WO2020029753A1 - 一种2,2-二甲基-1,3-丙二醇的生产工艺 - Google Patents

一种2,2-二甲基-1,3-丙二醇的生产工艺 Download PDF

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WO2020029753A1
WO2020029753A1 PCT/CN2019/095988 CN2019095988W WO2020029753A1 WO 2020029753 A1 WO2020029753 A1 WO 2020029753A1 CN 2019095988 W CN2019095988 W CN 2019095988W WO 2020029753 A1 WO2020029753 A1 WO 2020029753A1
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dimethyl
propanediol
producing
reaction
distillation
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PCT/CN2019/095988
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English (en)
French (fr)
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王振凡
李春伟
于海龙
胡伟光
黄珍妮
宋敬文
刘国新
顾天宇
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吉林市道特化工科技有限责任公司
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Priority to RU2020119135A priority Critical patent/RU2741574C1/ru
Publication of WO2020029753A1 publication Critical patent/WO2020029753A1/zh

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    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C45/00—Preparation of compounds having >C = O groups bound only to carbon or hydrogen atoms; Preparation of chelates of such compounds
    • C07C45/45—Preparation of compounds having >C = O groups bound only to carbon or hydrogen atoms; Preparation of chelates of such compounds by condensation
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C29/00—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring
    • C07C29/132—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring by reduction of an oxygen containing functional group
    • C07C29/136—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring by reduction of an oxygen containing functional group of >C=O containing groups, e.g. —COOH
    • C07C29/14—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring by reduction of an oxygen containing functional group of >C=O containing groups, e.g. —COOH of a —CHO group
    • C07C29/141—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring by reduction of an oxygen containing functional group of >C=O containing groups, e.g. —COOH of a —CHO group with hydrogen or hydrogen-containing gases
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C31/00—Saturated compounds having hydroxy or O-metal groups bound to acyclic carbon atoms
    • C07C31/18—Polyhydroxylic acyclic alcohols
    • C07C31/20—Dihydroxylic alcohols
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C47/00—Compounds having —CHO groups
    • C07C47/02—Saturated compounds having —CHO groups bound to acyclic carbon atoms or to hydrogen
    • C07C47/12—Saturated compounds having —CHO groups bound to acyclic carbon atoms or to hydrogen containing more than one —CHO group

Definitions

  • the invention relates to the field of fine chemical production, in particular to a production process of 2,2-dimethyl-1,3-propanediol.
  • 2,2-Dimethyl-1,3-propanediol is an important polyol organic chemical product and is widely used in functional materials such as resins and coatings.
  • the production of 2,2-dimethyl-1,3-propanediol can be achieved by using the organic tertiary amines such as trimethylamine and triethylamine to catalyze the condensation and hydrogenation of formaldehyde, isobutyraldehyde and other aldehyde raw materials. In the condensation reaction, it is inevitable to cause a disproportionation reaction of a small amount of aldehyde raw materials to generate by-products of organic acids and alcohols.
  • ester impurities such as methyl formate, isobutyl formate, and isobutyric acid Isobutyl ester, 2,2-dimethyl-1,3-propanediol mono (di) ester, 2,2-dimethyl-1,3-propanediol isobutyrate, 3-hydroxy-2,2- Dimethyl-3-hydroxy-2,2-dimethylpropylpropionate and the like.
  • ester impurities will follow the material and flow into the product refining unit. Some of these esters have a large difference between the boiling point of the target product and the boiling point of the target product or azeotropy with water.
  • esters can be easily removed from the crude product system, such as methyl formate, Butyl ester, isobutyl isobutyrate, 3-hydroxy-2,2-dimethyl-3-hydroxy-2,2-dimethylpropylpropionate.
  • Some esters can be hydrolyzed to alcohols which can be easily separated by the action of hydrogenation catalysts, such as 2,2-dimethyl-1,3-propanediol mono (di) ester.
  • Chinese patent CN101993351A discloses a method for eliminating esters using alkaline hydrolysis after hydrogenation, but the process is complicated and costly, and the sewage treatment burden is large.
  • Chinese patent CN107311840A discloses a method of first extracting after hydrogenation and then purifying by distillation under reduced pressure. However, since petroleum ether extraction is required first, solvent consumption is large, extraction efficiency is low, production efficiency is reduced, and secondary pollution is easy to occur.
  • the object of the present invention is to provide a production process of 2,2-dimethyl-1,3-propanediol, which can effectively control the occurrence of disproportionation side reactions in the condensation reaction process, and control the generation of ester impurities from the source. So as to effectively reduce the content of ester impurities in the product.
  • the invention provides a process for producing 2,2-dimethyl-1,3-propanediol, which comprises a 37% mass formaldehyde aqueous solution and a 99% mass isobutyraldehyde in a mass ratio of 1.05- Mix at 1.25: 1, adjust the pH of the reaction system to 6.0-8.5, and obtain the intermediate by condensation reaction under the conditions of temperature of 60-80 ° C and time of 0.5-2.0 hours. The intermediate is obtained after hydrogenation reaction. The crude 2,2-dimethyl-1,3-propanediol was then purified to obtain the final product of 2,2-dimethyl-1,3-propanediol.
  • an alkaline regulator is added to adjust the pH of the reaction system from 6.0 to 8.5, preferably, alkaline is added.
  • the pH value of the reaction system is adjusted by the regulator.
  • an acidity regulator and a basicity regulator are added to adjust the pH of the reaction system from 6.0 to 8.5, preferably .
  • the pH value of the reaction system is adjusted by adding an acidity regulator and a basicity regulator to 6.5-8.0.
  • the basicity regulator is a tertiary organic amine, and the organic tertiary amine is trimethylamine.
  • the acidity regulator is an organic acid, and the organic acid is any one or two of formic acid and isobutyric acid.
  • the reaction temperature is preferably 65-75 ° C, and more preferably 65-70 ° C.
  • the reaction time is preferably 0.8-1.5 hours, and more preferably 1.0-1.2 hours.
  • the method of the present invention is implemented in industrial production, and continuous production can be performed by using a plurality of condensation reactors in series.
  • the raw materials isobutyraldehyde and formaldehyde can be metered in the first-stage reactor, or they can be added separately in different amounts in the respective reactors.
  • the substances that adjust the pH of the reaction system can be added to the first-stage reactor, or can be added to the reactors of different stages in different quantities according to the pH of the reactors of each stage.
  • a pH detection device can be set on the reactor equipment or the pipeline connected to the equipment, and a substance that adjusts the pH of the reaction system can be added according to the pH control.
  • the condensation reaction is an exothermic reaction.
  • the heat released will at least increase the local temperature of the material. This local or small increase in temperature will lead to an increase in undesired side reactions.
  • the uniform temperature of the material system is also necessary.
  • forced heat exchange is required. Use forced circulation stirrer internal parts, static mixers, jet mixers and other enhanced mass transfer equipment for mixing, and use heat exchangers to remove heat.
  • the condensation product can be hydrogenated using a method similar to Chinese Patent CN103449970B to obtain a crude product containing the target product.
  • the hydrogenated crude product is sent to a distillation column to first remove the light components relative to the product components, and then remove the heavy components to obtain the target product.
  • the 2,2-dimethyl-1,3-propanediol formate was lower than 0.03% by mass after hydrogenation.
  • the hydrogenation reaction is an intermediate under the action of a hydrogenation catalyst at a temperature of 110-140 ° C and a pressure of 2.0-5.0MPa Under the conditions of catalytic hydrogenation, crude 2,2-dimethyl-1,3-propanediol was obtained.
  • the hydrogenated crude product is sent to a distillation column to separate light components, and then heavy components are removed to obtain a target product.
  • the boiling point of the ester is between alcohol and acid.
