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

Abstract

The invention provides a method for the co-production of cyclohexanol and ethanol. The method takes benzene as the starting material and co-produces cyclohexyl through selective hydrogenation of benzene, addition and esterification of cyclohexene, and hydrogenation of cyclohexyl acetate. alcohol and ethanol. The invention also provides a device for realizing the method. The characteristics of the present invention are: (1) both esterification and ester hydrogenation reactions have high selectivity and high atom utilization; (2) the process is environmentally friendly; (3) co-production of ethanol while producing cyclohexanol ; (4) Using reactive distillation for addition esterification can not only significantly improve the reaction efficiency, but also simplify the separation process of extractive distillation.

Description

A method and device for co-producing cyclohexanol and ethanol

technical field

The invention relates to a method and device for co-producing cyclohexanol and ethanol.

Background technique

Both cyclohexanol and ethanol are important chemical raw materials and solvents. Cyclohexanol is mainly used to produce nylon 6, nylon 66 and other products; ethanol is not only a raw material for synthetic esters and other chemical products, but also widely used as a fuel additive for gasoline.

The method of industrially synthesizing ethanol is mainly the direct hydration method of ethylene, but in some countries rich in agricultural and sideline products, the fermentation method is still the main method of ethanol production. Due to the large population and insufficient arable land in our country, and the problem of "competing with the mouth for food" in the production of ethanol by fermentation, the fermentation method does not conform to the national conditions of our country. In addition, the pollution of the fermentation method is also relatively serious. my country's oil resources are relatively insufficient, and the price of ethylene is greatly affected by fluctuations in international oil prices. Therefore, the application of ethylene hydration in my country will face certain pressure on raw material costs. In addition, the reaction conditions of the ethylene direct hydration method are relatively harsh and need to be carried out under high temperature and high pressure. In summary, the development of new ethanol synthesis process routes is an inevitable requirement for technological and economic development.

CN1022228831A discloses a catalyst for preparing ethanol by gas-phase hydrogenation of acetic acid. The catalyst is composed of three parts: main active component, auxiliary agent and carrier; the carrier is any one of activated carbon, graphite or multi-walled carbon nanotubes, and the main active The component is any one or two of metal W or Mo, and the auxiliary agent is one or more of Pd, Re, Pt, Rh or Ru; the content of the main active component is 0.1-30.0% of the weight of the catalyst, and the auxiliary agent is The content of the agent is 0.1-10.0% of the weight of the catalyst, and the balance is the carrier.

CN102149661A discloses a method for the direct and selective production of ethanol from acetic acid using a platinum/tin catalyst, comprising: a feed stream containing acetic acid and hydrogen is contacted with a hydrogenation catalyst at a relatively high temperature, and the hydrogenation catalyst is included in a suitable The group of platinum and tin on the catalyst support and optionally a third metal supported on the support, wherein the third metal is selected from the group consisting of the following metals: palladium, rhodium, ruthenium, rhenium. Iridium, Chromium, Copper, Molybdenum, Tungsten, Vanadium and Zinc.

Industrially, the production methods of cyclohexanol mainly include cyclohexane air oxidation method, phenol hydrogenation method and cyclohexene hydration method, among which cyclohexane oxidation method is the most common application.

Cyclohexane oxidation is currently the most important cyclohexanol production process. The process uses an oxidant (usually air) to oxidize cyclohexane to cyclohexyl hydroperoxide, and cyclohexyl hydroperoxide is decomposed to obtain a mixture of cyclohexanol and cyclohexanone (commonly known as KA oil). The advantages of this process are mild oxidation process conditions, less slagging, and long continuous operation period. The disadvantage is that the process route is long, the energy consumption is high, and the pollution is large. The conversion rate of cyclohexane in this process is only 3-5%; especially in the decomposition process of cyclohexyl hydroperoxide, the selectivity of cyclohexanol is poor, and the yield In addition, this process also produces a large amount of waste lye that is difficult to handle, which is still a worldwide environmental protection problem.

Phenol hydrogenation is a relatively clean technical route for the production of cyclohexanol, and has the advantages of short process flow and high product purity. The hydrogenation of phenol to cyclohexanol mainly adopts the gas phase hydrogenation method. This method usually adopts 3 to 5 reactors connected in series. Under the action of supported Pd catalyst, the yield of cyclohexanone and cyclohexanol can reach 90%-95% at 140-170°C and 0.1MPa. However, this process needs to vaporize phenol (heat of vaporization 69kJ·mol ^-1 ) and methanol (heat of vaporization 35.2kJ·mol ^-1 ), which requires high energy consumption, and the catalyst is prone to carbon deposition during use, resulting in a decrease in activity, and the shortage of phenol , expensive and the use of noble metal catalysts limit the industrial application of this method.

In the 1980s, Asahi Kasei Corporation of Japan developed a process for producing cyclohexene by partial hydrogenation of benzene and hydration of cyclohexene to produce cyclohexanol, and realized industrialization in 1990. Related Chinese patent applications include CN 1079727A and CN 1414933A and CN101796001A. Cyclohexene hydration is a relatively new production method of cyclohexanol. This method has high reaction selectivity and almost no three wastes in the process. However, there are disadvantages such as low reaction conversion rate and high requirements for cyclohexene purity. If high silicon ZSM-5 catalyst is used, the conversion rate of cyclohexene is only 12.5% after staying in two series slurry reactors for 2 hours.

The traditional methods for producing cyclohexene are dehydration of cyclohexanol and dehydrohalogenation of halocyclohexane. Partial hydrogenation of benzene and oxidative dehydrogenation of cyclohexane are two other methods for preparing cyclohexene. The partial hydrogenation of benzene to prepare cyclohexene mainly includes gas-phase method, liquid-phase method and homogeneous complexation hydrogenation method. The liquid-phase method can adopt a four-phase reaction system of gas (hydrogen), liquid (benzene), liquid (polar solvent), and solid (catalyst), such as the method of feeding benzene and hydrogen into a slurry containing catalyst and water . If a four-phase reaction system is adopted, the amount of water must be at least sufficient to form an oil-water two-phase, and metal salts are usually added to the water, and the preferred metal salts are zinc sulfate or cobalt sulfate. The reaction conditions of the liquid phase method are generally: reaction temperature 25-250°C; hydrogen partial pressure 0.1-20MPa, catalyst dosage 0.001-0.2 times the weight of water. Catalysts for the partial hydrogenation of benzene generally use one or more of metals such as Pt, Pd, Ru, Rh and Ni as the main catalyst component; in order to further improve hydrogenation activity and selectivity, K, One or more of metals such as Zr, Hf, Co, Cu, Ag, Fe, Mo, Cr, Mn, Au, la and Zn are used as additive components. The partial hydrogenation catalyst of benzene can be a supported or non-supported catalyst. The supported method can be ion exchange method, spray method, impregnation method, evaporative drying method, etc. The carrier used can be natural clay, sepiolite, ZrO ₂ , SiO ₂ , TiO ₂ , Al ₂ O ₃ , La ₂ O ₃ , activated carbon, insoluble sulfate, insoluble phosphate or molecular sieve, etc. Taking ruthenium-based catalysts as an example, supported catalysts can be prepared by impregnating ruthenium salts alone or together with other metal salts on the carrier, followed by drying and reducing; unsupported catalysts can be prepared by ruthenium salts alone or together with other metal salts It can be prepared by precipitation, followed by drying and reduction, or by direct reduction of ruthenium compounds or mixtures of ruthenium compounds and other metal compounds. The ruthenium-based catalyst and the method and device for hydrogenation of benzene are described in detail in CN102264471A.

