WO2022166540A1 - 一种制备马来酸二醇酯的工艺方法 - Google Patents

一种制备马来酸二醇酯的工艺方法 Download PDF

Info

Publication number
WO2022166540A1
WO2022166540A1 PCT/CN2022/071265 CN2022071265W WO2022166540A1 WO 2022166540 A1 WO2022166540 A1 WO 2022166540A1 CN 2022071265 W CN2022071265 W CN 2022071265W WO 2022166540 A1 WO2022166540 A1 WO 2022166540A1
Authority
WO
WIPO (PCT)
Prior art keywords
reaction
esterification
temperature
distillation column
methanol
Prior art date
Application number
PCT/CN2022/071265
Other languages
English (en)
French (fr)
Inventor
窦和瑞
Original Assignee
窦和瑞
Priority date (The priority date is an assumption and is not a legal conclusion. Google has not performed a legal analysis and makes no representation as to the accuracy of the date listed.)
Filing date
Publication date
Application filed by 窦和瑞 filed Critical 窦和瑞
Publication of WO2022166540A1 publication Critical patent/WO2022166540A1/zh

Links

Images

Classifications

    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C67/00—Preparation of carboxylic acid esters
    • C07C67/08—Preparation of carboxylic acid esters by reacting carboxylic acids or symmetrical anhydrides with the hydroxy or O-metal group of organic compounds
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C67/00—Preparation of carboxylic acid esters
    • 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