  • the ester in the material is raised to the distillation section of the distillation column.
  • the temperature of the bottom of the distillation column during the treatment process is preferably higher than The boiling point of the polyol under process pressure conditions. In order to maintain the temperature of the process, heating can be used.
  • the purification process is to add crude 2,2-dimethyl-1,3-propanediol to a distillation column at a temperature of 139. -155 °C, pressure is 1-50kPa under the conditions of one-time reduced pressure distillation for 0.5-8.0 hours, the light components are discharged, and then the heavy components are removed by secondary distillation to obtain the finished product of 2,2-dimethyl-1,3-propanediol .
  • the time of the one-time vacuum distillation is 1.0-4.0 hours.
  • the acidity of the material system in the tower kettle and the distillation section of the distillation tower is adjusted so that the pH value of the effluent from the tower kettle It is 4.0-7.0, preferably 4.5-6.5.
  • a catalyst is introduced to catalyze the ester impurities during the purification process.
  • the catalyst is an acid catalyst, and preferably, the catalyst is an organic acid, an inorganic acid, or a solid acid catalyst.
  • the catalyst is added in an amount to ensure the pH of the tower kettle system is 4.0-7.0.
  • the solid acid catalyst is a molecular sieve catalyst having a silicon-alumina ratio of not less than 200.
  • ester impurities in the materials in the distillation section and distillation section of the distillation column react, and the product is separated and removed with other impurities.
  • the amount of ester impurities in the liquid phase material will be correspondingly removed to an acceptable level, and the purity obtained by purification will be acceptable.
  • the product in order to maintain the temperature of the treatment process, heating can be used.
  • At least the residence time of the stream should not be less than the time for the corresponding amount of ester impurities to reach the chemical reaction equilibrium at the temperature of the processing process, and control the material in the distillation column.
  • the residence time is from 0.5 to 8.0 hours, preferably from 1.0 to 4.0 hours.
  • the ester impurities need to be sufficiently raised to the reaction in the packing, and an excessive reflux ratio should not be used, at least the reflux ratio of the overhead stream should be guaranteed to be 0.5- 2.0, preferably a reflux ratio of 0.5-1.0.
  • the reflux ratio refers to the ratio of the amount of liquid material condensed on the top of the tower returned to the inside and outside the tower.
  • the present invention can implement the distillation technology in the following specific ways.
  • the equipment of the process unit that can be used is one or a combination of structural forms such as tanks, towers, and evaporators.
  • the equipment can use internal components such as fillers and heating pipes
  • the filler can be made of regular or random metal or non-metallic materials.
  • the material of the equipment is stainless steel with corrosion resistance.
  • the equipment of the process unit can be one or several units in series. After the components having a lower relative volatility than the target product are separated in the distillation column according to the present invention, the product of the tower kettle is sent to another distillation column to remove the heavy components and obtain a final pure product.
  • the raw materials are continuously fed into and discharged from the process unit.
  • the tower is used for operation and treatment. Liquid materials can be sent into the equipment from the middle of the tower. Internal components such as distributors or spray nozzles can be used to enhance the uniformity of liquid distribution in the tower packing. After the gaseous material leaving the tower from the upper part is condensed by the condenser, part of it is returned to the tower, and the other part is sent out of the equipment unit. Part of the liquid material leaving the tower from the lower part is heated and boiled by the heater, part of the liquid material is vaporized and returned to the tower, and the other part is sent out of the equipment unit. Heaters, condensers and enhanced insulation methods are used to maintain the temperature of the process unit system. Each component in the material undergoes mass and heat transfer in the tower based on the "gas-liquid equilibrium” principle, and is based on the "chemical reaction equilibrium” principle in Chemical reactions occurred in the tower.
  • a forced circulation pump can be used to return the lower part of the tank to the upper part of the equipment and send it back into the equipment again to enhance the processing effect.
  • distilled materials such materials usually contain a sufficient amount of water. If the amount of water is insufficient, add the required water evenly before the raw materials are sent to the process unit.
  • the required amount of liquid catalyst can be added uniformly before the raw materials are sent to the process unit, or the liquid catalyst is separately added elsewhere in the equipment.
  • the solid catalyst needs to be packed into the equipment in advance. The bed.
  • ester reaction products are separated according to the relative volatility of each product.
  • the production process of 2,2-dimethyl-1,3-propanediol provided by the present invention is controlled by controlling the mass ratio of 37% formaldehyde aqueous solution and 99% isobutyraldehyde in a mass ratio of 1.05: 1 and Between 1.25: 1, the utilization of isobutyraldehyde condensation to produce the target product is fully improved, which not only ensures that the main reaction of the condensation can be carried out as far as possible to the chemical equilibrium, but also does not excessively form isobutyraldehyde ester, which can improve the target product 2
  • the purity of 2,2-dimethyl-1,3-propanediol, reducing the content of ester impurities such as 2,2-dimethyl-1,3-propanediol isobutyrate in its products, and by controlling the pH of the reaction system The value is between 6.0 and 8.5, which further promotes the rapid progress of the main reaction of the condensation reaction while minimizing the occurrence of side reactions such as disproportionation
  • the pH of the tower effluent is adjusted by adjusting the acidity of the material in the tower kettle and the distillation section of the distillation tower.
  • the value is 4.0-7.0, which makes the ester impurities accelerate the decomposition at the temperature of 139-155 ° C and the pressure of 1-50kPa, not only does not need petroleum ether extraction, reduces the purification cost, enhances the purification effect, but also reduces the pressure and high safety.
  • the addition of an acidic catalyst can further promote the decomposition of ester impurities, and the products of the ester impurities are separated and removed with other impurities, thereby reducing the content of ester impurities in the target product.
  • the condensation reaction in the following examples of the present application uses a condensation reaction system composed of three series condensation reactors for continuous production, and is equipped with a heat exchanger, a circulation pump, a stirrer internals, a jet mixer, and a temperature measurement thermoelectric Couple, pH meter and other devices.
  • the raw materials isobutyraldehyde and formaldehyde are metered into the condensation reaction system in the first-stage reactor.
  • the purification process in the embodiment of the present application is performed in a distillation column under reduced pressure, and the lower part of the distillation column uses a heater to heat the liquid phase. The material is boiled.
  • a condenser is used at the top of the distillation column to condense the gas phase material. The part of the condensed material at the top is returned to the column as reflux, and part of it is collected outside the product delivery equipment.
  • Devices such as trays or packings and distributors are installed in the distillation column.
  • step (2) The condensation reaction product obtained in step (1) is subjected to a hydrogenation reaction by recovering unreacted raw materials through an evaporator.
  • a fixed-bed reactor equipped with a nickel catalyst is used as a catalytic hydrogenation reactor.
  • the liquid space velocity of the condensation product was controlled to be 0.1h -1 and the gas space velocity of H 2 was 300h -1 .
  • the reaction was performed at a temperature of 110 ° C and a pressure of 2.0MPa to obtain 2,2-dimethyl-1,3. -Crude propylene glycol.
  • step (3) The crude 2,2-dimethyl-1,3-propanediol obtained in step (2) was purified by distillation under reduced pressure at a pressure of 10 kPa and a tower temperature of 142 ° C for 0.5 hours to remove light components.
  • the column effluent is sent to another distillation column to continue the distillation to obtain the finished product of 2,2-dimethyl-1,3-propanediol.
  • the test column effluent uses 20% of the relative mass of the material.
  • an appropriate amount of 5% phosphoric acid aqueous solution is added as needed to adjust the pH value of the distillation tower tower kettle effluent to 6.0.
  • the reflux ratio of the distillation tower is 0.5.
  • the mass content of the obtained 2,2-dimethyl-1,3-propanediol was 99.5%.
