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CN103664531A - Method and device for co-producing cyclohexanol and ethanol - Google Patents

Method and device for co-producing cyclohexanol and ethanol Download PDF

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Publication number
CN103664531A
CN103664531A CN201210560665.9A CN201210560665A CN103664531A CN 103664531 A CN103664531 A CN 103664531A CN 201210560665 A CN201210560665 A CN 201210560665A CN 103664531 A CN103664531 A CN 103664531A
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reactor
hydrogenation
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logistics
catalyst
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CN103664531B (en
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宗保宁
马东强
温朗友
孙斌
杨克勇
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Sinopec Research Institute of Petroleum Processing
China Petroleum and Chemical Corp
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Sinopec Research Institute of Petroleum Processing
China Petroleum and Chemical Corp
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Priority to CN201210560665.9A priority Critical patent/CN103664531B/en
Priority to US14/429,189 priority patent/US9561991B2/en
Priority to PCT/CN2013/001100 priority patent/WO2014044020A1/en
Priority to TW102133691A priority patent/TWI612031B/en
Priority to KR1020157010078A priority patent/KR102008352B1/en
Publication of CN103664531A publication Critical patent/CN103664531A/en
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    • C—CHEMISTRY; METALLURGY
    • 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
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C5/00—Preparation of hydrocarbons from hydrocarbons containing the same number of carbon atoms
    • C07C5/02—Preparation of hydrocarbons from hydrocarbons containing the same number of carbon atoms by hydrogenation
    • C07C5/10—Preparation of hydrocarbons from hydrocarbons containing the same number of carbon atoms by hydrogenation of aromatic six-membered rings
    • 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
    • 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
    • Y02P20/00—Technologies relating to chemical industry
    • Y02P20/10—Process efficiency

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  • Organic Chemistry (AREA)
  • Organic Low-Molecular-Weight Compounds And Preparation Thereof (AREA)
  • Low-Molecular Organic Synthesis Reactions Using Catalysts (AREA)

Abstract

The invention provides a method for co-producing cyclohexanol and ethanol. The method comprises the following steps: on the basis that benzene is taken as an initial material, co-producing the cyclohexanol and ethanol by selective hydrogenation of benzene, addition esterification of cyclohexene, and hydrogenation of cyclohexyl acetate. The invention further provides a device for achieving the method. The method has the characteristics that (1) the esterification and ester hydrogenation reaction have very high selectivity, and the atom use ratio is very high; (2) the process is environmental friendly; (3) the ethanol is co-produced while the cyclohexanol is produced; (4) the addition esterification is performed by using reactive distillation, not only is the reaction efficiency obviously improved, but also the quenching rectification separation process can be simplified.

Description

Method and the device of a kind of coproduction hexalin and ethanol
Technical field
The present invention relates to method and the device of a kind of coproduction hexalin and ethanol.
Background technology
Hexalin and ethanol are important industrial chemicals and solvent.Hexalin is mainly for the production of products such as nylon 6, nylon 66; And ethanol is not only the raw material of the multiple Chemicals such as synthesizing ester, be also widely used as the fuel dope of gasoline.
The method of industrial synthesizing alcohol is mainly ethene direct hydration method, but in the abundant country of some agricultural byproducts, fermentation method is still the main method of producing ethanol.Because China is populous and cultivated area is not enough, and fermentation legal system ethanol exists the problem of " striving grain with mouth ", so fermentation method does not meet the national conditions of China.In addition, the pollution of fermentation method is also more serious.China's oil relative inadequacy of resources, and ethene price is subject to the influence of fluctuations of international oil price very large, therefore at China's application ethylene hydration method, can face certain raw materials cost pressure.In addition, the reaction conditions of ethene direct hydration method is harsher, need under High Temperature High Pressure, carry out.In sum, developing new ethanol synthesis route is the inevitable requirement of technology and Economic development.
CN1022228831A discloses a kind of catalyzer of acetic acid gas phase hydrogenation ethanol production, and this catalyzer is comprised of main active ingredient, auxiliary agent and carrier three parts; Carrier is any one in gac, graphite or multiple-wall carbon nanotube, and main active ingredient is any one or two kinds of metal W or Mo, and auxiliary agent is one or more of Pd, Re, Pt, Rh or Ru; The content of main active ingredient is 0.1～30.0% of catalyst weight, and the content of auxiliary agent is catalyst weight 0.1～10.0%, and surplus is carrier.
CN102149661A discloses a kind of platinum/tin catalyst that uses and by the direct selectivity of acetic acid, has been prepared the method for ethanol, comprise: the incoming flow containing acetic acid and hydrogen contacts with hydrogenation catalyst under comparatively high temps, described hydrogenation catalyst be included in platinum in applicable support of the catalyst and tin group and and optional load on the 3rd metal on described carrier, wherein said the 3rd metal is selected from the group that following metal forms: palladium, rhodium, ruthenium, rhenium.Iridium, chromium, copper, molybdenum, tungsten, vanadium and zinc.
Industrial, the production method of hexalin mainly contains air oxidation of cyclohexane method, phenol hydrogenation method and cyclohexene hydration method, and wherein the application of cyclohexane oxidation process is the most general.
Cyclohexane oxidation process is current topmost cyclohexanol production technique.This technology utilization oxygenant (being generally air) is cyclohexyl hydroperoxide by cyclohexane oxidation, and decomposing cyclohexyl hydrogen peroxide obtains the mixture (being commonly called as KA oil) of hexalin and pimelinketone.The advantage of this technique is that oxidizing process condition relaxes, slagging scorification is less, cycle of operation is long.Shortcoming is that operational path is long, energy consumption is high, pollution is large, and the cyclohexane conversion of this technique only has 3～5%; Particularly, in the decomposition course of cyclohexyl hydroperoxide, the selectivity of hexalin is poor, and yield is low; In addition, this technique also produces a large amount of unmanageable waste lyes, is still so far global the difficult problem of environmental protection.
Phenol hydrogenation method is to produce the comparatively clean technological line of hexalin, and has that technical process is short, product purity advantages of higher.Phenol hydrogenation is produced hexalin and is mainly adopted gas phase hydrogenation method.The method adopts 3～5 reactors in series conventionally.Under the effect of Supported Pd-Catalyst, under 140～170 ℃ and 0.1MPa, the yield of pimelinketone and hexalin can reach 90%～95%.Yet, this technique phenol (vaporization heat 69kJmol that need to vaporize -1) and methyl alcohol (vaporization heat 35.2kJmol -1), energy consumption is higher, and catalyzer in use easily carbon deposit cause active decline, short, the expensive and use noble metal catalyst of phenol, is restricted the industrial application of the method in addition.
The eighties in 20th century, Japan Asahi Kasei Corporation has developed the technique by partial hydrogenation of benzene cyclohexene, cyclohexene hydration hexalin processed, and having realized industrialization in nineteen ninety, relevant Chinese patent application has CN 1079727A, CN 1414933A and CN101796001A.Cyclohexene hydration method is relatively new cyclohexanol production method, and the reaction preference of the method is high, and process does not almost have three waste discharge, but exist reaction conversion ratio very low, to the more high deficiency of tetrahydrobenzene purity requirement.As adopt high silica ZSM-5 catalyzer, and in two series connection slurry reactors, stopping 2h, cyclohexene conversion rate only has 12.5%.
