WO2014044020A1 - 一种联产环己醇和链烷醇的方法和装置 - Google Patents

一种联产环己醇和链烷醇的方法和装置 Download PDF

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WO2014044020A1
WO2014044020A1 PCT/CN2013/001100 CN2013001100W WO2014044020A1 WO 2014044020 A1 WO2014044020 A1 WO 2014044020A1 CN 2013001100 W CN2013001100 W CN 2013001100W WO 2014044020 A1 WO2014044020 A1 WO 2014044020A1
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reaction
cyclohexene
hydrogenation
catalyst
reactor
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PCT/CN2013/001100
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English (en)
French (fr)
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宗保宁
马东强
温朗友
孙斌
杨克勇
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中国石油化工股份有限公司
中国石油化工股份有限公司石油化工科学研究院
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Priority claimed from CN201210347463.6A external-priority patent/CN103657658B/zh
Priority claimed from CN201210559981.4A external-priority patent/CN103880598B/zh
Priority claimed from CN201210560215.XA external-priority patent/CN103664530A/zh
Priority claimed from CN201210560237.6A external-priority patent/CN103880599B/zh
Priority claimed from CN201210560665.9A external-priority patent/CN103664531B/zh
Priority claimed from CN201210559160.0A external-priority patent/CN103880597A/zh
Priority claimed from CN201210559915.7A external-priority patent/CN103664529B/zh
Priority claimed from CN201310001078.0A external-priority patent/CN103910602A/zh
Priority claimed from CN201310001152.9A external-priority patent/CN103910603B/zh
Priority to US14/429,189 priority Critical patent/US9561991B2/en
Priority to KR1020157010078A priority patent/KR102008352B1/ko
Application filed by 中国石油化工股份有限公司, 中国石油化工股份有限公司石油化工科学研究院 filed Critical 中国石油化工股份有限公司
Publication of WO2014044020A1 publication Critical patent/WO2014044020A1/zh

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    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C29/00—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring
    • C07C29/132—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring by reduction of an oxygen containing functional group
    • C07C29/136—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring by reduction of an oxygen containing functional group of >C=O containing groups, e.g. —COOH
    • C07C29/147—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring by reduction of an oxygen containing functional group of >C=O containing groups, e.g. —COOH of carboxylic acids or derivatives thereof
    • C07C29/149—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring by reduction of an oxygen containing functional group of >C=O containing groups, e.g. —COOH of carboxylic acids or derivatives thereof with hydrogen or hydrogen-containing gases
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    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
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    • B01J23/00—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group B01J21/00
    • B01J23/70—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group B01J21/00 of the iron group metals or copper
    • B01J23/76—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group B01J21/00 of the iron group metals or copper combined with metals, oxides or hydroxides provided for in groups B01J23/02 - B01J23/36
    • B01J23/84—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group B01J21/00 of the iron group metals or copper combined with metals, oxides or hydroxides provided for in groups B01J23/02 - B01J23/36 with arsenic, antimony, bismuth, vanadium, niobium, tantalum, polonium, chromium, molybdenum, tungsten, manganese, technetium or rhenium
    • B01J23/85—Chromium, molybdenum or tungsten
    • B01J23/86—Chromium
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    • B01J37/00—Processes, in general, for preparing catalysts; Processes, in general, for activation of catalysts
    • B01J37/02—Impregnation, coating or precipitation
    • B01J37/0201—Impregnation
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
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    • B01J37/00—Processes, in general, for preparing catalysts; Processes, in general, for activation of catalysts
    • B01J37/02—Impregnation, coating or precipitation
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    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
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    • C07C45/002—Preparation of compounds having >C = O groups bound only to carbon or hydrogen atoms; Preparation of chelates of such compounds by dehydrogenation
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C67/00—Preparation of carboxylic acid esters
    • C07C67/04—Preparation of carboxylic acid esters by reacting carboxylic acids or symmetrical anhydrides onto unsaturated carbon-to-carbon bonds
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07D—HETEROCYCLIC COMPOUNDS
    • C07D223/00—Heterocyclic compounds containing seven-membered rings having one nitrogen atom as the only ring hetero atom
    • C07D223/02—Heterocyclic compounds containing seven-membered rings having one nitrogen atom as the only ring hetero atom not condensed with other rings
    • C07D223/06—Heterocyclic compounds containing seven-membered rings having one nitrogen atom as the only ring hetero atom not condensed with other rings with hetero atoms or with carbon atoms having three bonds to hetero atoms with at the most one bond to halogen, e.g. ester or nitrile radicals, directly attached to ring carbon atoms
    • C07D223/08—Oxygen atoms
    • C07D223/10—Oxygen atoms attached in position 2
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C2601/00—Systems containing only non-condensed rings
    • C07C2601/12—Systems containing only non-condensed rings with a six-membered ring
    • C07C2601/14—The ring being saturated
    • Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
    • Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
    • Y02P—CLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
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    • Y02P20/10—Process efficiency
    • Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
    • Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
    • Y02P—CLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
    • Y02P20/00—Technologies relating to chemical industry
    • Y02P20/50—Improvements relating to the production of bulk chemicals
    • Y02P20/52—Improvements relating to the production of bulk chemicals using catalysts, e.g. selective catalysts

Definitions

  • the present invention relates to a process for the manufacture of cyclohexanol. More specifically, it relates to a method of co-producing cyclohexanol and an alkanol. The present invention also relates to a process for further producing cyclohexanone or caprolactam starting from the process for producing cyclohexanol. The invention also relates to a device for the co-production of cyclohexanol and an alkanol. Background technique
  • Cyclohexanol and alkanols such as ethanol are important organic chemical raw materials and organic solvents. Cyclohexanol is mainly used in the dehydrogenation process to produce cyclohexanone, while cyclohexanone is the main intermediate for the further production of nylon 6 and nylon 66. Since the advent of nylon, the world's major chemical companies have been working to develop industrial sources of cyclohexanol (ketone). In the 1980s, Asahi Kasei Corporation of Japan developed a method for producing cyclohexanol by using cyclohexene hydrate (cyclohexene hydration).
  • the method uses a complex reaction system including an aqueous phase, an oil phase and a solid catalyst phase, It is necessary to rely on strong agitation to form an emulsion system in which water droplets and oil droplets are well dispersed, so that cyclohexene can be adsorbed on the surface of the catalyst, and the catalyst and the oil phase need to be well separated in the sedimentation zone, and the operation is complicated. The losses are also serious. In countries with abundant products, sputum is still the main method of producing ethanol. The main disadvantage of the fermentation method is that the pollution is more serious.
  • the fermentation method still has the problem of “competing grain with the mouth.” It is not applicable to countries with large population and small cultivated land area.
  • the reaction conditions of ethylene direct water law are harsh. It needs to be carried out under high temperature and high pressure.
  • the price of ethylene is greatly affected by fluctuations in international oil prices. For countries with insufficient petroleum resources, the use of ethylene hydration to produce ethanol will face certain raw material cost pressures.
  • the inventors have conducted painstaking research on the basis of the prior art and found that by at least the two steps of the cyclohexene esterification step and the cyclohexyl ester hydrogenation step, it is possible to, for example, a simpler manufacturing process and more.
  • the present invention has been accomplished by co-production of cyclohexanol and an alkanol such as ethanol at an inexpensive production cost and solving the aforementioned problems of the prior art.
  • the present invention relates to the following aspects:
  • a method for co-producing cyclohexanol and an alkanol comprising the steps of: (1) subjecting a cyclohexene source to at least one carboxylic acid to form an ester in the presence of an addition esterification catalyst a reaction to form an addition esterification product containing cyclohexyl carboxylate, wherein the at least one carboxylic acid is represented by the formula R-COOH, and the group R is hydrogen or C 1-23 linear or a branched alkyl group, preferably a d 6 straight or branched alkyl group, more preferably a d.
  • the cyclohexene source contains 20 mol% or more and 35 mol% or more. 20 to 80 mol%, 20 to 60 mol%, 40 to 80 mol%, 80 to 95 mol% or more of cyclohexene; and
  • step (B) further separating the hydrogenation product obtained in the step (A) to obtain a step of using cyclohexene, a mixture of cyclohexene and benzene or a mixture of cyclohexene and cyclohexane as the source of cyclohexene;
  • (C) a step of subjecting cyclohexane to a partial dehydrogenation reaction in the presence of a partial dehydrogenation catalyst to obtain a partial dehydrogenation product containing cyclohexene as the cyclohexene source;
  • step (D) further separating the partially dehydrogenated product obtained in the step (C) to obtain a cyclohexene or a mixture of cyclohexene and cyclohexane as the cyclohexene source.
  • step (I) recovering benzene and / or hydrogen separated from any of the steps of the process for co-producing cyclohexanol and alkanol, and recycling the benzene and / or hydrogen to the step (A);
  • step (III) recovering cyclohexane separated from any of the steps of the method for co-producing cyclohexanol and an alkanol, and dehydrogenating the cyclohexane in the presence of a dehydrogenation catalyst to obtain benzene and hydrogen, The benzene and/or hydrogen are recycled to the step (A).
  • the addition esterification catalyst is selected from one or more of the solid acid catalysts, preferably selected from the group consisting of an acid strength function (Hammett function) H0 of -8 or less (preferably - One or more of 12 or less, more preferably 13 or less) of the solid acid catalyst, more preferably one or more selected from the group consisting of a strong acid type ion exchange resin (preferably selected from the group consisting of sulfonic acid type ion exchange resins, more preferably One or more selected from the group consisting of a macroporous sulfonic acid type ion exchange resin and a halogen modified sulfonic acid type ion exchange resin), a heteropoly acid (such as a heteropoly acid selected from a keggin structure, a heteropoly acid of a Dawson structure, One or more of the heteropolyacid of the Anderson structure, the heteropoly acid of the Silverton structure, the acid salt of the aforementioned heteropol
  • an acid strength function H0 of -8 or
  • zeolite molecular sieves preferably selected from ⁇ zeolite molecular sieves, fluorine and/or phosphorus modified ⁇ zeolite molecular sieves, yttrium zeolite molecules
  • zeolite molecular sieves preferably selected from ⁇ zeolite molecular sieves, fluorine and/or phosphorus modified ⁇ zeolite molecular sieves, yttrium zeolite molecules
  • zeolite molecular sieves preferably selected from ⁇ zeolite molecular sieves, fluorine and/or phosphorus modified ⁇ zeolite molecular sieves, yttrium zeolite molecules
  • the hydrogenation catalyst is selected from a copper-based catalyst (more preferably one or more selected from the group consisting of a zinc-based copper-based catalyst and a chromium-containing copper-based catalyst), ruthenium.
  • a catalyst preferably selected from one or more of Ru/Al 2 0 3 and Ru-Sn/Al 2 0 3
  • a noble metal catalyst preferably selected from the group consisting of Pt/Al 2 O 3 , Pd-Pt/Al 2
  • the copper-based catalyst comprises the following components (preferably consisting of: copper oxide); (b) zinc oxide; (c) selected from the group consisting of aluminum and gallium An oxide of one or more metals selected from the group consisting of tin, titanium, zirconium, chromium, molybdenum, tungsten, manganese, lanthanum, lanthanide and lanthanide metals, preferably selected from the group consisting of aluminum, gallium, tin, titanium, zirconium, chromium And an oxide of one or more metals of molybdenum, tungsten, manganese, lanthanum, cerium and lanthanum; and (d) one or more selected from the group consisting of alkali metal hydroxides and alkaline earth metal hydroxides, preferably One or more selected from the group consisting of potassium hydroxide, sodium hydroxide and barium hydroxide, wherein in parts by mass, component (a): component (b):
  • the composite metal oxide comprises the following components (preferably composed of the following components): (a) copper oxide; (b) zinc oxide; c) an oxide of one or more metals selected from the group consisting of aluminum, gallium, tin, titanium, zirconium, chromium, molybdenum, tungsten, manganese, lanthanum, lanthanide metals and lanthanide metals, preferably selected from the group consisting of aluminum, gallium, An oxide of one or more metals of tin, titanium, zirconium, chromium, molybdenum, tungsten, manganese, lanthanum, cerium and lanthanum, in parts by mass, component (a): component (b): Component (c) is 5 to 60: 10 to 50: 5 to 60, preferably 10 to 50: 15 to 45: 15 to 55, more preferably 30 to 45: 20 to 35: 20 to 50;
  • the molar ratio of the at least one carboxylic acid to the cyclohexene source in terms of cyclohexene is from 0.2 to 20:1, preferably from 1.2 to 4:1, more Preferably, 1.2 to 3:1, and the step (1) is performed in the following manners (1), (2), or any combination thereof, and preferably in combination (2) or in combination of the modes (1) and (2), More preferably, the mode (1) is followed by the combination of the modes (2):
  • the reactor is a tank reactor, a fixed bed reactor, a fluidized bed reactor, an ebullated bed reactor or any combination thereof in parallel, preferably a tubular fixed bed reactor, more preferably a tubular tube tubular reaction
  • the reaction temperature is 50 to 200 ° C, preferably 60 to 120 ° C
  • the reaction pressure is from atmospheric pressure to 10 MPa, preferably from atmospheric pressure to 1 MPa
  • the addition esterification reaction is carried out in a continuous manner, the liquid feed space velocity 5 ⁇ 20 ⁇ , ⁇ ⁇ 0.5 ⁇ 5 ⁇ , More preferably 1 ⁇ 5h _1 , when the addition esterification reaction is carried out in a batch mode, the reaction time is 0.1 ⁇ 10h, preferably 0.2 ⁇ 2h;
  • the reactor is a reaction fine column, preferably selected from a plate column, a packed column or any parallel combination thereof, the number of theoretical plates is 10 to 150, preferably 30 to 100, and the operating pressure is -0.0099 MPa to 5 MPa.
  • the atmospheric pressure to the IMPa the temperature of the addition esterification catalyst bed loading zone is 40 to 200 ° C, preferably 50 to 20 CTC, more preferably 60 to 120 ° C, and the reflux ratio is 0.1:1 to total reflux, preferably 0.1 to 100:1, more preferably 0.5 to 10:1, 5 to 30 plates (preferably 8 to 20 plates) are selected between the 1/3 to 2/3 positions of the number of theoretical plates to arrange the addition ester catalyst, and an addition of an esterification catalyst with respect to the total volume of packing, the liquid feed space velocity of from 0.1 to 2011-1, 2011-1 to preferably 0.2, more preferably 0.5 to 511
  • the reactor is a tank reactor, a fixed bed reactor, an ebullated bed reactor, a fluidized bed reactor or Any combination in parallel, preferably a tubular fixed bed reactor, more preferably a tube-and-tube tubular reactor
  • the reaction temperature is 150 to 400 ° C, preferably 200 to 300 ° C
  • the reaction pressure is from atmospheric pressure to 20 MPa, preferably atmospheric pressure to lOMPa, more preferably 4 to 10 MPa
  • the molar ratio of hydrogen to the addition esterified product based on cyclohexyl carboxylate is from 1 to 1000:1, preferably from 5 to 100:1, the hydrogenation reaction is in a continuous manner
  • the liquid feed space velocity is from 0.1 to 20 H" 1 , preferably from 0.2 to 2 h
  • the reaction time is from 0.2 to 20 h, preferably from 0.5 to 5 h
  • the carboxylic acid hydrogenation catalyst is composed of a main active component of 0.1 to 30% by weight, an auxiliary agent of 0.1 to 25% by weight, and a balance of a carrier, wherein the main active component One or more selected from the group consisting of platinum, palladium, rhodium, tungsten, molybdenum and cobalt, the auxiliary agent being selected from the group consisting of tin, chromium, aluminum, zinc, calcium, magnesium, nickel, titanium, zirconium, hafnium, tantalum, niobium And one or more of the gold, the carrier is selected from one or more of the group consisting of silica, alumina, titania, zirconia, activated carbon, graphite, carbon nanotubes, calcium silicate, zeolite, and aluminum silicate.
  • the main active component One or more selected from the group consisting of platinum, palladium, rhodium, tungsten, molybdenum and cobalt
  • the auxiliary agent being selected
  • the carboxylic acid hydrogenation reaction is carried out under the following reaction conditions: the reactor is a tank reactor, a fixed bed reactor, an ebullated bed reactor, a fluidized bed reactor or any combination thereof in parallel, preferably a tubular fixed bed reactor More preferably, the shell-and-tube tubular reactor has a reaction temperature of 100 to 400 ° C, preferably 180 to 300 ° C, a reaction pressure of 0.1 to 30 MPa, preferably 2 to 1 OMPa, and hydrogen with the free carboxylic acid.
  • the molar ratio is from 1 to 500:1, preferably from 5 to 50:1.
  • the liquid feed space velocity is from 0.1 to 5, preferably from 0.2 to 2 h, and the carboxylic acid is added.
  • the reaction time is from 0.5 to 20 h, preferably from 1 to 5 h.
  • a method of producing cyclohexanone characterized by comprising:
  • Cyclohexanone is produced using the cyclohexanol.
  • a method of producing caprolactam characterized by comprising:
  • Caprolactam is produced using the cyclohexanone.
  • a device for co-producing cyclohexanol and an alkanol characterized by comprising a hydrogenation reaction Unit A, an optional hydrogenation product separation unit A, an addition esterification reaction unit, an optional addition esterification product separation unit, a hydrogenation reaction unit B, and a hydrogenation product separation unit B, wherein In the hydrogen reaction unit A, partial hydrogenation reaction of benzene with hydrogen in the presence of a partial hydrogenation catalyst to obtain a hydrogenated product containing cyclohexene;
  • the hydrogenation product from the hydrogenation reaction unit A is separated to obtain cyclohexene, a mixture of cyclohexene and benzene, or cyclohexene and cyclohexane. mixture;
  • the hydrogenation product from the hydrogenation reaction unit A and/or a mixture of cyclohexene, cyclohexene and benzene from the hydrogenation product separation unit A Or an addition and esterification reaction of a mixture of cyclohexene and cyclohexane with a carboxylic acid in the presence of an addition esterification catalyst to form an addition esterification product containing cyclohexyl carboxylate;
  • the addition esterification product from the addition esterification reaction unit is separated to obtain a cyclohexyl carboxylate or a cyclohexyl carboxylate and a carboxylic acid.
  • the addition esterification product from the addition esterification reaction unit and/or the cyclohexyl carboxylate or carboxylic acid from the addition esterification product separation unit a mixture of cyclohexyl ester and a carboxylic acid is hydrogenated with hydrogen in the presence of a hydrogenation catalyst to form a hydrogenation product comprising cyclohexanol and an alkanol;
  • the hydrogenation product from the hydrogenation reaction unit B is separated to obtain cyclohexanol and an alkanol.
  • the one-pass selectivity and the single-pass conversion of the cyclohexene esterification step and the cyclohexyl ester hydrogenation step are both high, and there are almost no by-products, especially cyclohexyl ester.
  • the single-pass selectivity and single-pass conversion of the hydrogenation step are close to 100%, so the overall single-pass selectivity and single-pass conversion of the method are very high, such as much higher than the cyclohexanol production method developed by Asahi Kasei, compared with the prior art. It has the characteristics of low manufacturing cost and high atomic economy of cyclohexanol.
  • the addition and esterification reaction of cyclohexene is carried out by using a reactive distillation column alone or in combination, and the single-pass conversion rate of cyclohexene is significantly improved (up to 99%).
  • the above greatly simplifies the separation operation of the esterified product, and has the characteristics of significantly lower manufacturing cost and high atomic economy of cyclohexanol compared with the prior art.
  • the cyclohexene esterification step has a simple reaction history, less side reactions, and less influence by impurities, so that the purity requirement of the cyclohexene raw material is higher.
  • the crude product the lowest content of cyclohexene, for example, 20 mol%
  • the present invention has found for the first time in the art that the process for co-producing cyclohexanol and alkanol can even directly use the product stream of benzene partial hydrogenation as a starting material without necessarily requiring complicated or expensive pre-purification or separation. Compared with the prior art, the production cost is significantly reduced.
  • cyclohexane which is partially hydrogenated by partial hydrogenation is subjected to dehydrogenation reaction with very high conversion rate and selection. Re-conversion to benzene, the carbon utilization of the benzene partial hydrogenation, cyclohexene esterification, cyclohexyl ester hydrogenation process is close to 100%.
  • an alkanol having a higher economic value than a carboxylic acid raw material, particularly ethanol, is co-produced while producing cyclohexanol.
  • the reaction system is simple, the operation is simple and convenient, and the manufacturing cost of the cyclohexanol and the maintenance cost of the manufacturing apparatus can be remarkably reduced as compared with the prior art.
  • Fig. 1 is a general flow chart schematically showing a method of co-producing cyclohexanol and an alkanol of the present invention.
  • 2 through 16 are flow diagrams schematically showing various embodiments of the process for the co-production of cyclohexanol and alkanol of the present invention. detailed description
  • hydrocarbons or hydrocarbon derivative groups of more than 3 carbon atoms such as propyl, propoxy, butyl) , butane, butene, butenyl, hexane, etc.
  • propyl is generally understood to be n-propyl
  • butyl is generally understood to be n-butyl unless otherwise specified.
  • selectivity does not refer to single pass selectivity
  • conversion rate refers to single pass conversion
  • cyclohexene source refers to a ring that can be in the present invention.
  • Any reaction raw material as a cyclohexene source (ie, providing cyclohexene) in the hexene esterification step including cyclohexene industrial pure product (cyclohexene content such as 95 mol% or more), cyclohexene industrial crude product (ring
  • the hexene content is, for example, at least 80 mol%, up to 95 mol% or an industrial mixed product containing cyclohexene (such as a cyclohexene content of at least 20 mol% and a maximum of 80 mol%).
  • the reaction course is simple, the side reaction is less, and the esterification reaction is less affected by the impurities, so that the purity requirement of the cyclohexene source is lower.
  • the term "cyclohexene source” also includes industrial waste or industrial by-products containing cyclohexene, such as cyclohexene-containing waste gas (such as cyclohexene from the prior art). And cyclohexene-containing tail gas (for example, as a by-product of the chemical synthesis industry), etc.
  • the present invention does not have a significant effect (such as reducing the single pass conversion of cyclohexene by no more than 5%) and can be used directly without prior purification or impurity removal.
  • the cyclohexene esterification step of the invention is chemically inert, and examples thereof include nitrogen, a rare gas, carbon dioxide, benzene, hydrogen or cyclohexane, etc., which are referred to herein as inert diluents.
  • Technicians can confirm whether an industrial waste or industrial by-product contains a simple test (such as by measuring the degree of reduction of cyclohexene single-pass conversion). Either excessively contains impurities which have a significant influence on the cyclohexene esterification step of the present invention, thereby confirming whether or not it can be directly applied to the co-production method of the present invention. Further, those skilled in the art can also be known by conventionally as needed.
  • a process for co-producing cyclohexanol and an alkanol comprising at least (1) a cyclohexene esterification step and (2) a cyclohexyl ester hydrogenation step.
  • an esterification reaction of a cyclohexene source with at least one carboxylic acid in the presence of an addition esterification catalyst is carried out to form a cyclohexyl carboxylate-containing compound. Addition of the esterification product.
  • the "addition esterification reaction” refers to a reaction in which an olefin bond of a carboxylic acid to cyclohexene is added to form a cyclohexyl carboxylate.
  • said at least one carboxylic acid may be represented by the formula R-COOH, wherein R is hydrogen or a group d_ 23 straight or branched chain alkyl group.
  • the group R is preferably a linear or branched alkyl group, preferably a d. 6 straight or branched alkyl group, more preferably a C 1-3 linear or branched alkyl group, most preferably a methyl group.
  • the carboxylic acid may be used singly or in combination of two or more.
  • one or more of citric acid, acetic acid, propionic acid and n-butyric acid are preferably used, and acetic acid is most preferably used as the carboxylic acid.
  • the content of cyclohexanone in the cyclohexene source is generally selected from the range of 20 mol% or more, 35 mol% or more, 20 to 80 mol%, 20 to 60 mol%, 40 to 80 mol%, 80 to 95 mol% or 95 mol. %the above.
  • the aforementioned inert diluent may also be included in the cyclohexene source.
  • the inert diluent benzene, cyclohexane or a combination thereof in any ratio is preferred.
  • a cyclohexene source a cyclohexene industrial pure product (having a cyclohexene content of, for example, 95 mol% or more) or a cyclohexene industrial crude product (a cyclohexene content such as a minimum of 80 mol%, a maximum of 95 mol%) or an industrial mixed product containing cyclohexene (cyclohexene content such as a minimum of 20 mol%, up to 80 mol%), especially the industrial mixed product containing cyclohexene.
  • These cyclohexene sources are conveniently produced in the form of industrial products or are commercially available.
  • the industrial mixed product containing cyclohexene for example, (1) hydrogenation of cyclohexene obtained by partial hydrogenation of benzene with hydrogen in the presence of a partial hydrogenation catalyst can be exemplified.
  • a product also referred to as a benzene partial hydrogenated product stream
  • (2) a partial dehydrogenation product containing cyclohexene obtained by partial dehydrogenation of cyclohexane in the presence of a partial dehydrogenation catalyst also referred to as a product stream of partial dehydrogenation of cyclohexane
  • step (A) a step of partially hydrogenating benzene with hydrogen in the presence of a partial hydrogenation catalyst (partial hydrogenation of benzene) to obtain a hydrogenated product containing cyclohexene as a source of the cyclohexene;
  • step (B) further separating the hydrogenation product obtained in the step (A) to obtain cyclohexene, a mixture of cyclohexene and benzene or a mixture of cyclohexene and cyclohexane as the cyclohexene source Step
  • (C) a step of subjecting cyclohexane to a partial dehydrogenation reaction (cyclohexane partial dehydrogenation) in the presence of a partial dehydrogenation catalyst to obtain a partial dehydrogenation product containing cyclohexene as the cyclohexene source ;
  • step (D) further separating the partially dehydrogenated product obtained in the step (C) to obtain a mixture of cyclohexene or a mixture of cyclohexane and cyclohexane as the source of the cyclohexene.
  • the benzene partial hydrogenation method (step (A)) and the cyclohexane partial dehydrogenation method (step (C)) and subsequent separation methods (step (B) or step (D) And the like are not particularly limited, and those known in the art can be directly used.
  • the benzene partial hydrogenation method is carried out by a liquid phase method in a manner conventionally known in the art.
  • a partial hydrogenation catalyst used in the step (A) a ruthenium-based catalyst is preferable, and a ruthenium-based catalyst containing cobalt and/or zinc is more preferable.
  • Such catalysts can be produced by methods such as coprecipitation or impregnation of the same support in a manner known in the art.
  • the cyclohexene-containing hydrogenation product (also referred to as a partial hydrogenated product stream) is generally a mixture of cyclohexane, cyclohexene and benzene, which can be directly used as the present invention. Use as a cyclohexene source.
  • cyclohexene (also referred to as pure cyclohexene) can be separated from the benzene partially hydrogenated product stream by any method known in the art.
  • a mixture of cyclohexene and benzene or a mixture of cyclohexene and cyclohexane is used directly as the cyclohexene source of the present invention.
  • the separation method for example, an extractive distillation method or an azeotropic distillation method may be mentioned, and an extractive distillation method is preferred.
  • N-methyl-2-pyrrolidone, N,N-dimercaptoacetamide, adiponitrile, dinonyl malonate, dinonyl succinate, ethylene glycol or sulfolane can be used.
  • Etc. as an extractant.
  • a product stream in which the benzene is partially hydrogenated is fed from the middle to the extractive distillation column, and N, N-dimethylacetamide is introduced from the upper portion of the column.
  • the top of the column is obtained as a cyclohexane stream (ie pure cyclohexane) or cyclohexane and ring
  • a mixture of hexene i.e., a mixture of cyclohexane and cyclohexene).
  • a solution elbow containing cyclohexene, benzene and N, N-dimethylacetamide is obtained at the bottom of the column, and the solution is sent to a rectification column for further separation, and the ring can be obtained from the top of the column.
  • a mixture of a mixture of alkene and benzene ie, a mixture of cyclohexene and benzene
  • a solution containing benzene and N,N-dimethylacetamide at the bottom of the column the solution is sent to a rectification column for further separation, At the top of the column, a benzene stream (i.e., pure benzene) is obtained, and N, N-dimethylacetamide is obtained from the bottom of the distillation column.
  • the mixture of cyclohexane and cyclohexene is fed to a rectification column, N, N-dimethylacetamide is introduced into the upper part of the column, and a cyclohexane stream (ie, pure cyclohexane) is obtained at the top of the column, and the bottom of the column is obtained.
  • a mixture of cyclohexene and N,N-dimethylacetamide the bottoms stream is sent to the next-stage rectification column to separate cyclohexene, and the top of the column is obtained as a cyclohexene stream (ie, pure cyclohexene).
  • the bottom of the column gives a stream of N,N-dimercaptoacetamide.
  • the mixture of cyclohexene and benzene is fed to a rectification column, N, N-dimethylacetamide is introduced at the top of the column, and the cyclohexene stream (ie, pure cyclohexene) is obtained at the top of the column, and benzene is obtained at the bottom of the column.
  • the bottoms stream is sent to the next-stage rectification column to separate benzene, the top of the column is obtained to obtain a benzene stream (ie, pure benzene), and the bottom of the column is N, N-dimethyl Acetamide logistics.
  • the benzene stream can be recycled as part of the reaction feed to step (A), and the cyclohexane stream is withdrawn as a by-product or recycled as part of the step (C) reaction feed.
  • the benzene partially hydrogenated product stream, the cyclohexene and benzene mixture or the mixture of cyclohexene and cyclohexane
  • the content of the cyclohexene is generally at least 20 mol%, 35 mol% or 40 mol%, up to 60 mol% or 80 mol%, and the content of the cyclohexene in the pure cyclohexene is generally 95 mol% or more.
  • the cyclohexane partial dehydrogenation is carried out in a manner conventionally known in the art such as cyclohexane oxidative dehydrogenation.
  • cyclohexene also known as cyclohexene pure product
  • cyclohexene can be separated from the partially dehydrogenated product stream of cyclohexene by any method known in the art. Or a mixture of cyclohexene and cyclohexane and used directly as a source of cyclohexene of the present invention.
  • an acid catalyst such as a liquid acid catalyst may be mentioned, specifically, an inorganic acid such as phosphoric acid or sulfuric acid, or an organic acid such as methylbenzenesulfonic acid or aminosulfonic acid, or a solid. Acid catalyst.
  • solid acid is preferred from the viewpoints of reducing equipment corrosion, improving separation of catalyst and esterification product, and the like.
  • the chemical agent in particular, an acid strength function (Hammett function) H0 is a solid acid catalyst of -8 or less (preferably -12 or less, more preferably -13 or less).
  • a strong acid type ion exchange resin, a heteropoly acid or a zeolite molecular sieve is more preferable. These solid acid catalysts may be used alone or in combination of two or more.
  • the strong acid type ion exchange resin a sulfonic acid type ion exchange resin is preferable, and a macroporous sulfonic acid type ion exchange resin (macroporous sulfonic acid type polystyrene-divinyl benzene resin) or [3 ⁇ 4) is more preferable.
  • a modified sulfonic acid type ion exchange resin are readily available from the market and can be obtained by the methods described in the classical literature.
  • the method for producing a macroporous sulfonic acid type polystyrene-divinylbenzene resin is usually a method in which a mixture of styrene and divinylbenzene is dropped into a water phase containing a dispersing agent, an initiator, and a porogen under high-speed stirring.
  • the suspension copolymerization is carried out in the system, and the obtained polymer beads (white spheres) are separated from the system, and the porogen is removed by using a solvent, and then dichloroethane is used as a solvent and concentrated sulfuric acid is used as a sulfonating agent.
  • the sulfonation reaction is finally carried out by filtration, washing, and the like, and finally the product is obtained.
  • halogen-modified sulfonic acid type ion exchange resin can be obtained by at least two routes, one of which is to introduce a sulphur atom on a benzene ring of a sulfonated styrene resin skeleton, for example, a chlorine atom, due to strong electron absorption of a halogen element.
  • the function not only stabilizes the benzene ring, but also increases the acidity of the sulfonic acid group on the benzene ring, so that the acid strength function of the resin catalyst (Hammett function) H0 ⁇ -8, and can be used at a temperature of 150 ° C or more for a long time.
  • resins are readily available on the market, such as Amberlyst 45 resin produced by ROHM & HASS in foreign countries, D008 resin produced by the domestic chemical plant in Hebei Wingzhong, etc.; another way is to replace all hydrogen on the resin skeleton with fluorine. Due to the strong electron absorption of fluorine, it has superior acidity and super high thermal stability.
  • the acid strength function (Hammett function) H0 can be less than -12, and the heat resistant temperature reaches 250 °C or higher.
  • a typical example of a strongly acidic resin is Nafion resin manufactured by DuPont. These strong acid type ion exchange resins may be used singly or in combination of two or more.
  • the heteropoly acid (the acid strength function H0 is generally less than -13.15), for example, a heteropoly acid of a keggin structure, a heteropoly acid of a Dawson structure, a heteropoly acid of an Anderson structure, and a Silverton structure may be mentioned.
  • keggin structure of heteropolyacids/carriers and keggin The acid salt/carrier of the heteropolyacid of the structure.
  • these heteropolyacids can be used at a temperature of 300 ° C or more for a long period of time, and their BET specific surface area is generally 100 m 2 /g or more.
  • These heteropolyacids may be used alone or in combination of two or more.
  • Load ## obtained by loading ## on the carrier.
  • the carrier to be used at this time for example, silica, activated carbon or a combination thereof can be mentioned.
  • heteropoly acid specifically, for example, dodecaphosphoric acid or dodecyl tungstic acid/carrier, dodecyl tungstic acid or dodecyl tungstic acid/carrier, dodecaphosphomolybdic acid or Dodecaphosphonic acid/carrier, dodecaphosphomovanadic acid or dodecaphosphoric vanadic acid/carrier, the acid salt of the aforementioned heteropolyacid and the acid salt/carrier of the aforementioned heteropoly acid, wherein acid phosphotungsten is preferred Acid strontium salt (CS2.5H0.5P12WO40) and acid phosphotungstate strontium salt/carrier.
  • These heteropolyacids may be used singly or in combination of two or more.
  • zeolite molecular sieve for example, an ⁇ zeolite molecular sieve, a fluorine and/or phosphorus modified ⁇ zeolite molecular sieve, a cerium zeolite molecular sieve, a fluorine and/or phosphorus modified cerium zeolite molecular sieve, and a HZSM-5 zeolite may be mentioned.
  • Molecular sieves or fluorine and/or phosphorus modified HZSM-5 zeolite molecular sieves preferably fluorine and/or phosphorus modified A ⁇ zeolite molecular sieves, fluorine and/or phosphorus modified zeolite zeolite molecular sieves or fluorine and/or phosphorus modified HZSM-5 zeolite molecular sieve.
  • These zeolite molecular sieves may be used singly or in combination of two or more.
  • the molar ratio of the at least one carboxylic acid to the cyclohexene source in terms of cyclohexene is generally from 0.2 to 20:1, preferably from 1.2 to 4:1, more preferably 1.2 to 3:1, but sometimes it is not limited to this.
  • the following modes (1) and (2) may be mentioned, wherein from the viewpoint of remarkably improving the single-pass conversion ratio of cyclohexene, it is preferred.
  • the addition esterification reaction is carried out in one or more addition esterification reactors.
  • the addition esterification reactor for example, a tank reactor, a fixed bed reactor, a fluidized bed reactor, an ebullated bed reactor or any parallel combination thereof may be mentioned, and a tubular fixed bed reactor is preferred, More preferred is a shell and tube reactor.
  • the operation mode of the addition esterification reactor may be either a batch mode or a continuous mode.
  • the tubular fixed bed reactor is a preferred reactor of the present invention because of its advantages of low manufacturing cost, simple operation, and the like.
  • Fixed bed reactors can be operated in an adiabatic or isothermal manner.
  • Adiabatic reactor can use barrel reactor, catalyst It is fixed in the reactor, and the outer wall of the reactor is insulated and insulated. Since the addition esterification reaction is an exothermic reaction, it is necessary to control the concentration of the reactants to control the temperature rise of the reactor bed, or to recycle to the reactor after cooling some of the reaction products.
  • the inlet is used to dilute the reactant concentration.
  • the isothermal reactor can be a tube-and-tube tubular reactor in which the catalyst is fixed in a column tube, and the shell side passes through cooling water to remove the heat released from the reaction.
  • the reaction temperature of the addition esterification reaction is usually from 50 to 200 ° C, preferably from 60 to 120 ° C, but is sometimes not limited thereto.
  • the reaction pressure should ensure that the reaction is in a liquid phase state.
  • the reaction pressure is atmospheric pressure - lOMPa
  • the optimized pressure is normal pressure ⁇ lMPa, but sometimes it is not limited thereto.
  • the liquid feed space velocity is generally for 20 h to 0.5, preferably 0.5 to 511-1, more preferably 1 to 511-1, but sometimes Not limited to this.
  • the reaction time is usually 0.1 to 10 h, preferably 0.2 to 2 h, but it is not limited thereto.
  • the single pass conversion of the cyclohexene of the addition esterification reaction can generally reach 80% or more, and the single pass selectivity of the cyclohexyl carboxylate such as cyclohexyl acetate can reach 99% or more.
  • the addition esterification product obtained by the addition esterification reaction mainly contains unreacted at least one of a carboxylic acid and cyclohexene, and a reaction product of cyclohexyl carboxylate, and further includes Various inert diluents from cyclohexene sources such as cyclohexane and benzene, etc., depend on the original composition of the cyclohexene source.
  • the addition esterification product can be directly introduced into the cyclohexyl ester hydrogenation step of the present invention as an addition esterification product without any separation or purification.
  • the addition esterification product can be subjected to rectification separation by an esterification product separation system.
  • the addition esterification product separation system may be provided with a rectification separation unit and/or an extractive rectification separation unit in a manner known in the art, and the specific separation scheme is the original composition of the cyclohexene source. related.
  • the general rule is that the addition esterification product is separated into a carboxylic acid stream and a cyclohexyl carboxylate stream (or a mixture of a carboxylic acid and a cyclohexyl carboxylate), and a C6 hydrocarbon by the esterification product separation system. Logistics.
  • the C6 hydrocarbon may represent benzene, cyclohexane, cyclohexene, a mixture of cyclohexane and benzene, a mixture of cyclohexane and cyclohexene, a mixture of cyclohexene and benzene, or a mixture of cyclohexane, cyclohexene and benzene, optionally also Contains other inert diluents.
  • the esterification product separation system may be provided with one or more de C6 hydrocarbon/carboxylic acid columns.
  • the addition esterification product is separated into the de C6 hydrocarbon/carboxylic acid column, and the column can be operated at normal pressure, and is controlled from the top of the column by controlling the heating amount of the column, the reflux ratio, the top of the column and the amount of the column.
  • the resulting mixture of C6 hydrocarbons and carboxylic acid can be further separated into a C6 hydrocarbon stream and a carboxylic acid stream by a de C6 hydrocarbon column as described below.
  • the esterification product separation system may be provided with one or more de C6 hydrocarbon columns and one or more decarboxylation (e.g., acetic acid) columns.
  • the addition esterification product first enters the de-C6 hydrocarbon column for separation.
  • the column can be operated at atmospheric pressure, and the C6 hydrocarbon is recovered from the top of the column by controlling the heating amount of the column, the reflux ratio, the top of the column and the amount of the column.
  • the stream, the stream recovered from the C6 hydrocarbon tower enters the decarboxylation column for separation, and the column can also be operated at atmospheric pressure by controlling the heating amount of the column, the reflux ratio, the top of the column and the amount of the tower.
  • the esterification product separation system may further comprise one or more extraction fine condensing towers, and further separate the obtained C6 hydrocarbon stream into a mixture flow of cyclohexane and benzene, a mixture of cyclohexene and benzene, and a cyclohexane.
  • the esterification product separation system may further comprise one or more de-heavy component columns, and the cyclohexyl carboxylate stream enters the de-heavy-component column, and the heavy components in the stream are further removed by distillation separation.
  • the cyclohexyl carboxylate stream from which the heavy component is removed is obtained, and the separated heavy component is discharged as a by-product.
  • the obtained C6 hydrocarbon stream is a mixture, it can be hydrogenated to produce cyclohexane, or the cyclohexane can be separated therefrom according to the manner described herein, for example, by extractive distillation.
  • the obtained cyclohexylcarboxylate stream or a mixture of a carboxylic acid and a cyclohexyl carboxylate stream (preferably a cyclohexyl carboxylate stream) is introduced into the present invention as the addition esterification product.
  • the cyclohexyl ester hydrogenation step, and the obtained carboxylic acid stream can be recycled as part of the process (1) reaction feed.
  • the obtained mixture of cyclohexane and benzene or a mixture of cyclohexene and benzene may be further separated into a cyclohexane stream, a benzene stream and a cyclohexene stream by extractive rectification, and the obtained ring is obtained.
  • a hexene stream a mixture of cyclohexene and benzene or a mixture of cyclohexene and cyclohexane can be used as the cyclohexyl ester of the present invention
  • the cyclohexene source of the step of use, the obtained benzene stream or a mixture of cyclohexene and benzene can be recycled as part of the reaction feed of step (A), and the obtained cyclohexane stream is used as a by-product It is vented or recycled as part of the reaction feed to step (C).
  • the esterification product separation system may Providing only one or more rectification columns for the removal of carboxylic acids (such as acetic acid) and cyclohexene; one or more rectification columns for removing carboxylic acid and cyclohexene and one or more A rectification column for removing heavy components.
  • the addition esterification product first enters the deacidified olefin column for rectification separation, and the column can be operated at normal pressure, by controlling the heating amount of the column, the reflux ratio, the top of the column and the amount of the column, and the The reacted cyclohexene and carboxylic acid are recovered from the top of the column, recycled to the reaction system, and the cyclohexyl carboxylate is recovered from the column. If the cyclohexyl carboxylic acid product produced from the deacidified olefinic tray contains more heavy components, the cyclohexyl carboxylate product also needs to enter the de-weighting component column to remove heavy components, from the de-weighting component tower. The high purity carboxylic acid cyclohexyl ester product is obtained, and the bottom heavy component is discharged as a by-product.
  • the addition esterification reaction is carried out in one or more reactive distillation columns.
  • the reactive rectification column is identical in form to a conventional rectification column, and generally consists of a column body, a column top condenser, a reflux tank, a reflux pump, a column reactor, and a reboiler.
  • the type of the tower may be a plate column, a packed column, or a parallel combination of the two. Types of plate towers that can be used include floating towers, sieve trays, bubble columns, and the like.
  • the packing used in the packed tower may be a random packing such as a Pall ring, an annulus ring, a saddle type packing, a step ring packing or the like; a structured packing such as a corrugated board packing or a corrugated wire packing may also be used.
  • the addition esterification catalyst (such as the solid acid catalyst) is disposed in the reaction rectification column.
  • the arrangement of the catalyst in the reactive rectification column should meet the following two requirements: (1) to provide sufficient channels for vapor-liquid two-phase passage, or to have relatively large bed voids. Rate (generally required to be at least 50% or more) to ensure that the vapor-liquid two phases can convect through without causing flooding; (2) to have good mass transfer performance, the reactants are transferred from the fluid phase to the catalyst for reaction. At the same time, the reaction product is transferred from the catalyst.
  • the arrangement of various catalysts in the reactive rectification column has been disclosed in the prior literature, and these arrangements can be employed in the present invention.
  • the arrangement of the existing catalysts in the reaction column can be divided into the following three types: (1) The catalyst is directly arranged in the column in the manner of rectifying the filler, and the main method is to have a certain size and shape of the catalyst particles and the rectification packing. Mechanically mixing, or sandwiching the catalyst between the structured packing and the structured packing to form a monolithic packing, or directly forming the catalyst into a rectified packing shape; (2) loading the catalyst into a gas-liquid permeable small container and arranging it On the tray of the reaction tower, or arrange the catalyst in the downcomer of the reaction tower; (3) The catalyst is directly charged into the reaction tower in a fixed bed manner, and the liquid phase flows directly through the catalyst bed, and is set up for the gas phase.
  • a dedicated channel in this way, where the catalyst is placed, alternately arranged by the catalyst bed and the rectification tray, the liquid on the tray passes through the downcomer and the redistributor into the next catalyst bed, in the bed The addition reaction is carried out, and the liquid in the lower portion of the catalyst bed passes through the liquid collector to the next tray.
  • the number of theoretical plates of the reaction condensate column is generally from 10 to 150, preferably from 30 to 100.
  • the addition esterification catalyst such as the solid acid catalyst described
  • the liquid feed space velocity of from 0.1 to ⁇ ⁇ 1, preferably for 20 h to 0.2, more preferably 0.5 to 0.5 to 5h or 21T 1, but Sometimes it is not limited to this.
  • the operating pressure of the reactive rectification column can be operated under negative pressure, normal pressure and pressure.
  • the reactive distillation column has an operating pressure of from -0.0099 MPa to 5 MPa, preferably from atmospheric pressure to IMPa, but is sometimes not limited thereto.
  • the operating temperature of the reactive rectification column is related to the pressure of the reactive rectification column.
  • the temperature distribution of the reaction column can be adjusted by adjusting the operating pressure of the reaction column so that the temperature of the catalyst loading zone is within the active temperature range of the catalyst.
  • the temperature of the addition esterification catalyst bed packing zone is generally from 40 to 200 ° C, preferably from 50 to 200 ° C or from 50 to 180 ° C, more preferably from 60 to 120 ° C or from 60 to 150. °C, but sometimes it is not limited to this.
  • the reflux ratio of the reaction column should meet the requirements of separation and reaction. Under normal circumstances, increasing the reflux ratio is beneficial to improve the separation capacity and reaction conversion rate, but at the same time increase the process energy consumption.
  • this mode (2) if a pure cyclohexene product and a carboxylic acid (such as acetic acid) are used as a reaction raw material, it is theoretically possible to achieve total reflux. However, when there is a small amount of light component impurities in the reaction raw material, a small amount of overhead stream needs to be taken out of the reaction fine column.
  • the reflux ratio of the reactive distillation column is generally from 0.1:1 to total reflux, preferably from 0.1 to 100:1, more preferably from 0.5 to 10:1, but sometimes it is not limited thereto.
