WO2017133357A1 - 乙二醇和1,2-丁二醇的分离方法、工艺及装置 - Google Patents

乙二醇和1,2-丁二醇的分离方法、工艺及装置 Download PDF

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WO2017133357A1
WO2017133357A1 PCT/CN2016/112840 CN2016112840W WO2017133357A1 WO 2017133357 A1 WO2017133357 A1 WO 2017133357A1 CN 2016112840 W CN2016112840 W CN 2016112840W WO 2017133357 A1 WO2017133357 A1 WO 2017133357A1
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ethylene glycol
ketone
column
acetal
butanediol
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PCT/CN2016/112840
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English (en)
French (fr)
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李鑫钢
高鑫
李洪
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天津大学
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    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C29/00—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring
    • C07C29/74—Separation; Purification; Use of additives, e.g. for stabilisation
    • 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
    • C07C29/74—Separation; Purification; Use of additives, e.g. for stabilisation
    • C07C29/76—Separation; Purification; Use of additives, e.g. for stabilisation by physical treatment
    • C07C29/80—Separation; Purification; Use of additives, e.g. for stabilisation by physical treatment by distillation
    • 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
    • C07C29/74—Separation; Purification; Use of additives, e.g. for stabilisation
    • C07C29/88—Separation; Purification; Use of additives, e.g. for stabilisation by treatment giving rise to a chemical modification of at least one compound
    • C07C29/92—Separation; Purification; Use of additives, e.g. for stabilisation by treatment giving rise to a chemical modification of at least one compound by a consecutive conversion and reconstruction
    • 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

Definitions

  • the invention belongs to the field of chemical rectification separation, and relates to a separation method, a process and a device for ethylene glycol and 1,2-butanediol.
  • ethylene glycol is mainly used in the production of polyester fiber, antifreeze, unsaturated polyester resin, lubricant, plasticizer, nonionic surfactant and explosive.
  • ethylene glycol Alcohol can also be used in coatings, photographic developing solutions, brake fluids and inks, as a solvent and medium for ammonium perborate, for the production of special solvent glycol ethers, etc., and is widely used.
  • studies have pointed out that ethylene glycol has great potential for development in the field of liquid fuel cells in the future. China is a big consumer of ethylene glycol. The domestic production of ethylene glycol is far from meeting its own needs. In 2014, China's ethylene glycol production was 3.5 million tons, and the apparent consumption reached 12.25 million tons. Therefore, China's The ethylene glycol industry has good prospects for development.
  • ethylene glycol production at home and abroad uses direct hydration or pressurized hydration process lines.
  • the process is to mix ethylene oxide and water into a mixed aqueous solution at a ratio of 1:20 to 22 (molar ratio) in a fixed bed.
  • the reactor is reacted at 130-180 ° C, 1.0-2.5 MPa for 18-30 minutes, and the ethylene oxide is completely converted into a mixed alcohol, and then dehydrated and concentrated by a multi-effect evaporator and vacuum distillation to obtain ethylene glycol.
  • Evaporative dehydration requires a lot of energy.
  • some major domestic and foreign companies producing ethylene glycol and research institutes have begun to study catalytic water law.
  • the process of obtaining ethylene glycol by hydrogenation of oxalate is considered to be the most industrially promising route. It is currently put into operation and under construction in China. More than 20 sets of coal-based ethylene glycol units.
  • the reaction product of hydrogenation of oxalate to ethylene glycol contains a small amount of 1,2-propanediol, 1,2-butanediol and the like in addition to a lower boiling point such as methanol or glycolic acid ester.
  • the boiling point of alcohol is close to that which is difficult to separate by ordinary rectification.
  • 1,2-butanediol has the closest boiling point to ethylene glycol and forms azeotropy at the same time, so it is most difficult to separate.
  • the ethylene glycol refining tower purifies the ethylene glycol by means of the top of the tower.
  • many theoretical plates and high reflux ratios have to be used, resulting in a pressure drop. If it is too high, the temperature of the column is too high, so it is easy to affect the quality of the ethylene glycol in the column.
  • How to effectively separate ethylene glycol and 1,2-butanediol reduce the number of ethylene glycol refining trays and reflux ratio, increase the yield of ethylene glycol, and recover high value-added 1,2-butanediol, which is lower The key to production costs and economic efficiency.
  • Patent CN 101928201 discloses a purification process for a crude ethylene glycol product, and the purification process is reversed by saponification.
  • High-purity ethylene glycol products are obtained from coal-based ethylene glycol crude products by methanol, hydrogenation reaction, three-column distillation and adsorption treatment.
  • 1,2-butanediol only passes through ethylene glycol.
  • Azeotropic removal not only causes a high value-added product loss of 1,2-butanediol, but also affects the yield of ethylene glycol.
  • Patent CN 102276418 proposes the use of a mixture of ethylene glycol and 1,2-butanediol for a process for the hydrogenation of ethylene glycol products from oxalates, by using at least ZSM-5, Y zeolite or beta zeolite.
  • An adsorbent that adsorbs 1,2-butanediol to obtain purification of ethylene glycol.
  • the recovery of ethylene glycol is not indicated, and the recycling of 1,2-butanediol is not involved.
  • Patent CN 103193594 discloses a method for separating and purifying ethylene glycol from a liquid phase product of oxalate hydrogenation by azeotropic distillation using a dioxoane compound which is immiscible with ethylene glycol as an azeotroping agent.
  • the liquid mixture is removed from the low boiling point compound such as methanol and methyl glycolate through two de-lighting towers, and then enters the azeotropic distillation column.
  • the azeotrope and the ethylene glycol azeotrope are produced at the top of the column, and after phase separation
  • the azeotrope is refluxed at the top of the column, and the ethylene glycol enters the ethylene glycol refining tower.
  • the method varies with the composition of the raw materials, and the recovery rate of ethylene glycol varies greatly from 99.1% to 89.65%, while 1,2-butyl The recovery of alcohol is also poorly considered.
  • Patent CN 102372596 proposes a separation process for separating ethylene glycol, propylene glycol and butanediol by using azeotropic distillation, and the synthesis gas ethylene glycol product is removed from the light component by the fraction cutting tower, and then passed through an azeotropic distillation column ( 2), the top of the tower removes propylene glycol by extracting a solution of ethylene glycol rich in 1,2-propanediol, and the stream of the tower is transferred to the next azeotropic distillation column (3), which is rich in 1,2-butyl
  • the ethylene glycol solution of the alcohol, the azeotropic agent is returned to the top of the column after the top phase separation, and the ethylene glycol stream is taken into the ethylene glycol refined tower.
  • the object of the present invention is to provide a separation and refining method, a process and a device for ethylene glycol and 1,2-butanediol, which are applied in a coal-to-ethylene glycol product separation process or a biomass-catalyzed ethylene glycol process.
  • a method for separating and purifying a mixture of ethylene glycol and 1,2-butanediol firstly reacting ethylene glycol and 1,2-butanediol with a reactant aldehyde/ketone by an acetal or a ketal to form a corresponding acetal/ketone
  • the product liquid mixture recovers the aldehyde/ketone in the liquid mixture of the acetal/ketone product of ethylene glycol and 1,2-butanediol, and the liquid mixture is passed through a separation column to separate the different acetal/ketone products by rectification, and then
  • the acetal/ketone product is hydrolyzed to obtain an initial product of ethylene glycol and 1,2-butanediol, and finally the respective preliminary products are purified by rectification to obtain ethylene glycol and 1,2-butanediol products.
  • the reactant aldehyde/ketone is one of an aldehyde having 1 to 8 carbon atoms or a ketone having 3 to 8 carbon atoms.
  • the reactant is acetaldehyde or acetone.
  • hydrolyzed ethylene glycol and the 1,2-butanediol preliminary product were further purified by ethylene glycol and 1,2-butanediol to further purify ethylene glycol and 1,2-butanediol.
  • a separation and refining process for a mixture of ethylene glycol and 1,2-butanediol comprising the steps of:
  • the liquid mixed reaction product (S03) enters the aldehyde/ketone recovery column (T11), and the overhead discharge mainly contains an unreacted aldehyde/ketone stream (S09), which is recycled, and the acetal/ketone reaction product is produced in the column kettle. And unreacted diol mixture stream S10;
  • the column reactor stream (S10) enters the ethylene glycol acetal/ketone product separation column (T12), and an azeotrope (S16) of the ethylene glycol acetal/ketone product and water is produced at the top of the column, and the tower kettle contains 1 a stream of 2-butanediol acetal/ketone product (S17);
  • the tower kettle recovery stream (S17) enters the 1,2-butanediol acetal/ketone product separation column (T13), and a mixture of butanediol acetal/ketone product and water is produced at the top of the column (S37).