The traditional method of producing tetrahydrobenzene is dehydration of cyclohexanol method and cyclohexane halide dehydrohalogenation method.Partial hydrogenation of benzene and oxidative dehydrogenation of cyclonexane are other two kinds of methods of preparing tetrahydrobenzene.Partial hydrogenation of benzene is prepared tetrahydrobenzene and is mainly contained vapor phase process, liquid phase method and homogeneous phase complexing hydrogenation method.Liquid phase method can adopt gas (hydrogen), liquid (benzene), liquid (polar solvent), solid (catalyzer) four phase reaction systems, as adopted the method that passes into benzene and hydrogen in the slurry that comprises catalyzer and water.As adopt four phase reaction systems, and the consumption of water at least will meet and can form profit two-phase, conventionally adds metal-salt in water, and metal-salt is zinc sulfate or rose vitriol preferably.The reaction conditions of liquid phase method is generally: 25～250 ℃ of temperature of reaction; Hydrogen partial pressure 0.1～20MPa, catalyst levels is 0.001～0.2 times of water weight.It is main catalyst component that partial hydrogenation of benzene catalyzer generally be take one or more in the metals such as Pt, Pd, Ru, Rh and Ni; In order further to improve hydrogenation activity and selectivity, conventionally also can in catalyzer, introduce one or more in the metals such as K, Zr, Hf, Co, Cu, Ag, Fe, Mo, Cr, Mn, Au, la and Zn as adjuvant component.Partial hydrogenation of benzene catalyzer can be loading type or unsupported catalyst, and the method for load can adopt ion exchange method, spraying process, impregnation method, evaporation drying method etc., and the carrier of employing can be natural clay, sepiolite, ZrO 2, SiO 2, TiO 2, Al 2o 3, La 2o 3, gac, insoluble vitriol, insoluble phosphoric acid salt or molecular sieve etc.Take ruthenium catalyst as example, and loaded catalyst can be by by ruthenium salt separately or be jointly immersed on carrier then drying, also original preparation with other metal-salts; Unsupported catalyst can by ruthenium salt separately or with other metal-salt coprecipitations, then drying, also original preparation, also can prepare by the mixture of direct-reduction ruthenium compound or ruthenium compound and other metallic compounds.The method and apparatus of ruthenium catalyst and benzene hydrogenation is described later in detail in CN102264471A.
Hexalin acetate is a kind of liquid with banana or apple aroma, with the fruit flavour type essence of its preparation, is widely used in the industries such as food, beverage and makeup.In addition, hexalin acetate has good solubility energy to resin, is also often used as the environment-friendly type solvent of high-grade paint, paint.Recently, the inventor also finds, its hydrogenation can be produced to hexalin co-producing ethanol, can predict, and hexalin acetate will become a kind of important organic synthesis intermediate.
At present, the synthetic method of industrial hexalin acetate is acetic acid and hexalin esterification.Esterification reaction need to could carry out smoothly under the effect of an acidic catalyst.Song Guijia, Wu Xionggang (chemical propellant and macromolecular material, 2009, V0l.7 (2): P31-33), summarized the progress of synthesis situation of acetic acid and hexalin esterification Synthesis of Cyclohexyl Acetate.
JPA254634/1989 discloses the preparation method of a kind of hexalin and hexalin acetate, and employing strong-acid ion exchange resin is catalyzer, by aqueous acetic acid and tetrahydrobenzene reaction synthesizing cyclohexanol and hexalin acetate.The best result of mentioning in document example is, cyclohexene conversion rate 62.7%, hexalin yield 18.4%, hexalin acetate yield 43.7%.
CN1023115C, JP flat-313447 discloses a kind of preparation method of hexalin, and adopting ZSM5 or supersiliceous zeolite is catalyzer, under water exists, by acetic acid and tetrahydrobenzene reaction synthesizing cyclohexanol and hexalin acetate.In the document, at 120 ℃ of reaction 4h, the output of hexalin and hexalin acetate only has respectively 12.5% and 65%.
EP0461580A2, USP5254721 disclose a kind of method of being prepared hexalin acetate by acetic acid and tetrahydrobenzene.The method adopts heteropoly acid containing tungsten catalyzer, and in heteropolyacid molecule, crystal water content is preferably 0～3.The best result providing in document is, at 370 ℃ of roasting 3h, obtain completely not containing 12 silicotungstic acid catalysts of crystal water, in 200mL autoclave pressure, add 61.5g acetic acid, 13.5g tetrahydrobenzene, 5g catalyzer reacts 0.5h under 0.5MPa, the condition of 130 ℃, cyclohexene conversion rate is 95.2%, and the selectivity of hexalin acetate is 99.2%.As can be seen here, under the condition of very high sour alkene ratio, tetrahydrobenzene can not transform completely.
From existing disclosed document, existing document discloses the various solid acid catalysts of acetic acid and tetrahydrobenzene addition esterification, addition esterification generally adopts tank reactor, reaction raw materials is pure tetrahydrobenzene, even if adopt very high sour alkene ratio, be also difficult to realize the conversion completely of tetrahydrobenzene.
Reactive distillation has been widely used in the processes such as alfin etherificate, alcoholic acid esterification, transesterify, ester hydrolysis, aldolization, but up to now, has no the report for acetic acid and tetrahydrobenzene addition esterification process by reactive distillation.
CN86105765A proposes a kind of carboxylicesters Hydrogenation that passes through for the method for alcohol, the method is under the existence of solid copper containing catalyst that has reduction activation, under high temperature, normal pressure or high pressure by carboxylicesters hydrogenation, outside this catalyzer copper removal, also contain magnesium, at least one in lanthanide series metal or actinide metals.Catalyzer represented with following general formula before reduction activation: Cu am 1m 2 ba co x, M 1magnesium, at least one in lanthanide series metal or actinide metals, M 2be selected from Ca, Mo, Rh, Pt, Cr, Zn, Al, Ti, V, Ru, Re, Pd, Ag and Au; A is a kind of basic metal; A is 0.1-4; B is 0-1.0; C is 0-0.5; X is for can meet the numeral of other element to the total valence mumber requirement of oxygen.Basic metal in this catalyzer is a kind of selection component, and its form by an alkali metal salt is introduced catalyzer.The acyclic monobasic that the applicable carboxylicesters of the method and catalyzer is C1-C24 or binary, saturated or unsaturated, straight or branched carboxylicesters, do not relate to the preparation of the cycloalkanol as hexalin in document.
CN1075048C proposes a kind of method and catalyzer of direct hydrogenation of carboxylic esters, comprise and make one or more esters contact and react under following catalyzer exists with hydrogen, this catalyzer contains a kind of copper compound, a kind of zn cpds and at least one and is selected from the compound of aluminium, zirconium, magnesium, a kind of rare earth element or its mixture as its component, by these catalyst components are made to this catalyzer 200 to being less than roasting in the temperature range of 400 ℃, the method is under liquid phase, at 170～250 ℃ and 20.7～138 bar tables, depresses and carries out.The applicable carboxylicesters of the method and catalyzer is to carry out C6～C22 dimethyl ester, C6-C66 natural glycerin three esters that transesterify makes or the C6～C44 compound making do transesterify for natural glycerin three esters by natural oil.
US4939307 proposes the technique of a kind of ester through hydrogenation alcohol processed.By general formula, be R 1-CO-OR 2or R 4o-CO-R 3-CO-OR 2(R wherein 1for H or C 1～C 20alkyl, R 2and R 4for C 1～C 20alkyl, R 3for-(CH 2) n-group, n=1～10) ester and H 2mix with CO, carry out hydrogenation reaction and generate alcohol under 30～150 ℃, 5～100 bar pressures, its catalyzer is composed of the following components: (a) VIII family metal ion compound in a kind of periodictable; (b) alkoxide of a kind of basic metal or alkaline-earth metal; (c) a kind of alcohol.
US4113662 and USP4149021 disclose a kind of ester through hydrogenation catalyzer, this catalyzer is comprised of element, oxide compound, oxyhydroxide or the carbonate of cobalt, zinc, copper, the most applicable carboxylicesters of this catalyzer is polyglycolide, the preparation of not mentioned cycloalkanol in document.
US4611085 discloses a kind of C 1-C 20the method of carboxylicesters gas phase hydrogenation alcohol processed, its catalyzer is comprised of a kind of VIII family element, a kind of auxiliary agent and charcoal carrier, wherein said VIII family element comprises Ru, Ni, Rh, auxiliary agent comprises IA (except Li), IIA family (except Be and Mg), group of the lanthanides and actinide elements, and the BET specific surface area of charcoal carrier is greater than 100m 2/ g.Hydrogenation reaction is at 100～400 ℃, gas space velocity 100～120000h -1under condition, carry out.Basic metal in this catalyzer is to introduce with the form of an alkali metal salt, as alkali-metal nitrate, carbonate or acetate.The method is applicable to the carboxylicesters that can vaporize under reaction conditions, and the alcohol derivative moiety carbon number in carboxylicesters is preferably less than 5 and the carbon that is connected with oxygen primary carbon preferably.
GB2250287A discloses a kind of method of fatty acid ester Hydrogenation alcohol, and the feature of the method is that hydrogenation adopts copper containing catalyst and in ester raw material, adds a certain amount of water to maintain the activity of catalyzer.The fatty acid methyl ester that its applicable carboxylicesters is C12～C18.
From open source literature, in prior art, without any the information disclosure about ethyl cyclohexyl ester through hydrogenation energy coproduction hexalin and ethanol, more by benzene selective hydrogenation, tetrahydrobenzene addition esterification, ethyl cyclohexyl ester through hydrogenation, do not prepare the information disclosure of hexalin and ethanol.