Cyclohexyl acetate is a liquid with banana or apple aroma, and the fruity flavor prepared with it is widely used in food, beverage and cosmetic industries. In addition, cyclohexyl acetate has good solubility for resins, and is often used as an environmentally friendly solvent for high-grade paints and paints. Recently, the present inventors also found that cyclohexanol can be produced by hydrogenation and co-produce ethanol. It can be predicted that cyclohexyl acetate will become an important organic synthesis intermediate.

At present, the synthetic method of cyclohexyl acetate in industry is the esterification reaction of acetic acid and cyclohexanol. The alkyd esterification reaction needs to be carried out smoothly under the action of an acidic catalyst. Song Guijia, Wu Xionggang (Chemical Propellants and Polymer Materials, 2009, V0l.7(2):P31-33), reviewed the research progress on the synthesis of cyclohexyl acetate through the esterification of acetic acid and cyclohexanol.

JPA254634/1989 discloses a preparation method of cyclohexanol and cyclohexyl acetate, using a strongly acidic ion exchange resin as a catalyst to synthesize cyclohexanol and cyclohexyl acetate by the reaction of aqueous acetic acid and cyclohexene. The best result mentioned in this literature example is that the conversion rate of cyclohexene is 62.7%, the yield of cyclohexanol is 18.4%, and the yield of cyclohexyl acetate is 43.7%.

CN1023115C and JP Ping-313447 disclose a method for preparing cyclohexanol, using ZSM5 or high silica zeolite as a catalyst to synthesize cyclohexanol and cyclohexyl acetate by reacting acetic acid and cyclohexene in the presence of water. In this document, the yields of cyclohexanol and cyclohexyl acetate were only 12.5% and 65% respectively at 120°C for 4 hours.

EP0461580A2 and USP5254721 disclose a method for preparing cyclohexyl acetate from acetic acid and cyclohexene. The method adopts a tungsten-containing heteropolyacid catalyst, and the crystal water content in the heteropolyacid molecule is preferably 0-3. The best result given in the literature is that the dodecasilic tungstic acid catalyst completely free of crystal water obtained by roasting at 370°C for 3 hours, in a 200mL autoclave, add 61.5g of acetic acid, 13.5g of cyclohexene, and 5g of catalyst, Under the conditions of 0.5MPa and 130℃ for 0.5h, the conversion of cyclohexene was 95.2%, and the selectivity of cyclohexyl acetate was 99.2%. It can be seen that under the condition of very high acid-to-ene ratio, cyclohexene cannot be completely converted.

As can be known from the existing published literature, the existing literature has disclosed various solid acid catalysts for the addition esterification of acetic acid and cyclohexene. The addition esterification reaction generally adopts a tank reactor, and the reaction raw material is pure cyclohexene , even with a high acid-to-ene ratio, it is difficult to achieve complete conversion of cyclohexene.

Reactive distillation has been widely used in processes such as alcohol-ene etherification, alcohol esterification, transesterification, ester hydrolysis, acetal reaction, etc., but so far, no reactive distillation has been used for the addition and esterification of acetic acid and cyclohexene process reports.

CN86105765A proposes a method for preparing alcohol by hydrogenation of carboxylic acid ester, which method is to hydrogenate carboxylic acid ester at high temperature, normal pressure or high pressure in the presence of a reduction-activated solid copper-containing catalyst, and the catalyst removes copper In addition, it also contains magnesium, at least one of lanthanide metals or actinide metals. The catalyst is represented by the following general formula before reduction activation: Cu _a M ¹ M ² _b A _c O _x , M ¹ is at least one of magnesium, lanthanide metal or actinide metal, M ² is selected from Ca, Mo, Rh , Pt, Cr, Zn, Al, Ti, V, Ru, Re, Pd, Ag and Au; A is an alkali metal; a is 0.1-4; b is 0-1.0; C is 0-0.5; x is A number that can meet the requirements of other elements for the total valence of oxygen. The alkali metal in the catalyst is an optional component which is introduced into the catalyst in the form of an alkali metal salt. The applicable carboxylate of this method and catalyzer is C1-C24 acyclic monovalent or divalent, saturated or unsaturated, linear or branched chain carboxylate, does not relate to cycloalkanols like cyclohexanol in the literature preparation.

CN1075048C proposes a method and catalyst for direct hydrogenation of carboxylic acid esters, comprising contacting and reacting one or more esters with hydrogen in the presence of a catalyst containing a copper compound, a zinc compound and at least one selected The catalyst is prepared from aluminum, zirconium, magnesium, a compound of a rare earth element or a mixture thereof as its components by calcining these catalyst components at a temperature ranging from 200 to less than 400°C, the method being in a liquid phase , Carried out at 170-250°C and 20.7-138 bar gauge pressure. The carboxylate suitable for the method and catalyst is C6-C22 dimethyl ester obtained by transesterification of natural oil, C6-C66 natural triglyceride or C6-C44 compound obtained by transesterification of natural triglyceride .

US4939307 proposes a process for ester hydrogenation to alcohol. The general formula is R ₁ -CO-OR ₂ or R ₄ O-CO-R ₃ -CO-OR ₂ (wherein R ₁ is H or C ₁ ~ C ₂₀ hydrocarbon group, R ₂ and R ₄ are C ₁ ~ C ₂₀ Hydrocarbon group, R ₃ is -(CH ₂ )n-group, n=1~10) ester mixed with H ₂ and CO, hydrogenation reaction is carried out at 30~150°C, 5~100 bar pressure to form alcohol, which The catalyst is composed of the following components: (a) a group VIII metal ion compound in the periodic table; (b) an alkoxide of an alkali metal or an alkaline earth metal; (c) an alcohol.

US4113662 and USP4149021 disclose a kind of ester hydrogenation catalyst, and this catalyst is made up of the element of cobalt, zinc, copper, oxide, hydroxide or carbonate, and the most suitable carboxylate of this catalyst is polyglycolide, The preparation of cycloalkanols is not mentioned in the literature.