  • step (2) The condensation reaction product obtained in step (1) is subjected to a hydrogenation reaction by recovering unreacted raw materials through an evaporator.
  • a fixed-bed reactor equipped with a nickel catalyst is used as a catalytic hydrogenation reactor.
  • the liquid space velocity of the condensation product is controlled to be 0.1h -1
  • the gas space velocity of H 2 is 300h -1
  • the reaction is performed at a temperature of 110 ° C. and a pressure of 2.0 MPa to obtain 2,2-dimethyl-1,3. -Crude propylene glycol.
  • step (3) The crude 2,2-dimethyl-1,3-propanediol obtained in step (2) was purified by distillation under reduced pressure at a pressure of 10 kPa and a tower temperature of 145 ° C for 1.0 hour to remove light components.
  • the column effluent is sent to another distillation column to continue the distillation to obtain the finished product of 2,2-dimethyl-1,3-propanediol.
  • the test column effluent uses 20% of the relative mass of the material.
  • the mass content of the obtained 2,2-dimethyl-1,3-propanediol was 99.5%.
  • step (2) The condensation reaction product obtained in step (1) is subjected to a hydrogenation reaction by recovering unreacted raw materials through an evaporator.
  • a fixed-bed reactor equipped with a nickel catalyst is used as a catalytic hydrogenation reactor.
  • the liquid space velocity of the condensation product was controlled to be 0.1h -1 and the gas space velocity of H 2 was 300h -1 .
  • the reaction was performed at a temperature of 110 ° C and a pressure of 2.0MPa to obtain 2,2-dimethyl-1,3. -Crude propylene glycol.
  • step (3) The crude 2,2-dimethyl-1,3-propanediol obtained in step (2) was purified by distillation under reduced pressure at a pressure of 10 kPa and a tower temperature of 150 ° C for 2.3 hours to remove light components.
  • the column effluent is sent to another distillation column to continue the distillation to obtain the finished product of 2,2-dimethyl-1,3-propanediol.
  • the test column effluent uses 20% of the relative mass of the material.
  • the mass content of the obtained 2,2-dimethyl-1,3-propanediol was 99.5%.
  • step (2) The condensation reaction product obtained in step (1) is subjected to a hydrogenation reaction by recovering unreacted raw materials through an evaporator.
  • a fixed-bed reactor equipped with a nickel catalyst is used as a catalytic hydrogenation reactor.
  • the liquid space velocity of the condensation product was controlled to be 0.1h -1 and the gas space velocity of H 2 was 300h -1 .
  • the reaction was performed at a temperature of 110 ° C and a pressure of 2.0MPa to obtain 2,2-dimethyl-1,3. -Crude propylene glycol.
  • step (3) The crude 2,2-dimethyl-1,3-propanediol obtained in step (2) was purified by distillation under reduced pressure for 3.0 hours at a pressure of 10 kPa absolute pressure and a tower temperature of 150 ° C to remove light components.
  • the column effluent is sent to another distillation column to continue the distillation to obtain the finished product of 2,2-dimethyl-1,3-propanediol.
  • the test column effluent uses 20% of the relative mass of the material.
  • the mass content of the obtained 2,2-dimethyl-1,3-propanediol was 99.3%.
  • step (2) The condensation reaction product obtained in step (1) is subjected to a hydrogenation reaction by recovering unreacted raw materials through an evaporator.
  • a fixed-bed reactor equipped with a nickel catalyst is used as a catalytic hydrogenation reactor.
  • the liquid space velocity of the condensation product was controlled to be 0.1h -1 and the gas space velocity of H 2 was 300h -1 .
  • the reaction was performed at a temperature of 110 ° C and a pressure of 2.0MPa to obtain 2,2-dimethyl-1,3. -Crude propylene glycol.
  • step (3) The crude 2,2-dimethyl-1,3-propanediol obtained in step (2) was purified by distillation under reduced pressure at a pressure of 10 kPa and a tower temperature of 150 ° C for 4.0 hours to remove the light components.
  • the column effluent is sent to another distillation column to continue the distillation to obtain the finished product of 2,2-dimethyl-1,3-propanediol.
  • the test column effluent uses 20% of the relative mass of the material.
  • the mass content of the finished 2,2-dimethyl-1,3-propanediol was 99.4%.
  • step (2) The condensation reaction product obtained in step (1) is subjected to a hydrogenation reaction by recovering unreacted raw materials through an evaporator.
  • a fixed-bed reactor equipped with a nickel catalyst is used as a catalytic hydrogenation reactor.
  • the liquid space velocity of the condensation product was controlled to be 0.1h -1 and the gas space velocity of H 2 was 300h -1 .
  • the reaction was performed at a temperature of 110 ° C and a pressure of 2.0MPa to obtain 2,2-dimethyl-1,3. -Crude propylene glycol.
  • step (3) The crude 2,2-dimethyl-1,3-propanediol obtained in step (2) was purified by distillation under reduced pressure at a pressure of 10 kPa and a tower temperature of 155 ° C for 1.0 hour to remove light components.
  • the column effluent is sent to another distillation column to continue the distillation to obtain the finished product of 2,2-dimethyl-1,3-propanediol.
  • a hydrogen-type ZSM-5 molecular sieve with a silicon-aluminum ratio of 220 is filled in the packing layer of the distillation column.
  • Catalyst while depressurizing distillation, measure the pH value of the tower effluent after dissolving the material at room temperature using 20% relative mass of deionized water, and add an appropriate amount of 5% phosphoric acid aqueous solution as needed to adjust the pH of the distillation tower effluent The value was 5.5, and the reflux ratio of the distillation column was adjusted to 0.8.
  • the mass content of the finished 2,2-dimethyl-1,3-propanediol was 99.4%.
  • step (2) The condensation reaction product obtained in step (1) is subjected to a hydrogenation reaction by recovering unreacted raw materials through an evaporator.
  • a fixed-bed reactor equipped with a nickel catalyst is used as a catalytic hydrogenation reactor.
  • the liquid space velocity of the condensation product is controlled to be 0.1h -1
  • the gas space velocity of H 2 is 300h -1
  • the reaction is performed at a temperature of 110 ° C. and a pressure of 2.0 MPa to obtain 2,2-dimethyl-1,3. -Crude propylene glycol.
  • step (3) The crude 2,2-dimethyl-1,3-propanediol obtained in step (2) was purified by distillation under reduced pressure at a pressure of 10 kPa and a tower temperature of 142 ° C for 0.5 hours to remove light components.
  • the column effluent is sent to another distillation column to continue the distillation to obtain the finished product of 2,2-dimethyl-1,3-propanediol.
  • the test column effluent uses 20% of the relative mass of the material. After deionized water dissolves the material at room temperature, an appropriate amount of 5% phosphoric acid aqueous solution is added as needed to adjust the pH value of the distillation column tower effluent to 3.0.
  • the reflux ratio of the distillation column is 0.5.
  • the mass content of the 2,2-dimethyl-1,3-propanediol product obtained after testing was 99.1%.
  • the gas chromatography hydrogen flame method was used to analyze the 2,2-dimethyl-1,3-propanediol products obtained in Examples 1-6 and Comparative Examples 1-2.
  • 2,2-dimethyl-1 The average contents of 3-propanediol and 2,2-dimethyl-1,3-propanediol formate and 2,2-dimethyl-1,3-propanediol isobutyrate are shown in the following table.
  • step (2) The condensation reaction product obtained in step (1) is subjected to a hydrogenation reaction by recovering unreacted raw materials through an evaporator.
  • a fixed-bed reactor equipped with a nickel catalyst is used as a catalytic hydrogenation reactor.
  • the liquid space velocity of the condensation product was controlled to be 0.1h -1 and the gas space velocity of H 2 was 300h -1 .
