, available through ASTM International (West Conchehoken Bar Harbor Drive 100, Pennsylvania, USA). According to this procedure, the sample is dissolved in a suitable solvent containing a catalyst that aids in the titration reaction. Potentiometric titration of this solution is carried out using a 0.01M bromide-bromate solution at room temperature. This titration uses a platinum detection electrode and a glass standard electrode with a recording potentiometric titrator. The bromine index is calculated in mg of bromine consumed per 100 grams of sample under test conditions. For the bromine index see ASTM D 1492 "Bromine Index of Aromatic Hydrocarbons by Coulometric Titration"; "Bromine Index of Aromatic Hydrocarbons by Electrometric Titration" by ASTM D 5776; It should be noted that there are at least three different standard test methods including ASTM D 2710, "Bromine Index of Petroleum Hydrocarbons by Electrometric Titration." Information on these ASTM methods can also be obtained from ASTM International. UOP methods 304-90 do not correspond to each of those methods and other methods of measuring the bromine index, and therefore, it should be understood that the numerical values for the bromine index herein are only measured by the UOP method.

In order to provide a suitable feed stream to the xylene loop, the feedstock of the xylene production plant comprising olefins and C8 aromatics needs to be treated to reduce the olefin content. According to an embodiment of the present invention, a feedstock containing C8 aromatics is mixed with a C9 + aromatic stream, which is preferably a C9 + aromatic containing high boiling fraction from a xylene column. Such mixtures are exposed to transalkylation conditions to provide transalkylation products with higher amounts of xylenes and reduced bromine index. The mass ratio of C8 aromatic containing feedstock to C9 + aromatic stream in the mixture can vary widely. Often, the ratio is 0.01: 1 to 3: 1 or more. In some embodiments, the ratio is 0.1: 1 to 2: 1, and in other embodiments, 0.3: 1 to 1.2: 1. The mixture will usually have a lower xylene concentration than the feedstock and often less than 35% by mass of the total aromatics in the mixture is xylene. In another embodiment, less than 30% by mass of the total aromatics in the mixture is xylene, and in yet other embodiments, 5-25% by mass of the total aromatics in the mixture is xylene. When light aromatics (benzene and / or toluene) are present in the mixture, the molar ratio of (benzene + toluene) to C9 + aromatics in the mixture is greater than 0.01: 1 in one embodiment. In another embodiment, the molar ratio of (benzene + toluene) to C9 + aromatics is greater than 0.5: 1, and in another embodiment, the ratio is 0.5: 1 to 2: 1.

Thus, the transalkylation process of the present invention can be used to produce xylene and reduce the olefin content of the feed. The term "alkyl exchange reaction" as used herein encompasses alkyl exchange reactions between alkyl aromatics, between benzene and alkyl aromatics, between benzene and alkyl naphthenic compounds such as methylcyclopentane and methylcyclohexane, Examples include the disproportionation of toluene to benzene and xylene. Thus, in one embodiment, the transalkylation feed includes an aromatic compound having at least six carbon atoms. The transalkylation process of the present invention may process a feed having a bromine index of at least 50. In one embodiment, the bromine index of the transalkylation feed is at least 100, in other embodiments at least 300, and in another embodiment the bromine index of the transalkylation feed is at least 500. The transalkylation conditions are sufficient to provide a transalkylation product having a higher xylene concentration than the transalkylation feed and at least 60% lower the bromine index than the transalkylation feed. In some embodiments, the bromine index of the product is at least 80% lower than the feed. In one embodiment, the bromine index of the product is less than 30 when the bromine index of the feed is at least 200. In another embodiment, the bromine index of the product is less than 20. In another embodiment, the bromine index of the product is less than 10. The transalkylation process may be part of the xylene production center, part of another configuration of the process units, or may be a single unit. The aromatic composition of the transalkylation feed may vary significantly. The transalkylation feed includes an aromatic compound having at least six carbon atoms. In one embodiment, the transalkylation feed comprises at least one of toluene and C9 + aromatics. This feed may optionally include one or both of benzene and C8 aromatics. In another embodiment, the transalkylation feed comprises at least one of benzene and toluene and a C9 + aromatic. In one embodiment, the final boiling point of the transalkylation feed is at least 190 ° C. In another embodiment, the final boiling point of the transalkylation feed is at least 220 ° C. and in still other embodiments at least 240 ° C. In one embodiment, the final boiling point of the transalkylation feed is from 240 ° C. to 280 ° C. and in another embodiment from 340 ° C. to 360 ° C.

Particular transalkylation conditions will depend in part on the composition of the transalkylation feed as well as the catalyst and its activity. Generally, the transalkylation conditions include, for example, high temperatures of 100 ° C to 425 ° C, preferably 200 ° C to 400 ° C. In commercial installations, the transalkylation temperature is often elevated to compensate for any decrease in activity in the catalyst. The feed to the transalkylation reactor is usually first heated by indirect heat exchange with respect to the effluent of the reactor and then heated to the reaction temperature by heat exchange with a hot stream or steam or by a furnace. The feed then passes through the reaction zone, which may comprise one or more individual reactors containing the catalyst of the present invention. The reactor may be in any suitable form or configuration. While the use of a single reaction vessel having a cylindrical fixed catalyst bed is preferred, other reaction configurations using moving catalyst beds or radial flow reactors may be used if desired. The transalkylation conditions include a pressure in the range of 100 kPa to 10 MPa (absolute pressure), preferably in the range of 0.5 MPa to 5 MPa (absolute pressure). The transalkylation reaction can take place over a wide range of space velocities. Is: (WHSV weight hourly space velocity) usually ranges 0.1hr ^-1 to 30hr ^-1, preferably from 0.5hr ^-1 to 20hr ^-1 range, usually it is 1hr ^-1 to 5hr ^-1 range weight space velocity.

By means of the present invention, various combinations of transalkylation conditions or operating parameters, including catalysts, still provide increased concentrations of xylenes while providing the desired reduction of olefins in the transalkylation feed without causing excessive aromatic ring loss. It was confirmed that the transalkylation reaction can be achieved. A preferred combination of transalkylation conditions is a combination such that the net make of methane is less than 0.5 mass%. In one embodiment, the net production of methane is less than 0.2 mass%. The net production of methane is the increment of the methane concentration from the methane concentration of the feed to the methane concentration of the product. The ring loss, represented by the mole% loss of monocyclic aromatic rings in the product as compared to the feed, is preferably less than 2 mole%. In one embodiment, the ring loss is less than 1.5 mole percent.

The transalkylation reaction can be in the gas phase or in the liquid phase in the presence of hydrogen. For liquid phase transalkylation, the addition of hydrogen is optional but preferred. When the feed is alkylated in the gas phase, hydrogen is usually added in an amount of from 0.1 mole per mole of hydrocarbon to 10 mole per mole of hydrocarbon in the mixture fed to the transalkylation reaction unit. This operating parameter is referred to as the hydrogen to hydrocarbon ratio and is often referred to as 5.5 H ₂ / HC, for example, meaning 5.5 moles of hydrogen per mole of hydrocarbon. If the transalkylation reaction is in the liquid phase, the transalkylation reaction can be carried out substantially without hydrogen except that it may already be present and dissolved in the usual liquid aromatic feedstock. In the case of partial liquid phase, hydrogen may be added in an amount of less than 1 mole per mole of hydrocarbon. Preferably, the minority to hydrocarbon ratio in the liquid, partial liquid or gaseous transalkylation mode is at least 1.5 H ₂ / HC. In one embodiment, the ratio is at least 2 H ₂ / HC, and in other embodiments from 2 H ₂ / HC to 8 H ₂ / HC.