  • the addition esterification product mainly contains unreacted at least one carboxylic acid and cyclohexene, and a reaction product of cyclohexyl carboxylate, and further includes various inert diluents derived from a cyclohexene source. For example, cyclohexane and benzene, etc., depending on the original composition of the cyclohexene source.
  • a cyclohexane carboxylate stream or a mixture of a carboxylic acid and a cyclohexyl carboxylate is obtained from the bottom of the reactive distillation column, and a C6 hydrocarbon stream is obtained from the top of the reactive distillation column according to the separation of the bottom of the column. , or a mixture of a carboxylic acid and a C6 hydrocarbon.
  • the C6 hydrocarbon may represent benzene, cyclohexane, cyclohexene, a mixture of cyclohexane and benzene, a mixture of cyclohexane and cyclohexene, Mixtures of cyclohexene with benzene, or mixtures of cyclohexane, cyclohexene and benzene, optionally with other inert diluents.
  • a mixture of carboxylic acid and C6 hydrocarbons it can be separated by distillation into a C6 hydrocarbon stream and a carboxylic acid stream.
  • the mixture is passed through a de-C6 hydrocarbon column for rectification separation.
  • the column can be operated at atmospheric pressure by controlling the heating amount of the column, the reflux ratio, the top of the column and the amount of the column, and is taken from the top of the column.
  • a C6 hydrocarbon stream (such as a cyclohexane stream or a mixture of cyclohexane and benzene) is recovered from the column.
  • the obtained C6 hydrocarbon stream is a mixture, it can be hydrogenated to produce cyclohexane, or the cyclohexane can be separated therefrom according to the method described herein, for example, by extractive distillation.
  • the obtained cyclohexylcarboxylate stream or a mixture of a carboxylic acid and a cyclohexyl carboxylate stream (preferably a cyclohexyl carboxylate stream) is introduced into the present invention as the addition esterification product.
  • the cyclohexyl ester hydrogenation step, and the obtained carboxylic acid stream can be recycled as part of the process (2) reaction feed.
  • the obtained mixture of cyclohexene and cyclohexane or the mixture of cyclohexene and benzene may be further separated into a cyclohexane stream, a benzene stream and a cyclohexene stream by extractive distillation.
  • a cyclohexene stream a mixture of cyclohexene and benzene or a mixture of cyclohexene and cyclohexane may be used as the cyclohexene source of the cyclohexene esterification step of the present invention, or the obtained benzene stream or
  • the mixed stream of cyclohexene and benzene can be recycled as part of the feed to step (A), and the obtained cyclohexane stream is either withdrawn as a by-product or recycled as part of the reaction feed of step (C).
  • the mode (1) in order to carry out the cyclohexene esterification step, the mode (1) may be carried out first, followed by the mode (2).
  • an addition esterification reaction may be carried out according to the mode (1) as described above (at this time, the inverse of the mode (1)
  • the obtained addition esterification product is directly (or after separating some or all of the cyclohexyl carboxylate) as a starting material, as in the past, as follows
  • An addition esterification reaction hereinafter referred to as mode (3)
  • mode (1) and mode (2) can be carried out in the same manner as described above, and the addition esterification catalysts used in each case can be the same or different, independently from the above. Selected in the addition esterification catalyst.
  • the (additional) addition esterification product obtained as described above is hydrogenated with hydrogen in the presence of a hydrogenation catalyst to simultaneously produce cyclohexanol and an alkanol.
  • the alkanol of formula R-CH 2 -OH where the radicals R are the same as for the definition of at least a carboxylic acid, most preferably methyl.
  • the alkanol ethanol is most preferred.
  • examples of the hydrogenation catalyst include a copper-based catalyst, a ruthenium-based catalyst, and a noble metal-based catalyst.
  • a copper-based catalyst is preferable. These catalysts may be used alone or in combination of two or more.
  • examples of the ruthenium-based catalyst include Ru/Al 2 O 3 and Ru-Sn/Al 2 O 3 .
  • examples of the noble metal-based catalyst include Pt/Al 2 O 3 , Pd-Pt/Al 2 O 3 and Pd /C.
  • examples of the copper-based catalyst include a copper-based catalyst containing zinc and a copper-based catalyst containing chromium.
  • a copper-based catalyst comprising the following components (more preferably consisting of the following components) is preferable.
  • component (a): component (b): component (c): component (d) has a ratio of 5 to 60: 10 to 50: 5 to 60: 0.2 to 2, preferably 10 to 50: 15 to 45: 15 to 55: 0.2 to 2, more preferably Choose 30 to 45: 20 to 35: 20 to 50: 0.5 to 1.5.
  • the copper-based catalyst generally needs to be reduced prior to use. Moreover, it should be understood that in the art, the catalyst is generally traded and stored in the form of a precursor, and although the catalyst precursor does not directly catalyze the reaction, the catalyst precursor is conventionally referred to as a "catalyst.”
  • the copper-based catalyst of the present invention (actually a copper-based catalyst precursor) is catalytically activated after reduction, and this is usually done by an operator of an industrial plant, which is well known to those skilled in the art, and the present invention is hereby No longer.
  • the copper-based catalyst can be formed into various desired shapes, such as a molded pellet, or a state before molding, such as a powder.
  • the copper-based catalyst can be produced by a production method comprising the following steps (la) and (2a).
  • the coprecipitation method refers to reacting two or more metal cations in a solution in a solution with a precipitating agent to precipitate a metal cation in the solution to obtain a uniform precipitation of various components, or a precipitate mixture formed or
  • a solid solution precursor is subjected to filtration, washing, and calcination (thermal decomposition of a precipitation mixture or a solid solution precursor) to obtain a composite metal oxide.
  • the coprecipitation method can be carried out in different ways, either by adding a solution containing a metal cation to the precipitant solution, or by adding the precipitant solution to the solution containing the metal cation, or The solution containing the metal cation and the precipitant solution are simultaneously added to the solvent.
  • the coprecipitation method may comprise, for example, the following steps:
  • (I) a (water) solution prepared by mixing a metal soluble salt in a predetermined ratio, the metal being referred to as (a) copper; (b) zinc; and (c) being selected from the group consisting of aluminum, gallium, tin, titanium, zirconium, One or more metals of chromium, molybdenum, tungsten, manganese, lanthanum, lanthanide and lanthanide metals, preferably selected from aluminum, One or more metals of gallium, tin, titanium, zirconium, chromium, molybdenum, tungsten, manganese, lanthanum, cerium and lanthanum, said ratio being converted to the corresponding metal oxide, in parts by mass, metal ( a): metal (b): metal (c) ratio is 5 to 60: 10 to 50: 5 to 60, preferably 10 to 50: 15 to 45: 15 to 55, more preferably 30 to 45: 20 to 35: 20 to 50;
  • step (II) adding a precipitant (water) solution to the mixed solution obtained in the step (I) at 15 ° C to 80 ° C, adjusting the pH to 6 ⁇ 9, to form a mixed precipitate, by selecting the precipitate Agent, such that the resulting mixed precipitate can be pyrolyzed into a metal oxide;
  • step (III) After the precipitation system obtained in the step (II) is maintained at 30 ° C to 80 ° C for 1 to 48 hours, the precipitate is filtered and washed until the metal cation concentration in the filtrate is less than 100 ug / g. Then, it is dried at 100 ° C ⁇ 20 CTC for 3 ⁇ 48 h, and then calcined at 250 ° C ⁇ 400 ° C for 3 ⁇ 48 h to obtain a powdery composite metal oxide.
  • the meaning of "soluble” means that the concentration of the metal salt in the solution (for example, solubility in water) can satisfy the composition requirements of the produced composite metal oxide, and for example, the A metal nitrate, sulfate, hydrochloride, acetate or hydrate.
  • step (?) it is preferred to add the (aqueous) solution of the precipitating agent to the mixed solution obtained in the step (I) under stirring, which is advantageous in improving the uniformity of the catalyst.
  • the precipitating agent for example, one or more of sodium hydroxide, potassium hydroxide, sodium carbonate, potassium carbonate, ammonium carbonate, ammonia, urea, sodium oxalate, potassium oxalate and ammonium oxalate may be mentioned. .
  • the concentration of the metal-soluble salt and the precipitant in the solution thereof is not particularly limited as long as the purpose of producing the composite metal oxide can be achieved, and those skilled in the art can appropriately select them as needed.
  • the metal soluble salt and/or precipitant may be dissolved in water or dissolved in Among the aqueous solvents, such as ethanol, or a mixture of any ratio of water to ethanol, those skilled in the art can arbitrarily select them as appropriate.
  • the ratio is 5 to 60: 10 to 50: 5 to 60: 0.2 to 2, preferably 10 to 50: 15 to 45: 15 to 55: 0.2 to 2, more preferably 30 to 45: 20 to 35: 20 to 50: 0.5 to 1.5.
  • the (metal) hydroxide and/or alkaline earth metal hydroxide (water) of the composite metal oxide obtained in the step (la) may be mentioned.
  • the solution concentration is, for example, 0.5 to 5 wt%) is impregnated, filtered, dried, and calcined to obtain a copper-based catalyst (precursor) of the present invention.
  • the immersion temperature is 30 ° C ⁇ 80 ° C
  • the immersion time is l ⁇ 48 h
  • the drying temperature is 100 ° C ⁇ 200 ° C
  • the drying time is 3 ⁇ 48 h
  • the calcination temperature is from 250 ° C to 400 ° C and the calcination time is from 3 to 48 h.
  • the alkali metal hydroxide and/or alkaline earth metal hydroxide is preferably one or more of potassium hydroxide, sodium hydroxide and barium hydroxide.
  • the obtained copper-based catalyst is in the form of a powder, which can be molded into a shape desired by a user as required by the user.
  • the copper-based catalyst exhibits hydrogenation activity after reduction in a hydrogen atmosphere, whether it is a powdery product or a molded product.
  • the reduction process may be either an additional step prior to use of the catalyst or during the hydrogenation step of the cyclohexyl ester of the present invention, preferably as an additional step prior to use of the catalyst.
  • the hydrogenation reaction is carried out in one or more hydrogenation reactors.
  • the hydrogenation reactor for example, a tank reactor, a fixed bed reactor, an ebullated bed reactor, a fluidized bed reactor or any parallel combination thereof may be mentioned, and a tubular fixed bed reactor is preferred, and more preferably Shell-and-tube tubular reactor.
  • the hydrogenation reaction can be carried out in a batch manner or in a continuous manner.
  • the batch reaction generally adopts a reaction vessel as a reactor, and the addition esterification product and the hydrogenation catalyst are put into a reaction vessel, and hydrogen is introduced under a certain temperature and pressure to carry out a reaction.
  • the reaction product is taken from the kettle.
  • the product is discharged, the product is separated, and the next batch of material is put into the reaction.
  • the continuous hydrogenation reaction may be carried out in a shell-and-tube type tubular reactor in which a hydrogenation catalyst is fixed in a column tube, and cooling water is passed through the shell side to remove the released heat of the reaction.
  • the reaction temperature of the hydrogenation reaction is generally from 150 to 400 ° C, preferably from 200 to 300 ° C.
  • the reaction pressure of the hydrogenation reaction is usually from atmospheric pressure to 20 MPa, preferably from atmospheric pressure to 10 MPa, more preferably from 4 to 10 MPa.
  • the molar ratio of the esterified product is from 1 to 1000:1, preferably from 5 to 100:1.
  • the liquid feed space velocity is from 0.1 to 20, preferably from 0.2 to 2 h.
  • the reaction time is from 0.2 to 20 h, preferably from 0.5 to 5 h, more preferably from 1 to 5 h.
  • the hydrogenated product obtained by the hydrogenation reaction is sent to a hydrogenation product separation system for separation.
  • the hydrogenation product is introduced into a gas-liquid separation tank for gas-liquid separation, the gas phase is mainly hydrogen, and may also include various inert diluents such as cyclohexane and benzene from a cyclohexene source, depending on The original composition of the cyclohexene source in the cyclohexene esterification step.
  • the separated hydrogen is compressed by a compressor and recycled to the hydrogenation reactor.
  • the liquid phase product mainly contains alkanols such as ethanol and cyclohexanol, and may also contain a certain amount of alkaloids of carboxylic acid (such as ethyl carboxylate) and cyclohexanone, and may also contain a certain amount of unreacted carboxylic acid. Acid cyclohexyl ester, and a small amount of reboil (dimer ketone).
  • the liquid phase product can be separated by rectification and/or extraction separation.
  • the rectification process may employ a batch scheme or a continuous scheme, wherein continuous distillation is preferably employed to separate the hydrogenation product.
  • This continuous rectification process requires the use of a series of columns to separate the various components.
  • Various separation schemes can be designed in accordance with the order of separation of the components, and the present invention is preferably a sequential separation scheme.
  • the addition of the addition esterification product to hydrogen in the presence of a carboxylic acid hydrogenation catalyst is optionally subjected to a carboxylic acid hydrogenation reaction (referred to as a carboxylic acid hydrogenation) prior to the cyclohexyl ester hydrogenation step.
  • a carboxylic acid hydrogenation reaction referred to as a carboxylic acid hydrogenation
  • Step) to pre-convert the free carboxylic acid (and possibly benzene) that may be present in the addition esterification product to an alkanol (and cyclohexane).
  • the reaction product obtained by the carboxylic acid hydrogenation step (optionally after separating some or all of the alkanol and/or cyclohexane) is still regarded as the addition esterification product, and can be carried out in exactly the same manner as described above.
  • the cyclohexyl ester hydrogenation step is optionally subjected to a carboxylic acid hydrogenation reaction (referred to as a carboxylic acid hydrogenation) prior
  • the carboxylic acid hydrogenation catalyst may use any catalyst conventionally used in the art for hydrogenating a carboxylic acid to produce a corresponding alcohol, but it is preferred that the carboxylic acid hydrogenation catalyst consists of a main active component of 0.1. It is composed of a carrier of 30% by weight, an auxiliary agent of 0.1 to 25% by weight and the balance.
  • the primary active component is selected from one or more of platinum, palladium, rhodium, tungsten, molybdenum and cobalt.
  • the adjuvant is selected from one or more of the group consisting of tin, chromium, aluminum, zinc, calcium, magnesium, nickel, titanium, zirconium, hafnium, tantalum, niobium and gold.
  • the support is selected from one or more of the group consisting of silica, alumina, titania, zirconia, activated carbon, graphite, carbon nanotubes, calcium silicate, zeolite and aluminum silicate.
  • the carboxylic acid hydrogenation step and the cyclohexyl ester hydrogenation step may be carried out in different reactors or in different regions of the same reactor.
  • the carboxylic acid hydrogenation step can be carried out in a separate carboxylic acid hydrogenation reactor.
  • a carboxylic acid hydrogenation reactor for example, a tank reactor, a fixed bed reactor, an ebullated bed reactor, a fluidized bed reactor or any parallel combination thereof may be mentioned, and a tubular fixed bed reactor is preferred, More preferred is a shell and tube reactor.
  • the reaction temperature of the carboxylic acid hydrogenation reaction is from 100 to 400 ° C, preferably from 180 to 300. C.
  • the reaction pressure of the carboxylic acid hydrogenation reaction is from 0.1 to 30 MPa, preferably from 2 to 10 MPa.
  • the molar ratio of hydrogen to the free carboxylic acid is from 1 to 500:1, preferably from 5 to 50:1.
  • the carboxylic acid hydrogenation reaction is carried out in a continuous manner, the liquid feed space velocity of from 0.1 to 5H-], preferably 0.2 to 2h.
  • the reaction time is from 0.5 to 20 h, preferably from 1 to 5 h.
  • the method of co-producing cyclohexanol and an alkanol optionally further comprises one of the following steps (1), (II) and (III), or any combination thereof. These steps can be carried out in a manner conventionally known in the art, and the description thereof is omitted here.
  • step (III) the reaction of dehydrogenation of cyclohexane to benzene is quite easy, cyclohexane It is only necessary to produce benzene at a high conversion rate and high selectivity in the presence of a catalyst having a single dehydrogenation function and under suitable reaction conditions.
  • a person skilled in the art can select a suitable implementation method by referring to JP285001/87, WO2009/131769, CN1038273.
  • bifunctional or multifunctional catalysts such as catalytic reforming catalysts having both dehydrogenation and acid bifunctionality.
  • the invention can dehydrogenate cyclohexane to benzene in at least two ways, one is to establish a separate cyclohexane dehydrogenation device, and perform cyclohexane dehydrogenation reaction in the presence of a monofunctional or multifunctional dehydrogenation catalyst; The other is to use the existing catalytic reforming unit to treat the same ring. It should be understood that, based on prior knowledge, the presence of benzene does not adversely affect the reaction of dehydrogenation of cyclohexane to benzene, and thus the cyclohexane may contain benzene.
  • the method of co-producing cyclohexanol and an alkanol optionally further comprises the steps described below or a combination thereof. These steps can be carried out in a manner conventionally known in the art, and the description thereof is omitted here.
  • Step (v h recovers hydrogen separated from any step of the method of co-producing cyclohexanol and alkanol, and recycles the hydrogen to the cyclohexyl ester hydrogenation step.
  • cyclohexanone can be produced by using cyclohexanol previously produced as a raw material. Accordingly, the present invention also relates to a process for producing cyclohexanone comprising the steps of producing cyclohexanol according to the aforementioned method of the present invention, and the step of producing cyclohexanone using the cyclohexanol as a starting material.
  • the step of producing cyclohexanone using cyclohexanol as a starting material can be carried out in a manner conventionally known in the art, and the description thereof is omitted here.
  • caprolactam can be produced by using the cyclohexanone produced as described above as a raw material. Accordingly, the present invention also relates to a process for producing caprolactam comprising the steps of producing cyclohexanone according to the aforementioned production method of the present invention, and the step of producing caprolactam using the cyclohexanone as a starting material. ,
  • the step of producing caprolactam using cyclohexanone as a starting material can be carried out in a manner conventionally known in the art, and the description thereof is omitted here.
  • the apparatus for co-producing cyclohexanol and an alkanol comprises a hydrogenation reaction unit A (benzene hydrogenation reactor), an optional hydrogenation product separation unit A, and an addition esterification reaction unit (esterification reaction). And an optional addition esterification product separation unit, a hydrogenation reaction unit B (ester hydrogenation reactor), and a hydrogenation product separation unit B (ester hydrogenation product separation unit), wherein
  • the hydrogenation product from the hydrogenation reaction unit A is separated to obtain cyclohexene, a mixture of cyclohexene and benzene, or cyclohexene and cyclohexane. a mixture (corresponding to the aforementioned step (B));
  • the hydrogenation product from the hydrogenation reaction unit A and/or a mixture of cyclohexene, cyclohexene and benzene from the hydrogenation product separation unit A Or an addition and esterification reaction of a mixture of cyclohexene and cyclohexane with a carboxylic acid in the presence of an addition esterification catalyst to form an addition esterification product containing cyclohexyl carboxylate (corresponding to the aforementioned cyclohexene) Esterification step);
  • the addition esterification product from the addition esterification reaction unit is separated to obtain a cyclohexyl carboxylate or a cyclohexyl carboxylate and a carboxylic acid.
  • the addition esterification product from the addition esterification reaction unit and/or the cyclohexyl carboxylate or carboxylic acid from the addition esterification product separation unit a mixture of cyclohexyl ester and a carboxylic acid is hydrogenated with hydrogen in the presence of a hydrogenation catalyst to form a hydrogenation product containing cyclohexanol and an alkanol (corresponding to the aforementioned cyclohexyl ester hydrogenation step);
  • the hydrogenation product from the hydrogenation reaction unit B is separated to obtain cyclohexanol and an alkanol.
  • the hydrogenation reaction unit A is provided with one or more reactors in parallel, the reactor type being selected from fixed bed reactors and/or tank reactors.
  • the addition esterification reaction unit is provided with one or more reactors X connected in parallel, the reactor X being selected from the group consisting of a tank reactor, a fixed bed reactor, an ebullated bed reactor and a stream.
  • the reactor X being selected from the group consisting of a tank reactor, a fixed bed reactor, an ebullated bed reactor and a stream.
  • One or more of the chemical bed reactors for implementation in the manner described above (1) (1) A cyclohexene esterification step.
  • the addition esterification reaction unit is provided with at least one reactive distillation column for carrying out the cyclohexene esterification step according to the above mode (2).
  • one or more of the reactors X (also referred to as pre-addition esterification reaction units) connected in parallel are arranged in series before the reactive distillation column for carrying out the above-mentioned manner (3).
  • a cyclohexene esterification step is preferred.
  • the addition esterification product separation unit is provided with at least one rectification column.
  • the hydrogenation reaction unit B is provided with one or more reactors connected in parallel, the reactor type being selected from the group consisting of a tank reactor, a fixed bed reactor, an ebullated bed reactor and a fluidized bed reactor.
  • the reactor type being selected from the group consisting of a tank reactor, a fixed bed reactor, an ebullated bed reactor and a fluidized bed reactor.
  • One or more of them are preferably provided with one or more shell-and-tube tubular reactors connected in parallel.
  • the hydrogenation product separation unit B is provided with at least one rectification column.
  • the apparatus for co-producing cyclohexanol and alkanol also optionally includes at least one of the following circulation devices.
  • Circulating Apparatus A Benzene and/or hydrogen from any unit of the apparatus for co-producing cyclohexanol and alkanol is recovered, and the benzene and/or hydrogen are recycled to the hydrogenation reaction unit A.
  • Circulating device B recovering carboxylic acid and/or cyclohexene from any unit of the apparatus for co-producing cyclohexanol and alkanol, and recycling the carboxylic acid and/or cyclohexene to the addition esterification Reaction unit.
  • Circulating device C Hydrogen gas is recovered from any unit of the apparatus for co-producing cyclohexanol and alkanol, and the hydrogen is recycled to the hydrogenation reaction unit B.
  • the addition esterification reaction unit is selected from the group consisting of a reactor (such as a tank reactor, a fixed bed reactor, a fluidized bed reactor, and boiling).
  • One of the bed reactors or any combination thereof and the reactive distillation column (the number of theoretical plates is, for example, 10 to 150, preferably 30 to 100, preferably a plate column or a packed column) or any combination thereof, preferably a reaction rectification column (for carrying out the cyclohexene esterification step according to the above mode (2)) or a series combination of the reactor and the reactive rectification column, more preferably the reactor and the reaction refinement
  • a series combination of distillation columns, most preferably the reactor is a series combination upstream of the reactive rectification column (for carrying out the cyclohexane according to the aforementioned mode (3) Enesterification step).
  • cyclohexene and acetic acid enter the esterification reactor 1, and an addition esterification reaction is carried out under the action of an addition esterification catalyst, and the esterification product stream is added to the ester hydrogenation reaction via line 11.
  • the ester hydrogenation reaction is carried out by contacting the hydrogen gas under the action of the ester hydrogenation catalyst, and the ester hydrogenation product stream 22 enters the ester hydrogenation product separation unit 3, and is separated to obtain a cyclohexanol stream 31 and an ethanol stream 32.
  • the benzene and cyclohexane-containing stream and the acetic acid enter the reaction fine column 1, and the addition esterification reaction is carried out under the action of the addition esterification catalyst, and the reaction product is separated at the same time.
  • the distillation section obtains a benzene + cyclohexane stream 13 and an acetic acid stream 14, the acetic acid stream is recycled back to the reactive rectification column 1, the bottom of the column is obtained to obtain a cyclohexyl acetate stream, and the ester 12 is fed to the ester hydrogenation reactor 2 via the line 12, in the ester hydrogenation catalyst.
  • the ester hydrogenation reaction is carried out by contact with hydrogen, and the ester hydrogenation product stream 22 is passed to the ester hydrogenation product separation unit 3, and is separated to obtain a cyclohexanol stream 31, an ethanol stream 32, a high boiler stream 33 and an acetate ring.
  • the hexyl ester stream 34, the cyclohexanol stream 31 and the ethanol stream 32 are used as product discharge units, the high boiler stream 33 is used as a by-product unit, and the cyclohexyl acetate stream 34 is recycled to the ester hydrogenation reactor 2.
  • cyclohexene and acetic acid enter an addition esterification reaction system, and an esterification reaction is carried out under the action of an addition esterification catalyst, and the reaction product is sent to an esterification product separation system to obtain an acetic acid ring.
  • a hexyl ester stream, a cyclohexene stream and an acetic acid stream, a cyclohexene stream and an acetic acid stream are recycled back to the addition esterification reaction system, and the cyclohexyl acetate stream enters the ester hydrogenation reaction system, under the action of the ester hydrogenation catalyst, Hydrogen contact, hydrogenation reaction of cyclohexyl acetate occurs, and the hydrogenation product is sent to the hydrogenation product separation system to obtain a cyclohexanol stream, an ethanol stream, a cyclohexyl acetate stream and a hydrogen stream, a cyclohexanol stream and ethanol.
  • the logistics as a product discharge device, the cyclohexyl acetate stream and the hydrogen stream are recycled to the ester hydrogenation reaction. System.
  • benzene and hydrogen enter the benzene hydrogenation reactor 1, and under the action of the benzene hydrogenation catalyst, a hydrogenation reaction is carried out, and the benzene hydrogenation product stream enters the addition esterification reactor 2 via line 11 with
  • the acetic acid entering through the line 21 is mixed, and the addition esterification reaction is carried out under the action of the solid acid catalyst, and the addition esterification product stream is introduced into the hydrogenation reactor 3 via the line 22, and is contacted with hydrogen under the action of the noble metal catalyst.
  • the hydrogenation reaction of benzene and carboxylic acid is carried out, and the hydrogenation product stream enters the hydrogenation reactor 4, and under the action of the ester hydrogenation catalyst, the ester hydrogenation reaction is carried out by contact with hydrogen, and the hydrogenation product stream enters the product separation unit 5,
  • the cyclohexanol stream 51, the ethanol stream 52, and the cyclohexane stream 53 are separated.
  • benzene and hydrogen enter the benzene hydrogenation reactor 1, and under the action of the benzene hydrogenation catalyst, a hydrogenation reaction is carried out, and the benzene hydrogenation product stream enters the addition esterification reactor 2 via line 11,
  • the acetic acid entering via line 21 is mixed, and the addition esterification reaction is carried out under the action of the solid acid catalyst, and the addition esterification product stream is passed through line 22 to the addition esterification product separation unit 3, and the cyclohexane stream 31 is separated.
  • the by-product effluent device, the acetic acid/cyclohexyl acetate stream enters the carboxylic acid hydrogenation reactor 4, and under the action of the carboxylic acid hydrogenation catalyst, the carboxylic acid hydrogenation reaction is carried out by contact with hydrogen, and the carboxylic acid hydrogenation product stream enters the ester hydrogenation reaction.
  • the ester hydrogenation reaction is carried out by contacting with hydrogen, and the ester hydrogenation reaction product stream enters the hydrogenation product separation unit 6, and is separated to obtain cyclohexanol.
  • benzene and hydrogen enter the benzene hydrogenation reactor 1, and under the action of the benzene selective hydrogenation catalyst, the hydrogenation reaction is carried out, and the benzene hydrogenation product stream enters the reaction distillation column 2 via the line 11,
  • the acetic acid entering through the line 21 is mixed, and under the action of the solid acid catalyst, the addition and esterification reaction is carried out, and the esterification product is separated at the same time, and the cyclohexane/benzene stream is obtained from the top of the reaction distillation column 2, from the reaction fine
  • the bottom of the distillation column 2 obtains an acetic acid/cyclohexyl acetate stream; the cyclohexane/benzene stream enters the addition esterification product separation unit 3 via line 22, and is separated to obtain a cyclohexane stream and a benzene stream, and the benzene stream is recycled to the benzene group.
  • Hydrogen reactor 1 cyclohexane stream as a by-product device, acetic acid / cyclohexyl acetate stream into the carboxylic acid hydrogenation reactor 4, under the action of a carboxylic acid hydrogenation catalyst, hydrogenation reaction with hydrogen
  • the carboxylic acid hydrogenation product stream enters the ester hydrogenation reactor 5, and under the action of the ester hydrogenation catalyst, the ester hydrogenation reaction is carried out by contact with hydrogen, and the ester hydrogenation reaction product stream enters the hydrogenation product separation unit 6
  • the cyclohexanol stream 61 and the ethanol stream 62 are separated.
  • benzene and hydrogen enter the benzene hydrogenation reactor 1, and under the action of the benzene hydrogenation catalyst, a hydrogenation reaction is carried out, and the benzene hydrogenation product stream enters the addition esterification reactor 2 via line 11,
  • the acetic acid entering via line 21 is mixed, and the addition esterification reaction is carried out under the action of the solid acid catalyst, and the addition esterification product stream is passed through line 22 to the addition esterification product separation unit 3, and the cyclohexane stream 31 is separated.
  • the cyclohexyl acetate stream enters the ester hydrogenation reactor 4, and the ester hydrogenation reaction is carried out by contact with hydrogen under the action of the ester hydrogenation catalyst, and the ester hydrogenation product stream enters the ester hydrogenation product separation unit 5, and is separated to obtain a ring.
  • benzene and hydrogen enter the benzene hydrogenation reactor 1, and under the action of the benzene hydrogenation catalyst, a hydrogenation reaction is carried out, and the benzene hydrogenation product stream enters the reaction condensate column 2 via the pipeline 11, and the pipeline 21 entering the acetic acid mixture, under the action of the solid acid catalyst, performing the addition esterification reaction, simultaneously performing the separation of the addition esterification product, and obtaining the cyclohexane/benzene/acetic acid stream from the top of the reaction rectification column 2,
  • the bottom of the reaction rectification column 2 is obtained to obtain a cyclohexyl acetate stream; the cyclohexane/benzene/acetic acid stream enters the addition esterification product separation unit 3 via line 22, and is separated to obtain a cyclohexane stream, a benzene stream and an acetic acid stream, and a ring
  • the hexane stream is discharged as a by-product
  • ester hydrogenation reaction is carried out in contact with hydrogen, and the ester hydrogenation product stream enters the ester hydrogenation product separation unit 5, and is separated to obtain a cyclohexanol stream 52, an ethanol stream 53, and an acetic acid ring.
  • Ester stream 51 stream 54 and high boilers, cyclohexanol ethanol stream as a product stream and a means, as by-products of high boiling stream out means, cyclohexyl acetate ester stream is recycled back to the hydrogenation reactor 4.
  • benzene and hydrogen enter the benzene hydrogenation reactor 1, and under the action of the benzene hydrogenation catalyst, the hydrogenation reaction is carried out, and the benzene hydrogenation product stream enters the benzene hydrogenation product separation unit 2 via the line 11
  • the cyclohexane stream 21 and the cyclohexene/benzene stream 22 are separated, the cyclohexane stream 21 is used as a by-product output unit, and the cyclohexanyl/benzene stream is passed via line 22 to the addition esterification reactor 3, and enters via line 31.
  • the acetic acid is mixed, and the addition esterification reaction is carried out under the action of the solid acid catalyst, and the addition esterification product stream is added to the esterification product via line 32.
  • the separation unit 4 after separation, obtains a cyclohexene/benzene stream 41, an acetic acid stream 42 and a cyclohexyl acetate stream 43 , and the cyclohexene/benzene stream 41 is recycled to the benzene hydrogenation reactor 1 or further separated into a cyclohexene stream.
  • the ester hydrogenation product stream 52 is passed to an ester hydrogenation product separation unit 6 which is separated to provide a cyclohexanol stream 62, an ethanol stream 63, a cyclohexyl acetate stream 61 and a high boiler stream 64, a cyclohexanol stream 62 and an ethanol stream 63.
  • a high boiler stream 64 is used as a by-product discharge unit, and a cyclohexyl acetate stream 61 is recycled to the ester hydrogenation reactor 5.
  • benzene and hydrogen enter the benzene hydrogenation reactor 1, and under the action of the benzene hydrogenation catalyst, a hydrogenation reaction is carried out, and the benzene hydrogenation product stream enters the benzene hydrogenation product separation unit 2 via the line 11
  • the cyclohexane stream 21 and the cyclohexene/benzene stream 22 are separated, and the cyclohexane stream is passed as a by-product.
  • the cyclohexene/benzene stream enters the reactive rectification column 3 via line 22 and is mixed with acetic acid entering via line 31.
  • the addition esterification reaction is carried out, and the addition esterification product is separated, the acetic acid/benzene stream is obtained from the top of the reaction rectification column 3, and the acetic acid ring is obtained from the bottom of the reaction rectification column 3
  • the hexyl ester stream, the acetic acid/benzene stream enters the addition esterification product separation unit 4 via line 33, and the benzene stream 41 and the acetic acid stream 42 are separated, the benzene stream 41 is recycled to the benzene hydrogenation reactor 1, and the acetic acid stream 42 is recycled.
  • the esterification reactor 3, the cyclohexyl acetate stream 32 enters the ester hydrogenation reactor 5, and the ester hydrogenation reaction is carried out by contact with hydrogen under the action of the ester hydrogenation catalyst, and the ester hydrogenation product is obtained.
  • Stream 52 enters the ester hydrogenation product separation unit 6, which is separated to obtain a cyclohexanol stream 62, an ethanol stream 63, a cyclohexyl acetate stream 61 and a high boiler stream 64, a cyclohexanol stream and an ethanol stream as product exit means,
  • the boiler stream is passed as a by-product and the cyclohexyl acetate stream is recycled to the ester hydrogenation reactor 5.
  • benzene and hydrogen enter the benzene hydrogenation reactor 1, and under the action of the benzene hydrogenation catalyst, a hydrogenation reaction is carried out, and the benzene hydrogenation product stream enters the benzene hydrogenation product separation unit 2 via line 11
  • the benzene stream is recycled to the benzene hydrogenation reactor 1 via line 21, and the cyclohexene/cyclohexane stream is passed through line 22 to the addition esterification reactor 3
  • the acetic acid entering through the line 31 is mixed, and the addition esterification reaction is carried out under the action of the solid acid catalyst, and the addition esterification product stream is passed through the line 32 to the addition esterification product separation unit 4, and after separation, the cyclohexane stream is obtained.
  • Cyclohexene stream 42, acetic acid stream 43 and cyclohexyl acetate stream 44, cyclohexane stream 41 as a by-product The apparatus, the acetic acid stream 43 and the cyclohexene stream 42 are recycled back to the addition esterification reactor 3, and the cyclohexyl acetate stream 44 is fed to the ester hydrogenation reactor 5, and is contacted with hydrogen under the action of the ester hydrogenation catalyst.
  • ester hydrogenation product stream 51 is passed to ester hydrogenation product separation unit 6 and is separated to provide cyclohexanol stream 62, ethanol stream 63, cyclohexyl acetate stream 61 and high boiler stream 64, cyclohexanol stream 62.
  • the ethanol stream 63 is used as a product discharge unit
  • the high boiler stream 64 is used as a by-product discharge unit
  • the cyclohexyl acetate stream 61 is recycled to the ester hydrogenation reactor 5.
  • benzene and hydrogen enter the benzene hydrogenation reactor 1, and under the action of the benzene hydrogenation catalyst, a hydrogenation reaction is carried out, and the benzene hydrogenation product stream enters the benzene hydrogenation product separation unit 2 via line 11 Separating to obtain a cyclohexene/cyclohexane stream and a benzene stream, the benzene stream is recycled to the benzene hydrogenation reactor 1 via line 21, and the cyclohexene/cyclohexane stream is passed through line 22 to the reaction rectification column 3, 31 entering the acetic acid mixture, under the action of the solid acid catalyst, performing the addition esterification reaction, simultaneously performing the separation of the addition esterification product, and obtaining the acetic acid/cyclohexane stream from the top of the reaction rectification column 3, from the reaction fine
  • the bottom of the distillation column 3 is subjected to a cyclohexyl acetate stream, and the ace
  • the ester hydrogenation product stream enters the ester hydrogenation product separation unit 6 and is separated to obtain a cyclohexanol stream 62, an ethanol stream 63, a cyclohexyl acetate stream 61 and a high boiler stream 64, a cyclohexanol stream.
  • the ethanol stream is used as a product discharge device
  • the high boiler stream is used as a by-product discharge device
  • the cyclohexyl acetate stream is recycled to the ester hydrogenation reactor 5.
  • benzene and hydrogen enter the benzene hydrogenation reactor 1, and under the action of the benzene hydrogenation catalyst, the hydrogenation reaction is carried out, and the benzene hydrogenation product stream enters the benzene hydrogenation product separation unit 2 via the line 11
  • the cyclohexane stream and the cyclohexene stream and the benzene stream are separated, the cyclohexane stream 21 is used as a by-product discharge unit, the benzene stream is recycled to the benzene hydrogenation reactor 1 via line 22, and the cyclohexene stream is passed to the addition ester via line 23.
  • the reactor 3 is mixed with acetic acid entering via line 31, and subjected to an addition esterification reaction under the action of a solid acid catalyst, and the addition esterification product stream is passed through line 32 to the addition esterification product separation unit 4, and is separated.
  • a cyclohexene/acetic acid stream 41 and a cyclohexyl acetate stream 42 are obtained, the cyclohexene/acetic acid stream is recycled back to the addition esterification reactor 3, and the cyclohexyl acetate stream is passed to the ester hydrogenation reactor 5 in the ester hydrogenation catalyst.
  • the ester hydrogenation reaction is carried out in contact with hydrogen, and the ester hydrogenation product stream enters the ester addition.
  • the hydrogen product separation unit 6 is separated to obtain an ethanol stream 63, a cyclohexanol stream 62, a cyclohexyl acetate stream 61 and a high boiler stream 64, a cyclohexanol stream and an ethanol stream as product outlets, and a high boiler stream as a by-product.
  • the cyclohexyl acetate stream is recycled to the ester hydrogenation reactor 5.
  • benzene and hydrogen enter the benzene hydrogenation reactor 1, and under the action of the benzene hydrogenation catalyst, a hydrogenation reaction is carried out, and the benzene hydrogenation product stream enters the benzene hydrogenation product separation unit 2 via line 11
  • the cyclohexane stream, the cyclohexene stream and the benzene stream are separated, the cyclohexane 21 stream is used as a by-product discharge device, the benzene stream 22 is recycled to the benzene hydrogenation reactor 1, and the cyclohexene stream is passed through the line 23 to the reaction distillation column. 3.
  • the ester hydrogenation product stream enters the ester hydrogenation product separation unit 5, and is separated to obtain an ethanol stream 53, a cyclohexanol stream 52, a cyclohexyl acetate stream 51, and a high boiler stream 54, a ring.
  • the hexanol stream and the ethanol stream are used as product outlets, the high boiler stream is used as a by-product discharge unit, and the cyclohexyl acetate stream is recycled to the ester hydrogenation reactor 5 .
  • benzene and hydrogen enter the benzene hydrogenation reactor 1, and under the action of the benzene selective hydrogenation catalyst, the hydrogenation reaction is carried out, and the benzene hydrogenation product stream enters the addition esterification reactor via line 11.
  • ester hydrogenation product stream Entering the ester hydrogenation reactor 4, under the action of the ester hydrogenation catalyst, contacting with hydrogen to carry out ester hydrogenation reaction, the ester hydrogenation product stream enters the ester hydrogenation product separation unit 5, and is separated to obtain a cyclohexanol stream 51, ethanol Stream 52, high boiler stream 53 and cyclohexyl acetate stream 54, cyclohexanol stream and ethanol stream as product outlets, high boiler stream as a by-product unit, acetic acid loop Ester stream was recycled back to hydrogenation reactor 4.
  • Example 1 Example 1
  • the catalysts of Examples 1 to 6 were prepared according to the following procedure: A certain amount of soluble metal salt was weighed according to the formula of Table 1, placed in a 2000 mL three-necked flask, dissolved in water to prepare a solution of about 1000 mL, and the flask was equipped with a stirrer and a pH meter. And the thermometer, and the flask is placed in a constant temperature water bath with adjustable temperature, the stirring is turned on, the temperature of the constant temperature water bath is adjusted, a certain concentration of the precipitant solution is gradually dropped into the flask, the dropping speed of the aqueous solution of the precipitating agent is controlled, and the temperature of the solution is raised. High control is within 1 °C.
  • the solution precipitates and gradually increases with increasing pH.
  • the aqueous solution of the precipitant is stopped. Then maintain a certain temperature for a certain period of time while continuing to stir. Stirring was stopped, and the mixture was naturally cooled to room temperature.
  • the precipitate was centrifugally filtered on a high-speed centrifuge and washed 5 times with deionized water. The resulting precipitate was dried in an oven and transferred to a muffle furnace for calcination to obtain a mixed metal oxide.
  • the metal oxide is immersed in a certain concentration of an alkali solution at room temperature, and the impregnation liquid is removed by vacuum filtration, and the mixture is dried in an oven, transferred to a muffle furnace for firing, and finally a mixed metal oxide is obtained.
  • the composition of the obtained sample was analyzed by ICP method. The specific manufacturing conditions and results are shown in Table 1.
  • Examples 7 to 15 are hydrogenation tests of cyclohexyl acetate obtained by carrying out the catalysts obtained in Examples 1 to 6 in an autoclave.
  • the test procedure is as follows: A certain amount of the catalyst powder is placed in a 500 mL autoclave, and 250 g of acetic acid is added. Cyclohexyl ester, the reactor was closed, replaced with nitrogen three times, hydrogen was introduced to a certain pressure, and the temperature was gradually increased. At about 80 ° C, the pressure in the autoclave began to decrease, indicating that the catalyst in the autoclave began to reduce and began the ester hydrogenation reaction.
  • the hydrogen is supplied in a timely manner to maintain a certain pressure of the reaction vessel, and finally the temperature is raised to a given temperature, and after maintaining the pressure reaction at this temperature for a certain period of time, the reaction is stopped, and after cooling to room temperature, the reaction product and the catalyst are discharged.
  • the product composition was analyzed by gas chromatography, and the one-pass conversion of cyclohexyl acetate and the one-way selectivity of cyclohexanol were calculated according to the analysis results according to the following formula.
  • Cyclohexanol single pass selectivity [mole of cyclohexanol / (moles of cyclohexanol + moles of cyclohexane + moles of ethylcyclohexyl ether)] ⁇ 100%
  • Example 16 ester hydrogenation in a fixed bed
  • the catalyst powder obtained in Example 3 was tableted, and the sieved 40-60 mesh particles were crushed, and 40 g of the catalyst particles were placed in the middle of a (x) 20 x 2.5 x 800 mm jacketed stainless steel tube reactor. Fill both ends with a certain amount of quartz sand. After passing hydrogen (500 mL/min) at 280 °C and 6 MPa for 24 h, it was lowered to the reaction temperature. The cyclohexyl acetate is pumped into the reactor by a metering pump, and the hydrogen gas enters the reaction system through the mass flow controller for hydrogenation reaction. The heat transfer oil is introduced into the outer jacket of the reaction tube to control the reaction temperature, and the back pressure valve is passed through the reactor outlet. Control reactor pressure.
  • the reaction product was sampled by on-line chromatography on the in-line sampling valve at the rear of the reactor.
  • the reaction conditions and results are shown in Table 3.
  • the results show that the single pass conversion of cyclohexyl acetate hydrogenation can reach more than 99.5%, and the single-pass selectivity of ester products is more than 99.0%. After 1000 hours of operation, the single-pass conversion rate and single-pass selectivity are not obvious.
  • the catalyst was prepared as in Example 1 but was not treated with NaOH solution and was not subjected to the second drying and baking treatment.
  • the manufacturing conditions and the composition of the catalyst prepared are shown in Comparative Example 1 in Table 1.
  • the catalyst prepared by the above method was subjected to evaluation of hydrogenation of cyclohexyl acetate using a high pressure reactor.
  • the evaluation conditions and results are shown in Comparative Example 2 in Table 2.
  • This example is intended to illustrate the results of preparing cyclohexyl acetate using a fixed bed reactor using cyclohexene and acetic acid as raw materials.
  • the fixed bed reactor is a (
  • Will be 500mL big hole strong Acidic hydrogen type ion exchange resin (the laboratory is synthesized according to the classical literature method, the styrene solution containing 15% divinylbenzene is suspended and copolymerized to form a white ball, and then sulfonated with concentrated sulfuric acid to obtain the exchange capacity. 5.2 mmol H+/g dry basis) was placed in the middle of the fixed bed reactor with a certain amount of quartz sand filled at both ends.
  • This example illustrates the results of hydrogenation of a mixture stream of acetic acid and cyclohexyl acetate.
  • the reaction system consisted of a single fixed bed reactor with a jacketed titanium steel tube measuring (
  • the catalyst was charged to the reactor in two layers.
  • the upper layer is charged with 20 g of silica-supported platinum palladium tin acetate hydrogenation catalyst (laboratory synthesis, composition of Pt(10%)-Pd(5%)-Sn(5%)/SiO 2 , from 20 to 40 Porous silica support (BET specific surface area 400 m 2 /g, pore volume 0.35 mL/g) impregnated with chloroplatinic acid, palladium chloride and stannous chloride Easy to mix, dried at 120 °C, calcined at 500 °C);
  • the lower layer is charged with 20g of copper aluminide hydrogenation catalyst (laboratory synthesis, composition is CuO 40%, ZnO 29.6 %, A1 2 0 3 30.4%)
  • the catalyst is charged into the central constant temperature zone of the reactor.
  • the two layers of catalyst are separated by a glass fiber cloth.
  • the reactor is filled with a certain amount of quartz sand as a raw material to heat the gasification zone or the filler.
  • a heat transfer oil can be introduced into the reactor jacket to control the reaction temperature.
  • a mixture of acetic acid and cyclohexyl acetate (reactor outlet stream of Example 1) is pumped into the reactor by a metering pump, and hydrogen is introduced into the reaction system through a mass flow controller for hydrogenation reaction, through the outer jacket of the reaction tube.
  • the heat transfer oil was introduced to control the reaction temperature, and the reactor pressure was controlled through a reactor outlet back pressure valve.
  • the reaction product was sampled by a linear sampling valve at the rear of the reactor for in-line color analysis. The reaction conditions and results are shown in Table 2.
  • This example is intended to illustrate the results of preparing cyclohexyl acetate using a reactive distillation column using cyclohexene and acetic acid as raw materials.
  • the test was carried out in a reactive distillation mode test apparatus of the following specifications:
  • the main body of the mode apparatus was a stainless steel tower having a diameter (inner diameter) of 50 mm and a height of 3 m, and a lower connecting body of the tower.