Summary of the invention
The invention provides a kind of method of coproduction hexalin and ethanol, the method be take benzene as starting raw material, by benzene selective hydrogenation, tetrahydrobenzene addition esterification, ethyl cyclohexyl ester through hydrogenation, comes coproduction hexalin and ethanol.The present invention also provides the device that can realize aforesaid method.
In the present invention, for logistics, with " A/B ", represent the mixture of A and B; For catalyzer, " A/B " representative " active ingredient/carrier ".
A method for coproduction hexalin and ethanol, comprising:
(1), under the condition of preparing cyclohexene from benzene added with hydrogen, benzene and hydrogen generation hydrogenation reaction, obtain containing tetrahydrobenzene logistics;
(2) what step (1) is obtained contacts with acetic acid containing tetrahydrobenzene logistics, and addition esterification occurs under the effect of the first catalyzer; Reaction product is carried out to separation, obtain hexalin acetate logistics;
(3) hexalin acetate logistics step (2) being obtained contacts with hydrogen, and ester through hydrogenation reaction occurs under the effect of the second catalyzer; Reaction product is carried out to separation, obtain hexalin and ethanol.
Above-mentioned three steps are below described respectively.
One, preparing cyclohexene from benzene added with hydrogen
To the method for preparing cyclohexene from benzene added with hydrogen and catalyzer, there is no particular limitation in the present invention, and existing to utilize prepared from benzene and hydrogen all to can be the present invention for the method for tetrahydrobenzene and benzene hydrogenating catalyst used.The present invention preferably adopts liquid phase method technique.Benzene hydrogenating catalyst preferably adopts ruthenium catalyst, more preferably adopts the ruthenium catalyst containing cobalt and/or zinc.Ruthenium catalyst containing cobalt and/or zinc can be prepared by the method for co-precipitation or dipping identical carrier.
In step (1), the mixture mainly being formed by hexanaphthene, tetrahydrobenzene and unreacted benzene containing tetrahydrobenzene logistics that benzene hydrogenation obtains, this logistics directly, as the raw material of next step, does not need its separation.
Two, tetrahydrobenzene addition esterification
In the present invention, " addition esterification " refers to that carboxylic acid generates the reaction of ester to olefinic double bonds addition.
In step (2), the first described catalyzer is acid catalyst, can be both liquid acid catalyst, can be also solid acid catalyst.Described liquid acid catalyst can be both mineral acid, as sulfuric acid, phosphoric acid etc.; Also can be organic acid, as toluene sulfonic acide, amidosulfonic acid etc.The present invention preferably adopts solid acid catalyst.Optional one or more in strong acid ion exchange resin catalyzer, heteropolyacid catalyst and molecular sieve catalyst of described solid acid catalyst.
Described strong acid ion exchange resin catalyzer had both comprised common macropore sulfonic acid type polystyrene-divinylbenzene resin, also comprised the sulfonic resin after halogen atom modification.This resinoid is easy to buy from market, and the method that also can record by classical documents is produced.The preparation method of macropore sulfonic acid type polystyrene-divinylbenzene resin normally splashes into the mixture of vinylbenzene and Vinylstyrene and in the aqueous phase system that contains dispersion agent, initiator, pore-creating agent, carries out suspension copolymerization under the condition of high-speed stirring, resulting polymer globules (Archon) is separated from system, with solvent, pump pore-creating agent wherein, take again ethylene dichloride as solvent, the vitriol oil be sulphonating agent, carry out sulfonation reaction, finally by operations such as filtration, washings, finally make product.In the skeleton of common strong acid ion exchange resin, introduce halogen atom, as fluorine, chlorine, bromine etc., can further improve heat resistance and the strength of acid of resin.This halogen-containing strongly-acid fire resistant resin at least can obtain by following two kinds of approach, a kind of approach is to introduce halogen atom on the phenyl ring of sulfonated styrol resin skeleton, chlorine atom for example, because the strong electron attraction of halogens not only can make phenyl ring stable, but also can improve the acidity of sulfonic acid group on phenyl ring, can make like this strength of acid function (Hammett function) H0≤-8 of resin catalyst, and can be at more than 150 ℃ life-time service, this resinoid can conveniently buy from the market, such as Amberlyst 45 resins that abroad ROHM & HASS company produces, the D008 resin that Ji Zhong chemical plant, domestic Hebei produces etc., another kind of approach all replaces the hydrogen on resin matrix with fluorine, strong electrophilic due to fluorine, make it have superpower acidity and the thermostability of superelevation, strength of acid function (Hammett function) H0 can be less than-12, and heat resisting temperature reaches more than 250 ℃, the exemplary of this class fire resistant and highly acidic resin is the Nafion resin that DuPont company produces.
Described heteropolyacid catalyst can be both heteropolyacid and/or heteropolyacid acid salt, can be also the catalyzer of carried heteropoly acid and/or heteropolyacid acid salt.The strength of acid function H0 of heteropolyacid and acid salt thereof can be less than-13.15, and can be up to more than 300 ℃ life-time service.Described heteropolyacid and acid salt thereof comprise heteropolyacid and the acid salt thereof of Keggin structure, Dawson, Anderson structure, Silverton structure.Heteropolyacid and the acid salt thereof of preferred keggin structure, as 12 phospho-wolframic acid (H 3pW 12o 40xH 2o), 12 silicotungstic acid (H 4siW 12o 40xH 2o), 12 phosphomolybdate (H 3pMo 12o 40xH 2o), 12 molybdovanaphosphoric acid (H 3pMo 12-yv yo 40xH 2o) etc.The described preferred acid Tricesium dodecatungstophosphate of heteropolyacid acid salt salt (Cs 2.5h 0.5p 12wO 40), its strength of acid function H0 is less than-13.15, and specific surface area can reach 100m 2more than/g.Described heteropolyacid catalyst can be selected from one or more in above-mentioned preferred heteropolyacid and heteropolyacid acid salt.In the catalyzer of described carried heteropoly acid and/or heteropolyacid acid salt, carrier is generally SiO 2and/or gac.
Described solid acid catalyst can also be molecular sieve catalyst.Described molecular sieve can be one or more in H β, HY and HZSM-5, preferably by H β, the HY of fluorine or phosphorus modification and one or more in HZSM-5.These molecular sieves after fluorine, phosphorus modification, the acidity of the molecular sieve that can further improve and catalytic performance.
Below illustrate two kinds of embodiments of step (2).
The first embodiment:
Catalyzer adopts solid acid catalyst.In step (2), can adopt the reactor of one or more parallel connections, type of reactor can be selected from one or more in tank reactor, fixed-bed reactor, ebullated bed reactor and fluidized-bed reactor.Preferably adopt the tubular fixed-bed reactor of one or more parallel connections.More preferably adopt the shell shell and tube reactor of one or more parallel connections.The operating method of reactor can be both mode intermittently, can be also continuous mode, preferably adopted operate continuously mode.Fixed-bed reactor can adopt adiabatic or isothermal mode to operate.Adiabatic reactor can adopt cartridge reactor, catalyzer is fixed in reactor, reactor outer wall is incubated adiabatic, because addition esterification is thermopositive reaction, therefore need to control reactant concn to control reactor bed temperature rise, or be circulated to reactor inlet with diluting reaction substrate concentration after adopting partial reaction product cooling.Isothermal reactor can adopt shell shell and tube reactor, and catalyzer is fixed in tubulation, at shell side by water coolant to remove the liberated heat of reaction.
Temperature of reaction is generally 50～200 ℃, and optimizing temperature of reaction is 60～120 ℃.
The pressure of described addition esterification is relevant with temperature of reaction.Because addition esterification is carried out in liquid phase, so reaction pressure should guarantee that reaction is in liquid phase state.In general, reaction pressure is normal pressure～10MPa, and optimization pressure is normal pressure～1MPa.
The sour alkene mol ratio of described addition esterification is generally 0.2～20:1, and optimal conditions is 1.2～3:1.
In described addition esterification, liquid feeding air speed is generally 0.5～20h -1, optimal conditions is 0.5～5h -1.
Under these conditions, the cyclohexene conversion of addition esterification generally can reach more than 80%, and the selectivity of esterification can reach more than 99%.