The transalkylation process provides xylenes, whereby the transalkylation product contains a higher concentration of xylenes relative to the feed. When light aromatics (benzene and toluene) are included in the feed, the molar ratio of light aromatics to C9 + aromatics in the feed is greater than 0.01: 1, in one embodiment the ratio is from 0.1: 1 to 10: 1. In another embodiment, the molar ratio of (benzene + toluene) to C9 + aromatics in the feed is 0.9: 1 to 5: 1.

In one embodiment, the transalkylation feed comprises at least 1 mass% polycyclic aromatic. During the transalkylation reaction, the polycyclic aromatics are converted to a significant degree. In one embodiment, at least 50 mole percent of the polycyclic aromatics are converted to monocyclic aromatics. Thus, according to embodiments of the present invention, it is possible to convert indene and naphthalene to alkyl substituted monocyclic aromatics with or without alkyl substitution. Lower products of inyne and naphthalene during the conversion to monocyclic aromatics can provide a source of alkyl family. Importantly, the process of the present invention can convert polycyclic aromatics to monocyclic aromatics without excessive loss of monocyclic half groups, providing high selectivity for the desired monocyclic alkyl aromatics.

In one embodiment, the transalkylation reaction is carried out for a time sufficient for at least 10 mole% of heavy alkyl aromatics (C9 +) to be consumed and under other conditions. In another embodiment, at least 20 mole percent heavy alkyl aromatics are consumed and in another embodiment 20 to 90 mole percent heavy alkyl aromatics are consumed. In one embodiment, at least 70 mole percent of the heavy alkyl aromatics consumed are converted to low molecular weight aromatics. In another embodiment, at least 75 mole percent of the heavy alkyl aromatics consumed are converted to low molecular weight aromatics. The transalkylation conditions are preferably sufficient to provide an transalkylation product having a final boiling point at least 5 ° C. lower than the transalkylation feed. In another embodiment, the transalkylation product has a final boiling point at least 10 ° C. lower than the transalkylation feed.

The transalkylation conditions include the presence of an transalkylation catalyst comprising a metal component and an acidic molecular sieve component. In one embodiment, the catalyst comprises an acidic molecular sieve, a metal component and an inorganic oxide component. Metal components typically have a hydrogenation function. The metal component includes at least one of precious metals and base metals. The precious metal is a platinum group metal and is selected from the group consisting of platinum, palladium, rhodium, ruthenium, osmium, iridium and mixtures thereof. Base metals are selected from the group consisting of rhenium, tin, germanium, lead, iron, cobalt, nickel, indium, gallium, zinc, uranium, dysprosium, thallium and mixtures thereof. Accelerators or modifiers may also be used in the catalyst formulation. Such promoters or modifiers are at least one of base metals, IUPAC groups 1, 2, 5, 6, 7, 11, 12, 13, 14, 15, 16 and 17. In one embodiment, the metal component comprises platinum and tin. In another embodiment, the metal component comprises rhenium.

Stable amounts of metal components reduce the negative effects of polycyclic aromatics on catalysis. That is, the stability of the catalyst is improved. In one embodiment, the deactivation rate of the catalyst is less than 200% of what is observed for the same catalyst under substantially the same conditions except that the transalkylation feed contains less than 0.5 mass% C11 + aromatics. In another embodiment, less than 150%. The degradation is measured as the temperature rise required to maintain a constant aromatic conversion under other constant conditions. The rate of deactivation must be determined after the catalyst has been lined-out, for example after operation for 100 hours. Since it is impractical to perform comparative tests in commercial plants, stability determinations are preferably made in pilot plants using synthetic feeds.

In order to determine whether the catalyst contains a stable amount of metal components, evaluation conditions can be conveniently used. These evaluation conditions are only used for catalyst evaluation and do not necessarily form the typical alkyl exchange reaction conditions used in accordance with the present invention. Evaluation conditions were weight average bed temperature 1.6hr ^-1 , pressure 2760 kPa (gauge pressure), hydrogen to hydrocarbon ratio 6: 1, and weight average bed temperature sufficient to convert 45 mole percent aromatics in the aromatic containing feed. bed temperature) involves the operation of the pilot plant. Suitable polycyclic aromatic feeds for the evaluation conditions are included in the ranges shown in Table 2.

Feed Ingredient Polycyclic Feeds (mol%) toluene 28-32 C9 monocyclic aromatic 55-60 C10 monocyclic aromatic 5-7 Indein 0.3-0.7 naphthalene 0.25-0.4 C11 + aromatic 4-7

For the reference feed, the polycyclic aromatic feed (toluene not included) is distilled to contain less than 0.5 mole percent C11 + aromatics and then toluene is added to keep the molar ratio of toluene to C9 + aromatics the same. C11 + aromatics include monocyclic and polycyclic aromatics.

In catalysts that are advantageous for use in the process of the present invention, the amount of metal components is less than the amount that results in excessive aromatic ring loss. Under the evaluation conditions, the ring loss is preferably less than 2 mol% based on the total moles of monocyclic aromatic compounds in the feed. Appropriate metal amounts in the transalkylation catalyst will depend on the metal present as well as any promoter or modifier. In one embodiment, the amount of gold component is in the range of 0.01 to 10 mass% of the catalyst. In another embodiment, the amount of metal component is in the range of 0.1 to 3% by mass of the catalyst, in another embodiment in the range of 0.1 to 1% by mass.

Acidic molecular sieve components include one or more molecular sieves. Molecular sieves include BEA, MTW, FAU (including zeolite Y in both cubic and hexagonal forms, zeolite X), MOR, LTL, ITH, ITW, MEL, FER, TON, IWW, MFI, EUO, MTT, HEU, MFS, CHA, ERI, and LTA are included, but are not limited to these. Molecular sieves with known forms of structure have been classified according to the three-letter notation (available from the Internet website www.iza-structure.org/databases) according to the Structural Commission of the International Zeolite Association, and such codes Is used herein. In one embodiment, the acidic molecular sieve component comprises at least one of MOR and MFI. The molecular sieve component is preferably at least partially in hydrogen form in the finished catalyst. The acidity of the molecular sieve may be the acidity of the molecular sieve to be used when preparing the catalyst of the present invention, or may be obtained during the preparation of the catalyst. In one embodiment, the acidic molecular sieve component has a total acidity of at least 0.15 as determined by Ammonia Temperature Programmed Desorption (Ammonia TPD). In another embodiment, the total acidity of the molecular sieve component is at least 0.25, in another embodiment the total acidity of the molecular sieve component is at least 0.4, and in still other embodiments, the total acidity of the molecular sieve is in the range of 0.4 to 0.8. .