  • the tower kettle is produced as a reaction mixture of diol mixture (S38), recycled to the reactor (R1);
  • Step (3) The overhead production stream (S16) and the supplementary stream water (S18) enter the acetal/ketone hydrolysis rectification column (T21), the reaction section is set in the middle of the column, and the hydrolyzate aldehyde/ketone (S23) is discharged from the top of the column. Recycling, the tower kettle produces an aqueous glycol stream (S24);
  • Step (5) The tower kettle production stream (S24) enters the ethylene glycol refining tower (T22), the top of the tower excludes water (S30), is recycled, and the tower kettle extracts the ethylene glycol product (S31);
  • Step (4) The overhead production stream (S37) and the supplementary stream water (S40) enter another acetal/ketone hydrolysis rectification column (T31), the reaction section is set in the middle of the column, and the hydrolyzate aldehyde/ketone is discharged from the top of the column ( S45), recycled, the tower kettle produces an aqueous 1,2-butanediol stream (S46);
  • Step (7) The tower kettle production stream (S46) enters the 1,2-butanediol refining tower (T32), the top of the tower excludes water (S52), is recycled, and the tower kettle produces 1,2-butanediol product (S53) ).
  • reaction conditions of the above steps are as follows:
  • the reactor R1 in the step (1) adopts one of a tank reactor and a fixed bed reactor, and the operating pressure is 0.5 to 10 atm in absolute pressure, and the reaction temperature is 40 to 200 ° C, and a catalyst using a solid acid or a solid alkali resin catalyst is catalyzed. Or mixing, the molar ratio of the aldehyde / ketone feed to the total amount of diol feed into the reactor is 1 ⁇ 10;
  • the aldehyde/ketone recovery column T11 of the step (2) has an operating pressure of 0.1 to 10 atm in absolute pressure, a reflux ratio of 0.01 to 15, or direct gas phase extraction, and no reflux;
  • the operating pressure of the butanediol acetal/ketone product separation column T13 in step (4) is 0.01 to 5 atm and the reflux ratio is 0.1 to 15;
  • the acetal/ketone hydrolyzed rectification column T21 of the step (5) has an operating pressure of 0.1 to 10 atm in absolute pressure, a reflux ratio of 0.01 to 20, or direct gas phase extraction at the top of the column, without reflux, and water in the replenished waters S18 and S16.
  • the molar ratio of the total feed amount to the total amount of acetal/ketone feed is from 1 to 10;
  • the operating pressure of the ethylene glycol refining column T22 of the step (6) is 0.01 to 2 atm and the reflux ratio is 0.01 to 10;
  • the acetal/ketone hydrolyzed rectification column T31 of the step (7) has an operating pressure of 0.1 to 10 atm in absolute pressure, a reflux ratio of 0.01 to 20, or a direct gas phase extraction at the top of the column, and a total feed of water in the replenished waters S40 and S37. Molar ratio of the amount to the total amount of acetal/ketone 1 to 10;
  • the 1,2-butanediol refining column T32 of the step (8) has an operating pressure of 0.01 to 2 atm and a reflux ratio of 0.01 to 10 in terms of absolute pressure.
  • the hydrolysis of the acetal/ketone product of the step (5) and the step (7) is carried out by means of separation after the first reaction, and the streams (S16) and (S46) are respectively introduced into the respective hydrolysis units, and the hydrolysis unit comprises a reactor (R2 or R3), an aldehyde/ketone separation column (T51 or T61), an acetal/ketone product recovery column (T52 or T62), and a product refining column (T53 or T63).
  • the reactor (R2 or R3) operating pressure is 0.5 to 10 atm in terms of absolute pressure, and the reaction temperature is 40 to 200 ° C, catalyzing one or a mixture of solid acid or solid alkali resin catalysts; and an aldehyde/ketone separation column ( T51, T61)
  • the operating pressure is 0.5-1010m under absolute pressure, the reflux ratio is 0.01-15, or the gas phase is directly produced at the top of the column;
  • the operating pressure of the acetal/ketone product recovery tower (T52, T62) is 0.1 ⁇ by absolute pressure. 10atm, reflux ratio of 0.1 to 15; product refining tower (T53, T63) operating pressure is 0.01 ⁇ 2atm absolute, reflux ratio of 0.01 ⁇ 10.
  • the acetal/ketone product of 1,2-butanediol is not hydrolyzed, and the acetal/ketone product is directly isolated, such as the column T41 shown in Fig. 2, and the phase separator is arranged at the top of the column, and the operating temperature is 20-150. °C, the aqueous phase is discharged, and the organic phase is refluxed.
  • the column T41 is 0.05 to 5 atm in terms of absolute pressure.
  • a separation and refining device for a mixture of ethylene glycol and 1,2-butanediol comprising a reactor (R1), an aldehyde/ketone recovery column (T11), an ethylene glycol acetal/ketone product separation column (T12), and Glycol acetal/ketone product separation column (T13), acetal/ketone hydrolysis rectification column (T21, T31), ethylene glycol refining column (T22), 1,2-butanediol refining column (T32), condensation , reboilers, pumps and associated feed lines and lines connecting the above equipment; condensers and reboilers are provided at the top of the distillation column and at the bottom of the column.
  • the reactor (R1) is discharged into the aldehyde/ketone recovery column (T11), and the unreacted aldehyde/ketone pump is sent to the reactor (R1) at the top of the column.
  • the ethylene glycol acetal/ketone product separation column (T12) The ethylene glycol acetal/ketone product separation column (T12) is topped with a mixture of ethylene glycol acetal/ketone product and water into the ethylene glycol acetal/ketone product hydrolysis distillation column (T21).
  • the outlet stream enters the butanediol acetal/ketone product separation column (T13); the butanediol acetal/ketone product separation column (T13) produces a mixture of butanediol acetal/ketone product and water into the dibutyl Alcohol acetal/ketone product hydrolyzed rectification column (T31), the column kettle is used to recycle unreacted diol mixture; ethylene glycol acetal/ketone product hydrolyzed rectification column (T21) is set in the middle of the reaction section, The aldehyde/ketone product is used for recycling, and the mixture of ethylene glycol and water is taken into the ethylene glycol refining tower (T22); the top of the ethylene glycol refining tower (T22) is dehydrated, and the bottom of the column is B.
  • the diol product; the butanediol acetal/ketone product hydrolyzed rectification column (T31) is provided with a reaction section in the middle, and the aldehyde/ketone product is produced at the top of the column for recycling, and the mixture of the butanediol and water is taken in the column kettle.
  • the separation and purification method provided by the invention can efficiently separate the ethylene glycol and the 1,2-butanediol in the azeotropic state, and completely solves the problem that the ethylene glycol and the 1,2-butanediol cannot be completely separated due to the close boiling point.
  • the problem is to obtain a high purity ethylene glycol product, a 1,2-butanediol product or an acetal/ketone product.
  • the most significant two points of the separation and refining method of the present invention are energy saving and product recovery product quality, and the present invention converts to acetal by comparison with the existing separation of ethylene glycol and 1,2-butanediol by a separate rectification method. After the diol is formed, the energy consumption is saved by 50-60%, the quality of the recovered product is high, and the recovered ethylene glycol is polyester grade, which can be directly used for polyester production, and the purity of the produced 1,2-butanediol. Up to 99% or more, it can also be used directly for subsequent production.
  • the condensation process can be applied to the reaction rectification method, and the organic reaction is coupled with the rectification process, which not only saves equipment investment, but also reduces the operating cost of subsequent separation.
  • the purity of the main product ethylene glycol product in the invention can reach above 99.9%, the recovery rate can reach above 99.5%, and the purity of the 1,2-butanediol product can reach above 98.5% (1,2-butanediol acetal) / ketone product purity can reach 99.6% or more).
  • the invention has the advantages of adopting the method of reactive distillation, the separation problem of the acetal/ketone product which is easy to separate due to the close boiling point, relatively low volatility and azeotropy of ethylene glycol and 1,2-butanediol. The separation of ethylene glycol and 1,2-butanediol is effectively achieved.
  • the separation and purification method provided by the invention is applied to the oxalate process to produce ethylene glycol, which can reduce the load of the ethylene glycol separation and purification tower, improve the quality and yield of ethylene glycol, and recover high value-added products1. 2-butanediol, achieving good annual economic benefits.
  • the separation and refining method provided by the invention can well realize the separation of the azeotropic polyethylene glycol and the 1,2-butanediol mixture, and the whole separation process is simple in operation and low in energy consumption, and Ethylene glycol, butanediol (or butanediol acetal / ketone) products have high purity and high ethylene glycol recovery.
  • Figure 1 is a schematic diagram of a process for separation and purification of a mixture of ethylene glycol and 1,2-butanediol.
  • FIG. 2 is a schematic view showing a process of separation and purification of a modified mixture of ethylene glycol and 1,2-butanediol.