The reaction product of step (2) is mainly comprised of hexanaphthene, benzene, acetic acid and hexalin acetate, also have a certain amount of unreacted tetrahydrobenzene, the separation of this product can be carried out in the addition esterification products separating unit that is provided with rectifying separation part and/or separation of extractive distillation part.A kind of optional separate mode is, esterification products logistics is carried out to rectifying separation, obtains by the logistics of hexanaphthene/tetrahydrobenzene/benzene, acetic acid stream and hexalin acetate logistics.The logistics of hexanaphthene/tetrahydrobenzene/benzene can be carried out Hydrogenation for hexanaphthene in the hydrogenation unit arranging in addition, and acetic acid stream can be used as a part for step (2) reaction feed, and hexalin acetate logistics is as the raw material of step (3) reaction.The present invention also can be further to hexanaphthene/tetrahydrobenzene/benzene logistics carry out separation of extractive distillation, a kind of optional separation scheme is, adopt extracting rectifying to be isolated into the logistics of hexanaphthene/tetrahydrobenzene and benzene logistics, benzene logistics can be used as a part for step (1) reaction feed, and the logistics of hexanaphthene/tetrahydrobenzene can be carried out Hydrogenation for hexanaphthene; Another kind of optional separation scheme is, adopt extracting rectifying to be isolated into hexanaphthene logistics, tetrahydrobenzene logistics and benzene logistics, benzene logistics can be used as a part for step (1) reaction feed, tetrahydrobenzene logistics can be used as a part for step (2) reaction feed, and hexanaphthene logistics goes out device as byproduct.The present invention can also remove the heavy constituent in above-mentioned hexalin acetate logistics by rectifying separation, the raw material that the hexalin acetate logistics removing after heavy constituent is reacted as step (3), and isolated heavy constituent logistics is as byproduct discharger.
Specifically, addition esterification products separating unit can arrange a decarburization six component towers and a desacetoxy tower.First addition esterification products enters decarburization six component towers and carries out separation, this tower can adopt atmospheric operation, by controlling tower reactor heating amount, reflux ratio, tower top and tower reactor produced quantity, from the logistics of overhead extraction hexanaphthene/tetrahydrobenzene/benzene, from the logistics of decarburization six component tower tower reactor extraction, enter desacetoxy tower and carry out separation, this tower also can adopt atmospheric operation, by controlling tower reactor heating amount, reflux ratio, tower top and tower reactor produced quantity, from overhead extraction acetic acid stream, from the logistics of tower reactor extraction hexalin acetate.Described esterification products separating unit can arrange an extractive distillation column again, and the logistics of hexanaphthene/tetrahydrobenzene/benzene is further separated into the logistics of hexanaphthene/tetrahydrobenzene and benzene logistics; Or two extractive distillation columns are set again, the logistics of hexanaphthene/tetrahydrobenzene/benzene is further separated into hexanaphthene logistics, tetrahydrobenzene logistics and benzene logistics.Addition esterification products separating unit can also arrange a de-heavy oil column again, from the hexalin acetate logistics of desacetoxy tower tower reactor extraction, enters de-heavy oil column, further removes the heavy constituent in logistics, thereby obtains the hexalin acetate logistics that removes heavy constituent.
The second embodiment:
Catalyzer adopts solid acid catalyst.In step (2), adopt the reactive distillation column of one or more parallel connections, when carrying out addition esterification, carry out the separation of reaction product, at the bottom of reactive distillation column tower, obtain hexalin acetate logistics, from reactive distillation column overhead, obtain hexanaphthene/benzene/acetic acid stream.
The theoretical plate number of described reactive distillation column is 10～150, between 1/3 to 2/3 position of theoretical plate number, arranges solid acid catalyst; With respect to the total fill able volume of catalyzer, liquid feeding air speed is 0.2～20h -1; The working pressure of reactive distillation column is-0.0099MPa to 5MPa; The temperature of catalyst filling zone is between 50～200 ℃; Reflux ratio is 0.1～100:1.
Described reactive distillation column is identical with common rectifying tower in form, generally tower body, overhead condenser, return tank, reflux pump, tower reactor and reboiler etc., consists of.The type of tower can be tray column, can be also packing tower, can also be both combinations.Adoptable tray column type comprises valve tray column, sieve-tray tower, bubble-plate column etc.The filler that packing tower uses can adopt random packing, as Pall ring, θ ring, Berl saddles, ladder ring packing etc.; Also can adopt structured packing, as corrugated plate packing, ripple silk net filler etc.
According to method provided by the present invention, in reactive distillation column, be furnished with solid acid catalyst.Know with those skilled in the art know that, catalyst arrangement mode in reactive distillation column should meet following 2 requirements: (1) wants to provide enough passages passing through for vapour-liquid two-phase, or there is a larger bed voidage (general requirement is more than at least 50%), with assurance vapour-liquid two-phase, can convection current pass through, and do not cause liquid flooding; (2) will have good mass-transfer performance, reactant will be delivered in catalyzer and react from fluid-phase, and simultaneous reactions product will transmit out from catalyzer.In existing document, disclose the decoration form of multiple catalysts in reactive distillation column, these decoration forms all can be the present invention and adopt.The decoration form of existing catalyzer in reaction tower can be divided into following three kinds: (1) is directly arranged in tower by catalyzer in the mode of fractional distillation filling-material, major way is by by a certain size and shape granules of catalyst and fractional distillation filling-material mechanically mixing or catalyzer is clipped between structured packing and forms whole filler with structured packing, or catalyzer is directly made to fractional distillation filling-material shape; (2) catalyzer packed in the permeable small vessels of gas-liquid and be arranged on the column plate of reaction tower, or by catalyst arrangement in the downtake of reaction tower; (3) catalyzer is directly packed in reaction tower in fixed bed mode, liquid phase directly flows through beds, and be that gas phase is set up special-purpose passage, adopt in this way at the position that catalyzer is housed, by beds and rectifying tower tray, be arranged alternately, liquid on tower tray enters next beds through downtake and redistributor, carries out addition reaction in bed, and the liquid of beds bottom enters next tower tray by liquid header.
Described reactive distillation column must have enough theoretical plate numbers and reaction stage number could meet reaction and separation processes requirement.The theoretical plate number of described reactive distillation column is preferably 30～100, between 1/3 to 2/3 position of theoretical plate number, arranges solid acid catalyst.
In the present invention, need to guarantee that reactant has enough residence time, to realize the conversion completely of tetrahydrobenzene.With respect to the total fill able volume of catalyzer, liquid feeding air speed is preferably 0.5～5h -1.
In the present invention, the working pressure of reactive distillation column can operate under negative pressure, normal pressure and pressurized conditions.The working pressure of reactive distillation column is preferably normal pressure to 1MPa.
The service temperature of reactive distillation column is relevant with the pressure of reactive distillation column, can, by regulating the working pressure of reaction tower to regulate the temperature distribution of reaction tower, make the temperature of catalyst filling zone within the scope of the active temperature of catalyzer.The temperature of catalyst filling zone is preferably between 60～120 ℃.
The reflux ratio of reactive distillation column should meet requirement separated and reaction simultaneously, generally, increases reflux ratio and is conducive to improve separating power and reaction conversion ratio, but can increase process energy consumption simultaneously.Described reflux ratio is preferably 0.5～10:1.
According to method provided by the invention, the raw material from the hexalin acetate logistics that obtains at the bottom of reactive distillation column as step (3) reaction.Hexanaphthene/benzene/the acetic acid stream obtaining for reactive distillation tower top, can adopt rectifying to be isolated into the logistics of hexanaphthene/benzene and acetic acid stream, and acetic acid stream can be used as a part for step (2) reaction feed; The logistics of hexanaphthene/benzene can be in the hydrogenation unit arranging in addition Hydrogenation for hexanaphthene, also can be isolated into hexanaphthene logistics and benzene logistics by extracting rectifying, benzene logistics can be used as a part for step (1) reaction feed, and hexanaphthene logistics goes out device as byproduct.Specifically, hexanaphthene/benzene/acetic acid stream can be carried out rectifying separation at a decarburization six component towers, and this tower can adopt atmospheric operation, by controlling tower reactor heating amount, reflux ratio, tower top and tower reactor produced quantity, from the logistics of overhead extraction hexanaphthene/benzene, from tower reactor extraction acetic acid stream; The logistics of hexanaphthene/benzene can be entered the hydrogenation unit of other setting, or enters extractive distillation column, is further separated into hexanaphthene logistics and benzene logistics.