Definitions

  • the invention relates to the technical field of fine chemicals and polymer monomer synthesis, in particular to a synthetic process method and route for preparing maleic acid glycol ester.
  • Dimethyl maleate (DMM), Diethyl Maleate (DEM) and Dipropyl Maleate (DPM) are the most representative of maleic acid glycol esters.
  • the first two are important fine chemical raw materials.
  • the production process and methods of other maleic acid glycol esters are basically similar to the first two.
  • DMM and DEM are basically used as representatives to introduce their process routes and methods.
  • Dimethyl maleate (DMM) is a colorless viscous liquid with a boiling point (1.0bar) of 200.4°C, a melting point of 19.0°C and a relative density (25°C) of 1.1462g/cm 3 . It is used for the production of paints, coatings and adhesives. , pesticides, water purifiers, optical materials, anti-shrinkage finishing agents and anti-rust additives and other important chemical raw materials, the domestic and foreign demand for dimethyl maleate is increasing.
  • the traditional production method of dimethyl maleate mostly adopts maleic anhydride as raw material, and carries out esterification reaction with methanol under the catalysis of sulfuric acid or p-toluenesulfonic acid.
  • the production process has high catalytic activity and low price of the catalyst , but there are many side reactions, the product is easily isomerized into dimethyl fumarate DMF, the equipment is seriously corroded, and the follow-up treatment is complicated.
  • the traditional process is semi-continuous, generally a reaction kettle and a matching reactive distillation column are used as the reaction device, and the water and excess alcohol produced by the reaction are distilled off while the reaction is completed, and the corresponding dimethyl maleate and the catalyst are obtained.
  • the mixture, distilled water and excess alcohol are rectified to recover valuable methanol.
  • the mixed solution is then subjected to alkaline washing or water extraction to obtain an intermediate product after removing the catalyst.
  • the intermediate product is further subjected to double-column rectification to achieve light and weight removal to obtain pure dimethyl maleate.
  • This traditional production process will lose a considerable amount of dimethyl maleate during water washing or alkali washing; at least four rectification processes are involved in the traditional process, and the energy consumption is also quite high. Even so, this traditional process is still the mainstream process for the production of DMM in China because of its low technological difficulty and small equipment investment.
  • Dimethyl maleate is also an important organic chemical raw material, especially through catalytic hydrogenation, it can be used to produce 1,4-butanediol, tetrahydrofuran and ⁇ -butyrolactone, so that maleic anhydride and methanol can react to produce maleic acid.
  • the method of dimethyl acid has received great attention. This has been described in detail in many patents, such as CN87105388A, CN200410032595.5, CN87105338A, CN87105388A, EU0255399A2, WO90/08127, US4795824, US4751334, WO88/00937, US4584419, etc.
  • Davy Mckee has developed a maleic anhydride esterification process using methanol as an esterification agent.
  • the advantages of this process are: the use of catalyst reactive distillation process, the separation of methanol and water after esterification becomes easy; the volatility of dimethyl maleate is increased, and the operating range of gas-phase hydrogenation is widened; The esterification conversion rate of enedioic anhydride is as high as 99.5%, no need to recycle unreacted maleic anhydride and monomethyl maleate, only methanol is recycled; the process is simplified, and the total project investment is reduced by about 15%. Since 1987, a series of patents have been published for the aforementioned route developed by the British company Davy Mckee. Although this technology has certain advanced nature, it also has its limitations.
  • the British Davy Mckee Company transferred several sets of equipment from several Chinese companies, none of which are still driving.
  • the defects of some of the technology are also one of the reasons for its failure.
  • the specific reasons are: the process of the esterification process of maleic anhydride (MAH) and methanol is not described in depth, and the reaction process and heat release of the two-step esterification reaction are not deeply understood, resulting in the operation of the equipment during the operation.
  • the sulfuric acid was removed, and then the sulfuric acid corroded the catalyst packaging mesh bag and the subsequent filter screen made of 316L stainless steel, causing the acidic resin with polystyrene as the skeleton to enter the next section, causing tens of millions of economic losses to the factory within a week. .
  • patent EU0255399A2 Represented by patent EU0255399A2, this patent describes their process route in detail by taking maleic anhydride (MAH) esterification to produce diethyl maleate as a specific example.
  • the purpose is to produce diethyl maleate by esterification of maleic anhydride, and then hydrogenation to produce BDO, tetrahydrofuran and ⁇ -butyrolactone.
  • maleic anhydride and excess ethanol are mixed to carry out esterification reaction to first generate maleic acid monoethyl ester.
  • the monoesterification reaction is carried out under the conditions of 0.1 MPa and 50-80° C. without catalyst, and the yield is 99%.
  • the monoethyl ester is further reacted with ethanol for diesterification to generate diethyl maleate.
  • a catalyst distillation process is adopted, and a solid acid ion exchange resin is used as a catalyst.
  • Excess ethanol and water are removed from the reactor by rectification, and the generated diethyl maleate gas is further purified by distillation, and the unreacted monoethyl ester is removed and recycled to the reactor.
  • By-products which account for 1% of the total product, can be removed by combustion.
  • the excess ethanol recovered from the reaction part is rectified to remove water, and then mixed with the ethanol recycled from the product rectification and returned to the esterification reactor.
  • the Chinese patent has CN103360252B, and the first paragraph of its technology is modeled on the technology of Davy Mckee Company of the United Kingdom.
  • the first monoester reactor has a temperature of 110-130° C., a pressure of 0.85-0.95 Mpa, the feeding molar ratio of maleic anhydride/methanol is 2-3, and the residence time is 0.01-0.09h. This process route is in fact inappropriate and cannot solve the most fundamental and fatal problems accompanying the Davy process.
  • the double esterification reactor proposed in this patent is divided into multiple sections, 1-4 fixed beds, the top temperature is 80-130°C, the pressure is 0.15-0.25Mpa; the bottom temperature is 160-180°C, the pressure is 0.2-0.4Mpa, and the residence time is 0.02 ⁇ 0.08h.
  • the fixed bed reaction temperature is 80 ⁇ 180°C, and the pressure is 0.1 ⁇ 0.5Mpa.
  • the second stage seems reasonable and is different from the British Davy Mckee company, but for a two-stage reversible esterification dehydration reaction, this design is not practical at all: only a four-stage reaction equilibrium is impossible achieve its stated reaction progress (conversion rate) and effect.
  • the patent does not have a practical case of industrialization in China or even in the world.
  • the Chinese patent CN107473966A proposes to absorb the maleic anhydride in the middle of the n-butane process and the benzene process maleic anhydride reaction mixture with DMM, the final reaction product of maleic anhydride and methanol, and then learn from the technology of Davy Mckee Company in the United Kingdom, with its mixture and Methanol and other alcohols react to generate dimethyl maleate, DMM, etc.
  • the technology described in this patent seems reasonable, but in fact there is a fatal problem: in traditional processes, dibutyl phthalate DBP (or dibutyl cyclohexanedioate) is used as an absorbent in the production of maleic anhydride.
  • DBP The biggest feature of DBP is that it can dissolve the maleic anhydride produced by the partial oxidation of n-butane and benzene, and is incompatible with the by-product water. Because the production of maleic anhydride, n-butane and benzene are also partially oxidized to generate a large amount of water.
  • the DMM and water used as the absorbent in this patent are partially miscible.
  • DMM will be layered with the water phase, the DMM dissolved in the water layer is not only difficult to recover, but also increases the system wastewater discharge and treatment difficulty.
  • DMM oil The phase also contains water, which increases the difficulty of rectification. Therefore, the patent is not engineering feasible.
  • the reaction temperature is 60 ⁇ 90°C
  • the pressure is 0.2 ⁇ 0.8Mpa
  • the acid anhydride/methanol (molar ratio) 2 ⁇ 6
  • the maleic anhydride space velocity is 0.5 ⁇ 1h -1
  • the conversion rate of maleic anhydride reaches over 95%
  • the conversion rate of DMM the representative product of the reaction process
  • the route is (1) esterification of maleic anhydride and lower aliphatic alcohol to obtain diol maleate; (2) selective hydrogenation of the diol maleate obtained in step (1), obtaining diol succinate; (3) performing transesterification and polycondensation of the diol succinate obtained in step (2) and aliphatic diol to obtain the desired ultra-high molecular weight PBS.
  • Maleic anhydride and BDO are in a state of overcapacity in China and even in the world, and there are abundant sources of raw materials. Therefore, the process route of producing PBS with maleic anhydride, hydrogen and BDO as raw materials is the most feasible and can solve the white pollution caused by traditional plastics in China and even the world, and is the most promising process route.
  • maleic anhydride and lower aliphatic alcohol are used for esterification to obtain diol maleate, which is the first stage of the latter two-step reaction and the basis of the latter two steps. Optimized for large-scale industrial production.
  • the technical problem to be solved by the present invention is to provide a process method for preparing diol maleate with little process pollution, low energy consumption and high product yield, aiming at the current situation of the prior art.
  • the technical scheme adopted by the present invention to solve the technical problem is "a process method for preparing diol maleate", which is characterized in that it comprises the following steps:
  • maleic anhydride and lower fatty alcohol carry out pre-esterification reaction in the pre-esterification reactor to obtain the mixture of corresponding water, alcohol, maleic acid monoalcohol ester, maleic acid glycol ester and a small amount of maleic anhydride;
  • Both the two steps of the esterification reaction are reversible reactions, in which both the reaction (1) and the reaction (2) can occur without a catalyst, but usually the reaction (2) is performed with a catalyst and a reactive distillation process. It can be done quickly and thoroughly.
  • Reaction (2) needs to use the process of reactive distillation to distill out the water and remaining alcohol produced by the reaction to obtain a higher content of dimethyl maleate (DMM).
  • DDM dimethyl maleate
  • the esterification of alcohols such as ethanol and propanol with maleic anhydride and methanol is similar to the esterification reaction of maleic anhydride and methanol, and the corresponding maleic diol esters such as diethyl maleate (DEM) and dipropyl maleate (DPM) can also be obtained. Wait.
  • the traditional processes of these esterification reactions generally use sulfuric acid, p-toluenesulfonic acid and other acid catalysts, the reaction temperature is 90-140 ° C, the reaction pressure is 0.1-1.6 Mpa absolute pressure, and the reaction device includes an esterification pre-reactor, an intermediate flash evaporation Rectification column, reactive rectification column, alcohol-water separation rectification column and product rectification column, etc.
  • the traditional process generally takes sulfuric acid or p-toluenesulfonic acid as a catalyst, and takes a reaction kettle and a reaction distillation column that is allocated as a reaction device to complete the reaction (1) and the reaction (2), and simultaneously steam the reaction (2) to produce
  • the water and excess alcohol are obtained to obtain the corresponding mixture of diol maleate and catalyst, and the mixture of distilled water and excess alcohol is rectified to recover valuable fatty alcohol.
  • the mixed solution is then subjected to alkaline washing or water extraction, and then an intermediate product is obtained.
  • the intermediate product is further de-lighted and de-weighted by double-column rectification to obtain pure maleic acid glycol ester.
  • This traditional production process generally uses more polluting and corrosive sulfuric acid or p-toluenesulfonic acid as a catalyst; a considerable amount of maleic acid glycol will be lost during water washing or alkali washing; at least four catalysts are involved in the traditional process. Distillation process, energy consumption is quite high. Although the traditional process can obtain pure products, it is obviously not the optimal process path. While learning from and summarizing traditional processes, we developed a new process path for the production of diol maleate through experiments and process simulation, realizing the continuous production of diol maleate and solving the pollution of traditional processes. Large, high energy consumption and low product yield problems.
  • the first solution is: in the step (2) A reactive distillation column with a flash section is used for flash distillation and reactive distillation.
  • the reactive distillation column is divided into three sections. The upper section is a stripping section for methanol and water, and the temperature is controlled between 80 and 100 °C.
  • the middle section is the flashing section of the reaction mixture, the temperature is controlled and distributed between 220 and 100 ° C from top to bottom, and the lower section is a plurality of trays arranged in sequence from top to bottom, and the temperature is controlled and distributed between 100 and 130 ° C, At least 6 trays are filled with macroporous acid resin catalyst, in this reactive distillation column, the mixture produced by pre-esterification flows from top to bottom, and the water produced by lower aliphatic alcohol and diesterification is in gaseous state from the bottom. And upstream flow.
  • the second scheme is: in the step (2), a flash distillation column and a reactive distillation column are used to carry out flash distillation and reactive distillation respectively, and the reactive distillation column is divided into two sections, and the lower section is composed of a plurality of The lower trays are arranged in sequence, and the temperature is controlled between 100 and 130°C. At least 6 trays are filled with macroporous acidic resin catalysts.
  • the pre-esterification reaction mixture flows from top to bottom.
  • the lower aliphatic alcohol and the water produced by the further esterification reaction flow countercurrently from bottom to top in a gaseous state.
  • the operating pressure of the flash distillation column is 0.1-0.5Mpa, and the temperature is controlled between 90-130°C.
  • Both schemes are beneficial to make the whole process simple and easy to operate in the process of producing diol maleate, and further enable each section of the process to operate smoothly and avoid uncontrollable fluctuations.
  • the second scheme adds a flash distillation column. Although the investment is increased, the total number of trays of the reactive distillation column is reduced, the height of the column is reduced, and the reaction concentration is reduced. The cost and difficulty of making the distillation column. At the same time, due to the low temperature of the material entering the reactive distillation column, the difficulty of operation is reduced, the operation of the entire column is simpler and more stable, the possibility of over-temperature of the catalyst tray is reduced, and the deacidification of the catalyst due to over-temperature is further reduced. risk, which is beneficial to prolong the life of the catalyst.
  • the operating pressure of the reactive distillation column is 0.1-0.5Mpa.
  • the corresponding lower aliphatic alcohol and the water produced by the diesterification reaction are distilled from the top of the reactive distillation column, and then separated in the alcohol recovery column, and the obtained lower aliphatic alcohol is recycled.
  • the reactant obtained at the bottom of the reactive rectification tower is subjected to double-column rectification to achieve delight and weight removal to obtain maleic acid glycol ester with a purity of more than 99.5%.
  • the pre-esterification reactor is a shell-and-tube reactor, and the tube side of the shell-and-tube reactor is used for the pre-esterification reaction.
  • the process is used for the passage of lower aliphatic alcohols.