  • the 5L tower is equipped with a 10KW electric heating rod.
  • the heating rod is controlled by the intelligent controller through the thyristor (SCR) to control the heating capacity of the tower.
  • the top of the tower is connected to a condenser with a heat exchange area of 0.5 m 2 , and the overhead steam is condensed into a liquid through the condenser and then enters a 2 L reflux tank.
  • the liquid in the reflux tank is partially refluxed to the reaction column via a reflux pump and partially recovered as a light component.
  • the operating parameters of the tower are displayed and controlled by intelligent automated control instruments.
  • the tower return flow is controlled by a reflux regulator valve, and the overhead output is controlled by the level controller of the return tank.
  • the amount of tower kettle produced is controlled by the tower tank level controller to adjust the tower tank discharge valve.
  • the high temperature sulfonic acid type ion exchange resin (brand Amberlyst 45, manufactured by Rhom & Haas) is pulverized into a powder having a particle size of less than 200 mesh (0.074 mm) by a multi-stage high-speed pulverizer, and added to a pore former, a lubricant, and an antioxidant.
  • the agent and the binder are uniformly mixed on a high-speed mixer, and then kneaded at 180 ° C for 10 minutes on the internal mixer to completely plasticize the material, and then injected into the mold to make a diameter of 5 mm, a height of 5 mm, and a wall thickness of 1 mm.
  • Ring type resin catalyst filler brand Amberlyst 45, manufactured by Rhom & Haas
  • This example illustrates the results of hydrogenation of a mixture stream of acetic acid and cyclohexyl acetate.
  • the reaction system consisted of a single fixed bed reactor with a jacketed titanium steel tube measuring (
  • the catalyst was charged to the reactor in two layers.
  • the upper layer was charged with 20 g of silica-supported platinum palladium tin acetate hydrogenation catalyst (laboratory synthesis, composition of Pt(10m%)-Pd(5m%)-Sn(5m%)/Si0 2 , from 20 ⁇ 40 Porous silica support (BET specific surface area 400 m 2 /g, pore volume 0.35 mL / g) impregnated mixed solution of chloroplatinic acid, palladium chloride and stannous chloride, dried at 120 ° C, 500 ⁇ roasting
  • the lower layer is charged with 20g copper zinc aluminum ester hydrogenation catalyst (laboratory synthesis, composition is CuO 40m%, ZnO 29.6m %, ⁇ 1 2 0 3 30.4m%.
  • a mixture of acetic acid and cyclohexyl acetate (the bottoms stream of Example 1) is pumped into the reactor by a metering pump, and hydrogen is introduced into the reaction system through the mass flow controller for hydrogenation reaction, and is passed through the outer jacket of the reaction tube.
  • the heat transfer oil controls the reaction temperature and the reactor pressure is controlled by a reactor outlet back pressure valve. The conditions and results are shown in the test data of the hydrogenation of the mixture of acetic acid and cyclohexyl acetate.
  • Examples 1 to 4 are for explaining a method for producing cyclohexyl acetate using a reaction fine.
  • the tests carried out in Examples 1 to 4 were carried out in a reaction-precision mode test apparatus of the following specifications:
  • the main body of the mode apparatus was a stainless steel tower having a diameter (inner diameter) of 50 mm and a height of 3 m, and the lower portion of the tower was connected.
  • the column is 5L in volume, and the kettle is equipped with a 10KW electric heating rod.
  • the heating rod is controlled by the intelligent controller through the thyristor (SCR) to control the heating amount of the tower.
  • the top of the tower is connected to a condenser with a heat exchange area of 0.5 m 2 , and the overhead steam is condensed into a liquid through the condenser and then enters a 2 L reflux tank.
  • the liquid in the reflux tank is partially refluxed to the reaction column via a reflux pump and partially recovered as a light component.
  • the operating parameters of the tower are displayed and controlled by intelligent automated control instruments.
  • the tower return flow is controlled by a reflux regulator valve, and the overhead output is controlled by the level controller of the return tank.
  • the amount of tower kettle produced is controlled by the tower tank level controller to adjust the tower tank discharge valve.
  • the raw material acetic acid and cyclohexene are respectively charged into a 30L storage tank, and are pumped into a corresponding preheater through a metering pump to preheat to a certain temperature and then enter the reaction tower.
  • the feed rate is controlled by the metering pump.
  • the electronic scale is accurately metered.
  • the high temperature sulfonic acid type ion exchange resin (brand Amberlyst 45, manufactured by Rhom & Haas) is pulverized into a powder having a particle size of less than 200 mesh (0.074 mm) by a multi-stage high-speed pulverizer, and added to a pore former, a lubricant, and an antioxidant.
  • the agent and the binder are uniformly mixed on a high-speed mixer, and then kneaded at 180 ° C for 10 minutes on the internal mixer to completely plasticize the material, and then injected into the mold to make a diameter of 5 mm, a height of 5 mm, and a wall thickness of 1 mm.
  • Ring type resin catalyst filler brand Amberlyst 45, manufactured by Rhom & Haas
  • the configuration of the reaction column and the catalyst was the same as in Example 1.
  • the test was carried out by simply replacing cyclohexene with a mixture of cyclohexene, cyclohexane and benzene, and the reaction column was operated at 0.3 MPa. The heating amount of the column and the reflux flow at the top of the column were continuously reacted.
  • the reaction conditions and reaction results under stable operation are shown in Table 2.
  • the spherical shape of ⁇ 3 ⁇ 4 is H 0 . 5 Cs 2 . 5 PW 12 O 4 . /SiO 2 catalyst (from H 0 5 Cs 2 5 PW 12 O 40 powder and coarse pore silica gel powder with a particle size of less than 200 mesh, after thoroughly mixing in a blender, using a silica sol as a bonding machine for ball forming in a sugar-coating machine , after drying and roasting, is sandwiched into a titanium mesh wave plate to form a cylindrical structured packing having a diameter of 50 mm and a height of 50 mm.
  • the packed catalyst L was charged into the middle of the mode reactor (high lm, equivalent to 12 theoretical plates).
  • Example 3 The configuration of the reaction column and the catalyst was the same as in Example 3. The test was carried out by simply replacing the cyclohexene with a mixture of cyclohexene, cyclohexane and benzene, and the reaction column was operated at 0.2 MPa. The heating of the tower and the reflux of the column were continuously carried out, and the reaction conditions and reaction results under stable operation are shown in Table 4. Examples 5 to 8 are used to illustrate a method for producing cyclohexyl acetate using pre-esterification and reactive distillation.
  • the tests carried out in Examples 5 to 8 were carried out in a cyclohexyl acetate mode test apparatus.
  • the mode unit consists of a fixed bed preesterification reactor and a reactive distillation esterification column.
  • the pre-esterification reactor is a (
  • the reactive distillation esterification column is a titanium steel (TA2) column having a diameter (inner diameter) of 50 mm and a height of 3 m.
  • the lower part of the tower is connected to a 5L tower.
  • the kettle is equipped with a 10KW electric heating rod.
  • the heating rod is controlled by an intelligent controller through a thyristor (SCR) to control the heating capacity of the tower.
  • the top of the tower is connected to a condenser having a heat exchange area of 0.5 m 2 and a reflux tank having a volume of 2 L.
  • the raw materials acetic acid and cyclohexene are separately charged into a 30 L storage tank, and are fed into a pre-esterification reactor through a metering pump to carry out a reaction, and the pre-esterified product enters the reaction rectification column to further carry out a reaction.
  • the amount of heating of the reaction column is adjusted by adjusting the heating power of the column.
  • the reflux ratio of the column is adjusted by an overhead reflux ratio regulator. Light components are taken from the top of the tower. The cyclohexyl acetate product was recovered from the bottom of the column.
  • 500mL macroporous strong acid hydrogen type ion exchange resin (laboratory according to the classical literature method, suspension polymerization of styrene solution containing 15% divinylbenzene to make white ball, then sulfonated with concentrated sulfuric acid, measured It has a exchange capacity of 5.2 mmolH+/g dry basis. It is charged into the middle of the pre-reactor and filled with a certain amount of quartz sand at both ends.
  • a high-temperature sulfonic acid type ion exchange resin brand Amberlyst 45, manufactured by Rhom & Haas
  • a pore-forming agent, a lubricant, and an anti-wear agent are added.
  • the oxygen agent and the binder are mixed and hooked on a high-speed mixer, and then kneaded at 180 ° C for 10 minutes on the internal mixer to completely plasticize the material, and then injected into the mold to make a diameter of 5 mm, a height of 5 mm, and a wall thickness of 1 mm.
  • Raschig ring type resin catalyst filler 1950mL of this packing was placed in the middle of the mode reaction column (high lm, equivalent to 8 theoretical plates). The top and bottom were filled with 1950mL glass spring packing with a diameter of 3mm and a length of 6mm (filling height is lm, equivalent to 10 theoretical towers) board).
  • the cyclohexene and acetic acid are respectively fed into the pre-reactor by a metering pump, and the pre-reaction product is further reacted by entering the reaction column.
  • the pre-reaction temperature is adjusted by adjusting the pre-reactor jacket hot water temperature.
  • the reaction of the heating of the tower and the reflux of the column were continuously carried out, and the reaction conditions and reaction results of the stable operating conditions are shown in Table 5.
  • the configuration of the reaction column and the catalyst was the same as in Example 5.
  • cyclohexene cyclohexane
  • the test was carried out by replacing the cyclohexene with a mixture of benzene, and the pre-reactor pressure was 2.0 MPa, and the reaction column was operated under normal pressure conditions.
  • the reaction of the heating of the tower and the reflux of the column were continuously carried out, and the reaction conditions and reaction results of the stable operating conditions are shown in Table 6.
  • the cyclohexene and acetic acid are respectively driven into the pre-reactor by a metering pump, and the pre-reaction product is further reacted in the reaction column.
  • the pre-reaction temperature is adjusted by adjusting the temperature of the pre-reactor jacket hot water. The heating of the tower kettle and the reflux flow of the column were continuously carried out, and the reaction conditions and reaction results of the stable operating conditions are shown in Table 7.
  • the configuration of the reaction column and the catalyst was the same as in Example 7.
  • the test was carried out by replacing the cyclohexene with a mixture of cyclohexene, cyclohexane and benzene, and the pre-reaction pressure 2.0 MPa reaction column was operated at 0.2 MPa.
  • the heating of the tower and the reflux of the column were continuously carried out, and the reaction conditions and reaction results of the stable operating conditions are shown in Table 8.
  • Examples 9 to 10 are used to illustrate a method of hydrogenating cyclohexyl acetate.
  • the hydrogenation feedstock was cyclohexyl acetate having a purity of 99.6%.
  • the heat transfer oil is introduced into the outer jacket of the reaction tube to control the reaction temperature, and the back pressure valve is passed through the reactor outlet.
  • Line chromatography analysis The reaction conditions and results are shown in Table 9.
  • the results in Table 9 show that with the copper-zinc-aluminum catalyst, the single-pass conversion rate of cyclohexyl acetate hydrogenation can reach more than 99.0%, the cyclohexanol single-pass selectivity is greater than 99.9 %, the operation is 1000 hours, and the single-pass conversion rate and single-pass selectivity are both Not falling.
  • the hydrogenation feedstock was cyclohexyl acetate having a purity of 99.6%.
  • the heat transfer oil is introduced into the outer jacket of the reaction tube to control the reaction temperature, and the back pressure valve is passed through the reactor outlet. Control reactor pressure.
  • the reaction product was sampled by in-line chromatography through a linear sampling valve at the rear of the reactor.
  • the reaction conditions and results are shown in Table 10.
  • the results in Table 10 show that with the copper-zinc-aluminum catalyst, the single-pass conversion rate of cyclohexyl acetate hydrogenation reaction can reach more than 98.0%, the single-pass selectivity of cyclohexanol is more than 99.9%, and the operation rate is not decreased after one-hour conversion. .
  • the single pass conversion of cyclohexene was 98.8%
  • the single pass selectivity of cyclohexyl acetate was 98.0%.
  • the single pass conversion of cyclohexene was calculated to be 98.7%, and the single pass selectivity of cyclohexyl acetate was 99.43%.
  • the single pass conversion of cyclohexene was calculated to be 99.35%, and the single pass selectivity of cyclohexyl acetate was 99.6%.
  • composition (mass score)
  • the single pass conversion of cyclohexene was calculated to be 98.38%, and the single pass selectivity of cyclohexyl acetate was 99.11%.
  • the single pass conversion of cyclohexene was calculated to be 99.9%, and the single pass selectivity of cyclohexyl acetate was 99.35%.
  • the single pass conversion of cyclohexene was calculated to be 99.02%, and the single pass selectivity of cyclohexyl acetate was 99.19%.
  • This example is intended to illustrate the test results of catalytic dehydrogenation of cyclohexane to benzene.
  • the raw material for the catalytic dehydrogenation reaction is a C6 hydrocarbon mixture obtained by distillation of the overhead stream of Example 4.
  • the mixture is washed with water to remove a small amount of acetic acid and analyzed by gas phase colorimetry, containing 67.5 m% of cyclohexane, 32.3 m% of benzene and 0.2 m% of cyclohexene.
  • the reactor was a tubular fixed bed reactor, and the reactor was a jacketed titanium steel pipe having a size of ⁇ 20 ⁇ 2.5 ⁇ 800 mm.
  • the catalyst uses a supported platinum rhodium catalyst (Pt content) 0.3m%, Rh content 0.1m%).
  • the reaction conditions are: temperature 480 ° C, pressure 0.7 MPa, weight hourly space velocity 5 h-1.
  • cyclohexane in the reaction starting material was quantitatively converted to benzene.
  • Example 2 Test methods and apparatus employed in Example 1 were esterified acid with cyclohexene test, except that is Cs 2. 5 H 0. 5 PW 12 O 40 / SiO 2 catalyst (referred to as a PW / Si0 2, The same below). The reaction conditions and results are shown in Table 2. It can be seen from Table 2 that the conversion per pass of cyclohexene and acetic acid can reach 95%, the single-pass selectivity of ester product is more than 99%, and the activity of the catalyst and the selectivity of one-pass are stable.
  • the test apparatus and method are the same as those in the first embodiment, except that the catalyst is a phosphorus-modified ⁇ molecular sieve catalyst (the ⁇ molecular sieve having a silica-alumina ratio of 50 is modified by 85% phosphoric acid, and then kneaded with alumina to form a strip. Dry at 120 ° C, calcined at 500 ° C, the phosphorus content is 2%).
  • the reaction conditions and results are shown in Table 3. It can be seen from Table 3 that the conversion of cyclohexanyl with acetic acid is 90% per pass, the selectivity of the ester product is more than 99%, and the activity of the catalyst and the selectivity of the single pass are stable.
  • the addition esterification products of Examples 1 to 3 were collected and subjected to a rectification separation test.
  • the rectification adopts a glass tower rectification unit with a height of 2 m and a diameter of 40 mm.
  • the column column is equipped with a D3mm stainless steel crucible ring high-efficiency rectification packing, and the column is a volume 5 L glass flask with a charge of 4L.
  • the electric heating jacket heats the tower kettle, and the heating capacity of the tower kettle is adjusted by a pressure regulator.
  • the reflux of the column is controlled by a reflux ratio regulator.
  • Table 4 The results of distillation separation are shown in Table 4.
  • the test apparatus, catalyst and method were the same as in Example 1, except for the cyclohexene starting material (benzene 53.3%, cyclohexene 35.4%, cyclohexane 1 1.3%).
  • the reaction conditions and results are shown in Table 5. It can be seen from Table 5 that the strong acid type ion exchange resin catalyst catalyzes the reaction of the cyclohexene raw material with acetic acid, the single pass conversion of cyclohexene is more than 80%, the single pass selectivity of the ester product is more than 99%, the operation is 600 hours, the catalyst activity and the single pass selectivity. Stable.
  • the hydrogenation feedstock was a cyclohexyl acetate having a purity of 99.6%.
  • the cyclohexyl acetate is pumped into the reactor by a metering pump, and the hydrogen gas enters the reaction system through the mass flow controller for hydrogenation reaction.
  • the heat transfer oil is introduced into the outer jacket of the reaction tube to control the reaction temperature, and the back pressure valve is passed through the reactor outlet. Line chromatography analysis.
  • the reaction conditions and results are shown in Table 6. The results in Table 6 show that with the copper-aluminum catalyst, the single-pass conversion of cyclohexyl acetate hydrogenation can reach more than 99%, the cyclohexanol single-pass selectivity is greater than 99.9%, the operation is 1000 hours, and the single-pass conversion and single-pass selectivity are both Not falling.
  • the hydrogenation feedstock was a cyclohexyl acetate having a purity of 99.6%.
  • the system performs a hydrogenation reaction, and the reaction temperature is controlled by introducing a heat transfer oil through the outer jacket of the reaction tube, and passing through the outlet of the reactor The pressure valve controls the reactor pressure.
  • the reaction product was sampled by in-line chromatography through a linear sampling valve at the rear of the reactor.
  • the reaction conditions and results are shown in Table 7.
  • the results in Table 7 show that with the copper-zinc-aluminum catalyst, the single-pass conversion rate of cyclohexyl acetate hydrogenation reaction can reach over 98%, the single-pass selectivity of cyclohexanol is more than 99.9%, and the single-pass conversion rate and selection are not decreased after 500 hours of operation.
  • Example 6-7 The reaction product of Example 6-7 was collected in 4000 g, and a rectification separation test was carried out.
  • the high-quality 2m glass tower is used.
  • the column is equipped with a stainless steel ⁇ mesh ring with high efficiency rectification packing of 0>3mm.
  • the tower is a 5L glass flask, which is heated by an electric heating sleeve, and the heating capacity of the tower is adjusted by a pressure regulator.
  • the reflux of the column is controlled by a reflux ratio regulator.
  • Table 8 Strong acid ion exchange resin catalyzes the test data of acetic acid and cyclohexene esterification
  • Copper-zinc-aluminum catalyst catalyzed hydrogenation of cyclohexyl acetate
  • This example illustrates a method for the selective hydrogenation of benzene to cyclohexene.
  • the benzene and hydrogen molar ratio of 1:3 was injected into the hydrogenation reactor packed with the ruthenium particle catalyst, and the benzene hydrogenation reaction was carried out under the conditions of a reaction temperature of 135 ° C, a pressure of 4.5 MPa, and a residence time of 15 min, and the reaction product separated hydrogen. After that, the liquid product was collected and run continuously for 1000 h. After the end of the test, the collected liquid product was analyzed by gas chromatography, and its composition was: benzene 53.3%. Cyclohexene 35.4%, cyclohexane 11.3%.
  • the acetic acid and cyclohexene raw materials (benzene 53.3%, cyclohexene 35.4%, cyclohexane 1 1.3%) were respectively fed into the reactor by a metering pump at a certain flow rate, and the reaction was carried out in the outer jacket of the reaction tube. Hot water is used to control the reaction temperature and the reactor pressure is controlled by a reactor outlet back pressure valve. The reactor outlet product was sampled by an on-line sampling valve for in-line color analysis, and the cyclohexene single pass conversion and cyclohexyl acetate one-pass selectivity were calculated from the product composition. The reaction conditions and results are shown in Table 1.
  • the strong acid type ion exchange resin catalyst catalyzes the reaction of the cyclohexene raw material with acetic acid
  • the single pass conversion of cyclohexene is more than 80%
  • the single pass selectivity of the ester product is more than 99%
  • the operation is 600 hours
  • the test apparatus, method and raw materials are the same as those in Example 2.
  • the catalyst is a phosphorus-modified ⁇ molecular sieve catalyst (the ⁇ molecular sieve with a silica-alumina ratio of 50 is modified by 85% phosphoric acid, and then kneaded with alumina to form a strip. It is dried at 120 ° C and calcined at 500 ° C with a phosphorus content of 2%).
  • the reaction conditions and results are shown in Table 2. It can be seen from Table 2 that the conversion per pass of cyclohexene and acetic acid is more than 80%, the single-pass selectivity of the ester product is more than 99%, and the activity of the catalyst and the selectivity of one-pass are stable.
  • Table 2 ⁇ molecular sieve catalyst catalyzed acetic acid and cyclohexane / cyclohexene / phenyl esterification test data
  • This example illustrates a method of simultaneous hydrogenation of an esterification reaction mixture.
  • the hydrogenation feedstock is a mixture of cyclohexene and acetonitrile containing benzene and cyclohexane (7.4% cyclohexane, 35.4% benzene, 4.6% cyclohexene, 20.5% acetic acid, 32.2% cyclohexyl acetate, polymerization). 0.2%).
  • the reaction system consisted of two fixed bed reactors connected in series. Both reactors are jacketed titanium steel tubes measuring (
  • the former reactor is a hydrogenation reactor of acetic acid, benzene and cyclohexene, and contains 40 g of a silica-supported platinum palladium tin acetate hydrogenation catalyst (synthesized in a chamber, the composition is Pt (10 m%)-Pd (5 m%) - Sn(5m%)/SiO 2 , impregnated with chloroplatinic acid, palladium chloride and stannous chloride by a 20 ⁇ 40 mesh macroporous silica support (BET specific surface area 400 m 2 /g, pore volume 0.35 mL/g) The solution is mixed and dried at 120 ° C and calcined at 500 ° C).
  • the reaction system After the catalyst is installed in the reactor, the reaction system is connected, and after the system is airtightly tested, hydrogen (100 mL/min) is passed through the mixture at 300 ° C and 6 MPa for 24 hours, and then the temperature and pressure of the hydrogenation reaction are lowered.
  • the mixture of acetic acid and cyclohexyl acetate is pumped into the reactor by a metering pump, and the hydrogen gas enters the reaction system through the mass flow controller for hydrogenation reaction, and the reaction temperature is controlled by passing the heat transfer oil through the outer jacket of the reaction tube, and the reactor is passed through the reactor.
  • the outlet back pressure valve controls the reactor pressure.
  • the reaction product was subjected to on-line chromatography by linear sampling at the rear of the reactor. The reaction conditions and results are shown in Table 3.
  • This example illustrates a method of simultaneous hydrogenation of an esterification reaction mixture.
  • the hydrogenation feedstock is a mixture of cyclohexene and acetonitrile containing benzene and cyclohexane, which constitutes 7.4% cyclohexane, 35.4% benzene, 4.6% cyclohexene, 20.5% acetic acid, cyclohexyl acetate. 32.2%, polymer 0.2%).
  • the reaction system consisted of a single fixed bed reactor, which was a jacketed titanium steel tube measuring (J) 20 x 2.5 x 800 mm.
  • the catalyst was charged to the reactor in two layers.
  • the upper layer was charged with 20 g of silica supported platinum palladium tin acetate hydrogenation catalyst (laboratory synthesis, composition of Pt(10%)-Pd(5%)-Sn(5%)/SiO 2 , from 20 to 40 Porous silica carrier (BET specific surface area 400 m 2 /g, pore volume 0.35 mL/g) impregnated with chloroplatinic acid, palladium chloride and stannous chloride, easily mixed, dried at 120 ° C, baked at 500 ° C
  • the lower layer is charged with 20g of copper zinc aluminum ester hydrogenation catalyst (laboratory synthesis, composition is CuO 40.5%, ZnO 29.6 %, A1 2 0 3 30.4%.
  • the catalyst is charged into the central constant temperature zone of the reactor.
  • the two layers of catalyst are separated by a glass fiber cloth.
  • the reactor is filled with a certain amount of quartz sand as a raw material to heat the gasification zone or the filler.
  • a heat transfer oil can be introduced into the reactor jacket to control the reaction temperature.
  • the mixture of acetic acid and cyclohexyl acetate is pumped into the reactor by a metering pump, and the hydrogen gas enters the reaction system through the mass flow controller for hydrogenation reaction, and the reaction temperature is controlled by passing the heat transfer oil through the outer jacket of the reaction tube, and the reactor is passed through the reactor.
  • the outlet back pressure valve controls the reactor pressure.
  • the reaction product was sampled by in-line chromatography through a linear sampling valve at the rear of the reactor. The reaction conditions and results are shown in Table 4.
  • the reaction product of Examples 5 to 6 was collected in 4000 g (gas chromatography analysis composition: ethanol 27.4 m%, ethyl acetate 0.2 m%, cyclohexane 40.2 m%, water 6.4 m%, acetic acid 0.2 m%, cyclohexanol 25.0 m %, cyclohexyl acetate 0.3 m%, other 0.3 m%), subjected to fine separation test.
  • the rectification adopts a high 2m glass tower, the column is equipped with a stainless steel crucible ring of ⁇ 3 ⁇ , and the column is a 5L glass flask, which is heated by an electric heating sleeve, and the heating amount of the tower kettle is adjusted by a pressure regulator.
  • the reflux of the column is controlled by a reflux ratio regulator. After separation, 845 g of cyclohexanol product was obtained, and the purity by gas chromatography was 99.4%.
  • Example 1 This example illustrates a method for the selective hydrogenation of benzene to cyclohexene.
  • the benzene and hydrogen molar ratio of 1:3 was injected into the hydrogenation reactor packed with the ruthenium particle catalyst, and the benzene hydrogenation reaction was carried out under the conditions of a reaction temperature of 135 ° C, a pressure of 4.5 MPa, and a residence time of 15 min, and the reaction product separated hydrogen. After that, the liquid product was collected and run continuously for 1000 h. After the end of the test, the collected liquid product was subjected to gas chromatography analysis, and its composition was: benzene 53.3%, cyclohexene 35.4%, and cyclohexane 11.3%.
  • the acetic acid and cyclohexene raw materials (benzene 53.3%, cyclohexene 35.4%, cyclohexane 1 1.3%) were respectively fed into the reactor by a metering pump at a certain flow rate, and the reaction was carried out in the outer jacket of the reaction tube. Hot water is used to control the reaction temperature and the reactor pressure is controlled by a reactor outlet back pressure valve. The reactor outlet product was sampled by an on-line sampling valve for on-line chromatographic analysis to calculate the cyclohexene single pass conversion and cyclohexyl acetate one-pass selectivity from the product composition. The reaction conditions and results are shown in Table 1.
  • the strong acid type ion exchange resin catalyst catalyzes the reaction of the cyclohexene raw material with acetic acid
  • the single pass conversion of cyclohexene is more than 80%
  • the single pass selectivity of the ester product is more than 99%
  • the operation is 600 hours
  • the test apparatus, method and raw materials are the same as those in Example 2.
  • the catalyst is a phosphorus-modified ⁇ molecular sieve catalyst (the ⁇ molecular sieve with a silica-alumina ratio of 50 is modified by 85% phosphoric acid, and then kneaded with alumina to form a strip. , dried at 120 ° C, calcined at 500 ° C, the phosphorus content is 2m%).
  • the reaction conditions and results are shown in Table 2. It can be seen from Table 2 that the single pass conversion of cyclohexene and acetic acid is greater than 80%, the single pass selectivity of the ester product is greater than 99%, and the operation is 480 hours. The activity of the agent and the one-way selectivity were stable.
  • Table 2 ⁇ molecular sieve catalyst catalyzed the test data of acetic acid and cyclohexane/cyclohexene/phenyl esterification
  • the addition esterification products of Examples 2 and 3 were collected and subjected to a rectification separation test.
  • the rectification adopts a glass tower rectification unit with a height of 2 m and a diameter of 40 mm.
  • the column column is equipped with a 3 mm stainless steel crucible ring high-efficiency rectification packing, and the tower is a volume 5 L glass flask with a charge of 4 L.
  • the tower kettle is heated, and the amount of heating of the tower kettle is adjusted by a pressure regulator.
  • the reflux of the column is controlled by a reflux ratio regulator.
  • the results of distillation separation are shown in Table 3. Addition esterification product rectification separation test results
  • Examples 5 to 6 are used to illustrate a method for producing cyclohexyl acetate by reactive distillation.
  • the tests carried out in Examples 5 to 6 were carried out in a reactive distillation mode test apparatus of the following specifications:
  • the main body of the mode apparatus was a stainless steel tower having a diameter (inner diameter) of 50 mm and a height of 3 m, and the lower connection volume of the tower was
  • the 5L tower is equipped with a 10KW electric heating rod.
  • the heating rod is controlled by the intelligent controller through the thyristor (SCR) to control the heating capacity of the tower.
  • the top of the tower is connected to a condenser with a heat exchange area of 0.5 m 2 , and the overhead steam is condensed into a liquid through the condenser and then enters a 2 L reflux tank.
  • the liquid in the reflux tank is partially refluxed to the reaction column via a reflux pump and partially recovered as a light component.
  • the operating parameters of the tower are displayed and controlled by intelligent automated control instruments.
  • the tower return flow is controlled by a reflux regulator valve, and the overhead output is controlled by the level controller of the return tank.
  • the amount of tower kettle produced is controlled by the tower tank level controller to adjust the tower tank discharge valve.
  • the acetic acid and cyclohexene raw materials are respectively charged into a 30L storage tank, and are pumped into a corresponding preheater through a metering pump to be preheated to a certain temperature and then enter the reaction tower.
  • the feed rate is controlled by the metering pump and accurately measured by the electronic scale.
  • the high temperature sulfonic acid type ion exchange resin (brand Amberlyst 45, manufactured by Rhom & Haas) is pulverized into a powder having a particle size of less than 200 mesh (0.074 mm) by a multi-stage high-speed pulverizer, and added to a pore former, a lubricant, and an antioxidant.
  • the agent and the binder are uniformly mixed on a high-speed mixer, and then kneaded at 180 ° C for 10 minutes on the internal mixer to completely plasticize the material, and then injected into the mold to make a diameter of 5 mm, a height of 5 mm, and a wall thickness of 1 mm.
  • Ring type resin catalyst filler brand Amberlyst 45, manufactured by Rhom & Haas
  • the single pass conversion of cyclohexene was calculated to be 99%, and the single pass selectivity of cyclohexyl acetate was 99.2%.
  • the ball H ⁇ 3 ⁇ 4 to 0. 5 Cs 2. 5 PW 12 O 40 / SiO 2 catalyst manufactured by 12 O 40 powder and the particle size of H 0. 5 Cs 2. 5 PW less than 200 mesh coarse pore silica powder, in the mix
  • the silica sol is used as a bonding machine in a sugar-coated machine, and then baked and baked to be sandwiched into a titanium mesh wave plate to form a cylinder having a diameter of 50 mm and a height of 50 mm.
  • Type structured packing The packed catalyst L was charged into the middle of the mode reactor (high lm, equivalent to 12 theoretical plates).
  • the single pass conversion of cyclohexene was calculated to be 98.3 %, and the single pass selectivity of cyclohexyl acetate was 99.5 %.
  • This example illustrates a method of hydrogenating a mixture of acetic acid and cyclohexyl acetate.
  • the hydrogenation feedstock is a mixture of acetic acid and cyclohexyl acetate (acetic acid 39.5%, cyclohexyl acetate)
  • the reaction system consisted of two fixed bed reactors connected in series. Both reactors are jacketed titanium steel tubes measuring ⁇ 1) 20 x 2.5 x 800 mm.
  • the former reactor is a acetic acid hydrogenation reactor containing 40 g of silica-supported platinum palladium tin acetate hydrogenation catalyst (laboratory synthesis, composition Pt (10 m%) - Pd (5 m%) - Sn (5 m%) / SiO 2 , a 20 ⁇ 40 mesh macroporous silica carrier (BET specific surface area 400m 2 /g, pore volume 0.35mL / g) impregnated mixed solution of chloroplatinic acid, palladium chloride and stannous chloride, and then passed through 120 Dry at °C, calcined at 500 °C).
  • the reaction system After the catalyst is installed in the reactor, the reaction system is connected, and after the system is airtightly tested, the temperature and pressure of the hydrogenation reaction are reduced after hydrogenation (100 mL/min) is passed at 300 ° (:, 6 MPa for 24 h).
  • the mixture of acetic acid and cyclohexyl acetate is pumped into the reactor by a metering pump, and the hydrogen gas enters the reaction system through the mass flow controller for hydrogenation reaction, and the reaction temperature is controlled by introducing a heat transfer oil through the outer jacket of the reaction tube.
  • the reactor outlet back pressure valve controls the reactor pressure.
  • the reaction product is sampled by a linear sampling valve at the rear of the reactor for online chromatographic analysis. The reaction conditions and results are shown in Table 6.
  • This example illustrates a method of hydrogenating a mixture of acetic acid and cyclohexyl acetate.
  • the hydrogenation feedstock was a mixture of acetic acid and cyclohexyl acetate (acetic acid 39.5 %, cyclohexyl acetate 60.5%).
  • the reaction system consisted of a single fixed bed reactor with a jacketed titanium steel tube measuring ()) 20 x 2.5 x 800 mm.
  • the catalyst was charged to the reactor in two layers.
  • the upper layer is charged with 20 g of silica-supported platinum palladium tin acetate hydrogenation catalyst (laboratory synthesis, composition of Pt(10%)-Pd(5%)-Sn(5%)/SiO 2 , from 20 to 40 Porous silica carrier (BET Specific surface area 400m 2 /g, pore volume 0.35mL / g) impregnation of chloroplatinic acid, palladium chloride and stannous chloride is easy to mix, then dried at 120 ° C, 500 ° C roasting); the lower layer is loaded with 20g Copper-aluminum ester hydrogenation catalyst (laboratory synthesis, composition of CuO 40%, ZnO 29.6 %, A1 2 0 3 30.4%.
  • the catalyst is charged into the central constant temperature zone of the reactor.
  • the two layers of catalyst are separated by a glass fiber cloth.
  • the reactor is filled with a certain amount of quartz sand as a raw material to heat the gasification zone or the filler.
  • a heat transfer oil can be introduced into the reactor jacket to control the reaction temperature.
  • This example illustrates a method for the selective hydrogenation of benzene to cyclohexene.
  • the benzene and hydrogen molar ratio of 1:3 was injected into the hydrogenation reactor packed with the ruthenium particle catalyst, and the benzene hydrogenation reaction was carried out under the conditions of a reaction temperature of 135 ° C, a pressure of 4.5 MPa, and a residence time of 15 min, and the reaction product separated hydrogen. After that, the liquid product was collected and run continuously for 1000 h. After the end of the test, the collected liquid product was subjected to gas chromatography analysis, and its composition was: benzene 53.3%, cyclohexene 35.4%, and cyclohexane 1 1.3%.
  • the acetic acid and cyclohexene raw materials (benzene 53.3%, cyclohexene 35.4%, cyclohexane 1 1.3%) were respectively fed into the reactor by a metering pump at a certain flow rate, and the reaction was carried out in the outer jacket of the reaction tube. Hot water is used to control the reaction temperature and the reactor pressure is controlled by a reactor outlet back pressure valve. The reactor outlet product was sampled by an on-line sampling valve for on-line chromatographic analysis, and the cyclohexene single pass conversion and cyclohexyl acetate one-pass selectivity were calculated from the product composition. The reaction conditions and results are shown in Table 1.
  • the strong acid type ion exchange resin catalyst catalyzes the reaction of the cyclohexene raw material with acetic acid
  • the single pass conversion of cyclohexene is more than 80%
  • the single pass selectivity of the ester product is more than 99%
  • the operation is 600 hours
  • Stable. Strong Acidic Ion Exchange Resin Catalyzed Esterification of Acetic Acid with Cyclohexane/Cyclohexyl/Phenyl Ester
  • the test apparatus, method and raw materials are the same as those in Example 2.
  • the catalyst is a phosphorus-modified ⁇ molecular sieve catalyst (the ⁇ molecular sieve with a silica-alumina ratio of 50 is modified by 85% phosphoric acid, and then kneaded with alumina. Molding, drying at 120 ° C, calcination at 500 ° C, phosphorus content of 2%).
  • the reaction conditions and results are shown in Table 2. It can be seen from Table 2 that the single pass conversion of cyclohexene with acetic acid is greater than 80%, the single pass selectivity of the ester product is greater than 99%, the operation is 480 hours, the catalyst activity and The single pass selectivity is stable.
  • Table 2 ⁇ molecular sieve catalyst catalyzed test data of acetic acid and cyclohexane / cyclohexene / phenyl esterification
  • the addition esterification products of Examples 2 and 3 were collected and subjected to a rectification separation test.
  • the rectification adopts a glass tower rectification unit with a height of 2 m and a diameter of 40 mm.
  • the column column is equipped with a ⁇ 3 ⁇ stainless steel crucible ring high-efficiency rectification packing, and the tower is a volume of 5 L glass flask, and the charging amount is 4L, and the electric heating sleeve is passed.
  • the tower kettle is heated, and the amount of heating of the tower kettle is adjusted by a pressure regulator.
  • the reflux of the column is controlled by a reflux ratio regulator.
  • the results of distillation separation are shown in Table 3. Addition esterification product rectification separation test results
  • Examples 5 to 6 are for explaining a method for producing cyclohexyl acetate by reactive distillation.
  • the tests carried out in Examples 5 to 6 were carried out in a reactive distillation mode test apparatus of the following specifications:
  • the main body of the mode apparatus was a stainless steel tower having a diameter (inner diameter) of 50 mm and a height of 3 m, and the lower connection volume of the tower was
  • the 5L tower is equipped with a 10KW electric heating rod.
  • the heating rod is controlled by the intelligent controller through the thyristor (SCR) to control the heating capacity of the tower.
  • the top of the tower is connected to a condenser with a heat exchange area of 0.5 m 2 , and the overhead steam is condensed into a liquid through the condenser and then enters a 2 L reflux tank.
  • the liquid in the reflux tank is partially refluxed to the reaction column via a reflux pump and partially recovered as a light component.
  • the operating parameters of the tower are displayed and controlled by intelligent automated control instruments.
  • the tower return flow is controlled by a reflux regulator valve, and the overhead output is controlled by the level controller of the return tank.
  • the amount of tower kettle produced is controlled by the tower tank level controller to adjust the tower tank discharge valve.
  • the acetic acid and cyclohexene raw materials are respectively charged into a 30L storage tank, and are pumped into a corresponding preheater through a metering pump to be preheated to a certain temperature and then enter the reaction tower.
  • the feed rate is controlled by the metering pump and accurately measured by the electronic scale.
  • the high temperature sulfonic acid type ion exchange resin (brand Amberlyst 45, manufactured by Rhom & Haas Company) is pulverized into a powder having a particle size of less than 200 mesh (0.074 mm) by a multi-stage high-speed pulverizer, and added to a pore former, a lubricant, and an antioxidant.
  • the agent and the binder are uniformly mixed on a high-speed mixer, and then kneaded at 180 ° C for 10 minutes on the internal mixer to completely plasticize the material, and then injected into the mold to make a diameter of 5 mm, a height of 5 mm, and a wall thickness of 1 mm.
  • Ring type resin catalyst filler brand Amberlyst 45, manufactured by Rhom & Haas Company
  • the single pass conversion of cyclohexene was calculated to be 98.8 %, and the single pass selectivity of cyclohexyl acetate was 98.0 %.
  • the single pass conversion of cyclohexene was calculated to be 99.35 %, and the single pass selectivity of cyclohexyl acetate was 99.6 %.
  • Examples 7 to 8 are used to illustrate a method of hydrogenating cyclohexyl acetate.
  • the hydrogenation feedstock was a cyclohexyl acetate having a purity of 99.6%.
  • the cyclohexyl acetate is pumped into the reactor by a metering pump, and the hydrogen gas enters the reaction system through the mass flow controller for hydrogenation reaction.
  • the heat transfer oil is introduced into the outer jacket of the reaction tube to control the reaction temperature, and the back pressure valve is passed through the reactor outlet. Control reactor pressure.
  • the reaction product was subjected to online chromatographic analysis by linear sampling at the rear of the reactor. The reaction conditions and results are shown in Table 6.
  • Table 6 shows that copper-zinc-aluminum catalysis
  • the conversion rate of cyclohexyl acetate hydrogenation reaction can reach up to 99.0%
  • the single-pass selectivity of cyclohexanol is more than 99.9 %
  • the one-way conversion rate and single-pass selectivity are not decreased after 1000 hours of operation.
  • Table 6 Hydrogenation test data of cyclohexyl acetate of copper zinc aluminide hydrogenation catalyst
  • the hydrogenation feedstock was a cyclohexyl acetate having a purity of 99.6%.
  • the heat transfer oil is introduced into the outer jacket of the reaction tube to control the reaction temperature, and the back pressure valve is passed through the reactor outlet. Control reactor pressure.
  • the reaction product was sampled by in-line chromatography through a linear sampling valve at the rear of the reactor.
  • the reaction conditions and results are shown in Table 7.
  • the results in Table 7 show that with the copper-zinc-aluminum catalyst, the single-pass conversion rate of cyclohexyl acetate hydrogenation reaction can reach more than 98.0%, the single-pass selectivity of cyclohexanol is more than 99.9%, and the single-pass conversion rate and selection are not decreased. .
  • This example is intended to illustrate a method for the selective hydrogenation of benzene to cyclohexene.
  • the benzene and hydrogen molar ratio of 1:3 was injected into the hydrogenation reactor packed with the ruthenium particle catalyst, and the benzene hydrogenation reaction was carried out under the conditions of a reaction temperature of 135 ° C, a pressure of 4.5 MPa, and a residence time of 15 min, and the reaction product separated hydrogen. After that, the liquid product was collected and run continuously for 1000 h. After the test, the collected liquid product was subjected to gas chromatography analysis, and its composition was: benzene 53.3 m%, cyclohexene 35.4 m%, and cyclohexane 11.3 m%. The above liquid product is subjected to extraction and separation using N, N-dimethylacetamide as an extractant to obtain a mixture of cyclohexene and benzene.
  • the acetic acid and cyclohexene raw materials (obtained by the method of Example 1, the composition is: benzene 60m%, cyclohexene 40m%) are respectively driven into the reactor by a metering pump at a certain flow rate, and the reaction is jacketed outside the reaction tube.
  • the hot water is passed through to control the reaction temperature, and the reactor pressure is controlled through a reactor outlet back pressure valve.
  • the reactor outlet product was sampled by an on-line sampling valve for online chromatographic analysis.
  • the single-pass conversion of cyclohexene and the one-way selectivity of cyclohexyl acetate were calculated from the product composition. Selective.
  • the reaction conditions and results are shown in Table 1.
  • the test apparatus, method and raw materials are the same as those in Example 2.
  • the catalyst is a phosphorus-modified ⁇ molecular sieve catalyst (the ⁇ molecular sieve with a silica-alumina ratio of 50 is modified by 85% phosphoric acid, and then kneaded with alumina to form a strip. It is dried at 120 ° C and calcined at 500 ° C with a phosphorus content of 2%).
  • the reaction conditions and results are shown in Table 2. It can be seen from Table 2 that the single-pass conversion of cyclohexene is 80%, the selectivity of the ester product is more than 99%, and the activity of the catalyst and the selectivity of one-pass are stable.
  • the addition esterification products of Examples 2 and 3 were collected and subjected to a rectification separation test.
  • the rectification adopts a glass tower rectification unit with a height of 2 m and a diameter of 40 mm.
  • the column column is equipped with a D3mm stainless steel crucible ring and a high-efficiency fine crucible packing.
  • the tower is a volumetric L glass flask with a volume of 4L, and the electric heating sleeve is passed.
  • the tower kettle is heated, and the amount of heating of the tower kettle is adjusted by a pressure regulator.
  • the reflux of the column is controlled by a reflux ratio regulator.
  • Table 3 The results of distillation separation are shown in Table 3.
  • Examples 5 to 6 are for explaining a method for producing a cyclohexyl acetate by reactive distillation.
  • the tests carried out in Examples 5 to 6 were carried out in a reactive distillation mode test apparatus of the following specifications:
  • the main body of the mode apparatus was a stainless steel tower having a diameter (inner diameter) of 50 mm and a height of 3 m, and the lower connection volume of the tower was
  • the 5L tower is equipped with a 10KW electric heating rod.
  • the heating rod is controlled by the intelligent controller through the thyristor (SCR) to control the heating capacity of the tower.
  • the top of the tower is connected to a condenser with a heat exchange area of 0.5 m 2 , and the overhead steam is condensed into a liquid through the condenser and then enters a 2 L reflux tank.
  • the liquid in the reflux tank is partially refluxed to the reaction column via a reflux pump and partially recovered as a light component.
  • the operating parameters of the tower are displayed and controlled by intelligent automated control instruments.
  • the tower return flow is controlled by a reflux regulator valve, and the overhead output is controlled by the level controller of the return tank.
  • the amount of tower kettle produced is controlled by the tower tank level controller to adjust the tower tank discharge valve.
  • the acetic acid and cyclohexene raw materials (the same as in the second embodiment) are respectively charged into a 30L storage tank, and are pumped into a corresponding preheater through a metering pump to be preheated to a certain temperature and then enter the reaction tower, and the feed rate is controlled by a metering pump. , electronic scales accurately metered.
  • High temperature resistant sulfonic acid type ion exchange resin brand Amberlyst 45, by Produced by Rhom & Hass
  • the machine was immersed at 180 ° C for 10 min to completely plasticize the material, and then injected into a mold to prepare a Raschig ring type resin catalyst filler having a diameter of 5 mm, a height of 5 mm, and a wall thickness of 1 mm.
  • Examples 7 to 8 are used to illustrate a method of hydrogenating cyclohexyl acetate.
  • a cyclohexyl acetate having a purity of 99.6% was used as a hydrogenation raw material.
  • the cyclohexyl acetate is pumped into the reactor by a metering pump, and the hydrogen gas enters the reaction system through the mass flow controller for hydrogenation reaction.
  • the heat transfer oil is introduced into the outer jacket of the reaction tube to control the reaction temperature, and the reactor outlet is pressed back. Line chromatography analysis.
  • Table 6 Table " ⁇ results show that copper zinc aluminide plus Hydrogen catalyst, cyclohexyl acetate hydrogenation reaction single pass conversion rate of up to 99%, cyclohexanol single pass selectivity greater than 99%, 1000 hours of operation, single pass conversion and single pass selectivity did not decrease.
  • a cyclohexyl acetate having a purity of 99.6% was used as a hydrogenation raw material.
  • the heat transfer oil is introduced into the outer jacket of the reaction tube to control the reaction temperature, and the back pressure valve is passed through the reactor outlet. Control reactor pressure.
  • the reaction product was subjected to online chromatographic analysis by linear sampling at the rear of the reactor.
  • the reaction conditions and results are shown in Table 7.