  • the reaction was performed at a temperature of 110 ° C and a pressure of 2.0MPa to obtain 2,2-dimethyl-1,3. -Crude propylene glycol.
  • step (3) The crude 2,2-dimethyl-1,3-propanediol obtained in step (2) was purified by distillation under reduced pressure for 0.3 hours at a pressure of 10 kPa absolute pressure and a tower temperature of 139 ° C to remove light components.
  • the column effluent is sent to another distillation column to continue the distillation to obtain the finished product of 2,2-dimethyl-1,3-propanediol.
  • a hydrogen-type ZSM-5 molecular sieve with a silicon-aluminum ratio of 220 is filled in the packing layer of the distillation column.
  • Catalyst while depressurizing distillation, measure the pH value of the tower effluent after dissolving the material at room temperature using 20% relative mass of deionized water, and add an appropriate amount of 5% phosphoric acid aqueous solution as needed to adjust the pH of the distillation tower effluent The value was 7.8, and the reflux ratio of the distillation column was adjusted to 0.4.
  • the mass content of the obtained 2,2-dimethyl-1,3-propanediol was 99.2%.
  • step (2) The condensation reaction product obtained in step (1) is subjected to a hydrogenation reaction by recovering unreacted raw materials through an evaporator.
  • a fixed-bed reactor equipped with a nickel catalyst is used as a catalytic hydrogenation reactor.
  • the liquid space velocity of the condensation product was controlled to be 0.1h -1 and the gas space velocity of H 2 was 300h -1 .
  • the reaction was performed at a temperature of 110 ° C and a pressure of 2.0MPa to obtain 2,2-dimethyl-1,3. -Crude propylene glycol.
  • step (3) The crude 2,2-dimethyl-1,3-propanediol obtained in step (2) was purified by distillation under reduced pressure for 2.0 hours under a pressure of 10 kPa absolute pressure and a tower temperature of 145 ° C to remove light components.
  • the column effluent is sent to another distillation column to continue the distillation to obtain the finished product of 2,2-dimethyl-1,3-propanediol.
  • a hydrogen-type ZSM-5 molecular sieve with a silicon-aluminum ratio of 220 is filled in the packing layer of the distillation column.
  • Catalyst while depressurizing distillation, measure the pH value of the tower effluent after dissolving the material at room temperature using 20% relative mass of deionized water, and add an appropriate amount of 5% phosphoric acid aqueous solution as needed to adjust the pH of the distillation tower effluent
  • the value was 6.0, and the reflux ratio of the distillation column was adjusted to 2.5.
  • the mass content of the obtained 2,2-dimethyl-1,3-propanediol was 98.9%.

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

Abstract

本发明提供了一种2,2-二甲基-1,3-丙二醇的生产工艺，包括，将质量含量为37％的甲醛水溶液和质量含量99％的异丁醛按质量比为1.05-1.25:1混合，调节反应体系的pH值为6.0-8.5，在反应温度为60-80℃，反应时间为0.5-2.0小时的条件下经缩合反应得到中间体，所述中间体经加氢反应后得到2,2-二甲基-1,3-丙二醇粗品，然后经提纯得到2,2-二甲基-1,3-丙二醇成品，能够有效控制缩合反应过程中歧化副反应的发生，提高最终产品的收率，有效降低难降解的酯类杂质的生成，提高最终产品的质量，从而有利于下游产品性能的提高。

Description

一种2,2-二甲基-1,3-丙二醇的生产工艺 技术领域
本发明涉及精细化工生产领域，具体涉及2,2-二甲基-1,3-丙二醇的生产工艺。
背景技术