  • the mixture produced by the pre-esterification reaction in the tube process of the shell-and-tube reactor is sent to the upper part of the reactive distillation column, or after flash distillation and then enters the corresponding reactive distillation column.
  • the lower aliphatic alcohol in the shell side of the shell reactor is vaporized to remove the heat released by the pre-esterification reaction and then sent to the lower part of the reactive distillation column.
  • the molar ratio of maleic anhydride and fatty alcohol in the tube side of the shell-and-tube reactor is 1:1.5-4, the reaction temperature is 60-220°C, the reaction pressure is 0.1-1.6MPa, and the reaction time is 0.1-8h;
  • the temperature of the shell fatty alcohol of the shell-and-tube reactor is 80-200°C, the reaction pressure is 0.1-1.6Mpa, and the vapor-liquid equilibrium vaporization of the corresponding alcohol at this temperature can be maintained, and the corresponding amount is sufficient to remove the reaction generated The heat shall prevail.
  • the lower aliphatic alcohol is at least one of methanol, ethanol, propanol and butanol.
  • the technical route of the present invention is to use maleic anhydride as the initial raw material to produce diol maleate, and the raw material of maleic anhydride can be n-butane or pure benzene, and the source of raw materials is abundant, and now maleic anhydride is already a hundred and one in China.
  • the current production capacity of 10,000-ton-level bulk chemicals in China is over 2 million tons per year, and the annual output is over 1 million tons.
  • Various large companies are actively expanding production; their sources are also very rich;
  • the DMM demand required for PBS with more than one million tons provides a basic guarantee;
  • the addition ratio of alcohol in the maleic anhydride first esterification section is improved, the reaction temperature and the reaction pressure are improved, the reaction time is prolonged, and the two-step reaction is carried out lower, so as to be beneficial to reach as far as possible In its equilibrium state, it is possible to avoid the excessive occurrence of the monoesterification reaction with a large exothermic heat and the diesterification reaction with the same exothermic heat on the trays in the subsequent reactive distillation column, reducing the excess heat caused by the reaction in the later stage.
  • Fig. 1 is the structural representation of the reaction device of the maleic acid glycol ester adopted in the embodiment of the present invention C2;
  • FIG. 2 is a schematic structural diagram of the reaction device of the maleic acid glycol ester adopted in Example C1 of the present invention.
  • the present invention uses maleic anhydride and methanol as reaction raw materials to produce dimethyl maleate (DMM) as an example to introduce our newly invented process route.
  • DDM dimethyl maleate
  • the reaction is divided into two steps; the pre-esterification of the first step includes the monoesterification reaction (1) and the diesterification reaction (2), both of which are exothermic and reversible reactions,
  • the reaction equation is:
  • Pipeline 1 is the liquid maleic anhydride from the production system, the temperature is 60°C ⁇ 220°C, the pipe 2 is the methanol recovered by the alcohol recovery tower, the temperature is 35 ⁇ 60°C, the mixed material 4, the temperature is 60 ⁇ 160°C, The pressure is 0.1-1.6Mpa, and the molar ratio of maleic anhydride and methanol is between 1:1.5-4, preferably 1:2-3.
  • the initial stage of the reaction uses methanol from the methanol storage tank via the pump 10 via the branch line 7 of the line 5 .
  • the methanol used in this system is all methanol with purity >99.5% and water content ⁇ 0.5%, including methanol after rectification in alcohol recovery rectification column.
  • the material 4 enters the tubular reactor 9, and the mixture of maleic anhydride and methanol is carried on the tube side; the shell side is methanol, and the pressure is 0.3-1.6Mpa.
  • the two-step reaction of producing dimethyl maleate with maleic anhydride and methanol as reaction raw materials is an exothermic and reversible reaction, and the exothermic heat is removed by the vaporization of methanol in the shell side.
  • the reaction temperature is 60-220°C
  • the pressure is 0.1-1.6Mpa
  • the reaction residence time is 0.1-8h, in order to facilitate the reaction to proceed completely and consume as much maleic anhydride as possible
  • the reaction (2) is also As close as possible to the chemical reaction equilibrium point.
  • the catalyst used in the reactive distillation column is generally a sulfonic acid resin catalyst
  • the sulfonic acid resin catalyst will react with water to remove some sulfonic acid groups to generate sulfuric acid. A large amount of sulfuric acid is removed from the inside, causing the catalyst to deactivate, and the removed sulfuric acid will also corrode the equipment.
  • the flow rates of the reaction lines 2, 3, 5, 6 and 7 in this process section are controlled cooperatively to facilitate the adjustment of the molar ratio of maleic anhydride and methanol and the required amount of vaporized methanol for the reaction exotherm.
  • the mixture containing water, methanol, monomethyl maleate and dimethyl maleate and a small amount of maleic anhydride enters the upper part of the reactive distillation column 12 via the pipeline 8, and the vaporized methanol passes through the pipeline 11. After adiabatic expansion, it enters the lower part of the reactive distillation column 12 .
  • the reactive distillation column is divided into three sections, the upper section is the stripping section for methanol and water, the middle section is the flashing section of the reaction mixture in pipeline 8, and the lower section is a plurality of catalyst trays filled with macroporous acid resin catalysts.
  • the liquids whose main components are monomethyl maleate and dimethyl maleate flow down from the top, methanol and the water produced by the reaction flow countercurrently from bottom to top in a gaseous state, and the reaction of the second step occurs on each tray,
  • the residence time of the liquid reaction is 0.5-4h, and the specific time is determined according to the type of the macroporous acid resin catalyst, the relevant reaction kinetic data and the design of the reactive distillation column.
  • reaction progress of the entire reactive distillation column material under the action of the catalyst reaches more than 99.5%, the content of the liquid material dimethyl maleate is more than 95%, and the rest are methanol, reaction by-products and traces of maleic anhydride and malic acid. Monomethyl methanoate.
  • the pressure of the reactive distillation column 12 is controlled at 0.1 ⁇ 0.5Mpa, the temperature of the stripping section of methanol and water is controlled at 80 ⁇ 100°C, and the top distillate contains 10 ⁇ 20% of water, 80 ⁇ 90% of methanol and a small amount of The reaction by-products of dimethyl ether and traces of dimethyl maleate.
  • the distillate is condensed in the heat exchanger 18 through the pipeline 14, and then sent to the liquid collection tank 19. Part of the distillate is refluxed to the column 12 through the pump 20, and the rest is passed through the pump 21 and then through the pipeline 22 to the methanol recovery column 23.
  • the middle of the reactive distillation column 12 is the flashing section of the reaction mixture on several trays, and the temperature is controlled to be distributed at 220-100°C.
  • the lower section of reactive distillation tower 12 is a catalyst tray, and the temperature is controlled at 100 to 130° C.
  • the 2 to 4 trays at the bottom are not loaded with catalysts, so as to facilitate the upward flow of superheated methanol steam and maleic acid glycol esters in the lower part.
  • the liquid material reacts and exchanges heat to avoid damage to the catalyst due to overheating of the tray.
  • the bottom of the reactive distillation column 12 is equipped with a reboiler 17 (temperature 130-150° C.), which uses medium-pressure steam to control the temperature of the entire column body.
  • the bottom liquid is the main mixture of dimethyl maleate, the molar content of DMM is more than 95%, and the rest are dissolved methanol, reaction by-products and trace amounts of maleic anhydride and monomethyl maleate.
  • the mixed liquor is pumped out of the system via line 13 to heat exchanger 15 and then to pump 16.
  • the mixture is subjected to classic double-column rectification (not shown in the figure) to obtain dimethyl maleate DMM with a purity of >99.5%.
  • the light components are returned to the methanol recovery tower through pipeline 36. After the heavy components are taken and partially discarded, Partial reuse is returned to the system via pipe 1.
  • the materials in pipeline 36 and pipeline 22 enter methanol recovery tower 23, the temperature is controlled at 60 ⁇ 120°C, and the pressure is controlled at 0.1 ⁇ 0.2Mpa.
  • the top of the tower is the steam of methanol and by-product dimethyl ether DME, which is sent to the heat exchanger 29 through the pipeline 25 and to the sump 30, and part of the reflux liquid is formed through the pump 31, and the non-condensable gas (mainly dimethyl ether and tetrahydrofuran, etc.) ), recovered through 32 after freezing and condensing.
  • the methanol at the top of the column is extracted through the pipeline 26, and is further divided into the pipeline 2 and the pipeline 3 through the heat exchanger 27 to the pump 28 to form the methanol reuse.
  • the liquid at the bottom of the kettle is waste water with a concentration of more than 99.5%, and contains ⁇ 0.5% of dimethyl maleate DMM, after the heat exchanger 34, the pump 35 is sent to the biochemical treatment pool.
  • Dimethyl maleate DMM is a biodegradable organic matter, and the wastewater treatment pressure is small.
  • Fig. 2 also utilizes maleic anhydride and methanol as reaction raw materials to produce dimethyl maleate process route, and the difference from Fig. 1 is that Fig. 2 adds a flash distillation column 37.
  • the water and methanol mixture at the top of the flash distillation column 37 is mixed with the water and methanol mixture from the pump 21 through the pipeline 39, and then enters the methanol recovery column through the pipeline 22.
  • the content of water and methanol in the pipeline 39 varies according to the ratio of maleic anhydride and methanol in the tubular reactor 9 and the reaction progress, and contains 20 to 35% of water, 65 to 80% of methanol and a trace amount of maleic anhydride. Dimethyl acid.
  • the flash distillation column 37 has the dual properties of flash distillation and rectification. The insufficient heat is provided by the low pressure steam of the bottom reboiler 40.
  • the pressure of the flash distillation column 37 is 0.1-0.5Mpa, and the temperature of each tray is 0.1-0.5Mpa. The change is large, and the entire tower is controlled between 100 and 220 °C.
  • the bottom material of the flash distillation column 37 after passing through the cooler 41, enters the line 38 to the reactive distillation column 12.
  • the reactive distillation column 12 of the process shown in Fig. 2 and the process shown in Fig. 1 is also different in structure from Fig. 1: the reactive distillation column 12 of Fig. 2 is divided into two sections, and the upper section is the stripping section of methanol and water, The lower section is a plurality of catalyst trays filled with macroporous acidic resin.
  • the material composition of pipeline 38 entering into reactive distillation column 12 is also different from that in the flow process of Fig. 1.
  • the pipeline 38 entering the middle and upper part of the tower is a mixture of monomethyl maleate, dimethyl maleate and a small amount of maleic anhydride. mixture at a temperature of 100-120°C.
  • the pressure of the reactive distillation column 12 is controlled at 0.1 to 0.5 Mpa
  • the temperature of the stripping section of methanol and water is controlled at 80 to 100°C
  • the overhead distillate contains 10 to 20% of water and 80 to 90% of methanol.
  • a small amount of reaction by-products dimethyl ether and traces of dimethyl maleate is controlled.
  • the temperature of the multiple catalyst trays in the lower section of the reactive distillation column 12 is controlled at 100-130°C, and the 2-4 trays at the bottom are not loaded with catalysts, so as to facilitate the upper part of the superheated methanol steam and the maleic acid glycol ester as the catalyst.
  • the main liquid material reacts and exchanges heat to avoid damage to the catalyst due to overheating of the tray.
  • the process of FIG. 2 adds a flash distillation column 37, although the investment is increased, the total number of trays of the reactive distillation column 12 is reduced, the height of the column is reduced, and the The manufacturing cost and difficulty of the reactive distillation column 12.
  • the difficulty of operation is reduced, the operation of the entire tower is simpler and more stable, the possibility of overheating of the catalyst tray is reduced, and the overtemperature of the catalyst is further reduced.
  • the risk of deacidification is beneficial to further prolong the life of the catalyst.
  • Embodiments A1-A5 Exploration of maleic anhydride/methanol molar ratio
  • Example A1 A2 A3 A4 A5 Component/molar ratio 1/1.2 1/1.5 1/2.0 1/2.5 1/3.0 Methanol (%) 5.31 4.56 5.23 14.5 22.67 water(%) 2.03 6.8 7.57 7.8 6.33 Maleic anhydride (%) 12.49 6.57 4.01 0.30 0.55 Monomethyl maleate (%) 63.89 29.12 36.50 14.8 19.76 Dimethyl maleate (%) 16.27 54.4 56.63 62.4 50.69 other(%) 0 0 0 0.20 0
  • the first-step esterification reaction is carried out in a shell-side ⁇ 108X4 (SS304) and 7 ⁇ 25X3 (SS316L) tubular reactors.
  • the tube-side is filled with ⁇ 5X5 ceramic Raschig rings (bulk density 1.2, pores rate 0.46), the circulating heat transfer oil with a fixed temperature is used as the heat transfer medium on the shell side; the maleic anhydride is melted and kept at a constant temperature of 90 °C, the feeding pump head and the pipeline are heated at 110 °C with electric tracing; methanol is kept at 50 °C, The methanol pump head and pipeline are also kept at 50°C; the pressure of the two product collection tanks of the product collection system is maintained at 1.0Mpa in the mode of nitrogen back pressure, so that the reaction can be carried out continuously.
  • the molar ratio of maleic anhydride and methanol to be fed is 1:2.5; the reaction was carried out at 110°C and 130°C, respectively; the residence time of the reactants in the shell-and-tube reactor was 1h, 2h and 4h, and the compositions of the reaction products were listed in Table 2 together;
  • the bottom of the flash distillation column is equipped with a 2500ml three-necked flask, which is heated by an oil bath; the whole flash distillation column is equipped with upper reflux and extraction, The feed in the middle and the discharge at the bottom are carried out by a peristaltic pump equipped with a silicone tube; the entire flash distillation column is operated at a slight positive pressure, and the temperature of the column body is 80-120 ° C. The temperature was controlled at 140°C. After steaming most of the water and methanol in the reaction mixture, finally obtaining the mixture 3100g that the main component is monomethyl maleate and dimethyl maleate is for subsequent use, and its composition is shown in Table 3;
  • Table 3 Composition (w%) of the product after flash distillation at 110°C-4h
  • (2) second-stage esterification reaction the 130 °C-2h gained in the above-mentioned steps is sent into the reactive distillation column without rectifying product 4000g; the reaction conditions are basically the same as the step (3) of embodiment C1, and different places There are: the reactive distillation column is divided into three sections, the second section of the rectification column is not filled with catalyst, and the ceramic Raschig ring of ⁇ 5mm is filled with 100mm; 4000g of methanol, water, monomethyl maleate and dimethyl maleate mixture Add dropwise from the top of the second section of the rectifying column at a speed of 4ml/min, and the methanol is pumped into the bottom of the three-necked flask at the bottom of the rectifying column through a long glass dropper with a flow rate of 1.2ml/min; after 20 hours of reaction, This embodiment also obtains the reactant of 3200g at last, and wherein dimethyl maleate content is 96%;
  • (1) first-stage esterification reaction select the same shell-and-tube reactor with embodiment B, the molar ratio of maleic anhydride and n-propanol is 1:2.5, and the temperature of reaction is 140 °C, and the dual product of the product collection system The pressure of the collection tank is maintained at 0.8Mpa with the mode of nitrogen back pressure, and the retention time of the material in the reactor is 3h; finally, 5000g of product is obtained, and the composition is shown in Table 4;
  • Table 4 Composition (w%) of the first-stage esterification section reaction product of DPM synthesis at 140°C-3h
  • Table 5 Composition (w%) of DPM synthesis product after rectification at 140°C-3h