Three, ethyl cyclohexyl ester through hydrogenation
According to method provided by the present invention, the hexalin acetate logistics being obtained by the separation of addition esterification products is admitted to ester through hydrogenation reactor and carries out hydrogenation reaction.In step (3), the reactor of one or more parallel connections can be set, type of reactor is selected from one or more in tank reactor, fixed-bed reactor, ebullated bed reactor and fluidized-bed reactor.In step (3), the tubular fixed-bed reactor of one or more parallel connections is preferably set.In step (3), the shell shell and tube reactor of one or more parallel connections is more preferably set, ester through hydrogenation catalyzer is fixed in tubulation, at shell side, by heat-eliminating medium, removes reaction liberated heat.
The second described catalyzer is ester through hydrogenation catalyzer.Although existing disclosed document is mainly the hydrogenation about carboxylate methyl ester or carboxylic acid, ethyl ester, employing fatty acid methyl ester hydrogenation as usual is produced higher alcohols, maleic acid methyl ester hydrogenation produces 1,4-butyleneglycol, 1,6-dimethyl adipate hydrogenation produces 1,6-hexylene glycol etc., there are no any report about the derivative carboxylicesters hydrogenation reaction of cycloalkanol, but inventor's discovery, the hydrogenation of hexalin acetate can adopt existing ester through hydrogenation catalyzer.The hydrogenation of ester generally adopts Cu-series catalyst, and ruthenium catalyst and precious metal series catalysts are the most conventional with Cu-series catalyst.Copper is that ester through hydrogenation catalyzer be take copper as Primary Catalysts, then one or more components of adding chromium, aluminium, zinc, calcium, magnesium, nickel, titanium, zirconium, tungsten, molybdenum, ruthenium, platinum, palladium, rhenium, lanthanum, thorium, gold etc. are as promotor or binder component.Copper is that ester through hydrogenation catalyzer can conveniently be buied from market, also can adopt coprecipitation method to produce.Common preparation method puts into the soluble salt solution of each metal and still, at certain temperature and stir speed (S.S.), add alkaline solution (sodium hydroxide, sodium carbonate, ammoniacal liquor, urea etc.) to neutralize PH8～12 growth precipitation, precipitation through aging, filter, the operation such as washing, dry, roasting, moulding forms, last in hydrogen atmosphere reduction can be made into final ester through hydrogenation catalyzer.The composition that ruthenium catalyst is general: Ru/Al 2o 3or Ru-Sn/Al 2o 3.The composition that precious metal series catalysts is general: Pt/Al 2o 3, Pd-Pt/Al 2o 3or Pd/C.
In the present invention, ester through hydrogenation catalyzer can be selected from one or more in Cu-series catalyst, ruthenium catalyst and precious metal series catalysts, is preferably Cu-series catalyst, more preferably contains the Cu-series catalyst of zinc and/or chromium.
Ester through hydrogenation reaction member both can mode intermittently operate, and also can carry out in a continuous manner.Intermittent reaction generally adopts reactor to make reactor, hexalin acetate and hydrogenation catalyst are dropped in reactor, pass into hydrogen and react under certain temperature and pressure, after reaction finishes, reaction product is adopted and drawn off from still, isolate product, then drop into next batch material and react.Continuous hydrogenation reaction can adopt shell shell and tube reactor, and hydrogenation catalyst is fixed in tubulation, at shell side by water coolant to remove the liberated heat of reaction.
Hexalin acetate hydrogenation reaction temperature is relevant with the hydrogenation catalyst of selection, and for copper series hydrocatalyst, general hydrogenation reaction temperature is 150～400 ℃, and optimizing temperature of reaction is 200～300 ℃.Reaction pressure is normal pressure～20MPa, and optimization pressure is 4～10MPa.
The control of the hydrogen ester mol ratio of hexalin acetate hydrogenation reaction is very important.High hydrogen ester is than the hydrogenation that is conducive to ester, but too high hydrogen ester is than the energy consumption that will increase hydrogen compression cycle.General hydrogen ester ratio is 1～1000:1, and optimal conditions is 5～100:1.
In hydrogenation reaction, the size of the charging air speed of ester with select the activity of catalyzer relevant.High activated catalyst can adopt higher air speed.For selected catalyzer, reaction conversion ratio reduces with the increase of reaction velocity.In order to meet certain transformation efficiency, air speed must be limited within the specific limits.The liquid feeding air speed of general ester is 0.1～20h -1, optimal conditions is 0.2～2h -1.If employing intermittent reaction, the reaction times is 0.5～20h, is preferably 1～5h.
Ester through hydrogenation reaction product mainly contains ethanol and hexalin, also may contain a certain amount of ethyl acetate and pimelinketone, also may contain a certain amount of unreacted hexalin acetate simultaneously, and a small amount of high boiling material (two polyketone), these mixtures can adopt the method for rectifying and/or extracting rectifying to carry out separation, preferably adopt rectifying separation.Rectifying separation can adopt intermittently scheme, also can adopt continuous flow scheme.Batch fractionating, is about to ester through hydrogenation product and drops in rectifying tower reactor, steams ethanol, ethyl acetate, hexalin, pimelinketone, hexalin acetate successively, the residual a small amount of high boiling material of tower reactor from tower top.The present invention preferably adopts continuous rectification to carry out separated ester through hydrogenation product.Continuous rectification need to utilize the separated various components of a series of rectifying tower.Can design various separation process according to the separated sequencing of each component, the flow scheme of preferred sequence separation of the present invention, ester through hydrogenation product separation unit sets gradually for de-ethanol tower, decylization hexanol tower, hexalin acetate recovery tower.Ester through hydrogenation product is introduced into de-ethanol tower separation and obtains ethanol, then enters the separation of decylization hexanol tower and obtains hexalin, finally enters hexalin acetate recovery tower and reclaims unreacted hexalin acetate, the residual a small amount of high boiling material carrying device of tower reactor.The present invention also can only arrange a rectifying tower, only separation of ethanol is gone out, and obtains the hexalin that contains a small amount of impurity.
In step (3), after reaction product separation, the hexalin acetate logistics obtaining can be used as a part for step (3) reaction feed.
The invention provides the device of a kind of coproduction hexalin and ethanol, comprise the benzene hydrogenation unit, addition esterification unit, addition esterification products separating unit, ester through hydrogenation reaction member and the ester through hydrogenation product separation unit that are connected successively.
Those skilled in the art, according to method provided by the invention, can easily determine the pipeline mode of connection of each unit and equipment room.
Described benzene hydrogenation unit is provided with the reactor of one or more parallel connections, and type of reactor is selected from fixed-bed reactor and/or tank reactor.
Described addition esterification unit is provided with the reactor of one or more parallel connections, and type of reactor is selected from one or more in tank reactor, fixed-bed reactor, ebullated bed reactor and fluidized-bed reactor.
Described addition esterification unit at least arranges a reactive distillation column.
Before described reactive distillation column, also series connection arranges a pre-esterification reactor device, and described pre-esterification reactor device is tank reactor, fixed-bed reactor, fluidized-bed reactor or ebullated bed reactor.
Described addition esterification products separating unit at least arranges a rectifying tower.
Described ester through hydrogenation reaction member arranges the reactor of one or more parallel connections, and type of reactor is selected from one or more in tank reactor, fixed-bed reactor, ebullated bed reactor and fluidized-bed reactor.Described ester through hydrogenation reaction member preferably arranges the shell shell and tube reactor of one or more parallel connections.
Described ester through hydrogenation product separation unit at least arranges a rectifying tower.
The present invention also provides the device of another kind of coproduction hexalin and ethanol, comprises the benzene hydrogenation unit, addition esterification unit, ester through hydrogenation reaction member and the ester through hydrogenation product separation unit that are connected successively; Described addition esterification unit arranges the reactive distillation column of one or more parallel connections.
Those skilled in the art, according to method provided by the invention, can easily determine the pipeline mode of connection of each unit and equipment room.
Described benzene hydrogenation unit is provided with the reactor of one or more parallel connections, and type of reactor is selected from fixed-bed reactor and/or tank reactor.
Before described reactive distillation column, also series connection arranges a pre-esterification reactor device; Described pre-esterification reactor device is tank reactor, fixed-bed reactor, fluidized-bed reactor or ebullated bed reactor.
Described ester through hydrogenation reaction member arranges the reactor of one or more parallel connections, and type of reactor is selected from one or more in tank reactor, fixed-bed reactor, ebullated bed reactor and fluidized-bed reactor.