  • Fig. 3 is a schematic view showing the process of separation and purification of the modified mixture of ethylene glycol and 1,2-butanediol.
  • Fig. 4 is a schematic view showing the process of separation and purification of the modified ethylene glycol and 1,2-butanediol mixture by reactive distillation.
  • the feed is derived from a coal-to-ethylene glycol process or a biomass-catalyzed ethylene glycol process, firstly, ethylene glycol and 1,2-butane
  • the diol is reacted by acetal or ketal to form a corresponding acetal/ketone product liquid mixture, and the acetal/ketone product liquid mixture containing ethylene glycol and 1,2-butanediol is passed through a recovery column to recover unreacted aldehyde.
  • the liquid mixture is passed through a separation column to separate the different acetal/ketone products by rectification according to the difference in boiling point of the acetal or ketal, and then the acetal/ketone product is hydrolyzed to obtain ethylene glycol and 1,2-butanediol. Further, ethylene glycol and 1,2-butanediol are refined.
  • the reaction to be used is an acetal or ketal reaction, and the reactant is an aldehyde such as formaldehyde, acetaldehyde, propionaldehyde or butyraldehyde or a ketone such as acetone, methyl ethyl ketone or cyclohexanone.
  • acetaldehyde and acetone work well.
  • the separation problem is transformed into an acetal/ketone product that is easy to separate to achieve their effective separation.
  • the method comprises the following steps:
  • the stream S03 enters the aldehyde/ketone recovery column T11, and the overhead discharge mainly contains the unreacted aldehyde/ketone stream S09, which is recycled, and the tower occupant produces the acetal/ketone reaction product and the unreacted glycol mixture stream S10. ;
  • Stream S17 enters the butanediol acetal/ketone product separation column T13, and a mixture of butanediol acetal/ketone product and water is produced at the top of the stream, and the diol mixture stream S38 is recovered. Circulation to reactor R1;
  • the present application also provides a better, higher yield process parameter of the above steps according to a specific scheme, as follows:
  • the reactor R1 in the step (1) adopts one of a tank reactor and a fixed bed reactor, and the operating pressure is 0.5 to 10 atm in absolute pressure, and the reaction temperature is 40 to 200 ° C, and a catalyst using a solid acid or a solid alkali resin catalyst is catalyzed.
  • the molar ratio of the amount of aldehyde/ketone feed to the total amount of glycol feed entering the reactor is from 1 to 10.
  • the aldehyde/ketone recovery column T11 of the step (2) has an operating pressure of 0.1 to 10 atm in absolute pressure, a reflux ratio of 0.01 to 15, or direct gas phase extraction without reflux.
  • the ethylene glycol acetal/ketone product separation column T12 of the step (3) has an operating pressure of 0.5 to 10 atm and a reflux ratio of 0.1 to 15 in terms of absolute pressure.
  • the butanediol acetal/ketone product separation column T13 of the step (4) has an operating pressure of 0.01 to 5 atm and a reflux ratio of 0.1 to 15 in terms of absolute pressure.
  • the acetal/ketone hydrolyzed rectification column T21 of the step (5) has an operating pressure of 0.1 to 10 atm in absolute pressure, a reflux ratio of 0.01 to 20, or direct gas phase extraction at the top of the column, without reflux, and water in the replenished waters S18 and S16.
  • the molar ratio of total feed to total acetal/ketone feed is from 1 to 10.
  • the ethylene glycol refining column T22 of the step (6) has an operating pressure of 0.01 to 2 atm and a reflux ratio of 0.01 to 10 in terms of absolute pressure.
  • the acetal/ketone hydrolyzed rectification column T31 of the step (7) has an operating pressure of 0.1 to 10 atm in absolute pressure, a reflux ratio of 0.01 to 20, or a direct gas phase extraction at the top of the column, and a total feed of water in the replenished waters S40 and S37.
  • the molar ratio of the amount to the total amount of acetal/ketone is from 1 to 10.
  • the 1,2-butanediol refining column T32 of the step (8) has an operating pressure of 0.01 to 2 atm and a reflux ratio of 0.01 to 10 in terms of absolute pressure.
  • the acetal/ketone product of 1,2-butanediol can also be directly hydrolyzed to obtain an acetal/ketone product, such as column T41 shown in Fig. 2, and a phase separator is arranged at the top of the column, and the operating temperature is 20 ⁇ . At 150 ° C, the aqueous phase was discharged, and the organic phase was refluxed.
  • the column T41 was 0.05 to 5 atm in terms of absolute pressure.
  • the invention also provides a separation and refining device for the mixture of ethylene glycol and 1,2-butanediol, which can be carried out according to the following device connection operation:
  • the apparatus mainly comprises a reactor (R1), an aldehyde/ketone recovery tower (T11), an ethylene glycol acetal/ketone product separation tower (T12), a butanediol acetal/ketone product separation tower (T13), and an acetal/ Ketone hydrolysis distillation column (T21, T31), ethylene glycol refining tower (T22), 1,2-butanediol refining tower (T32), condenser, reboiler, pump and related feed lines and pipelines connecting the above equipment; all rectification tower tops and towers are equipped with condensers and reboilers Device.
  • the reactor (R1) is discharged into the aldehyde/ketone recovery column (T11), and the unreacted aldehyde/ketone pump is sent to the reactor (R1) at the top of the column.
  • the ethylene glycol acetal/ketone product separation column (T12) The ethylene glycol acetal/ketone product separation column (T12) is topped with a mixture of ethylene glycol acetal/ketone product and water into the ethylene glycol acetal/ketone product hydrolysis distillation column (T21).
  • the outlet stream enters the butanediol acetal/ketone product separation column (T13); the butanediol acetal/ketone product separation column (T13) produces a mixture of butanediol acetal/ketone product and water into the dibutyl Alcohol acetal/ketone product hydrolyzed rectification column (T31), the column kettle is used to recycle unreacted diol mixture; ethylene glycol acetal/ketone product hydrolyzed rectification column (T21) is set in the middle of the reaction section, The aldehyde/ketone product is used for recycling, and the mixture of ethylene glycol and water is taken into the ethylene glycol refining tower (T22); the top of the ethylene glycol refining tower (T22) is dehydrated, and the bottom of the column is B.
  • the diol product; the butanediol acetal/ketone product hydrolyzed rectification column (T31) is provided with a reaction section in the middle, and the aldehyde/ketone product is produced at the top of the column for recycling, and the mixture of the butanediol and water is taken in the column kettle.
  • a mixture of ethylene glycol, 1,2-butanediol (S02) and a reaction raw material aldehyde or a ketone (S01) enters an acetal/ketone reactor (R1); S03) obtaining a liquid phase (S04) through a feed pump (P1) as a feed for the aldehyde/ketone recovery column (T11); and condensing the aldehyde/ketone recovery column (T11) at the top gas phase (S05) through a condenser (E2)
  • the post-liquid phase (S06) is passed through a reflux recovery pump (P3) to obtain a liquid phase (S07), a portion of which is refluxed as an aldehyde/ketone recovery column (S08), and the other portion is returned to the acetal/ketone reaction as a cyclic aldehyde/ketone (S09).
  • the liquid phase (S35), a part of which is the reflux of the acetal/ketone rectification column (S36), and the other part of the liquid phase (S37) is taken as the feed of the butanediol acetal/ketone product hydrolysis tower (T31).
  • the unreacted glycol stream (S38) is taken out through the pump (P9) to obtain a liquid phase (S39), and the cycle is reversed.
  • liquid phase water (S40) enters the butanediol acetal/ketone product hydrolyzed rectification column (T31); butanediol acetal/ketone product hydrolyzed rectification column (T31) overhead gas phase (S41)
  • the liquid phase (S42) after condensation of the condenser (E12) is passed through a reflux recovery pump (P12) to obtain a liquid phase (S43), a part of which is refluxed at the top of the hydrolyzed rectification column (S44), and the other part is an produced aldehyde or a ketone stream (S45) for recycling, a liquid phase (S46) of the column is taken through a feed pump (P11) to obtain a liquid phase (S47) as a feed for a butanediol refining column (T32); butanediol
  • the refining tower (T32) overhead gas phase (S48) is condensed by the condenser (
  • the reactor adopts a full-mixed reactor and a fixed-bed reactor, the operating pressure is 0.5-10 atm, the reaction temperature is 40-200 ° C, and the catalyst uses solid acid or solid alkali resin.
  • the catalyst the molar ratio of the amount of the aldehyde/ketone entering the reactor to the total amount of the diol is 1 to 10; the operating pressure of the aldehyde/ketone recovery column is 0.1 to 10 atm, and the reflux ratio is 0.01 to 15, Or the gasification of the top of the column is direct; the operating pressure of the ethylene glycol acetal/ketone product separation column is 0.5-1010m under absolute pressure, and the reflux ratio is 0.1-15; the operating pressure of the butanediol acetal/ketone product separation column is absolute.