Ammonia TPD first involves heating a sample of molecular sieve (250 milligrams) to a temperature of 550 ° C. at a rate of 5 ° C. per minute in a helium atmosphere (100 ml flow rate per minute) in which 20 volume percent oxygen is present. After holding for 1 hour, the system is flushed with helium (15 minutes) and the sample is cooled to 150 ° C. The sample is then saturated by 40 ml of helium diluted ammonia pulses per minute. The total amount of ammonia used is an amount that greatly exceeds the amount needed to saturate all acid sites in the sample. The sample is purged with helium (40 ml / min) for 8 hours to remove physisorbed ammonia. The temperature is raised to a final temperature of 600 ° C. at a rate of 10 ° C. per minute while continuing helium purging. The amount of ammonia desorbed is monitored using a calibrated thermal conductivity detector. The total amount of ammonia is obtained by integration. The total acidity is obtained by dividing the total amount of this ammonia by the dry weight of the molecular sieve sample. As used herein, the value of total acidity is expressed in millimoles of ammonia per gram of dry molecular sieve.

When mordenite is a component of the catalyst, mordenite is preferably less than 40: 1 molar ratio of Si / Al2. In one embodiment, the Si / Al2 molar ratio of mordenite is less than 25: 1, and in another embodiment, the Si / Al2 molar ratio of mordenite is 15: 1 to 25: 1. Mordenite may be synthesized in a Si / Al 2 molar ratio of 10: 1 to 20: 1. Such mordenite can be used by itself as synthesized or dealuminated before or after inclusion in the catalyst.

The MFI molecular sieve used in the catalyst of the invention preferably has a Si / Al2 molar ratio of less than 80. In one embodiment, the Si / Al2 molar ratio of MFI is less than 40: 1 and in other embodiments less than 25: 1. In yet another embodiment, the Si / Al2 molar ratio of MFI is from 15: 1 to 25: 1. MFI can be used by itself or dealuminated. When dealuminated, the activity of the catalyst is improved, but excessive dealumination causes the transalkylation product to contain more components that boil together with benzene. While not wishing to be bound by theory, dealumination can cause some mesoporosity to be introduced into the MFI structure. It is believed that mesopores in this MFI structure can improve the overall conversion.

The dealumination may be accomplished by any suitable technique, such as acid treatment and / or steaming. In the case of using the vaporized molecular sieve, 2 to 50% by volume of a pressure of 100 kPa to 2 MPa and a temperature of less than 650 ° C, for example, 500 ° C to 600 ° C, more preferably 550 ° C to 600 ° C. Preferably, the steam is gently steamed using 5 to 30% by volume of steam. Steam calcination may take place before or after shaping the molecular sieve into the desired catalyst form. Preferred catalysts were obtained when the steaming took place after catalyst shaping.

In one embodiment, the acidic molecular sieve component comprises MOR and MFI. In one embodiment, the mass ratio of MFI to mordenite is in the range from 1:10 to 5: 1 and in another embodiment from 1:10 to 1: 1. Preferably, the mordenite is at least partially in hydrogen form in the finished catalyst. In one embodiment, the catalyst comprises rhenium with or without other metal components. See, for example, US Patent Application Nos. 60 / 825,306 and 60 / 825,313, both of which are filed on September 12, 2006, which are incorporated by reference in their entirety. Is cited.

Preferably an appropriate inorganic oxide component (heat resistant binder or matrix) can be used to facilitate the preparation of the catalyst, to provide strength and to reduce the manufacturing cost. Suitable binders include inorganic oxides such as at least one of alumina, magnesia, zirconia, chromia, titania, boria, toria, aluminum phosphate and zinc oxide. Preferred inorganic oxide binders include alumina, in particular transition alumina and gamma alumina. Particularly useful aluminas are available under the trade names CATAPAL B and VERSAL 250. In one embodiment, the molecular sieve component is present in the range of 5 to 99 mass% of the catalyst and the heat resistant inorganic oxide may be present in the range of 1 to 95 mass% of the catalyst when used.

Treatment techniques for preparing the catalyst can affect the performance of the catalyst. For example, occlusion of catalytically active sites can occur. Therefore, care should be taken not to overly impair the activity of the catalyst. The metal component may be included in the catalyst in any suitable manner such as co-kneading with the carrier material, co-precipitation, or co-gelling, ion exchange, impregnation and the like. This component may be present as a compound, such as an oxide, sulfide, halide, or oxyhalide, in the final catalyst composite, in chemical association with at least one of the other components of the complex, or as an elemental metal. have. One method of preparing the catalyst involves impregnating the molecular sieve containing carrier with a water soluble or solvent soluble degradable compound of the metal in question. Alternatively, metal compounds may be added when synthesizing the molecular sieve component and the binder. Another method of preparing the catalyst involves kneading the molecular sieve component, the metal component and the binder together to provide an extrudable mixture to be shaped into the desired catalyst shape. One shape of the catalyst of the present invention is cylindrical. Such cylinders can be formed by extrusion methods known in the art. Another shape of the catalyst is to have a cross-section in the form of trilobal or wash clover which can be formed by extrusion. Another shape is a sphere, which can be formed using oil dropping or other molding methods known in the art.

Preferably, the preparation of the catalyst comprises at least one oxidation or calcination step, especially when the metal is rhenium. The oxidation step is usually at a temperature of 370 ° C to 650 ° C. Usually, an oxygen atmosphere containing air is used. In general, the oxidation step is carried out for substantially the time necessary for all of the metal components to be converted to their corresponding oxide forms, usually from 0.5 to 10 hours or more. This time will of course depend on the metal component used to make the catalyst, the oxidation temperature used, and the oxygen content in the atmosphere used. Sometimes steam is present in the calcining in an amount of 5 to 70% by volume, for example 5 to 40% by volume.

If a rhenium containing catalyst is used, the catalyst is preferably sulfided. Rhenium is preferably in at least partial oxide form upon sulfiding. In some embodiments, the catalyst is partially reduced before or during sulfidation. The degree of reduction depends on the reducing atmosphere, the reducing temperature and the reducing time. In excess reducing conditions, especially under conditions involving high temperatures, the dispersion of rhenium in the catalyst can have a negative effect. When the reduction is carried out prior to sulfidation, the reduction temperature is preferably below 400 ° C., in one embodiment in the range of 100 ° C. to 350 ° C. The reduction time is preferably such that rhenium has an oxidation state of at least +4 so that excessive aggregation of rhenium does not occur in the catalyst. Thus, the reduction is done for a time of less than 24 hours, and shorter times are used for higher temperatures. In one embodiment at 12 ° C. or less for 12 or less, in another embodiment at 5 ° C. for 5-6 hours.

The sulfidation may occur simultaneously with at least a portion of the reduction step or after the reduction. Preferably, the reduction is in an environment substantially free of moisture. Preferably, the reducing gas is substantially pure dry hydrogen (ie less than 20 ppm by weight of moisture). However, other gases such as hydrocarbons, CO, nitrogen and the like may also be present. The reduction step can be carried out at atmospheric or higher pressure. Preferred pressures are 50 kPa (absolute pressure) to 10 MPa (absolute pressure), and in one embodiment, the pressure ranges from 200 kPa to 5000 kPa (absolute pressure).