2,2-二甲基-1,3-丙二醇是重要的多元醇类有机化工产品，广泛地应用于树脂、涂料等功能性材料领域。2,2-二甲基-1,3-丙二醇的生产可以使用三甲胺、三乙胺等有机叔胺催化甲醛、异丁醛等醛类原料缩合加氢的工艺技术实现。在缩合反应中，不可避免地导致少量醛类原料的歧化反应，生成副产物有机酸和醇，有机酸与醇会进一步反应生成酯类杂质，如甲酸甲酯、甲酸异丁酯、异丁酸异丁酯、甲酸2,2-二甲基-1,3-丙二醇单(双)酯、异丁酸2,2-二甲基-1,3-丙二醇酯、3-羟基-2,2-二甲基-3-羟基-2,2-二甲基丙基丙酸酯等。以上这些酯类杂质会跟随物料流入产物精制单元，其中有些酯的沸点与目标产物的沸点差距较大或者与水共沸，很容易地由粗产物体系中分离除去，如甲酸甲酯、甲酸异丁酯、异丁酸异丁酯、3-羟基-2,2-二甲基-3-羟基-2,2-二甲基丙基丙酸酯。有些酯可以通过加氢催化剂的作用而氢解为容易分离的醇，如甲酸2,2-二甲基-1,3-丙二醇单(双)酯。但是有些酯类杂质与目标产物的相对挥发度不大，沸点接近2,2-二甲基-1,3-丙二醇，如甲酸2,2-二甲基-1,3-丙二醇单(双)酯、异丁酸2,2-二甲基-1,3-丙二醇酯，尤其是异丁酸2,2-二甲基-1,3-丙二醇酯不易发生氢解，因此不能容易地从2,2- 二甲基-1,3-丙二醇酯粗产品中分离除去，最终造成异丁酸2,2-二甲基-1,3-丙二醇酯成为产品中的含量最多的主要关键杂质，杂质的存在不仅会影响到最终产品的纯度和收率，还会因生成酯类，而消耗造成生产目标产品多元醇原料消耗的上升，甚至影响下游产品的性能。
中国专利CN101993351A公开了可以在加氢后使用碱水解消除酯的方法，但其过程是复杂和高成本的，且污水处理负担较大。中国专利CN107311840A公开了加氢后先萃取后减压蒸馏提纯的方法，但是由于需要先经过石油醚萃取，溶剂消耗大、萃取效率低，降低生产效益，易产生二次污染。
中国专利CN103449970B公开了酯类物质可以通过加氢分解为醇，但是由于加氢反应进行的难易程度不同，物质在催化剂孔道内传递扩散性能的不同，甲酸酯类产物会相对更彻底地加氢氢解，因而对于异丁酸2,2-二甲基-1,3-丙二醇酯、3-羟基-2,2-二甲基-3-羟基-2,2-二甲基丙基丙酸酯等酯类物质的降解分离作用有限，造成成品中含有较多的酯类杂质，影响产品质量和性能。
发明内容
因此，本发明的目的在于提供一种2,2-二甲基-1,3-丙二醇的生产工艺，能够有效控制缩合反应过程中歧化副反应的发生，从源头上控制酯类杂质的生成，从而有效降低产品中酯类杂质的含量。
本发明提供了一种2,2-二甲基-1,3-丙二醇的生产工艺，包括，将质量含量为37％的甲醛水溶液和质量含量为99％的异丁醛按质量比为1.05-1.25:1混合，调节反应体系的pH值为6.0-8.5，在温度为60-80℃，时间为0.5-2.0 小时的条件下经缩合反应得到中间体，所述中间体经加氢反应后得到2,2-二甲基-1,3-丙二醇粗品，然后经提纯得到2,2-二甲基-1,3-丙二醇成品。
本发明的2,2-二甲基-1,3-丙二醇的生产工艺中，所述缩合反应过程中，加入碱性调节剂调节反应体系的pH值为6.0-8.5，优选地，加入碱性调节剂调节反应体系的pH值为6.5-8.0。
本发明的2,2-二甲基-1,3-丙二醇的生产工艺中，所述缩合反应过程中，加入酸性调节剂和碱性调节剂调节反应体系的pH值为6.0-8.5，优选地，加入酸性调节剂和碱性调节剂调节反应体系的pH值为6.5-8.0。
本发明的2,2-二甲基-1,3-丙二醇的生产工艺中，所述碱性调节剂为有机叔胺，所述有机叔胺为三甲胺。
本发明的2,2-二甲基-1,3-丙二醇的生产工艺中，所述酸性调节剂为有机酸，所述有机酸为甲酸和异丁酸中的任意一种或者两种。
本发明的2,2-二甲基-1,3-丙二醇的生产工艺中，在所述缩合反应过程中，反应温度优选为65-75℃，更优选为65-70℃。
本发明的2,2-二甲基-1,3-丙二醇的生产工艺中，在所述缩合反应过程中，反应时间优选为0.8-1.5小时，更优选为1.0-1.2小时。
本发明的2,2-二甲基-1,3-丙二醇的生产工艺中，在所述缩合反应过程中，质量含量为37％的甲醛水溶液和质量含量为99％的异丁醛的质量比为1.05-1.15:1。
本发明的方法在工业化生产中实施可以使用多个缩合反应器串联进行连续化生产。原料异丁醛、甲醛可以在第一级反应器中计量加入，也可以 以不同数量在各级反应器中分别加入。
调节反应体系的酸碱度的物质，可以在第一级反应器中加入，也可以根据各级反应器的酸碱度情况，以不同数量在各级反应器中分别加入。在反应器的设备上或与设备相连接的管道上可以设置pH值检测设备，依据pH值控制加入调节反应体系的酸碱度的物质。
缩合反应是放热反应，放出的热量至少会使物料局部温度升高，这种温度的局部或小幅度升高会导致不应有的副反应增多，物料体系温度均一也是必要的，所以在工业生产过程中需强制换热，使用强制循环搅拌器内件、静态混合器、喷射混合器等强化传质设备进行混合，并使用换热器移除热量。
缩合产物在回收未反应的原料后，可以使用与中国专利CN103449970B相似的方法进行加氢获得含有目标产品的粗产物。加氢的粗产物送入蒸馏塔首先蒸馏除去相对于产品组分而言的轻组分，然后再除去重组分，获得目标产品。其中甲酸2,2-二甲基-1,3-丙二醇酯在加氢过程后均低于0.03％的质量含量。
本发明的2,2-二甲基-1,3-丙二醇的生产工艺中，所述加氢反应为中间体在加氢催化剂作用下，在温度为110-140℃，压力为2.0-5.0MPa条件下经催化氢化得到2,2-二甲基-1,3-丙二醇粗品。
本发明中，将加氢的粗产物送入蒸馏塔中分离轻组分，然后再除去重组分，获得目标产品。一般酯的沸点介于醇和酸之间，依据气液平衡原理，通过多级的蒸馏过程，使物料中部分的酯上升至蒸馏塔的提馏段内，处理 过程的蒸馏塔底部温度优选高于工艺压力条件下的多元醇的沸点。为了保持处理过程的温度，可以使用加热的方法实现。
本发明的2,2-二甲基-1,3-丙二醇的生产工艺中，所述提纯过程为将2,2-二甲基-1,3-丙二醇粗品加入蒸馏塔中，在温度为139-155℃，压力为1-50kPa条件下经一次减压蒸馏0.5-8.0小时，排出轻组分，然后经二次蒸馏除去重组分，得到2,2-二甲基-1,3-丙二醇成品。
本发明的2,2-二甲基-1,3-丙二醇的生产工艺中，所述提纯过程中，所述一次减压蒸馏时间为1.0-4.0小时。
本发明的2,2-二甲基-1,3-丙二醇的生产工艺中，所述提纯过程中，调节蒸馏塔内塔釜和提馏段物料体系的酸度，使塔釜流出物的pH值为4.0-7.0，优选为4.5-6.5。
本发明的2,2-二甲基-1,3-丙二醇的生产工艺中，所述提纯过程中，酯类杂质反应时不另外引入催化剂催化。
本发明的2,2-二甲基-1,3-丙二醇的生产工艺中，所述提纯过程中，酯类杂质反应时引入催化剂催化。
本发明的2,2-二甲基-1,3-丙二醇的生产工艺中，所述提纯过程中，催化剂为酸性催化剂，优选地，所述催化剂为有机酸、无机酸或者固体酸催化剂中的一种或几种的混合物，所述催化剂的添加量以保证塔釜体系的pH值为4.0-7.0。
本发明的2,2-二甲基-1,3-丙二醇的生产工艺中，所述固体酸催化剂为硅铝比不低于200的分子筛催化剂。
蒸馏塔提馏段、精馏段物料中的部分酯类杂质发生反应，产物随同其它杂质分离除去，液相物料中酯类杂质的数量会相应的去除到可以接受的程度，提纯得到纯度可以接受的产品。在本发明中，为了保持处理过程的温度，可以使用加热的方法实现。
在本发明中，为了保证酯类杂质除去效果，至少保证物流的停留时间不低于在处理过程的温度下相应处理数量的酯类杂质反应达到化学反应平衡的时间，控制物料在蒸馏塔内的停留时间为0.5-8.0小时，优选地停留时间为1.0-4.0小时。