  • reflux ratio of 0.01 to 10; butanediol acetal / ketone product hydrolysis distillation column operating pressure The absolute pressure gauge is 0.1-1010m, the reflux ratio is 0.01-20, or the gas is directly produced at the top of the column, and the molar ratio of water to the acetal/ketone product in the two feed streams is controlled to be 1 to 10, and the top is controlled to produce aldehyde.
  • the molar purity of the ketone is greater than 0.85; the operating pressure of the butanediol refining column is 0.01 to 2 atm and the reflux ratio is 0.01 to 10.
  • Feed S37 can be fed to the refining column (T41) of the 1,2-butanediol acetal/ketone product from a column at the top of the column or at one of the columns of the rectification column.
  • the refining column (T41) of the 1,2-butanediol acetal/ketone product was operated at an absolute pressure of 0.05 to 5 atm, and the column was used to control the purity of the acetal/ketone product.
  • the phase separator (D1) operates at a temperature of 20 to 150 °C.
  • Modification 2 the above process for separating a mixture of ethylene glycol and 1,2-butanediol, one or all of the ethylene glycol and butanediol acetal/ketone products are replaced by a hydrolysis process after separation, as shown in the attached drawing 3 is shown.
  • the ethylene glycol acetal/ketone product is hydrolyzed in the reactor (R1) at a constant temperature, and then enters the aldehyde/ketone separation column (T51); the aldehyde/ketone separation column separates a certain purity of the aldehyde/ketone S25, which can be recycled to the acetal.
  • / ketone reactor (R1) is used, the column reactor stream is taken from the stream S26 and enters the ethylene glycol acetal/ketone product recovery column (T52); the ethylene glycol acetal/ketone product recovery tower is topped without hydrolysis.
  • the ethylene glycol acetal/ketone product and the water azeotrope S32 are recycled to the reactor R2, and the mixture is taken from the mixed stream S33 of ethylene glycol and water to the ethylene glycol refining tower (T53);
  • the top of the refining tower is drained from water S39, and the bottom of the tower is obtained from the ethylene glycol product S40.
  • the reactor adopts a full-mixed reactor and a fixed-bed reactor.
  • the operating pressure is 0.5-10 atm
  • the reaction temperature is 40-200 ° C
  • the catalyst is controlled to enter the reactor using solid acid or solid alkali resin catalyst.
  • the molar ratio of the amount of water to the total amount of the acetal/ketone product is 1 to 10; the operating pressure of the aldehyde/ketone separation column is 0.5 to 10 atm at an absolute pressure, the reflux ratio is 0.01 to 15, or the gas phase at the top of the column is directly vaporized. Production; ethylene glycol acetal / ketone product recovery tower operating pressure is 0.1 ⁇ 10atm absolute, reflux ratio of 0.1 ⁇ 15; glycol refining tower operating pressure of 0.01 ⁇ 2atm absolute pressure, reflux ratio of 0.01 ⁇ 10.
  • a similar method can be used for the butanediol acetal/ketone product, which will not be described here.
  • Modification 3 the process for separating the mixture of ethylene glycol and 1,2-butanediol, the improvement 1, the improvement 2 combination use.
  • the method of the invention is applied to the reaction separation process of ethylene glycol and 1,2-butanediol mixture, as shown in Fig. 1, including acetal reactor (R1), aldehyde/ketone recovery tower (T11), ethylene glycol acetal /ketone product separation column (T12), butanediol acetal / ketone product separation column (T13), acetal / ketone hydrolysis distillation column (T21, T31), ethylene glycol refining tower (T22), 1,2- Butanediol refining column (T32), condenser, reboiler, pump and associated feed lines and lines connecting the above equipment.
  • R1 acetal reactor
  • T11 aldehyde/ketone recovery tower
  • T12 ethylene glycol acetal /ketone product separation column
  • T13 butanediol acetal / ketone product separation column
  • T21, T31 ethylene glycol refining tower
  • T32 1,
  • the diol mixed raw material S02 (ethylene glycol and 1,2-butanediol molar ratio 0.55: 0.45) and the reactant acetaldehyde S01 are added to the fully mixed tank reactor (R1), and the acetaldehyde and diol are fed into the reactor.
  • the molar ratio of the total amount of the material was 1, the reaction was carried out at 60 ° C, 1.5 atm absolute pressure, and the reaction residence time was controlled to 40 minutes, using a macroporous acidic cationic resin NKC-9 catalyst.
  • the reacted stream S04 enters the acetaldehyde recovery tower (T11).
  • the total theoretical number of the column is 13, the absolute pressure is 1.3 atm, the reflux ratio is 0.17, and the removed acetaldehyde is recycled to the reactor (R1).
  • the outflow stream S10 enters the ethylene glycol acetal product separation column (T12).
  • the ethylene glycol acetal product separation tower (T12) has a total theoretical plate number of 24, the operating pressure is normal pressure, and the reflux ratio is 5.7.
  • the azeotrope S16 of the ethylene glycol acetal product and water is taken from the top of the column to enter the acetal hydrolyzed distillation.
  • the column recovery stream S17 enters the butanediol acetal product separation column (T13).
  • the butanediol acetal product separation column (T13) has a total theoretical plate number of 11, the operating pressure is 0.1 atm in absolute pressure, and the reflux ratio is 3.06.
  • the azeotrope S37 of the butanediol acetal product and water is taken into the bottom.
  • An aldehyde hydrolysis rectification column (T31), and a diol stream S38 which is recovered as a reaction is recycled to the reactor (R1).
  • the acetal hydrolysis distillation column (T21) has a total theoretical plate number of 19, 7-14, and the added water S18 and the stream S16 are added from the ninth theoretical plate to control the total feed water and acetal product.
  • the molar ratio is 1.6:1, the operating pressure is 1atm in absolute pressure, and the reflux ratio is 0.5.
  • the purity of acetaldehyde is recovered by recycling at the top of the tower to control the content of acetal product in the tower, and the glycol is produced in the tower.
  • the mixed stream S24 with water enters the ethylene glycol refining column (T22).
  • the acetal hydrolysis distillation column (T31) has a total theoretical plate number of 19, 7-14 to set up the reaction section, and the stream S37 is added from the ninth theoretical plate. No additional water supply S40 is required, and the operating pressure is 1 atm, and the reflux is measured.
  • Ratio 0.4 the top of the tower controls the purity of acetaldehyde to be recycled, the tower kettle controls the content of the acetal product, and the tower kettle produces a mixed stream of 1,2-butanediol and water, S46, into the butanediol refining tower ( T32).
  • the total theoretical plate number of the ethylene glycol refining tower (T22) is 8, the operating pressure is 0.1 atm under the absolute pressure, and the reflux ratio is 0.2, and the purity of the ethylene glycol product is controlled by the tower.
  • the total theoretical plate number of the 1,2-butanediol refining tower (T32) was 8, the operating pressure was 0.1 atm in absolute pressure, and the reflux ratio was 0.1.
  • the column kettle controlled the purity of the 1,2-butanediol product.
  • ethylene glycol and 1,2-butanediol can be separated well.
  • the final product ethylene glycol has a mass content of 99.91%, and the 1,2-butanediol has a mass content of 99.0%.
  • the rate is over 99.5%.
  • the process of the invention is applied to a reactive distillation separation process of a mixture of ethylene glycol and 1,2-butanediol, as shown in Figure 2, for the improved procedure of Example 1, considering the acetal product of 1,2-butanediol Separated as a product without hydrolysis, including reactor (R1), aldehyde/ketone recovery column (T11), ethylene glycol acetal/ketone product separation column (T12), butanediol acetal/ketone product separation column ( T13), ethylene glycol acetal/ketone hydrolysis rectification column (T21), ethylene glycol refining tower (T22), 1,2-butanediol acetal product refining tower (T41), Phase separator (D1), condenser, reboiler, pump and associated feed lines and lines connecting the above equipment.
  • reactor (R1) aldehyde/ketone recovery column (T11), ethylene glycol acetal/ketone product separation column (T12), butan
  • the other operations of this embodiment 2 are the same as those of the first embodiment.
  • the butanediol acetal product separation column (T13) overhead production stream S37 enters the 1,2-butanediol acetal product refining column (T41) from the phase separator (D1).
  • the phase separator (D1) was phase-separated at 1 atm, 65 ° C, the aqueous phase was taken out, and the organic phase was passed from the top of the column to the 1,2-butanediol acetal product refining column (T41).
  • the total theoretical plate number of the 1,2-butanediol acetal refining tower is 9, the absolute pressure is 1 atm, the azeotrope of the 1,2-butanediol acetal product and water is produced at the top of the column, and the tower is produced. , 2-butanediol acetal product, controlling the purity of the 1,2-butanediol acetal product.