The sulfur component can be included in the catalyst by any known technique. Any one or combination of in situ and / or ex-situ sulfur treatment methods may be used. Efficient treatment is achieved by contacting the catalyst with a sulfur source at a temperature in the range of 0 ° C to 500 ° C. The sulfur source may be in direct contact with the catalyst or via a carrier gas such as, for example, hydrogen or silane. In one embodiment, the sulfur source is hydrogen sulfide. Other sulfur sources may be used. Examples are alkyl sulfides or carbon disulfides such as methyl sulfide, dimethyl sulfide, dimethyl disulfide, diethyl sulfide, and dibutyl sulfide. Preferably the catalyst is sulfided in the presence of carbenium ions. Since carbenium ions are intermediate chemical reaction products, the most convenient way to introduce carbenium ions is to cause hydrocarbon cracking during sulfiding. The hydrocarbon can be any suitable compound that can be cracked under sulfiding conditions, and in one embodiment the hydrocarbon comprises at least one of ethylbenzene, methylethylbenzene and propylbenzene. The acidic molecular sieve component helps to produce carbenium ions. In one embodiment, the sulfidation is in the range of 250 ° C. to 500 ° C., and in another embodiment in the range of 250 ° C. to 400 ° C. If the rhenium containing catalyst does not have sufficient cracking activity, higher temperatures may be required.

In one embodiment, sulfidation is achieved by passing a sulfur containing gas over the catalyst at a weight space velocity of at least 0.5 hr ⁻¹ . The catalyst composition may be sulfided in situ, in which case the sulfur source is catalyzed by adding a hydrocarbon feed stream of sulfur concentration in the range of 1 to 5,000 or 10,000 volume ppm, preferably in the range of 5 to 500 volume ppm to the catalyst composition. Contact with the composition. The need to add a sulfur source to the hydrocarbon feed stream can be reduced or completely eliminated depending on the actual content of sulfur that may already be present in some hydrocarbon streams.

Depending on the concentration of sulfur in the feed, sulfidation can be achieved in less than one hour, or over a longer time period, such as for example one or more days. Sulfur treatment can be monitored by measuring the concentration of sulfur in the resulting off gas. The time calculated for the sulfur treatment will depend on the actual concentration of sulfur in the feed and the amount of sulfur desired to be achieved for the catalyst. In some instances, it has been found that even if more sulfur is provided than is required for the targeted sulfur to rhenium atomic ratio, the ratio of sulfur to rhenium seems to reach a certain level and rhenium is not excessively sulfided.

The sulfidation is usually made such that 0.2 to 0.7 sulfur atoms are provided per rhenium atom, and in another embodiment 0.25 to 0.5 sulfur atoms per rhenium atom. If the catalyst comprises other components that can adsorb or react with sulfur, the total amount of sulfur provided should be sufficient to ensure that the desired amount of sulfur is provided for rhenium.

The catalyst can be regenerated by calcination. Regeneration conditions are generally controlled in a controlled manner to carbon the carbonaceous deposits on the catalyst at a temperature of 370 ° C. to 650 ° C. for 0.5 to 24 hours in an oxygen containing atmosphere such as, for example, air or air with added nitrogen and / or steam. It involves burning up.

As such, in one embodiment, the present invention provides a process for transalkylation and olefin removal from a feedstock comprising C8 aromatics and having a bromine index of at least 100, wherein the feedstock and a stream comprising C9 + aromatics have a bromine index of at least Mix so that the final boiling point is 50 and at least 210 ° C. Such a mixture is contacted with an alkyl exchange catalyst comprising an acidic molecular sieve and at least one metal component under alkyl exchange reaction conditions to increase the xylene concentration and have an bromine index of at least 60% of the bromine index of the mixture. To provide the product.

The transalkylation process can be integrated with other processes by various flow techniques. In one embodiment, the transalkylation product is sent to the xylene distillation column as xylene loop feed. At least one desired xylene isomer is recovered from the xylene column product stream comprising xylene in a xylene isomer separation apparatus. The disproportionate xylene isomer stream produced in the xylene isomer separation unit is sent to a xylene isomerization reaction zone that produces an isomerization product with a balanced distribution of xylene isomers. The isomerization product is recycled to the xylene column. Other reactions such as transalkylation and cracking may occur under isomerization conditions. Thus, isomerization can be a source of high molecular weight aromatics, such C9 + aromatics will also be contained in the heavy aromatic containing fractions from the xylene column.

In some embodiments, a purge is taken in the heavy aromatic containing fraction from the xylene column to prevent any excessive generation of polycyclic aromatics. The purge can be used in any suitable manner and can be used as fuel if desired. Alternatively, the purge is separated by distillation to recover C9 and C10 aromatics that can be used as by-products or sent as transalkylation reactors.

Detailed description of the drawings

The present invention will be described in more detail with reference to FIG. 1, which schematically illustrates a xylene production apparatus 100. The illustrated apparatus is not intended to limit the scope of the process of the invention but to aid in understanding the principles of the invention.

A feedstock comprising C8 aromatics and olefins is provided to xylene production equipment via line 102. The feedstock can be obtained from any suitable source. The feedstock can also provide a significant proportion of light aromatics (benzene and toluene) useful for the alkyl exchange reaction of C9 + aromatics in the mixed stream. In the illustrated embodiment, the feedstock is sent to feed distillation apparatus 158 to provide a C4- hydrocarbon stream by fractional distillation. The bottoms stream from feed column 158 is sent through line 162 to merge with a portion of the high boiling fraction from the xylene column described below in line 142. The mixing of these streams may be in line 142 as shown, or in an alkyl exchange reactor 144. In some embodiments, the bottoms stream of the feedstock and feed column containing C8 aromatics also comprises light aromatics (benzene and / or toluene), and thus benzene to be recovered by distillation and recycled to the transalkylation zone. Operational savings in that the amount of toluene can be reduced or even eliminated. Advantageously, the bottoms stream of the feed column comprising C8 aromatics and olefins can be sent directly to the transalkylation reaction zone. The expression “immediately” herein means that the transalkylation feed does not undergo an olefin removal step such as hydration, hydrotreatment, or treatment with clay and / or molecular sieves prior to the transalkylation reaction. However, the transalkylation feed may be subjected to other treatments, such as drying, if necessary to meet other feed specifications. Likewise, a custom-made feedstock for xylene production apparatus 100 comprising an aromatic compound having at least six carbon atoms and olefins bypasses the feed column, so that it is not processed in the olefin removal zone but directly to the transalkylation reaction zone. Can also be sent. That is, as is known in the art, the feedstock for the xylene production apparatus is prepared by fractionation, non-aromatic removal, and other processes or processing steps, prior to introduction into the xylene production apparatus 100, to distill the feed. Device 158 may be unnecessary. As is also known, feed distillation apparatus 158 can generate a variety of effluent streams, for example, by side cuts and / or using a plurality of columns as described below. . Non-limiting illustrative examples of such effluent streams or “cuts” include light end rich materials, benzene, toluene, C8 aromatics and C9 + aromatics.