US4611085 discloses a method for gas-phase hydrogenation of C ₁ -C ₂₀ carboxylic acid esters to produce alcohols. The catalyst is composed of a VIII group element, an auxiliary agent and a carbon carrier, wherein the VIII group element includes Ru, Ni, Rh, additives include IA (except Li), IIA group (except Be and Mg), lanthanide and actinide elements, and the BET specific surface area of the carbon support is greater than 100m ² /g. The hydrogenation reaction is carried out under the conditions of 100-400°C and gas space velocity of 100-120000h ^-1 . The alkali metals in the catalyst are introduced in the form of alkali metal salts, such as alkali metal nitrates, carbonates or acetates. The method is applicable to carboxylic acid esters that can be vaporized under reaction conditions. The carbon number of the alcohol-derived moiety in the carboxylic acid esters is preferably less than 5 and the carbon connected to oxygen is preferably a primary carbon.

GB2250287A discloses a method for hydrogenating fatty acid esters to produce alcohols. The method is characterized in that a copper-containing catalyst is used for hydrogenation and a certain amount of water is added to the ester raw material to maintain the activity of the catalyst. The applicable carboxylate is C12-C18 fatty acid methyl ester.

It can be seen from the published literature that there is no information disclosure about the co-production of cyclohexanol and ethanol through the hydrogenation of cyclohexyl acetate in the prior art, and there is no selective hydrogenation of benzene, cyclohexene addition esterification, cyclohexyl acetate Information on the hydrogenation of esters to produce cyclohexanol and ethanol.

Contents of the invention

The invention provides a method for the co-production of cyclohexanol and ethanol. The method takes benzene as the starting material and co-produces cyclohexyl through selective hydrogenation of benzene, addition and esterification of cyclohexene, and hydrogenation of cyclohexyl acetate. alcohol and ethanol. The invention also provides a device capable of realizing the above method.

In the present invention, for stream, "A/B" represents the mixture of A and B; for catalyst, "A/B" represents "active component/carrier".

A method for the coproduction of cyclohexanol and ethanol, comprising:

(1) Under the conditions of hydrogenation of benzene to cyclohexene, benzene and hydrogen undergo a hydrogenation reaction to obtain a stream containing cyclohexene;

(2) contacting the cyclohexene-containing stream obtained in step (1) with acetic acid, and an addition esterification reaction occurs under the action of the first catalyst; separating the reaction product to obtain a cyclohexyl acetate stream;

(3) The cyclohexyl acetate stream obtained in step (2) is contacted with hydrogen, and ester hydrogenation reaction occurs under the action of the second catalyst; the reaction product is separated to obtain cyclohexanol and ethanol.

The above three steps will be described respectively below.

1. Benzene hydrogenation to cyclohexene

The method and catalyst for hydrogenation of benzene to cyclohexene are not particularly limited in the present invention, and the existing method for preparing cyclohexene by hydrogenation of benzene and the catalyst for hydrogenation of benzene can be used in the present invention. The present invention preferably adopts liquid phase process. The benzene hydrogenation catalyst is preferably a ruthenium-based catalyst, more preferably a ruthenium-based catalyst containing cobalt and/or zinc. Ruthenium-based catalysts containing cobalt and/or zinc can be prepared by co-precipitation or impregnation of the same carrier.

In step (1), the cyclohexene-containing stream obtained by benzene hydrogenation is mainly a mixture of cyclohexane, cyclohexene and unreacted benzene, which is directly used as the raw material for the next step without separation .

2. Addition and esterification of cyclohexene

In the present invention, "addition esterification reaction" refers to a reaction in which a carboxylic acid is added to an olefin double bond to form an ester.

In step (2), the first catalyst is an acid catalyst, either a liquid acid catalyst or a solid acid catalyst. The liquid acid catalyst can be an inorganic acid, such as sulfuric acid, phosphoric acid, etc.; it can also be an organic acid, such as toluenesulfonic acid, sulfamic acid, etc. The present invention preferably uses a solid acid catalyst. The solid acid catalyst can be selected from one or more of strong acid ion exchange resin catalysts, heteropolyacid catalysts and molecular sieve catalysts.

The strong acid ion exchange resin catalyst includes not only ordinary macroporous sulfonic acid polystyrene-divinylbenzene resin, but also sulfonic acid resin modified by halogen atoms. This kind of resin is easy to buy from the market, and can also be prepared according to the methods recorded in classic literature. The preparation method of macroporous sulfonic acid polystyrene-divinylbenzene resin is usually to drop the mixture of styrene and divinylbenzene into the water phase containing dispersant, initiator and porogen under the condition of high-speed stirring. Carry out suspension copolymerization in the system, separate the obtained polymer balls (white balls) from the system, use a solvent to remove the porogen, and then use dichloroethane as a solvent and concentrated sulfuric acid as a sulfonating agent to carry out Sulfonation reaction, and finally through filtration, washing and other processes, the final product is obtained. Introducing halogen atoms, such as fluorine, chlorine, bromine, etc., into the skeleton of ordinary strong acid ion exchange resin can further improve the temperature resistance and acid strength of the resin. This halogen-containing strongly acidic high-temperature resistant resin can be obtained at least through the following two ways. One way is to introduce a halogen atom, such as a chlorine atom, into the benzene ring of the sulfonated styrene resin skeleton. Due to the strong electron-withdrawing effect of the halogen element It can not only stabilize the benzene ring, but also increase the acidity of the sulfonic acid group on the benzene ring, so that the acid strength function (Hammett function) of the resin catalyst can be H0≤-8, and it can be used for a long time above 150°C. Resins can be easily purchased from the market, such as Amberlyst 45 resin produced by foreign ROHM & HASS companies, D008 resin produced by Hebei Jizhong Chemical Factory in China, etc.; another way is to replace all the hydrogen on the resin skeleton with fluorine. Electron-absorbing properties make it super acidic and super thermally stable. The acid strength function (Hammett function) H0 can be less than -12, and the heat-resistant temperature can reach above 250 ° C. This type of high-temperature-resistant strong acid resin is typical An example is Nafion resin produced by DuPont Corporation.

The heteropolyacid catalyst can be heteropolyacid and/or heteropolyacid acid salt, or a catalyst supporting heteropolyacid and/or heteropolyacid acid salt. The acid strength function H0 of the heteropoly acid and its acid salt can be less than -13.15, and can be used for a long time up to 300°C. The heteropolyacid and its acid salt include Keggin structure, Dawson, Anderson structure, Silverton structure heteropoly acid and its acid salt. Heteropolyacids with keggin structure and their acid salts are preferred, such as dodecaphosphotungstic acid (H ₃ PW ₁₂ O ₄₀ ·xH ₂ O), dodecasilicatetungstic acid (H ₄ SiW ₁₂ O ₄₀ ·xH ₂ O), Dodecaphosphomolybdic acid (H ₃ PMo ₁₂ O ₄₀ ·xH ₂ O), dodecaphosphomolybdovanadate (H ₃ PMo _12- yV _y O ₄₀ ·xH ₂ O), etc. The heteropolyacid acid salt is preferably cesium phosphotungstate (Cs _2.5 H _0.5 P ₁₂ WO ₄₀ ), whose acid strength function H0 is less than -13.15, and the specific surface area can reach more than 100 m ² /g. The heteropolyacid catalyst can be selected from one or more of the above-mentioned preferred heteropolyacids and heteropolyacid acid salts. In the catalyst supporting heteropolyacid and/or heteropolyacid acid salt, the carrier is generally _SiO2 and/or activated carbon.