  • ethylene glycol and 1,2-butanediol can be separated well, and the production of 1,2-butanediol products can be realized.
  • the final product ethylene glycol has a mass content of 99.91%, 1, 2
  • the mass ratio of the butanediol acetal product (4-ethyl-2-methyl-1,3-dioxolane) was 99.94%.
  • the method of the present invention is applied to a reactive distillation separation process of a mixture of ethylene glycol and 1,2-butanediol, as shown in FIG. 3, which is a modified procedure of Example 1, and the acetal/ketone product is considered to be a method of separation after hydrolysis first.
  • the other operations of the third embodiment are the same as those of the first embodiment.
  • the reactors R2 and R3 were all carried out at 75 ° C, 2 atm absolute pressure, and the reaction residence time was controlled to 40 minutes, using a basic ion resin catalyst; after hydrolysis, the total number of theoretical plates of the aldehyde recovery tower (T51) was 13, and the operating absolute pressure was 1.3 atm.
  • ethylene glycol and 1,2-butanediol can be separated well.
  • the final product ethylene glycol has a mass content of 99.91%
  • 1,2-butanediol has a mass content of 99.52%
  • ethylene glycol and The recovery of butanediol was over 94%.
  • reaction separation method and the device using the ethylene glycol and 1,2-butanediol mixture of the present invention are feasible, and the boiling point is close, the relative volatility is small, and the azeotropic ethylene glycol and the azeotrope are formed.
  • the separation of the 2-butanediol mixture, the ethylene glycol, butanediol (or butanediol acetal/ketone) product has high purity and high ethylene glycol recovery rate throughout the separation process.

Abstract

涉及一种乙二醇和1,2-丁二醇的分离方法、工艺及装置，首先将乙二醇和1,2-丁二醇通过缩醛或者缩酮反应形成相应的缩醛/酮产物液体混合物，再将含乙二醇和1,2-丁二醇的缩醛/酮产物液体混合物经过一系列精馏塔进行分离，根据缩醛或者缩酮沸点差异，通过精馏实现不同缩醛/酮产物的分离，然后分别水解缩醛/酮产物获得乙二醇和1,2-丁二醇初品，最后分别经过精馏对各自初品进行提纯，得到乙二醇和1,2-丁二醇产品。主要产品乙二醇产品纯度可以达到99.9％以上，回收率可以达到99.5％以上，1,2-丁二醇产品纯度可以达到98.5％以上，采用可逆反应转化的方法，变因沸点接近、相对挥发度小且有共沸的乙二醇和1,2-丁二醇分离难题为容易分离的缩醛/酮产物的分离问题。

Description

乙二醇和1,2-丁二醇的分离方法、工艺及装置 技术领域
本发明属于化工精馏分离领域，涉及一种乙二醇和1,2-丁二醇的分离方法、工艺及装置。
背景技术
乙二醇作为一种重要的基本有机化工原料，主要用于生产聚酯纤维、防冻剂、不饱和聚酯树脂、润滑剂、增塑剂、非离子表面活性剂以及炸药等，此外，乙二醇还可用于涂料、照相显影液、刹车液以及油墨等行业，用作过硼酸铵的溶剂和介质，用于生产特种溶剂乙二醇醚等，用途十分广泛。而且有研究指出，乙二醇在未来液体燃料电池领域有很大的发展潜力。我国是乙二醇的消费大国，国内的乙二醇的产量远远不能满足自身需求，2014年，我国乙二醇产量为350万吨，而表观消费量达1225万吨，因此，我国的乙二醇产业具有良好的发展前景。
目前，国内外大型乙二醇生产都采用直接水合法或加压水合法工艺线路，该工艺是将环氧乙烷和水按1：20～22(摩尔比)配成混合水溶液，在固定床反应器中于130～180℃，1.0～2.5MPa下反应18～30分钟，环氧乙烷全部转化为混合醇，然后经过多效蒸发器脱水浓缩和减压精馏分离得到乙二醇。蒸发脱水需要消耗大量的能量，为了降低成本，国内外的一些主要的生产乙二醇的大公司以及科研院所开始研究催化水合法。同时，也有研究针对碳酸乙烯酯法，由环氧乙烷和二氧化碳合成碳酸乙烯酯，再以碳酸乙烯酯水解得到乙二醇，专利US 4508927、US 4500559和JP 571006631针对碳酸乙烯酯法提出了不同的工艺线路。随着石油资源的日趋紧张，世界油价波动较大，结合我国贫油、少气、富煤的资源情况，在我国发展C1化工具有非常重要的意义，从而可以减少对石油进口的依赖，减轻环境压力。以煤或天然气为原料制备合成气，合成气通过偶联制草酸酯，草酸酯通过加氢获得乙二醇这一工艺被认为是最具有工业前景的线路，国内目前建成投产和在建的煤制乙二醇装置超过20套。在草酸酯加氢制乙二醇的反应产物中，除了含有甲醇、乙醇酸酯等沸点较低的物质外，还含有少量1,2-丙二醇，1,2-丁二醇等与乙二醇沸点接近、通过普通精馏难以分离的物质，其中，1,2-丁二醇与乙二醇沸点最为接近，同时形成共沸，因此最难分离。时下煤制乙二醇工艺中乙二醇精制塔通过塔顶甩料的方式提纯乙二醇，为了保障高纯乙二醇的回收率，不得不采用很多的理论板和高回流比，从而造成压降过高，塔釜温度过高，因而容易影响塔釜乙二醇的品质。如何有效的分离乙二醇和1,2-丁二醇，降低乙二醇精制塔塔板数和回流比，提高乙二醇收率，回收高附加值的1,2-丁二醇，是降低生产成本和提高经济效益的关键。另外，作为可续发展的能源线路，以生物质为原料生产乙二醇等二元醇路线也越来越受到各国的广泛关注，如中科院大连化学物理研究所张涛院士课题组所研究的生物质催化法制备乙二醇。要想获得高纯度的二醇化工产品，也会涉及到乙二醇和1,2-丁二醇等的分离。
据检索，发现如下与本申请相关的专利文献，具体公开内容如下：