在本发明中，为了使产物中的酯类杂质更多地除去，酯类杂质需充分地上升至填料中反应，不应使用过大的回流比，至少保证塔顶物流的回流比在0.5-2.0，优选回流比在0.5-1.0。回流比是指送回塔内和送出塔外的塔顶冷凝后的液体物料的数量的比值。
本发明可以按下面的具体方式实施蒸馏技术，可以采用的工艺单元的设备是罐、塔、蒸发器等结构形式中的一种或几种的组合，设备中可以使用填料和加热管等内件，填料可以是规整或散堆的金属或非金属材料制造。设备的材质选用具有抗腐蚀性的不锈钢材料。
工艺单元的设备可以是一台或几台串联组合。在本发明涉及的蒸馏塔中分离出比目标产品相对挥发度低的组分后，塔釜产物送入另一蒸馏塔脱除重组分并获得最终纯净的产品。
物料原料连续地送入和排出工艺单元。使用塔进行操作处理，液体物料可以由塔的中部送入设备内，可以使用分布器或喷雾喷头等内件强化液体在塔内填料中的分布均匀性。由上部离开塔的气体物料经冷凝器冷凝后， 部分回流回塔内，另一部分送出至设备单元外。由下部离开塔的液体物料的一部分经加热器加热沸腾部分气化后返回塔内，另一部分送出至设备单元外。使用加热器、冷凝器及强化保温的方法维持物工艺单元体系的温度，物料中的各组分依据“气液平衡”原理在塔内发生传质和传热，依据“化学反应平衡”原理在塔内发生化学反应。
如果需要，可以使用强制循环泵使罐的下部物料重新送回设备的上部，再次送入设备内，强化处理效果。
对蒸馏的物料，通常此类物料都含有足够数量的水，如果水的数量不足，在原料送入工艺单元前均匀地加入需要的水。
如果需要催化剂强化反应效果，可以在原料送入工艺单元前均匀地加入需要量的液体催化剂，或者在设备中其他位置另行加入液体催化剂，固体催化剂需要事先装填到设备内，优先安装固定在填料的床层中。
在设备的填料中，酯反应产物依据各产物相对挥发度的不同而分离。
本发明技术方案，具有如下优点：
1、本发明提供的2,2-二甲基-1,3-丙二醇的生产工艺，通过控制质量含量为37％的甲醛水溶液和质量含量为99％的异丁醛质量比在1.05：1与1.25:1之间，使异丁醛缩合生成目标产物的利用率充分地提高，不仅保证缩合主反应能够尽量进行到化学平衡，而且不会过多生成异丁醛酯，因而能够提高目标产物2,2-二甲基-1,3-丙二醇的纯度，减少其产物中异丁酸2,2-二甲基-1,3-丙二醇酯等酯类杂质的含量，同时通过控制反应体系的pH值在6.0-8.5之间，进一步促进缩合反应主反应正向快速进行的同时尽量减少歧化副反应和酯化反应等副反应的发生或者减缓这些副反应的速度，从而使 因歧化反应产生的有机酸尽量减少，从源头上控制酯的生成，并避免副反应产生的原料额外损耗，提高产品的纯度和收率，加快反应效率，缩短反应时间，而且有利于产物的纯化精制，降低了对后续提纯工艺的要求，无需石油醚萃取，提高生产的经济效益。
2、本发明2,2-二甲基-1,3-丙二醇的生产工艺中，在提纯过程中，通过调节蒸馏塔内塔釜和提馏段物料体系的酸度，使塔釜流出物的pH值为4.0-7.0，使得酯类杂质在温度为139-155℃，压力为1-50kPa条件下加快分解，不仅无需石油醚萃取，降低提纯成本，增强提纯效果，而且压力降低，安全性高，其中，通过加入酸性催化剂能够进一步促进酯类杂质的分解，酯类杂质分解后产物随同其它杂质分离除去，从而降低目标产品中酯类杂质的含量。
具体实施方式
本申请下述实施例的缩合反应采用三台串联的缩合反应器组成的缩合反应系统中进行连续化生产，并装有换热器、循环泵、搅拌器内件、喷射混合器、测温热电偶、pH值计等器件。为便于对比，原料异丁醛、甲醛均在第一级反应器中计量加入缩合反应系统，本申请实施例的提纯过程在减压条件下蒸馏塔中进行，蒸馏塔下部使用加热器加热液相物料沸腾，蒸馏塔顶部使用冷凝器冷凝气相物料，顶部冷凝的物料部分作为回流返回塔内，部分作为产品送出设备以外收集，在蒸馏塔中安装有塔板或填料及分布器等器件。
实施例1
(1)调节并维持三台缩合反应器组成的缩合反应系统的温度为60℃，向缩合反应系统中第一台反应器中连续地以1000千克/小时的速度送入质量含量为99％的异丁醛，同时向反应器中连续地以1250千克/小时的速度送入质量含量为37％的甲醛水溶液。根据pH值计的测量数据，向反应器中连续地送入质量含量为30％的三甲胺水溶液调节第三台反应器出口物料的pH值为8.5。调节反应器内的液位和出料速度，以使物料的总保留时间保持在2.0小时，得到缩合反应产物。
(2)将步骤(1)得到的缩合反应产物通过蒸发器回收未反应的原料后进行加氢反应，采用装有镍催化剂的固定床反应器为催化加氢反应器，在H 2存在下，控制缩合产物的液体空速为0.1h -1，H 2的气体空速为300h -1，在温度为110℃，压力为2.0MPa条件下反应制得2,2-二甲基-1,3-丙二醇粗品。
(3)将步骤(2)得到的2,2-二甲基-1,3-丙二醇粗品在压力为绝对压力10kPa，塔釜温度为142℃条件下减压蒸馏提纯0.5小时，除去轻组分，塔釜流出物送入另一蒸馏塔继续蒸馏，得到2,2-二甲基-1,3-丙二醇成品，其中，减压蒸馏的同时，检测塔釜流出物使用相对物料质量20％的去离子水室温溶解物料后的pH值，根据需要加入适量5％的磷酸水溶液以调节蒸馏塔塔釜流出物pH值为6.0，在减压蒸馏的过程中，蒸馏塔的回流比为0.5。
经检测得到的2,2-二甲基-1,3-丙二醇成品的质量含量为99.5％。
实施例2
(1)调节并维持三台缩合反应器组成的缩合反应系统的温度为65℃， 向缩合反应系统中第一台反应器中连续地以1000千克/小时的速度送入质量含量为99％的异丁醛，同时向反应器中连续地以1100千克/小时的速度送入质量含量为37％的甲醛水溶液。根据pH值计的测量数据，向第一台反应器中连续地送入质量含量为30％的三甲胺水溶液，调节第一台反应器出口物料的pH值为8.0，并在第二台反应器入口处连续地送入甲酸，调节第三台反应器出口物料的pH值为6.0。调节反应器内的液位和出料速度，以使物料的总保留时间保持在1.5小时，得到缩合反应产物。
(2)将步骤(1)得到的缩合反应产物通过蒸发器回收未反应的原料后进行加氢反应，采用装有镍催化剂的固定床反应器为催化加氢反应器，在H 2存在下，控制缩合产物的液体空速为0.1h -1，H 2的气体空速为300h -1，在温度为110℃，压力为2.0MPa条件下反应制得2,2-二甲基-1,3-丙二醇粗品。
(3)将步骤(2)得到的2,2-二甲基-1,3-丙二醇粗品在压力为绝对压力10kPa，塔釜温度为145℃条件下减压蒸馏提纯1.0小时，除去轻组分，塔釜流出物送入另一蒸馏塔继续蒸馏，得到2,2-二甲基-1,3-丙二醇成品，其中，减压蒸馏的同时，检测塔釜流出物使用相对物料质量20％的去离子水室温溶解物料后的pH值，根据需要加入适量5％的磷酸水溶液以调节蒸馏塔塔釜流出物的pH值为5.7，调节蒸馏塔的回流比为2.0。
经检测得到的2,2-二甲基-1,3-丙二醇成品的质量含量为99.5％。
实施例3
(1)调节并维持三台缩合反应器组成的缩合反应系统的温度为70℃， 向缩合反应系统中第一台反应器中连续地以1000千克/小时的速度送入质量含量为99％的异丁醛，同时向反应器中连续地以1250千克/小时的速度送入质量含量为37％的甲醛水溶液。根据pH值计的测量数据，向反应器中连续地送入质量含量为30％的三甲胺水溶液调节第三台反应器出口物料的pH值为8.5。调节反应器内的液位和出料速度，以使物料的总保留时间保持在1.2小时，得到缩合反应产物。