The resulting mixture is sent to an alkyl exchange reaction zone. As described below, the transalkylation reaction zone comprises a variety of configurations, thus more generally referred to as reaction zone 144. Additional light aromatics (benzene and / or toluene) may be added to the mixture via line 146. Benzene and toluene may be obtained from any suitable source, such as other feed streams for xylene production equipment and / or lines 106, 112, 132, 138 as described below. The feedstock may also provide a significant proportion of light aromatics (benzene and / or toluene) useful for the alkyl exchange reaction of C9 + aromatics in the mixed stream. The transalkylation reactor may comprise one or more individual reactors containing the catalyst of the invention. These reactors can be of any suitable form and configuration. While it is preferable to use a single reaction vessel with a cylindrical fixed catalyst bed, other reaction structures using moving catalyst beds or radial flow reactors may be used if desired.

As mentioned above, the transalkylation reaction according to the process of the present invention provides an transalkylation reaction product containing increased concentrations of xylenes and reduced concentrations of olefins. Thus, the reaction zone comprises at least one inlet which accepts xylene production equipment feedstock directly, or immediately after fractionation in the feed column, ie the transalkylation feed is not passed through the olefin removal zone. The transalkylation or reaction zone product is discharged to the reaction zone product distillation unit at the outlet of the reaction zone, where the reaction zone product is separated into several streams which are recycled in the xylene production plant and / or discharged as product. In the illustrated embodiment, the reaction zone product distillation apparatus includes a stripper column 150, a benzene column 104 and a toluene column 110. The transalkylation or reaction zone product is sent via line 148 to stripper column 150 to remove light material separated by fractional distillation via line 152. The stripper column 150 also provides a high boiling point C6 + fraction that is sent to the benzene column 104. Benzene column 104 separates benzene from the C7 + aromatics to be recovered as bottoms fraction by fractional distillation and recovers it as top fraction in line 106. The bottom fraction from the benzene column 104 is sent to the toluene column 110 via line 108. Toluene column 110 provides the top fraction of toluene removed via line 112 by fractional distillation and the bottom fraction C8 + aromatics sent to line xylene column 116 via line 114.

In the embodiment not shown, toluene may be converted to xylene and benzene in an optional toluene disproportionation reaction zone. The operation of the toluene disproportionation unit depends on the feed and economics for a given plant. The disproportionation reaction takes place via a catalytic reaction, and any such process known in the art can be used in connection with the present invention. Preferred disproportionation reaction conditions for increasing the resulting para-xylene include precoked catalysts as disclosed in US Pat. No. 6,429,347, which is incorporated herein by reference in its entirety.

Xylene column 116 separates by means of fractional distillation a low boiling point C8 aromatic stream containing at least two xylene isomers. In another embodiment, the effluent stream of the xylene column comprising C8 aromatics is a high boiling product stream. At least one xylene isomer is recovered from this C8 aromatic stream. If the desired isomer is ortho-xylene, separation is usually by fractional distillation. Para-xylene isomers may be recovered by selective crystallization fractionation or by selective adsorption such as simulated mobile bed adsorptive separation. Recovery of meta-xylene by sorption process is also known. Xylene isomer separation processes are known to those skilled in the art and any known process may be used individually or in combination to obtain one or more desired isomers. Thus, the C8 aromatic stream is sent from the first outlet of the xylene column 116 via line 118 to the inlet of the optional xylene isomer separation device 120.

Optionally, an olefin removal device 117 is provided in line 118 to remove any olefins. By the process of the present invention, the transalkylation reaction removes almost all olefins from the feedstock of the xylene production plant. Thus, the low boiling point C8 aromatic stream in line 118 will only contain the remaining amount contained in the xylene loop and the transalkylation reaction product with the small amount of olefin produced, for example, by isomerization. In some embodiments, the olefin removal apparatus 117 may be completely removed when there is little production of olefins in the isomerization reaction zone. In other embodiments, such as where the xylene isomerization reaction zone produces higher amounts of olefins, olefin removal apparatus 117 may be required. Nevertheless, the olefin removal apparatus 117 can be relatively compact and have a long service life, thereby saving capital and operating costs. In addition, the olefin removal zone may preferably only be located in one place in the xylene production apparatus 100 to treat the C8 aromatic stream sent from the xylene column outlet. In this embodiment, the optional xylene isomer separation device 120 is a selective sorption unit and para-xylene is withdrawn via line 122 via a first outlet. Such disproportionate isomer streams with reduced para-xylene, also referred to as raffinate, are sent from the xylene separator 120 via a second outlet to the isomerization reaction zone 126 via line 124.

Isomerization of xylenes serves to rebalance such an imbalanced mixture of xylenes. For example, when para-xylene is removed as the desired product, additional para-xylene is produced by bringing the xylene isomers back into balanced or near equilibrium. Any ethylbenzene in the raffinate of the para-xylene separation unit is converted to additional xylenes or to benzene by dealkylation depending on the type of isomerization catalyst and other conditions. In one embodiment, conversion of ethylbenzene by dealkylation is preferred in order to reduce the complexity in handling large amounts of naphthenes usually involved in the conversion of ethylbenzene to xylene. In one such embodiment, the total naphthenes in the feed to the isomerization reactor 126 is less than 1 mass%, preferably less than 0.7 mass%. Under favorable conditions for the isomerization of ethylbenzene, hydrogen and hydrocarbons are present in a molar ratio of hydrogen to hydrocarbons of at least 0.1: 1, for example 0.5: 1 to 6: 1, preferably 1.5: 1 to 5: 1. Preferably, the feed stream contains naphthenes, more preferably naphthenes are provided in the feed stream at a concentration sufficient to enhance the conversion of ethylbenzene, for example 2 to 20 mass%. Preferably the isomerization is at least partially done in gaseous conditions. The isomerization section may comprise the operation of the hydrogenation unit followed by the operation of the dehydrogenation unit.

Xylene isomerization involves the use of an isomerization catalyst under isomerization reaction conditions to provide isomerates. Isomerization catalysts usually consist of molecular sieve components, metal components and inorganic oxide components. The choice of molecular sieve component allows control of catalyst performance between isomerization of ethylbenzene and dealkylation of ethylbenzene depending on the overall requirements for benzene. Thus, the molecular sieve may be zeolite aluminosilicate or non-zeolitic molecular sieve. The zeolite aluminosilicate (or zeolite) component can be pentasil zeolite, MWW, beta zeolite, or mordenite including the structures of MFI, MEL, MTW, MFS, MTF and FER (IUPAC regulations for zeolite nomenclature). . The non-zeolitic molecular sieve is usually at least one of the AEL structure forms according to "Atlas of Zeolite Structure Types" (Butterworth-Heineman, Boston, Mass., 3rd ed. 1992), in particular SAPO-11, at least one of the ATO structure forms, In particular MAPSO-31.