The solid acid catalyst can also be a molecular sieve catalyst. The molecular sieve can be one or more of Hβ, HY and HZSM-5, preferably one or more of Hβ, HY and HZSM-5 modified with fluorine or phosphorus. After these molecular sieves are modified by fluorine and phosphorus, the acidity and catalytic performance of the molecular sieve can be further improved.

Two implementations of step (2) are described in detail below.

The first implementation mode:

The catalyst adopts solid acid catalyst. In step (2), one or more parallel reactors can be used, and the reactor type can be selected from one or more of tank reactors, fixed bed reactors, ebullating bed reactors and fluidized bed reactors . Preference is given to using one or more parallel tubular fixed-bed reactors. More preferably one or more shell-and-tube reactors connected in parallel are used. The operation mode of the reactor can be a batch mode or a continuous mode, preferably a continuous operation mode. Fixed bed reactors can be operated adiabatically or isothermally. The adiabatic reactor can be a cylinder reactor, the catalyst is fixed in the reactor, and the outer wall of the reactor is kept 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 use part of the reaction product to cool and then circulate to the reactor inlet to dilute the reactant concentration. The isothermal reactor can be a shell-and-tube reactor. The catalyst is fixed in the tube, and cooling water is passed through the shell to remove the heat released by the reaction.

The reaction temperature is generally 50-200°C, and the optimal reaction temperature is 60-120°C.

The pressure of the addition esterification reaction is related to the reaction temperature. Since the addition esterification reaction is carried out in the liquid phase, the reaction pressure should ensure that the reaction is in the liquid phase state. Generally speaking, the reaction pressure is normal pressure ~ 10MPa, and the optimal pressure is normal pressure ~ 1MPa.

The acid-to-ene molar ratio of the addition esterification reaction is generally 0.2-20:1, and the optimal condition is 1.2-3:1.

In the above-mentioned addition esterification reaction, the liquid feed space velocity is generally 0.5-20h ^-1 , and the optimum condition is 0.5-5h ^-1 .

Under the above conditions, the conversion of cyclohexene in the addition esterification reaction can generally reach more than 80%, while the selectivity of the esterification reaction can reach more than 99%.

The reaction product of step (2) is mainly composed of cyclohexane, benzene, acetic acid and cyclohexyl acetate, and there is a certain amount of unreacted cyclohexene. It is carried out in the addition esterification product separation unit of the distillation separation part. An optional separation method is to carry out rectification separation on the esterification product stream to obtain cyclohexane/cyclohexene/benzene stream, acetic acid stream and cyclohexyl acetate stream. The cyclohexane/cyclohexene/benzene stream can be hydrogenated to prepare cyclohexane in an additional hydrogenation unit, the acetic acid stream can be used as a part of the reaction feed in step (2), and the cyclohexyl acetate stream can be used as a step (3) ) raw materials for the reaction. The present invention can also carry out extractive distillation separation to cyclohexane/cyclohexene/benzene stream further, a kind of optional separation scheme is, adopt extractive distillation to separate it into cyclohexane/cyclohexene stream and benzene stream , the benzene stream can be used as a part of the reaction feed in step (1), and the cyclohexane/cyclohexene stream can be hydrogenated to produce cyclohexane; another optional separation scheme is to separate it into Cyclohexane stream, cyclohexene stream and benzene stream, benzene stream can be used as a part of step (1) reaction feed, cyclohexene stream can be used as a part of step (2) reaction feed, and cyclohexane stream is exported as by-product device. The present invention can also remove the heavy component in the above-mentioned cyclohexyl acetate stream through rectification separation, and use the cyclohexyl acetate stream after removing the heavy component as the raw material for the step (3) reaction, and the separated heavy component stream as a by-product discharge device.

Specifically, the addition esterification product separation unit can be provided with a decarburization six-component tower and a deacetic acid tower. The addition esterification product first enters the decarburization six-component tower for separation. This tower can be operated under normal pressure. By controlling the heating capacity of the tower bottom, the reflux ratio, the output of the tower top and the tower bottom, cyclohexene is extracted from the top of the tower. The alkane/cyclohexene/benzene stream, the stream extracted from the decarburization six-component tower tank enters the deacetic acid tower for separation. This tower can also be operated under normal pressure. And tower still extraction, extract acetic acid stream from tower top, extract cyclohexyl acetate stream from tower still. The esterification product separation unit can be provided with an extractive rectification tower, and the cyclohexane/cyclohexene/benzene stream is further separated into a cyclohexane/cyclohexene stream and a benzene stream; or two extractive distillation towers can be set again The tower further separates the cyclohexane/cyclohexene/benzene stream into cyclohexane stream, cyclohexene stream and benzene stream. The addition esterification product separation unit can also be provided with a deheavy component tower, and the cyclohexyl acetate stream extracted from the deacetate tower tank enters the deheavy component tower to further remove the heavy component in the stream, thereby obtaining the removal Heavy component cyclohexyl acetate stream.

The second implementation mode:

The catalyst adopts solid acid catalyst. In step (2), adopt one or more parallel reactive distillation towers, while carrying out addition esterification reaction, carry out the separation of reaction product, obtain cyclohexyl acetate stream from the bottom of reactive distillation tower, from reaction The cyclohexane/benzene/acetic acid stream is obtained from the top of the rectifying tower.

The theoretical plate number of the reactive distillation column is 10-150, and the solid acid catalyst is arranged between 1/3 to 2/3 of the theoretical plate number; relative to the total loading volume of the catalyst, the liquid feed space velocity 0.2-20h ^-1 ; the operating pressure of the reactive distillation tower is -0.0099MPa to 5MPa; the temperature of the catalyst loading area is between 50-200°C; the reflux ratio is 0.1-100:1.

The reactive distillation column is the same in form as an ordinary distillation column, and generally consists of a column body, a top condenser, a reflux tank, a reflux pump, a column kettle, and a reboiler. The type of column can be a tray column, a packed column, or a combination of both. The types of tray columns that can be used include valve columns, sieve tray columns, bubble cap columns, etc. The packing used in the packed tower can be random packing, such as Pall ring, θ ring, saddle packing, stepped ring packing, etc.; it can also be structured packing, such as corrugated plate packing, corrugated wire mesh packing, etc.