(2)将步骤(1)得到的缩合反应产物通过蒸发器回收未反应的原料后进行加氢反应，采用装有镍催化剂的固定床反应器为催化加氢反应器，在H 2存在下，控制缩合产物的液体空速为0.1h -1，H 2的气体空速为300h -1，在温度为110℃，压力为2.0MPa条件下反应制得2,2-二甲基-1,3-丙二醇粗品。
(3)将步骤(2)得到的2,2-二甲基-1,3-丙二醇粗品在压力为绝对压力10kPa，塔釜温度为150℃条件下减压蒸馏提纯2.3小时，除去轻组分，塔釜流出物送入另一蒸馏塔继续蒸馏，得到2,2-二甲基-1,3-丙二醇成品，其中，减压蒸馏的同时，检测塔釜流出物使用相对物料质量20％的去离子水室温溶解物料后的pH值，根据需要加入适量5％的磷酸水溶液以调节蒸馏塔塔釜流出物的pH值为5.3，调节蒸馏塔的回流比为1.0。
经检测得到的2,2-二甲基-1,3-丙二醇成品的质量含量为99.5％。
实施例4
(1)调节并维持三台缩合反应器组成的缩合反应系统的温度为75℃，向缩合反应系统中第一台反应器中连续地以1000千克/小时的速度送入质 量含量为99％的异丁醛，同时向反应器中连续地以1050千克/小时的速度送入质量含量为37％的甲醛水溶液。根据pH值计的测量数据，向反应器中连续地送入质量含量为30％的三甲胺水溶液调节第三台反应器出口物料的pH值为8.0。调节反应器内的液位和出料速度，以使物料的总保留时间保持在1.0小时，得到缩合反应产物。
(2)将步骤(1)得到的缩合反应产物通过蒸发器回收未反应的原料后进行加氢反应，采用装有镍催化剂的固定床反应器为催化加氢反应器，在H 2存在下，控制缩合产物的液体空速为0.1h -1，H 2的气体空速为300h -1，在温度为110℃，压力为2.0MPa条件下反应制得2,2-二甲基-1,3-丙二醇粗品。
(3)将步骤(2)得到的2,2-二甲基-1,3-丙二醇粗品在压力为绝对压力10kPa，塔釜温度为150℃条件下减压蒸馏提纯3.0小时，除去轻组分，塔釜流出物送入另一蒸馏塔继续蒸馏，得到2,2-二甲基-1,3-丙二醇成品，其中，减压蒸馏的同时，检测塔釜流出物使用相对物料质量20％的去离子水室温溶解物料后的pH值，根据需要加入适量5％的磷酸水溶液以调节蒸馏塔塔釜流出物pH值为6.2，调节蒸馏塔的回流比为0.8。
经检测得到的2,2-二甲基-1,3-丙二醇成品的质量含量为99.3％。
实施例5
(1)调节并维持三台缩合反应器组成的缩合反应系统的温度为80℃，向缩合反应系统中第一台反应器中连续地以1000千克/小时的速度送入质量含量为99％的异丁醛，同时向反应器中连续地以1150千克/小时的速度 送入质量含量为37％的甲醛水溶液。根据pH值计的测量数据，向第一台反应器中连续地送入质量含量为30％的三甲胺水溶液，调节第一台反应器出口物料的pH值为8.0，并在第二台反应器入口处连续地送入异丁酸，调节第三台反应器出口物料的pH值为6.5。调节反应器内的液位和出料速度，以使物料的总保留时间保持在0.8小时，得到缩合反应产物。
(2)将步骤(1)得到的缩合反应产物通过蒸发器回收未反应的原料后进行加氢反应，采用装有镍催化剂的固定床反应器为催化加氢反应器，在H 2存在下，控制缩合产物的液体空速为0.1h -1，H 2的气体空速为300h -1，在温度为110℃，压力为2.0MPa条件下反应制得2,2-二甲基-1,3-丙二醇粗品。
(3)将步骤(2)得到的2,2-二甲基-1,3-丙二醇粗品在压力为绝对压力10kPa，塔釜温度为150℃条件下减压蒸馏提纯4.0小时，除去轻组分，塔釜流出物送入另一蒸馏塔继续蒸馏，得到2,2-二甲基-1,3-丙二醇成品，其中，减压蒸馏的同时，检测塔釜流出物使用相对物料质量20％的去离子水室温溶解物料后的pH值，根据需要加入适量5％的磷酸水溶液以调节蒸馏塔塔釜流出物pH值为4.1，调节蒸馏塔的回流比为1.0。
经检测得到的2,2-二甲基-1,3-丙二醇成品的质量含量为99.4％。
实施例6
(1)调节并维持三台缩合反应器组成的缩合反应系统的温度为80℃，向缩合反应系统中第一台反应器中连续地以1000千克/小时的速度送入质量含量为99％的异丁醛，同时向反应器中连续地以1250千克/小时的速度 送入质量含量为37％的甲醛水溶液。根据pH值计的测量数据，向反应器中连续地送入质量含量为30％的三甲胺水溶液调节第三台反应器出口物料的pH值为8.0。调节反应器内的液位和出料速度，以使物料的总保留时间保持在0.5小时，得到缩合反应产物。
(2)将步骤(1)得到的缩合反应产物通过蒸发器回收未反应的原料后进行加氢反应，采用装有镍催化剂的固定床反应器为催化加氢反应器，在H 2存在下，控制缩合产物的液体空速为0.1h -1，H 2的气体空速为300h -1，在温度为110℃，压力为2.0MPa条件下反应制得2,2-二甲基-1,3-丙二醇粗品。
(3)将步骤(2)得到的2,2-二甲基-1,3-丙二醇粗品在压力为绝对压力10kPa，塔釜温度为155℃条件下减压蒸馏提纯1.0小时，除去轻组分，塔釜流出物送入另一蒸馏塔继续蒸馏，得到2,2-二甲基-1,3-丙二醇成品，其中，蒸馏塔填料层中装入硅铝比220的氢型ZSM-5分子筛催化剂，减压蒸馏的同时，检测塔釜流出物使用相对物料质量20％的去离子水室温溶解物料后的pH值，根据需要加入适量5％的磷酸水溶液以调节蒸馏塔塔釜流出物的pH值为5.5，调节蒸馏塔的回流比为0.8。
经检测得到的2,2-二甲基-1,3-丙二醇成品的质量含量为99.4％。
实施例7
(1)调节并维持三台缩合反应器组成的缩合反应系统的温度为60℃，向缩合反应系统中第一台反应器中连续地以1000千克/小时的速度送入质量含量为99％的异丁醛，同时向反应器中连续地以1250千克/小时的速度 送入质量含量为37％的甲醛水溶液。根据pH值计的测量数据，向反应器中连续地送入质量含量为30％的三甲胺水溶液调节第三台反应器出口物料的pH值为8.5。调节反应器内的液位和出料速度，以使物料的总保留时间保持在2.0小时，得到缩合反应产物。
(2)将步骤(1)得到的缩合反应产物通过蒸发器回收未反应的原料后进行加氢反应，采用装有镍催化剂的固定床反应器为催化加氢反应器，在H 2存在下，控制缩合产物的液体空速为0.1h -1，H 2的气体空速为300h -1，在温度为110℃，压力为2.0MPa条件下反应制得2,2-二甲基-1,3-丙二醇粗品。
(3)将步骤(2)得到的2,2-二甲基-1,3-丙二醇粗品在压力为绝对压力10kPa，塔釜温度为142℃条件下减压蒸馏提纯0.5小时，除去轻组分，塔釜流出物送入另一蒸馏塔继续蒸馏，得到2,2-二甲基-1,3-丙二醇成品，其中，减压蒸馏的同时，检测塔釜流出物使用相对物料质量20％的去离子水室温溶解物料后的pH值，根据需要加入适量5％的磷酸水溶液以调节蒸馏塔塔釜流出物pH值为3.0，在减压蒸馏的过程中，蒸馏塔的回流比为0.5。
经检测得到的2,2-二甲基-1,3-丙二醇成品的质量含量为99.1％。
使用气相色谱氢火焰法对实施例1-6和对比例1-2得到的2,2-二甲基-1,3-丙二醇成品进行分析，目标成品中2,2-二甲基-1,3-丙二醇和甲酸2,2-二甲基-1,3-丙二醇酯、异丁酸2,2-二甲基-1,3-丙二醇酯平均含量如下表所示。
对比例1(对比参数：甲醛用量和pH值)