The metal component is usually a noble metal component and may include an optional nonmetal modifier component in addition to or instead of this noble metal. The precious metal is a platinum group metal selected from platinum, palladium, rhodium, ruthenium, osmium, and iridium. Base metals are selected from the group consisting of rhenium, tin, germanium, lead, iron, cobalt, nickel, indium, gallium, zinc, uranium, dysprosium, thallium and mixtures thereof. Base metals may be mixed with other base metals or precious metals. In one embodiment, the total metal in the isomerization catalyst is in the range of 0.01 to 10% by mass and in other embodiments the amount of metal is in the range of 0.1 to 3% by mass. Suitable zeolite amounts in the catalyst range from 1 to 99 mass%, in one embodiment from 10 to 90 mass%, and in another embodiment from 25 to 75 mass%. The remainder of the catalyst consists of an inorganic oxide binder, usually alumina. In some instances, it may be desirable to modify the catalyst by sulfiding in situ or ex situ or the like. One isomerization catalyst for use in the present invention is disclosed in US Pat. No. 4,899,012, which is incorporated herein by reference in its entirety.

Typical isomerization reaction conditions include temperatures of 0 ° C. to 600 ° C. and pressures of 100 kPa to 6 MPa (absolute pressure). The hydrocarbon weight space velocity of the feedstock relative to the weight of the catalyst is 0.1 to 30 hr ⁻¹ . The hydrocarbon is contacted to the catalyst in a mixed state with hydrogen in gaseous phase at a molar ratio of hydrogen to hydrocarbon of 0.5: 1 to 15: 1 or more, in another embodiment H ₂ / HC of 0.5 to 10. When liquid phase conditions are used for isomerization, hydrogen is usually not added. The isomerization reactor may comprise one or more individual reactors containing the catalyst and may be in any suitable form and configuration. While the use of a single reaction vessel having a cylindrical fixed catalyst bed is preferred, other reaction configurations using moving catalyst beds or radial flow reactors may be used if desired.

In the illustrated embodiment, the isomerate from the isomerization reaction zone 126 is sent to the lights column 130 via line 128, which is passed through line 132, for example benzene. And low boiling oil fractions such as hard components. A stream containing high boiling point C7 + aromatics is sent from light column 130 to deheptanizer column 136 via line 134. Toluene fraction is provided in the deheptan column 136 by fractional distillation and removed via line 138. The bottoms stream containing the C8 + aromatics is sent to line xylene column 116 via line 140. Similar to the reaction zone product distillation apparatus, various device configurations can be used to separate the C8 + aromatics from the isomerates for recycling to the xylene column. Thus, the construction of columns such as light column 130 and deheptane column 136 used to achieve this separation is referred to as isomerate distillation apparatus. Thus, the isomerate produced in the xylene isomerization reactor 126 is indirectly sent to the xylene column 116. The fraction comprising benzene and toluene from the isomerate distillation unit can be sent to another distillation column in the xylene production plant for further separation.

Looking back at the xylene column 116, the feed stream for the xylene column consisting of a bottom fraction from the toluene column 110 via line 114 and a bottom fraction from the deheptan column 136 through line 140. Will all contain C9 + aromatics. The former will contain heavy components from the feed and the latter will usually contain heavy components from isomerization which together produce the heavy components. Fractional distillation of xylene column 116 provides one or more high boiling fractions. As shown, one high boiling fraction containing C9 + aromatics is withdrawn from xylene column 116 via line 142. However, within the broader aspects of the present invention, the xylene column can be operated to provide at least two high boiling fractions, one of which mainly contains C9 and C10 aromatics and some C11 + aromatics, while the other is lower. Concentrations of C9 and C10 and higher concentrations of C11 +. The first fraction preferably contains at least 0.5% by mass polycyclic aromatics, and the second fraction preferably contains at least 50% by mass C11 + aromatics. Since the separation is approximate, a small reboiler duty is required as compared to those typically used in heavy distillation columns to remove almost all C11 + aromatics in the high boiling fraction from xylene column 116. As shown in the embodiment, at least a portion of the C9 + aromatics produced in the xylene column is sent back to the reaction zone via line 142. In general, higher concentrations of C9 + aromatics in the transalkylation feed are advantageous for the production of xylene during the transalkylation reaction. The concentration of C9 + aromatics should not be raised to such an extent that an excessive load on the xylene column or a large amount of purge should be taken. In one embodiment, the concentration of C9 + aromatics in the transalkylation product is less than 10 mass%. In one embodiment, the transalkylation product comprises less than 5 mass% C9 + aromatics, and in other embodiments, the C9 + aromatics comprise 0.01-5 mass% of the transalkylation product.

Optionally, a purge of some of the C9 + aromatics can be used to prevent unwanted rise in concentration. In one embodiment, less than 50 mass% of the C9 + aromatics produced in the xylene column are purged. In another embodiment, less than 20% by mass of the C9 + aromatics in the xylene column is purged and in another embodiment the range of 5-20% by mass is purged. The use of a purge stream may be used such that the transalkylation conditions do not have to be harsh enough to consume sufficient polycyclic aromatics to maintain steady state conditions. The purge may be used in any suitable manner, but it is desirable to recover at least a portion of the C9 + monocyclic aromatic contained therein via distillation as described below.

Line 142 provides fluid communication between the xylene column 116 and the transalkylation reaction zone 144 (reaction zone). If the feedstock in line 102 contains benzene and / or toluene, additional light aromatics (benzene and / or toluene) may be provided via line 146 to reactor 144 in the transalkylation reaction zone if desired. have. Light aromatics can be obtained from any convenient source, including but not limited to lines 106, 112, 132, 138. Although the apparatus of FIG. 1 has been described as having both toluene column 110 and deheptane column 136, for example, to separate toluene from C8 + aromatics, a single column providing both functions may be used. In addition, other columns that perform the same separation can be combined in a similar manner. In one embodiment, the isomerate from the outlet of the isomerization reaction zone is sent to the inlet of the reaction zone product distillation apparatus for treatment with the reaction zone product. This embodiment reduces capital costs by eliminating redundant equipment. In the embodiment not shown, the isomerization reaction zone and the reaction zone share a common hydrogen gas circulation system. For example, the effluent from the isomerization reactor and the transalkylation reactor can be introduced into two separate or one combined product separators, and the gas from one or more separators can be sequentially and / or simultaneously desired. May be combined to recycle to separate reactors.

Since the transalkylation reaction according to the process of the present invention can operate to convert some but not all of the monocyclic aromatics, accumulation of polycyclic aromatics can be achieved. Although steady state conditions can be achieved without the need for purging, the process of the present invention is, in one embodiment, from xylene column 116 to maintain the desired polycyclic aromatic concentration for the transalkylation reactor 144. Some of the heavy oil is removed continuously or intermittently. This purge may be taken as the second high boiling fraction from xylene column 116 as described above, or may be an aliquot portion of the high boiling fraction as shown by line 159.