Landscapes

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

Abstract

本发明公开了一种制备马来酸二醇酯的工艺方法，其特征在于包括有以下步骤：(1)将顺酐和低级脂肪醇在预酯化反应器中进行预酯化反应，得到相应水、醇、马来酸单醇酯、马来酸二醇酯以及少量顺酐的混合物；(2)将预酯化反应产生的反应混合物先进行闪蒸和精馏，再在酸性树脂的催化下和相应的低级脂肪醇通过反应精馏完成酯化，得到所需的马来酸二醇酯。与现有技术相比，本发明的制备马来酸二醇酯的工艺方法工艺污染小、能耗低和产品收率高，整个工艺操作平稳，避免了反应精馏塔塔盘上的超温，延长了催化剂的使用寿命。

Description

一种制备马来酸二醇酯的工艺方法 技术领域
本发明涉及精细化学品和聚合物单体合成技术领域，具体指一种制备马来酸二醇酯的合成工艺方法和路线。
背景技术
马来酸二醇酯中最具代表性的是马来酸二甲酯(Dimethyl Maleate，DMM)、马来酸二乙酯(Diethyl Maleate，DEM)和马来酸二丙酯(Dipropyl Maleate，DPM)，前二者都是重要的精细化工原材料。其他马来酸二醇酯的生产工艺流程和方法基本类似于前二者，现有文章和专利中基本以DMM、DEM为代表介绍其工艺路线和方法。马来酸二甲酯(DMM)为无色粘稠液体，沸点(1.0bar)200.4℃、熔点：19.0℃、相对密度(25℃)1.1462g/cm 3，为生产油漆、涂料、粘结剂、杀虫剂、净水剂、光学材料、防缩整理剂和防锈添加剂和等产品的重要化工原料，国内外对马来酸二甲酯的需求量日趋增长。
传统的马来酸二甲酯的生产方法多采用以顺酐为原料，在以硫酸或者对甲苯磺酸的催化作用下与甲醇进行酯化反应，该生产工艺虽然催化剂的催化活性高、价格便宜，但存在副反应多、产品容易异构化成富马酸二甲酯DMF、设备腐蚀严重、后续处理复杂等缺点。传统的工艺为半连续式，一般以一个反应釜和搭配的反应精馏塔为反应装置，完成反应同时蒸出反应产生的水和多余的醇，得到相应的马来酸二甲酯和催化剂的混合液，蒸出的水和多余的醇的混合物进行精馏，回收有价值的甲醇。混合液接下来进行碱洗或者水萃取，除去催化剂后得到中间产品。中间产品再进一步地进行双塔精馏实现脱轻、脱重，得到纯品马来酸二甲酯。这种传统生产工艺在水洗或者碱洗的时候会损失相当量的马来酸二甲酯；传统工艺中至少涉及到4个精馏过程，能耗也相当高。即便如此这种传统工艺因为工艺难度小，设备投资小，在中国当下依然是生产DMM的主流工艺。
马来酸二甲酯还是一种重要的有机化工原料，特别是通过催化加氢可用来生产1,4-丁二醇、四氢呋喃和γ-丁内酯，使得马来酸酐和甲醇反应生产马来酸二甲酯的方法受到了极大的关注。这在很多专利中已经进行了详细说明，如CN87105388A、CN200410032595.5、CN87105338A、CN87105388A、EU0255399A2、WO90/08127、US4795824、US4751334、WO88/00937、US4584419等。
Davy Mckee公司开发出了以甲醇为酯化剂的顺丁烯二酸酐酯化工艺。该工艺的优点是：采用催化剂反应精馏工艺，酯化后的甲醇和水分离变得容易；增加了马来酸二甲酯的挥发度，使之气相加氢的操作范围变宽；顺丁烯二酸酐的酯化转化率高达99.5％， 不需要未反应的顺丁烯二酸酐和顺丁烯二酸单甲酯循环，只有甲醇循环；简化了流程，工程总投资比以前减少约15％。自1987年以来，由英国Davy Mckee公司开发公布了前面所述路线的一系列专利。该技术虽然有一定的先进性，但是也有它的局限性。具体来说，英国Davy Mckee公司转让中国多家公司的数套的设备，没有一套现在依然开车的。这除了市场变化，该工艺失去竞争力的原因外，该技术的某些技术存在的缺陷，也是其失败的原因之一。具体原因有：其工艺过程对顺酐(MAH)和甲醇的酯化过程没有深入描述，对其两步酯化反应的反应进程、放热量没有深刻的理解，导致其设备在运行过程中可操作性差；催化剂床层因反应放热出现的局部热点不能有效的控制，酸性树脂催化剂容易因超温脱酸失活，其使用的酸性树脂催化剂寿命较短(4～5个月)；催化剂脱下的硫酸在高温下对反应精馏塔的塔内件、催化剂网状包装袋产生强烈腐蚀，导致最终的产品质量出现重大问题。中国某公司采用Davy Mckee公司的技术的装置上运行时发现，因其部分流程的设计不合理，反应精馏塔的温度极难操控，导致装有催化剂塔盘的超温，塔盘上酸性树脂脱下了硫酸，接着硫酸腐蚀了316L不锈钢材料的催化剂包装网袋和后续过滤网，导致以聚苯乙烯为骨架的酸性树脂进入下个工段，一周内对该工厂造成了数千万的经济损失。
以专利EU0255399A2为代表，该专利以顺丁烯二酸酐(MAH)酯化生产马来酸二乙酯为具体实施例详细描述他们的工艺路线。目的是以顺丁烯二酸酐酯化生产马来酸二乙酯，然后加氢生产BDO、四氢呋喃和γ-丁内酯。该工艺是由顺丁烯二酸酐和过量乙醇混合进行酯化反应先生成马来酸单乙酯。单酯化反应在0.1MPa、50～80℃条件下进行，不需催化剂，收率为99％。单乙酯再进一步和乙醇进行双酯化反应生成马来酸二乙酯。这一步采用催化剂精馏工艺，固体酸离子交换树脂为催化剂，反应温度100～130℃、压力为0.1MPa、双酯收率可达98％。通过精馏把过量乙醇和水从反应器中去除，生成的马来酸二乙酯气体进一步蒸馏提纯，除去未反应的单乙酯使之循环回反应器中。占生成物总量1％的副产物可燃烧去除。从反应部分回收的过量乙醇经精馏除水后再与从产物精馏处循环回的乙醇混合返回酯化反应器。该技术虽然有一定的先进性，但是也有像该公司其他一系列专利同样的局限性。
中国的专利有CN103360252B，其技术的第一段仿照英国Davy Mckee公司的技术。其第一个单酯反应器，温度110～130℃，压力0.85～0.95Mpa，顺酐/甲醇的加料摩尔比为2～3，停留时间0.01-0.09h。这种工艺路线事实上是不合适的，根本无法解决戴维工艺伴随的最根本、最致命的问题。该专利所提出的双酯化反应器分多段，固定床1-4个，顶部温度80～130℃，压力0.15～0.25Mpa；塔底温度160～180℃，压力0.2～0.4Mpa，停留时间0.02～0.08h。固定床反应温度80～180℃，压力0.1～0.5Mpa。其第二段看似合理，并与英国Davy Mckee公司有了不同，但是对一个两段均可逆的酯化脱水反应来说，这种 设计根本不符合实际：仅仅四级的反应平衡是不可能达到其所述的反应进程(转化率)和效果的。该专利在中国甚至全世界范围内没有一个工业化的实际案例。
中国某公司的专利CN102908955B和专利CN102911053B在中国也获得了授权，但是这个专利中所叙述的流程和反应状况，和真实顺酐和甲醇的反应相去甚远。这两个专利都没有提到顺酐和甲醇酯化反应第一步和第二步并不是孤立存在的，均有其特定的平衡存在。其单酯化反应器的设计并不合理，反应的放热计算与设计与真实情况不符，亦无法避免酯化反应过度放热在装有酸性树脂的反应精馏塔的塔盘上过多发生，因此也无法避免酸性树脂催化剂因超温脱酸失活、催化剂脱下的稀硫酸在高温下对反应精馏塔中的塔内件和催化剂网状不锈钢包装袋的强烈腐蚀。这两个专利均对反应精馏塔的塔盘设计，催化剂和催化剂的装填提出了新的改进，但由于其根本性的问题没有解决，其应用也受到了限制。事实情况是：该专利在中国甚至全世界范围内没有一个工业化的实际工业化案例。
中国的专利CN107473966A，提出了以顺酐和甲醇的最终反应产物DMM来吸收正丁烷法和苯法顺酐反应混合气中间的顺酐，然后再借鉴英国Davy Mckee公司的技术，以其混合物和甲醇等醇类反应，生成马来酸二甲酯DMM等。该专利所述技术看似合理，其实存在着致命的问题：传统工艺中都是使用邻苯二甲酸二丁酯DBP(或者环己烷二酸二丁酯)作为顺酐生产过程中的吸收剂，DBP最大特点是能溶解正丁烷和苯法部分氧化产生的顺酐，而和副产的水不相溶。因为生产顺酐的同时，正丁烷和苯也部分氧化生成了大量的水。该专利所用作为吸收剂的DMM和水是部分互溶的，虽然DMM会和水相分层，而溶解在水层中的DMM，不仅难以回收，更增加了系统废水排放量和处理难度，DMM油相也含水，精馏难度加大。因此，该专利是工程上不可行的。
中国专利CN102070448B，“一种制备丁二酸二甲酯的方法”，他们第一步合成马来酸二甲酯DMM段的工艺基本仿照了Davy Mckee公司的工艺。但是他们改动的是第一步预酯化用了固定床和酸性树脂催化剂，与Davy Mckee公司的工艺过程不同，其第二步反应精馏段和戴维工艺基本相同。我们按照他们的实施例重复了他们的实验，发现实施例的结果不能重复。另外：顺酐和甲醇第一步预酯化的反应，其平衡比例也与其实施例中不符；并且第一步的顺酐甲醇预酯化放热量非常大。该专利中第一段预酯化使用了固定床反应器，其所述：反应温度60～90℃，压力0.2～0.8Mpa，酸酐/甲醇(摩尔比)＝2～6，顺酐空速0.5～1h -1；顺酐的转化率达到了95％以上，反应进程的代表产物DMM的转化率在87％以上。根据我们的实验数据和理论模拟，这种转化率和选择性不仅不可能，其固定床中反应器的绝热温升也会在120℃以上，反应混合物的温度在没有移热和物料汽化的情况下将会飞温到180℃以上。根据该专利提到的、其实施例所用的几种酸性树脂，没有一种能耐温140℃以上。因此可以确定，其专利描述的，由顺酐和甲醇生产马来酸 二甲酯DMM，再选择性加氢的工艺是不可行的。
本专利申请人的在原先申请的“一种超高分子量PBS及其制备方法”申请公布号“CN110563933 A”的提出了以顺酐(MAH)、氢气(H 2)和1、4丁二醇(BDO)为主原料三步法生产聚丁二酸丁二醇酯PBS的新路线。其路线是(1)将顺酐和低级脂肪醇进行酯化反应，得到马来酸二醇酯；(2)将步骤(1)制得的马来酸二醇酯进行选择性加氢反应，得到丁二酸二醇酯；(3)将步骤(2)制得的丁二酸二醇酯和脂肪族二醇进行酯交换和缩聚反应，得到所需的超高分子量的PBS。顺酐和BDO在中国国内甚至是世界范围内都处于产能过剩的状态，原料来源丰富。因此以顺酐、氢气和BDO为原料来生产PBS的的工艺路线是最有可行性、能解决中国甚至全世界范围内因传统塑料带来的白色污染，是最有发展前途的工艺路径。该工艺路径中，以顺酐和低级脂肪醇进行酯化反应，得到马来酸二醇酯，是后面两步反应的前段，也是后两步的基础，最需要在以前的工艺过程的基础上进行优化，以利于大规模的工业化生产。
发明内容
本发明所要解决的技术问题是针对现有技术的现状，提供一种工艺污染小、能耗低和产品收率高的制备马来酸二醇酯的工艺方法。
本发明解决技术问题所采用的技术方案为“一种制备马来酸二醇酯的工艺方法”，其特征在于包括有以下步骤：
(1)顺酐和低级脂肪醇在预酯化反应器中进行预酯化反应，得到相应水、醇、马来酸单醇酯、马来酸二醇酯以及少量顺酐的混合物；
(2)将预酯化反应产生的混合物先进行闪蒸、精馏，再在酸性树脂的催化下和相应的低级脂肪醇通过反应精馏完成酯化，得到所需的马来酸二醇酯。