  • the results in Table 7 show that with the copper-zinc-aluminide hydrogenation catalyst, the single-pass conversion of cyclohexyl acetate hydrogenation can reach more than 98%, the cyclohexanol one-pass selectivity is greater than 99%, the operation is 500 hours, the single pass conversion rate and one-way conversion. The selectivity has not decreased.
  • the single pass conversion of cyclohexene was calculated to be 99.4%, and the single pass selectivity of cyclohexyl acetate was 99.6%.
  • This example is intended to illustrate a method for the selective hydrogenation of benzene to cyclohexene.
  • the benzene and hydrogen molar ratio of 1:3 is injected into the hydrogenation reactor packed with the ruthenium particle catalyst, and the benzene hydrogenation reaction is carried out under the conditions of a reaction temperature of 135 ° C, a pressure of 4.5 MPa, and a residence time min, and the reaction product separates hydrogen. After that, the liquid product was collected and run continuously for 1000 h. After the end of the test, the collected liquid product was analyzed by gas chromatography, and its composition was: benzene 53.3 m%, cyclohexene 35.4 m%, cyclohexane 11.3 m%. The liquid product is subjected to extraction and separation using N, N-dimethylacetamide as an extractant to obtain a mixture of cyclohexane and cyclohexene.
  • the acetic acid and cyclohexene raw materials (obtained by the method of Example 1, the composition is: cyclohexene 75m%, cyclohexane 25m%) are respectively driven into the reactor by a metering pump at a certain flow rate, and are reacted outside the reaction tube. Hot water is introduced into the jacket to control the reaction temperature, and the reactor pressure is controlled through a reactor outlet back pressure valve. The reactor outlet product was sampled by an on-line sampling valve for on-line chromatographic analysis to calculate the cyclohexene single pass conversion and cyclohexyl acetate one-pass selectivity from the product composition. The reaction conditions and results are shown in Table 1.
  • the test apparatus, method and raw materials are the same as those in Example 2.
  • the catalyst is a phosphorus-modified ⁇ molecular sieve catalyst (the ⁇ molecular sieve with a silica-alumina ratio of 50 is modified by 85% phosphoric acid, and then kneaded with alumina to form a strip. It is dried at 120 ° C and calcined at 500 ° C with a phosphorus content of 2%.
  • the reaction conditions and results are shown in Table 2. It can be seen from Table 2 that the cyclohexene single-pass conversion rate is 90%, the ester product single-pass selectivity is greater than 99%, and the catalyst activity and single-pass selectivity are stable and stable for 480 hours.
  • the addition esterification products of Examples 2 and 3 were collected and subjected to a rectification separation test.
  • the rectification adopts a glass tower rectification unit with a height of 2 m and a diameter of 40 mm.
  • the column is equipped with a ⁇ 3 mm stainless steel crucible ring high-efficiency rectification packing, and the column is a volume 5 L glass flask with a charge of 4 L.
  • the tower is heated by a set of tubes, and the amount of heating of the tower is adjusted by a pressure regulator.
  • the reflux of the column is controlled by a reflux ratio regulator.
  • Table 3 The results of distillation separation are shown in Table 3.
  • Examples 5 to 6 are for explaining a method for producing a cyclohexyl acetate by reactive distillation.
  • the tests carried out in Examples 5 to 6 were carried out in a reactive distillation mode test apparatus of the following specifications:
  • the main body of the mode apparatus was 50 mm in diameter (inner diameter) and 3 m in height.
  • the stainless steel tower, the lower part of the tower is connected with a 5L column kettle.
  • the kettle is equipped with a 10KW electric heating rod.
  • the heating rod is controlled by the intelligent controller through the thyristor (SCR) to control the heating capacity of the tower.
  • the top of the tower is connected to a condenser with a heat exchange area of 0.5 m 2 , and the overhead steam is condensed into a liquid through the condenser and then enters a 2 L reflux tank.
  • the liquid in the reflux tank is partially refluxed to the reaction column via a reflux pump and partially recovered as a light component.
  • the operating parameters of the tower are displayed and controlled by intelligent automated control instruments.
  • the tower return flow is controlled by a reflux regulator valve, and the overhead output is controlled by the level controller of the return tank.
  • the amount of tower kettle produced is controlled by the tower tank level controller to adjust the tower tank discharge valve.
  • the acetic acid and cyclohexene raw materials (the same as in the second embodiment) are respectively charged into a 30L storage tank, and are pumped into a corresponding preheater through a metering pump to be preheated to a certain temperature and then enter the reaction tower, and the feed rate is controlled by a metering pump. , electronic scales accurately metered.
  • the high temperature sulfonic acid type ion exchange resin (brand Amberlyst 45, manufactured by Rhom & Haas) is pulverized into a powder having a particle size of less than 200 mesh (0.074 mm) by a multi-stage high-speed pulverizer, and added to a pore former, a lubricant, and an antioxidant.
  • the agent and the binder are uniformly mixed on a high-speed mixer, and then kneaded at 180 ° C for 10 minutes on the internal mixer to completely plasticize the material, and then injected into the mold to make a diameter of 5 mm, a height of 5 mm, and a wall thickness of 1 mm.
  • Ring type resin catalyst filler brand Amberlyst 45, manufactured by Rhom & Haas
  • the ball H ⁇ 3 ⁇ 4 to 0. 5 Cs 2. 5 PW 12 O 40 / SiO 2 catalyst manufactured by 12 O 40 powder and the particle size of H 0. 5 Cs 2. 5 PW less than 200 mesh coarse pore silica powder, in the mix
  • the silica sol is used as a bonding machine in a sugar-coated machine, and then baked and baked to be sandwiched into a titanium mesh wave plate to form a cylinder having a diameter of 50 mm and a height of 50 mm.
  • Type structured packing The packed catalyst L was charged into the middle of the mode reactor (high lm, equivalent to 12 theoretical plates).
  • Examples 7 to 8 are used to illustrate a method of hydrogenating cyclohexyl acetate.
  • a cyclohexyl acetate having a purity of 99.6% was used as a hydrogenation raw material.
  • the cyclohexyl acetate is pumped into the reactor by a metering pump, and the hydrogen gas enters the reaction system through the mass flow controller for hydrogenation reaction.
  • the heat transfer oil is introduced into the outer jacket of the reaction tube to control the reaction temperature, and the back pressure valve is passed through the reactor outlet. Line chromatography analysis.
  • the reaction conditions and results are shown in Table 6.
  • the results in Table 6 show that with the copper-zinc-aluminide hydrogenation catalyst, the single-pass conversion of cyclohexyl acetate hydrogenation can be up to 99%, the cyclohexanol single-pass selectivity is greater than 99%, run for 1000 hours, single pass conversion and one-way The selectivity has not decreased.
  • a cyclohexyl acetate having a purity of 99.6% was used as a hydrogenation raw material.
  • the heat transfer oil is introduced into the outer jacket of the reaction tube to control the reaction temperature, and the back pressure valve is passed through the reactor outlet. Control reactor pressure.
  • the reaction product was sampled by a linear sampling valve at the rear of the reactor for in-line color analysis.
  • the reaction conditions and results are shown in Table 7.
  • the results in Table 7 show that with the copper-zinc-aluminide hydrogenation catalyst, the single-pass conversion of cyclohexyl acetate hydrogenation can reach more than 98%, the cyclohexanol one-pass selectivity is greater than 99%, the operation is 500 hours, the single pass conversion rate and one-way conversion. The selectivity has not decreased.
  • ⁇ molecular sieve catalyst catalyzes the experimental data of acetic acid and cyclohexene esterification
  • the single pass conversion of cyclohexene was calculated to be 98.66%, and the single pass selectivity of cyclohexyl acetate was 99.3%.
  • This example is used to illustrate the test method for selective hydrogenation of benzene to cyclohexene.
  • the benzene and hydrogen molar ratio of 1:3 was injected into the hydrogenation reactor packed with the ruthenium particle catalyst, and the benzene hydrogenation reaction was carried out under the conditions of a reaction temperature of 135 ° C, a pressure of 4.5 MPa, and a residence time of 15 min, and the reaction product separated hydrogen. After that, the liquid product was collected and run continuously for 1000 h. After the end of the test, the collected liquid product was analyzed by gas chromatography, and its composition was: benzene 53.3%. Cyclohexene 35.4%, cyclohexane 1 1.3%. The above liquid product is subjected to extraction and separation using N, N-dimercaptoacetamide as an extractant to obtain cyclohexene.
  • the back pressure is used to control the reactor pressure.
  • the reactor outlet product was sampled by an on-line sampling valve for on-line chromatographic analysis, and the cyclohexene single pass conversion and cyclohexyl acetate one-pass selectivity were calculated from the product composition.
  • the reaction conditions and results are shown in Table 1.
  • the strong acid type ion exchange resin catalyst catalyzes the reaction of the cyclohexene raw material with acetic acid, the single pass conversion of cyclohexene is more than 90%, the single pass selectivity of the ester product is more than 99%, the operation is 600 hours, the catalyst activity and the single pass selectivity. Stable.
  • Example 2 The test apparatus, method and raw materials were the same as in Example 2 except that Cs 25 H 05 PW 12 O 40 /SiO 2 was used as a catalyst (referred to as PW/Si0 2 , the same applies hereinafter).
  • the reaction conditions and results are shown in Table 2. It can be seen from Table 2 that the conversion of cyclohexene with acetic acid is 95% in one pass, the selectivity of the ester product is 99% in one-pass, and the operation is 480 hours, and the catalyst activity and single-pass selectivity are stable.
  • Cs 2 . 5 H. 5 PW 12 0 4 . /Si0 2 catalyzes the test data of acetic acid and cyclohexene esterification
  • the test apparatus, method and raw materials are the same as those in Example 2.
  • the catalyst is a phosphorus-modified ⁇ molecular sieve catalyst (the ⁇ molecular sieve with a silica-alumina ratio of 50 is modified by 85% phosphoric acid, and then kneaded with alumina to form a strip. It is dried at 120 ° C and calcined at 500 ° C with a phosphorus content of 2%).
  • the reaction conditions and results are shown in Table 3. It can be seen from Table 3 that the conversion per pass of cyclohexene and acetic acid is 90%, the single-pass selectivity of the ester product is more than 99%, and the activity of the catalyst and the selectivity of one-pass are stable.
  • Table 3 Experimental data of ⁇ molecular sieve catalyst catalyzed esterification of acetic acid with cyclohexene
  • the addition esterification products of Examples 2 to 4 were collected and subjected to a rectification separation test.
  • the rectification crucible uses a glass tower rectification unit with a height of 2 m and a diameter of 40 mm.
  • the column column is equipped with a ⁇ D3mm stainless steel crucible ring high-efficiency rectification packing, and the column is a volume 5 L glass flask with a charge of 4L.
  • the electric heating jacket heats the tower kettle, and the heating capacity of the tower kettle is adjusted by a pressure regulator.
  • the reflux of the column is controlled by a reflux ratio regulator.
  • Table 4 Addition esterification product distillation separation test data
  • Examples 6 to 7 are used to illustrate a method for producing a cyclohexyl acetate by reaction.
  • the tests carried out in Examples 6 to 7 were carried out in a reactive distillation mode test apparatus of the following specifications:
  • the main body of the mode apparatus was a stainless steel tower having a diameter (inner diameter) of 50 mm and a height of 3 m, and the lower connection volume of the tower was
  • the 5L tower is equipped with a 10KW electric heating rod.
  • the heating rod is controlled by the intelligent controller through the thyristor (SCR) to control the heating capacity of the tower.
  • SCR thyristor
  • the top of the tower is connected to a condenser with a heat exchange area of 0.5 m 2 , and the overhead steam is condensed into a liquid through the condenser and then enters a 2 L reflux tank.
  • the liquid in the reflux tank is partially refluxed to the reaction column via a reflux pump and partially recovered as a light component.
  • the operating parameters of the tower are displayed and controlled by intelligent automated control instruments.
  • the tower return flow is controlled by the reflux adjustment, and the overhead output is controlled by the liquid level controller of the reflux tank.
  • the amount of tower kettle produced is controlled by the tower tank level controller to adjust the tower tank discharge valve.
  • Acetic acid and cyclohexene (obtained by the method of Example 1) are respectively charged into a 30L storage tank, and are pumped into a corresponding preheater through a metering pump to preheat to a certain temperature and then enter the reaction tower.
  • the feed rate is controlled by a metering pump. , electronic scales accurately metered.
  • High temperature resistant sulfonic acid type ion exchange resin brand Amberlyst 45, by
  • the single pass conversion of cyclohexene was calculated to be 99%, and the single pass selectivity of cyclohexyl acetate was 99.72%.
  • the ball H ⁇ 3 ⁇ 4 to 0. 5 Cs 2. 5 PW 12 O 40 / SiO 2 catalyst manufactured by 12 O 40 powder and the particle size of H 0. 5 Cs 2. 5 PW less than 200 mesh coarse pore silica powder, in the mix
  • the silica sol is used as a bonding machine in a sugar-coated machine, and then baked and baked to be sandwiched into a titanium mesh wave plate to form a cylinder having a diameter of 50 mm and a height of 50 mm.
  • Type structured packing The packed catalyst L was charged into the middle of the mode reactor (high lm, equivalent to 12 theoretical plates).
  • the single pass conversion of cyclohexene was calculated to be 98.7 %, and the single pass selectivity of cyclohexyl acetate was 99.43%.
  • Examples 8 to 9 are used to illustrate the results of hydrogenation test of cyclohexyl acetate.
  • a cyclohexyl acetate having a purity of 99.6% was used as a hydrogenation raw material.
  • the cyclohexyl acetate is pumped into the reactor by a metering pump, and the hydrogen gas enters the reaction system through the mass flow controller for hydrogenation reaction.
  • the heat transfer oil is introduced into the outer jacket of the reaction tube to control the reaction temperature, and the back pressure valve is passed through the reactor outlet. Control reactor pressure.
  • the reaction product was sampled by in-line chromatography through a linear sampling valve at the rear of the reactor. The reaction conditions and results are shown in Table 7.
  • a cyclohexyl acetate having a purity of 99.6% was used as a hydrogenation raw material.
  • 40g copper chromium ester hydrogenation catalyst (commercially produced by Taiyuan Xinjida Chemical Co., Ltd., grade Cl-XH-1, CuO content 55%, 5mm diameter tablet, broken into 10 ⁇ 20 mesh particles)
  • ) 20x2.5x800mm jacketed stainless steel tube reactor filled with a certain amount of quartz sand at both ends. After passing hydrogen (500mL/min) at 280° (:, 6MPa for 24h, the temperature and pressure are reduced.
  • the cyclohexyl acetate is pumped into the reactor by the metering pump, and the hydrogen enters the mass flow controller.
  • the reaction system is subjected to a hydrogenation reaction, the reaction temperature is controlled by introducing a heat transfer oil through the outer jacket of the reaction tube, and the reactor pressure is controlled by a reactor outlet back pressure valve.
  • the reaction product is sampled by a linear sampling valve at the rear of the reactor for online chromatographic analysis.
  • the reaction conditions and results are shown in Table 8. The results show that with the copper chromate hydrogenation catalyst, the single-pass conversion of cyclohexyl acetate hydrogenation can reach more than 98%, the cyclohexanol single-pass selectivity is greater than 99.9 %, and the operation is 500 hours. , Single pass conversion rate and selection have not decreased. Copper chromium ester catalyst catalyzed hydrogenation test data of cyclohexyl acetate
  • This example is intended to illustrate the results of the rectification separation test of the hydrogenated product of cyclohexyl acetate.
  • This example illustrates a method for the selective hydrogenation of benzene to cyclohexene.
  • the benzene and hydrogen molar ratio of 1:3 was injected into the hydrogenation reactor packed with the ruthenium particle catalyst, and the benzene hydrogenation reaction was carried out under the conditions of a reaction temperature of 135 ° C, a pressure of 4.5 MPa, and a residence time of 15 min, and the reaction product separated hydrogen. After that, the liquid product was collected and run continuously for 1000 h. After the test, the collected liquid product was analyzed by gas chromatography, and the composition was: benzene 53.3%, cyclohexene 35.4%, and cyclohexane 1 1.3%.
  • Acetic acid and cyclohexene raw materials (composition: cyclohexene 75m%, cyclohexane 25m%; obtained by extractive distillation with the reaction product of Example 1, extractant using N, N-dimercaptoacetamide)
  • the flow rate is separately driven into the reactor by the metering pump for reaction, hot water is introduced into the jacket outside the reaction tube to control the reaction temperature, and the reactor pressure is controlled through the reactor outlet back pressure valve.
  • the reactor outlet product is sampled by an on-line sampling valve for online chromatographic analysis.
  • the composition calculates the single pass conversion of cyclohexene and the single pass selectivity of cyclohexyl acetate.
  • the reaction conditions and results are shown in Table 1.
  • the test apparatus and method are the same as those in the second embodiment.
  • the catalyst is a phosphorus-modified ⁇ molecular sieve catalyst (the ⁇ molecular sieve having a silica-alumina ratio of 50 is modified by 85% citric acid, and then kneaded with alumina to form a strip. After drying at 120 ° C, calcined at 500 ° C, the phosphorus content is 2%); cyclohexene raw material (composition: benzene 60m%, cyclohexene 40m%; using Example 1 reaction product obtained by extractive distillation , the extractant uses N, N-dimercaptoacetamide). The reaction conditions and results are shown in Table 2.
  • Examples 4 to 5 are for explaining a method for producing cyclohexyl acetate by reactive distillation.
  • the tests carried out in Examples 4 to 5 were carried out in a reactive distillation mode test apparatus of the following specifications:
  • the main body of the mode apparatus was a stainless steel tower having a diameter (inner diameter) of 50 mm and a height of 3 m, and the lower connection volume of the tower was
  • the 5L tower is equipped with a 10KW electric heating rod.
  • the heating rod is controlled by the intelligent controller through the thyristor (SCR) to control the heating capacity of the tower.
  • the top of the tower is connected to a condenser with a heat exchange area of 0.5 m 2 , and the overhead steam is condensed into a liquid through the condenser and then enters a 2 L reflux tank.
  • the liquid in the reflux tank is partially refluxed to the reaction column via a reflux pump and partially recovered as a light component.
  • the operating parameters of the tower are displayed and controlled by intelligent automated control instruments.
  • the tower return flow is controlled by a reflux regulator valve, and the overhead output is controlled by the level controller of the return tank.
  • the amount of tower kettle produced is controlled by the tower tank level controller to adjust the tower tank discharge valve.