The purge may be discarded, for example, by use of fuel, or preferably fractionally distilled in the purge column 161. Purge column 161 provides a top fraction enriched in C9 and C10 aromatics, which may be sent via line 163 to line 142 and then to the transalkylation reactor 144. Some or all of this stream may be recovered via line 164 for use as a by-product, for example as a solvent. The bottom fraction in the purge column 161 is rich in C 11+ aromatics and can be discarded and used as a fuel, for example. Since the purge is only a portion of the high boiling fraction of the xylene column 116, for example, often less than 50% by mass, preferably less than 20% by mass, of the high boiling fraction, the purge column is substantially smaller, thus reducing the overall high boiling fraction. Less heat duty may be required than the heavy distillation column used to process.

2 and 3 show various reaction zones including an integrated transalkylation and xylene isomerization reactor system. 3 shows a reactor system 200 having two zones within a single vessel 202. As shown, vessel 202 houses two separate catalyst beds, i.e., an alkyl exchange catalyst bed 204 and a xylene isomerization catalyst bed 206. In the non-illustrated examples, a single catalyst bed is used that is active for transalkylation, olefin removal and xylene isomerization. Feed to reactor vessel 202 is provided by line 208 and one or both of lines 210 and 210A. C8 aromatic and olefin containing feedstock from line 168 is provided via line 208 when lie 102 or an additional sorting step is used, and the C9 + aromatic containing high boiling fraction from xylene column is line 210. Is provided through If additional benzene and / or toluene are fed to the catalyst bed 204, one of the lines shown or a separate dedicated line may be used. If desired, all or part of the unbalanced xylene mixture obtained from line 124 is introduced into catalyst bed 204 via one or a dedicated line shown, and / or all or that mixture from line 124 or Some may be introduced between catalyst beds 204 and 206 via line 210A. Isomerates or reaction zone products may be withdrawn via line 212 in fluid communication with the reaction zone product distillation apparatus, similar to line 148.

The system shown in FIG. 2, indicated generally at 300, involves using two reactors in sequential flow order. The transalkylation reactor 302 therein is equipped with a catalyst bed 304. Feed to the transalkylation reactor 302 is provided via line 306. Although only one line is shown, it is also conceivable to provide the feed components by separate lines to form the reaction mixture in reactor 302. The feed components provided by line 306 are the feedstock from line 102 or line 162 and the C9 + aromatic containing high boiling fraction from the xylene column. If desired, all or part of the unbalanced xylene mixture obtained from line 124 of FIG. 1 may be introduced into alkyl exchange reactor 302 via line 306. Line 308 is adapted to send the transalkylation reaction product from reactor 302 to xylene isomerization reactor 310. As shown, line 314 is provided to introduce all or a portion of the unbalanced xylene mixture obtained from line 124 into xylene isomerization reactor 310. The xylene isomerization reactor contains a catalyst bed 312. The isomerate or reaction zone product may be drawn through line 316 in fluid communication with the reaction zone product distillation apparatus similar to line 148.

Accordingly, the present invention enables and includes a variety of flow techniques that can remove olefins while using an transalkylation reaction zone. In some embodiments, the reaction zone is an alkyl exchange reaction zone, and in other embodiments the reaction zone comprises a transalkylation and isomerization reaction zone, either sequentially or combined. In another embodiment, not shown, the isomerate in line 128 from the outlet of the isomerization reaction zone is sent to the reaction zone product distillation unit for separation with the transalkylation reaction product prior to recycling a portion to the xylene distillation column. Lose. In another embodiment, not shown, the reaction zone product distillation apparatus comprises a single distillation column having at least one outlet to provide a C8 + aromatic stream to the inlet of the xylene column.

In one embodiment, almost all of the feedstock for the xylene production facility defined for the xylene loops is treated in the transalkylation reaction zone to produce xylene and reduce the olefin content. In some embodiments, one or more portions of the feedstock are not sent to the xylene loops and such portions are not necessary but may be treated in an alkyl exchange reaction zone such that excess benzene, toluene or other product may be released from the plant. have. In another embodiment, a stream comprising a C8 aromatic compound is obtained from a separate tailored feedstock and / or obtained by separating the feedstock in a feedstock distillation unit and sent to the xylene column, while the feed defined for the xylene column is The remainder of the feed can first be sent to an alkyl exchange reaction zone to produce xylene and remove olefins.

As shown in FIG. 4, a feedstock comprising olefins and aromatics is sent through conduit 102 to feedstock distillation unit 158 where a stream comprising olefins and C9 + aromatics is produced, which stream exits. And through conduits 162 and 142 to reaction zone 144. Feedstock distillation apparatus 158 also includes a stream comprising at least one of benzene and toluene through outlet and conduit 160 to increase reaction of xylene as desired and to reduce olefin content as desired. Is sent). In addition, a stream comprising C8 aromatics is produced in feedstock distillation apparatus 158 and sent to the xylene column 116 via an outlet and conduit 165. In another embodiment, not shown, the feedstock distillation apparatus comprises a plurality of columns and a stream comprising C8 aromatics is obtained as the top or bottom stream of the column. Thus, reaction zone 144 reduces the amount of olefins in the feed fuel while producing more xylenes per pass because of reduced xylene in the transalkylation feed. The portion of the C8 aromatics rich feedstock sent to the xylene column through conduit 165 may contain olefins. However, the olefin removal capacity required in xylene production equipment is significantly reduced by the removal of olefins from other parts of the feedstock treated in the transalkylation reaction zone. As other aromatic compounds are directed to the alkyl exchange reaction zone for xylene conversion and olefin reduction, it is desirable to treat the xylene rich stream using olefin removal zone 117 as described above.

From the foregoing description, further non-limiting embodiments will be readily understood. The olefin removal apparatus 117 may be a clay treater. In another embodiment, the outlet of the xylene isomerization reaction zone is in fluid communication with the inlet of the reaction zone product distillation apparatus.

In one embodiment, the isomerate distillation apparatus comprises (a) an inlet, a first outlet and a second outlet in fluid communication with the xylene isomerization reaction zone 126, from which the benzene containing stream A light column (130) withdrawing the C7 + aromatic containing stream from the second outlet; And (b) an inlet, a first outlet and a second outlet in fluid communication with the second outlet of the hard column, from which the toluene containing stream is discharged and the C8 + aromatic containing stream is discharged from the second outlet. Deheptane column 136. Optionally, at least one of the first outlet 132 of the hard column and the first outlet 138 of the deheptane column is in fluid communication with the inlet of the reaction zone 144.

In one embodiment, the reaction zone product distillation apparatus comprises: (a) a stripper column 150 having an inlet in fluid communication with the outlet of the reaction zone 144 and an outlet therefrom which discharge the C6 + aromatic containing stream. ; (b) a benzene column comprising an inlet, a first outlet and a second outlet in fluid communication with the outlet of the stripper column, withdrawing a benzene containing stream from the first outlet and a C7 + aromatic containing stream from the second outlet; 104; And (c) an inlet, a first outlet, and a second outlet in fluid communication with the second outlet of the benzene column, from which the first outlet discharges the toluene containing stream and is in fluid communication with the xylene distillation column 116. From the two outlets there is a toluene column 110 which withdraws the C8 + aromatic containing stream. Optionally, at least one of the first outlet 112 of the hard column and the first outlet 106 of the deheptane column is in fluid communication with the inlet of the reaction zone 144.