其中，以顺酐和甲醇的反应为例，上述顺酐和低级脂肪醇进行单酯化反应和双酯化反应的反应方程式如下：
(1)
Figure PCTCN2022071265-appb-000001
(2)
Figure PCTCN2022071265-appb-000002
该酯化反应两步的均是可逆反应，其中反应(1)和反应(2)均可在没有催化剂的情况下发生，但通常反应(2)都在有催化剂、使用反应精馏工艺的情况下才得以快速、彻底地进行。反应(2)需要用反应精馏的工艺蒸出反应产生的水和剩余的醇，得到较高含量的马来酸二甲酯(DMM)。乙醇和丙醇等醇类与顺酐和甲醇的酯化反应类似，也可以得到相应的马来酸二醇酯，如马来酸二乙酯(DEM)和马来酸二丙酯(DPM)等。这些酯化反应传统的工艺一般用硫酸、对甲苯磺酸和等酸催化剂，反应温度为90～140℃，反应压力为绝压0.1～1.6Mpa，反应装置包括酯化预反应器、中间闪蒸精馏塔、反应精馏塔、醇水分离精馏塔和产品精馏塔等。
传统的工艺一般是以硫酸或者对甲苯磺酸为催化剂，以一个反应釜和所配的反应精馏塔为反应装置，完成反应(1)和反应(2)，同时蒸出反应(2)产生的水和多余的醇，得到相应的马来酸二醇酯和催化剂的混合液，蒸出的水和多余的醇的混合物进行精馏，回收有价值的脂肪醇。混合液接下来进行碱洗或者水萃取，然后得到中间产品。中间产品在进一步的进行双塔精馏实现脱轻脱重，得到纯品马来酸二醇酯。这种传统生产工艺一般使用的污染和腐蚀性较大硫酸或者对甲苯磺酸做催化剂；在水洗或者碱洗的时候会损失相当量的马来酸二醇酯；传统工艺中至少涉及到4个精馏过程，能耗相当高。传统工艺虽然能得到纯品，但显然不是最优的工艺路径。我们借鉴总结传统工艺的同时，通过实验以及过程模拟，开发了一种全新的生产马来酸二醇酯的工艺路径，实现了马来酸二醇酯连续化的生产，同时解决了传统工艺污染大、能耗高和产品收率低的问题。
为了实现物料降温，避免物料因过热在接下来的塔盘上继续反应放热伤害树脂催化剂，需要先对预酯化反应的产物进行闪蒸，第一个方案为：所述步骤(2)中采用带有闪蒸段的反应精馏塔进行闪蒸精馏和反应精馏，该反应精馏塔分三段，上段为甲醇和水的提馏段，温度控制分布在80～100℃之间；中段为反应混合物的闪蒸段，温度自上而下控制分布在220～100℃之间，下段为多个自上而下依次布置的塔盘，温度控制分布在100～130℃之间，至少6个塔盘上装填有大孔酸性树脂催化剂，在该反应精馏塔中，预酯化反应产生的混合物自上而下流动，低级脂肪醇和双酯化反应产生的水以气体状态自下而上逆流流动。第二个方案为：所述步骤(2)中采用闪蒸精馏塔和反应精馏塔分别进行闪蒸精馏和反应精馏，该反应精馏塔分两段，下段为多个自上而下依次布置的塔盘，温度控制分布在100～130℃之间，至少6个塔盘上装填有大孔酸性树脂催化剂，该反应 精馏塔中，预酯化反应混合物自上而下流动，低级脂肪醇和进一步酯化反应产生的水以气体状态自下而上逆流流动，闪蒸精馏塔的操作压力为0.1～0.5Mpa，温度控制分布在90～130℃之间。
两个方案均有利于在生产马来酸二醇酯的工艺中让整个流程简便、易操作，进一步使每一段工艺能平稳操作，避免非可控的波动。第二个方案和第一个方案相比，增加了一个闪蒸精馏塔，虽然增加了投资，但是减少了反应精馏塔的总塔盘数，降低了该塔的高度，减少了反应精馏塔的制作成本和难度。同时，由于进入反应精馏塔的物料温度较低，降低了操作难度，使整个塔的操作更简单平稳，降低了催化剂塔盘超温的可能性，同时也进一步降低了催化剂因超温脱酸的风险，有利于延长催化剂的寿命。
为了利于下部上行过热甲醇蒸汽和以马来酸二醇酯为主的液体物料反应并换热，避免塔盘超温伤害催化剂，所述塔盘中，至少两个位于底部的塔盘不装填有催化剂。
优选地，所述反应精馏塔的操作压力为0.1～0.5Mpa。
为了对低级脂肪醇进行回收，所述反应精馏塔的塔顶蒸出相应的低级脂肪醇和双酯化反应产生的水，然后在醇回收塔中分离，得到的低级脂肪醇回收利用。
为了对产物进行进一步提纯，所述反应精馏塔塔底得到的反应物经过双塔精馏，实现脱轻、脱重，得到纯度在99.5％以上的马来酸二醇酯。
为了充分利用预酯化反应所放的热量，所述的预酯化反应器为管壳式反应器，管壳式反应器的管程用于预酯化反应进行，管壳式反应器的壳程用于供低级脂肪醇通过，管壳式反应器的管程中预酯化反应产生的混合物送入反应精馏塔上部，或经闪蒸精馏后再进入相应的反应精馏塔，管壳式反应器的壳程中的低级脂肪醇汽化移走预酯化反应所放热量后送入反应精馏塔下部。
优选地，所述管壳式反应器的管程中顺酐和脂肪醇的摩尔比为1：1.5～4，反应温度为60～220℃，反应压力为0.1～1.6MPa，反应时间为0.1～8h；所述管壳式反应器的壳层脂肪醇的温度为80～200℃，反应压力为0.1～1.6Mpa，维持相应醇的该温度下的气液平衡汽化即可，相应量足以移走反应产生的热为准。
优选地，所述的低级脂肪醇为甲醇、乙醇、丙醇和丁醇中的至少一种。
与现有技术相比，本发明的优点在于：
(1)本发明的技术路线是以顺酐为最初原料生产马来酸二醇酯，而顺酐的原料可以是正丁烷，也可以纯苯，原料来源丰富，现在顺酐在中国已经是百万吨级的大宗化工品，全国的当前产能在200万吨/年以上，年产量在100万吨以上，各家大公司还在积极扩产；其来源也非常丰富；这就对中国以后发展百万吨以上的PBS所需的DMM需求提供了基本保障；
(2)通过工艺过程优化，让预酯化反应放热对后续酯化反应需要的醇类进行了预 热和汽化，有利于整个流程的进能量综合利用，实现能量的最大化利用；
(3)本发明的工艺路线中提高了顺酐第一酯化段中醇的添加比例，提高了反应温度和反应压力，延长了反应时间，使两步反应进行的更低，以利于尽量达到其平衡状态，因而可以避免放热量较大的单酯化反应和同样放热的双酯化反应在随后的反应精馏塔中的塔盘上过多发生，降低了后面时因反应放热超温导致塔盘上酸性树脂脱酸的可能性，减少了催化剂失活，有利于延长催化剂的寿命，同时减少了催化剂超温脱掉的硫酸对设备的腐蚀。
附图说明
图1为本发明实施例C2所采用的马来酸二醇酯的反应装置的结构示意图；
图2为本发明实施例C1所采用的马来酸二醇酯的反应装置的结构示意图。
具体实施方式
以下结合附图实施例对本发明作进一步详细描述。
如图1所示，本发明利用以顺酐和甲醇为反应原料生产马来酸二甲酯(DMM)为例来介绍我们新发明的工艺路线。
以顺酐和甲醇的反应为例，其反应分两步；第一步的预酯化包括单酯化反应(1)，和双酯化反应(2)，均为放热、可逆的反应，反应方程式为：
(1)
Figure PCTCN2022071265-appb-000003
(2)
Figure PCTCN2022071265-appb-000004
管道1中为生产系统来的液体顺酐，温度为60℃～220℃，管道2中为醇回收塔回收的甲醇，温度35～60℃，混合后为物料4，温度为60～160℃，压力为0.1～1.6Mpa，顺酐和甲醇的摩尔比在1：1.5～4之间，优选1：2～3。反应起始阶段使用由甲醇储罐经泵10经管道5的分支管道7过来的甲醇。该系统所用的甲醇均为纯度>99.5％，含水<0.5％的 甲醇，包括醇回收精馏塔精馏后的甲醇。物料4进入列管式反应器9，管程走顺酐和甲醇的混合物；壳程为甲醇，压力0.3～1.6Mpa。以顺酐和甲醇为反应原料生产马来酸二甲酯的两步反应均为放热、可逆的反应，所放热由壳程中甲醇汽化移走。物料进入反应器9管程后反应温度为60～220℃，压力为0.1～1.6Mpa，反应停留时间为0.1～8h，以利于反应进行彻底，消耗掉尽量多的顺酐，反应(2)也尽量接近于化学反应平衡点。
本工艺段加了过量的甲醇，提高了反应温度和反应压力，延长了反应时间，目的是使反应进行的更快、更彻底。因而可以避免放热量较大的单酯化反应(1)和同样放热的双酯化反应(2)在随后的反应精馏塔中过多的发生。这也进一步降低了后面使用催化剂时因反应放热超温导致酸性树脂脱酸的可能，减少了催化剂失活，有利于延长催化剂的寿命。因为反应精馏塔中使用的催化剂一般为磺酸树脂催化剂，在130℃以上时，磺酸树脂催化剂就会和水反应脱掉部分磺酸基团产生硫酸，极端高温情况下，会在短时间内脱出大量硫酸，造成催化剂失活，脱掉的硫酸还会腐蚀设备。该工艺段中反应管路2、3、5、6和7的流量协同控制，以利于调节顺酐和甲醇的摩尔比和反应放热的所需的汽化甲醇量。
经由反应器9后的反应混合物，含有水、甲醇、马来酸单甲酯和马来酸二甲酯以及少量顺酐的混合物经由管道8进入反应精馏塔12上部，汽化的甲醇经管道11绝热膨胀后进入反应精馏塔12下部。反应精馏塔分三段，上段为甲醇和水的提馏段，中段为管道8反应混合物的闪蒸段，下段为装填有大孔酸性树脂催化剂的多个催化剂塔盘。主成分为马来酸单甲酯和马来酸二甲酯的液体自上流下，甲醇和反应产生的水以气体状态自下而上逆流流动，第二步的反应在各个塔盘上发生，液体反应停留时间为0.5～4h，具体时间根据大孔酸性树脂催化剂的类型、相关反应动力学数据和反应精馏塔的设计确定。反应液到塔12最底部未装催化剂的塔盘时，接触的为纯的甲醇蒸汽。整个反应精馏塔物料在催化剂的作用下反应进程达到99.5％以上，液体物料马来酸二甲酯的含量已经在95％以上，其余为甲醇、反应副产物以及痕量的马来酸酐和马来酸单甲酯。
反应精馏塔12的压力控制在0.1～0.5Mpa，甲醇和水的提馏段温度控制在80～100℃，塔顶馏出物含有10～20％的水、80～90％的甲醇和少量的反应副产物二甲醚和痕量的马来酸二甲酯。该馏出物经管道14到换热器18冷凝，再到集液罐19，部分经泵20回流到塔12，其余的经泵21后经管道22到甲醇回收塔23。反应精馏塔12的中部为数个塔板的反应混合物的闪蒸段，温度控制分布在220～100℃，从上到下因甲醇和水的蒸发实现混合液温度的降低。反应精馏塔12下段为催化剂塔盘，温度控制在100～130℃，最底部的2～4个塔盘不装催化剂，以利于下部上行过热甲醇蒸汽和以马来酸二醇酯为主的液体物料反应并换热，避免塔盘超温伤害催化剂。反应精馏塔12底部配有再沸器17(温度130～150℃)，使用中压蒸汽进性整个塔体的温度控制。塔底液为马来酸二甲酯为主混合 物，DMM的摩尔含量已经在95％以上，其余为溶解的甲醇、反应副产物以及痕量的顺酐和马来酸单甲酯。该混合液经管道13到换热器15，再到泵16打出该系统。该混合物经经典的双塔精馏(图中未示出)后得到纯度>99.5％的马来酸二甲酯DMM，轻组分经管道36回到甲醇回收塔，重组分采取部分弃置后，再部分回用经管道1回到该系统。
管道36和管道22的物料进入甲醇回收塔23，温度控制在60～120℃，压力控制0.1～0.2Mpa，塔23带有杂质侧线采出装置(图中未示出)。塔顶为甲醇和副产物二甲醚DME的蒸汽，经管道25到换热器29，并到集液槽30，部分经泵31形成回流液，不凝气(主要为二甲醚和四氢呋喃等)，冷冻冷凝后经32采出回收。塔顶甲醇经管道26采出，并经换热器27到泵28再分为管道2和管道3形成甲醇的回用。塔底有再沸器33，使用低压蒸汽；釜底液体为99.5％以上浓度的废水，并含有<0.5％的马来酸二甲酯DMM，经换热器34后由泵35打到生化处理池。马来酸二甲酯DMM为可生化降解的有机物，废水处理压力较小。