According to the method provided by the present invention, a solid acid catalyst is arranged in the reactive distillation column. Those skilled in the art know clearly that the catalyst arrangement in the reactive distillation column should meet the following two requirements: (1) It should be able to provide enough channels for the passage of vapor-liquid two-phase, or have a relatively large bed gap rate (generally at least 50%) to ensure that the gas-liquid two-phase can flow through without causing flooding; (2) to have good mass transfer performance, the reactant should be transferred from the fluid phase to the catalyst for reaction, At the same time, the reaction product is transferred from the catalyst. Various arrangements of catalysts in the reactive distillation column have been disclosed in the existing literature, and all of these arrangements can be adopted in the present invention. The arrangement of existing catalysts in the reaction tower can be divided into the following three types: (1) The catalyst is directly arranged in the tower in the form of rectification packing. The main method is to mechanically mix catalyst particles of a certain size and shape with rectification packing , or the catalyst is sandwiched between the structured packing and the structured packing to form an integral packing, or the catalyst is directly made into the shape of the rectification packing; (2) The catalyst is placed in a small gas-liquid permeable container and placed in the reaction The catalyst is placed on the tray of the tower, or the catalyst is arranged in the downcomer of the reaction tower; (3) The catalyst is directly loaded into the reaction tower in the form of a fixed bed, and the liquid phase flows directly through the catalyst bed layer, and a dedicated gas phase is set up. In this way, the catalyst bed and the rectification tray are alternately arranged at the position where the catalyst is installed. The liquid on the tray enters the next catalyst bed through the downcomer and the redistributor, and is carried out in the bed. Addition reaction, the liquid in the lower part of the catalyst bed enters the next tray through the liquid collector.

The reactive distillation column must have enough theoretical plate numbers and reaction plate numbers to meet the reaction and separation requirements. The theoretical plate number of the reactive distillation column is preferably 30-100, and the solid acid catalyst is arranged between 1/3 to 2/3 of the theoretical plate number.

In the present invention, it is necessary to ensure that the reactants have sufficient residence time to realize the complete conversion of cyclohexene. The space velocity of the liquid feed is preferably 0.5 to 5 h ^-1 relative to the total loading volume of the catalyst.

In the present invention, the operating pressure of the reactive distillation column can be operated under negative pressure, normal pressure and pressurized conditions. The operating pressure of the reactive distillation column is preferably from normal pressure to 1 MPa.

The operating temperature of the reactive distillation tower is related to the pressure of the reactive distillation tower. The temperature distribution of the reaction tower can be adjusted by adjusting the operating pressure of the reaction tower so that the temperature of the catalyst loading area is within the active temperature range of the catalyst. The temperature of the catalyst loading zone is preferably between 60°C and 120°C.

The reflux ratio of the reactive distillation column should meet the requirements of separation and reaction at the same time. In general, increasing the reflux ratio is conducive to improving the separation ability and reaction conversion rate, but at the same time it will increase the energy consumption of the process. The reflux ratio is preferably 0.5˜10:1.

According to the method provided by the present invention, the cyclohexyl acetate stream obtained from the bottom of the reactive distillation column is used as the raw material for the reaction in step (3). For the hexanaphthene/benzene/acetic acid flow that the reactive distillation tower top obtains, can adopt rectification to separate it into hexanaphthene/benzene flow and acetic acid flow, the acetic acid flow can be used as a part of step (2) reaction feed; The hexane/benzene stream can be hydrogenated in an additional hydrogenation unit to prepare cyclohexane, or it can be separated into cyclohexane stream and benzene stream by extractive distillation, and the benzene stream can be used as the reaction feed of step (1) Part of the cyclohexane stream exits the unit as a by-product. Specifically, the cyclohexane/benzene/acetic acid stream can be rectified and separated in a decarburization six-component tower. This tower can be operated at normal pressure. output, the cyclohexane/benzene stream is extracted from the top of the tower, and the acetic acid stream is extracted from the tower bottom; the cyclohexane/benzene stream can enter an additional hydrogenation unit, or enter an extractive distillation tower, and be further separated into cyclohexane alkanes and benzene streams.

3. Hydrogenation of cyclohexyl acetate

According to the method provided by the present invention, the cyclohexyl acetate stream obtained by separating the addition esterification product is sent to an ester hydrogenation reactor for hydrogenation reaction. In step (3), one or more parallel reactors can be set, and the reactor type is selected from one or more of tank reactors, fixed bed reactors, ebullating bed reactors and fluidized bed reactors. In step (3), one or more parallel tubular fixed-bed reactors are preferably set up. In step (3), it is more preferable to set up one or more parallel shell-and-tube reactors, the ester hydrogenation catalyst is fixed in the tubes, and the heat released by the reaction is removed through the cooling medium at the shell side.

The second catalyst is an ester hydrogenation catalyst. Although the existing published literature is mainly about the hydrogenation of methyl carboxylate or ethyl carboxylate, as usual, hydrogenation of fatty acid methyl esters is used to produce higher alcohols, and methyl maleate is hydrogenated to produce 1,4 -butanediol, 1,6-hexanedioic acid dimethyl ester hydrogenation produces 1,6-hexanediol etc., do not see any report about the carboxylic acid ester hydrogenation reaction of naphthenic alcohol derivation, but the present invention It has been found that existing ester hydrogenation catalysts can be used for the hydrogenation of cyclohexyl acetate. The hydrogenation of esters generally uses copper-based catalysts, ruthenium-based catalysts and noble metal-based catalysts, with copper-based catalysts being the most commonly used. Copper-based ester hydrogenation catalysts use copper as the main catalyst, and add chromium, aluminum, zinc, calcium, magnesium, nickel, titanium, zirconium, tungsten, molybdenum, ruthenium, platinum, palladium, rhenium, lanthanum, thorium, gold, etc. One or more components are used as co-catalysts or additive components. Copper-based ester hydrogenation catalysts can be easily purchased from the market, and can also be prepared by co-precipitation. The usual preparation method is to put the soluble salt solution of each metal into the neutralization kettle, and add alkali solution (sodium hydroxide, sodium carbonate, ammonia water, urea, etc.) at a certain temperature and stirring speed to neutralize to pH8~ 12 Growth and precipitation, the precipitation is formed through aging, filtration, washing, drying, roasting, molding and other processes, and finally reduced in a hydrogen atmosphere to make the final ester hydrogenation catalyst. The general composition of ruthenium catalysts: Ru/Al ₂ O ₃ or Ru-Sn/Al ₂ O ₃ . The general composition of noble metal catalysts: Pt/Al ₂ O ₃ , Pd-Pt/Al ₂ O ₃ or Pd/C.

In the present invention, the ester hydrogenation catalyst can be selected from one or more of copper-based catalysts, ruthenium-based catalysts and noble metal-based catalysts, preferably copper-based catalysts, more preferably copper-based catalysts containing zinc and/or chromium.