In one embodiment, the reaction zone 144 further comprises a xylene isomerization reaction zone 126, wherein the inlet of the xylene isomerization reaction zone is in fluid communication with the outlet of the first reactor and the outlet of the xylene isomerization reaction zone Is in fluid communication with the inlet of the reaction zone product distillation apparatus. In another embodiment, reaction zone 144 further includes a xylene isomerization reaction zone 126, which contains an alkyl exchange and olefin conversion catalyst in the first bed and a xylene isomerization catalyst in the second bed. The inlet of the reaction zone is in fluid communication with the bed, the inlet of the xylene isomerization reaction zone is in fluid communication with the second bed, the outlet of the first bed is in fluid communication with the second bed, and the xylene isomerization reaction zone. The outlet of is in fluid communication with the inlet of the second bed and reaction zone product distillation apparatus. In another embodiment, feed stream distillation apparatus 158 produces a stream comprising at least one of benzene and toluene and sends it directly to the inlet of the reaction zone.

Claims

A process for transalkylation and olefin removal from a feed stream comprising an aromatic compound having a Bromine Index of at least 50 and having at least 6 carbon atoms,
An alkyl exchange catalyst comprising an acidic molecular sieve and at least one metal component under alkyl exchange reaction conditions to provide an alkyl exchange reaction product having an increased concentration of xylene and having a bromine index at least 60% lower than the bromine index of the feed stream; Contacting the feed stream
Alkyl exchange and olefin removal method comprising a.

The method of claim 1, wherein the transalkylation conditions include a temperature in the range of 100 ° C. to 425 ° C., a pressure in the range of 100 Pa to 10 MPa (absolute pressure), and a weight space velocity in the range of 0.1 hr ⁻¹ to 30 hr ^−1. an alkyl exchange and olefin removal method comprising hourly space velocity (WHSV).

The method of claim 1, wherein the acidic molecular sieve is BEA, MTW, FAU, MOR, LTL, ITH, ITW, MEL, FER, TON, IWW, MFI, EUO, MTT, HEU, MFS, CHA, ERI, and LTA At least one metal, and the at least one metal is platinum, palladium, rhodium, ruthenium, osmium, iridium, rhenium, tin, germanium, lead, iron, cobalt, nickel, indium, gallium, zinc, uranium, dysprosium, And at least one of thallium.

The process of claim 1 wherein the feed stream comprises at least one of toluene and C9 + aromatics and the bromine index of the feed stream is at least 100.

The alkyl of claim 1, wherein the feed stream comprises C9 + aromatics and has an end boiling point of at least 190 ° C., wherein 20 to 75 mol% of the total C9 + aromatics in the feed stream are converted. Exchange and olefin removal method.

2. The process of claim 1, wherein Ring Loss is less than 2 mol% based on the number of moles of monocyclic aromatics in the feed stream.

The method of claim 1 wherein the metal component is present in an amount such that the rate of activity reduction of the catalyst is less than 200% when measured under evaluation conditions.

2. The process of claim 1 wherein (a) the xylene loop feed comprising xylene and C9 + aromatics is distilled to provide a low boiling fraction comprising xylenes and at least one high boiling fraction comprising C9 + aromatics. step;
(b) separating the low boiling fraction in at least one xylene isomer separation zone to recover the desired xylene isomers and provide a C8 aromatic stream with reduced xylene isomers;
(c) isomerizing the xylene isomer reduced C8 aromatic stream under isomerization reaction conditions to provide an isomerate having a distribution of xylene isomers approximating a balanced state;
(d) separating the recycle stream comprising xylenes from the isomerate by fractional distillation;
(e) sending the recycle stream to step (a);
(f) separating C7 and lower hydrocarbons by fractional distillation from the transalkylation product to provide a C8 containing fraction;
(g) sending at least a portion of the C8 containing fraction to step (a) as at least a portion of a xylene loop feed; And
(h) mixing a feed stream with a portion of at least one high boiling fraction comprising the C9 + aromatic from step (a) to provide a mixture having a bromine index of at least 50 and a final boiling point of at least 210 ° C.
Wherein the mixture of step (h) is contacted with an alkyl exchange catalyst, the mixture further comprises at least one of benzene and toluene, wherein the ratio of the total moles of benzene and toluene to the total moles of C9 + aromatics is 0.01 : Greater than 1 and wherein the ratio of said feed stream to said at least one high boiling fraction comprising said C9 + aromatics in said mixture is from 0.01: 1 to 3: 1.

An apparatus for the transalkylation and olefin removal from a feed stream comprising an aromatic compound having a bromine index of at least 50 and having at least six carbon atoms,
(a) a feed stream distillation apparatus having an inlet and an outlet and withdrawing at the outlet a stream comprising aromatic compounds and olefins having at least six carbon atoms;
(b) a reaction zone comprising at least a first reactor containing an alkyl exchange and olefin conversion catalyst comprising an acidic molecular sieve and at least one metal component and having at least one inlet and at least one outlet;
(c) a conduit in fluid communication with the outlet of said feed stream distillation apparatus and no fluid communication to said olefin removal zone;
(d) a reaction zone product distillation apparatus comprising at least one distillation column and having at least one inlet in fluid communication with the outlet of said reaction zone and an outlet therefrom, wherein said outlet discharges a stream comprising C8 aromatics;
(e) a xylene distillation column having an inlet, a first outlet and a second outlet in fluid communication with said reaction zone product distillation apparatus, withdrawing from said first outlet a stream comprising at least two xylene isomers;
(f) an inlet, a first outlet, and a second outlet in fluid communication with the first outlet of the xylene distillation column, wherein the predetermined xylene isomer is discharged from the first outlet and the predetermined xylene isomer is reduced from the second outlet A selective xylene isomer separation device for venting an unbalanced xylene isomer stream;
(g) a xylene isomerization zone containing a xylene isomerization catalyst and having at least one inlet and at least one outlet, the inlet being in fluid communication with a second outlet of the optional xylene isomer separation apparatus; And
(i) a purge column having an inlet in fluid communication with a second outlet of the xylene distillation column, a first outlet in fluid communication with the inlet of the reaction zone, and a second outlet for discharging the purge stream
Alkyl exchange and olefin removal apparatus comprising a.

10. The system of claim 9, comprising (a) at least one distillation column, at least one inlet in fluid communication with the outlet of the xylene isomerization zone, and an outlet in fluid communication with the inlet of the xylene distillation column, An isomerate distillation apparatus for withdrawing a stream comprising C8 aromatics from this outlet;
(b) an olefin removal zone having an inlet in fluid communication with the first outlet of the xylene distillation column and an outlet in fluid communication with the inlet of the optional xylene isomer separation device;
(c) a second outlet of said feed stream distillation apparatus; And
(d) a second conduit providing fluid communication between the second outlet of the feed stream distillation unit and the inlet of the xylene distillation column and not providing fluid communication with the olefin removal zone
Alkyl exchange and olefin removal device further comprising.