1、专利CN 101928201公开了一种煤制乙二醇粗产品的提纯工艺，提纯工艺通过皂化反 应、去甲醇、加氢反应、三塔精馏以及吸附处理，从煤制乙二醇粗产品中获得高纯度的乙二醇产品，然而1,2-丁二醇仅仅通过与乙二醇的共沸去除，不仅造成1,2-丁二醇这一高附加值的产品损失，同时还影响了乙二醇的收率。
2、专利CN 102276418提出使用针对来自草酸酯加氢制乙二醇产品的工艺的乙二醇和1,2-丁二醇的混合液，通过采用ZSM-5、Y型沸石或β沸石的至少一种吸附剂，吸附1,2-丁二醇从而获得乙二醇的纯化，然而没有指出乙二醇的回收率，对于1,2-丁二醇的回收利用也没有涉及。
3、专利CN 103193594公开了一种使用与乙二醇不互溶的二氧五环类化合物作为共沸剂，采用共沸精馏的方法从草酸酯加氢液相产物中分离提纯乙二醇的方法，液相混合物经过两个脱轻塔脱除甲醇、乙醇酸甲酯等低沸点化合物后进入共沸精馏塔，塔顶采出共沸剂与乙二醇共沸物，分相后共沸剂回流塔顶，乙二醇进入乙二醇精制塔，该方法随原料的组成不同，乙二醇的回收率变化较大，从99.1％下降到89.65％，同时1,2-丁二醇的回收也欠佳考虑。
4、专利CN 102372596提出通过使用共沸精馏分离乙二醇、丙二醇、丁二醇的分离工艺，合成气制乙二醇产物经过馏分切割塔脱出轻组分后，经过共沸精馏塔(2)，塔顶通过采出富含1,2-丙二醇的乙二醇溶液除去丙二醇，塔釜物流进入下一个共沸精馏塔(3)，塔釜采出富含1,2-丁二醇的乙二醇溶液，塔顶分相后共沸剂回到塔顶，采出乙二醇物流进入乙二醇精致塔，同样可以看到，通过甩料的方式脱出丙二醇和丁二醇，不能保障乙二醇的高回收率，同时乙二醇作为含量大的组分，与共沸剂从塔顶共沸分离的操作对节能不利。
综上所述，目前在本领域中急需一种能够有效分离乙二醇和1,2-丁二醇的方法，回收高附加值的1,2-丁二醇，提高乙二醇的回收率和品质，操作简单，成本低廉。
发明内容
本发明的目的是提供一种乙二醇和1,2-丁二醇的分离精制方法、工艺及装置，应用该方法于煤制乙二醇产物分离工艺中或者生物质催化制乙二醇工艺中，具有分离能耗低，目标产物乙二醇回收率高、纯度高，同时可以回收高附加值的1,2-丁二醇等优点，尤其是针对形成共沸的乙二醇和1,2-丁二醇混合物的分离，整个分离工艺过程操作简单、能耗较小，而且乙二醇，丁二醇(或者丁二醇缩醛/酮)产品纯度高，乙二醇回收率高。
本发明解决其技术问题所采用的技术方案是：
一种乙二醇和1,2-丁二醇混合物的分离精制方法，首先将乙二醇和1,2-丁二醇与反应剂醛/酮通过缩醛或者缩酮反应形成相应的缩醛/酮产物液体混合物，回收乙二醇和1,2-丁二醇的缩醛/酮产物液体混合物中的醛/酮，液体混合物再通过分离塔，通过精馏实现不同缩醛/酮产物的分离，然后水解缩醛/酮产物获得乙二醇和1,2-丁二醇初品，最后分别经过精馏对各自初品进行提纯，得到乙二醇和1,2-丁二醇产品。
而且，所述反应剂醛/酮为含1～8个碳原子的醛的一种，或者含3～8个碳原子的酮的一种。
而且，所述反应剂为乙醛或丙酮。
而且，水解后的乙二醇和1,2-丁二醇初品再进行乙二醇和1,2-丁二醇精制，进一步提纯乙二醇和1,2-丁二醇。
一种乙二醇和1,2-丁二醇混合物的分离精制工艺，方法包含以下步骤：
⑴含乙二醇和1,2-丁二醇的物流(S01)与反应剂醛或者酮(S02)进入到反应器(R1)中反应，采出液体混合反应产物(S03)；
⑵液体混合反应产物(S03)进入醛/酮回收塔(T11)，塔顶排出主要包含未反应的醛/酮的物流(S09)，循环使用，塔釜采出缩醛/酮反应的生成物和未反应的二醇混合物流S10；
⑶塔釜物流(S10)进入乙二醇缩醛/酮产物分离塔(T12)，塔顶采出乙二醇缩醛/酮产物与水的共沸物(S16)，塔釜采出含1,2-丁二醇缩醛/酮产物的流股(S17)；
⑷塔釜采出物流(S17)进入1,2-丁二醇缩醛/酮产物分离塔(T13)，塔顶采出丁二醇缩醛/酮产物与水的混合物流股(S37)，塔釜采出为反应的二醇混合物流股(S38)，循环至反应器(R1)；
⑸步骤⑶塔顶采出物流(S16)与补充的流股水(S18)进入缩醛/酮水解精馏塔(T21)，塔中部设置反应段，塔顶排出水解产物醛/酮(S23)，循环使用，塔釜采出含水的乙二醇流股(S24)；
⑹步骤⑸塔釜采出物流(S24)进入乙二醇精制塔(T22)，塔顶排除水(S30)，循环使用，塔釜采出乙二醇产品(S31)；
⑺步骤⑷塔顶采出物流(S37)与补充的流股水(S40)进入另外一个缩醛/酮水解精馏塔(T31)，塔中部设置反应段，塔顶排出水解产物醛/酮(S45)，循环使用，塔釜采出含水的1,2-丁二醇流股(S46)；
⑻步骤⑺塔釜采出物流(S46)进入1,2-丁二醇精制塔(T32)，塔顶排除水(S52)，循环使用，塔釜采出1,2-丁二醇产品(S53)。
而且，上述步骤的反应条件如下：
步骤⑴中的反应器R1采用釜式反应器、固定床反应器的一种，操作压力以绝对压力计为0.5～10atm，反应温度40～200℃，催化采用固体酸或固体碱树脂催化剂的一种或者混合，进入反应器中的醛/酮进料量与二醇进料总量的摩尔比为1～10；
步骤⑵的醛/酮回收塔T11以绝对压力计的操作压力为0.1～10atm，回流比0.01～15，或者直接气相采出，不回流；
步骤⑶的乙二醇缩醛/酮产物分离塔T12以绝对压力计的操作压力为0.5～10atm，回流比0.1～15；
步骤⑷的丁二醇缩醛/酮产物分离塔T13以绝对压力计的操作压力为0.01～5atm，回流比0.1～15；
步骤⑸的缩醛/酮水解精馏塔T21以绝对压力计的操作压力为0.1～10atm，回流比0.01～20，或者塔顶直接气相采出，不回流，补充的水S18与S16中的水总进料量与缩醛/酮的总进料量的摩尔比为1～10；
步骤⑹的乙二醇精制塔T22以绝对压力计的操作压力为0.01～2atm，回流比0.01～10；
步骤⑺的缩醛/酮水解精馏塔T31以绝对压力计的操作压力为0.1～10atm，回流比0.01～20，或者塔顶直接气相采出，补充的水S40与S37中的水总进料量与缩醛/酮的总进料量的摩尔比 为1～10；
步骤⑻的1,2-丁二醇精制塔T32以绝对压力计的操作压力为0.01～2atm，回流比0.01～10。
而且，所述步骤⑸和步骤⑺的缩醛/酮产物的水解采用先反应后分离的方式进行，物流(S16)与(S46)分别进入各自的水解单元，水解单元包含一个反应器(R2或R3)，一个醛/酮分离塔(T51或T61)，一个缩醛/酮产物回收塔(T52或T62)，产品精制塔(T53或T63)。
而且，所述反应器(R2或R3)操作压力以绝对压力计为0.5～10atm，反应温度40～200℃，催化采用固体酸或固体碱树脂催化剂的一种或者混合；醛/酮分离塔(T51，T61)操作压力以绝压计为0.5～10atm，回流比0.01～15，或者塔顶直接气相采出；缩醛/酮产物回收塔(T52，T62)操作压力以绝压计为0.1～10atm，回流比0.1～15；产品精制塔(T53，T63)操作压力以绝压计为0.01～2atm，回流比0.01～10。
而且，1,2-丁二醇的缩醛/酮产物不进行水解，直接分离获得缩醛/酮产品，如附图2所示的塔T41，塔顶设置分相器，操作温度20～150℃，水相排出，有机相回流，塔T41以绝对压力计为0.05～5atm。
一套乙二醇和1,2-丁二醇混合物的分离精制装置，装置包含反应器(R1)、醛/酮回收塔(T11)、乙二醇缩醛/酮产物分离塔(T12)、丁二醇缩醛/酮产物分离塔(T13)，缩醛/酮水解精馏塔(T21、T31)、乙二醇精制塔(T22)、1,2-丁二醇精制塔(T32)、冷凝器、再沸器、泵以及相关的进料管线和连接以上设备的管线；所有精馏塔塔顶和塔釜均设置冷凝器和再沸器。反应器(R1)出料进入、醛/酮回收塔(T11)，塔顶采出未反应的醛/酮泵送回反应器(R1)，乙二醇缩醛/酮产物分离塔(T12)；乙二醇缩醛/酮产物分离塔(T12)塔顶采出乙二醇缩醛/酮产物与水的混合物进入乙二醇缩醛/酮产物水解精馏塔(T21)，塔釜采出流股进入丁二醇缩醛/酮产物分离塔(T13)；丁二醇缩醛/酮产物分离塔(T13)塔顶采出丁二醇缩醛/酮产物与水的混合物进入丁二醇缩醛/酮产物水解精馏塔(T31)，塔釜采出未反应的二醇混合物循环使用；乙二醇缩醛/酮产物水解精馏塔(T21)中部设置反应段，塔顶采出醛/酮产物，用于循环使用，塔釜采出乙二醇与水的混合物进入乙二醇精制塔(T22)；乙二醇精制塔(T22)塔顶脱除水，塔釜为乙二醇产品；丁二醇缩醛/酮产物水解精馏塔(T31)中部设置反应段，塔顶采出醛/酮产物，用于循环使用，塔釜采出丁二醇与水的混合物进入丁二醇精制塔(T32)；丁二醇精制塔(T32)塔顶脱除水，塔釜为1,2-丁二醇产品。