The ester hydrogenation reaction unit can be operated either batchwise or continuously. The batch reaction generally uses a reactor as a reactor, puts cyclohexyl acetate and a hydrogenation catalyst into the reactor, feeds hydrogen to react at a certain temperature and pressure, and discharges the reaction product from the reactor after the reaction is completed. , The product is separated, and then put into the next batch of materials for reaction. The shell-and-tube reactor can be used for the continuous hydrogenation reaction. The hydrogenation catalyst is fixed in the tube, and cooling water is passed through the shell to remove the heat released by the reaction.

The hydrogenation reaction temperature of cyclohexyl acetate is related to the selected hydrogenation catalyst. For copper-based hydrogenation catalysts, the general hydrogenation reaction temperature is 150-400°C, and the optimal reaction temperature is 200-300°C. The reaction pressure is from normal pressure to 20MPa, and the optimal pressure is from 4 to 10MPa.

The control of the hydrogen ester molar ratio in the cyclohexyl acetate hydrogenation reaction is very important. A high hydrogen-to-ester ratio is beneficial to the hydrogenation of esters, but too high a hydrogen-to-ester ratio will increase the energy consumption of the hydrogen compression cycle. Generally, the ratio of hydrogen to ester is 1-1000:1, and the optimal condition is 5-100:1.

In the hydrogenation reaction, the feed space velocity of the ester is related to the activity of the selected catalyst. Higher activity catalysts can use higher space velocities. For the selected catalyst, the reaction conversion decreases with the increase of the reaction space velocity. In order to meet a certain conversion rate, the space velocity must be limited within a certain range. Generally, the liquid feeding space velocity of the ester is 0.1-20h ^-1 , and the optimum condition is 0.2-2h ^-1 . If a batch reaction is adopted, the reaction time is 0.5-20 h, preferably 1-5 h.

The ester hydrogenation reaction product mainly contains ethanol and cyclohexanol, may also contain a certain amount of ethyl acetate and cyclohexanone, and may also contain a certain amount of unreacted cyclohexyl acetate, and a small amount of high boilers (dimerization Ketones), these mixtures can be separated by rectification and/or extractive distillation, preferably by rectification. Distillation and separation can adopt a batch scheme or a continuous process scheme. Batch rectification, that is, put the ester hydrogenation product into the rectification column, and distill ethanol, ethyl acetate, cyclohexanol, cyclohexanone, and cyclohexyl acetate from the top of the column in turn, leaving a small amount of high boilers remaining in the column. The present invention preferably adopts continuous rectification to separate ester hydrogenation products. Continuous distillation requires the separation of the various components using a series of distillation columns. Various separation processes can be designed according to the sequence of separation of each component. The process scheme of the preferred sequence separation of the present invention, that is, the ester hydrogenation product separation unit is sequentially arranged for the de-ethanol tower, the de-cyclohexanol tower, the cyclohexyl acetate recovery tower. The ester hydrogenation product first enters the ethanol removal tower to separate to obtain ethanol, then enters the decyclohexanol tower to obtain cyclohexanol, and finally enters the cyclohexyl acetate recovery tower to recover unreacted cyclohexyl acetate, and a small amount of high boiling delivery device. In the present invention, only one rectifying tower can be arranged, and only ethanol is separated to obtain cyclohexanol containing a small amount of impurities.

In step (3), after the reaction product is separated, the obtained cyclohexyl acetate stream can be used as a part of the reaction feed of step (3).

The invention provides a device for the co-production of cyclohexanol and ethanol, comprising a sequentially connected benzene hydrogenation reaction unit, an addition esterification reaction unit, an addition esterification product separation unit, an ester hydrogenation reaction unit and an ester hydrogenation reaction unit Product separation unit.

According to the method provided by the present invention, those skilled in the art can easily determine the pipeline connection mode between each unit and equipment.

The benzene hydrogenation reaction unit is provided with one or more parallel reactors, and the reactor type is selected from fixed bed reactors and/or tank reactors.

The addition esterification reaction unit is provided with one or more parallel reactors, and the reactor type is selected from one or more of tank reactors, fixed bed reactors, ebullating bed reactors and fluidized bed reactors Several kinds.

The addition esterification reaction unit is provided with at least one reactive distillation column.

A pre-esterification reactor is arranged in series before the reactive distillation tower, and the pre-esterification reactor is a tank reactor, a fixed-bed reactor, a fluidized-bed reactor or an ebullated-bed reactor.

The addition esterification product separation unit is provided with at least one rectification tower.

The ester hydrogenation reaction unit is provided with one or more parallel reactors, and the reactor type is selected from one or more of tank reactors, fixed bed reactors, ebullating bed reactors and fluidized bed reactors . The ester hydrogenation reaction unit is preferably provided with one or more parallel shell-and-tube reactors.

The ester hydrogenation product separation unit is provided with at least one rectification tower.

The present invention also provides another device for the co-production of cyclohexanol and ethanol, comprising a sequentially connected benzene hydrogenation reaction unit, an addition esterification reaction unit, an ester hydrogenation reaction unit and an ester hydrogenation product separation unit; The addition esterification reaction unit is provided with one or more parallel reactive distillation columns.

According to the method provided by the present invention, those skilled in the art can easily determine the pipeline connection mode between each unit and equipment.

The benzene hydrogenation reaction unit is provided with one or more parallel reactors, and the reactor type is selected from fixed bed reactors and/or tank reactors.

A pre-esterification reactor is arranged in series before the reactive distillation tower; the pre-esterification reactor is a tank reactor, a fixed-bed reactor, a fluidized-bed reactor or an ebullated-bed reactor.

The ester hydrogenation reaction unit is provided with one or more parallel reactors, and the reactor type is selected from one or more of tank reactors, fixed bed reactors, ebullating bed reactors and fluidized bed reactors .

The ester hydrogenation reaction unit is provided with one or more parallel shell-and-tube reactors.

The ester hydrogenation product separation unit is provided with at least one rectification tower.

The invention provides a high-efficiency, low-cost new technology route for producing cyclohexanol. The characteristics of the present invention are: (1) both esterification and ester hydrogenation reactions have high selectivity, so the utilization rate of atoms is high; (2) the process is environmentally friendly; (3) co-production while producing cyclohexanol Ethanol, which converts cheap acetic acid into ethanol with high price and huge market capacity through an indirect method, which greatly increases the economy of the process; (4) Addition esterification by reactive distillation can not only significantly improve the reaction efficiency, but also The extractive distillation separation process can be simplified, and the investment and operation costs can be greatly reduced.

Description of drawings

Fig. 1 is a schematic flow chart of the first embodiment of the present invention.

Fig. 2 is a schematic flowchart of a second embodiment of the present invention.

Detailed ways

Two preferred process flows of the present invention are briefly described below in conjunction with the accompanying drawings.