Described ester through hydrogenation reaction member arranges the shell shell and tube reactor of one or more parallel connections.
Described ester through hydrogenation product separation unit at least arranges a rectifying tower.
The present invention provides a high-level efficiency, new technology path cheaply for producing hexalin.Feature of the present invention is: (1) esterification and ester through hydrogenation reaction all have very high selectivity, so atom utilization is very high; (2) process environment is friendly; (3) co-producing ethanol when producing hexalin, changes into the high and ethanol that market capacity is huge of price by mode indirectly by cheap acetic acid, greatly increases the economy of process; (4) adopt reactive distillation to carry out addition esterification, not only can significantly improve reaction efficiency, can also simplify separation of extractive distillation process, greatly reduce investment and operation cost.
Accompanying drawing explanation
Fig. 1 is the block diagram of the first embodiment of the present invention.
Fig. 2 is the block diagram of the second embodiment of the present invention.
Embodiment
Below in conjunction with the preferred two kinds of technical process of Brief Description Of Drawings the present invention.
In the first technical process, addition esterifier adopts tank reactor, tubular fixed-bed reactor, ebullated bed reactor or fluidized-bed reactor.As shown in Figure 1: benzene and hydrogen enter benzene hydrogenation device 1, under the effect of benzene hydrogenating catalyst, carry out hydrogenation reaction, the logistics of benzene hydrogenation product enters addition esterifier 2 through pipeline 11, mix with the acetic acid entering through pipeline 21, under the effect of solid acid catalyst, carry out addition esterification, the logistics of addition esterification products enters addition esterification products separating unit 3 through pipeline 22, through separation, obtain hexanaphthene logistics 31, tetrahydrobenzene logistics 32, benzene logistics 33, acetic acid stream 34 and hexalin acetate logistics 35, benzene logistics loops back benzene hydrogenation device 1, tetrahydrobenzene logistics and acetic acid stream loop back addition esterifier 2, hexalin acetate logistics enters ester through hydrogenation reactor 4, under the effect of ester through hydrogenation catalyzer, contact with hydrogen and carry out ester through hydrogenation reaction, the logistics of ester through hydrogenation product enters ester through hydrogenation product separation unit 5, through separation, obtain hexalin logistics 52, ethanol stream 53, hexalin acetate logistics 51 and high boiling material logistics 54, hexalin logistics and ethanol stream go out device as product, high boiling material logistics goes out device as byproduct, hexalin acetate logistics loops back ester through hydrogenation reactor 4.
In the second technical process, addition esterifier adopts reactive distillation column.As shown in Figure 2: benzene and hydrogen enter benzene hydrogenation device 1, under the effect of benzene hydrogenating catalyst, carry out hydrogenation reaction, the logistics of benzene hydrogenation product enters reactive distillation column 2 through pipeline 11, mixes, with the acetic acid entering through pipeline 21 under the effect of solid acid catalyst, carry out addition esterification, carry out the separation of addition esterification products simultaneously, from reactive distillation column 2 tower tops, obtain hexanaphthene/benzene/acetic acid stream, at the bottom of reactive distillation column 2 towers, obtain hexalin acetate logistics, hexanaphthene/benzene/acetic acid stream enters addition esterification products separating unit 3 through pipeline 22, through separation, obtain hexanaphthene logistics, benzene logistics and acetic acid stream, hexanaphthene logistics goes out device as byproduct, benzene logistics loops back benzene hydrogenation device 1, acetic acid stream loops back addition esterifier 2, hexalin acetate logistics enters ester through hydrogenation reactor 4, under the effect of ester through hydrogenation catalyzer, contact with hydrogen and carry out ester through hydrogenation reaction, the logistics of ester through hydrogenation product enters ester through hydrogenation product separation unit 5, through separation, obtain hexalin logistics 52, ethanol stream 53, hexalin acetate logistics 51 and high boiling material logistics 54, hexalin logistics and ethanol stream go out device as product, high boiling material logistics goes out device as byproduct, hexalin acetate logistics loops back ester through hydrogenation reactor 4.
Further illustrate by the following examples the present invention, but not thereby limiting the invention.
Embodiment 1
The method of the present embodiment explanation benzene selective hydrogenation to prepare cyclohexene.
By benzene and hydrogen in molar ratio 1:3 inject the hydrogenator that is filled with ruthenium beaded catalyst, under the condition of 135 ℃ of temperature of reaction, pressure 4.5MPa, residence time 15min, carry out benzene hydrogenation, reaction product is isolated after hydrogen, collects product liquid, moves continuously 1000h.After off-test, the product liquid of collecting is carried out to gas chromatographic analysis, it consists of: benzene 53.3%, tetrahydrobenzene 35.4%, hexanaphthene 11.3%.
Embodiment 2
By 100mL macropore strong acid Hydrogen ion exchange resin, (laboratory is synthetic by classical literature method, the styrene solution that contains 15% Vinylstyrene is carried out to suspension copolymerization and make Archon, then make through concentrated acid sulfonation, and recording its exchange capacity is 5.2mmolH +/ g butt) pack φ 32 * 4 * 1000mm into the middle part in the stainless steel tube reactor of chuck, a certain amount of quartz sand is filled at two ends.By acetic acid and tetrahydrobenzene raw material (benzene 53.3%, tetrahydrobenzene 35.4%, hexanaphthene 11.3%) by certain flow, by volume pump, squeeze in reactor and react respectively, in reaction tubes external jacket, pass into hot water to control temperature of reaction, by reactor outlet back pressure valve, control reactor pressure.Reactor outlet product carries out on-line chromatograph analysis by online sampling valve sampling, by product composition, calculates cyclohexene conversion rate and hexalin acetate selectivity.Reaction conditions and the results are shown in Table 1.
As shown in Table 1, adopt strong acid ion exchange resin catalyst tetrahydrobenzene raw material and acetic acidreaction, cyclohexene conversion rate is greater than 80%, and ester products selectivity is greater than 99%, moves 600 hours, and catalyst activity and selectivity is stablized constant.
Table 1 strong-acid ion exchange resin catalysis acetic acid and hexanaphthene/tetrahydrobenzene/phenyl ester testing data
Figure BDA00002625566000191
Figure BDA00002625566000201
Embodiment 3
Testing apparatus, method and raw material are with embodiment 2, difference is that catalyzer is that (the H beta-molecular sieve that is 50 by silica alumina ratio is through 85% phosphoric acid modification, then mediates extruded moulding with aluminum oxide, through 120 ℃ of oven dry for the H beta-molecular sieve catalyzer of phosphorus modification, 500 ℃ of roastings make, and phosphorus content is 2%).Reaction conditions and the results are shown in Table 2.From table 2, tetrahydrobenzene and acetic acidreaction transformation efficiency are greater than 80%, and ester products selectivity is greater than 99%, move 480 hours, and catalyst activity and selectivity is stablized constant.
Table 2 H beta-molecular sieve catalyst acetic acid and hexanaphthene/tetrahydrobenzene/phenyl ester testing data
Figure BDA00002625566000202
Embodiment 4
The addition esterification products of collecting embodiment 2 and 3, carries out rectifying separation test.Rectifying adopts the glass tower rectifier unit that high 2m diameter is 40mm, king-post is equipped with the stainless steel Dixon ring highly efficient distilling filler of Ф 3mm, and tower reactor is the L glass flask of volume 5, and charge amount is 4L, by electric mantle, tower reactor is heated, by voltate regulator, regulate tower reactor heating to measure.The backflow of tower adopts reflux ratio setter to control.Rectifying separation the results are shown in Table 3.
Table 3 addition esterification products rectifying separation test-results
Embodiment 5～6 is for illustrating the method that adopts reactive distillation to prepare hexalin acetate.
The test of carrying out in embodiment 5～6 is all that the reactive distillation model test device in following specification carries out: the main body of mode device is that diameter (internal diameter) is 50mm, height is the stainless head tower of 3m, the bottom of tower connects the tower reactor that volume is 5L, in still, dispose the electrically heated rod of 10KW, this heating rod is controlled tower reactor heating amount by intelligent controller by silicon controlled rectifier (SCR).It is 0.5m that tower top is connected with heat interchanging area 2condenser, overhead vapours enters the return tank that a volume is 2L after this condenser condenses becomes liquid.Liquid in return tank is partly refluxed to reaction tower through reflux pump, and part extraction is as light constituent.The operating parameters of tower is shown and is controlled by intelligent Automation Control Instruments.Tower quantity of reflux is controlled by return valve, and overhead extraction amount is controlled by the fluid level controller of return tank.Tower reactor produced quantity regulates tower reactor blow-off valve to control by tower bottoms level controller.Acetic acid and tetrahydrobenzene raw material are respectively charged in 30L storage tank, and are squeezed into after being preheating to certain temperature in corresponding preheater and entered reaction tower by volume pump, input speed by volume pump control, electronic scales accurate measurement.