图2也是利用以顺酐和甲醇为反应原料生产马来酸二甲酯工艺路线，与图1不同的地方在于图2增加了一个闪蒸精馏塔37。经由管式反应器9管程反应后的含有水、甲醇、马来酸单甲酯和马来酸二甲酯以及少量顺酐的反应混合物，经由管道8进入的闪蒸精馏塔37进行闪蒸和精馏。闪蒸精馏塔37塔顶的水和甲醇混合物经管线39和泵21出来的水和甲醇混合物混合后经管道22进入甲醇回收塔。管线39的水和甲醇和含量，是根据管式反应器9中顺酐和甲醇比例以及反应进程的不同变化的，含有20～35％的水、65～80％的甲醇和痕量的马来酸二甲酯。闪蒸精馏塔37具有闪蒸和精馏的双重属性，不足的热量由塔底再沸器40的低压蒸汽提供热量，闪蒸精馏塔37的压力0.1～0.5Mpa，各个塔板的温度变化较大，整个塔控制在100～220℃之间。闪蒸精馏塔37的底部物料，经由冷却器41后进入管线38到反应精馏塔12。
图2所示流程与图1所示流程的反应精馏塔12在结构上也与图1中有所不同：图2反应精馏塔12分两段，上段为甲醇和水的提馏段，下段为多个装填有大孔酸性树脂的催化剂塔盘。进入到反应精馏塔12的管线38物料成分也与图1流程中不同，图2中进入塔中上部的管线38中为马来酸单甲酯、马来酸二甲酯和少量顺酐的混合物，温度为100～120℃。
图2中反应精馏塔12的压力控制0.1～0.5Mpa，甲醇和水的提馏段温度控制在80～100℃，塔顶馏出物含有10～20％的水、80～90％的甲醇和少量的反应副产物二甲醚和痕量的马来酸二甲酯。反应精馏塔12下段的多个催化剂塔盘，温度控制在100～130℃，最底部的2-4个塔盘不装催化剂，以利于下部上行过热甲醇蒸汽和以马来酸二醇酯为主的液体物料反应并换热，避免塔盘超温伤害催化剂。与图1的流程相比，图2的流程增加了一个闪蒸精馏塔37，虽然增加了投资，但是减少了反应精馏塔12的总塔盘数，降 低了该塔的高度，减少了反应精馏塔12制作成本和难度。同时，由于进入反应精馏塔12的物料38温度较低，降低了操作难度，使整个塔的操作更简单平稳，降低了催化剂塔盘超温的可能性，同时也进一步降低了催化剂因超温脱酸的风险，有利于进一步延长催化剂的寿命。
实施例A1～A5：顺酐/甲醇摩尔配比探究
在2.5L的高压釜中，加入4mol(392g)的顺酐MAH，再加入6mol(192g)的甲醇，经3次5公斤氮气置换后，升高反应温度到110℃并维持搅拌，经间隔取样，并用气相色谱检测，反应8小时后，体系内的各组分组成不再变化，可证明反应达到了平衡状态。同样的反应条件下，反应再进性三次，不同的是顺酐和甲醇的摩尔比例分别从1：1.2、变化到了1：1.5、1：2、1：2.5和1：3，也就是加入的甲醇量分别为4.8mol(153.6g)、58mol(256g)、10mol(320g)和12mol(384g)，反应过程相同，最后都达到平衡状态，以上不同MAH/甲醇配比和反应的平衡组成(w％)见表1。
表1：不同顺酐/甲醇摩尔配比的反应物的平衡组成(w％)
实施例 A1 A2 A3 A4 A5
组分/摩尔配比 1/1.2 1/1.5 1/2.0 1/2.5 1/3.0
甲醇(％) 5.31 4.56 5.23 14.5 22.67
水(％) 2.03 6.8 7.57 7.8 6.33
顺酐(％) 12.49 6.57 4.01 0.30 0.55
马来酸单甲酯(％) 63.89 29.12 36.50 14.8 19.76
马来酸二甲酯(％) 16.27 54.4 56.63 62.4 50.69
其他(％) 0 0 0 0.20 0
从上表可见，顺酐和甲醇的摩尔比例为1：2.5时，马来酸二甲酯的转化率最佳。
实施例B1～B5：第一阶段酯化反应条件探究
第一步酯化反应在一个壳程为Φ108X4(SS304)，管程为7根Φ25X3(SS316L)列管式反应器中实施，管程中填装Φ5X5的陶瓷拉西环(堆密度1.2，孔隙率0.46)，壳程以固定温度的循环导热油为传热介质；顺酐融化并恒温在90℃，进料泵泵头和管线均用电伴热保温在110℃；甲醇保温在50℃，甲醇泵泵头和管线也保温在50℃；产品收集系统的两个产品收集罐压力以氮气背压的模式维持在1.0Mpa，使反应得以连续进行，进料的顺酐和甲醇的摩尔比例为1：2.5；反应分别在110℃和130℃的温度下进行；反应物在管壳式反应器的停留时间为1h，2h和4h，反应产物的组成一并列在表2中；
表2：顺酐/甲醇不同温度条件和时间下的反应物组成(w％)
实施 条件\组成 甲醇 水 顺酐 MMM DMM 其他
例 (w％) (％) (％) (％) (％) (％) (％)
B1 110℃-1h 19.74 5.55 0.39 29.75 44.41 0.16
B2 110℃-2h 14.9 6.73 0.36 23.93 53.82 0.26
B3 110℃-4h 15.62 7.74 0.33 16.60 59.74 0.24
B4 130℃-1h 13.77 7.26 0.96 19.62 58.12 0.27
B5 130℃-2h 11.52 7.71 0.97 17.86 61.64 0.30
B6 130℃-4h 12.28 8.01 0.84 14.50 64.06 0.20
实施例C1：马来酸二甲酯DMM的合成
(1)第一阶段酯化反应：同实施例B3；
(2)闪蒸精馏：把上个步骤中所得的110℃-4h条件下的反应混合物4000g进行闪蒸精馏；所用闪蒸精馏塔为中部连续进料，总长度为一米，内径为25mm、内部装有Φ5钛环丝网填料的玻璃精馏塔，闪蒸精馏塔底部配有2500ml的三口烧瓶，采用油浴加热；整个闪蒸精馏塔配有上部回流和采出、中部进料和底部出料口，进料和出料均用配有硅胶管的蠕动泵实行；整个闪蒸精馏塔为微正压操作，塔体温度为80～120℃，底部三口烧瓶内的温度控制在140℃。在蒸出反应混合物中的大部分水和甲醇后，最后得到主要成分为马来酸单甲酯和马来酸二甲酯的混合物3100g备用，其组成见表3；
表3：110℃-4h产物闪蒸精馏后产品的组成(w％)
组成(w％) 甲醇(％) 水(％) 顺酐(％) MMM(％) DMM(％) 其他(％)
110℃-4h精馏后 0.13 2.28 0.42 21.06 75.80 0.30
(3)第二阶段酯化反应：将上个步骤中所得的110℃-4h产物精馏后产品3100g送入反应精馏塔，该反应精馏塔为自制的、两段精馏柱，总长度1.5米，内径为25mm的玻璃精馏塔，釜底为2500ml的三口烧瓶，采用油浴加热；整个反应精馏塔配有上部回流和采出、中上部第二段进料和底部三口烧瓶出料口，进料和出料均用配有硅胶管的蠕动泵实行；使用的DZH酸性磺酸树脂先用甲醇浸泡置换4次，再用80目的316不锈钢丝网包裹紧密，制成约直径约25mm、高25mm的柱状包；玻璃精馏柱分量段，上段装填Φ5钛环丝网填料；DZH酸性磺酸树脂包填入下段精馏柱中，每个催化剂包中间装填Φ5mm的陶瓷拉西环25mm；精馏柱最下段100mm装填Φ5mm的陶瓷拉西环，替代没有装催化剂的塔板；两段精馏柱分别用电加热带加热保温控温，上、下温度分别80～100℃，110～130℃；3100g的马来酸单甲酯和马来酸二甲酯混合物从第二段精馏塔上部以3ml/min的速度滴加；下部用泵把甲醇经长玻璃滴管泵入此三口烧瓶底部，流速为1.2ml/min；三口烧瓶内温度维持在130-140℃之间，整个塔维持在微正压状态；通过实验中测试2500ml的三口烧瓶中的物料含量比例，再通过微调精馏柱的温度和上下进料 的流量速度，以保证马来酸单甲酯转化为马来酸二甲酯的反应完成99.9％以上；反应20小时后，最后得到3200g的反应物，其中马来酸二甲酯含量为97％；
(4)精制：3200g的上述反应混合液分两次，在装配有简单蒸镏柱的2500ml的三口烧瓶，用油浴加热；先把反应液加热到120℃并维持搅拌，蒸出轻组分；再提高温度到180-200℃，减压蒸馏出3000g的纯度>99.7％的产品马来酸二甲酯DMM。
实施例C2：马来酸二甲酯DMM的合成
(1)第一阶段酯化段反应：同实施例B5；
(2)第二阶段酯化反应：把上述步骤中所得的130℃-2h未经精馏产品4000g送入反应精馏塔；反应条件与实施例C1的步骤(3)基本相同，不同的地方有：该反应精馏塔分三段，第二段精馏柱不装填催化剂，装填Φ5mm的陶瓷拉西环100mm；4000g的甲醇、水、马来酸单甲酯、马来酸二甲酯混合物从第二段精馏塔上部以4ml/min的速度滴加，甲醇用泵经长玻璃滴管泵入此精馏塔底部的三口烧瓶底部，流速为1.2ml/min；经过反应20小时后，该实施例最后也得到3200g的反应物，其中马来酸二甲酯含量为96％；
(3)精制：与实施例C1的步骤(4)相同，经过脱轻脱重后，也得到了3000g的纯度>99.8％的产品马来酸二甲酯DMM。
实施例D：马来酸二丙酯DPM的合成
(1)第一阶段酯化反应：选用与实施例B同样的管壳式反应器，顺酐和正丙醇的进料摩尔比例为1：2.5，反应温度为140℃，产品收集系统的双产品收集罐压力以氮气背压的模式维持在0.8Mpa，物料在反应器中的保留时间为3h；最后得到产品5000g，组成见表4；
表4：140℃-3h DPM合成第一阶段酯化段反应产物的组成(w％)
组成(w％) 丙醇(％) 水(％) 顺酐(％) MPM(％) DPM(％) 其他(％)
140℃-3h产品 16.44 5.91 0.19 12.46 64.81 0.18
(2)闪蒸精馏：DPM合成第一阶段酯化段反应产物的闪蒸精馏同样在于实施例C1的精馏装置中进行，所不同的就是精馏所需的温度不同：塔体温度为90～140℃，底部三口烧瓶内的温度控制在155℃；在蒸出反应混合物中的大部分水和丙醇后，最后得到主要成分为马来酸单丙酯和马来酸二丙酯的混合物3950g备用，其组成见表5；
表5：140℃-3h DPM合成产物精馏后的组成(w％)
组成(w％) 丙醇(％) 水(％) 顺酐(％) MPM(％) DPM(％) 其他(％)
140℃-3h精馏后 0.31 2.39 0.24 15.66 81.48 0.23
(3)第二阶段酯化反应：把上述所得的110℃-4h产物精馏后产品3950g在实施例C1使用的反应精馏塔实现产品的第二步酯化；两段精馏柱的温度自上而下分别为 80～100℃，110～130℃，反应混合物从第二段精馏塔上部以5ml/min的速度滴加，丙醇泵进流速为2ml/min；三口烧瓶内温度维持在150～160℃之间，整个塔维持在微正压状态；通过实验中测试2500ml的三口烧瓶中的物料含量比例，再通过微调精馏柱的温度和上下进料的流量速度，以保证马来酸单丙酯转化为马来酸二丙酯的反应完成99.9％以上，反应20小时后，最后得到4100g的反应物，其中马来酸二丙酯含量约为98％。
(4)精制：与实施例C1的步骤(4)相同，经过脱轻脱重后，也得到了4000g的纯度>99.9％的产品马来酸二丙酯DPM。