In the first process flow, the addition esterification reactor adopts a tank reactor, a tubular fixed bed reactor, an ebullated bed reactor or a fluidized bed reactor. As shown in Figure 1: benzene and hydrogen enter benzene hydrogenation reactor 1, under the effect of benzene hydrogenation catalyst, carry out hydrogenation reaction, benzene hydrogenation product stream enters addition esterification reactor 2 through pipeline 11, and passes through The acetic acid entering in the pipeline 21 is mixed, and under the action of a solid acid catalyst, the addition esterification reaction is carried out, and the addition esterification product flow enters the addition esterification product separation unit 3 through the pipeline 22, and the cyclohexane flow 31, Cyclohexene stream 32, benzene stream 33, acetic acid stream 34 and cyclohexyl acetate stream 35, benzene stream circulates back to benzene hydrogenation reactor 1, cyclohexene stream and acetic acid stream circulate back to addition esterification reactor 2, acetic acid The cyclohexyl ester stream enters the ester hydrogenation reactor 4, and under the action of the ester hydrogenation catalyst, contacts with hydrogen to carry out the ester hydrogenation reaction, and the ester hydrogenation product stream enters the ester hydrogenation product separation unit 5, and cyclohexanol is obtained after separation Stream 52, ethanol stream 53, cyclohexyl acetate stream 51 and high-boiler stream 54, cyclohexanol stream and ethanol stream go out device as product, high-boiler stream goes out device as by-product, and cyclohexyl acetate stream loops back to ester processing Hydrogen Reactor 4.

In the second process flow, the addition esterification reactor adopts a reactive distillation tower. As shown in Figure 2: benzene and hydrogen enter benzene hydrogenation reactor 1, under the effect of benzene hydrogenation catalyst, carry out hydrogenation reaction, benzene hydrogenation product flow enters reactive distillation tower 2 through pipeline 11, and through pipeline 21 The acetic acid that enters mixes, and under the effect of solid acid catalyst, carries out addition esterification reaction, carries out the separation of addition esterification product simultaneously, obtains hexanaphthene/benzene/acetic acid flow from reaction rectification tower 2 tower top, from reaction The cyclohexyl acetate stream is obtained at the bottom of the rectifying tower 2; the cyclohexane/benzene/acetic acid stream enters the addition esterification product separation unit 3 through the pipeline 22, and the cyclohexane stream, the benzene stream and the acetic acid stream are obtained through separation. The alkane stream goes out of the device as a by-product, the benzene stream circulates back to the benzene hydrogenation reactor 1, the acetic acid stream circulates back to the addition esterification reactor 2, and the cyclohexyl acetate stream enters the ester hydrogenation reactor 4. Next, the 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, a cyclohexyl acetate stream 51, and a high boiler stream 54 , the cyclohexanol stream and the ethanol stream go out of the device as products, the high boiler stream goes out of the device as a by-product, and the cyclohexyl acetate stream is recycled back to the ester hydrogenation reactor 4.

The present invention is further illustrated by the following examples, but the present invention is not limited thereto.

Example 1

This example illustrates the method for the selective hydrogenation of benzene to cyclohexene.

Benzene and hydrogen are injected into the hydrogenation reactor filled with ruthenium particle catalyst at a molar ratio of 1:3, and the benzene hydrogenation reaction is carried out under the conditions of a reaction temperature of 135°C, a pressure of 4.5MPa, and a residence time of 15min. After the reaction product is separated from the hydrogen , Collect the liquid product and run continuously for 1000h. After the test, the collected liquid product was analyzed by gas chromatography, and its composition was: 53.3% benzene, 35.4% cyclohexene, and 11.3% cyclohexane.

Example 2

100mL of macroporous strongly acidic hydrogen-type ion exchange resin (synthesized in the laboratory according to the classic literature method, suspension copolymerization of styrene solution containing 15% divinylbenzene to make white balls, and then sulfonated with concentrated sulfuric acid, measured Its exchange capacity is 5.2mmolH ⁺ /g dry basis) into the middle part of a stainless steel tube reactor with a jacket of φ32×4×1000mm, and a certain amount of quartz sand is filled at both ends. The raw materials of acetic acid and cyclohexene (53.3% benzene, 35.4% cyclohexene, and 11.3% cyclohexane) are pumped into the reactor by a metering pump at a certain flow rate for reaction, and heat is introduced into the external jacket of the reaction tube. Water is used to control the reaction temperature, and the reactor pressure is controlled through the reactor outlet back pressure valve. The outlet product of the reactor was sampled by an online sampling valve for online chromatographic analysis, and the conversion of cyclohexene and the selectivity of cyclohexyl acetate were calculated from the composition of the product. The reaction conditions and results are shown in Table 1.

It can be seen from Table 1 that the strong acid ion exchange resin catalyst is used to catalyze the reaction of cyclohexene raw material and acetic acid, the conversion rate of cyclohexene is greater than 80%, the selectivity of ester products is greater than 99%, and the catalyst activity and selectivity are stable after 600 hours of operation. .

Table 1 Strong acidic ion exchange resin catalyzed acetic acid and cyclohexane/cyclohexene/phenyl esterification test data

Example 3

The test device, method and raw materials are the same as in Example 2, except that the catalyst is a phosphorus-modified Hβ molecular sieve catalyst (the Hβ molecular sieve with a silicon-aluminum ratio of 50 is modified with 85% phosphoric acid, and then kneaded with alumina to form , obtained by drying at 120°C and roasting 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 reaction conversion rate of cyclohexene and acetic acid is greater than 80%, the selectivity of ester product is greater than 99%, and the catalyst activity and selectivity are stable after running for 480 hours.

Table 2 Hβ molecular sieve catalyst catalyzed esterification test data of acetic acid and cyclohexane/cyclohexene/phenyl

Example 4

Collect the addition esterification product of embodiment 2 and 3, carry out rectification separation test. The rectification adopts a glass tower rectification device with a height of 2m and a diameter of 40mm. The tower column is equipped with a Ф3mm stainless steel θ mesh ring high-efficiency rectification filler. The tower kettle is heated, and the heating amount of the tower kettle is adjusted through a pressure regulator. The reflux of the tower is controlled by a reflux ratio regulator. The distillation separation results are shown in Table 3.

Table 3 Addition esterification product distillation separation test result

Embodiment 5～6 are used for illustrating the method adopting reactive distillation to prepare cyclohexyl acetate.

The tests carried out in Examples 5 to 6 are all carried out in the reactive distillation model test device of the following specifications: the main body of the model device is a stainless steel tower with a diameter (inner diameter) of 50mm and a height of 3m, and the lower connecting volume of the tower is The 5L tower kettle is equipped with a 10KW electric heating rod in the kettle. The heating rod is controlled by an intelligent controller through a silicon controlled rectifier (SCR). A condenser with a heat exchange area of ^0.5m2 is connected to the top of the tower, and the steam at the top of the tower is condensed into a liquid by the condenser and then enters a reflux tank with a volume of 2L. The liquid in the reflux tank is partially refluxed to the reaction tower through the reflux pump, and part of it is extracted as light components. The operating parameters of the tower are displayed and controlled by intelligent automatic control instruments. The reflux flow of the tower is controlled by the reflux regulating valve, and the output at the top of the tower is controlled by the liquid level controller of the reflux tank. The production volume of the tower kettle is controlled by adjusting the tower kettle discharge valve by the tower kettle liquid level controller. The raw materials of acetic acid and cyclohexene are put into 30L storage tanks respectively, and are pumped into corresponding preheaters through metering pumps to preheat to a certain temperature before entering the reaction tower. The feeding speed is controlled by metering pumps and accurately measured by electronic scales.