Embodiment 5
By high temperature resistant sulfonic acid ion exchange resin, (trade mark is Amberlyst 45, by Rhom & Hass company, produced) with multistage high speed disintegrator, be ground into the powder that granularity is less than 200 orders (0.074mm), add perforating agent, lubricant, oxidation inhibitor and tackiness agent to mix on high-speed mixer, again on Banbury mixer in 180 ℃ of banburying 10min, material is plastified completely, injecting afterwards mould, to make diameter be 5mm, high 5mm, wall thickness is 1mm Raschig ring type resin catalyst filler.Respectively pack up and down the middle part of this filler 1950mL loading pattern reaction tower (high 1m is equivalent to 8 blocks of theoretical trays) into the glass spring filler 1950mL that diameter is 3mm, long 6mm (loading height is 1m, is equivalent to 10 blocks of theoretical trays).After tetrahydrobenzene raw material and acetic acid are squeezed into preheater preheating by volume pump respectively, enter reaction tower, regulate tower reactor heating amount and trim the top of column amount to react continuously, the reaction conditions under stable operation and reaction result are in Table 4.
The reactive distillation testing data of table 4 is high temperature resistant sulfonic acid ion exchange resin catalyzer
Figure BDA00002625566000221
According to the transformation efficiency 98.8% of testing data ring hexene, hexalin acetate selectivity 98.0%.
Embodiment 6
By the ball-type H of φ 3～4 0.5cs 2.5pW 12o 40/ SiO 2catalyzer is (by H 0.5cs 2.5pW 12o 40powder and granularity are less than 200 object silochrom powder, after fully mixing in mixer, in coater, take silicon sol as bonder roller forming, then drying, roasting form) sandwich in titanium wire network ripple plate, making diameter is the cylinder shape structured packing of 50mm, high 50mm.Respectively packing up and down the middle part of this packing type catalyzer L loading pattern reaction tower (high 1m is equivalent to 12 blocks of theoretical trays) into diameter is that 4mm, height are the 1950mL glass spring filler of 4mm (loading height is 1m, is equivalent to 15 blocks of theoretical trays).After tetrahydrobenzene raw material and acetic acid are squeezed into preheater preheating by volume pump respectively, enter reaction tower, regulate tower reactor heating amount and trim the top of column amount to react continuously, the reaction conditions under stable operation and reaction result are in Table 5.
Table 5 H 0.5cs 2.5pW 12o 40/ SiO 2the reactive distillation testing data of catalyzer
Figure BDA00002625566000231
According to the transformation efficiency 99.35% of testing data ring hexene, hexalin acetate selectivity 99.6%.
Embodiment 7～8 is for illustrating the method for hydrotreating of hexalin acetate.
Embodiment 7
Hydrogenating materials is the hexalin acetate of purity 99.6%.
By 40g copper zinc-aluminium ester through hydrogenation catalyzer, (laboratory is synthetic, consists of CuO 40.5%, ZnO29.6%, Al 2o 330.4%.By the nitrate solution of copper, zinc, chromium, add sodium hydroxide solution to be neutralized to PH=9.0, through centrifugation, washing, dry, compression molding, roasting makes) pack φ 20 * 2.5 * 800mm into the middle part in the stainless steel tube reactor of chuck, a certain amount of quartz sand is filled at two ends.Pass into hydrogen (500mL/min) and after reductase 12 4h, be down to the temperature and pressure of hydrogenation reaction under 280 ℃, 6MPa condition.Hexalin acetate is squeezed in reactor by volume pump, and hydrogen enters reactive system through mass flow controller and carries out hydrogenation reaction, by passing into thermal oil in reaction tubes external jacket, controls temperature of reaction, by reactor outlet back pressure valve, controls reactor pressure.Reaction product is carried out on-line chromatograph analysis by the straight line sampling valve sampling at reactor rear portion.Reaction conditions and the results are shown in Table 6.Table 6 result shows, adopts CuZnAl catalyst, and hexalin acetate hydrogenation reaction transformation efficiency can reach more than 99.0%, and hexalin selectivity is greater than 99.9%, moves 1000 hours, and transformation efficiency and selectivity all do not decline.
The hexalin acetate hydropyrolysis experiment data of table 6 copper zinc-aluminium ester through hydrogenation catalyzer
Figure BDA00002625566000251
Embodiment 8
Hydrogenating materials is the hexalin acetate of purity 99.6%.
40g copper chromium ester through hydrogenation catalyzer is (commercially available, Taiyuan City Xin Jida Chemical Co., Ltd. produces, the trade mark is C1-XH-1, CuO content is 55%, diameter 5mm tablet, be broken into 10～20 order particles) pack φ 20 * 2.5 * 800mm into the middle part in the stainless steel tube reactor of chuck, a certain amount of quartz sand is filled at two ends.Pass into hydrogen (500mL/min) after reductase 12 4h, is down to and reacts to obtain temperature and pressure under 280 ℃, 6MPa condition.Hexalin acetate is squeezed in reactor by volume pump, and hydrogen enters reactive system through mass flow controller and carries out hydrogenation reaction, by passing into thermal oil in reaction tubes external jacket, controls temperature of reaction, by reactor outlet back pressure valve, controls reactor pressure.Reaction product is carried out on-line chromatograph analysis by the straight line sampling valve sampling at reactor rear portion.Reaction conditions and the results are shown in Table 7.Table 7 result shows, adopts CuZnAl catalyst, and hexalin acetate hydrogenation reaction transformation efficiency can reach more than 98.0%, and hexalin selectivity is greater than 99.9%, moves 500 hours, and transformation efficiency and selection all do not decline.
The hexalin acetate hydropyrolysis experiment data of table 7 copper chromium ester through hydrogenation catalyzer
Embodiment 9
The reaction product 4000g that collects example 7～8, carries out rectifying separation test.Rectifying adopts high 2m glass tower, and king-post is equipped with the stainless steel Dixon ring highly efficient distilling filler of Ф 3mm, and tower reactor is 5L glass flask, by electric mantle, heats, and by voltate regulator, regulates tower reactor heating to measure.The backflow of tower adopts reflux ratio setter to control.Rectifying separation the results are shown in Table 8.
The rectifying separation testing data of table 8 hexalin acetate hydrogenation products

Claims (50)

1. a method for coproduction hexalin and ethanol, comprising:
(1), under the condition of preparing cyclohexene from benzene added with hydrogen, benzene and hydrogen generation hydrogenation reaction, obtain containing tetrahydrobenzene logistics;
(2) what step (1) is obtained contacts with acetic acid containing tetrahydrobenzene logistics, and addition esterification occurs under the effect of the first catalyzer; Reaction product is carried out to separation, obtain hexalin acetate logistics;
(3) hexalin acetate logistics step (2) being obtained contacts with hydrogen, and ester through hydrogenation reaction occurs under the effect of the second catalyzer; Reaction product is carried out to separation, obtain hexalin and ethanol.
2. in accordance with the method for claim 1, it is characterized in that, in step (1), the benzene hydrogenating catalyst of use is ruthenium catalyst.
3. in accordance with the method for claim 2, it is characterized in that, in step (1), the benzene hydrogenating catalyst of use is the ruthenium catalyst containing cobalt and/or zinc.
4. in accordance with the method for claim 1, it is characterized in that, in step (2), the first described catalyzer is solid acid catalyst.
5. in accordance with the method for claim 4, it is characterized in that, described solid acid catalyst is selected from one or more in strong acid ion exchange resin catalyzer, heteropolyacid catalyst and molecular sieve catalyst.
6. in accordance with the method for claim 5, it is characterized in that, described strong acid ion exchange resin is macropore sulfonic acid type polystyrene-divinylbenzene resin or the sulfonic resin after halogen atom modification.
7. in accordance with the method for claim 5, it is characterized in that, described heteropolyacid catalyst is heteropolyacid and/or heteropolyacid acid salt, or the catalyzer of carried heteropoly acid and/or heteropolyacid acid salt.