Claims (9)

  1. 一种制备马来酸二醇酯的工艺方法，其特征在于包括有以下步骤：
    (1)将顺酐和低级脂肪醇在预酯化反应器中进行预酯化反应，得到相应水、醇、马来酸单醇酯、马来酸二醇酯以及少量顺酐的混合物；
    (2)将预酯化反应产生的反应混合物先进行闪蒸和精馏，再在酸性树脂的催化下和相应的低级脂肪醇通过反应精馏完成酯化，得到所需的马来酸二醇酯。
  2. 根据权利要求1所述的制备马来酸二醇酯的工艺方法，其特征在于：所述步骤(2)中采用带有闪蒸段的反应精馏塔进一步反应完成酯化，该反应精馏塔分三段，上段为甲醇和水的提馏段，温度控制分布在80～100℃之间；中段为反应混合物的闪蒸段，温度自上而下控制分布在220～100℃之间，下段为多个自上而下依次布置的塔盘，温度控制分布在100～130℃之间，至少6个塔盘上装填有大孔酸性树脂催化剂，在该反应精馏塔中预酯化反应产生的混合物自上而下流动，低级脂肪醇和进一步酯化反应产生的水以气体状态自下而上逆流流动。
  3. 根据权利要求1所述的制备马来酸二醇酯的工艺方法，其特征在于：所述步骤(2)中采用闪蒸精馏塔和反应精馏塔，进行闪蒸精馏后进一步进行反应精馏，该反应精馏塔分两段，下段为多个自上而下依次布置的塔盘，温度控制分布在100～130℃之间，至少6个塔盘上装填有大孔酸性树脂催化剂，在该反应精馏塔中，预酯化反应产生的混合物自上而下流动，低级脂肪醇和进一步酯化反应产生的水以气体状态自下而上逆流流动，所述闪蒸精馏塔的操作压力为0.1～0.5Mpa，温度控制分布在90～130℃之间。
  4. 根据权利要求2或3所述的制备马来酸二醇酯的工艺方法，其特征在于：所述塔盘中，至少两个位于底部的塔盘不装填有催化剂。
  5. 根据权利要求2或3所述的制备马来酸二醇酯的工艺方法，其特征在于：所述反应精馏塔的操作压力为0.1～0.5Mpa。
  6. 根据权利要求2或3所述的制备马来酸二醇酯的工艺方法，其特征在于：所述反应精馏塔塔底得到的反应物经过双塔精馏，实现脱轻脱重，得到纯度在99.5％以上的马来酸二醇酯。
  7. 根据权利要求2或3所述的制备马来酸二醇酯的工艺方法，其特征在于：所述的预酯化反应器为管壳式反应器，管壳式反应器的管程用于供预酯化反应进行，管壳式反应器的壳程用于供低级脂肪醇通过，管壳式反应器的管程中预酯化反应产生的混合物送入反应精馏塔上部，或经闪蒸精馏后再进入相应的反应精馏塔，管壳式反应器的壳程中的低级脂肪醇汽化移走预酯化反应所放热量后送入反应精馏塔下部。
  8. 根据权利要求7所述的制备马来酸二醇酯的工艺方法，其特征在于：所述管壳 式反应器的管程中顺酐和脂肪醇的摩尔比为1：1.5～4，反应温度为60～220℃，反应压力为0.1～1.6MPa，反应时间为0.1～8h；所述管壳式反应器的壳层中脂肪醇的温度为80～200℃，反应压力为0.3～1.6MPa，维持相应醇的该温度下的气液平衡汽化即可，相应量足以移走反应产生的热为准。
  9. 根据权利要求1所述的制备马来酸二醇酯的工艺方法，其特征在于：所述的低级脂肪醇为甲醇、乙醇、丙醇和丁醇中的至少一种。
PCT/CN2022/071265 2021-02-08 2022-01-11 一种制备马来酸二醇酯的工艺方法 WO2022166540A1 (zh)