  • the acetic acid and cyclohexene raw materials are respectively charged into a 30L storage tank, and are pumped into a corresponding preheater through a metering pump to be preheated to a certain temperature and then enter the reaction tower.
  • the feed rate is controlled by the metering pump and accurately measured by the electronic scale.
  • the high temperature sulfonic acid type ion exchange resin (brand Amberlyst 45, manufactured by Rhom & Haas) is pulverized into a powder having a particle size of less than 200 mesh (0.074 mm) by a multi-stage high-speed pulverizer, and added to a pore former, a lubricant, and an antioxidant.
  • the agent and the binder are uniformly mixed on a high-speed mixer, and then kneaded at 180 ° C for 10 minutes on the internal mixer to completely plasticize the material. Thereafter, it was injected into a mold to prepare a Raschig ring type resin catalyst filler having a diameter of 5 mm, a height of 5 mm, and a wall thickness of 1 mm.
  • the ball H ⁇ 3 ⁇ 4 to 0. 5 Cs 2. 5 PW 12 O 40 / SiO 2 catalyst manufactured by 12 O 40 powder and the particle size of H 0. 5 Cs 2. 5 PW less than 200 mesh coarse pore silica powder, in the mix
  • the silica sol is used as a bonding machine in a sugar-coated machine, and then baked and baked to be sandwiched into a titanium mesh wave plate to form a cylinder having a diameter of 50 mm and a height of 50 mm.
  • Type structured packing The packed catalyst L was charged into the middle of the mode reactor (high lm, equivalent to 12 theoretical plates).
  • Examples 6 to 7 are used to illustrate a method of hydrogenating cyclohexyl acetate.
  • the hydrogenation feedstock was a cyclohexyl acetate having a purity of 99.6%.
  • the cyclohexyl acetate is pumped into the reactor by a metering pump, and the hydrogen gas enters the reaction system through the mass flow controller for hydrogenation reaction.
  • the heat transfer oil is introduced into the outer jacket of the reaction tube to control the reaction temperature, and the back pressure valve is passed through the reactor outlet. Control reactor pressure.
  • the reaction product was sampled by in-line chromatography through a linear sampling valve at the rear of the reactor. The reaction conditions and results are shown in Table 5.
  • the hydrogenation feedstock was a cyclohexyl acetate having a purity of 99.6%.
  • the heat transfer oil is introduced into the outer jacket of the reaction tube to control the reaction temperature, and the back pressure valve is passed through the reactor outlet. Control reactor pressure.
  • the reaction product was sampled by in-line chromatography through a linear sampling valve at the rear of the reactor.
  • the reaction conditions and results are shown in Table 6.
  • the results in Table 6 show that the conversion rate of single-pass hydrogenation of cyclohexyl acetate can reach up to 98.0%, the single-pass selectivity of cyclohexanol is more than 99.9 %, and the operation of 500 hours, the single pass conversion rate and selection are not decreased. .
  • the single pass conversion of cyclohexene was calculated to be 99.35%, and the single pass selectivity of cyclohexyl acetate was 99.6%.

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Abstract

一种联产环己醇和链烷醇的方法，包括环己烯酯化步骤和环己酯加氢步骤。一种以该联产方法为起始，进一步制造环己酮或己内酰胺的方法，以及一种联产环己醇和链烷醇的装置。

Description

一种联产环己醇和链烷醇的方法和装置 技术领域
本发明涉及一种制造环己醇的方法。 更具体而言， 涉及一种联产 环己醇和链烷醇的方法。 本发明还涉及一种以该环己醇的制造方法为 起始， 进一步制造环己酮或己内酰胺的方法。 本发明还涉及一种联产 环己醇和链烷醇的装置。 背景技术
环己醇和链烷醇比如乙醇均为重要的有机化工原料和有机溶剂。 环己醇主要用于脱氢法制环己酮，而环己酮则是进一步生产尼龙 6 和尼龙 66的主要中间体。 自尼龙问世以来， 世界各大化工公司一直致 力于开发环己醇 （酮） 的工业来源。 20 世纪 80 年代， 日本旭化成公 司开发了利用环己烯水合制造环己醇的方法 （环己烯水合法） 。 但是， 该方法存在以下的不足： ①环己烯直接水合存在着热力学限制， 并且 环己烯在水中的溶解度很小， 水合反应主要发生在两相界面上， 这导 致了环己烯水合反应的速率较慢、单程转化率 ί艮低，如采用高硅 ZSM-5 催化剂， 在两个串联浆态反应器中停留 2h， 环己烯的单程转化率也只 有 12.5%,如此低的单程转化率不得不将大量未反应的环己烯从产物物 流中分离出来循环回用， 这大大增加了过程的能耗； ②需要用高纯度 的环己烯作为原料， 否则由于其他组分的稀释作用， 将导致更多的循 环物料及更低的反应效率， 而环己烯是由苯部分加氢制造， 其产物物 流中除环己烯外还含有相当多的环己烷和苯， 这三者的沸点非常接近， 这使得环己烯的提純非常困难， 并且提纯成本极其高昂； ③该方法采 用了包括水相、 油相和固体催化剂相的复杂反应体系， 既需要依靠强 力搅拌， 形成水滴、 油滴良好分散的乳化液体系， 以使环己烯能够被 催化剂表面吸附， 又需要在沉降区将催化剂与油相很好地分离， 操作 起来比较复杂， 催化剂损失也比较严重。 品丰富的国家， ϋ法仍是生产乙醇的主要方法。 发酵法的主要缺点 是污染比较严重， 此外发酵法还存在着 "与口争粮，，的问题，对于人口多 且耕地面积少的国家并不适用。 乙烯直接水合法的反应条件比较苛刻， 需要在高温、 高压下进行， 此外乙烯价格受国际油价波动的影响很大， 对于石油资源不足的国家， 采用乙烯水合法制乙醇会面临一定的原料 成本压力。
因此， 现有技术目前的现状是， 仍旧需要一种环己醇的制造方法， 尤其是一种可以联产环己醇和链烷醇比如乙醇的方法， 其能够比如以 更为简单的制造工艺和更为低廉的生产成本联产环己醇和链烷醇比如 乙醇， 并且可以克服现有技术中存在的那些问题。 发明内容
本发明人在现有技术的基础上经过刻苦的研究发现， 通过至少包 括环己烯酯化步骤和环己酯加氢步骤这两个步骤， 就能够比如以更为 简单的制造工艺和更为低廉的生产成本联产环己醇和链烷醇比如乙 醇， 并且可以解决现有技术存在的前述问题， 由此完成了本发明。
具体而言， 本发明涉及以下方面的内容：
1. 一种联产环己醇和链烷醇的方法， 其特征在于， 包括以下步骤： ( 1 )使环己烯源与至少一种羧酸在加成酯化催化剂的存在下发生 加成酯化反应， 生成含有羧酸环己酯的加成酯化产物的步骤， 其中所 述至少一种羧酸用式 R-COOH表示，并且所述基团 R是氢或 C1-23直链 或支链烷基， 优选 d _6直链或支链烷基， 更优选 d.3直链或支链烷基， 最优选甲基， 所述环己烯源含有 20mol%以上、 35mol%以上、 20 至 80mol%、 20至 60mol%、 40至 80mol%、 80至 95mol%或者 95mol%以 上的环己烯； 和
( 2 )使所述加成酯化产物与氢气在加氢催化剂的存在下发生加氢 反应， 同时生成环己醇和链烷醇的步骤， 其中所述链烷醇用式 R-CH2-OH表示， 并且所述基团 R与所述至少一种羧酸中的定义相同， 最优选曱基。
2. 前述任一方面所述的方法， 进一步包括下述的步骤 （A )、 步骤 ( A ) +步骤 （B )、 步骤 （C )、 步骤 （C ) +步骤 （D )之一或其任意的 组合：
( A )使苯与氢气在部分加氢催化剂的存在下发生部分加氢反应， 以获得含有环己烯的加氢产物作为所述环己烯源的步骤；
( B ) 对步骤 （A ) 获得的所述加氢产物进行进一步分离， 以获得 环己烯、 环己烯与苯的混合物或环己烯与环己烷的混合物作为所述环 己烯源的步骤；
( C )使环己烷在部分脱氢催化剂的存在下发生部分脱氢反应， 以 获得含有环己烯的部分脱氢产物作为所述环己烯源的步骤；
( D ) 对步骤 （C ) 获得的所述部分脱氢产物进行进一步分离， 以 获得环己烯或环己烯与环己烷的混合物作为所述环己烯源的步骤。
3. 前述任一方面所述的方法， 进一步包括下述的步骤 （1 )、 步骤 ( II ) 和步骤 （III ) 之一或其任意的组合：
( I ) 回收从所述联产环己醇和链烷醇的方法的任何步骤中分离的 苯和 /或氢气， 并将该苯和 /或氢气循环至所述步骤 （A );
( II )回收从所述联产环己醇和链烷醇的方法的任何步骤中分离的 环己烷， 并将该环己烷循环至所述步骤 （C );
( III ) 回收从所述联产环己醇和链烷醇的方法的任何步骤中分离 的环己烷， 使该环己烷在脱氢催化剂的存在下发生脱氢反应， 以获得 苯和氢气， 并将该苯和 /或氢气循环至所述步骤 （A )。
4. 前述任一方面所述的方法， 其中所述加成酯化催化剂选自固体 酸催化剂中的一种或多种， 优选选自酸强度函数（Hammett 函数） H0 为 -8 以下 （优选 -12 以下， 更优选 -13 以下） 的固体酸催化剂中的一种 或多种， 更优选选自强酸型离子交换树脂 （优选选自磺酸型离子交换 树脂中的一种或多种， 更优选选自大孔磺酸型离子交换树脂和卤素改 性磺酸型离子交换树脂中的一种或多种）、 杂多酸（比如选自 keggin结 构的杂多酸、 Dawson结构的杂多酸、 Anderson结构的杂多酸、 Silverton 结构的杂多酸、 前述杂多酸的酸式盐、 前述杂多酸 /载体和前述杂多酸 的酸式盐 /载体中的一种或多种，优选选自 keggin结构的杂多酸、 keggin 结构的杂多酸的酸式盐、 keggin结构的杂多酸 /载体和 keggin结构的杂 多酸的酸式盐 /载体中的一种或多种， 更优选选自十二磷钨酸或十二磷 钨酸 /载体、 十二硅钨酸或十二硅钨酸 /载体、 十二磷钼酸或十二磷钼酸 /载体、 十二磷钼钒酸或十二磷钼钒酸 /载体、 前述杂多酸的酸式盐和前 述杂多酸的酸式盐 /载体中的一种或多种， 更优选选自酸式磷钨酸铯盐
( Cs2.5Ho.5Pi2W04o ) 和酸式磷钨酸铯盐 /载体中的一种或多种； 所述载 体比如选自二氧化硅和活性炭中的一种或多种） 和沸石分子筛 （优选 选自 Ηβ沸石分子筛、 氟和 /或磷改性的 Ηβ沸石分子筛、 ΗΥ沸石分子 筛、 氟和 /或磷改性的 HY沸石分子筛、 HZSM-5沸石分子筛以及氟和 / 或磷改性的 HZSM-5沸石分子筛中的一种或多种） 中的一种或多种。
5. 前述任一方面所述的方法， 其中所述加氢催化剂选自铜系催化 剂 （更优选选自含锌的铜系催化剂和含铬的铜系催化剂中的一种或多 种）、钌系催化剂 (优选选自 Ru/Al203和 Ru-Sn/Al203中的一种或多种 ) 和贵金属系催化剂 （优选选自 Pt/Al203、 Pd-Pt/Al203和 Pd/C中的一种 或多种） 中的一种或多种， 优选铜系催化剂中的一种或多种。
6. 前述任一方面所述的方法， 其中所述铜系催化剂包括以下组分 (优选由以下组分构成）： （a)氧化铜； （b)氧化锌； （c)选自铝、 镓、 锡、 钛、 锆、 铬、 钼、 钨、 锰、 铼、 镧系金属和锕系金属中的一种或 多种金属的氧化物， 优选选自铝、 镓、 锡、 钛、 锆、 铬、 钼、 钨、 锰、 铼、 镧和铈中的一种或多种金属的氧化物； 和 （d) 选自碱金属氢氧化 物和碱土金属氢氧化物中的一种或多种， 优选选自氢氧化钾、 氢氧化 钠和氢氧化钡中的一种或多种， 其中以质量份数计， 组分 （a): 组分 (b): 组分 （c): 组分 （d) 为 5至 60: 10至 50: 5至 60: 0.2至 2， 优选 10至 50: 15至 45: 15至 55: 0.2至 2, 更优选 30至 45: 20至 35: 20至 50: 0.5至 1.5。
7. 前述任一方面所述的方法， 其中所述铜系催化剂是通过包括以 下步骤的制造方法制造的：
( 1 ) 通过共沉淀法， 制造复合金属氧化物的步驟， 其中所述复合 金属氧化物包括以下组分（优选由以下组分构成）： （a)氧化铜； （b) 氧化锌； 和 （c) 选自铝、 镓、 锡、 钛、 锆、 铬、 钼、 钨、 锰、 铼、 镧 系金属和锕系金属中的一种或多种金属的氧化物， 优选选自铝、 镓、 锡、 钛、 锆、 铬、 钼、 钨、 锰、 铼、 镧和铈中的一种或多种金属的氧 化物， 其中以质量份数计， 组分 （a) ： 组分 （b) ： 组分 （c) 为 5至 60: 10至 50: 5至 60, 优选 10至 50: 15至 45: 15至 55, 更优选 30 至 45: 20至 35: 20至 50; 和
(2) 通过浸渍法， 向所述复合金属氧化物中引入 （d) 选自碱金 属氢氧化物和碱土金属氢氧化物中的一种或多种（优选选自氢氧化钾、 氢氧化钠和氢氧化钡中的一种或多种） 的步骤， 使得以质量份数计， 组分 （ a ) ： 组分 （ b ) ： 组分 （ c ) ： 组分 （ d ) 为 5至 60: 10至 50: 5至 60: 0.2至 2, 优选 10至 50: 15至 45: 15至 55: 0.2至 2， 更优 选 30至 45: 20至 35: 20至 50: 0.5至 1.5。
8. 前述任一方面所述的方法， 其中所述至少一种羧酸与以环己烯 计的所述环己烯源的摩尔比为 0.2至 20:1, 优选 1.2至 4:1， 更优选 1.2 至 3:1, 并且所述步骤 （ 1 )按照以下的方式 （ 1 )、 方式 （2) 或其任意 组合进行， 优选方式 （2) 或方式 （ 1 ) 与方式 （2) 的组合， 更优选先 进行方式 （ 1 ) 然后再进行方式 （2) 的组合：
方式 （ 1 ): 反应器为釜式反应器、 固定床反应器、 流化床反应器、 沸腾床反应器或其任意并联组合， 优选管式固定床反应器， 更优选管 壳列管式反应器， 反应温度为 50至 200 °C, 优选 60至 120°C, 反应压 力为常压至 lOMPa, 优选常压至 lMPa， 所述加成酯化反应按照连续方 式进行时， 液体进料空速为 0.5至 20^， 优选 0.5至 5h , 更优选 1至 5h_1, 所述加成酯化反应按照间歇方式进行时， 反应时间为 0.1 ~ 10h， 优选 0.2~2h;
方式 （2): 反应器为反应精镏塔， 优选选自板式塔、 填料塔或其 任意并联组合， 理论塔板数为 10至 150, 优选 30至 100， 操作压力为 -0.0099MPa至 5MPa, 优选常压至 IMPa, 加成酯化催化剂床层装填区 的温度为 40至 200°C, 优选 50至 20CTC, 更优选 60至 120°C, 回流比 为 0.1:1至全回流， 优选 0.1至 100:1, 更优选 0.5至 10:1, 在所述理论 塔板数的 1/3至 2/3位置之间选择 5至 30块板（优选 8至 20块板）布 置所述加成酯化催化剂， 并且相对于加成酯化催化剂的总装填体积， 液体进料空速为 0.1至 2011-1, 优选 0.2至 2011-1， 更优选 0.5至 511
9. 前述任一方面所述的方法， 其中所述步骤（2)在以下反应条件 下进行： 反应器为釜式反应器、 固定床反应器、 沸腾床反应器、 流化 床反应器或其任意并联组合， 优选管式固定床反应器， 更优选管壳列 管式反应器， 反应温度为 150至 400°C, 优选 200至 300°C, 反应压力 为常压至 20MPa, 优选常压至 lOMPa, 更优选 4至 lOMPa, 氢气与以 羧酸环己酯计的所述加成酯化产物的摩尔比为 1 至 1000:1, 优选 5至 100: 1, 所述加氢反应按照连续方式进行时， 液体进料空速为 0.1 至 20H"1, 优选 0.2至 2h , 所述加氢反应按照间歇方式进行时， 反应时间 为 0.2至 20h, 优选 0.5至 5h， 更优选 1至 5h。
10. 前述任一方面所述的方法，还包括在使所述加成酯化产物与氢 气发生加氢反应之前， 对所述加成酯化产物进行分离的步骤， 以获得 羧酸环己酯或羧酸环己酯与所述至少一种羧酸的混合物作为所述加成 酯化产物， 优选获得羧酸环己酯作为所述加成酯化产物，
和 /或，
在使所述加成酯化产物与氢气发生加氢反应之前， 使所述加成酯 化产物与氢气在羧酸加氢催化剂的存在下发生羧酸加氢反应的步骤， 以将所述加成酯化产物中含有的游离羧酸转化为链烷醇。
1 1. 前述任一方面所述的方法，其中所述羧酸加氢催化剂由主活性 组分 0.1至 30wt%、 助剂 0.1至 25wt%和余量的载体构成， 其中所述主 活性组分选自铂、 钯、 钌、 钨、 钼和钴中的一种或多种， 所述助剂选 自锡、 铬、 铝、 锌、 钙、 镁、 镍、 钛、 锆、 铼、 镧、 钍和金中的一种 或多种， 所述载体选自氧化硅、 氧化铝、 氧化钛、 氧化锆、 活性炭、 石墨、 纳米炭管、 硅酸钙、 沸石和硅酸铝中的一种或多种，
和 /或，
所述羧酸加氢反应在以下反应条件下进行： 反应器为釜式反应器、 固定床反应器、 沸腾床反应器、 流化床反应器或其任意并联组合， 优 选管式固定床反应器， 更优选管壳列管式反应器， 反应温度为 100 至 400 °C ,优选 180至 300 °C ,反应压力为 0. 1至 30MPa,优选 2至 l OMPa, 氢气与所述游离羧酸的摩尔比为 1 至 500: 1 , 优选 5至 50: 1 , 所述 羧酸加氢反应按照连续方式进行时， 液体进料空速为 0.1至 5^， 优选 0.2 至 2h ， 所述羧酸加氢反应按照间歇方式进行时， 反应时间为 0.5 至 20h, 优选 1至 5h。
12. 前述任一方面所述的方法，其中回收从所述联产环己醇和链烷 醇的方法的任何步骤中分离的羧酸和 /或环己烯源， 并将该羧酸和 /或环 己烯源循环至所述步骤 （ 1 ), 和 /或， 回收从所述联产环己醇和链烷醇 的方法的任何步骤中分离的氢气， 并将该氢气循环至所述步骤 （2 )。
13. 一种制造环己酮的方法， 其特征在于， 包括：
按照前述任一方面所述的方法制造环己醇， 和
使用所述环己醇制造环己酮。
14. 一种制造己内酰胺的方法， 其特征在于， 包括：
按照前述任一方面所述的方法制造环己酮， 和
使用所述环己酮制造己内酰胺。
15. 一种联产环己醇和链烷醇的装置， 其特征在于， 包括加氢反应 单元 A、 任选的加氢产物分离单元 A、 加成酯化反应单元、 任选的加 成酯化产物分离单元、 加氢反应单元 B和加氢产物分离单元 B, 其中， 在所述加氢反应单元 A中， 使苯与氢气在部分加氢催化剂的存在 下发生部分加氢反应， 以获得含有环己烯的加氢产物；
在所述加氢产物分离单元 A中， 对来自所述加氢反应单元 A的所 述加氢产物进行分离， 以获得环己烯、 环己烯与苯的混合物或者环己 烯与环己烷的混合物；
在所述加成酯化反应单元中， 使来自所述加氢反应单元 A的所述 加氢产物和 /或来自所述加氢产物分离单元 A的环己烯、 环己烯与苯的 混合物或者环己烯与环己烷的混合物与羧酸在加成酯化催化剂的存在 下发生加成酯化反应， 生成含有羧酸环己酯的加成酯化产物；
在所述加成酯化产物分离单元中， 对来自所述加成酯化反应单元 的所述加成酯化产物进行分离， 以获得羧酸环己酯或者羧酸环己酯与 羧酸的混合物；
在所述加氢反应单元 B 中， 使来自所述加成酯化反应单元的所述 加成酯化产物和 /或来自所述加成酯化产物分离单元的羧酸环己酯或者 羧酸环己酯与羧酸的混合物与氢气在加氢催化剂的存在下发生加氢反 应， 生成含有环己醇和链烷醇的加氢产物； 和
在所述加氢产物分离单元 B中， 对来自所述加氢反应单元 B的所 述加氢产物进行分离， 以获得环己醇和链烷醇。 技术效果
根据本发明， 比如可以取得如下的技术效果。
根据本发明的联产环己醇和链烷醇的方法， 环己烯酯化步骤和环 己酯加氢步骤的单程选择性和单程转化率均很高， 几乎没有副产物， 尤其是环己酯加氢步骤的单程选择性和单程转化率均接近百分之百， 因此该方法的总体单程选择性和单程转化率均非常高， 比如远高于旭 化成开发的环己醇制造方法， 与现有技术相比， 具有环己醇制造成本 低和原子经济性高的特点。
根据本发明的联产环己醇和链烷醇的方法， 氢耗低， 全过程几乎 没有三废排放， 与现有技术相比， 具有环保性强和原子经济性高的特 点。 根据本发明的联产环己醇和链烷醇的方法， 全过程的反应条件温 和， 与现有技术相比， 生产安全性较高。
根据本发明的联产环己醇和链烷醇的方法， 通过单独或组合使用 反应精馏塔来进行环己烯的加成酯化反应， 在显著提高环己烯单程转 化率 （可达 99%以上） 的同时， 大大简化了酯化产物的分离操作， 与 现有技术相比， 具有环己醇制造成本显著低和原子经济性高的特点。
根据本发明的联产环己醇和链烷醇的方法， 环己烯酯化步骤的反 应历程简单， 副反应较少， 受杂质影响较小， 由此该方法对环己烯原 料的纯度要求较低， 可以直接使用其粗产品 （环己烯含量最低比如可 以为 20mol% )作为起始原料。 比如， 本发明在本领域第一次发现， 该 联产环己醇和链烷醇的方法甚至可以直接使用苯部分加氢的产物物流 作为起始原料而不必然需要复杂或昂贵的预先提纯或分离， 与现有技 术相比， 生产成本显著降低。
根据本发明的联产环己醇和链烷醇的方法， 在通过苯部分加氢获 得环己烯源时， 苯部分加氢副产的环己烷通过脱氢反应以非常高的转 化率和选择性重新转化为苯， 使苯部分加氢、 环己烯酯化、 环己酯加 氢全过程的碳利用率接近 100%。
根据本发明的联产环己醇和链烷醇的方法， 在制造环己醇的同时， 联产与羧酸原料相比经济附加值更高的链烷醇， 尤其是乙醇。 在联产 这些经济附加值更高的链烷醇比如乙醇时， 不需要将羧酸与羧酸环己 酯分离， 将未反应的羧酸与羧酸环己酯共同加氢可获得环己醇和更多 的这些链烷醇， 尤其是采用反应精馏塔进行酯化反应时， 这样做还具 有比如降低反应精镏塔操作温度的好处。
根据本发明的联产环己醇和链烷醇的方法， 反应体系简单， 操作 简单方便， 与现有技术相比， 可以显著降低环己醇的制造成本和制造 装置的维护成本。 附图说明
图 1 为示意性地表示本发明的联产环己醇和链烷醇的方法的总体 流程图。
图 2至图 16为示意性地表示本发明的联产环己醇和链烷醇的方法 的各种具体实施方案的流程图。 具体实施方式
下面对本发明的具体实施方式进行详细说明， 但是需要指出的是， 本发明的保护范围并不受这些具体实施方式的限制， 而是由附录的权 利要求书来确定。
在本发明的上下文中， 除非另有明确定义， 或者该含义超出了本 领域技术人员的理解范围， 3个碳原子以上的烃或烃衍生物基团（比如 丙基、 丙氧基、 丁基、 丁烷、 丁烯、 丁烯基、 己烷等）在未冠以词头 "正" 时均具有与冠以词头 "正"时相同的含义。比如，丙基一般理解为正丙基， 而丁基一般理解为正丁基， 除非另有明确。
本说明书提到的所有出版物、 专利申请、 专利和其它参考文献全 都引于此供参考。 除非另有定义， 本说明书所用的所有技术和科学术 语都具有本领域技术人员常规理解的含义。 在有冲突的情况下， 以本 说明书的定义为准。
当本说明书以词头"本领域技术人员公知"、 "现有技术"或其类似用 语来导出材料、 物质、 方法、 步骤、 装置或部件等时， 该词头导出的 对象涵盖本申请提出时本领域常规使用的那些， 但也包括目前还不常 用， 却将变成本领域公认为适用于类似目的的那些。
在本说明书的上下文中， 除了明确说明的内容之外， 未提到的任 何事宜或事项均直接适用本领域已知的那些而无需进行任何改变。 而 且， 本文描述的任何实施方式均可以与本文描述的一种或多种其他实 施方式自由结合， 由此而形成的技术方案或技术思想均视为本发明原 始公开或原始记载的一部分， 而不应被视为是本文未曾披露或预期过 的新内容， 除非本领域技术人员认为该结合是明显不合理的。
此外， 本说明书提到的各种范围均包括它们的端点在内， 除非另 有明确说明。 此外， 当对量、 浓度或其它值或参数给出范围、 一个或 多个优选范围或很多优选上限值与优选下限值时， 应把它理解为具体 公开了由任意对任意范围上限值或优选值与任意范围下限值或优选值 所形成的所有范围， 不论是否一一公开了这些数值对。
在本说明书的上下文中， 在没有明确指明的情况下， 选择性指的 是单程选择性， 而转化率则指的是单程转化率。
在本说明书的上下文中，术语"环己烯源 "指的是可以在本发明的环 己烯酯化步骤中作为环己烯来源 （即提供环己烯） 的任何反应原料， 包括环己烯工业纯品 （环己烯含量比如是 95mol%以上）、 环己烯工业 粗产品 （环己烯含量比如最低为 80mol%, 最高为 95mol% ) 或含环己 烯的工业混合产品（环己烯含量比如最低为 20mol%, 最高为 80mol% ) 等。 根据本发明的该环己烯酯化步骤， 反应历程筒单， 副反应较少， 酯化反应受杂质影响较小， 由此该步骤对环己烯源的纯度要求较低。 鉴于此， 在本说明书的上下文中， 术语"环己烯源，，还包括含有环己烯的 工业废料或工业副产品， 比如含环己烯废气 （比如来自现有技术的环 己烯水合法）和含环己烯尾气（比如作为化学合成工业的副产物） 等。 一般而言， 只要该环己烯源， 尤其是该工业废料或副产品中除环己烯 以外的杂质的种类或含量对本发明的环己烯酯化步骤不产生显著影响 (比如使环己烯单程转化率的降低不超过 5% ) 即可直接使用， 而无需 对其进行预先的纯化或除杂处理。 这类杂质对于本发明的环己烯酯化 步骤而言是化学惰性的， 比如可以举出氮气、 稀有气体、 二氧化碳、 苯、 氢气或者环己烷等， 在本申请中被称为惰性稀释剂。 当然， 本领 域技术人员通过简单的试验（比如通过测定环己烯单程转化率的降低 程度）， 就能够确认某一工业废料或工业副产品是否含有或者过量地含 有对本发明的环己烯酯化步骤产生显著影响的杂质， 由此确认其是否 可以直接应用于本发明的联产方法。 另外， 根据需要， 本领域技术人 员也可以通过常规已知的技术手段， 将某一工业废料或工业副产品中 含有的这类杂质降低至不显著影响本发明的联产方法实施的水平， 以 及根据需要， 将某一工业废料或工业副产品中环己烯的浓度浓缩至更 适宜本发明的联产方法实施的水平 （比如通过浓缩， 将环己烯的含量 提高至占该工业废料或工业副产品总量的 20mol%以上）。
最后， 在没有明确指明的情况下， 本说明书内所提到的所有百分 数、 份数、 比率等都是以重量为基准的， 除非以重量为基准时不符合 本领域技术人员的常规认识。
根据本发明， 涉及一种联产环己醇和链烷醇的方法， 该方法至少 包括 （ 1 ) 环己烯酯化步骤和 （2 ) 环己酯加氢步骤。
以下对该环己烯酯化步骤进行具体的说明。
根据本发明的该环己烯酯化步骤， 使环己烯源与至少一种羧酸在 加成酯化催化剂的存在下发生加成酯化反应， 生成含有羧酸环己酯的 加成酯化产物。
在本发明中， 所谓"加成酯化反应，，， 指的是羧酸对环己烯的烯键加 成而生成羧酸环己酯的反应。
根据本发明， 所述至少一种羧酸可以用式 R-COOH表示， 其中基 团 R是氢或 d_23直链或支链烷基。
根据本发明， 所述基团 R优选 直链或支链烷基， 优选 d.6直 链或支链烷基， 更优选 C1-3直链或支链烷基， 最优选甲基。
根据本发明， 所述羧酸可以单独使用一种， 也可以两种或多种组 合使用。 其中， 优选使用曱酸、 乙酸、 丙酸和正丁酸中的一种或多种， 最优选使用乙酸作为所述羧酸。
根据本发明， 所述环己烯源中环己浠的含量一般选自以下范围： 20mol%以上、 35mol%以上、 20至 80mol%、20至 60mol%、40至 80mol%、 80至 95mol%或者 95mol%以上。
根据本发明， 所述环己烯源中还可以包括余量的前述惰性稀释剂。 作为所述惰性稀释剂， 优选苯、 环己烷或其任意比例的组合。
根据本发明， 作为所述环己烯源， 优选环己烯工业纯品 （环己烯 含量比如是 95mol%以上）、 环己烯工业粗产品 （环己烯含量比如最低 为 80mol%, 最高为 95mol% ) 或含环己烯的工业混合产品 （环己烯含 量比如最低为 20mol%, 最高为 80mol% ), 尤其是所述含环己烯的工业 混合产品。 这些环己烯源均可以以工业产品的形式方便地制造或市售 获得。
根据本发明， 作为所述含环己烯的工业混合产品， 比如可以举出 ( 1 )使苯与氢气在部分加氢催化剂的存在下发生部分加氢反应而获得 的含有环己烯的加氢产物 （也称为苯部分加氢的产物物流）， 或者对所 述加氢产物进行进一步分离而获得的环己烯纯品、 环己烯与苯的混合 物或者环己烯与环己烷的混合物， 或者 （2 )使环己烷在部分脱氢催化 剂的存在下发生部分脱氢反应而获得的含有环己烯的部分脱氢产物 (也称为环己烷部分脱氢的产物物流）， 或者对所述部分脱氢产物进行 进一步分离而获得的环己烯纯品或环己烯与环己烷的混合物。
鉴于此， 根据本发明的联产环己醇和链烷醇的方法， 在所述环己 烯酯化步骤之前， 还任选包括下述的步骤（A )、 步骤（A ) +步骤（B )、 步骤 （C )、 步骤 （C ) +步骤 （D )之一或其任意的组合： ( A ) 使苯与氢气在部分加氢催化剂的存在下发生部分加氢反应 (苯部分加氢法）， 以获得含有环己烯的加氢产物作为所述环己烯源的 步骤；
( B ) 对步骤 （A ) 获得的所述加氢产物进行进一步分离， 以获得 环己烯、 环己烯与苯的混合物或环己烯与环己烷的混合物作为所述环 己烯源的步骤；
( C )使环己烷在部分脱氢催化剂的存在下发生部分脱氢反应（环 己烷部分脱氢法）， 以获得含有环己烯的部分脱氢产物作为所述环己烯 源的步骤；
( D ) 对步骤 （C ) 获得的所述部分脱氢产物进行进一步分离， 以 获得环己烯或环己浠与环己烷的混合物作为所述环己烯源的步骤。
根据本发明， 对所述苯部分加氢法 （步骤 （A ) ) 和所述环己烷部 分脱氢法 （步骤 （C ) ) 及其后续的分离方法 （步骤 （B ) 或步骤 （D ) ) 等均没有特别的限定， 可以直接使用现有技术已知的那些。
根据本发明， 作为一个优选的实施方式， 在所述步骤（A ) 中， 按 照本领域常规已知的方式采用液相法来进行所述苯部分加氢法。 作为 步骤（A ) 中使用的所述部分加氢催化剂， 优选钌系催化剂， 更优选含 钴和 /或锌的钌系催化剂。 这类催化剂可以按照现有技术已知的方式通 过比如共沉淀法或浸渍同一载体的方法来制造。
根据该步骤 （A ), 所述含有环己烯的加氢产物 （也称为苯部分加 氢的产物物流） 一般是环己烷、 环己烯和苯的混合物， 它可以直接作 为本发明的环己烯源使用。
根据本发明， 在所述步骤（B ) 中， 可以通过现有技术已知的任何 方法从所述苯部分加氢的产物物流中分离出环己烯 （也称为环己烯純 品）、 环己烯与苯的混合物或环己烯与环己烷的混合物， 并直接作为本 发明的环己烯源使用。 作为所述分离方法， 比如可以举出萃取精馏法 或共沸精馏法， 其中优选萃取精馏法。 根据该萃取精馏法， 可以使用 N-甲基 -2-吡咯烷酮、 N, N-二曱基乙酰胺、 己二腈、 丙二酸二曱酯、 琥珀酸二曱酯、 乙二醇或环丁砜等作为萃取剂。
根据本发明， 作为该萃取精镏法的一例， 比如可以举出将所述苯 部分加氢的产物物流从中部送入萃取精馏塔， 将 N , N-二甲基乙酰胺 从塔上部引入， 塔顶得到环己烷物流 （即环己烷纯品） 或环己烷与环 己烯的混合物流（即环己烷和环己烯的混合物）。 根据塔顶的具体分离 情况， 在塔底得到含有环己烯、 苯和 N, N-二甲基乙酰胺的溶液肘， 将该溶液送入精馏塔进一步分离， 由塔顶可得到环己烯和苯的混合物 流（即环己烯和苯的混合物）， 或者在塔底得到含有苯和 N, N-二甲基 乙酰胺的溶液时， 将该溶液送入精馏塔进一步分离， 由塔顶可得到苯 物流（即苯纯品）， 由精馏塔底得到 N, N-二甲基乙酰胺。 所述环己烷 与环己烯的混合物流送入精馏塔， 塔上部引入 N, N-二甲基乙酰胺， 塔顶得到环己烷物流（即环己烷纯品）， 塔底得到环己烯和 N, N-二甲 基乙酰胺的混合物流， 塔底物流送入下一级精馏塔分离出环己烯， 塔 顶得到环己烯物流（即环己烯纯品）， 塔底得到 N, N-二曱基乙酰胺物 流。 所述环己烯与苯的混合物流送入精馏塔， 塔上部引入 N, N-二甲 基乙酰胺， 塔顶得到环己烯物流 （即环己烯纯品）， 塔底得到苯和 N, N-二曱基乙酰胺的混合物流， 塔底物流送入下一级精馏塔分离出苯， 塔顶得到苯物流（即苯纯品）， 塔底得到 N, N-二甲基乙酰胺物流。 此 时， 所述苯物流可以作为步骤（A )反应进料的一部分循环利用， 而所 述环己烷物流则作为副产品排出或者作为步骤（ C )反应进料的一部分 循环利用。
根据本发明， 根据步骤 （A ) 或步骤 （B ) 的具体操作情况， 所述 苯部分加氢的产物物流、 所述环己烯与苯的混合物或者所述环己烯与 环己烷的混合物中环己烯的含量一般最低为 20mol%、 35mol%或者 40mol%, 最高为 60mol%或者 80mol%， 而所述环己烯纯品的环己烯的 含量一般为 95mol%以上或者更高。
根据本发明， 作为一个实施方式， 在所述步骤（C ) 中， 按照本领 域常规已知的方式比如环己烷氧化脱氢进行所述环己烷部分脱氢。
根据本发明， 在所述步骤（D ) 中， 可以通过现有技术已知的任何 方法从所述环己烯部分脱氢的产物物流中分离出环己烯 （也称为环己 烯纯品） 或环己烯与环己烷的混合物， 并直接作为本发明的环己烯源 使用。
根据本发明， 作为所述加成酯化催化剂， 比如可以举出酸催化剂， 比如液体酸催化剂， 具体比如磷酸、 硫酸等无机酸或者甲基苯磺酸、 胺基磺酸等有机酸， 或者固体酸催化剂。 作为所述加成酯化催化剂， 从降低设备腐蚀、 改善催化剂与酯化产物分离等角度， 优选固体酸催 化剂， 尤其是酸强度函数 （Hammett函数） H0为 -8以下 （优选 -12以 下， 更优选 -13以下） 的固体酸催化剂。 作为所述固体酸催化剂， 更优 选强酸型离子交换树脂、 杂多酸或者沸石分子筛。 这些固体酸催化剂 可以单独使用一种， 也可以两种或多种组合使用。
根据本发明， 作为所述强酸型离子交换树脂， 优选磺酸型离子交 换树脂， 更优选大孔磺酸型离子交换树脂 （大孔磺酸型聚苯乙烯-二 乙烯基苯树脂） 或者 [¾素改性磺酸型离子交换树脂。 这些强酸型离子 交换树脂很容易从市场中购得， 也可以按经典文献记载的方法制取。 大孔磺酸型聚苯乙烯 -二乙烯基苯树脂的制造方法通常是将苯乙烯和 二乙烯基苯的混合物在高速搅拌的条件下滴入含有分散剂、 引发剂、 致孔剂的水相体系中进行悬浮共聚， 将所得到的聚合物小球 （白球） 从体系中分离出来， 用溶剂抽去其中的致孔剂， 再以二氯乙烷为溶剂、 浓硫酸为磺化剂， 进行磺化反应， 最后经过滤、 洗涤等工序， 最后制 得产品。 另外， 在强酸型离子交换树脂的骨架中引入卤素原子， 如氟、 氯、 溴等， 可进一步提高树脂的耐温性能和酸强度。 这种卤素改性磺 酸型离子交换树脂至少可以通过以下两种途径获得， 一种途径是在磺 化苯乙烯树脂骨架的苯环上引入 素原子， 例如氯原子， 由于卤素元 素的强吸电子作用不仅可使苯环稳定、 而且还可以提高苯环上磺酸基 团的酸性， 这样可使树脂催化剂的酸强度函数（ Hammett函数） H0≤-8, 而且可以在 150°C以上长期使用， 此类树脂可从市场上方便购买到， 比 如国外 ROHM & HASS公司生产的 Amberlyst 45树脂， 国内河北翼中 化工厂生产的 D008树脂等； 另一种途径将树脂骨架上的氢全部用氟取 代， 由于氟的强吸电子性， 使其具有超强的酸性和超高的热稳定性， 酸强度函数（Hammett函数） H0可小于 -12 , 而耐热温度达到 250°C以 上， 这类耐高温强酸性树脂的典型例子是 DuPont公司生产的 Nafion 树脂。 这些强酸型离子交换树脂可以单独使用一种， 也可以两种或多 种组合使用。
根据本发明， 作为所述杂多酸 （酸强度函数 H0—般小于 -13.15 ), 比如可以举出 keggin结构的杂多酸、 Dawson结构的杂多酸、 Anderson 结构的杂多酸、 Silverton 结构的杂多酸、 前述杂多酸的酸式盐、 前述 杂多酸 /载体和前述杂多酸的酸式盐 /载体，其中优选 keggin结构的杂多 酸、 keggin结构的杂多酸的酸式盐、 keggin结构的杂多酸 /载体和 keggin 结构的杂多酸的酸式盐 /载体。 优选的是， 这些杂多酸可以在 300°C以 上长期使用， 并且其 BET法比表面积一般为 100m2/g以上。 这些杂多 酸可以单独使用一种， 也可以两种或多种组合使用。
在本申请的上下文中， 与该杂多酸相关的表述" ##/载体"的含义是
"将 ##负载在该载体上而获得的负载型 ##"。作为此时所使用的载体， 比 如可以举出二氧化硅、 活性炭或其组合。
根据本发明， 作为所述杂多酸， 具体比如可以举出十二磷钨酸或 十二磷钨酸 /载体、 十二硅钨酸或十二硅钨酸 /载体、 十二磷钼酸或十二 磷钼酸 /载体、 十二磷钼钒酸或十二磷钼钒酸 /载体、 前述杂多酸的酸式 盐和前述杂多 酸的酸式盐 /载体， 其中优选酸式磷钨酸铯盐 ( CS2.5H0.5P12WO40 ) 和酸式磷钨酸铯盐 /载体。 这些杂多酸可以单独使 用一种， 也可以两种或多种组合使用。
根据本发明，作为所述沸石分子筛， 比如可以举出 Ηβ沸石分子筛、 氟和 /或磷改性的 Ηβ沸石分子筛、 ΗΥ 沸石分子筛、 氟和 /或磷改性的 ΗΥ 沸石分子筛、 HZSM-5 沸石分子筛或者氟和 /或磷改性的 HZSM-5 沸石分子筛， 其中优选氟和 /或磷改性的 Ηβ沸石分子筛、 氟和 /或磷改 性的 ΗΥ沸石分子筛或者氟和 /或磷改性的 HZSM-5沸石分子筛。 这些 沸石分子筛可以单独使用一种， 也可以两种或多种组合使用。
根据本发明， 在进行所述环己烯酯化步骤时， 所述至少一种羧酸 与以环己烯计的所述环己烯源的摩尔比一般为 0.2至 20: 1 , 优选 1.2至 4: 1 , 更优选 1.2至 3: 1 , 但有时并不限于此。
根据本发明， 作为所述环己烯酯化步骤的进行方式， 比如可以举 出以下的方式 （ 1 ) 和方式 （2 ), 其中从显著提高环己烯的单程转化率 的角度而言， 优选方式 （2 ) 或方式 （ 1 ) 与方式 （2 ) 的组合。
根据方式 （ 1 ), 所述加成酯化反应在一个或多个加成酯化反应器 中进行。 作为所述加成酯化反应器， 比如可以举出釜式反应器、 固定 床反应器、 流化床反应器、 沸腾床反应器或其任意的并联组合， 其中 优选管式固定床反应器， 更优选管壳列管式反应器。
根据该方式 （ 1 ), 所述加成酯化反应器的操作方式既可以是间歇 的方式， 也可以是连续的方式。 由于管式固定床反应器具有制造费用 低、 操作简单等优点， 因此是本发明优选的反应器。 固定床反应器可 采用绝热的或等温方式操作。 绝热反应器可采用筒式反应器， 催化剂 固定在反应器中， 反应器外壁进行保温绝热， 由于加成酯化反应为放 热反应， 因此需要控制反应物浓度以控制反应器床层温升， 或采用部 分反应产物冷却后循环至反应器入口以稀释反应物浓度。 等温反应器 可采用管壳列管式反应器， 催化剂固定在列管中， 在壳程通过冷却水 以移走反应的放出的热量。
根据该方式（ 1 )，所述加成酯化反应的反应温度一般为 50至 200°C， 优选 60至 120°C , 但有时并不限于此。
根据该方式 （ 1 )， 由于加成酯化反应在液相中进行， 因此反应压 力应保证反应处于液相状态。 一般来说， 反应压力为常压- lOMPa, 优 化压力为常压 ~ lMPa, 但有时并不限于此。
根据该方式 （ 1 )， 在所述加成酯化反应按照连续方式进行时， 液 体进料空速一般为 0.5至 20h , 优选 0.5至 511-1 , 更优选 1至 511-1 , 但 有时并不限于此。
根据该方式 （ 1 ), 在所述加成酯化反应按照间歇方式进行时， 反 应时间一般为 0.1 ~ 10h， 优选 0.2 ~ 2h, 但有时并不限于此。
根据该方式 （ 1 ), 所述加成酯化反应的环己烯单程转化率一般能 达到 80 %以上， 而羧酸环己酯比如乙酸环己酯的单程选择性也可达到 99 %以上。
根据该方式 （ 1 ), 所述加成酯化反应获得的加成酯化产物主要含 有未反应的所述至少一种羧酸和环己烯， 以及反应产物羧酸环己酯， 并且还包括来自环己烯源的各种惰性稀释剂比如环己烷和苯等， 这取 决于所述环己烯源的原始组成。 该加成酯化产物可以不经任何分离或 纯化而直接作为加成酯化产物进入本发明的环己酯加氢步骤。
或者优选的是， 根据该方式 （ 1 )， 该加成酯化产物可以通过酯化 产物分离系统进行精馏分离。 此时， 所述加成酯化产物分离系统可按 照本领域已知的方式设置精馏分离单元和 /或萃取精馏分离单元， 而具 体的分离方案则与所述环己烯源的原始组成有关。 一般原则是， 通过 该酯化产物分离系统， 将所述加成酯化产物分离成羧酸物流和羧酸环 己酯物流 （或者羧酸与羧酸环己酯的混合物流）， 以及 C6烃物流。 在 此， 取决于所使用的环己烯源的原始组成， 所述 C6烃可能代表苯、 环 己烷、 环己烯、 环己烷与苯的混合物、 环己烷与环己烯的混合物、 环 己烯与苯的混合物、 或者环己烷、 环己烯与苯的混合物， 任选还可能 含有其他的惰性稀释剂。
具体地说， 该酯化产物分离系统可设置一个或多个脱 C6 烃 /羧酸 塔。 该加成酯化产物进入该脱 C6 烃 /羧酸塔进行分离， 此塔可采用常 压操作， 通过控制塔釜加热量、 回流比、 塔顶和塔釜釆出量， 从塔顶 采出 C6烃与羧酸的混合物流， 从塔釜采出羧酸环己酯物流。 所述获得 的 C6烃与羧酸的混合物流可以通过下述的脱 C6烃塔进一步分离为 C6 烃物流和羧酸物流。 或者， 该酯化产物分离系统可设置一个或多个脱 C6烃塔和一个或多个脱羧酸 （比如乙酸）塔。 该加成酯化产物首先进 入该脱 C6烃塔进行分离 ,此塔可采用常压操作，通过控制塔釜加热量、 回流比、 塔顶和塔釜采出量， 从塔顶采出 C6烃物流， 从脱 C6烃塔塔 釜采出的物流进入脱羧酸塔进行分离， 此塔也可采用常压操作， 通过 控制塔釜加热量、 回流比、 塔顶和塔釜采出量， 从塔顶采出羧酸物流， 从塔釜采出羧酸环己酯物流。 所述酯化产物分离系统可以再设置一个 或多个萃取精镏塔，将获得的 C6烃物流根据情况进一步分离成环己烷 与苯的混合物流、 环己烯与苯的混合物流、 环己烷与环己烯的混合物 流和苯物流等， 或者再设置两个或更多个萃取精馏塔， 将获得的 C6烃 物流根据情况进一步分离成环己烷物流、 环己烯物流和苯物流。 任选 地， 所述酯化产物分离系统还可以再设置一个或多个脱重组分塔， 羧 酸环己酯物流进入该脱重组分塔， 通过精馏分离， 进一步脱除物流中 的重组分， 获得脱除重組分的羧酸环己酯物流， 而分离出来的重组分 作为副产物排出系统。
根据该方式（ 1 ), 如果所述获得的 C6烃物流为混合物， 可以对其 进行加氢制造环己烷， 也可以按照本文前述方式比如采用萃取精馏法， 根据情况从其中分离出环己烯物流、 环己烯与苯的混合物流、 环己烯 与环己烷的混合物流、 苯物流和环己烷物流等。
根据该方式 （ 1 ), 所述获得的羧酸环己酯物流或羧酸与羧酸环己 酯的混合物流 （优选羧酸环己酯物流） 作为所述加成酯化产物进入本 发明的环己酯加氢步骤， 而所述获得的羧酸物流则可以作为该方式（ 1 ) 反应进料的一部分而循环利用。 另外， 所述获得的环己烷与苯的混合 物流或环己烯与苯的混合物流可以进一步采用萃取精馏法分离为环己 烷物流、 苯物流和环己烯物流， 所述获得的环己烯物流、 环己烯与苯 的混合物流或环己烯与环己烷的混合物流可以作为本发明的环己浠酯 化步骤的环己烯源使用， 所述获得的苯物流或环己烯与苯的混合物流 可以作为步骤（A )反应进料的一部分循环利用， 而所述获得的环己烷 物流则作为副产品排出或者作为步骤（ C )反应进料的一部分循环利用。
根据该方式（ 1 ),特别地，在所述环己烯源的环己烯含量为 95mol% 以上或者使用环己烯纯品作为所述环己烯源时， 所述酯化产物分离系 统可只设置一个或多个用于脱除羧酸（比如乙酸）和环己烯的精馏塔； 也可以设置一个或多个用于脱除羧酸和环己烯的精馏塔和一个或多个 用于脱除重组分的精馏塔。 为此， 所述加成酯化产物首先进入脱酸烯 塔进行精馏分离， 此塔可采用常压操作， 通过控制塔釜加热量、 回流 比、 塔顶和塔釜采出量， 使未反应的环己烯和羧酸从塔顶采出， 循环 回反应系统， 从塔釜采出羧酸环己酯。 如从脱酸烯塔塔釜采出的羧酸 环己酯产物中含有较多的重组分， 则羧酸环己酯产物还需要进入脱重 组分塔脱除重组分， 从脱重組分塔塔顶获得高纯度的羧酸环己酯产物， 而塔底重组分作为副产物排出系统。