本发明的优点和有益效果如下：
1、本发明提供的分离精制方法能够将处于共沸状态乙二醇和1,2-丁二醇进行高效的分离，彻底解决了乙二醇和1,2-丁二醇因沸点接近而不能彻底分离的难题，获得高纯度的乙二醇产品、1,2-丁二醇产品或者其缩醛/酮产品。
2、本发明分离精制方法最显著的两点是节能和产品的回收产品品质，与现有通过单独精馏方法分离乙二醇和1,2-丁二醇相比，本发明通过转化成缩醛后再生成二醇，其能耗节省50-60％，回收产品品质高，回收的乙二醇为聚酯级，可以直接去用于聚酯生产，生产的1,2-丁二醇的纯度高达99％以上，也可以直接用于后续生产。
3、本发明提供的分离精制方法中缩合过程可以应用反应精馏方法，有机的耦合了反应与精馏过程，既节省了设备投资，又减少了后续分离的操作费用。
4、本发明中主要产品乙二醇产品纯度可以达到99.9％以上，回收率可以达到99.5％以上，1,2-丁二醇产品纯度可以达到98.5％以上(1,2-丁二醇缩醛/酮产品纯度可以达到99.6％以上)。本发明的优点是采用反应精馏的方法，变因沸点接近、相对挥发度小且有共沸的乙二醇和1,2-丁二醇分离难题为容易分离的缩醛/酮产物的分离问题，有效的实现了乙二醇和1,2-丁二醇的分离。
5、本发明提供的分离精制方法应用于草酸酯法制乙二醇工艺中，可以减轻乙二醇分离纯化塔的负荷，提高乙二醇的品质和收率，同时回收高附加值产品1,2-丁二醇，实现很好的年经济效益。
6、本发明提供的分离精制方法可以很好的实现沸点很接近，且形成共沸的乙二醇和1,2-丁二醇混合物的分离，整个分离工艺过程操作简单、能耗较小，而且乙二醇，丁二醇(或者丁二醇缩醛/酮)产品纯度高，乙二醇回收率高。
附图说明
图1为乙二醇和1,2-丁二醇混合物反应分离提纯工艺流程示意图。
图2为改进的乙二醇和1,2-丁二醇混合物反应分离提纯工艺流程示意图。
图3为改进的乙二醇和1,2-丁二醇混合物反应分离提纯工艺流程示意图。
图4为改进的乙二醇和1,2-丁二醇混合物反应精馏分离提纯工艺流程示意图。
具体实施方式
下面结合附图并通过具体实施例对本发明作进一步详述，以下实施例只是描述性的，不是限定性的，不能以此限定本发明的保护范围。
本发明提供的乙二醇和1,2-丁二醇混合物的分离精制方法，进料来源于煤制乙二醇工艺或者生物质催化制乙二醇工艺，首先将乙二醇和1,2-丁二醇通过缩醛或者缩酮反应形成相应的缩醛/酮产物液体混合物，再将含乙二醇和1,2-丁二醇的缩醛/酮产物液体混合物经过回收塔，回收未反应的醛/酮，液体混合物再通过分离塔，根据缩醛或者缩酮沸点差异，通过精馏实现不同缩醛/酮产物的分离，然后水解缩醛/酮产物获得乙二醇和1,2-丁二醇，再对乙二醇和1,2-丁二醇进行精制。使用的反应为缩醛或者缩酮反应，反应剂使用甲醛、乙醛、丙醛、丁醛等醛或者丙酮、丁酮、环己酮等酮的一种。尤其是乙醛和丙酮效果较好。
一种乙二醇和1,2-丁二醇混合物的分离精制工艺，通过缩醛或者缩酮反应，把因沸点接近、相对挥发度小且有共沸的乙二醇和1,2-丁二醇分离难题转化为容易实现分离的缩醛/酮产物，从而实现他们的有效分离，方法包含以下步骤：
⑴含乙二醇和1,2-丁二醇的物流S01与反应剂醛或者酮S02进入到反应器R1中反应，采出反应产物S03；
⑵物流S03进入醛/酮回收塔T11，塔顶排出主要包含未反应的醛/酮的物流S09，循环使用，塔釜采出缩醛/酮反应的生成物和未反应的二醇混合物流S10；
⑶物流S10进入乙二醇缩醛/酮产物分离塔T12，塔顶采出乙二醇缩醛/酮产物与水的共沸 物S16，塔釜采出含1,2-丁二醇缩醛/酮产物的流股S17；
⑷物流S17进入丁二醇缩醛/酮产物分离塔T13，塔顶采出丁二醇缩醛/酮产物与水的混合物流股S37，塔釜采出为反应的二醇混合物流股S38，循环至反应器R1；
⑸物流S16与补充的流股水S18进入缩醛/酮水解精馏塔T21，塔中部设置反应段，塔顶排出水解产物醛/酮S23，循环使用，塔釜采出含水的乙二醇流股S24；
⑹物流S24进入乙二醇精制塔T22，塔顶排除水S30，可以循环使用，塔釜采出乙二醇产品S31；
⑺物流S37与补充的流股水S40进入缩醛/酮水解精馏塔T31，塔中部设置反应段，塔顶排出水解产物醛/酮S45，循环使用，塔釜采出含水的1,2-丁二醇流股S46；
⑻物流S46进入乙二醇精制塔T32，塔顶排除水S52，可以循环使用，塔釜采出乙二醇产品S53。
本申请还根据具体方案，提供上述步骤的较好的、收率较高的工艺参数，具体如下：
步骤⑴中的反应器R1采用釜式反应器、固定床反应器的一种，操作压力以绝对压力计为0.5～10atm，反应温度40～200℃，催化采用固体酸或固体碱树脂催化剂的一种或者混合，进入反应器中的醛/酮进料量与二醇进料总量的摩尔比为1～10。
步骤⑵的醛/酮回收塔T11以绝对压力计的操作压力为0.1～10atm，回流比0.01～15，或者直接气相采出，不回流。
步骤⑶的乙二醇缩醛/酮产物分离塔T12以绝对压力计的操作压力为0.5～10atm，回流比0.1～15。
步骤⑷的丁二醇缩醛/酮产物分离塔T13以绝对压力计的操作压力为0.01～5atm，回流比0.1～15。
步骤⑸的缩醛/酮水解精馏塔T21以绝对压力计的操作压力为0.1～10atm，回流比0.01～20，或者塔顶直接气相采出，不回流，补充的水S18与S16中的水总进料量与缩醛/酮的总进料量的摩尔比为1～10。
步骤⑹的乙二醇精制塔T22以绝对压力计的操作压力为0.01～2atm，回流比0.01～10。
步骤⑺的缩醛/酮水解精馏塔T31以绝对压力计的操作压力为0.1～10atm，回流比0.01～20，或者塔顶直接气相采出，补充的水S40与S37中的水总进料量与缩醛/酮的总进料量的摩尔比为1～10。
步骤⑻的1,2-丁二醇精制塔T32以绝对压力计的操作压力为0.01～2atm，回流比0.01～10。
还可以将1,2-丁二醇的缩醛/酮产物不进行水解，直接分离获得缩醛/酮产品，如附图2所示的塔T41，塔顶设置分相器，操作温度20～150℃，水相排出，有机相回流，塔T41以绝对压力计为0.05～5atm。
本发明还提供一种乙二醇和1,2-丁二醇混合物的分离精制设备，可以按照下述设备连接操作进行：
其装置主要包含反应器(R1)、醛/酮回收塔(T11)、乙二醇缩醛/酮产物分离塔(T12)、丁二醇缩醛/酮产物分离塔(T13)，缩醛/酮水解精馏塔(T21、T31)、乙二醇精制塔(T22)、 1,2-丁二醇精制塔(T32)、冷凝器、再沸器、泵以及相关的进料管线和连接以上设备的管线；所有精馏塔塔顶和塔釜均设置冷凝器和再沸器。反应器(R1)出料进入、醛/酮回收塔(T11)，塔顶采出未反应的醛/酮泵送回反应器(R1)，乙二醇缩醛/酮产物分离塔(T12)；乙二醇缩醛/酮产物分离塔(T12)塔顶采出乙二醇缩醛/酮产物与水的混合物进入乙二醇缩醛/酮产物水解精馏塔(T21)，塔釜采出流股进入丁二醇缩醛/酮产物分离塔(T13)；丁二醇缩醛/酮产物分离塔(T13)塔顶采出丁二醇缩醛/酮产物与水的混合物进入丁二醇缩醛/酮产物水解精馏塔(T31)，塔釜采出未反应的二醇混合物循环使用；乙二醇缩醛/酮产物水解精馏塔(T21)中部设置反应段，塔顶采出醛/酮产物，用于循环使用，塔釜采出乙二醇与水的混合物进入乙二醇精制塔(T22)；乙二醇精制塔(T22)塔顶脱除水，塔釜为乙二醇产品；丁二醇缩醛/酮产物水解精馏塔(T31)中部设置反应段，塔顶采出醛/酮产物，用于循环使用，塔釜采出丁二醇与水的混合物进入丁二醇精制塔(T32)；丁二醇精制塔(T32)塔顶脱除水，塔釜为1,2-丁二醇产品。