Example 5

The high-temperature-resistant sulfonic acid ion exchange resin (the brand name is Amberlyst 45, produced by Rhom&Hass Company) was crushed into a powder with a particle size of less than 200 mesh (0.074mm) with a multi-stage high-speed pulverizer, and pore-forming agents, lubricants, and antioxidants were added. The agent and adhesive are mixed evenly on a high-speed mixer, and then internally kneaded on an internal mixer at 180°C for 10 minutes to make the material completely plasticized, and then injected into a mold to make a Raschig ring with a diameter of 5mm, a height of 5mm, and a wall thickness of 1mm. type resin catalyst filler. Put 1950mL of this packing into the middle of the model reaction tower (1m in height, equivalent to 8 theoretical trays) and 1950mL of glass spring packing with a diameter of 3mm and a length of 6mm (filling height is 1m, equivalent to 10 theoretical trays). plate). The cyclohexene raw material and acetic acid were pumped into the preheater by the metering pump to preheat and enter the reaction tower, and the reaction was carried out continuously by adjusting the heating capacity of the tower kettle and the reflux flow at the top of the tower. The reaction conditions and reaction results under stable operation are shown in Table 4.

The reactive distillation test data of table 4 high temperature resistant sulfonic acid type ion exchange resin catalyst

According to the experimental data, the conversion rate of cyclohexene is calculated as 98.8%, and the selectivity of cyclohexyl acetate is 98.0%.

Example 6

Spherical H _0.5 Cs _2.5 PW ₁₂ O ₄₀ /SiO ₂ catalyst with a diameter of 3-4 (composed of 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 meshes is fully mixed in a mixer, then In the sugar coating machine, silica sol is used as a bonding machine to roll the ball into shape, and then dried and roasted) sandwiched into a titanium wire mesh wave plate to make a cylindrical structured packing with a diameter of 50mm and a height of 50mm. Put this packed catalyst L into the middle of the model reaction tower (1m in height, equivalent to 12 theoretical trays), and put 1950mL glass spring packing with a diameter of 4mm and a height of 4mm in the upper and lower parts (the packing height is 1m, equivalent to 15 block theoretical plate). The cyclohexene raw material and acetic acid were pumped into the preheater by the metering pump to preheat and enter the reaction tower, and the heating capacity of the tower reactor and the reflux flow at the top of the tower were adjusted for continuous reaction. The reaction conditions and reaction results under stable operation are shown in Table 5.

Table 5 Reactive distillation test data of H _0.5 Cs _2.5 PW ₁₂ O ₄₀ /SiO ₂ catalyst

According to the experimental data, the conversion rate of cyclohexene is calculated to be 99.35%, and the selectivity of cyclohexyl acetate is 99.6%.

Embodiments 7-8 are used to illustrate the hydrogenation method of cyclohexyl acetate.

Example 7

The raw material for hydrogenation is cyclohexyl acetate with a purity of 99.6%.

40g copper zinc aluminum ester hydrogenation catalyst (laboratory synthesis, composition is CuO 40.5%, ZnO29.6%, Al ₂ O ₃ 30.4%. By the nitrate solution of copper, zinc, chromium, add sodium hydroxide solution to neutralize to PH=9.0, centrifuged, washed, dried, pressed into tablets, and roasted) into the middle of a φ20×2.5×800mm stainless steel tube reactor with a jacket, and a certain amount of quartz sand is filled at both ends. Pass in hydrogen (500mL/min) and reduce for 24h under the conditions of 280°C and 6MPa, then drop to the temperature and pressure of hydrogenation reaction. The cyclohexyl acetate is pumped into the reactor by the metering pump, the hydrogen gas enters the reaction system through the mass flow controller for hydrogenation reaction, the reaction temperature is controlled by introducing heat transfer oil into the outer jacket of the reaction tube, and the reaction temperature is controlled by the back pressure valve at the outlet of the reactor. Control the reactor pressure. The reaction product is sampled through the linear sampling valve at the rear of the reactor for online chromatographic analysis. The reaction conditions and results are shown in Table 6. The results in Table 6 show that, using the copper-zinc-aluminum catalyst, the conversion rate of cyclohexyl acetate hydrogenation reaction can reach more than 99.0%, and the selectivity of cyclohexanol is greater than 99.9%. After 1000 hours of operation, the conversion rate and selectivity have not decreased.

The cyclohexyl acetate hydrogenation test data of table 6 copper-zinc-aluminum ester hydrogenation catalyst

Example 8

The raw material for hydrogenation is cyclohexyl acetate with a purity of 99.6%.

Put 40g copper chromium ester hydrogenation catalyst (commercially available, produced by Taiyuan Xinjida Chemical Co., Ltd., brand name is C1-XH-1, CuO content is 55%, diameter 5mm tablet, crushed into 10-20 mesh particles) Enter the middle part of the φ20×2.5×800mm jacketed stainless steel tube reactor, and fill a certain amount of quartz sand at both ends. Pass in hydrogen (500mL/min) and reduce for 24h under the conditions of 280°C and 6MPa, then drop to the reaction temperature and pressure. The cyclohexyl acetate is pumped into the reactor by the metering pump, the hydrogen 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 reaction temperature is controlled by the back pressure valve at the outlet of the reactor. Control the reactor pressure. The reaction product is sampled through the linear sampling valve at the rear of the reactor for online chromatographic analysis. The reaction conditions and results are shown in Table 7. The results in Table 7 show that, using the copper-zinc-aluminum catalyst, the conversion rate of cyclohexyl acetate hydrogenation reaction can reach more than 98.0%, and the selectivity of cyclohexanol is greater than 99.9%. After 500 hours of operation, the conversion rate and selectivity have not decreased.

The cyclohexyl acetate hydrogenation test data of copper chromium ester hydrogenation catalyst of table 7

Example 9

Collect the reaction product 4000g of example 7～8, carry out rectifying separation test. The rectification adopts a glass tower with a height of 2m. The tower column is equipped with Ф3mm stainless steel θ mesh ring high-efficiency rectification packing. The tower kettle is a 5L glass flask, which is heated by an electric heating mantle and the heating capacity of the tower kettle is adjusted by a pressure regulator. The reflux of the tower is controlled by a reflux ratio regulator. The distillation separation results are shown in Table 8.

The rectification separation test data of cyclohexyl acetate hydrogenation product of table 8

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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