8. in accordance with the method for claim 7, it is characterized in that, described heteropolyacid catalyst is the heteropolyacid of keggin structure and/or the heteropolyacid acid salt of keggin structure, or the catalyzer of the heteropolyacid of load keggin structure and/or the heteropolyacid acid salt of keggin structure.
9. in accordance with the method for claim 7, it is characterized in that, in the catalyzer of described carried heteropoly acid and/or heteropolyacid acid salt, carrier is SiO 2and/or gac.
10. in accordance with the method for claim 5, it is characterized in that, described heteropolyacid catalyst is selected from one or more in 12 phospho-wolframic acids, 12 silicotungstic acids, 12 phosphomolybdate, 12 molybdovanaphosphoric acids and acid phospho-wolframic acid cesium salt.
11. in accordance with the method for claim 5, it is characterized in that, described molecular sieve is one or more in H β, HY and HZSM-5.
12. in accordance with the method for claim 11, it is characterized in that, described molecular sieve is by H β, the HY of fluorine or phosphorus modification and one or more in HZSM-5.
13. in accordance with the method for claim 4, it is characterized in that, in step (2), adopts the reactor of one or more parallel connections, and type of reactor is selected from one or more in tank reactor, fixed-bed reactor, ebullated bed reactor and fluidized-bed reactor.
14. in accordance with the method for claim 4, it is characterized in that, in step (2), adopts the shell shell and tube reactor of one or more parallel connections.
15. according to the method described in claim 13 or 14, it is characterized in that, in step (2), temperature of reaction is 50～200 ℃.
16. in accordance with the method for claim 15, it is characterized in that, in step (2), temperature of reaction is 60～120 ℃.
17. according to the method described in claim 13 or 14, it is characterized in that, in step (2), sour alkene mol ratio is 0.2～20:1.
18. in accordance with the method for claim 17, it is characterized in that, in step (2), sour alkene mol ratio is 1.2～3:1.
19. according to the method described in claim 13 or 14, it is characterized in that, in step (2), liquid feeding air speed is 0.5～20h -1.
20. in accordance with the method for claim 19, it is characterized in that, in step (2), liquid feeding air speed is 0.5～5h -1.
21. according to the method described in claim 13 or 14, it is characterized in that, in step (2), esterification products logistics is carried out to rectifying separation, obtain by the logistics of hexanaphthene/tetrahydrobenzene/benzene, acetic acid stream and hexalin acetate logistics, acetic acid stream is as a part for step (2) reaction feed.
22. in accordance with the method for claim 21, it is characterized in that, adopt extracting rectifying that the logistics of described hexanaphthene/tetrahydrobenzene/benzene is separated into hexanaphthene logistics, tetrahydrobenzene logistics and benzene logistics, benzene logistics is as a part for step (1) reaction feed, and tetrahydrobenzene logistics is as a part for step (2) reaction feed.
23. in accordance with the method for claim 4, it is characterized in that, in step (2), adopt the reactive distillation column of one or more parallel connections, when carrying out addition esterification, carry out the separation of reaction product, at the bottom of reactive distillation column tower, obtain hexalin acetate logistics, from reactive distillation column overhead, obtain hexanaphthene/benzene/acetic acid stream.
24. in accordance with the method for claim 23, it is characterized in that, the theoretical plate number of described reactive distillation column is 10～150, between 1/3 to 2/3 position of theoretical plate number, arranges solid acid catalyst; With respect to the total fill able volume of catalyzer, liquid feeding air speed is 0.2～20h -1; The working pressure of reactive distillation column is-0.0099MPa to 5MPa; The temperature of catalyst filling zone is between 50～200 ℃; Reflux ratio is 0.1～100:1.
25. in accordance with the method for claim 24, it is characterized in that, the theoretical plate number of described reactive distillation column is 30～100.
26. in accordance with the method for claim 24, it is characterized in that, with respect to the total fill able volume of catalyzer, liquid feeding air speed is 0.5～5h -1.
27. in accordance with the method for claim 24, it is characterized in that, the working pressure of reactive distillation column is that normal pressure is to 1MPa.
28. in accordance with the method for claim 24, it is characterized in that, the temperature of catalyst filling zone is between 60～120 ℃.
29. in accordance with the method for claim 24, it is characterized in that, the reflux ratio of reactive distillation column is 0.5～10:1.
30. in accordance with the method for claim 1, it is characterized in that, in step (3), the second described catalyzer is ester through hydrogenation catalyzer.
31. in accordance with the method for claim 30, it is characterized in that, described ester through hydrogenation catalyzer is selected from one or more in Cu-series catalyst, ruthenium catalyst and precious metal series catalysts.
32. according to the method described in claim 31, it is characterized in that, in step (3), described catalyzer is Cu-series catalyst, and hydrogenation reaction temperature is 150～400 ℃, and reaction pressure is normal pressure～20MPa, hydrogen ester mol ratio is 1～1000:1, and liquid feeding air speed is 0.1～20h -1.
33. according to the method described in claim 32, it is characterized in that, hydrogenation reaction temperature is 200～300 ℃, and reaction pressure is 4～10MPa, and hydrogen ester mol ratio is 5～100:1, and liquid feeding air speed is 0.2～2h -1.
34. according to the method described in claim 31, it is characterized in that, in step (3), described Cu-series catalyst is the Cu-series catalyst containing zinc and/or chromium.
35. in accordance with the method for claim 1, it is characterized in that, in step (3), after reaction product separation, also obtains hexalin acetate logistics, the part of (3) reaction feed using it as step.
The device of 36. 1 kinds of coproduction hexalin and ethanol, comprises the benzene hydrogenation unit, addition esterification unit, addition esterification products separating unit, ester through hydrogenation reaction member and the ester through hydrogenation product separation unit that are connected successively.
37. according to the device described in claim 36, it is characterized in that, described benzene hydrogenation unit is provided with the reactor of one or more parallel connections, and type of reactor is selected from fixed-bed reactor and/or tank reactor.
38. according to the device described in claim 36, it is characterized in that, described addition esterification unit is provided with the reactor of one or more parallel connections, and type of reactor is selected from one or more in tank reactor, fixed-bed reactor, ebullated bed reactor and fluidized-bed reactor.
39. according to the device described in claim 36, it is characterized in that, addition esterification unit at least arranges a reactive distillation column.
40. according to the device described in claim 39, it is characterized in that, before described reactive distillation column, also series connection arranges a pre-esterification reactor device; Described pre-esterification reactor device is tank reactor, fixed-bed reactor, fluidized-bed reactor or ebullated bed reactor.
41. according to the device described in claim 36, it is characterized in that, described addition esterification products separating unit at least arranges a rectifying tower.
42. according to the device described in claim 36, it is characterized in that, described ester through hydrogenation reaction member arranges the reactor of one or more parallel connections, and type of reactor is selected from one or more in tank reactor, fixed-bed reactor, ebullated bed reactor and fluidized-bed reactor.
43. according to the device described in claim 36, it is characterized in that, described ester through hydrogenation reaction member arranges the shell shell and tube reactor of one or more parallel connections.
44. according to the device described in claim 36, it is characterized in that, described ester through hydrogenation product separation unit at least arranges a rectifying tower.
The device of 45. 1 kinds of coproduction hexalin and ethanol, comprises the benzene hydrogenation unit, addition esterification unit, ester through hydrogenation reaction member and the ester through hydrogenation product separation unit that are connected successively; Described addition esterification unit arranges the reactive distillation column of one or more parallel connections.
46. according to the device described in claim 45, it is characterized in that, described benzene hydrogenation unit is provided with the reactor of one or more parallel connections, and type of reactor is selected from fixed-bed reactor and/or tank reactor.
47. according to the device described in claim 45, it is characterized in that, before described reactive distillation column, also series connection arranges a pre-esterification reactor device; Described pre-esterification reactor device is tank reactor, fixed-bed reactor, fluidized-bed reactor or ebullated bed reactor.
48. according to the device described in claim 45, it is characterized in that, described ester through hydrogenation reaction member arranges the reactor of one or more parallel connections, and type of reactor is selected from one or more in tank reactor, fixed-bed reactor, ebullated bed reactor and fluidized-bed reactor.
49. according to the device described in claim 45, it is characterized in that, described ester through hydrogenation reaction member arranges the shell shell and tube reactor of one or more parallel connections.
50. according to the device described in claim 45, it is characterized in that, described ester through hydrogenation product separation unit at least arranges a rectifying tower.
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