(1)调节并维持三台缩合反应器组成的缩合反应系统的温度为65℃， 向缩合反应系统中第一台反应器中连续地以1000千克/小时的速度送入质量含量为99％的异丁醛，同时向反应器中连续地以1000千克/小时的速度送入质量含量为37％的甲醛水溶液。根据pH值计的测量数据，向反应器中连续地送入质量含量为30％的三甲胺水溶液调节第三台反应器出口物料的pH值为9.2。调节反应器内的液位和出料速度，以使物料的总保留时间保持在1.0小时，得到缩合反应产物。
(2)将步骤(1)得到的缩合反应产物通过蒸发器回收未反应的原料后进行加氢反应，采用装有镍催化剂的固定床反应器为催化加氢反应器，在H 2存在下，控制缩合产物的液体空速为0.1h -1，H 2的气体空速为300h -1，在温度为110℃，压力为2.0MPa条件下反应制得2,2-二甲基-1,3-丙二醇粗品。
(3)将步骤(2)得到的2,2-二甲基-1,3-丙二醇粗品在压力为绝对压力10kPa，塔釜温度为139℃条件下减压蒸馏提纯0.3小时，除去轻组分，塔釜流出物送入另一蒸馏塔继续蒸馏，得到2,2-二甲基-1,3-丙二醇成品，其中，蒸馏塔填料层中装入硅铝比220的氢型ZSM-5分子筛催化剂，减压蒸馏的同时，检测塔釜流出物使用相对物料质量20％的去离子水室温溶解物料后的pH值，根据需要加入适量5％的磷酸水溶液以调节蒸馏塔塔釜流出物的pH值为7.8，调节蒸馏塔的回流比为0.4。
经检测得到的2,2-二甲基-1,3-丙二醇成品的质量含量为99.2％。
对比例2(对比参数：反应时间)
(1)调节并维持三台缩合反应器组成的缩合反应系统的温度为85℃， 向缩合反应系统中第一台反应器中连续地以1000千克/小时的速度送入质量含量为99％的异丁醛，同时向反应器中连续地以1100千克/小时的速度送入质量含量为37％的甲醛水溶液。根据pH值计的测量数据，向反应器中连续地送入质量含量为30％的三甲胺水溶液调节第三台反应器出口物料的pH值为8.5。调节反应器内的液位和出料速度，以使物料的总保留时间保持在6.0小时，得到缩合反应产物。
(2)将步骤(1)得到的缩合反应产物通过蒸发器回收未反应的原料后进行加氢反应，采用装有镍催化剂的固定床反应器为催化加氢反应器，在H 2存在下，控制缩合产物的液体空速为0.1h -1，H 2的气体空速为300h -1，在温度为110℃，压力为2.0MPa条件下反应制得2,2-二甲基-1,3-丙二醇粗品。
(3)将步骤(2)得到的2,2-二甲基-1,3-丙二醇粗品在压力为绝对压力10kPa，塔釜温度为145℃条件下减压蒸馏提纯2.0小时，除去轻组分，塔釜流出物送入另一蒸馏塔继续蒸馏，得到2,2-二甲基-1,3-丙二醇成品，其中，蒸馏塔填料层中装入硅铝比220的氢型ZSM-5分子筛催化剂，减压蒸馏的同时，检测塔釜流出物使用相对物料质量20％的去离子水室温溶解物料后的pH值，根据需要加入适量5％的磷酸水溶液以调节蒸馏塔塔釜流出物的pH值为6.0，调节蒸馏塔的回流比为2.5。
经检测得到的2,2-二甲基-1,3-丙二醇成品的质量含量为98.9％。
Figure PCTCN2019095988-appb-000001
显然，上述实施例仅仅是为清楚地说明所作的举例，而并非对实施方式的限定。对于所属领域的普通技术人员来说，在上述说明的基础上还可以做出其它不同形式的变化或变动。这里无需也无法对所有的实施方式予以穷举。而由此所引申出的显而易见的变化或变动仍处于本发明创造的保护范围之中。

Claims (16)

  1. 一种2,2-二甲基-1,3-丙二醇的生产工艺，包括，将质量含量为37％的甲醛水溶醛和质量含量99％的异丁醛按质量比为1.05-1.25:1混合，调节反应体系的pH值为6.0-8.5，在反应温度为60-80℃，反应时间为0.5-2.0小时的条件下经缩合反应得到中间体，所述中间体经催化加氢反应后得到2,2-二甲基-1,3-丙二醇粗品，然后经提纯得到2,2-二甲基-1,3-丙二醇成品。
  2. 根据权利要求1所述的2,2-二甲基-1,3-丙二醇的生产工艺，其特征在于，所述缩合反应过程中，加入碱性调节剂调节反应体系的pH值为6.0-8.5，优选地，加入碱性调节剂调节反应体系的pH值为6.5-8.0。
  3. 根据权利要求1所述的2,2-二甲基-1,3-丙二醇的生产工艺，其特征在于，所述缩合反应过程中，加入酸性调节剂和碱性调节剂调节反应体系的pH值为6.0-8.5，优选地，加入酸性调节剂和碱性调节剂调节反应体系的pH值为6.5-8.0。
  4. 根据权利要求2或者3所述的2,2-二甲基-1,3-丙二醇的生产工艺，其特征在于，所述碱性调节剂为有机叔胺，所述有机叔胺优选为三甲胺。
  5. 根据权利要求3所述的2,2-二甲基-1,3-丙二醇的生产工艺，其特征在于，所述酸性调节剂为有机酸，所述有机酸优选为甲酸或异丁酸中的任意一种或者两种。
  6. 根据权利要求1-5中任一项所述的2,2-二甲基-1,3-丙二醇的生产工艺，其特征在于，在所述缩合反应过程中，反应温度优选为65-75℃，更优选为65-70℃。
  7. 根据权利要求1-6中任一项所述的2,2-二甲基-1,3-丙二醇的生产工艺，其特征在于，在所述缩合反应过程中，反应时间优选为0.8-1.5小时，更优选为1.0-1.2小时。
  8. 根据权利要求1-7中任一项所述的2,2-二甲基-1,3-丙二醇的生产工艺，其特征在于，在所述缩合反应过程中，质量含量为37％的甲醛水溶液和质量含量为99％的异丁醛的质量比为1.05-1.15:1。
  9. 根据权利要求1-8中任一项所述的2,2-二甲基-1,3-丙二醇的生产工艺，其特征在于，所述提纯过程为将2,2-二甲基-1,3-丙二醇粗品加入蒸馏塔中，在蒸馏塔塔釜的温度为139-155℃，压力为1-50kPa条件下经一次减压蒸馏0.5-8.0小时，排出轻组分，然后经二次蒸馏除去重组分，得到2,2-二甲基-1,3-丙二醇成品。
  10. 根据权利要求9所述的2,2-二甲基-1,3-丙二醇的生产工艺，其特征在于，所述提纯过程中，2,2-二甲基-1,3-丙二醇粗品经一次减压蒸馏的时间优选为1.0-4.0小时。
  11. 根据权利要求9或者10中所述的2,2-二甲基-1,3-丙二醇的生产工艺，其特征在于，所述提纯过程中，调节蒸馏塔内塔釜和提馏段物料体系的酸度，使蒸馏塔塔釜流出物的pH值为4.0-7.0，优选pH值为4.5-6.5。
  12. 根据权利要求9-11中任一项所述的2,2-二甲基-1,3-丙二醇的生产工艺，其特征在于，所述提纯过程中，酯类杂质反应时不另外引入催化剂。
  13. 根据权利要求9-11中任一项所述的2,2-二甲基-1,3-丙二醇的生产工艺，其特征在于，所述提纯过程中，酯类杂质反应时引入催化剂。
  14. 根据权利要求13所述的2,2-二甲基-1,3-丙二醇的生产工艺，其特征在于，所述提纯过程中，所述催化剂为酸性催化剂，优选地，所述催化剂为有机酸、无机酸或者固体酸催化剂中的一种或几种的混合物，所述催化剂的添加量以保证塔釜体系的pH值为4.0-7.0。
  15. 根据权利要求14所述的2,2-二甲基-1,3-丙二醇的生产工艺，其特征在于，所述固体酸催化剂为硅铝比不低于200的分子筛催化剂。
  16. 根据权利要求9-15中任一项所述的2,2-二甲基-1,3-丙二醇的生产工艺，其特征在于，在所述减压蒸馏过程中，塔顶物流的回流比为0.5-2.0，优选所述回流比为0.5-1.0。
PCT/CN2019/095988 2018-08-06 2019-07-15 一种2,2-二甲基-1,3-丙二醇的生产工艺 WO2020029753A1 (zh)

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