根据方式 （2 ), 所述加成酯化反应在一个或多个反应精馏塔中进 行。
根据该方式 （2 )， 所述反应精馏塔在形式上与普通精馏塔相同， 一般由塔体、 塔顶冷凝器、 回流罐、 回流泵、 塔釜和再沸器等组成。 塔的类型可以是板式塔， 也可以是填料塔， 还可以是两者的并联组合。 可采用的板式塔类型包括浮岡塔、 筛板塔、 泡罩塔等。 填料塔所使用 的填料可采用散堆填料， 如鲍尔环、 Θ环、 马鞍型填料、 阶梯环填料等； 也可以采用规整填料， 如波纹板填料、 波纹丝网填料等。
根据该方式（2 ), 在反应精馏塔内布置有所述加成酯化催化剂（比 如所述固体酸催化剂）。 本领域技术人员清楚地知道， 反应精馏塔中的 催化剂布置方式应满足以下两点要求： （ 1 ) 要能提供足够的用于汽液 两相通过的通道，或有比较大的床层空隙率（一般要求至少 50 %以上）， 以保证汽液两相能够对流通过， 而不造成液泛； （2 ) 要有良好的传质 性能， 反应物要从流体相传递到催化剂内进行反应， 同时反应产物要 从催化剂中传递出来。 现有文献中已公开多种催化剂在反应精馏塔中 的布置方式， 这些布置方式均可为本发明所采用。 现有催化剂在反应 塔中的布置方式可分为以下三种： （ 1 ) 将催化剂以精馏填料的方式直 接布置在塔中， 主要方式有将一定大小和形状催化剂颗粒与精馏填料 机械混合、 或将催化剂夹在规整填料之间与规整填料组成整体填料， 或将催化剂直接制成精馏填料形状； （2 ) 将催化剂装入气液可透过的 小容器内并将其布置于反应塔的塔板上， 或将催化剂布置在反应塔的 降液管中； （3 ) 将催化剂直接以固定床方式装入反应塔中， 液相直接 流过催化剂床层， 而为气相设立专用的通道， 采用这种方式在装有催 化剂的部位， 由催化剂床层和精馏塔盘交替设置， 塔盘上的液体经降 液管和再分布器进入下一催化剂床层， 在床层中进行加成反应， 催化 剂床层下部的液体通过液体收集器进入下一塔盘。
根据该方式 （2 ), 所述反应精镏塔的理论塔板数一般为 10 - 150, 优选为 30 ~ 100。优选的是， 在所述理论塔板数的 1/3至 2/3位置之间， 比如在 10-120理论塔板之间， 优选在 10-80块理论塔板之间， 比如选 择 5至 30块板， 优选选择 8至 20块板， 布置所述加成酯化催化剂（比 如所述的固体酸催化剂）， 但有时并不限于此。
根据该方式 （2 )， 相对于所述加成酯化催化剂的总装填体积， 液 体进料空速为 0.1至 ΖΟΙι·1 , 优选 0.2至 20h , 更优选 0.5至 5h 或 0.5 至 21T1 , 但有时并不限于此。
根据该方式 （2 ), 反应精馏塔的操作压力可以在负压、 常压和加 压条件下操作。 一般而言， 该反应精馏塔的操作压力为 -0.0099MPa至 5MPa, 优选为常压至 IMPa, 但有时并不限于此。
反应精馏塔的操作温度与反应精馏塔的压力有关， 可通过调节反 应塔的操作压力来调节反应塔的温度分布， 使催化剂装填区的温度在 催化剂的活性温度范围内。 根据该方式 （2 )， 加成酯化催化剂床层装 填区的温度一般为 40至 200°C , 优选 50至 200°C或 50 - 180°C , 更优 选 60至 120°C或 60至 150°C , 但有时并不限于此。
反应精镏塔的回流比应同时满足分离和反应的要求， 一般情况下， 增大回流比有利于提高分离能力和反应转化率， 但同时会增大过程能 耗。 根据该方式 （2 ), 如果使用环己烯纯品和羧酸 （比如乙酸） 作为 反应原料， 理论上可以实现全回流。 但是， 当反应原料中有少量轻组 分杂质时， 需要将少量塔顶物流引出反应精镏塔。
根据本发明， 一般而言， 所述反应精馏塔的回流比一般为 0.1 : 1至 全回流， 优选 0.1至 100: 1 , 更优选 0.5至 10: 1 , 但有时并不限于此。
根据该方式 （2 ), 在进行加成酯化反应的同时， 进行加成酯化产 物的分离。 此时， 所述加成酯化产物主要含有未反应的所述至少一种 羧酸和环己烯， 以及反应产物羧酸环己酯， 并且还包括来自环己烯源 的各种惰性稀释剂比如环己烷和苯等， 这取决于所述环己烯源的原始 组成。 具体而言， 从反应精馏塔的塔底得到羧酸环己酯物流或者羧酸 与羧酸环己酯的混合物流， 并且根据塔底的分离情况， 从反应精馏塔 顶得到 C6烃物流， 或者羧酸与 C6烃的混合物流。 在此， 取决于所使 用的环己烯源的原始组成， 所述 C6烃可能代表苯、 环己烷、 环己烯、 环己烷与苯的混合物、 环己烷与环己烯的混合物、 环己烯与苯的混合 物、 或者环己烷、 环己烯与苯的混合物， 任选还可能含有其他的惰性 稀释剂。 对于羧酸与 C6烃的混合物流， 可采用精馏将其分离成 C6烃 物流和羧酸物流。 具体而言， 将该混合物流通过脱 C6烃塔进行精馏分 离， 此塔可采用常压操作， 通过控制塔釜加热量、 回流比、 塔顶和塔 釜采出量， 从塔顶采出 C6烃物流（比如环己烷物流或者环己烷与苯的 混合物流）， 从塔釜采出羧酸物流。
根据该方式（2 ), 如果所述获得的 C6烃物流为混合物， 可以对其 进行加氢制造环己烷， 也可以按照本文前述方式比如采用萃取精馏法， 根据情况从其中分离出环己烯物流、 环己烯与苯的混合物流、 环己烯 与环己烷的混合物流、 苯物流和环己烷物流等。
根据该方式 （2 ), 所述获得的羧酸环己酯物流或羧酸与羧酸环己 酯的混合物流 （优选羧酸环己酯物流） 作为所述加成酯化产物进入本 发明的环己酯加氢步骤， 而所述获得的羧酸物流则可以作为该方式（ 2 ) 反应进料的一部分而循环利用。 另外， 所述获得的环己烯与环己烷的 混合物流或环己烯与苯的混合物流可以进一步采用萃取精馏法分离为 环己烷物流、 苯物流和环己烯物流， 所述获得的环己烯物流、 环己烯 与苯的混合物流或环己烯与环己烷的混合物流可以作为本发明的环己 烯酯化步骤的环己烯源使用， 所述获得的苯物流或环己烯与苯的混合 物流可以作为步骤（A )反应进料的一部分循环利用， 而所述获得的环 己烷物流则作为副产品排出或者作为步骤（C )反应进料的一部分循环 利用。
根据本发明， 为了进行所述环己烯酯化步骤， 还可以先进行方式 ( 1 ), 然后再进行方式 （2 )。 作为其优选的组合方式， 比如可以举出 先如前所述按照方式 （ 1 ) 进行加成酯化反应 （此时方式 （ 1 ) 的该反 应也称为预酯化反应） 之后， 将所获得的加成酯化产物直接 （或者在 分离出一部分或全部的羧酸环己酯之后） 作为起始原料如前所迷继续 按照方式 （2 ) 进行加成酯化反应 （以下称为方式 （3 ) )。 此时， 方式 ( 1 ) 和方式 （2 ) 分别可以按照如前所述相同的方式进行， 并且各自 所使用的加成酯化催化剂可以相同， 也可以不同， 分别独立地从如前 所述的加成酯化催化剂中选择。
以下对该环己酯加氢步骤进行具体的说明。
根据该环己酯加氢步骤， 使前述获得的 （各种） 加成酯化产物与 氢气在加氢催化剂的存在下发生加氢反应， 同时生成环己醇和链烷醇。
根据本发明， 所述链烷醇用式 R-CH2-OH表示， 其中所述基团 R 与前述针对所述至少一种羧酸的定义相同， 最优选甲基。 作为所述链 烷醇， 最优选乙醇。
根据本发明， 作为所述加氢催化剂， 比如可以举出铜系催化剂、 钌系催化剂和贵金属系催化剂， 其中优选铜系催化剂。 这些催化剂可 以单独使用一种， 也可以两种或多种组合使用。
根据本发明， 作为所述钌系催化剂， 比如可以举出 Ru/Al203和 Ru-Sn/Al203。
根据本发明， 作为所述贵金属系催化剂， 比如可以举出 Pt/Al203、 Pd-Pt/Al203和 Pd /C。
根据本发明， 作为所述铜系催化剂， 比如可以举出含锌的铜系催 化剂和含铬的铜系催化剂， 其中优选包括以下组分 （更优选由以下组 分构成） 的铜系催化剂。
( a ) 氧化铜；
( b ) 氧化锌；
( c ) 选自铝、 镓、 锡、 钛、 锆、 铬、 钼、 钨、 锰、 铼、 镧系金属 和锕系金属中的一种或多种金属的氧化物， 优选选自铝、 镓、 锡、 钛、 锆、 铬、 钼、 钨、 锰、 铼、 镧和铈中的一种或多种金属的氧化物； 和
( d ) 选自碱金属氢氧化物和碱土金属氢氧化物中的一种或多种， 优选选自氢氧化钾、 氢氧化钠和氢氧化钡中的一种或多种。
根据本发明， 优选的是， 在所述铜系催化剂中， 以质量份数计， 组分（a ): 组分（b ): 组分（c ): 组分（d )之比为 5至 60: 10至 50: 5至 60: 0.2至 2, 优选 10至 50: 15至 45: 15至 55: 0.2至 2, 更优 选 30至 45 : 20至 35: 20至 50: 0.5至 1.5。
根据本发明， 该铜系催化剂在使用前一般需要经过还原。 而且， 应当理解到， 在本领域中， 催化剂一般以前体的形式进行交易和贮存， 虽然催化剂前体不能直接催化反应， 然而在习惯上将催化剂前体就称 为"催化剂"。 本发明的该铜系催化剂（实际上是铜系催化剂前体）经过 还原后才具有催化活性， 而这通常由工业装置的操作人员来完成， 本 领域技术人员熟知该还原过程， 本发明在此不再赘述。
根据本发明， 所述铜系催化剂 （前体） 可制成各种需要的形状， 如经过成型的小球， 也可以是成型前的状态， 如粉末。
根据本发明， 所述铜系催化剂可以通过包括以下步骤 （ la ) 和步 骤 （2a ) 的制造方法进行制造。
步骤 （ la ): 通过共沉淀法， 制造复合金属氧化物， 其中所述复合 金属氧化物包括以下组分 （优选由以下组分构成）： ( a ) 氧化铜； （b ) 氧化锌； 和 （c ) 选自铝、 镓、 锡、 钛、 锆、 铬、 钼、 钨、 锰、 铼、 镧 系金属和锕系金属中的一种或多种金属的氧化物， 优选选自铝、 镓、 锡、 钛、 锆、 铬、 钼、 钨、 锰、 铼、 镧和铈中的一种或多种金属的氧 化物， 其中以质量份数计， 组分（a ): 组分 （b ): 组分 （c ) 之比为 5 至 60: 10至 50: 5至 60, 优选 10至 50: 15至 45: 15至 55 , 更优选 30至 45 : 20至 35 : 20至 50。
根据本发明， 共沉淀法是指以均相存在于溶液中两种或多种金属 阳离子与沉淀剂反应， 使溶液中的金属阳离子沉淀下来， 得到各种成 分均一的沉淀， 生成的沉淀混合物或固溶体前驱体， 经过滤、 洗涤、 焙烧（使沉淀混合物或固溶体前驱体热分解）， 得到复合金属氧化物的 方法。
根据本发明， 所述的共沉淀法可以采取不同的方式实现， 既可以 将含有金属阳离子的溶液加入到沉淀剂溶液中， 也可以将沉淀剂溶液 加入到含有金属阳离子的溶液中， 还可以将含有金属阳离子的溶液和 沉淀剂溶液同时加入到溶剂中。
根据本发明， 所述共沉淀法比如可以包括以下步骤：
( I ) 配制并按预定比例混合金属可溶性盐的 （水） 溶液， 所迷金 属指的是（a )铜； （b )锌； 和（c )选自铝、 镓、 锡、 钛、 锆、 铬、 钼、 钨、 锰、 铼、 镧系金属和锕系金属中的一种或多种金属， 优选选自铝、 镓、 锡、 钛、 锆、 铬、 钼、 钨、 锰、 铼、 镧和铈中的一种或多种金属， 所述比例使得换算为相应的金属氧化物， 以质量份数计， 金属 （a ): 金属 （ b ): 金属 （ c ) 之比为 5至 60: 10至 50: 5至 60, 优选 10至 50: 15至 45: 15至 55 , 更优选 30至 45: 20至 35: 20至 50;
( II ) 在 15°C ~ 80°C下， 向步骤 （I ) 获得的混合溶液加入沉淀剂 的 （水） 溶液， 调节 pH值至 6 ~ 9, 生成混合沉淀物， 通过选择所述 的沉淀剂， 使得所生成的混合沉淀物能被热解为金属氧化物； 和
( III )将步骤（ II )得到的沉淀体系， 在 30°C ~ 80°C之间保持 1 ~ 48h后， 过滤混合沉淀物， 并洗涤之， 直至滤液中的金属阳离子浓度小 于 100ug/g,然后在 100 °C ~ 20CTC下干燥 3 ~ 48h,接着在 250°C ~ 400 °C 焙烧 3 ~ 48h， 得到粉末状的复合金属氧化物。
步骤 （I ) 中， 所述的 "可溶性"的含义是指该金属盐在其溶液中的 浓度 （比如在水中的溶解度） 能够满足制造的复合金属氧化物的组成 要求， 比如可以举出所述金属的硝酸盐、 硫酸盐、 盐酸盐、 醋酸盐或 水合物。
步骤 （Π ) 中， 最好在搅拌下， 向步骤 （I ) 获得的混合溶液加入 所述沉淀剂的 （水） 溶液， 这将有利于提高催化剂的均一性。
根据本发明， 作为所述沉淀剂， 比如可以举出氢氧化钠、 氢氧化 钾、 碳酸钠、 碳酸钾、 碳酸铵、 氨、 尿素、 草酸钠、 草酸钾和草酸铵 中的一种或几种。
根据本发明， 对所述金属可溶性盐和沉淀剂在其溶液中的浓度没 有特别的限制， 只要可以实现制造所述复合金属氧化物的目的即可， 本领域技术人员可以根据需要适当选择。
根据本发明， 为了制造各自的溶液， 所述金属可溶性盐和 /或沉淀 剂 （以及如下所述的碱金属氢氧化物和 /或碱土金属氢氧化物） 可以溶 解在水中， 也可以溶解在非水溶剂中， 比如乙醇， 或者水与乙醇的任 意比例混合物， 本领域技术人员可以任意适当选择。
步骤 （2a ): 通过浸渍法， 向步骤 （ la ) 制造的复合金属氧化物中 引入 （d ) 选自碱金属氢氧化物和碱土金属氢氧化物中的一种或多种， 优选选自氢氧化钾、 氢氧化钠和氢氧化钡中的一种或多种， 使得以质 量份数计， 前述的组分 （a ): 组分（b ): 组分 （c ): 组分 （d )之比为 5至 60: 10至 50: 5至 60: 0.2至 2 , 优选 10至 50: 15至 45: 15至 55: 0.2至 2 , 更优选 30至 45: 20至 35: 20至 50: 0.5至 1.5。
根据本发明， 作为所述步骤 （2a ) 的具体实施方式， 比如可以举 出： 将步骤 （ la ) 得到的复合金属氧化物用碱金属氢氧化物和 /或碱土 金属氢氧化物的 （水） 溶液（浓度比如为 0.5 ~ 5wt % ) 浸渍， 再经过 滤、 干燥、 焙烧， 得到本发明的铜系催化剂 （前体）。
根据本发明， 优选的情况下， 步骤 （2a ) 中， 浸渍温度为 30°C ~ 80 °C , 浸渍时间为 l ~ 48h; 干燥温度为 100°C ~ 200°C , 干燥时间为 3 ~ 48h; 焙烧温度为 250°C ~ 400°C , 焙烧时间为 3 ~ 48h。
根据本发明， 所述的碱金属氢氧化物和 /或碱土金属氢氧化物优选 为氢氧化钾、 氢氧化钠和氢氧化钡中的一种或几种。
根据本发明， 所述获得的铜系催化剂为粉末状， 其可以按照用户 的要求成型加工成用户需要的形状。 如前所述， 不管是粉末状的产品， 还是成型加工后的产品， 在氢气气氛中还原后， 该铜系催化剂才能表 现出加氢活性。 在此， 该还原过程既可以是该催化剂使用前的额外步 骤， 也可以在本发明的环己酯加氢步骤进行过程中实现， 其中优选作 为该催化剂使用前的额外步骤来进行。
根据本发明， 所述加氢反应在一个或多个加氢反应器中进行。 作 为所述加氢反应器， 比如可以举出釜式反应器、 固定床反应器、 沸腾 床反应器、 流化床反应器或其任意的并联组合， 其中优选管式固定床 反应器， 更优选管壳列管式反应器。
根据本发明， 所述加氢反应可以以间歇的方式操作， 也可以以连 续的方式进行。 间歇式反应一般采用反应釜作反应器， 将所述加成酯 化产物和加氢催化剂投入反应釜中， 通入氢气在一定的温度和压力下 进行反应， 反应结束后将反应产物采用从釜中卸出， 分离出产物， 再 投入下一批物料进行反应。 连续加氢反应可采用管壳式列管式反应器， 加氢催化剂固定在列管中， 在壳程通过冷却水以移走反应的放出的热 量。
根据本发明， 所述加氢反应的反应温度一般为 150至 400°C , 优选 200至 300 °C。
根据本发明， 所述加氢反应的反应压力一般为常压至 20MPa， 优 选常压至 10MPa， 更优选 4至 10MPa。
根据本发明， 所述加氢反应中， 氢气与以羧酸环己酯计的所述加 成酯化产物的摩尔比为 1至 1000: 1 , 优选 5至 100: 1。
根据本发明， 在所述加氢反应按照连续方式进行时， 液体进料空 速为 0. 1至 20^， 优选 0.2至 2h 。
根据本发明， 在所述加氢反应按照间歇方式进行时， 反应时间为 0.2至 20h, 优选 0.5至 5h, 更优选 1至 5h。
据本发明， 将该加氢反应获得的加氢产物送入加氢产物分离系统 进行分离。 具体而言， 使该加氢产物进入气液分离罐中进行气液分离， 气相主要是氢气， 还可能包括来自环己烯源的各种惰性稀释剂比如环 己烷和苯等， 这取决于环己烯酯化步骤中环己烯源的原始组成。 分离 出来的氢气经压缩机压缩后向所述加氢反应器循环使用。 液相产物主 要含有链烷醇比如乙醇以及环己醇， 还可能含有一定量的羧酸链烷醇 酯 （比如羧酸乙酯） 和环己酮， 同时还可能含有一定量的未反应的羧 酸环己酯， 以及少量重沸物 （二聚酮）。 该液相产物可以采用精馏和 / 或萃取分离的方法进行分离。
根据本发明， 优选用精馏法分离所述加氢产物。 该精馏法可以采 用间歇方案， 也可以采用连续的流程方案， 其中优选采用连续精馏法 来分离所述加氢产物。 该连续精馏法需要利用一系列塔分离各种组分。 可根据各组分的分离的先后顺序设计各种分离流程， 本发明优选顺序 分离的流程方案。
根据本发明， 在进行所述环己酯加氢步骤之前， 任选使所述加成 酯化产物与氢气在羧酸加氢催化剂的存在下发生羧酸加氢反应 （称为 羧酸加氢步骤）， 以将所述加成酯化产物中可能含有的游离羧酸（和可 能存在的苯）预先转化为链烷醇 （和环己烷）。 该羧酸加氢步骤获得的 反应产物 （任选在分离出一部分或全部的链烷醇和 /或环己烷之后）依 然视为所述加成酯化产物， 可以按照与前述完全相同的方式进行所述 环己酯加氢步骤。
根据本发明， 所述羧酸加氢催化剂可以使用本领域在使羧酸加氢 而制造相应醇时常规使用的任何催化剂， 但优选的是， 所述羧酸加氢 催化剂由主活性组分 0.1至 30wt%、助剂 0.1至 25wt%和余量的载体构 成。
根据本发明， 所述主活性组分选自铂、 钯、 钌、 钨、 钼和钴中的 一种或多种。 根据本发明， 所述助剂选自锡、 铬、 铝、 锌、 钙、 镁、 镍、 钛、 锆、 铼、 镧、 钍和金中的一种或多种。
根据本发明， 所述载体选自氧化硅、 氧化铝、 氧化钛、 氧化锆、 活性炭、 石墨、 纳米炭管、 硅酸钙、 沸石和硅酸铝中的一种或多种。
根据本发明， 所述羧酸加氢步骤和所述环己酯加氢步骤可以在不 同的反应器中进行， 也可以在同一反应器的不同区域中进行。 比如， 所述羧酸加氢步骤可以在单独的羧酸加氢反应器中进行。 作为所述羧 酸加氢反应器， 比如可以举出釜式反应器、 固定床反应器、 沸腾床反 应器、 流化床反应器或其任意的并联组合， 其中优选管式固定床反应 器， 更优选管壳列管式反应器。
根据本发明， 所述羧酸加氢反应的反应温度为 100至 400°C , 优选 180至 300。C。
根据本发明， 所述羧酸加氢反应的反应压力为 0.1至 30MPa,优选 2至 10MPa。
根据本发明， 所述羧酸加氢反应中， 氢气与所述游离羧酸的摩尔 比为 1至 500: 1 , 优选 5至 50: 1。
根据本发明， 所述羧酸加氢反应按照连续方式进行时， 液体进料 空速为 0.1至 5h— ] , 优选 0.2至 2h 。
根据本发明， 所述羧酸加氢反应按照间歇方式进行时， 反应时间 为 0.5至 20h, 优选 1至 5h。
根据本发明， 所述联产环己醇和链烷醇的方法还任选进一步包括 下述的步骤 （1 )、 步骤 （II ) 和步骤 （III )之一或其任意的组合。 这些 步骤可以按照本领域常规已知的方式进行， 在此省略其说明。
步骤（ I ): 回收从所述联产环己醇和链烷醇的方法的任何步骤中分 离的苯和 /或氢气， 并将该苯和 /或氢气循环至所述步骤 （A )。
步骤 （II ): 回收从所述联产环己醇和链烷醇的方法的任何步骤中 分离的环己烷， 并将该环己烷循环至所述步骤 （C )。
步骤（III ): 回收从所述联产环己醇和链烷醇的方法的任何步骤中 分离的环己烷， 比如按照本领域常规已知的方式， 使该环己烷在脱氢 催化剂的存在下发生脱氢反应， 以获得苯和氢气。 然后， 将该获得的 苯和 /或氢气循环至所述步骤 （A )。
对于步骤（III ), 环己烷脱氢生成苯的反应相当容易进行， 环己烷 只需要在具有单一脱氢功能的催化剂存在下及适合的反应条件下， 即 可高转化率和高选择性地生成苯。本领域技术人员可参照 JP285001/87、 WO2009/131769, CN1038273 选择适合的实施方法。 当然， 也可以使 用双功能或多功能的催化剂， 如具有脱氢和酸性双功能的催化重整催 化剂。 本发明至少可采用两种方式将环己烷脱氢生成苯， 一种是建立 单独的环己烷脱氢装置， 在单功能或多功能脱氢催化剂存在下， 进行 环己烷脱氢反应； 另一种是利用已有的催化重整装置， 既可对所述环 同处理。 应该理解到， 根据已有的知识， 苯的存在对于环己烷脱氢生 成苯的反应没有不利影响， 因此所述的环己烷中可以含有苯。
根据本发明， 所述联产环己醇和链烷醇的方法还任选进一步包括 下述的步骤或其组合。 这些步骤可以按照本领域常规已知的方式进行， 在此省略其说明。
步骤（IV ): 回收从所述联产环己醇和链烷醇的方法的任何步骤中 分离的羧酸和 /或环己烯源， 并将该羧酸和 /或环己烯源循环至所述环己 烯酯化步骤。
步骤 （v h 回收从所述联产环己醇和链烷醇的方法的任何步骤中 分离的氢气， 并将该氢气循环至所述环己酯加氢步骤。
根据本发明， 以前迷制造的环己醇作为原料， 可以制造环己酮。 因此， 本发明还涉及一种制造环己酮的方法， 包括按照本发明前述的 方法制造环己醇的步骤， 和使用所述环己醇作为起始原料制造环己酮 的步骤。
根据本发明， 以环己醇作为起始原料来制造环己酮的步骤可以按 照本领域常规已知的方式进行， 在此省略其说明。
根据本发明， 以前述制造的环己酮作为原料， 可以制造己内酰胺。 因此， 本发明还涉及一种制造己内酰胺的方法， 包括按照本发明前述 的制造方法制造环己酮的步骤， 和使用所述环己酮作为起始原料制造 己内酰胺的步骤。 ，
根据本发明， 以环己酮作为起始原料来制造己内酰胺的步骤可以 按照本领域常规已知的方式进行， 在此省略其说明。
根据本发明， 还涉及一种联产环己醇和链烷醇的装置， 该装置专 门用于实施本发明的联产环己醇和链烷醇的方法。 根据本发明，所述联产环己醇和链烷醇的装置包括加氢反应单元 A (苯加氢反应器）、任选的加氢产物分离单元 A、加成酯化反应单元（酯 化反应器）、 任选的加成酯化产物分离单元、 加氢反应单元 B (酯加氢 反应器） 和加氢产物分离单元 B (酯加氢产物分离单元）， 其中，
在所述加氢反应单元 A 中， 使苯与氢气在部分加氢催化剂的存在 下发生部分加氢反应， 以获得含有环己烯的加氢产物 （相应于前述的 步骤 （A ) );
在所述加氢产物分离单元 A中， 对来自所述加氢反应单元 A的所 述加氢产物进行分离， 以获得环己烯、 环己烯与苯的混合物或者环己 烯与环己烷的混合物 （相应于前述的步骤 （B ) );
在所述加成酯化反应单元中， 使来自所述加氢反应单元 A的所述 加氢产物和 /或来自所述加氢产物分离单元 A的环己烯、 环己烯与苯的 混合物或者环己烯与环己烷的混合物与羧酸在加成酯化催化剂的存在 下发生加成酯化反应， 生成含有羧酸环己酯的加成酯化产物 （相应于 前述的环己烯酯化步骤）；
在所述加成酯化产物分离单元中， 对来自所述加成酯化反应单元 的所述加成酯化产物进行分离， 以获得羧酸环己酯或者羧酸环己酯与 羧酸的混合物；
在所述加氢反应单元 B 中， 使来自所述加成酯化反应单元的所述 加成酯化产物和 /或来自所述加成酯化产物分离单元的羧酸环己酯或者 羧酸环己酯与羧酸的混合物与氢气在加氢催化剂的存在下发生加氢反 应， 生成含有环己醇和链烷醇的加氢产物 （相应于前述的环己酯加氢 步骤）； 和
在所述加氢产物分离单元 B中， 对来自所述加氢反应单元 B的所 述加氢产物进行分离， 以获得环己醇和链烷醇。
本领域技术人员根据本发明提供的方法， 可容易地确定各单元及 设备间的管线连接方式。
根据本发明优选的是， 所述加氢反应单元 A设置有一个或多个并 联的反应器， 反应器类型选自固定床反应器和 /或釜式反应器。
根据本发明优选的是， 所述加成酯化反应单元设置有一个或多个 并联的反应器 X, 所述反应器 X选自釜式反应器、 固定床反应器、 沸 腾床反应器和流化床反应器中的一种或几种， 用于实施按照前述方式 ( 1 ) 的环己烯酯化步骤。
根据本发明优选的是， 所述加成酯化反应单元设置至少一个反应 精馏塔， 用于实施按照前述方式 （2 ) 的环己烯酯化步骤。
根据本发明优选的是， 所述反应精馏塔前， 串联设置一个或多个 并联的所述反应器 X (也称为预加成酯化反应单元）， 用于实施按照前 述方式 （3 ) 的环己烯酯化步骤。
根据本发明优选的是， 所述的加成酯化产物分离单元设置至少一 个精馏塔。
根据本发明优选的是， 所述的加氢反应单元 B设置一个或多个并 联的反应器， 反应器类型选自釜式反应器、 固定床反应器、 沸腾床反 应器和流化床反应器中的一种或几种， 优选设置一个或多个并联的管 壳列管式反应器。
根据本发明优选的是， 所述的加氢产物分离单元 B设置至少一个 精馏塔。
根据本发明， 所述联产环己醇和链烷醇的装置还任选包括至少一 个下述的循环装置。
循环装置 A: 回收出自所述联产环己醇和链烷醇的装置的任何单 元的苯和 /或氢气， 并将该苯和 /或氢气循环至所述加氢反应单元 A。
循环装置 B：回收出自所述联产环己醇和链烷醇的装置的任何单元 的羧酸和 /或环己烯， 并将该羧酸和 /或环己烯循环至所述加成酯化反应 单元。
循环装置 C：回收出自所述联产环己醇和链烷醇的装置的任何单元 的氢气， 并将该氢气循环至所述加氢反应单元 B。
根据本发明， 在所述联产环己醇和链烷醇的装置中， 所述加成酯 化反应单元选自反应器 （比如釜式反应器、 固定床反应器、 流化床反 应器和沸腾床反应器之一或其任意的组合） 和反应精馏塔 （理论塔板 数比如为 10至 150， 优选 30至 100， 其中优选板式塔或填料塔）之一 或其任意的组合， 优选所述反应精馏塔（用于实施按照前述方式 （2 ) 的环己烯酯化步骤） 或者所述反应器与所述反应精馏塔的串联组合， 更优选所述反应器与所述反应精馏塔的串联组合， 最优选所述反应器 (此时也称为预加成酯化反应器） 处于所述反应精馏塔上游的串联组 合 （用于实施按照前述方式 （3 ) 的环己烯酯化步骤）。 以下结合附图简要说明本发明的具体实施方案， 但本发明并不限 于这些具体实施方案。
根据本申请发明， 各附图中的附图标记是彼此独立的， 即每个附 图使用自身独立编号的附图标记。 除非明确指出， 同一附图标记在不 同附图之间可能不具有同一技术含义或者不代表同一技术特征。 本领 域技术人员完全可以参考以下针对每个附图提供的简要说明， 清楚地 理解该附图所代表的具体实施方案， 而不必拘泥于该附图标记的编号 方式或者不同附图之间是否使用了相同或不同的附图标记。
根据图 2的实施方案， 环己烯和乙酸进入酯化反应器 1 , 在加成酯 化催化剂的作用下， 进行加成酯化反应， 加成酯化产物物流经管线 11 进入酯加氢反应器 2， 在酯加氢催化剂的作用下， 与氢气接触进行酯加 氢反应， 酯加氢产物物流 22进入酯加氢产物分离单元 3 , 经过分离， 得到环己醇物流 31、 乙醇物流 32、 高沸物物流 33、 乙酸环己酯物流 34 , 环己醇物流 31和乙醇物流 32作为产品出装置， 高沸物物流 33作 为副产品出装置， 乙酸环己酯物流 34循环回酯加氢反应器 2。
根据图 3 的实施方案， 含苯和环己烷的物流与乙酸进入反应精镏 塔 1 , 在加成酯化催化剂的作用下， 进行加成酯化反应， 同时进行反应 产物的分离， 塔精馏段得到苯 +环己烷物流 13和乙酸物流 14 , 乙酸物 流循环回反应精馏塔 1 , 塔底得到乙酸环己酯物流， 经管线 12进入酯 加氢反应器 2 ,在酯加氢催化剂的作用下，与氢气接触进行酯加氢反应， 酯加氢产物物流 22进入酯加氢产物分离单元 3 , 经过分离， 得到环己 醇物流 31、 乙醇物流 32、 高沸物物流 33和乙酸环己酯物流 34 , 环己 醇物流 31和乙醇物流 32作为产品出装置， 高沸物物流 33作为副产品 出装置， 乙酸环己酯物流 34循环回酯加氢反应器 2。
根据图 4 的实施方案， 环己烯和乙酸进入加成酯化反应系统， 在 加成酯化催化剂的作用下， 进行酯化反应， 反应产物送入酯化产物分 离系统， 经分离得到乙酸环己酯物流、 环己烯物流和乙酸物流， 环己 烯物流和乙酸物流循环回加成酯化反应系统， 乙酸环己酯物流进入酯 加氢反应系统， 在酯加氢催化剂的作用下， 与氢气接触， 发生乙酸环 己酯的加氢反应， 加氢产物送入加氢产物分离系统， 分离后得到环己 醇物流、 乙醇物流、 乙酸环己酯物流和氢气物流， 环己醇物流和乙醇 物流作为产品出装置， 乙酸环己酯物流和氢气物流循环回酯加氢反应 系统。
根据图 5的实施方案， 苯和氢气进入苯加氢反应器 1 , 在苯加氢催 化剂的作用下， 进行加氢反应， 苯加氢产物物流经管线 11进入加成酯 化反应器 2， 与经管线 21进入的乙酸混合， 在固体酸催化剂的作用下， 进行加成酯化反应， 加成酯化产物物流经管线 22进入加氢反应器 3 , 在贵金属催化剂的作用下， 与氢气接触， 同时进行苯和羧酸的加氢反 应， 加氢产物物流进入加氢反应器 4 , 在酯加氢催化剂的作用下， 与氢 气接触进行酯加氢反应， 加氢产物物流进入产物分离单元 5 , 经过分离 得到环己醇物流 51、 乙醇物流 52 , 以及环己烷物流 53。
根据图 6的实施方案， 苯和氢气进入苯加氢反应器 1 , 在苯加氢催 化剂的作用下， 进行加氢反应， 苯加氢产物物流经管线 11进入加成酯 化反应器 2 , 与经管线 21进入的乙酸混合， 在固体酸催化剂的作用下， 进行加成酯化反应， 加成酯化产物物流经管线 22进入加成酯化产物分 离单元 3 , 经过分离得到环己烷物流 31、 环己烯物流 32、 苯物流 33、 乙酸 /乙酸环己酯物流 34 , 苯物流循环回苯加氢反应器 1， 环己烯物流 循环回加成酯化反应器 2 , 环己烷物作为副产品流出装置， 乙酸 /乙酸 环己酯物流进入羧酸加氢反应器 4， 在羧酸加氢催化剂的作用下， 与氢 气接触进行羧酸加氢反应， 羧酸加氢产物物流进入酯加氢反应器 5， 在 酯加氢催化剂的作用下， 与氢气接触进行酯加氢反应， 酯加氢反应产 物物流进入加氢产物分离单元 6 , 经过分离得到环己醇物流 61、 乙醇 物流 62。
根据图 7的实施方案， 苯和氢气进入苯加氢反应器 1， 在苯选择性 加氢催化剂的作用下， 进行加氢反应， 苯加氢产物物流经管线 11进入 反应精馏塔 2 , 与经管线 21进入的乙酸混合， 在固体酸催化剂的作用 下， 进行加成酯化反应， 同时进行酯化产物的分离， 从反应精馏塔 2 塔顶得到环己烷 /苯物流 , 从反应精馏塔 2塔底得到乙酸 /乙酸环己酯物 流； 环己烷 /苯物流经管线 22进入加成酯化产物分离单元 3 , 经过分离 得到环己烷物流和苯物流， 苯物流循环回苯加氢反应器 1， 环己烷物流 作为副产品出装置， 乙酸 /乙酸环己酯物流进入羧酸加氢反应器 4 , 在 羧酸加氢催化剂的作用下， 与氢气接触进行羧酸加氢反应， 羧酸加氢 产物物流进入酯加氢反应器 5， 在酯加氢催化剂的作用下， 与氢气接触 进行酯加氢反应， 酯加氢反应产物物流进入加氢产物分离单元 6 , 经过 分离得到环己醇物流 61、 乙醇物流 62。
根据图 8的实施方案， 苯和氢气进入苯加氢反应器 1 , 在苯加氢催 化剂的作用下， 进行加氢反应， 苯加氢产物物流经管线 11进入加成酯 化反应器 2 , 与经管线 21进入的乙酸混合， 在固体酸催化剂的作用下， 进行加成酯化反应， 加成酯化产物物流经管线 22进入加成酯化产物分 离单元 3 , 经过分离得到环己烷物流 31、 环己烯物流 32、 苯物流 33、 乙酸物流 34和乙酸环己酯物流 35 , 苯物流循环回苯加氢反应器 1 , 环 己烯物流和乙酸物流循环回加成酯化反应器 2 , 乙酸环己酯物流进入酯 加氢反应器 4，在酯加氢催化剂的作用下，与氢气接触进行酯加氢反应， 酯加氢产物物流进入酯加氢产物分离单元 5 ,经过分离得到环己醇物流 52、 乙醇物流 53、 乙酸环己酯物流 51 和高沸物物流 54 , 环己醇物流 和乙醇物流作为产品出装置， 高沸物物流作为副产品出装置， 乙酸环 己酯物流循环回酯加氢反应器 4。
根据图 9的实施方案， 苯和氢气进入苯加氢反应器 1， 在苯加氢催 化剂的作用下， 进行加氢反应， 苯加氢产物物流经管线 11进入反应精 镏塔 2 , 与经管线 21进入的乙酸混合， 在固体酸催化剂的作用下， 进 行加成酯化反应， 同时进行加成酯化产物的分离， 从反应精馏塔 2 塔 顶得到环己烷 /苯 /乙酸物流，从反应精馏塔 2塔底得到乙酸环己酯物流； 环己烷 /苯 /乙酸物流经管线 22进入加成酯化产物分离单元 3 ,经过分离 得到环己烷物流、 苯物流和乙酸物流， 环己烷物流作为副产品出装置， 苯物流循环回苯加氢反应器 1， 乙酸物流循环回加成酯化反应器 2 , 乙 酸环己酯物流进入酯加氢反应器 4 , 在酯加氢催化剂的作用下， 与氢气 接触进行酯加氢反应， 酯加氢产物物流进入酯加氢产物分离单元 5 , 经 过分离得到环己醇物流 52、 乙醇物流 53、 乙酸环己酯物流 51 和高沸 物物流 54 , 环己醇物流和乙醇物流作为产品出装置， 高沸物物流作为 副产品出装置， 乙酸环己酯物流循环回酯加氢反应器 4。
根据图 10的实施方案， 苯和氢气进入苯加氢反应器 1 , 在苯加氢 催化剂的作用下， 进行加氢反应， 苯加氢产物物流经管线 11进入苯加 氢产物分离单元 2 , 经过分离得到环己烷物流 21和环己烯 /苯物流 22， 环己烷物流 21作为副产品出装置， 环己浠 /苯物流经管线 22进入加成 酯化反应器 3， 与经管线 31进入的乙酸混合， 在固体酸催化剂的作用 下， 进行加成酯化反应， 加成酯化产物物流经管线 32进入加成酯化产 物分离单元 4 , 经过分离得到环己烯 /苯物流 41、 乙酸物流 42和乙酸环 己酯物流 43 , 环己烯 /苯物流 41循环回苯加氢反应器 1或者进一步分 离成环己烯物流和苯物流， 乙酸物流 42循环回加成酯化反应器 3 , 乙 酸环己酯物流 43进入酯加氢反应器 5 , 在酯加氢催化剂的作用下， 与 氢气 51接触进行酯加氢反应， 酯加氢产物物流 52进入酯加氢产物分 离单元 6， 经过分离得到环己醇物流 62、 乙醇物流 63、 乙酸环己酯物 流 61和高沸物物流 64 ,环己醇物流 62和乙醇物流 63作为产品出装置， 高沸物物流 64作为副产品出装置， 乙酸环己酯物流 61 循环回酯加氢 反应器 5。
根据图 11 的实施方案， 苯和氢气进入苯加氢反应器 1 , 在苯加氢 催化剂的作用下， 进行加氢反应， 苯加氢产物物流经管线 11进入苯加 氢产物分离单元 2 , 经过分离得到环己烷物流 21和环己烯 /苯物流 22， 环己烷物流作为副产品出装置，环己烯 /苯物流经管线 22进入反应精馏 塔 3 , 与经管线 31进入的乙酸混合， 在固体酸催化剂的作用下， 进行 加成酯化反应， 同时进行加成酯化产物的分离， 从反应精馏塔 3 塔顶 得到乙酸 /苯物流， 从反应精馏塔 3 塔底得到乙酸环己酯物流， 乙酸 / 苯物流经管线 33进入加成酯化产物分离单元 4 , 经过分离得到苯物流 41和乙酸物流 42， 苯物流 41循环回苯加氢反应器 1， 乙酸物流 42循 环回加成酯化反应器 3， 乙酸环己酯物流 32进入酯加氢反应器 5 , 在 酯加氢催化剂的作用下， 与氢气接触进行酯加氢反应， 酯加氢产物物 流 52进入酯加氢产物分离单元 6， 经过分离得到环己醇物流 62、 乙醇 物流 63、 乙酸环己酯物流 61 和高沸物物流 64 , 环己醇物流和乙醇物 流作为产品出装置， 高沸物物流作为副产品出装置， 乙酸环己酯物流 循环回酯加氢反应器 5。
根据图 12的实施方案， 苯和氢气进入苯加氢反应器 1 , 在苯加氢 催化剂的作用下， 进行加氢反应， 苯加氢产物物流经管线 11进入苯加 氢产物分离单元 2， 经过分离， 得到环己烯 /环己烷物流和苯物流， 苯 物流经管线 21 循环回苯加氢反应器 1， 环己烯 /环己烷物流经管线 22 进入加成酯化反应器 3 , 与经管线 31进入的乙酸混合， 在固体酸催化 剂的作用下， 进行加成酯化反应， 加成酯化产物物流经管线 32进入加 成酯化产物分离单元 4 , 经过分离， 得到环己烷物流 41、 环己烯物流 42、 乙酸物流 43和乙酸环己酯物流 44 , 环己烷物流 41作为副产品出 装置， 乙酸物流 43和环己烯物流 42循环回加成酯化反应器 3 , 乙酸环 己酯物流 44进入酯加氢反应器 5 , 在酯加氢催化剂的作用下， 与氢气 接触进行酯加氢反应，酯加氢产物物流 51进入酯加氢产物分离单元 6 , 经过分离， 得到环己醇物流 62、 乙醇物流 63、 乙酸环己酯物流 61 和 高沸物物流 64 , 环己醇物流 62和乙醇物流 63作为产品出装置， 高沸 物物流 64作为副产品出装置， 乙酸环己酯物流 61循环回酯加氢反应 器 5。
根据图 13的实施方案， 苯和氢气进入苯加氢反应器 1 , 在苯加氢 催化剂的作用下， 进行加氢反应， 苯加氢产物物流经管线 11进入苯加 氢产物分离单元 2 , 经过分离， 得到环己烯 /环己烷物流和苯物流， 苯 物流经管线 21 循环回苯加氢反应器 1 , 环己烯 /环己烷物流经管线 22 进入反应精馏塔 3， 与经管线 31进入的乙酸混合， 在固体酸催化剂的 作用下， 进行加成酯化反应， 同时进行加成酯化产物的分离， 从反应 精馏塔 3塔顶得到乙酸 /环己烷物流， 从反应精馏塔 3塔底得到乙酸环 己酯物流， 乙酸 /环己烷物流经管线 32进入加成酯化产物分离单元 4 , 经过分离， 得到环己烷物流 42和乙酸物流 41 , 环己烷物流 42作为副 产品出装置， 乙酸物流 41循环回加成酯化反应器 3 , 乙酸环己酯物流 33 进入酯加氢反应器 5 , 在酯加氢催化剂的作用下， 与氢气接触进行 酯加氢反应， 酯加氢产物物流进入酯加氢产物分离单元 6 , 经过分离， 得到环己醇物流 62、 乙醇物流 63、 乙酸环己酯物流 61 和高沸物物流 64 , 环己醇物流和乙醇物流作为产品出装置， 高沸物物流作为副产品 出装置， 乙酸环己酯物流循环回酯加氢反应器 5。
根据图 14的实施方案， 苯和氢气进入苯加氢反应器 1 , 在苯加氢 催化剂的作用下， 进行加氢反应， 苯加氢产物物流经管线 11进入苯加 氢产物分离单元 2 , 经过分离得到环己烷物流和环己烯物流和苯物流， 环己烷物流 21 作为副产品出装置， 苯物流经管线 22循环回苯加氢反 应器 1， 环己烯物流经管线 23进入加成酯化反应器 3 , 与经管线 31进 入的乙酸混合， 在固体酸催化剂的作用下， 进行加成酯化反应， 加成 酯化产物物流经管线 32进入加成酯化产物分离单元 4 , 经过分离得到 环己烯 /乙酸物流 41和乙酸环己酯物流 42 , 环己烯 /乙酸物流循环回加 成酯化反应器 3 , 乙酸环己酯物流进入酯加氢反应器 5 , 在酯加氢催化 剂的作用下， 与氢气接触进行酯加氢反应， 酯加氢产物物流进入酯加 氢产物分离单元 6， 经过分离得到乙醇物流 63、 环己醇物流 62、 乙酸 环己酯物流 61和高沸物物流 64 ,环己醇物流和乙醇物流作为产品出装 置， 高沸物物流作为副产品出装置， 乙酸环己酯物流循环回酯加氢反 应器 5。
根据图 15的实施方案， 苯和氢气进入苯加氢反应器 1， 在苯加氢 催化剂的作用下， 进行加氢反应， 苯加氢产物物流经管线 11进入苯加 氢产物分离单元 2 , 经过分离得到环己烷物流、 环己烯物流和苯物流， 环己烷物 21流作为副产品出装置， 苯物流 22循环回苯加氢反应器 1 , 环己烯物流经管线 23进入反应精馏塔 3 ,与经管线 31进入的乙酸混合， 在固体酸催化剂的作用下， 进行加成酯化反应， 同时进行加成酯化产 物的分离，从反应精馏塔 3塔顶引出少量物流（带出原料中的轻杂质）， 从反应精馏塔 3 塔底得到乙酸环己酯物流， 乙酸环己酯物流进入酯加 氢反应器 4 , 在酯加氢催化剂的作用下， 与氢气接触进行酯加氢反应， 酯加氢产物物流进入酯加氢产物分离单元 5 , 经过分离得到乙醇物流 53、 环己醇物流 52、 乙酸环己酯物流 51 和高沸物物流 54 , 环己醇物 流和乙醇物流作为产品出装置， 高沸物物流作为副产品出装置， 乙酸 环己酯物流循环回酯加氢反应器 5。
根据图 16的实施方案， 苯和氢气进入苯加氢反应器 1 , 在苯选择 性加氢催化剂的作用下， 进行加氢反应， 苯加氢产物物流经管线 11进 入加成酯化反应器 2， 与经管线 21进入的乙酸混合， 在加成酯化固体 酸催化剂的作用下， 进行加成酯化反应， 加成酯化产物物流经管线 22 进入加成酯化产物分离单元 3 , 经过分离得到环己烷物流 31、 苯物流 32、 乙酸物流 33和乙酸环己酯物流 34，苯物流循环回苯加氢反应器 1 , 乙酸物流循环回加成酯化反应器 2 , 乙酸环己酯物流进入酯加氢反应器 4 , 在酯加氢催化剂的作用下， 与氢气接触进行酯加氢反应， 酯加氢产 物物流进入酯加氢产物分离单元 5 , 经过分离得到环己醇物流 51、 乙 醇物流 52、 高沸物物流 53和乙酸环己酯物流 54 , 环己醇物流和乙醇 物流作为产品出装置， 高沸物物流作为副产品出装置， 乙酸环己酯物 流循环回酯加氢反应器 4。 实施例
以下采用实施例进一步详细地说明本发明， 但本发明并不限于这 些实施例。
第一种实施方案
实施例 1 ~6 (催化剂的制造）
实施例 1 ~ 6的催化剂按下列程序进行制造： 按表 1 配方称取一 定量的可溶性金属盐， 置于 2000mL 三口烧瓶中， 加水溶解配制成约 lOOOmL溶液， 在烧瓶上装上搅拌器、 pH计和温度计， 并将烧瓶置于 温度可调的恒温水浴中， 开启搅拌， 调节恒温水浴温度,将一定浓度的 沉淀剂溶液逐渐滴入烧瓶中， 控制沉淀剂水溶液的滴加速度， 使溶液 的温度升高控制在 1 °C以内。 随溶液 pH升高， 溶液出现沉淀， 并随 pH 升高逐渐增多， 当溶液 pH到达规定值时停止滴加沉淀剂水溶液。 然后 在继续搅拌的条件下保持一定温度老化一定时间。 停止搅拌， 自然冷 却到室温， 将沉淀在高速离心机上离心过滤， 并用去离子水洗涤 5次， 将所得沉淀在烘箱中烘干， 转移到马福炉中进行焙烧， 得到混合金属 氧化物。 将此金属氧化物用一定浓度的碱溶液在室温下浸渍,经真空过 滤， 脱除浸渍液， 将混合物在烘箱中烘干， 转移到马福炉中焙烧， 最 终得到混合金属氧化物。 采用 ICP 法分析所得样品的组成。 具体制造 条件和结果见表 1。
实施例 7~15 (高压釜评价催化剂）
实施例 7 ~ 15为在高压釜中进行实施例 1 ~ 6所制得催化剂的乙酸 环己酯加氢试验， 试验程序如下： 取一定量得催化剂粉末置于 500mL 高压反应釜中， 加入 250g乙酸环己酯， 将反应釜封闭， 用氮气置换三 次， 通入氢气至一定压力， 逐渐升温， 在大约 80°C时釜内压力开始下 降， 表明釜内催化剂开始还原， 并开始进行酯加氢反应， 及时补充氢 气使反应釜维持一定压力， 最后升温到给定温度,并在此温度下维持压 力反应一定时间后， 停止反应， 降温到室温后， 卸出反应产物和催化 剂。 用气相色谱分析产物组成， 并根据分析结果按下列公式计算乙酸 环己酯单程转化率和环己醇的单程选择性。
乙酸环己酯单程转化率 =[ 1 -未反应醋酸环己酯摩尔数 /(未反应乙酸 环己酯摩尔数 +环己烷摩尔数 +环己醇摩尔数 + 乙基环己醚摩尔 数] x l 00%
环己醇单程选择性 = [环己醇摩尔数 /(环己醇摩尔数 +环己烷摩尔 数 +乙基环己基醚摩尔数 )]χ 100% 实施例 16 (固定床中进行酯加氢）
将实施例 3所得到的催化剂粉末压片成型， 并敲碎筛分 40 ~ 60 目 颗粒， 取 40g此催化剂颗粒装入 (|)20x2.5x800mm带有夹套的不锈钢管 反应器中的中部， 两端填充一定量的石英砂。 通入氢气 （ 500mL/min ) 在 280 °C、 6MPa条件下还原 24h后， 降至反应的温度。 将乙酸环己酯 由计量泵打入反应器中， 氢气经质量流量控制器进入反应系统进行加 氢反应， 通过反应管外部夹套中通入导热油控制反应温度， 通过反应 器出口背压阀控制反应器压力。 反应产物通过反应器后部的在线取样 阀取样进行在线色谱分析。 反应条件和结果见表 3。 结果显示， 乙酸环 己酯加氢反应单程转化率最高可达到 99.5%以上，酯产物单程选择性大 于 99.0 %， 运行 1000小时， 单程转化率和单程选择性下降不明显。
对比例 1 ~ 2
按实施例 1 配方制造催化剂， 但不用 NaOH溶液处理， 且不进行 第二次烘干和焙烧处理。 制造条件和所制得催化剂组成见表 1 中的对 比例 1。
将上述方法所制得的催化剂采用高压反应釜进行乙酸环己酯加氢 评价。 评价条件和结果见表 2中的对比例 2。
结果表明， 催化剂不用碱处理， 生成较多的环己烷和乙基环己醚 副产物， 加氢反应单程选择性相对较差。
表 1 催化剂制造结果汇总
Figure imgf000039_0001
高压釜催化剂评价结果汇总
Figure imgf000040_0001
实施例 16固定床乙酸环己酯加氢数据
Figure imgf000041_0001
第二种实施方案
实施例 1
本实施例用来说明以环己烯和乙酸为原料， 采用固定床反应器制 备乙酸环己酯的结果。
固定床反应器为 (|)48x4x l200mm的 316L不锈钢管， 反应管外部带 有热水夹套， 可在夹套中通入热水以控制反应温度。 将 500mL大孔强 酸性氢型离子交换树脂 （实验室按经典的文献方法合成， 将含有 15 % 二乙烯基苯的苯乙烯溶液进行悬浮共聚制成白球， 再经浓硫酸磺化制 得， 测得其交换容量为 5.2mmolH+/g干基） 装入固定床反应器的中部， 两端填充一定量的石英砂。 将环己烯和乙酸分别由计量泵打入固定床 反应器中反应， 通过调节固定床反应器夹套热水温度调节反应温度， 稳定操作条件的反应条件和反应结果见表 1。 表 1 大孔强酸性氢型离子交换树脂催化的固定床试验数据
Figure imgf000042_0001
实施例 2
本实施例说明对乙酸与乙酸环己酯的混合物物流进行加氢的结 果。
反应系统由单个固定床反应器组成， 反应器为带有夹套的钛质钢 管， 尺寸为 (|)20x2.5x800mm。 催化剂分两层装入反应器。 上层装入 20g 二氧化硅负载的铂钯锡乙酸加氢催化剂（实验室合成， 组成为 Pt(10%)-Pd(5%)-Sn(5%)/SiO2, 由 20~40 目的大孔二氧化硅载体 （BET 比表面积 400m2/g, 孔容 0.35mL/g )浸渍氯铂酸、 氯化钯和氯化亚锡的 混合容易， 再经 120 °C干燥， 500 °C焙烧制得)； 下层装入 20g铜辞铝酯 加氢催化剂（实验室合成， 组成为 CuO 40%, ZnO 29.6 % , A1203 30.4%。 由铜、 锌、 铬的硝酸盐溶液， 加入氢氧化钠溶液中和至 PH - 9.0, 经离 心分离， 洗涤， 干燥， 压片成型， 焙烧制得)。 催化剂装入反应器的中 部恒温区， 两层催化剂间由玻璃纤维布隔开， 反应器两端填充一定量 的石英沙， 作为原料加热气化区或填料。 反应器夹套中可通入导热油 控制反应温度。 反应器中装入催化剂后， 连接反应系统， 并完成系统 气密试验后， 通入氢气（ 500mL/min )在 300°C、 6MPa条件下还原 24h 后， 降至加氢反应的温度和压力。 将乙酸和乙酸环己酯的混合物 （实 施例 1 的反应器出口物流） 由计量泵打入反应器中， 氢气经质量流量 控制器进入反应系统进行加氢反应， 通过反应管外部夹套中通入导热 油控制反应温度， 通过反应器出口背压阀控制反应器压力。 反应产物 通过反应器后部的直线取样阀取样进行在线色语分析。 反应条件和结 果见表 2。 表 2乙酸与乙酸环己酯的混合物物流加氢的试验数据