Applications Claiming Priority (2)

Application Number Priority Date Filing Date Title
CN202110180196.7 2021-02-08
CN202110180196.7A CN112961058B (zh) 2021-02-08 2021-02-08 一种制备马来酸二醇酯的工艺方法

Publications (1)

Publication Number Publication Date
WO2022166540A1 true WO2022166540A1 (zh) 2022-08-11

Family

ID=76284628

Family Applications (1)

Application Number Title Priority Date Filing Date
PCT/CN2022/071265 WO2022166540A1 (zh) 2021-02-08 2022-01-11 一种制备马来酸二醇酯的工艺方法

Country Status (2)

Country Link
CN (1) CN112961058B (zh)
WO (1) WO2022166540A1 (zh)

Cited By (3)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN115322092A (zh) * 2022-09-14 2022-11-11 上海赛斯格恩化学技术有限公司 一种利用顺酐尾气制备马来酸二甲酯的方法
CN115518401A (zh) * 2022-10-26 2022-12-27 万华节能科技(烟台)有限公司 一种用于生产单组分聚脲所需潜固化剂的生产工艺
CN118718931A (zh) * 2024-08-27 2024-10-01 北京弗莱明科技有限公司 一种低能耗酯化合成乙酸乙酯的生产系统及生产方法

Families Citing this family (6)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN112961058B (zh) * 2021-02-08 2022-05-10 宁波窦氏化学科技有限公司 一种制备马来酸二醇酯的工艺方法
CN114100679A (zh) * 2021-12-03 2022-03-01 王义成 一种树脂催化剂及其制备方法和应用
CN114570052B (zh) * 2022-03-04 2023-04-11 德州绿霸精细化工有限公司 一种连续生产顺丁烯二酸二乙酯的方法及其装置
CN114984866B (zh) * 2022-06-02 2024-03-26 中化学科学技术研究有限公司 一种制备马来酸二甲酯的系统及方法
CN115745772A (zh) * 2022-09-30 2023-03-07 云南大为恒远化工有限公司 一种顺酐酯化加氢生产丁二酸的方法
CN115651175B (zh) * 2022-12-13 2023-04-04 中科国生(杭州)科技有限公司 一种含呋喃环的抗菌耐热高阻隔共聚酯及其制备方法

Citations (8)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN87105388A (zh) * 1986-08-01 1988-03-09 戴维麦基(伦敦)有限公司 马来酸二烷基酯的制备方法
CN102070448A (zh) * 2011-02-22 2011-05-25 湖南长岭石化科技开发有限公司 一种制备丁二酸二甲酯的方法
CN102911053A (zh) * 2011-08-01 2013-02-06 中国石油化工股份有限公司 催化剂分段装填方式合成马来酸二烷基酯的方法
CN103360252A (zh) * 2012-04-01 2013-10-23 中国石油化工股份有限公司 马来酸二甲酯的生产方法
WO2015082916A1 (en) * 2013-12-06 2015-06-11 Johnson Matthey Davy Technologies Limited Process for the preparation of succinic acid ester
WO2017064470A1 (en) * 2015-10-13 2017-04-20 Johnson Matthey Davy Technologies Limited Process for the co-production of dialkyl maleate and dialkyl succinate
CN107473966A (zh) * 2017-08-29 2017-12-15 南京雪郎化工科技有限公司 一种马来酸二甲酯的生产方法
CN112961058A (zh) * 2021-02-08 2021-06-15 宁波窦氏化学科技有限公司 一种制备马来酸二醇酯的工艺方法

Patent Citations (8)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN87105388A (zh) * 1986-08-01 1988-03-09 戴维麦基(伦敦)有限公司 马来酸二烷基酯的制备方法
CN102070448A (zh) * 2011-02-22 2011-05-25 湖南长岭石化科技开发有限公司 一种制备丁二酸二甲酯的方法
CN102911053A (zh) * 2011-08-01 2013-02-06 中国石油化工股份有限公司 催化剂分段装填方式合成马来酸二烷基酯的方法
CN103360252A (zh) * 2012-04-01 2013-10-23 中国石油化工股份有限公司 马来酸二甲酯的生产方法
WO2015082916A1 (en) * 2013-12-06 2015-06-11 Johnson Matthey Davy Technologies Limited Process for the preparation of succinic acid ester
WO2017064470A1 (en) * 2015-10-13 2017-04-20 Johnson Matthey Davy Technologies Limited Process for the co-production of dialkyl maleate and dialkyl succinate
CN107473966A (zh) * 2017-08-29 2017-12-15 南京雪郎化工科技有限公司 一种马来酸二甲酯的生产方法
CN112961058A (zh) * 2021-02-08 2021-06-15 宁波窦氏化学科技有限公司 一种制备马来酸二醇酯的工艺方法

Cited By (4)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN115322092A (zh) * 2022-09-14 2022-11-11 上海赛斯格恩化学技术有限公司 一种利用顺酐尾气制备马来酸二甲酯的方法
CN115322092B (zh) * 2022-09-14 2024-05-07 上海赛斯格恩化学技术有限公司 一种利用顺酐尾气制备马来酸二甲酯的方法
CN115518401A (zh) * 2022-10-26 2022-12-27 万华节能科技(烟台)有限公司 一种用于生产单组分聚脲所需潜固化剂的生产工艺
CN118718931A (zh) * 2024-08-27 2024-10-01 北京弗莱明科技有限公司 一种低能耗酯化合成乙酸乙酯的生产系统及生产方法

Also Published As

Publication number Publication date
CN112961058A (zh) 2021-06-15
CN112961058B (zh) 2022-05-10

Similar Documents

Publication Publication Date Title
WO2022166540A1 (zh) 一种制备马来酸二醇酯的工艺方法
KR101431488B1 (ko) 프로필렌글리콜로의 글리세롤의 수소첨가 방법
CN100537511C (zh) 连续酯化生产己二酸二甲酯的方法及设备
CN101357890B (zh) 采用热泵技术的碳酸二甲酯合成与精制工艺及其装置
US20220410104A1 (en) Gas-liquid bubbling bed reactor, reaction system, and process for synthesizing carbonate ester
CN104788294B (zh) 一种反应精馏合成乙二醇单丁醚的装置和工艺方法
CN101830807B (zh) 亚硝酸烷基酯的合成装置及工艺
CN105669379B (zh) 一种醋酸乙酯加氢制备乙醇的工艺
CN103965040B (zh) 一种制备二元酸二甲酯的方法
CN109748790B (zh) 生产己二酸二甲酯的方法
CN109748791B (zh) 生产己二酸二甲酯的节能方法
CN211871828U (zh) 一种煤制乙二醇副产甲酸甲酯加氢连续生产甲醇的系统
CN110613946B (zh) 一种合成草酸二乙酯的背包反应精馏装备及其工艺
CN113200853A (zh) 一种制备丁二酸二醇酯的工艺方法
CN101898930A (zh) 苯加氢生产环己烷的装置及合成工艺
CN110483282B (zh) 一种生产高浓度乙二醇双乙酸酯的装置和方法
CN106608865A (zh) 用于碳酸乙烯酯合成的节能方法
CN209338419U (zh) 一种环己烯节能生产系统
CN114984866B (zh) 一种制备马来酸二甲酯的系统及方法
TW201728565A (zh) 方法
CN109438167B (zh) 一种环己烯节能生产系统以及生产方法
CN109824469A (zh) 一种环己烯节能生产方法
CN214937125U (zh) 一种草酸酯生产乙醇的系统
CN104672065A (zh) 高浓度异丁烯采用混相床反应器生产mtbe的方法
CN111253213A (zh) 一种醋酸酯加氢制乙醇的工艺方法及系统

Legal Events

Date Code Title Description
121 Ep: the epo has been informed by wipo that ep was designated in this application

Ref document number: 22748819

Country of ref document: EP

Kind code of ref document: A1

NENP Non-entry into the national phase

Ref country code: DE

122 Ep: pct application non-entry in european phase

Ref document number: 22748819

Country of ref document: EP

Kind code of ref document: A1