Figure imgf000043_0001
第三种实施方案
实施例 1
本实施例用来说明以环己烯和乙酸为原料， 采用反应精馏塔制备 乙酸环己酯的结果。
试验是在如下规格的反应精馏模式试验装置进行的： 模式装置的 主体为直径 （内径） 为 50mm, 高为 3m的不锈钢塔， 塔的下部连接体 积为 5L的塔釜， 釜内配置有 10KW的电加热棒， 此加热棒由智能控制 器通过可控硅（SCR )控制塔釜加热量。 塔顶连接有换热面积为 0.5m2 的冷凝器， 塔顶蒸汽经此冷凝器冷凝成液体后进入一个体积为 2L的回 流罐。 回流罐中的液体经回流泵部分回流至反应塔， 部分采出作为轻 组分。 塔的操作参数由智能型自动化控制仪表显示和控制。 塔回流量 由回流调节阀控制， 塔顶采出量由回流罐的液位控制器控制。 塔釜采 出量由塔釜液位控制器调节塔釜排料阀进行控制。
将耐高温磺酸型离子交换树脂 （牌号为 Amberlyst 45 , 由 Rhom&Hass 公司生产） 用多级高速粉碎机粉碎成粒度小于 200 目 ( 0.074mm )的粉料， 加入制孔剂、 润滑剂、 抗氧剂和粘合剂在高速混 合机上混合均匀，再在密炼机上于 180°C密炼 lOmin,使物料完全塑化， 之后注入模具中制成直径为 5mm, 高 5mm， 壁厚为 1mm拉西环型树 脂催化剂填料。 将此填料 1950mL装入模式反应塔的中部 （高 lm, 相 当于 8块理论塔板）上下各装入直径为 3mm、 长 6mm的玻璃弹簧填料 1950mL ( 装填高度为 lm, 相当于 10块理论塔板)。 将环己烯和乙酸分 别由计量泵打入预热器预热后进入反应塔， 调节塔釜加热量和塔顶回 流量连续进行反应， 稳定操作下的反应条件和反应结果见表 1。
表 1 耐高温磺酸型离子交换树脂催化剂催化的反应精馏试验数据
Figure imgf000045_0001
实施例 2
本实施例说明对乙酸与乙酸环己酯的混合物物流进行加氢的结 果。
反应系统由单个固定床反应器组成， 反应器为带有夹套的钛质钢 管， 尺寸为 (|)20x2.5x800mm。 催化剂分两层装入反应器。 上层装入 20g 二氧化硅负载的铂钯锡乙酸加氢催化剂（实验室合成， 組成为 Pt( 10m%)-Pd(5m%)-Sn(5m%)/ Si02 , 由 20~40 目的大孔二氧化硅载体 ( BET 比表面积 400m2/g, 孔容 0.35mL/g ) 浸溃氯铂酸、 氯化钯和氯 化亚锡的混合溶液， 再经 120°C干燥， 500Ό焙烧制得）； 下层装入 20g 铜锌铝酯加氢催化剂（实验室合成， 组成为 CuO 40m%, ZnO 29.6m % , Α1203 30.4m%。 由铜、 锌、 铬的硝酸盐溶液， 加入氢氧化钠溶液中和 至 PH = 9.0 , 经离心分离， 洗涤， 干燥， 压片成型， 焙烧制得)。 催化 剂装入反应器的中部恒温区， 两层催化剂间由玻璃纤维布隔开， 反应 器两端填充一定量的石英沙， 作为原料加热气化区或填料。 反应器夹 套中可通入导热油控制反应温度。 反应器中装入催化剂后， 连接反应 系统，并完成系统气密试验后，通入氢气（ 500mL/min )在 300°C 、 6MPa 条件下还原 24h后， 降至加氢反应的温度和压力。 将乙酸和乙酸环己 酯的混合物 （实施例 1 的塔底物流） 由计量泵打入反应器中， 氢气经 质量流量控制器进入反应系统进行加氢反应， 通过反应管外部夹套中 通入导热油控制反应温度， 通过反应器出口背压阀控制反应器压力。 条件和结果见表 口 ^ 、 ， ， 、 乙酸与乙酸环己酯的混合物物流加氢的试^数据
Figure imgf000046_0001
第四种实施方案
实施例 1 ~ 4用于说明采用反应精镏制造乙酸环己酯的方法。
实施例 1 ~ 4中所进行的试验均是在:^下规格的反应精熘模式试验 装置进行的： 模式装置的主体为直径 （内径） 为 50mm， 高为 3m的不 锈钢塔， 塔的下部连接体积为 5L的塔釜， 釜内配置有 10KW的电加热 棒， 此加热棒由智能控制器通过可控硅（SCR )控制塔釜加热量。 塔顶 连接有换热面积为 0.5m2的冷凝器，塔顶蒸汽经此冷凝器冷凝成液体后 进入一个体积为 2L的回流罐。 回流罐中的液体经回流泵部分回流至反 应塔， 部分采出作为轻组分。 塔的操作参数由智能型自动化控制仪表 显示和控制。 塔回流量由回流调节阀控制， 塔顶采出量由回流罐的液 位控制器控制。 塔釜采出量由塔釜液位控制器调节塔釜排料阀进行控 制。 原料乙酸和环己烯分别装入 30L储罐中， 并通过计量泵打入相应 的预热器中预热到一定温度后进入反应塔， 进料速度由计量泵控制、 电子秤精确计量。
实施例 1
将耐高温磺酸型离子交换树脂 （牌号为 Amberlyst 45 , 由 Rhom&Hass 公司生产） 用多级高速粉碎机粉碎成粒度小于 200 目 ( 0.074mm )的粉料， 加入制孔剂、 润滑剂、 抗氧剂和粘合剂在高速混 合机上混合均匀，再在密炼机上于 180°C密炼 lOmin,使物料完全塑化， 之后注入模具中制成直径为 5mm, 高 5mm, 壁厚为 1mm拉西环型树 脂催化剂填料。 将此填料 1950mL装入模式反应塔的中部 （高 lm, 相 当于 8块理论塔板）上下各装入直径为 3mm、长 6mm的玻璃弹簧填料 1950mL ( 装填高度为 lm, 相当于 10块理论塔板）。 将环己烯和乙酸分 别由计量泵打入预热器预热后进入反应塔， 调节塔釜加热量和塔顶回 流量连续进行反应， 稳定操作下的反应条件和反应结果见表 1。
实施例 2
反应塔和催化剂的配置与实施例 1 相同。 只是用环己烯、 环己烷 和苯混合物代替环己烯进行试验， 且反应塔在 0.3MPa条件下操作。 调 节塔釜加热量和塔顶回流量连续进行反应， 稳定操作下的反应条件和 反应结果见表 2。
实施例 3
将 φ3 ~ 4的球型 H0.5Cs2.5PW12O4。/SiO2催化剂 (由 H0 5Cs2 5PW12O40 粉末和粒度小于 200 目的粗孔硅胶粉末， 在混料机中充分混合后， 在 糖衣机中以硅溶胶为粘合机滚球成型， 再经烘干、 焙烧而成） 夹入钛 丝网波板中， 制成直径为 50mm、 高 50mm的圆柱型规整填料。 将此填 料型催化剂 L装入模式反应塔的中部（高 lm, 相当于 12块理论塔板） 上下各装入直径为 4mm、 高为 4mm的 1950mL玻璃弹簧填料（ 装填高 度为 lm, 相当于 15块理论塔板)。 将环己烯和苯分别由计量泵打入预 热器预热后进入反应塔， 调节塔釜加热量和塔顶回流量连续进行反应， 稳定操作下的反应条件和反应结果见表 3。
实施例 4
反应塔和催化剂的配置与实施例 3 相同。 只是用环己烯、 环己烷 和苯混合物代替环己烯进行试验， 反应塔在 0.2MPa条件下操作。 调节 塔釜加热量和塔顶回流量连续进行反应， 稳定操作下的反应条件和反 应结果见表 4。 实施例 5 ~ 8用于说明采用预酯化和反应精馏制造乙酸环己酯的方 法。
实施例 5 ~ 8 所进行的试验均是在乙酸环己酯模式试验装置进行 的。 该模式装置由固定床预酯化反应器和反应精馏酯化塔组成。 预酯 化反应器为 (|)48x4x l200mm的 316L不锈钢管， 反应管外部带有热水夹 套， 可在夹套中通入热水以控制反应温度。 反应精馏酯化塔为直径（内 径） 为 50mm、 高为 3m 的钛钢 （ TA2 )塔。 塔的下部连接体积为 5L 的塔釜， 釜内配置有 10KW的电加热棒， 此加热棒由智能控制器通过 可控硅（SCR )控制塔釜加热量。 塔顶连接有换热面积为 0.5m2的冷凝 器和体积为 2L的回流罐。 原料醋酸和环己烯分别装入 30L储罐中， 并 通过计量泵打入到预酯化反应器中进行反应， 预酯化产物进入反应精 馏塔进一步进行反应。 通过调节塔釜的加热功率调节反应塔的加热量。 通过塔顶回流比调节器调节塔的回流比。 从塔顶采出轻组分。 从塔底 采出乙酸环己酯产物。
实施例 5
将 500mL大孔强酸性氢型离子交换树脂 （实验室按经典的文献方 法合成，将含有 15 %二乙烯基苯的苯乙烯溶液进行悬浮共聚制成白球， 再经浓硫酸磺化制得， 测得其交换容量为 5.2mmolH+/g 干基） 装入预 反应器的中部， 两端填充一定量的石英砂。 另将耐高温磺酸型离子交 换树脂 （牌号为 Amberlyst 45， 由 Rhom&Hass公司生产 ) 用多级高速 粉碎机粉碎成粒度小于 200 目 （ 0.074mm )的粉料， 加入制孔剂、 润滑 剂、 抗氧剂和粘合剂在高速混合机上混合均勾， 再在密炼机上于 180°C 密炼 lOmin, 使物料完全塑化， 之后注入模具中制成直径为 5mm, 高 5mm, 壁厚为 lmm拉西环型树脂催化剂填料。 将此填料 1950mL装入 模式反应塔的中部 （高 lm, 相当于 8块理论塔板） 上下各装入直径为 3mm, 长 6mm的玻璃弹簧填料 1950mL ( 装填高度为 lm， 相当于 10 块理论塔板）。 将环己烯和乙酸分别由计量泵打入预反应器中反应， 预 反应产物在进入反应塔进一步进行反应。 通过调节预反应器夹套热水 温度调节预反应温度。 调节塔釜加热量和塔顶回流量连续进行反应， 稳定操作条件的反应条件和反应结果见表 5。
实施例 6
反应塔和催化剂的配置与实施例 5 相同。 只是用环己烯、 环己烷 和苯混合物代替环己烯进行试验， 且预反应器压力为 2.0MPa, 反应塔 在常压条件下操作。 调节塔釜加热量和塔顶回流量连续进行反应， 稳 定操作条件的反应条件和反应结果见表 6。
实施例 Ί
将 500mL φ3 ~ 4的球型 H。 5CS2.5PW1204Q/SI02催化剂装入预反应器 的 中部 ， 两端填充一定量的石 英沙。 另 将 φ3 ~ 4 的球型 H0 5Cs2 5PW12O40/SiO2催化剂 (由 ¾.5Cs2.5PW1204o粉末和粒度小于 200 目的粗孔硅胶粉末， 在混料机中充分混合后， 在糖衣机中以硅溶胶为 粘合机滚球成型， 再经供干、 焙烧而成） 夹入钛丝网波板中， 制成直 径为 50mm、 高 50mm的圆柱型规整填料。 将此填料型催化剂 L装入 模式反应塔的中部 （高 lm, 相当于 12块理论塔板） 上下各装入直径 为 4mm、 高为 4mm的 1950mL玻璃弹簧填料（ 装填高度为 lm, 相当 于 15 块理论塔板)。 将环己烯和乙酸分别由计量泵打入预反应器中反 应， 预反应产物在进入反应塔进一步进行反应。 通过调节预反应器夹 套热水温度调节预反应温度。 调节塔釜加热量和塔顶回流量连续进行 反应， 稳定操作条件的反应条件和反应结果见表 7。
实施例 8
反应塔和催化剂的配置与实施例 7相同。 只是用环己烯、 环己烷 和苯混合物代替环己烯进行试验，预反应压力 2.0MPa反应塔在 0.2MPa 条件下操作。 调节塔釜加热量和塔顶回流量连续进行反应， 稳定操作 条件的反应条件和反应结果见表 8。
实施例 9 ~ 10用于说明乙酸环己酯的加氢方法。
实施例 9
加氢原料为純度为 99.6%的乙酸环己酯。
将 40g铜锌铝酯加氢催化剂（实验室合成，组成为 CuO 40.5%， ZnO
29.6 % , Α1203 30.4%。 由铜、 锌、 铬的硝酸盐溶液， 加入氢氧化钠溶 液中和至 ΡΗ = 9.0 , 经离心分离， 洗涤， 干燥， 压片成型， 焙烧制得） 装入 (|)20x2.5x800mm 带有夹套的不锈钢管反应器中的中部， 两端填充 一定量的石英沙。 通入氢气 （ 500mL/min ) 在 280°C、 6MPa条件下还 原 24h后， 降至加氢反应的温度和压力。 将乙酸环己酯由计量泵打入 反应器中， 氢气经质量流量控制器进入反应系统进行加氢反应， 通过 反应管外部夹套中通入导热油控制反应温度， 通过反应器出口背压阀 线色谱分析。 反应条件和结果见表 9。 表 9结果显示， 采用铜锌铝催化 剂， 乙酸环己酯加氢反应单程转化率最高可达到 99.0%以上， 环己醇单 程选择性大于 99.9 % , 运行 1000小时， 单程转化率和单程选择性均未 下降。
实施例 10
加氢原料为纯度为 99.6%的乙酸环己酯。
将 40g铜铬酯加氢催化剂（市售，太原市欣吉达化工有限公司生产， 牌号为 C 1 -XH-1 , CuO含量为 55%， 直径 5mm片剂， 破碎成 10 ~ 20 目颗粒） 装入 (|)20x2.5x800mm 带有夹套的不锈钢管反应器中的中部， 两端填充一定量的石英沙。 通入氢气 （ 500mL/min ) 在 280 °C、 6MPa 条件下还原 24h后， 降至反应得温度和压力。 将乙酸环己酯由计量泵 打入反应器中， 氢气经质量流量控制器进入反应系统进行加氢反应， 通过反应管外部夹套中通入导热油控制反应温度， 通过反应器出口背 压阀控制反应器压力。 反应产物通过反应器后部的直线取样阀取样进 行在线色谱分析。 反应条件和结果见表 10。 表 10结果显示， 采用铜锌 铝催化剂， 乙酸环己酯加氢反应单程转化率最高可达到 98.0%以上， 环 己醇单程选择性大于 99.9 %， 运行 500 小时， 单程转化率和选择均未 下降。
实施例 1 1
收集实施例 9 ~ 10的反应产物 4000g, 进行精馏分离试验。 精馏采 用高 2m玻璃塔， 塔柱装有 0>3mm的不锈钢 Θ网环高效精馏填料， 塔 釜为 5L玻璃烧瓶，通过电热套进行加热，通过调压器调节塔釜加热量。 塔的回流釆用回流比调节器进行控制。 精馏分离结果见表 1 1。 操作奈件
操作压力 常压
塔顶温度 117°C
催化剂段温度 120~ 145 °C
塔釜温度 184°C
回流比 2
物料 环己烯进料 乙酸进料 塔顶采出 塔釜采出 位置 催化剂层下端 催化剂层上端 塔顶回流泵出口 塔釜出料口 量 g/h 411 601 303 709 温度 °C 75 75 40 184 组成 （质量分数）
环己烯 100% 1.3%
苯
环己烷
乙酸 100% 98.7% 0.42% 乙酸环己酯 99.30% 叠合物 0.28% 根据试验数据计算环己烯的单程转化率 99% , 乙酸环己酯单程选 择性 99.72%。
表 2
Figure imgf000052_0001
根据试验数据计算环己烯的单程转化率 98.8% , 乙酸环己酯单程 选择性 98.0%。
表 3
Figure imgf000053_0001
根据试验数据计算环己烯的单程转化率 98.7% , 乙酸环己酯单程 选择性 99.43%。
表 4
Figure imgf000054_0001
根据试验数据计算环己烯的单程转化率 99.35% , 乙酸环己酯单程 选择性 99.6%。
预反应器操作条件
操作压力 常压
夹套热水温度 90 °C
催化剂床层温度 90 ~ 95 °C
反应塔操作条件
操作压力 常压
塔顶温度 1 17°C
催化剂段温度 130 ~ 165 °C
塔釜温度 184°C
回流比 1.5
物料 预反应器进料 预反应器出口 塔顶采出 塔釜采出
(反应塔入口）
流量 g/h 2433 2433 308 2125 温度 °C 25 95
组成（质量分数）
环己烯 50.6% 5.75% 0.97%
苯
环己烷
乙酸 49.4% 12.78% 99.03% 0.28% 乙酸环己酯 77.15% 99.15% 叠合物 0.32% 0.56% 根据试验数据计算环己烯的单程转化率 99.76 % , 乙酸环己酯单程 选择性 99.03 % 。
表 6
Figure imgf000056_0001
根据试验数据计算环己烯的单程转化率 98.38% , 乙酸环己酯单程 选择性 99.11 %。
表 7
Figure imgf000057_0001
根据试验数据计算环己烯的单程转化率 99.9% , 乙酸环己酯单程 选择性 99.35 %。
表 8
Figure imgf000058_0001
根据试验数据计算环己烯的单程转化率 99.02% , 乙酸环己酯单程 选择性 99.19%。
表 9 铜锌铝催化剂环己酯加氢数据
Figure imgf000059_0001
铜铬催化剂环己酯加氢试验数据
Figure imgf000060_0001
乙酸环己酯加氢产物精馏分离结果
Figure imgf000060_0002
实施例 12
本实施例用于说明环己烷催化脱氢制苯的试验结果。
催化脱氢反应的原料为实施例 4塔顶物流经蒸馏得到的 C6烃混合 物， 该混合物经水洗除去少量乙酸后用气相色语分析， 含有 67.5m%的 环己烷、 32.3m%的苯和 0.2m%的环己烯。
反应器采用管式固定床反应器， 反应器为带有夹套的钛质钢管， 尺寸为 φ20χ2.5χ800 mm。 催化剂采用负载型铂铑催化剂 （ Pt 含量 0.3m%, Rh含量 0.1m% ) 。 反应条件为： 温度 480°C， 压力 0.7MPa， 重时空速 5h-l。 经在线色谱分析， 反应原料中的环己烷定量转化为苯。
第五种实施方案
实施例 1
将 lOOmL大孔强酸性氢型离子交换树脂 （实验室按经典的文献方 法合成，将含有 15 %二乙烯基苯的苯乙烯溶液进行悬浮共聚制成白球， 再经浓硫酸磺化制得， 测得其交换容量为 5.2mmolH+/g 干基） 装入 (j)32x4x l000mm带有夹套的不锈钢管反应器中的中部，两端填充一定量 的石英沙。 将乙酸和环己烯按一定的流量分别由计量泵打入反应器中 进行反应， 在反应管外部夹套中通入热水以控制反应温度， 通过反应 器出口背压阀控制反应器压力。 反应器出口产物由在线取样阔取样进 行在线色谱分析。 反应条件和结果见表 1。
由表 1可见， 采用强酸型离子交换树脂催化剂环己烯与乙酸反应， 环己烯单程转化率大于 90%， 酯产物单程选择性大于 99 % , 运行 600 小时， 催化剂活性和单程选择性稳定不变。
实施例 2
采用实施例 1 中的试验装置和方法进行乙酸与环己烯酯化试验， 所不同的是以 Cs2.5H0.5PW12O40/SiO2为催化剂（记为 PW/Si02, 下同）。 反应条件和结果见表 2。 由表 2可见， 环己烯与乙酸反应单程转化率可 达 95%， 酯产物单程选择性大于 99 % , 运行 480小时， 催化剂活性和 单程选择性稳定不变。
实施例 3
试验装置和方法同实施例 1 , 所不同的是催化剂为磷改性的 Ηβ分 子筛催化剂 （由硅铝比为 50的 Ηβ分子筛经 85%的磷酸改性， 再与氧 化铝捏合挤条成型， 经 120°C烘干， 500°C焙烧制得， 磷含量为 2% ) 。 反应条件和结果见表 3。 由表 3 可见， 环己浠与乙酸反应单程转化率 90%, 酯产物单程选择性大于 99 % , 运行 480 小时， 催化剂活性和单 程选择性稳定不变。
实施例 4
收集实施例 1 ~ 3的加成酯化产物， 进行精馏分离试验。 精馏采用 高 2m直径为 40 mm的玻璃塔精馏装置， 塔柱装有 D3mm的不锈钢 Θ 网环高效精馏填料， 塔釜为体积 5的 L玻璃烧瓶， 装料量为 4L, 通过 电热套对塔釜进行加热， 通过调压器调节塔釜加热量。 塔的回流采用 回流比调节器进行控制。 精馏分离结果见表 4。
实施例 5
试验装置、 催化剂和方法同实施例 1， 不同的是环己烯原料 （苯 53.3%, 环己烯 35.4%， 环己烷 1 1.3% ) 。 反应条件和结果见表 5。 由 表 5 可知， 采用强酸型离子交换树脂催化剂催化环己烯原料与乙酸反 应， 环己烯单程转化率大于 80%， 酯产物单程选择性大于 99 %， 运行 600小时， 催化剂活性和单程选择性稳定不变。
实施例 6
加氢原料为纯度 99.6%的乙酸环己酯。
将 40g铜锌铝酯加氢催化剂（实验室合成，组成为 CuO 40.5%, ZnO 29.6 % , Α1203 30.4%。 由铜、 辞、 铝的硝酸盐溶液， 加入氢氧化钠溶 液中和至 ΡΗ = 9.0, 经离心分离， 洗涤， 干燥， 压片成型， 焙烧制得) 装入 ())20x2.5x800mm 带有夹套的不锈钢管反应器中的中部， 两端填充 一定量的石英沙。 通入氢气 （ 500mL/min ) 在 280°C、 6MPa条件下还 原 24h后， 降至加氢反应的温度和压力。 将乙酸环己酯由计量泵打入 反应器中， 氢气经质量流量控制器进入反应系统进行加氢反应， 通过 反应管外部夹套中通入导热油控制反应温度， 通过反应器出口背压阀 线色谱分析。 反应条件和结果见表 6。 表 6结果显示， 采用铜辞铝催化 剂， 乙酸环己酯加氢反应单程转化率最高可达到 99%以上， 环己醇单 程选择性大于 99.9 %， 运行 1000小时， 单程转化率和单程选择性均未 下降。
实施例 7
加氢原料为纯度 99.6%的乙酸环己酯。
将 40g铜铬酯加氢催化剂（市售，太原市欣吉达化工有限公司生产， 牌号为 Cl-XH-1 , CuO含量为 55%, 直径 5mm片剂， 破碎成 10 ~ 20 目颗粒） 装入 (|)20x2.5x800mm 带有夹套的不锈钢管反应器中的中部， 两端填充一定量的石英沙。 通入氢气 （ 500mL/min ) 在 280° ( 、 6MPa 条件下还原 24h后， 降至反应得温度和压力。 将乙酸环己酯由计量泵 打入反应器中， 氢气经质量流量控制器进入反应系统进行加氢反应， 通过反应管外部夹套中通入导热油控制反应温度， 通过反应器出口背 压阀控制反应器压力。 反应产物通过反应器后部的直线取样阀取样进 行在线色谱分析。 反应条件和结果见表 7。 表 7结果显示， 采用铜锌铝 催化剂， 乙酸环己酯加氢反应单程转化率可达 98%以上， 环己醇单程 选择性大于 99.9 % , 运行 500小时， 单程转化率和选择均未下降。
实施例 8
收集实施例 6 - 7的反应产物 4000g， 进行精馏分离试验。 精镏采 用高 2m玻璃塔， 塔柱装有 0>3mm的不锈钢 Θ网环高效精馏填料， 塔 釜为 5L玻璃烧瓶，通过电热套进行加热，通过调压器调节塔釜加热量。 塔的回流采用回流比调节器进行控制。 精镏分离结果见表 8。 强酸性离子交换树脂催化乙酸与环己烯酯化试验数据
Figure imgf000063_0001
表 2 Cs25H。.5PW1204。/Si02催化乙酸与环己烯酯化试验数据
Figure imgf000064_0001
Ηβ催化剂乙酸与环己烯酯化试验数据
Figure imgf000065_0001
表 4 加成酯化产物精馏分离试验结果
Figure imgf000065_0002
表 5 强酸性离子交换树脂催化乙酸与环己烷 /环己烯 /苯酯化试验 数据
Figure imgf000066_0001
铜锌铝催化剂催化乙酸环己酯加氢数据
Figure imgf000067_0001
铜铬催化剂催化乙酸环己酯加氢试验数据
Figure imgf000068_0001
表 8 乙酸环己酯加氢产物精馏分离试验结果
Figure imgf000068_0002
第六种实施方案
实施例 1
本实施例说明苯选择性加氢制环己烯的方法。
将苯和氢气按摩尔比 1 :3注入装填有钌颗粒催化剂的加氢反应器， 在反应温度 135 °C、 压力 4.5MPa、 停留时间 15min的条件下进行苯加 氢反应， 反应产物分离出氢气后， 收集液体产物， 连续运行 1000h。 试 验结束后，对收集的液体产物进行气相色谱分析，其组成为：苯 53.3%, 环己烯 35.4%, 环己烷 11.3%。
实施例 2
将 lOOmL大孔强酸性氢型离子交换树脂 （实验室按经典的文献方 法合成，将含有 15 %二乙烯基苯的苯乙烯溶液进行悬浮共聚制成白球， 再经浓硫酸磺化制得， 测得其交换容量为 5.2mmolH+/g 干基） 装入 φ32χ4χ 1000ηιιη带有夹套的不锈钢管反应器中的中部，两端填充一定量 的石英沙。 将乙酸和环己烯原料 （苯 53.3%, 环己烯 35.4% , 环己烷 1 1.3% )按一定的流量分别由计量泵打入反应器中进行反应， 在反应管 外部夹套中通入热水以控制反应温度， 通过反应器出口背压阀控制反 应器压力。反应器出口产物由在线取样阀取样进行在线色语分析， 由产 物组成计算出环己烯单程转化率和乙酸环己酯单程选择性。 反应条件 和结果见表 1。
由表 1 可知， 采用强酸型离子交换树脂催化剂催化环己烯原料与 乙酸反应， 环己烯单程转化率大于 80%, 酯产物单程选择性大于 99 % , 运行 600小时， 催化剂活性和单程选择性稳定不变。
强酸性离子交换树脂催化乙酸与环己烷 /环己烯 /苯酯化试验
Figure imgf000070_0001
实施例 3
试验装置、 方法和原料同实施例 2 , 所不同的是催化剂为磷改性的 Ηβ分子筛催化剂 （由硅铝比为 50的 Ηβ分子筛经 85%的磷酸改性， 再 与氧化铝捏合挤条成型，经 120°C烘干， 500 °C焙烧制得，磷含量为 2% )。 反应条件和结果见表 2。 由表 2可见， 环己烯与乙酸反应单程转化率大 于 80%， 酯产物单程选择性大于 99 % , 运行 480小时， 催化剂活性和 单程选择性稳定不变。 表 2 Ηβ分子筛催化剂催化乙酸与环己烷 /环己烯 /苯酯化试验数 据
Figure imgf000071_0001
实施例 4
本实施例说明酯化反应混合物同时加氢的方法。
加氢原料为含苯和环己烷的环己烯与醋酸酯化反应的混合物 （环 己烷 7.4 %， 苯 35.4 %, 环己烯 4.6%, 乙酸 20.5%, 乙酸环己酯 32.2%, 聚合物 0.2% ) 。
反应系统由两个串联的固定床反应器组成。 两个反应器均为带有 夹套的钛质钢管， 尺寸为 (|)20x2.5x800mm。 前反应器为乙酸、 苯和环己 烯加氢反应器， 内装 40g二氧化硅负载的铂钯锡乙酸加氢催化剂（实猃 室合成， 组成为 Pt(10m%)-Pd(5m%)-Sn(5m%)/SiO2， 由 20~40 目的大孔 二氧化硅载体 （BET比表面积 400m2/g， 孔容 0.35mL/g ) 浸渍氯铂酸、 氯化钯和氯化亚锡的混合溶液， 再经 120°C干燥， 500°C焙烧制得）。 后 反应器为酯加氢反应器， 内装 40g铜锌铝酯加氢催化剂（实验室合成， 组成为 CuO 40%, ZnO 29.6 % , Α1203 30.4%。 由铜、 锌、 铬的硝酸盐 溶液， 加入氢氧化钠溶液中和至 ΡΗ = 9.0， 经离心分离， 洗涤， 干燥， 压片成型， 焙烧制得)。 催化剂装入反应器的中部， 两端填充一定量的 石英沙， 作为原料加热气化区或填料。 反应器夹套中可通入导热油控 制反应温度。 反应器中装好催化剂后， 连接反应系统， 并完成系统气 密试验后， 通入氢气 （ lOOOmL/min ) 在 300°C、 6MPa条件下还原 24h 后， 降至加氢反应的温度和压力。 将乙酸和乙酸环己酯的混合物由计 量泵打入反应器中， 氢气经质量流量控制器进入反应系统进行加氢反 应， 通过反应管外部夹套中通入导热油控制反应温度， 通过反应器出 口背压阀控制反应器压力。 反应产物通过反应器后部的直线取样阔取 样进行在线色谱分析。 反应条件和结果见表 3。
表 3 加氢试验数据
Figure imgf000073_0001
实施例 5
本实施例说明酯化反应混合物同时加氢的方法。
加氢原料为含苯和环己烷的环己烯与醋酸酯化反应的混合物， 组 成环己烷 7.4 %, 苯 35.4 %, 环己烯 4.6%, 乙酸 20.5%, 乙酸环己酯 32.2%,聚合物 0.2% ) 。
反应系统由单个固定床反应器组成， 反应器为带有夹套的钛质钢 管， 尺寸为 (J)20x2.5x800mm。 催化剂分两层装入反应器。 上层装入 20g 二氧化硅负载的铂钯锡乙酸加氢催化剂（实验室合成， 组成为 Pt( 10%)-Pd(5%)-Sn(5%)/SiO2, 由 20~40 目的大孔二氧化硅载体 （BET 比表面积 400m2/g, 孔容 0.35mL/g )浸渍氯铂酸、 氯化钯和氯化亚锡的 混合容易， 再经 120°C干燥， 500°C焙烧制得)； 下层装入 20g铜锌铝酯 加氢催化剂（实验室合成，组成为 CuO 40.5%, ZnO 29.6 % , A1203 30.4%。 由铜、 锌、 铬的硝酸盐溶液， 加入氢氧化钠溶液中和至 PH = 9.0 , 经离 心分离， 洗涤， 干燥， 压片成型， 焙烧制得)。 催化剂装入反应器的中 部恒温区， 两层催化剂间由玻璃纤维布隔开， 反应器两端填充一定量 的石英沙， 作为原料加热气化区或填料。 反应器夹套中可通入导热油 控制反应温度。 反应器中装入催化剂后， 连接反应系统， 并完成系统 气密试臉后， 通入氢气（ 500mL/min )在 300°C、 6MPa条件下还原 24h 后， 降至加氢反应的温度和压力。 将乙酸和乙酸环己酯的混合物由计 量泵打入反应器中， 氢气经质量流量控制器进入反应系统进行加氢反 应， 通过反应管外部夹套中通入导热油控制反应温度， 通过反应器出 口背压阀控制反应器压力。 反应产物通过反应器后部的直线取样阀取 样进行在线色谱分析。 反应条件和结果见表 4。
表 4 加氢试验数据
Figure imgf000075_0001
实施例 6
收集实施例 5 ~ 6的反应产物 4000g(气相色谱分析组成： 乙醇 27.4 m%, 乙酸乙酯 0.2 m%， 环己烷 40.2m%, 水 6.4 m%， 乙酸 0.2m%， 环己醇 25.0 m%, 乙酸环己酯 0.3m%, 其它 0.3 m% ) , 进行精镏分离 试验。 精馏采用高 2m玻璃塔， 塔柱装有 Φ3πιηι的不锈钢 Θ网环高效 精馏填料， 塔釜为 5L玻璃烧瓶， 通过电热套进行加热， 通过调压器调 节塔釜加热量。 塔的回流采用回流比调节器进行控制。 经分离， 得到 环己醇产品 845g, 气相色谱分析纯度为 99.4%。
第七种实施方案
实施例 1 本实施例说明苯选择性加氢制环己烯的方法。
将苯和氢气按摩尔比 1 :3注入装填有钌颗粒催化剂的加氢反应器， 在反应温度 135 °C、 压力 4.5MPa、 停留时间 15min的条件下进行苯加 氢反应， 反应产物分离出氢气后， 收集液体产物， 连续运行 1000h。 试 验结束后，对收集的液体产物进行气相色谱分析，其组成为：苯 53.3%, 环己烯 35.4%, 环己烷 11.3%。
实施例 2
将 lOOmL大孔强酸性氢型离子交换树脂 （实验室按经典的文献方 法合成，将含有 15 %二乙烯基苯的苯乙烯溶液进行悬浮共聚制成白球， 再经浓硫酸磺化制得， 测得其交换容量为 5.2mmolH+/g 干基） 装入 φ32χ4χ 1000πιιη带有夹套的不锈钢管反应器中的中部，两端填充一定量 的石英沙。 将乙酸和环己烯原料 （苯 53.3%, 环己烯 35.4% , 环己烷 1 1.3% )按一定的流量分别由计量泵打入反应器中进行反应， 在反应管 外部夹套中通入热水以控制反应温度， 通过反应器出口背压阀控制反 应器压力。反应器出口产物由在线取样阀取样进行在线色谱分析， 由产 物组成计算出环己烯单程转化率和乙酸环己酯单程选择性。 反应条件 和结果见表 1。
由表 1 可知， 采用强酸型离子交换树脂催化剂催化环己烯原料与 乙酸反应， 环己烯单程转化率大于 80%， 酯产物单程选择性大于 99 % , 运行 600小时， 催化剂活性和单程选择性稳定不变。
强酸性离子交换树脂催化乙酸与环己烷 /环己烯 /苯酯化试验
Figure imgf000077_0001
实施例 3
试验装置、 方法和原料同实施例 2 , 所不同的是催化剂为磷改性的 Ηβ分子筛催化剂 （由硅铝比为 50的 Ηβ分子筛经 85%的磷酸改性， 再 与氧化铝捏合挤条成型， 经 120°C烘干， 500°C焙烧制得， 磷含量为 2m% )。 反应条件和结果见表 2。 由表 2可见， 环己烯与乙酸反应单程 转化率大于 80% , 酯产物单程选择性大于 99 % , 运行 480小时， 催化 剂活' ί和单程选择性稳定不变 表 2 Ηβ分子筛催化剂催化乙酸与环己烷 /环己烯 /苯酯化试验数 据
Figure imgf000078_0001
实施例 4
收集实施例 2和 3 的加成酯化产物， 进行精馏分离试验。 精馏采 用高 2m直径为 40 mm的玻璃塔精馏装置， 塔柱装有 3mm的不锈钢 Θ网环高效精馏填料， 塔釜为体积 5的 L玻璃烧瓶， 装料量为 4L， 通 过电热套对塔釜进行加热， 通过调压器调节塔釜加热量。 塔的回流采 用回流比调节器进行控制。 精馏分离结果见表 3。 加成酯化产物精馏分离试验结果
Figure imgf000079_0001
实施例 5 ~ 6用于说明釆用反应精馏制造乙酸环己酯的方法。
实施例 5 ~ 6中所进行的试验均是在如下规格的反应精馏模式试验 装置进行的： 模式装置的主体为直径 （内径） 为 50mm, 高为 3m的不 锈钢塔， 塔的下部连接体积为 5L的塔釜， 釜内配置有 10KW的电加热 棒， 此加热棒由智能控制器通过可控硅（SCR )控制塔釜加热量。 塔顶 连接有换热面积为 0.5m2的冷凝器，塔顶蒸汽经此冷凝器冷凝成液体后 进入一个体积为 2L的回流罐。 回流罐中的液体经回流泵部分回流至反 应塔， 部分采出作为轻组分。 塔的操作参数由智能型自动化控制仪表 显示和控制。 塔回流量由回流调节阀控制， 塔顶采出量由回流罐的液 位控制器控制。 塔釜采出量由塔釜液位控制器调节塔釜排料阀进行控 制。 乙酸和环己烯原料分别装入 30L储罐中， 并通过计量泵打入相应 的预热器中预热到一定温度后进入反应塔， 进料速度由计量泵控制、 电子秤精确计量。
实施例 5
将耐高温磺酸型离子交换树脂 （牌号为 Amberlyst 45 , 由 Rhom&Hass 公司生产） 用多级高速粉碎机粉碎成粒度小于 200 目 ( 0.074mm )的粉料， 加入制孔剂、 润滑剂、 抗氧剂和粘合剂在高速混 合机上混合均匀，再在密炼机上于 180°C密炼 lOmin,使物料完全塑化， 之后注入模具中制成直径为 5mm, 高 5mm, 壁厚为 1mm拉西环型树 脂催化剂填料。 将此填料 1950mL装入模式反应塔的中部 （高 lm, 相 当于 8块理论塔板）上下各装入直径为 3mm、 长 6mm的玻璃弹簧填料 1950mL ( 装填高度为 lm, 相当于 10块理论塔板)。将环己烯原料和乙 酸分别由计量泵打入预热器预热后进入反应塔， 调节塔釜加热量和塔 顶回流量连续进行反应， 稳定操作下的反应条件和反应结果见表 4。 表 4 耐高温磺酸型离子交换树脂催化剂的反应精馏试验数据
Figure imgf000080_0001
根据试验数据计算环己烯的单程转化率 99 %， 乙酸环己酯单程选 择性 99.2 % 。
实施例 6
将 φ3 ~ 4的球型 H0.5Cs2.5PW12O40/SiO2催化剂（由 H0.5Cs2.5PW12O40 粉末和粒度小于 200 目的粗孔硅胶粉末， 在混料机中充分混合后， 在 糖衣机中以硅溶胶为粘合机滚球成型， 再经烘干、 焙烧而成） 夹入钛 丝网波板中， 制成直径为 50mm、 高 50mm的圆柱型规整填料。 将此填 料型催化剂 L装入模式反应塔的中部（高 lm， 相当于 12块理论塔板） 上下各装入直径为 4mm、 高为 4mm的 1950mL玻璃弹簧填料（ 装填高 度为 lm, 相当于 15块理论塔板)。 将环己烯原料和乙酸分别由计量泵 打入预热器预热后进入反应塔， 调节塔釜加热量和塔顶回流量连续进 行反应， 稳定操作下的反应条件和反应结果见表 5。 表 5 H 5Cs2.5PW1204。/Si02催化剂的反应精馏试验数据
Figure imgf000081_0001
根据试验数据计算环己烯的单程转化率 98.3 %， 乙酸环己酯单程 选择性 99.5 % 。
实施例 7
本实施例说明乙酸和乙酸环己酯混合物的加氢方法。
加氢原料为乙酸与乙酸环己酯的混合物（乙酸 39.5%， 乙酸环己酯
60.5% ) 。
反应系统由两个串联的固定床反应器组成。 两个反应器均为带有 夹套的钛质钢管，尺寸为 <l)20x2.5x800mm。前反应器为乙酸加氢反应器， 内装 40g二氧化硅负载的铂钯锡乙酸加氢催化剂（实验室合成， 组成为 Pt(10m%)-Pd(5m%)-Sn(5m%)/SiO2 , 由 20~40 目的大孔二氧化硅载体 ( BET 比表面积 400m2/g, 孔容 0.35mL/g ) 浸溃氯铂酸、 氯化钯和氯 化亚锡的混合溶液， 再经 120°C干燥， 500°C焙烧制得）。 后反应器为酯 加氢反应器， 内装 40g铜锌铝酯加氢催化剂（实验室合成， 组成为 CuO 40%, ZnO 29.6 % , Α1203 30.4%。 由铜、 锌、 铬的硝酸盐溶液， 加入氢 氧化钠溶液中和至 PH = 9.0， 经离心分离， 洗涤， 干燥， 压片成型， 焙 烧制得)。 催化剂装入反应器的中部， 两端填充一定量的石英沙， 作为 原料加热气化区或填料。 反应器夹套中可通入导热油控制反应温度。 反应器中装好催化剂后， 连接反应系统， 并完成系统气密试验后， 通 入氢气（ lOOOmL/min )在 300° (：、 6MPa条件下还原 24h后， 降至加氢 反应的温度和压力。 将乙酸和乙酸环己酯的混合物由计量泵打入反应 器中， 氢气经质量流量控制器进入反应系统进行加氢反应， 通过反应 管外部夹套中通入导热油控制反应温度， 通过反应器出口背压阀控制 反应器压力。 反应产物通过反应器后部的直线取样阀取样进行在线色 谱分析。 反应条件和结果见表 6。
表 6 乙酸 /乙酸环己酯加氢试验数据
Figure imgf000083_0001
实施例 8
本实施例说明乙酸和乙酸环己酯混合物的加氢方法。
加氢原料为乙酸与乙酸环己酯的混合物 （乙酸 39.5 %, 乙酸环己 酯 60.5% ) 。
反应系统由单个固定床反应器组成， 反应器为带有夹套的钛质钢 管， 尺寸为 ())20x2.5x800mm。 催化剂分两层装入反应器。 上层装入 20g 二氧化硅负载的铂钯锡乙酸加氢催化剂（实验室合成， 组成为 Pt(10%)-Pd(5%)-Sn(5%)/SiO2, 由 20~40 目的大孔二氧化硅载体 （BET 比表面积 400m2/g, 孔容 0.35mL/g )浸渍氯铂酸、 氯化钯和氯化亚锡的 混合容易， 再经 120°C干燥， 500°C焙烧制得)； 下层装入 20g铜辞铝酯 加氢催化剂（实验室合成， 组成为 CuO 40%, ZnO 29.6 % , A1203 30.4%。 由铜、 辞、 铬的硝酸盐溶液， 加入氢氧化钠溶液中和至 PH = 9.0， 经离 心分离， 洗涤， 干燥， 压片成型， 焙烧制得)。 催化剂装入反应器的中 部恒温区， 两层催化剂间由玻璃纤维布隔开， 反应器两端填充一定量 的石英沙， 作为原料加热气化区或填料。 反应器夹套中可通入导热油 控制反应温度。 反应器中装入催化剂后， 连接反应系统， 并完成系统 气密试验后， 通入氢气（ 500mL/min )在 300°C、 6MPa条件下还原 24h 后， 降至加氢反应的温度和压力。 将乙酸和乙酸环己酯的混合物由计 量泵打入反应器中， 氢气经质量流量控制器进入反应系统进行加氢反 应， 通过反应管外部夹套中通入导热油控制反应温度， 通过反应器出 口背压阀控制反应器压力。 反应产物通过反应器后部的直线取样阀取 样进行在线色谱分析。 反应条件和结果见表 7。
表 7 乙酸 /乙酸环己酯加氢试验数据
Figure imgf000085_0001
实施例 10
收集实施例 8 ~ 9的加氢反应产物 4000g, 进行精馏分离试验。 精 馏采用高 2m玻璃塔， 塔柱装有（D3mm的不锈钢 Θ网环高效精馏填料， 塔釜为 5L玻璃烧瓶， 通过电热套进行加热， 通过调压器调节塔釜加热 量。 塔的回流采用回流比调节器进行控制。 精馏分离结果见表 8。 表 8 加氢产物的精馏分离试验数据
Figure imgf000086_0001
第八种实施方案
实施例 1
本实施例说明苯选择性加氢制环己烯的方法。
将苯和氢气按摩尔比 1 :3注入装填有钌颗粒催化剂的加氢反应器， 在反应温度 135 °C、 压力 4.5MPa、 停留时间 15min的条件下进行苯加 氢反应， 反应产物分离出氢气后， 收集液体产物， 连续运行 1000h。 试 验结束后，对收集的液体产物进行气相色谱分析，其组成为：苯 53.3%, 环己烯 35.4%, 环己烷 1 1.3%。
实施例 2
将 lOOmL大孔强酸性氢型离子交换树脂 （实验室按经典的文献方 法合成，将含有 15 %二乙烯基苯的苯乙烯溶液进行悬浮共聚制成白球， 再经浓硫酸磺化制得， 测得其交换容量为 5.2mmolH+/g 干基） 装入 (})32x4x l000mm带有夹套的不锈钢管反应器中的中部，两端填充一定量 的石英沙。 将乙酸和环己烯原料 （苯 53.3%, 环己烯 35.4%， 环己烷 1 1.3% )按一定的流量分别由计量泵打入反应器中进行反应， 在反应管 外部夹套中通入热水以控制反应温度， 通过反应器出口背压阀控制反 应器压力。反应器出口产物由在线取样阀取样进行在线色谱分析， 由产 物组成计算出环己烯单程转化率和乙酸环己酯单程选择性。 反应条件 和结果见表 1。
由表 1 可知， 采用强酸型离子交换树脂催化剂催化环己烯原料与 乙酸反应， 环己烯单程转化率大于 80%, 酯产物单程选择性大于 99 %， 运行 600小时， 催化剂活性和单程选择性稳定不变。 强酸性离子交换树脂催化乙酸与环己烷 /环己浠 /苯酯化试验
Figure imgf000087_0001
实施例 3
试猃装置、 方法和原料同实施例 2 , 所不同的是催化剂为磷改性的 Ηβ分子筛催化剂 （由硅铝比为 50的 Ηβ分子筛经 85%的磷酸改性， 再 与氧化铝捏合挤条成型，经 120°C烘干， 500 °C焙烧制得，磷含量为 2% )。 反应条件和结果见表 2。 由表 2可见， 环己烯与乙酸反应单程转化率大 于 80%, 酯产物单程选择性大于 99 %， 运行 480小时， 催化剂活性和 单程选择性稳定不变。 表 2 Ηβ分子筛催化剂催化乙酸与环己烷 /环己烯 /苯酯化试验数 据
Figure imgf000088_0001
实施例 4
收集实施例 2和 3 的加成酯化产物， 进行精馏分离试验。 精馏采 用高 2m直径为 40 mm的玻璃塔精馏装置， 塔柱装有 Φ3ιηιη的不锈钢 Θ网环高效精馏填料， 塔釜为体积 5的 L玻璃烧瓶， 装料量为 4L, 通 过电热套对塔釜进行加热， 通过调压器调节塔釜加热量。 塔的回流采 用回流比调节器进行控制。 精馏分离结果见表 3。 加成酯化产物精馏分离试验结果
Figure imgf000089_0001
实施例 5 ~ 6用于说明采用反应精馏制造乙酸环己酯的方法。
实施例 5 ~ 6中所进行的试验均是在如下规格的反应精馏模式试验 装置进行的： 模式装置的主体为直径 （内径） 为 50mm, 高为 3m的不 锈钢塔， 塔的下部连接体积为 5L的塔釜， 釜内配置有 10KW的电加热 棒， 此加热棒由智能控制器通过可控硅（SCR )控制塔釜加热量。 塔顶 连接有换热面积为 0.5m2的冷凝器，塔顶蒸汽经此冷凝器冷凝成液体后 进入一个体积为 2L的回流罐。 回流罐中的液体经回流泵部分回流至反 应塔， 部分采出作为轻组分。 塔的操作参数由智能型自动化控制仪表 显示和控制。 塔回流量由回流调节阀控制， 塔顶采出量由回流罐的液 位控制器控制。 塔釜采出量由塔釜液位控制器调节塔釜排料阀进行控 制。 乙酸和环己烯原料分别装入 30L储罐中， 并通过计量泵打入相应 的预热器中预热到一定温度后进入反应塔， 进料速度由计量泵控制、 电子秤精确计量。
实施例 5
将耐高温磺酸型离子交换树脂 （牌号为 Amberlyst 45， 由 Rhom&Hass 公司生产） 用多级高速粉碎机粉碎成粒度小于 200 目 ( 0.074mm )的粉料， 加入制孔剂、 润滑剂、 抗氧剂和粘合剂在高速混 合机上混合均匀，再在密炼机上于 180°C密炼 lOmin,使物料完全塑化， 之后注入模具中制成直径为 5mm, 高 5mm， 壁厚为 1mm拉西环型树 脂催化剂填料。 将此填料 1950mL装入模式反应塔的中部 （高 lm， 相 当于 8块理论塔板）上下各装入直径为 3mm、长 6mm的玻璃弹簧填料 1950mL ( 装填高度为 lm， 相当于 10块理论塔板)。将环己烯原料和乙 酸分别由计量泵打入预热器预热后进入反应塔， 调节塔釜加热量和塔 顶回流量连续进行反应， 稳定操作下的反应条件和反应结果见表 4。 耐高温磺酸型离子交换树脂催化剂的反应精馏试猃数据
Figure imgf000090_0001
根据试验数据计算环己烯的单程转化率 98.8 %， 乙酸环己酯单程 选择性 98.0 % 。
实施例 6
将 φ3 ~ 4的球型 H0 5Cs2.5PW12O40/SiO2催化剂（由 H0 5Cs2 5PW12O40 粉末和粒度小于 200 目的粗孔硅胶粉末， 在混料机中充分混合后， 在 糖衣机中以硅溶胶为粘合机滚球成型， 再经烘干、 焙烧而成） 夹入钛 丝网波板中， 制成直径为 50mm、 高 50mm的圆柱型规整填料。 将此填 料型催化剂 L装入模式反应塔的中部（高 lm, 相当于 12块理论塔板） 上下各装入直径为 4mm、 高为 4mm的 1950mL玻璃弹簧填料（ 装填高 度为 lm, 相当于 15块理论塔板)。 将环己烯原料和乙酸分别由计量泵 打入预热器预热后进入反应塔， 调节塔釜加热量和塔顶回流量连续进 行反应， 稳定操作下的反应条件和反应结果见表 5。 表 5 H。 5CS2.5PW1204Q/SI02催化剂的反应精馏试验数据
Figure imgf000091_0001
根据试验数据计算环己烯的单程转化率 99.35 % , 乙酸环己酯单程 选择性 99.6 %。
实施例 7 ~ 8用于说明乙酸环己酯的加氢方法。
实施例 Ί
加氢原料为纯度 99.6%的乙酸环己酯。
将 40g铜锌铝酯加氢催化剂（实验室合成，组成为 CuO 40.5%, ZnO 29.6 % , A1203 30.4%。 由铜、 锌、 铬的硝酸盐溶液， 加入氢氧化钠溶 液中和至 PH - 9.0 , 经离心分离， 洗涤， 干燥， 压片成型， 焙烧制得) 装入 (j)20x2.5x800mm 带有夹套的不锈钢管反应器中的中部， 两端填充 一定量的石英沙。 通入氢气 （ 500mL/min ) 在 280°C、 6MPa条件下还 原 24h后， 降至加氢反应的温度和压力。 将乙酸环己酯由计量泵打入 反应器中， 氢气经质量流量控制器进入反应系统进行加氢反应， 通过 反应管外部夹套中通入导热油控制反应温度， 通过反应器出口背压阀 控制反应器压力。 反应产物通过反应器后部的直线取样岡取样进行在 线色谱分析。 反应条件和结果见表 6。 表 6结果表明， 采用铜锌铝催化 剂， 乙酸环己酯加氢反应单程转化率最高可达到 99.0%以上， 环己醇单 程选择性大于 99.9 % , 运行 1000小时， 单程转化率和单程选择性均未 下降。 表 6 铜锌铝酯加氢催化剂的乙酸环己酯加氢试验数据
Figure imgf000092_0001
实施例 8
加氢原料为纯度 99.6%的乙酸环己酯。
将 40g铜铬酯加氢催化剂（市售，太原市欣吉达化工有限公司生产, 牌号为 Cl-XH-1, CuO含量为 55%, 直径 5mm片剂， 破碎成 10~20 目颗粒） 装入 φ20χ2.5χ800ππη 带有夹套的不锈钢管反应器中的中部， 两端填充一定量的石英沙。 通入氢气 （ 500mL/min) 在 280°C 6MPa 条件下还原 24h后， 降至反应得温度和压力。 将乙酸环己酯由计量泵 打入反应器中， 氢气经质量流量控制器进入反应系统进行加氢反应， 通过反应管外部夹套中通入导热油控制反应温度， 通过反应器出口背 压阀控制反应器压力。 反应产物通过反应器后部的直线取样阀取样进 行在线色谱分析。 反应条件和结果见表 7。 表 7结果表明， 采用铜锌铝 催化剂， 乙酸环己酯加氢反应单程转化率最高可达到 98.0%以上， 环己 醇单程选择性大于 99.9% , 运行 500 小时， 单程转化率和选择均未下 降。 铜铬酯加氢催化剂的乙酸环己酯加氢试验数据
Figure imgf000093_0001
实施例 9
收集实施例 7~ 8的反应产物 4000g, 进行精馏分离试验。 精馏采 用高 2m玻璃塔， 塔柱装有 Φ3πιιη的不锈钢 Θ网环高效精馏填料， 塔 釜为 5L玻璃烧瓶，通过电热套进行加热，通过调压器调节塔釜加热量。 塔的回流采用回流比调节器进行控制。 精馏分离结果见表 8。 表 8 乙酸环己酯加氢产物的精馏分离试脸数据
Figure imgf000094_0001
第九种实施方案
实施例 1
本实施例用于说明苯选择性加氢制环己烯的方法。
将苯和氢气按摩尔比 1 :3注入装填有钌颗粒催化剂的加氢反应器， 在反应温度 135 °C、 压力 4.5MPa、 停留时间 15min的条件下进行苯加 氢反应， 反应产物分离出氢气后， 收集液体产物， 连续运行 1000h。 试 验结束后， 对收集的液体产物进行气相色谱分析， 其组成为： 苯 53.3m%, 环己烯 35.4m%, 环己烷 11.3m%。 以 N, N-二甲基乙酰胺为 萃取剂， 对上述液体产物进行萃取分离， 得到环己烯和苯混合物。
实施例 2
将 lOOmL大孔强酸性氢型离子交换树脂 （实验室按经典的文献方 法合成，将含有 15 %二乙烯基苯的苯乙烯溶液进行悬浮共聚制成白球， 再经浓硫酸磺化制得， 测得其交换容量为 5.2mmolH+/g 干基） 装入 φ32χ4χ 1000ιηπι带有夹套的不锈钢管反应器中的中部，两端填充一定量 的石英沙。 将乙酸和环己烯原料（用实施例 1 方法获得， 组成为： 苯 60m%,环己烯 40m% )按一定的流量分别由计量泵打入反应器中进行反 应， 在反应管外部夹套中通入热水以控制反应温度， 通过反应器出口 背压阀控制反应器压力。 反应器出口产物由在线取样阀取样进行在线 色谱分析， 由产物组成计算出环己烯单程转化率和乙酸环己酯单程选 择性。 反应条件和结果见表 1。
由表 1可知， 采用强酸型离子交换树脂催化剂环己烯与乙酸反应， 环己烯单程转化率大于 80% , 酯产物单程选择性大于 99 % , 运行 600 小时， 催化剂活性和单程选择性稳定不变。
实施例 3
试验装置、 方法和原料同实施例 2 , 所不同的是催化剂为磷改性的 Ηβ分子筛催化剂 （由硅铝比为 50的 Ηβ分子筛经 85%的磷酸改性， 再 与氧化铝捏合挤条成型，经 120°C烘干， 500 °C焙烧制得，磷含量为 2% )。 反应条件和结果见表 2。 由表 2可见， 环己烯单程转化率 80% , 酯产物 单程选择性大于 99 % , 运行 480小时， 催化剂活性和单程选择性稳定 不变。
实施例 4
收集实施例 2和 3 的加成酯化产物， 进行精馏分离试验。 精馏采 用高 2m直径为 40 mm的玻璃塔精馏装置， 塔柱装有 D3mm的不锈钢 Θ网环高效精镏填料， 塔釜为体积 5的 L玻璃烧瓶， 装料量为 4L， 通 过电热套对塔釜进行加热， 通过调压器调节塔釜加热量。 塔的回流采 用回流比调节器进行控制。 精馏分离结果见表 3。
实施例 5 ~ 6用于说明反应精馏制造乙酸环己酯的方法。
实施例 5 ~ 6中所进行的试验均是在如下规格的反应精馏模式试验 装置进行的： 模式装置的主体为直径 （内径） 为 50mm， 高为 3m的不 锈钢塔， 塔的下部连接体积为 5L的塔釜， 釜内配置有 10KW的电加热 棒， 此加热棒由智能控制器通过可控硅（SCR )控制塔釜加热量。 塔顶 连接有换热面积为 0.5m2的冷凝器，塔顶蒸汽经此冷凝器冷凝成液体后 进入一个体积为 2L的回流罐。 回流罐中的液体经回流泵部分回流至反 应塔， 部分采出作为轻组分。 塔的操作参数由智能型自动化控制仪表 显示和控制。 塔回流量由回流调节阀控制， 塔顶采出量由回流罐的液 位控制器控制。 塔釜采出量由塔釜液位控制器调节塔釜排料阀进行控 制。 乙酸和环己烯原料 （与实施例 2相同） 分别装入 30L储罐中， 并 通过计量泵打入相应的预热器中预热到一定温度后进入反应塔， 进料 速度由计量泵控制、 电子秤精确计量。
实施例 5
将耐高温磺酸型离子交换树脂 （牌号为 Amberlyst 45 , 由 Rhom&Hass 公司生产） 用多级高速粉碎机粉碎成粒度小于 200 目 ( 0.074mm )的粉料， 加入制孔剂、 润滑剂、 抗氧剂和粘合剂在高速混 合机上混合均匀，再在密炼机上于 180°C密炼 lOmin,使物料完全塑化， 之后注入模具中制成直径为 5mm, 高 5mm, 壁厚为 1mm拉西环型树 脂催化剂填料。 将此填料 1950mL装入模式反应塔的中部 （高 lm, 相 当于 8块理论塔板）上下各装入直径为 3mm、长 6mm的玻璃弹簧填料 1950mL ( 装填高度为 lm, 相当于 10块理论塔板)。将环己烯原料和乙 酸分别由计量泵打入预热器预热后进入反应塔， 调节塔釜加热量和塔 顶回流量连续进行反应， 稳定操作下的反应条件和反应结果见表 4。
实施例 6
将 φ3 ~ 4的球型 H 5Cs2.5PW12O40/SiO2催化剂（由 H0.5Cs2 5PW12O40 粉末和粒度小于 200 目的粗孔硅胶粉末， 在混料机中充分混合后， 在 糖衣机中以硅溶胶为粘合机滚球成型， 再经烘干、 焙烧而成） 夹入钛 丝网波板中， 制成直径为 50mm、 高 50mm的圆柱型规整填料。 将此填 料型催化剂 L装入模式反应塔的中部（高 lm， 相当于 12块理论塔板） 上下各装入直径为 4mm、 高为 4mm的 1950mL玻璃弹簧填料（ 装填高 度为 lm, 相当于 15块理论塔板)。 将环己烯原料和乙酸分别由计量泵 打入预热器预热后进入反应塔， 调节塔釜加热量和塔顶回流量连续进 行反应， 稳定操作下的反应条件和反应结果见表 5。
实施例 7 ~ 8用于说明乙酸环己酯的加氢方法。
实施例 7
采用纯度为 99.6%的乙酸环己酯为加氢原料。
将 40g铜锌铝酯加氢催化剂（实验室合成，组成为 CuO 40.5%, ZnO 29.6 % , Α1203 30.4%。 由铜、 锌、 铬的硝酸盐溶液， 加入氢氧化钠溶 液中和至 ΡΗ - 9.0, 经离心分离， 洗涤， 干燥， 压片成型， 焙烧制得) 装入 ())20x2.5x800mm 带有夹套的不锈钢管反应器中的中部， 两端填充 一定量的石英沙。 通入氢气 （ 500mL/min ) 在 280°C、 6MPa条件下还 原 24h后， 降至加氢反应的温度和压力。 将乙酸环己酯由计量泵打入 反应器中， 氢气经质量流量控制器进入反应系统进行加氢反应， 通过 反应管外部夹套中通入导热油控制反应温度， 通过反应器出口背压岡 线色谱分析。 反应条件和结果见表 6;。 表" ^结果表明， 采用铜锌铝酯加 氢催化剂， 乙酸环己酯加氢反应单程转化率最高可达 99%, 环己醇单 程选择性大于 99 %， 运行 1000小时， 单程转化率和单程选择性均未下 降。
实施例 8
采用纯度为 99.6%的乙酸环己酯为加氢原料。
将 40g铜铬酯加氢催化剂（市售，太原市欣吉达化工有限公司生产， 牌号为 Cl-XH- 1 , CuO含量为 55%, 直径 5mm片剂， 破碎成 10 ~ 20 目颗粒） 装入 (|)20x2.5x800mm 带有夹套的不锈钢管反应器中的中部， 两端填充一定量的石英沙。 通入氢气 （ 500mL/min ) 在 280°C、 6MPa 条件下还原 24h后， 降至反应得温度和压力。 将乙酸环己酯由计量泵 打入反应器中， 氢气经质量流量控制器进入反应系统进行加氢反应， 通过反应管外部夹套中通入导热油控制反应温度， 通过反应器出口背 压阀控制反应器压力。 反应产物通过反应器后部的直线取样岡取样进 行在线色谱分析。 反应条件和结果见表 7。 表 7结果表明， 采用铜锌铝 酯加氢催化剂， 乙酸环己酯加氢反应单程转化率最高可达到 98%以上， 环己醇单程选择性大于 99 % , 运行 500小时， 单程转化率和单程选择 性均未下降。
实施例 9
收集实施例 7 ~ 8的反应产物 4000g, 进行精馏分离试验。 精馏采 用高 2m玻璃塔， 塔柱装有 cD3mm的不锈钢 Θ网环高效精馏填料， 塔 釜为 5L玻璃烧瓶，通过电热套进行加热，通过调压器调节塔釜加热量。 塔的回流采用回流比调节器进行控制。 精馏分离结果见表 8。
表 1 强酸性离子交换树脂催化乙酸与环己烯 /苯进行酯化反应试 验数据
Figure imgf000098_0001
Ηβ分子筛催化剂催化乙酸与环己烯 /苯进行酯化反应试验数
Figure imgf000099_0001
加成酯化产物精馏分离试验结果
Figure imgf000099_0002
表 4 耐高温磺酸型离子交换树脂催化剂的反应精馏试验数据
Figure imgf000100_0001
根据试验数据计算环己烯的单程转化率 99.5%, 乙酸环己酯单程 选择性 99.3 %。 表 5 H。.5Cs2.5PW1204。/Si02催化剂的反应精馏试验数据
Figure imgf000100_0002
根据试验数据计算环己烯的单程转化率 99.4%, 乙酸环己酯单程 选择性 99.6%。 铜锌铝催化剂催化乙酸环己酯加氢试验数据
Figure imgf000101_0001
表 7铜铬催化剂催化乙酸环己酯加氢试验数据
Figure imgf000102_0001
表 8 乙酸环己酯加氢产物精馏分离试验数据
Figure imgf000102_0002
第十种实施方案
实施例 1
本实施例用于说明苯选择性加氢制环己烯的方法。
将苯和氢气按摩尔比 1 :3注入装填有钌颗粒催化剂的加氢反应器， 在反应温度 135 °C、 压力 4.5MPa、 停留时间 min的条件下进行苯加 氢反应， 反应产物分离出氢气后， 收集液体产物， 连续运行 1000h。 试 验结束后， 对收集的液体产物进行气相色谱分析， 其组成为： 苯 53.3m%, 环己烯 35.4m%, 环己烷 11.3m%。 以 N, N-二甲基乙酰胺为 萃取剂， 对上述液体产物进行萃取分离， 得到环己烷和环己烯混合物。
实施例 2
将 lOOmL大孔强酸性氢型离子交换树脂 （实验室按经典的文献方 法合成，将含有 15 %二乙烯基苯的苯乙烯溶液进行悬浮共聚制成白球， 再经浓硫酸磺化制得， 测得其交换容量为 5.2mmolH+/g 干基） 装入 φ32χ4χ 1000ηιιτι带有夹套的不锈钢管反应器中的中部，两端填充一定量 的石英沙。 将乙酸和环己烯原料 （用实施例 1 方法获得， 组成为： 环 己烯 75m%,环己烷 25m% )按一定的流量分别由计量泵打入反应器中 进行反应， 在反应管外部夹套中通入热水以控制反应温度， 通过反应 器出口背压阀控制反应器压力。 反应器出口产物由在线取样阀取样进 行在线色谱分析， 由产物组成计算出环己烯单程转化率和乙酸环己酯 单程选择性。 反应条件和结果见表 1。
由表 1可知， 采用强酸型离子交换树脂催化剂环己烯与乙酸反应， 环己烯单程转化率大于 90%, 酯产物单程选择性大于 99 % , 运行 600 小时， 催化剂活性和单程选择性稳定不变。
实施例 3
试验装置、 方法和原料同实施例 2, 所不同的是催化剂为磷改性的 Ηβ分子筛催化剂 （由硅铝比为 50的 Ηβ分子筛经 85%的磷酸改性， 再 与氧化铝捏合挤条成型，经 120°C供干， 500°C焙烧制得，磷含量为 2% )。 反应条件和结果见表 2。 由表 2可见， 环己烯单程转化率 90%, 酯产物 单程选择性大于 99 % , 运行 480小时， 催化剂活性和单程选择性稳定 不变。
实施例 4
收集实施例 2和 3 的加成酯化产物， 进行精馏分离试验。 精馏采 用高 2m直径为 40 mm的玻璃塔精馏装置， 塔柱装有 < 3mm的不锈钢 Θ网环高效精馏填料， 塔釜为体积 5的 L玻璃烧瓶， 装料量为 4L, 通 过电热套对塔釜进行加热， 通过调压器调节塔釜加热量。 塔的回流采 用回流比调节器进行控制。 精馏分离结果见表 3。
实施例 5 ~ 6用于说明反应精馏制造乙酸环己酯的方法。
实施例 5 ~ 6中所进行的试验均是在如下规格的反应精馏模式试验 装置进行的： 模式装置的主体为直径 （内径） 为 50mm, 高为 3m的不 锈钢塔， 塔的下部连接体积为 5L的塔釜， 釜内配置有 10KW的电加热 棒， 此加热棒由智能控制器通过可控硅（SCR )控制塔釜加热量。 塔顶 连接有换热面积为 0.5m2的冷凝器，塔顶蒸汽经此冷凝器冷凝成液体后 进入一个体积为 2L的回流罐。 回流罐中的液体经回流泵部分回流至反 应塔， 部分采出作为轻组分。 塔的操作参数由智能型自动化控制仪表 显示和控制。 塔回流量由回流调节阀控制， 塔顶采出量由回流罐的液 位控制器控制。 塔釜采出量由塔釜液位控制器调节塔釜排料阀进行控 制。 乙酸和环己烯原料 （与实施例 2相同） 分别装入 30L储罐中， 并 通过计量泵打入相应的预热器中预热到一定温度后进入反应塔， 进料 速度由计量泵控制、 电子秤精确计量。
实施例 5
将耐高温磺酸型离子交换树脂 （牌号为 Amberlyst 45 , 由 Rhom&Hass 公司生产） 用多级高速粉碎机粉碎成粒度小于 200 目 ( 0.074mm )的粉料， 加入制孔剂、 润滑剂、 抗氧剂和粘合剂在高速混 合机上混合均匀，再在密炼机上于 180°C密炼 lOmin,使物料完全塑化， 之后注入模具中制成直径为 5mm, 高 5mm, 壁厚为 1mm拉西环型树 脂催化剂填料。 将此填料 1950mL装入模式反应塔的中部 （高 lm, 相 当于 8块理论塔板）上下各装入直径为 3mm、 长 6mm的玻璃弹簧填料 1950mL ( 装填高度为 lm, 相当于 10块理论塔板)。将环己烯原料和乙 酸分别由计量泵打入预热器预热后进入反应塔， 调节塔釜加热量和塔 顶回流量连续进行反应， 稳定操作下的反应条件和反应结果见表 4。
实施例 6
将 φ3 ~ 4的球型 H0.5Cs2.5PW12O40/SiO2催化剂（由 H0.5Cs2.5PW12O40 粉末和粒度小于 200 目的粗孔硅胶粉末， 在混料机中充分混合后， 在 糖衣机中以硅溶胶为粘合机滚球成型， 再经烘干、 焙烧而成） 夹入钛 丝网波板中， 制成直径为 50mm、 高 50mm的圆柱型规整填料。 将此填 料型催化剂 L装入模式反应塔的中部（高 lm, 相当于 12块理论塔板） 上下各装入直径为 4mm、 高为 4mm的 1950mL玻璃弹簧填料（ 装填高 度为 lm, 相当于 15块理论塔板)。 将环己烯原料和乙酸分别由计量泵 打入预热器预热后进入反应塔， 调节塔釜加热量和塔顶回流量连续进 行反应， 稳定操作下的反应条件和反应结果见表 5。
实施例 7 ~ 8用于说明乙酸环己酯的加氢方法。 实施例 7
采用纯度为 99.6%的乙酸环己酯为加氢原料。
将 40g铜锌铝酯加氢催化剂（实验室合成，组成为 CuO 40.5%, ZnO 29.6 % , Α1203 30.4%。 由铜、 锌、 铬的硝酸盐溶液， 加入氢氧化钠溶 液中和至 ΡΗ = 9.0 , 经离心分离， 洗涤， 干燥， 压片成型， 焙烧制得) 装入 (|)20x2.5x800mm 带有夹套的不锈钢管反应器中的中部， 两端填充 一定量的石英沙。 通入氢气 （500mL/min ) 在 280°C、 6MPa条件下还 原 24h后， 降至加氢反应的温度和压力。 将乙酸环己酯由计量泵打入 反应器中， 氢气经质量流量控制器进入反应系统进行加氢反应， 通过 反应管外部夹套中通入导热油控制反应温度， 通过反应器出口背压阀 线色谱分析。 反应条件和结果见表 6。 表 6结果表明， 采用铜锌铝酯加 氢催化剂， 乙酸环己酯加氢反应单程转化率最高可达 99%以上， 环己 醇单程选择性大于 99 %， 运行 1000小时， 单程转化率和单程选择性均 未下降。
实施例 8
采用纯度为 99.6%的乙酸环己酯为加氢原料。
将 40g铜铬酯加氢催化剂（市售，太原市欣吉达化工有限公司生产， 牌号为 Cl-XH-1 , CuO含量为 55%, 直径 5mm片剂， 破碎成 10 ~ 20 目颗粒） 装入 (|)20x2.5x800mm 带有夹套的不锈钢管反应器中的中部， 两端填充一定量的石英沙。 通入氢气 （ 500mL/min ) 在 280°C、 6MPa 条件下还原 24h后， 降至反应得温度和压力。 将乙酸环己酯由计量泵 打入反应器中， 氢气经质量流量控制器进入反应系统进行加氢反应， 通过反应管外部夹套中通入导热油控制反应温度， 通过反应器出口背 压阀控制反应器压力。 反应产物通过反应器后部的直线取样阀取样进 行在线色镨分析。 反应条件和结果见表 7。 表 7结果表明， 采用铜锌铝 酯加氢催化剂， 乙酸环己酯加氢反应单程转化率最高可达到 98%以上， 环己醇单程选择性大于 99 % , 运行 500小时， 单程转化率和单程选择 性均未下降。
实施例 9
收集实施例 7 ~ 8的反应产物 4000g， 进行精馏分离试^ r。 精馏采 用高 2m玻璃塔， 塔柱装有 Φ3ιηπι的不锈钢 Θ网环高效精馏填料， 塔 釜为 5L玻璃烧瓶，通过电热套进行加热，通过调压器调节塔釜加热量。 塔的回流采用回流比调节器进行控制。 精熘分离结果见表 8。
强酸性离子交换树脂催化乙酸与环己烷 /环己烯酯化试验数
Figure imgf000106_0001
Ηβ分子筛催化剂催化乙酸与环己烯酯化试验数据
Figure imgf000107_0001
加成酯化产物精馏分离试验数据
Figure imgf000107_0002
耐高温磺酸型离子交换树脂催化剂的反应精馏试验数据
Figure imgf000108_0001
根据试验数据计算环己烯的单程转化率 99.9 % , 乙酸环己酯单程 选择性 99.2 %。 表 5 H。 5Cs2.5PW1204。/Si02催化剂的反应精馏试验数据
Figure imgf000108_0002
根据试验数据计算环己烯的单程转化率 98.66 % , 乙酸环己酯单程 选择性 99.3 %。
表 6 铜锌铝酯加氢催化剂的乙酸环己酯加氢试^ ^数据
Figure imgf000109_0001
铜铬酯加氢催化剂的乙酸环己酯加氢试验数据
Figure imgf000110_0001
表 8 乙酸环己酯加氢产物的精馏分离试验数据
Figure imgf000110_0002
第十一种实施方案
实施例 1
本实施例用于说明苯选择性加氢制环己烯的试险方法。
将苯和氢气按摩尔比 1 :3注入装填有钌颗粒催化剂的加氢反应器， 在反应温度 135 °C、 压力 4.5MPa、 停留时间 15min的条件下进行苯加 氢反应， 反应产物分离出氢气后， 收集液体产物， 连续运行 1000h。 试 验结束后，对收集的液体产物进行气相色谱分析，其组成为：苯 53.3%， 环己烯 35.4%, 环己烷 1 1.3%。 以 N, N-二曱基乙酰胺为萃取剂， 对上 述液体产物进行萃取分离， 得到环己烯。
实施例 2
将 lOOmL大孔强酸性氢型离子交换树脂 （实验室按经典的文献方 法合成，将含有 15 %二乙烯基苯的苯乙烯溶液进行悬浮共聚制成白球， 再经浓硫酸磺化制得， 测得其交换容量为 5.2mmolH+/g 干基） 装入 φ32χ4χ 1000ιηιτι带有夹套的不锈钢管反应器中的中部，两端填充一定量 的石英沙。 将乙酸和环己烯 （采用实施例 1 方法获得） 按一定的流量 分别由计量泵打入反应器中进行反应， 在反应管外部夹套中通入热水 以控制反应温度， 通过反应器出口背压阔控制反应器压力。 反应器出 口产物由在线取样阀取样进行在线色谱分析， 由产物组成计算出环己 烯单程转化率和乙酸环己酯单程选择性。 反应条件和结果见表 1。
由表 1 可知， 采用强酸型离子交换树脂催化剂催化环己烯原料与 乙酸反应， 环己烯单程转化率大于 90%, 酯产物单程选择性大于 99 % , 运行 600小时， 催化剂活性和单程选择性稳定不变。
强酸性离子交换树脂催化乙酸与环己烯酯化试猃数据
Figure imgf000112_0001
实施例 3
试验装置 、 方法和原料同 实施例 2 , 所不 同 的是以 Cs25H05PW12O40/SiO2为催化剂（记为 PW/Si02, 下同）。 反应条件和结 果见表 2。 由表 2可见， 环己烯与乙酸反应单程转化率 95%， 酯产物单 程选择性 99%, 运行 480小时， 催化剂活性和单程选择性稳定不变。 Cs2.5H。5PW1204。/Si02催化乙酸与环己烯酯化试验数据
Figure imgf000113_0001
实施例 4
试验装置、 方法和原料同实施例 2, 所不同的是催化剂为磷改性的 Ηβ分子筛催化剂 （由硅铝比为 50的 Ηβ分子筛经 85%的磷酸改性， 再 与氧化铝捏合挤条成型，经 120°C烘干， 500°C焙烧制得，磷含量为 2% )。 反应条件和结果见表 3。 由表 3 可见， 环己烯与乙酸反应单程转化率 90% , 酯产物单程选择性大于 99 % , 运行 480 小时， 催化剂活性和单 程选择性稳定不变。 表 3 Ηβ分子筛催化剂催化乙酸与环己烯酯化试验数据
Figure imgf000114_0001
实施例 5
收集实施例 2 ~ 4的加成酯化产物， 进行精馏分离试验。 精馏釆用 高 2m直径为 40 mm的玻璃塔精馏装置， 塔柱装有 <D3mm的不锈钢 Θ 网环高效精馏填料， 塔釜为体积 5的 L玻璃烧瓶， 装料量为 4L, 通过 电热套对塔釜进行加热， 通过调压器调节塔釜加热量。 塔的回流采用 回流比调节器进行控制。 精馏分离结果见表 4 表 4 加成酯化产物精馏分离试验数据
Figure imgf000114_0002
实施例 6 ~ 7用于说明反应精镏制造乙酸环己酯的方法。 实施例 6 ~ 7中所进行的试验均是在如下规格的反应精馏模式试验 装置进行的： 模式装置的主体为直径 （内径） 为 50mm, 高为 3m的不 锈钢塔， 塔的下部连接体积为 5L的塔釜， 釜内配置有 10KW的电加热 棒， 此加热棒由智能控制器通过可控硅（SCR )控制塔釜加热量。 塔顶 连接有换热面积为 0.5m2的冷凝器，塔顶蒸汽经此冷凝器冷凝成液体后 进入一个体积为 2L的回流罐。 回流罐中的液体经回流泵部分回流至反 应塔， 部分采出作为轻组分。 塔的操作参数由智能型自动化控制仪表 显示和控制。 塔回流量由回流调节阔控制， 塔顶采出量由回流罐的液 位控制器控制。 塔釜采出量由塔釜液位控制器调节塔釜排料阀进行控 制。 乙酸和环己烯 （由实施例 1 方法获得） 分别装入 30L储罐中， 并 通过计量泵打入相应的预热器中预热到一定温度后进入反应塔， 进料 速度由计量泵控制、 电子秤精确计量。
实施例 6
将耐高温磺酸型离子交换树脂 （牌号为 Amberlyst 45， 由
Rhom&Hass 公司生产） 用多级高速粉碎机粉碎成粒度小于 200 目 ( 0.074mm )的粉料， 加入制孔剂、 润滑剂、 抗氧剂和粘合剂在高速混 合机上混合均匀，再在密炼机上于 180°C密炼 10min，使物料完全塑化， 之后注入模具中制成直径为 5mm, 高 5mm, 壁厚为 1mm拉西环型树 脂催化剂填料。 将此填料 1950mL装入模式反应塔的中部 （高 lm, 相 当于 8块理论塔板）上下各装入直径为 3mm、 长 6mm的玻璃弹簧填料 1950mL ( 装填高度为 lm, 相当于 10块理论塔板)。 将乙酸和环己烯分 别由计量泵打入预热器预热后进入反应塔， 调节塔釜加热量和塔顶回 流量连续进行反应， 稳定操作下的反应条件和反应结果见表 5。
表 5耐高温磺酸型离子交换树脂催化剂的反应精馏试验数据
Figure imgf000116_0001
根据试验数据计算环己烯的单程转化率 99 %， 乙酸环己酯单程选 择性 99.72 % 。
实施例 7
将 φ3 ~ 4的球型 H0.5Cs2.5PW12O40/SiO2催化剂（由 H0.5Cs2.5PW12O40 粉末和粒度小于 200 目的粗孔硅胶粉末， 在混料机中充分混合后， 在 糖衣机中以硅溶胶为粘合机滚球成型， 再经烘干、 焙烧而成） 夹入钛 丝网波板中， 制成直径为 50mm、 高 50mm的圆柱型规整填料。 将此填 料型催化剂 L装入模式反应塔的中部（高 lm, 相当于 12块理论塔板） 上下各装入直径为 4mm、 高为 4mm的 1950mL玻璃弹簧填料（ 装填高 度为 lm, 相当于 15块理论塔板）。 将乙酸和环己烯分别由计量泵打入 预热器预热后进入反应塔， 调节塔釜加热量和塔顶回流量连续进行反 应， 稳定操作下的反应条件和反应结果见表 6。 表 6 H。.5Cs2 5PW1204。/Si02催化剂的反应精馏试验数据
Figure imgf000117_0001
根据试验数据计算环己烯的单程转化率 98.7 % , 乙酸环己酯单程 选择性 99.43 %。
实施例 8 ~ 9用于说明乙酸环己酯的加氢试验结果。
实施例 8
采用纯度为 99.6%的乙酸环己酯为加氢原料。
将 40g铜锌铝酯加氢催化剂（实验室合成，组成为 CuO 40.5%, ZnO 29.6 % , A1203 30.4%。 由铜、 锌、 铝的硝酸盐溶液， 加入氢氧化钠溶 液中和至 PH - 9.0, 经离心分离， 洗涤， 干燥， 压片成型， 焙烧制得) 装入 (|)20x2.5x800mm 带有夹套的不锈钢管反应器中的中部， 两端填充 一定量的石英沙。 通入氢气（ 500mL/min ) 在 280°C、 6MPa条件下还 原 24h后， 降至加氢反应的温度和压力。 将乙酸环己酯由计量泵打入 反应器中， 氢气经质量流量控制器进入反应系统进行加氢反应， 通过 反应管外部夹套中通入导热油控制反应温度， 通过反应器出口背压阀 控制反应器压力。 反应产物通过反应器后部的直线取样阀取样进行在 线色谱分析。 反应条件和结果见表 7。 结果显示， 采用铜辞铝酯加氢催 化剂， 乙酸环己酯加氢反应单程转化率最高可达到 99%以上， 环己醇 单程选择性大于 99.9 % , 运行 1000小时， 单程转化率和单程选择性均 未下降。 表 7铜锌铝催化剂催化乙酸环己酯加氢试验数据
Figure imgf000118_0001
实施例 9
采用纯度为 99.6%的乙酸环己酯为加氢原料。 将 40g铜铬酯加氢催化剂（市售，太原市欣吉达化工有限公司生产， 牌号为 Cl-XH-1 , CuO含量为 55%, 直径 5mm片剂， 破碎成 10 ~ 20 目颗粒） 装入 (|)20x2.5x800mm 带有夹套的不锈钢管反应器的中部， 两 端填充一定量的石英沙。 通入氢气 （ 500mL/min ) 在 280° (：、 6MPa条 件下还原 24h后， 降至反应得温度和压力。 将乙酸环己酯由计量泵打 入反应器中， 氢气经质量流量控制器进入反应系统进行加氢反应， 通 过反应管外部夹套中通入导热油控制反应温度， 通过反应器出口背压 阀控制反应器压力。 反应产物通过反应器后部的直线取样阀取样进行 在线色谱分析。 反应条件和结果见表 8。 结果显示， 采用铜铬酯加氢催 化剂， 乙酸环己酯加氢反应单程转化率最高可达到 98%以上， 环己醇 单程选择性大于 99.9 % , 运行 500小时， 单程转化率和选择均未下降。 铜铬酯催化剂催化乙酸环己酯加氢试验数据
Figure imgf000119_0001
实施例 10
本实施例用于说明乙酸环己酯加氢产物的精馏分离试验结果。
收集实施例 8 ~ 9的反应产物 4000g, 进行精馏分离试验。 精馏采 用高 2m玻璃塔， 塔柱装有 3mm的不锈钢 Θ网环高效精馏填料， 塔 釜为 5L玻璃烧瓶，通过电热套进行加热，通过调压器调节塔釜加热量。 塔的回流采用回流比调节器进行控制。 精馏分离结果见表 9。 表 9 乙酸环己酯加氢产物的精馏分离试验数据
Figure imgf000120_0001
第十二种实施方案
实施例 1
本实施例说明苯选择性加氢制环己烯的方法。
将苯和氢气按摩尔比 1 :3注入装填有钌颗粒催化剂的加氢反应器， 在反应温度 135 °C、 压力 4.5MPa、 停留时间 15min的条件下进行苯加 氢反应， 反应产物分离出氢气后， 收集液体产物， 连续运行 1000h。 试 验结束后，对收集的液体产物进行气相色谱分析，其组成为：苯 53.3% , 环己烯 35.4% , 环己烷 1 1.3%。
实施例 2
将 l OOmL大孔强酸性氢型离子交换树脂 （实验室按经典的文献方 法合成，将含有 15 %二乙烯基苯的苯乙烯溶液进行悬浮共聚制成白球， 再经浓硫酸磺化制得， 测得其交换容量为 5.2mmolH+/g 干基） 装入 φ32χ4χ 1000mm带有夹套的不锈钢管反应器中的中部，两端填充一定量 的石英沙。 将乙酸和环己烯原料 （组成： 环己烯 75m%，环己烷 25m%; 用实施例 1 反应产物经萃取精馏获得， 萃取剂采用 N， N-二曱基乙酰 胺） 按一定的流量分别由计量泵打入反应器中进行反应， 在反应管外 部夹套中通入热水以控制反应温度， 通过反应器出口背压阀控制反应 器压力。反应器出口产物由在线取样阀取样进行在线色谱分析， 由产物 组成计算出环己烯单程转化率和乙酸环己酯单程选择性。 反应条件和 结果见表 1。
由表 1可知， 采用强酸型离子交换树脂催化剂环己烯与乙酸反应， 环己烯单程转化率大于 90%, 酯产物单程选择性大于 99 % , 运行 600 小时， 催化剂活性和单程选择性稳定不变。
实施例 3
试验装置、 方法同实施例 2 , 所不同的是催化剂为磷改性的 Ηβ分 子筛催化剂 （由硅铝比为 50的 Ηβ分子筛经 85%的璘酸改性， 再与氧 化铝捏合挤条成型， 经 120°C烘干， 500°C焙烧制得， 磷含量为 2% ) ； 环己烯原料（组成为： 苯 60m%,环己烯 40m%; 用实施例 1 反应产物 经萃取精馏获得， 萃取剂采用 N, N-二曱基乙酰胺） 。 反应条件和结 果见表 2。 由表 2可见， 环己烯与乙酸反应单程转化率大于 80%, 酯产 物单程选择性大于 99 % , 运行 480小时， 催化剂活性和单程选择性稳 定不变。
实施例 4 ~ 5用于说明采用反应精馏制造乙酸环己酯的方法。
实施例 4 ~ 5中所进行的试验均是在如下规格的反应精馏模式试验 装置进行的： 模式装置的主体为直径 （内径） 为 50mm, 高为 3m的不 锈钢塔， 塔的下部连接体积为 5L的塔釜， 釜内配置有 10KW的电加热 棒， 此加热棒由智能控制器通过可控硅（SCR )控制塔釜加热量。 塔顶 连接有换热面积为 0.5m2的冷凝器，塔顶蒸汽经此冷凝器冷凝成液体后 进入一个体积为 2L的回流罐。 回流罐中的液体经回流泵部分回流至反 应塔， 部分采出作为轻组分。 塔的操作参数由智能型自动化控制仪表 显示和控制。 塔回流量由回流调节阀控制， 塔顶采出量由回流罐的液 位控制器控制。 塔釜采出量由塔釜液位控制器调节塔釜排料阀进行控 制。 乙酸和环己烯原料分别装入 30L储罐中， 并通过计量泵打入相应 的预热器中预热到一定温度后进入反应塔， 进料速度由计量泵控制、 电子秤精确计量。
实施例 4
将耐高温磺酸型离子交换树脂 （牌号为 Amberlyst 45 , 由 Rhom&Hass 公司生产） 用多级高速粉碎机粉碎成粒度小于 200 目 ( 0.074mm )的粉料， 加入制孔剂、 润滑剂、 抗氧剂和粘合剂在高速混 合机上混合均匀，再在密炼机上于 180°C密炼 lOmin,使物料完全塑化， 之后注入模具中制成直径为 5mm, 高 5mm, 壁厚为 1mm拉西环型树 脂催化剂填料。 将此填料 1950mL装入模式反应塔的中部 （高 lm, 相 当于 8块理论塔板）上下各装入直径为 3mm、长 6mm的玻璃弹簧填料 1950mL ( 装填高度为 lm, 相当于 10块理论塔板)。将环己烯原料和乙 酸分别由计量泵打入预热器预热后进入反应塔， 调节塔釜加热量和塔 顶回流量连续进行反应， 稳定操作下的反应条件和反应结果见表 3。
实施例 5
将 φ3 ~ 4的球型 H0.5Cs2.5PW12O40/SiO2催化剂（由 H0.5Cs2.5PW12O40 粉末和粒度小于 200 目的粗孔硅胶粉末， 在混料机中充分混合后， 在 糖衣机中以硅溶胶为粘合机滚球成型， 再经烘干、 焙烧而成） 夹入钛 丝网波板中， 制成直径为 50mm、 高 50mm的圆柱型规整填料。 将此填 料型催化剂 L装入模式反应塔的中部（高 lm， 相当于 12块理论塔板） 上下各装入直径为 4mm、 高为 4mm的 1950mL玻璃弹簧填料（ 装填高 度为 lm, 相当于 15块理论塔板)。 将环己烯原料和乙酸分别由计量泵 打入预热器预热后进入反应塔， 调节塔釜加热量和塔顶回流量连续进 行反应， 稳定操作下的反应条件和反应结果见表 5。
实施例 6 ~ 7用于说明乙酸环己酯的加氢方法。
实施例 6
加氢原料为纯度 99.6%的乙酸环己酯。
将 40g铜锌铝酯加氢催化剂（实验室合成，组成为 CuO 40.5%， ZnO 29.6 % , A1203 30.4%。 由铜、 锌、 铬的硝酸盐溶液， 加入氢氧化钠溶 液中和至 PH = 9.0 , 经离心分离， 洗涤， 干燥， 压片成型， 焙烧制得) 装入 (])20x2.5x800mm 带有夹套的不锈钢管反应器中的中部， 两端填充 一定量的石英沙。 通入氢气 （ 500mL/min )在 280°C、 6MPa条件下还 原 24h后， 降至加氢反应的温度和压力。 将乙酸环己酯由计量泵打入 反应器中， 氢气经质量流量控制器进入反应系统进行加氢反应， 通过 反应管外部夹套中通入导热油控制反应温度， 通过反应器出口背压阀 控制反应器压力。 反应产物通过反应器后部的直线取样阀取样进行在 线色谱分析。 反应条件和结果见表 5。 表 5结果表明， 采用铜锌铝催化 剂， 乙酸环己酯加氢反应单程转化率最高可达到 99.0%以上， 环己醇单 程选择性大于 99.9 % , 运行 1000小时， 单程转化率和单程选择性均未 下降。 实施例 7
加氢原料为纯度 99.6%的乙酸环己酯。
将 40g铜铬酯加氢催化剂（市售 ,太原市欣吉达化工有限公司生产， 牌号为 Cl -XH-1 , CuO含量为 55%, 直径 5mm片剂， 破碎成 10 ~ 20 目颗粒） 装入 φ20χ2.5χ800ηιπι 带有夹套的不锈钢管反应器中的中部， 两端填充一定量的石英沙。 通入氢气 （ 500mL/min ) 在 280°C、 6MPa 条件下还原 24h后， 降至反应得温度和压力。 将乙酸环己酯由计量泵 打入反应器中， 氢气经质量流量控制器进入反应系统进行加氢反应， 通过反应管外部夹套中通入导热油控制反应温度， 通过反应器出口背 压阀控制反应器压力。 反应产物通过反应器后部的直线取样阀取样进 行在线色谱分析。 反应条件和结果见表 6。 表 6结果表明， 采用铜辞铝 催化剂， 乙酸环己酯加氢反应单程转化率最高可达到 98.0%以上， 环己 醇单程选择性大于 99.9 % , 运行 500 小时， 单程转化率和选择均未下 降。
强酸性离子交换树脂催化乙酸与环己烷 /环己烯酯化试验数
Figure imgf000124_0001
Ηβ分子筛催化剂催化乙酸与环己烯 /苯进行酯化反应试验数
Figure imgf000125_0001
表 3耐高温磺酸型离子交换树脂催化剂的反应精馏试验数据
Figure imgf000126_0001
根据试猃数据计算环己烯的单程转化率 98.8%, 乙酸环己酯单程 选择性 98.0%。 表 4 H。.5Cs2.5PW12O Si02催化剂的反应精馏试验数据
Figure imgf000126_0002
根据试验数据计算环己烯的单程转化率 99.35% , 乙酸环己酯单程 选择性 99.6%。
表 5 铜锌铝酯加氢催化剂的乙酸环己酯加氢试验数据
Figure imgf000127_0001
表 6 铜铬酯加氢催化剂的乙酸环己酯加氢试验数据
Figure imgf000128_0001
以上虽然已结合实施例对本发明的具体实施方式进行了详细的说 明， 但是需要指出的是， 本发明的保护范围并不受这些具体实施方式 的限制， 而是由附录的权利要求书来确定。 本领域技术人员可在不脱 离本发明的技术思想和主旨的范围内对这些实施方式进行适当的变

Claims

权 利 要 求
1. 一种联产环己醇和链烷醇的方法， 其特征在于， 包括以下步骤： ( 1 )使环己烯源与至少一种羧酸在加成酯化催化剂的存在下发生 加成酯化反应， 生成含有羧酸环己酯的加成酯化产物的步骤， 其中所 述至少一种羧酸用式 R-COOH表示，并且所述基团 R是氢或 C^23直链 或支链烷基， 优选 C1-6直链或支链烷基， 更优选 d_3直链或支链烷基， 最优选甲基， 所述环己烯源含有 20mol%以上、 35mol%以上、 20 至 80mol%、 20至 60mol%、 40至 80mol%、 80至 95mol%或者 95mol%以 上的环己烯； 和
(2)使所述加成酯化产物与氢气在加氢催化剂的存在下发生加氢 反应， 同时生成环己醇和链烷醇的步骤， 其中所述链烷醇用式 R-CH2-OH表示， 并且所述基团 R与所迷至少一种羧酸中的定义相同， 最优选甲基。
2. 权利要求 1所述的方法，进一步包括下述的步骤（A)、步骤（A)
+步骤 （B)、 步骤 （C)、 步骤 （C) +步骤 （D)之一或其任意的组合：
( A ) 使苯与氢气在部分加氢催化剂的存在下发生部分加氢反应， 以获得含有环己烯的加氢产物作为所述环己烯源的步骤；
(B)对步骤 （A) 获得的所述加氢产物进行进一步分离， 以获得 环己烯、 环己烯与苯的混合物或环己烯与环己烷的混合物作为所述环 己烯源的步驟；
(C)使环己烷在部分脱氢催化剂的存在下发生部分脱氢反应， 以 获得含有环己烯的部分脱氢产物作为所述环己烯源的步骤；
(D)对步骤 （C) 获得的所述部分脱氢产物进行进一步分离， 以 获得环己烯或环己烯与环己烷的混合物作为所述环己烯源的步骤。
3. 权利要求 2所述的方法， 进一步包括下述的步骤（1)、 步骤（Π) 和步骤 （III)之一或其任意的组合：
(I) 回收从所述联产环己醇和链烷醇的方法的任何步骤中分离的 苯和 /或氢气， 并将该苯和 /或氢气循环至所述步骤 （A);
( II )回收从所述联产环己醇和链烷醇的方法的任何步骤中分离的 环己烷， 并将该环己烷循环至所述步骤 （C);
(III) 回收从所述联产环己醇和链烷醇的方法的任何步骤中分离 的环己烷， 使该环己烷在脱氢催化剂的存在下发生脱氢反应， 以获得 苯和氢气， 并将该苯和 /或氢气循环至所述步骤 （A )。
4. 权利要求 1 所述的方法， 其中所述加成酯化催化剂选自固体酸 催化剂中的一种或多种， 优选选自酸强度函数（Hammett函数） H0为 -8以下 （优选 -12以下， 更优选 -13以下） 的固体酸催化剂中的一种或 多种， 更优选选自强酸型离子交换树脂 （优选选自磺酸型离子交换树 脂中的一种或多种， 更优选选自大孔磺酸型离子交换树脂和 素改性 磺酸型离子交换树脂中的一种或多种）、 杂多酸（比如选自 keggin结构 的杂多酸、 Dawson结构的杂多酸、 Anderson结构的杂多酸、 Silverton 结构的杂多酸、 前述杂多酸的酸式盐、 前述杂多酸 /载体和前述杂多酸 的酸式盐 /载体中的一种或多种，优选选自 keggin结构的杂多酸、 keggin 结构的杂多酸的酸式盐、 keggin结构的杂多酸 /载体和 keggin结构的杂 多酸的酸式盐 /载体中的一种或多种， 更优选选自十二磷钨酸或十二磷 钨酸 /载体、 十二硅钨酸或十二硅钨酸 /载体、 十二磷钼酸或十二磷钼酸 /载体、 十二磷钼钒酸或十二磷钼钒酸 /载体、 前述杂多酸的酸式盐和前 述杂多酸的酸式盐 /载体中的一种或多种， 更优选选自酸式磷钨酸铯盐
( CS2.5H0.5P12WO40 ) 和酸式磷钨酸铯盐 /载体中的一种或多种； 所述载 体比如选自二氧化硅和活性炭中的一种或多种） 和沸石分子筛 （优选 选自 Ηβ沸石分子筛、 氟和 /或磷改性的 Ηβ沸石分子筛、 ΗΥ沸石分子 筛、 氟和 /或磷改性的 ΗΥ沸石分子筛、 HZSM-5沸石分子筛以及氟和 / 或磷改性的 HZSM-5沸石分子筛中的一种或多种） 中的一种或多种。
5. 权利要求 1 所述的方法， 其中所述加氢催化剂选自铜系催化剂 (更优选选自含锌的铜系催化剂和含铬的铜系催化剂中的一种或多 种）、钌系催化剂（优选选自 Ru/Al203和 Ru-Sn/Al203中的一种或多种） 和贵金属系催化剂 （优选选自 Pt/Al203、 Pd-Pt/Al203和 Pd IC中的一种 或多种） 中的一种或多种， 优选铜系催化剂中的一种或多种。
6. 权利要求 5所述的方法，其中所述铜系催化剂包括以下组分（优 选由以下组分构成）： （a )氧化铜； （b )氧化锌； （c )选自铝、 镓、 锡、 钛、 锆、 铬、 钼、 钨、 锰、 铼、 镧系金属和锕系金属中的一种或多种 金属的氧化物， 优选选自铝、 镓、 锡、 钛、 锆、 铬、 钼、 钨、 锰、 铼、 镧和铈中的一种或多种金属的氧化物； 和 （d )选自碱金属氢氧化物和 碱土金属氢氧化物中的一种或多种， 优选选自氢氧化钾、 氢氧化钠和 氢氧化钡中的一种或多种， 其中以质量份数计， 组分（a): 组分 （b): 组分（ c ): 组分（ d ) 为 5至 60: 10至 50: 5至 60: 0.2至 2, 优选 10 至 50: 15至 45: 15至 55: 0.2至 2, 更优选 30至 45: 20至 35: 20 至 5( 0.5至 1.5。
7. 权利要求 6所述的方法， 其中所述铜系催化剂是通过包括以下 步骤的制造方法制造的：
( 1 ) 通过共沉淀法， 制造复合金属氧化物的步骤， 其中所述复合 金属氧化物包括以下组分（优选由以下组分构成）： （a)氧化铜； （b) 氧化锌； 和 （c) 选自铝、 镓、 锡、 钛、 锆、 铬、 钼、 钨、 锰、 铼、 镧 系金属和锕系金属中的一种或多种金属的氧化物， 优选选自铝、 镓、 锡、 钛、 锆、 铬、 钼、 钨、 锰、 铼、 镧和铈中的一种或多种金属的氧 化物， 其中以质量份数计， 组分 （a) ： 组分 （b) ： 组分 （c) 为 5至 60: 10至 50: 5至 60, 优选 10至 50: 15至 45: 15至 55, 更优选 30 至 45: 20至 35: 20至 50; 和
(2) 通过浸渍法， 向所述复合金属氧化物中引入 （d) 选自碱金 属氢氧化物和碱土金属氢氧化物中的一种或多种（优选选自氢氧化钾、 氢氧化钠和氢氧化钡中的一种或多种） 的步骤， 使得以质量份数计， 组分 （ a ) ： 组分 （ b ) ： 组分 （ c ) ： 组分 （ d ) 为 5至 60: 10至 50: 5至 60: 0.2至 2, 优选 10至 50: 15至 45: 15至 55: 0.2至 2, 更优 选 30至 45: 20至 35: 20至 50: 0.5至 1.5。
8. 权利要求 1 所述的方法， 其中所述至少一种羧酸与以环己烯计 的所述环己烯源的摩尔比为 0.2至 20:1， 优选 1.2至 4:1, 更优选 1.2 至 3:1, 并且所述步骤 （1 )按照以下的方式 （ 1)、 方式 （2) 或其任意 组合进行， 优选方式 （2) 或方式 （ 1 ) 与方式 （2) 的组合， 更优选先 进行方式 （ 1 ) 然后再进行方式 （ 2 ) 的组合：
方式 （ 1 ): 反应器为釜式反应器、 固定床反应器、 流化床反应器、 沸腾床反应器或其任意并联组合， 优选管式固定床反应器， 更优选管 壳列管式反应器， 反应温度为 50至 200 °C， 优选 60至 120°C, 反应压 力为常压至 lOMPa, 优选常压至 lMPa， 所述加成酯化反应按照连续方 式进行时， 液体进料空速为 0.5至 SOh , 优选 0.5至 51^, 更优选 1至 5h_1, 所述加成酯化反应按照间歇方式进行时， 反应时间为 0.1 ~ 10h， 优选 0.2~2h; 方式 （2 ): 反应器为反应精馏塔， 优选选自板式塔、 填料塔或其 任意并联组合， 理论塔板数为 10至 150, 优选 30至 100， 操作压力为 -0.0099MPa至 5MPa, 优选常压至 IMPa, 加成酯化催化剂床层装填区 的温度为 40至 200°C, 优选 50至 200°C, 更优选 60至 120°C, 回流比 为 0.1:1至全回流， 优选 0.1至 100:1， 更优选 0.5至 10:1， 在所述理论 塔板数的 1/3至 2/3位置之间选择 5至 30块板（优选 8至 20块板）布 置所述加成酯化催化剂， 并且相对于加成酯化催化剂的总装填体积， 液体进料空速为 0.1至 201-1, 优选 0.2至 201-1, 更优选 0.5至 511
9. 权利要求 1所述的方法， 其中所述步骤（2)在以下反应条件下 进行： 反应器为釜式反应器、 固定床反应器、 沸腾床反应器、 流化床 反应器或其任意并联组合， 优选管式固定床反应器， 更优选管壳列管 式反应器， 反应温度为 150至 400°C, 优选 200至 300°C, 反应压力为 常压至 20MPa， 优选常压至 10MPa， 更优选 4至 lOMPa, 氢气与以羧 酸环己酯计的所述加成酯化产物的摩尔比为 1至 1000:1,优选 5至 100: 1, 所述加氢反应按照连续方式进行时， 液体进料空速为 0.1至 201-1, 优选 0.2至 211-1, 所述加氢反应按照间歇方式进行时， 反应时间为 0.2 至 20h, 优选 0.5至 5h, 更优选 1至 5h。
10. 权利要求 1或 2所述的方法，还包括在使所述加成酯化产物与 氢气发生加氢反应之前， 对所述加成酯化产物进行分离的步骤， 以获 得羧酸环己酯或羧酸环己酯与所述至少一种羧酸的混合物作为所述加 成酯化产物， 优选获得羧酸环己酯作为所述加成酯化产物，
和 /或，
在使所述加成酯化产物与氢气发生加氢反应之前， 使所述加成酯 化产物与氢气在羧酸加氢催化剂的存在下发生羧酸加氢反应的步骤， 以将所述加成酯化产物中含有的游离羧酸转化为链烷醇。
11. 权利要求 10所述的方法， 其中所述羧酸加氢催化剂由主活性 组分 0.1至 30wt%、 助剂 0.1至 25wt%和余量的载体构成， 其中所述主 活性组分选自铂、 钯、 钌、 钨、 钼和钴中的一种或多种， 所述助剂选 自锡、 铬、 铝、 锌、 钙、 镁、 镍、 钛、 锆、 铼、 镧、 钍和金中的一种 或多种， 所述载体选自氧化硅、 氧化铝、 氧化钛、 氧化锆、 活性炭、 石墨、 纳米炭管、 硅酸钙、 沸石和硅酸铝中的一种或多种，
和 /或， 所述羧酸加氢反应在以下反应条件下进行： 反应器为釜式反应器、 固定床反应器、 沸腾床反应器、 流化床反应器或其任意并联组合， 优 选管式固定床反应器， 更优选管壳列管式反应器， 反应温度为 100 至 400 °C，优选 180至 300 °C，反应压力为 0.1至 30MPa,优选 2至 lOMPa, 氢气与所述游离羧酸的摩尔比为 1 至 500: 1 , 优选 5至 50: 1 , 所述 羧酸加氢反应按照连续方式进行时， 液体进料空速为 0.1至 5h- 优选 0.2 至 2h , 所述羧酸加氢反应按照间歇方式进行时， 反应时间为 0.5 至 20h, 优选 1至 5h。
12. 权利要求 1所述的方法，其中回收从所述联产环己醇和链烷醇 的方法的任何步骤中分离的羧酸和 /或环己烯源， 并将该羧酸和 /或环己 烯源循环至所述步骤 （ 1 )， 和 /或， 回收从所述联产环己醇和链烷醇的 方法的任何步骤中分离的氢气， 并将该氢气循环至所述步骤 （2 )。
13. 一种制造环己酮的方法， 其特征在于， 包括：
按照权利要求 1-12任一项所述的方法制造环己醇， 和
使用所述环己醇制造环己酮。 '
14. 一种制造己内酰胺的方法， 其特征在于， 包括：
按照权利要求 13所述的方法制造环己酮， 和
使用所述环己酮制造己内酰胺。
15. 一种联产环己醇和链烷醇的装置， 其特征在于， 包括加氢反应 单元 A、 任选的加氢产物分离单元 A、 加成酯化反应单元、 任选的加 成酯化产物分离单元、 加氢反应单元 B和加氢产物分离单元 B , 其中， 在所述加氢反应单元 A中， 使苯与氢气在部分加氢催化剂的存在 下发生部分加氢反应， 以获得含有环己烯的加氢产物；
在所述加氢产物分离单元 A中， 对来自所述加氢反应单元 A的所 述加氢产物进行分离， 以获得环己烯、 环己烯与苯的混合物或者环己 烯与环己烷的混合物；
在所述加成酯化反应单元中， 使来自所述加氢反应单元 A的所述 加氢产物和 /或来自所述加氢产物分离单元 A的环己烯、 环己烯与苯的 混合物或者环己烯与环己烷的混合物与羧酸在加成酯化催化剂的存在 下发生加成酯化反应， 生成含有羧酸环己酯的加成酯化产物；
在所述加成酯化产物分离单元中， 对来自所述加成酯化反应单元 的所述加成酯化产物进行分离， 以获得羧酸环己酯或者羧酸环己酯与 羧酸的混合物；
在所述加氢反应单元 B 中， 使来自所述加成酯化反应单元的所述 加成酯化产物和 /或来自所述加成酯化产物分离单元的羧酸环己酯或者 羧酸环己酯与羧酸的混合物与氢气在加氢催化剂的存在下发生加氢反 应， 生成含有环己醇和链烷醇的加氢产物； 和
在所述加氢产物分离单元 B中， 对来自所述加氢反应单元 B的所 述加氢产物进行分离， 以获得环己醇和链烷醇。
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