本发明的工艺方法，乙二醇、1,2-丁二醇的混合物(S02)和反应原料醛或者酮一种(S01)进入缩醛/酮反应器(R1)；反应后获得液相(S03)经过进料泵(P1)得到液相(S04)，作为醛/酮回收塔(T11)的进料；醛/酮回收塔(T11)塔顶气相(S05)经过冷凝器(E2)冷凝后液相(S06)经过回流采出泵(P3)得到液相(S07)，一部分作为醛/酮回收塔的回流(S08)，另外一部分作为循环醛/酮(S09)返回缩醛/酮反应器(R1)进料,塔釜采出液相(S10)经进料泵(P2)得到液相(S11)进入乙二醇缩醛/酮产物分离塔(T12)；乙二醇缩醛/酮产物分离塔(T12)塔顶气相(S12)经过冷凝器(E4)冷凝后液相(S13)经回流采出泵(P5)得到液相(S14),一部分作为精馏塔塔顶回流(S15)，另一部分作为乙二醇缩醛/酮产物水解精馏塔(T21)的进料(S16)，塔釜采出液相(S17)经进料泵(P4)得到液相(S32)进入丁二醇缩醛/酮产物分离塔(T13)；液相水(S18)进入乙二醇缩醛/酮产物水解精馏塔(T21)；乙二醇缩醛/酮产物水解精馏塔(T21)塔顶气相(S19)经过冷凝器(E8)冷凝后液相(S20)经回流采出泵(P7)得到液相(S21)，一部分作为水解精馏塔的塔顶回流(S22)，另一部分是采出的醛或者酮流股(S23)，用于循环使用，塔釜采出液相(S24)经进料泵(P6)得到液相(S25)作为乙二醇精制塔(T22)的进料；乙二醇精制塔(T22)塔顶气相(S26)经冷凝器(E10)冷凝后的液相(S27)经回流采出泵(P8)得到液相(S28)，一部分作为乙二醇精制塔的回流(S29)，另一部分液相(S30)采出，塔釜采出液相乙二醇产品流股(S31)；丁二醇缩醛/酮产物分离塔(T13)塔顶气相(S33)经冷凝器(E6)冷凝后的液相(S34)经回流采出泵(P10)得到液相(S35)，一部分作为缩醛/酮精馏塔的回流(S36)，另一部分液相(S37)采出作为丁二醇缩醛/酮产物水解塔(T31)的进料，塔釜采出未反应的二醇流股(S38)经过泵(P9)得到液相(S39)，循环至反应器(R1)；液相水(S40)进入丁二醇缩醛/酮产物水解精馏塔(T31)；丁二醇缩醛/酮产物水解精馏塔(T31)塔顶气相(S41)经冷凝器(E12)冷凝后的液相(S42)经回流采出泵(P12)得到液相(S43)，一部分作为水解精馏塔的塔顶回流(S44),另一部分是采出的醛或者酮流股(S45)，用于循环使用，塔釜采出液相(S46)经进料泵(P11)得到液相(S47)作为丁二醇精制塔(T32)的进料；丁二醇 精制塔(T32)塔顶气相(S48)经冷凝器(E14)冷凝后液相(S49)经回流采出泵(P13)得到液相(S50),一部分作为精制塔塔顶的回流(S51)，另一部分液相(S52)采出，塔釜采出产品1,2-丁二醇(S53)。
在上述技术方案中，反应器采用全混釜式反应器、固定床反应器的一种，操作压力以绝压计为0.5～10atm，反应温度40～200℃，催化剂使用固体酸或固体碱树脂催化剂，控制进入反应器中的醛/酮的量与二醇的总量的摩尔配比为1～10；醛/酮回收塔操作压力以绝压计为0.1～10atm，回流比0.01～15，或者塔顶直接气相采出；乙二醇缩醛/酮产物分离塔操作压力以绝压计为0.5～10atm，回流比0.1～15；丁二醇缩醛/酮产物分离塔操作压力以绝压计为0.01～5atm，回流比0.1～15；乙二醇缩醛/酮产物水解精馏塔操作压力以绝压计为0.1～10atm，回流比0.01～20，或者塔顶直接气相采出，控制两股进料流股中水和缩醛/酮产物的摩尔比1～10，塔顶控制采出醛或者酮的摩尔纯度大于0.85；乙二醇精制塔操作压力以绝压计为0.01～2atm，回流比0.01～10；丁二醇缩醛/酮产物水解精馏塔操作压力以绝压计为0.1～10atm，回流比0.01～20，或者塔顶直接气相采出，控制两股进料流股中水和缩醛/酮产物的摩尔比1～10，塔顶控制采出醛或者酮的摩尔纯度大于0.85；丁二醇精制塔操作压力以绝压计为0.01～2atm，回流比0.01～10。
改进1，上述所述分离乙二醇和1,2-丁二醇混合物的工艺方法，1,2-丁二醇的缩醛/酮产物不作水解，而是通过塔顶带分相器的简单精馏塔分离获得缩醛/缩酮产品，如附图2所示；液相流股(S37)作为1,2-丁二醇缩醛/酮产物的精制塔(T41)的进料，塔顶气相采出冷凝后进入分相器，有机相作为回流，采出水相(S45)，塔釜液相采出缩醛/酮产品(S46)。进料S37可以从塔顶分相器或精馏塔的某一塔板处进入1,2-丁二醇缩醛/酮产物的精制塔(T41)。1,2-丁二醇缩醛/酮产物的精制塔(T41)操作绝对压力0.05～5atm，塔釜控制缩醛/酮产品的纯度。分相器(D1)操作温度20～150℃。
改进2，上述所述分离乙二醇和1,2-丁二醇混合物的工艺方法，乙二醇和丁二醇缩醛/酮产物的一种或全部采用反应后分离的水解工艺代替，如附图3所示。乙二醇缩醛/酮产物分离塔(T12)塔顶采出的乙二醇缩醛/酮产物与水的混合物S16进入水解单元，该单元由一个反应器(R2)和三个精馏塔(T51，T52，T53)组成。乙二醇缩醛/酮产物在反应器(R1)中恒温水解，然后进入醛/酮分离塔(T51)；醛/酮分离塔顶分离出一定纯度的醛/酮S25，可以循环至缩醛/酮反应器(R1)使用，塔釜流股采出流股S26进入到乙二醇缩醛/酮产物回收塔(T52)；乙二醇缩醛/酮产物回收塔塔顶采出未水解的乙二醇缩醛/酮产物和水的共沸物S32，循环至反应器R2，塔釜采出乙二醇和水的混合流股S33进入到乙二醇精制塔(T53)；乙二醇精制塔塔顶排除水S39，塔釜获得乙二醇产品S40。反应器采用全混釜式反应器、固定床反应器的一种，操作压力以绝压计为0.5～10atm，反应温度40～200℃，催化剂使用固体酸或固体碱树脂催化剂，控制进入反应器中的水的量与缩醛/酮产物的总量的摩尔配比为1～10；醛/酮分离塔操作压力以绝压计为0.5～10atm，回流比0.01～15，或者塔顶直接气相采出；乙二醇缩醛/酮产物回收塔操作压力以绝压计为0.1～10atm，回流比0.1～15；乙二醇精制塔操作压力以绝压计为0.01～2atm，回流比0.01～10。丁二醇缩醛/酮产物可采用类似的方法，这里不再叙述。
改进3，述所述分离乙二醇和1,2-丁二醇混合物的工艺方法，改进1、改进2组合使用。
上述所有精馏塔、分离塔塔釜均安装有再沸器(E1、E3、E7、E9、E11、E7、13)，用于塔釜物料的再沸。
下面通过具体实施方式来说明本申请的具体工艺流程和推荐的操作参数。
实施例1
将本发明方法用于乙二醇和1,2-丁二醇混合物反应分离过程，如图1所示，包括缩醛反应器(R1)、醛/酮回收塔(T11)、乙二醇缩醛/酮产物分离塔(T12)、丁二醇缩醛/酮产物分离塔(T13)，缩醛/酮水解精馏塔(T21、T31)、乙二醇精制塔(T22)、1,2-丁二醇精制塔(T32)、冷凝器、再沸器、泵以及相关的进料管线和连接以上设备的管线。二醇混合原料S02(乙二醇和1,2-丁二醇摩尔配比0.55：0.45)与反应剂乙醛S01加入全混釜式反应器(R1)，反应器中乙醛与二醇的进料总量摩尔配比为1，反应在60℃，1.5atm绝对压力进行，反应停留时间控制为40分钟，采用大孔酸性阳离子树脂NKC-9催化剂。反应后的流股S04进入乙醛回收塔(T11)，该塔总理论板数13，操作绝对压力为1.3atm，回流比0.17，脱除的乙醛循环至反应器(R1),塔釜采出流股S10进入乙二醇缩醛产物分离塔(T12)。乙二醇缩醛产物分离塔(T12)总理论板数24，操作压力为常压，回流比5.7，塔顶采出乙二醇缩醛产物与水的共沸物S16进入缩醛水解精馏塔(T21)，塔釜采出流股S17进入丁二醇缩醛产物分离塔(T13)。丁二醇缩醛产物分离塔(T13)总理论板数11，操作压力以绝压计为0.1atm，回流比3.06，塔顶采出丁二醇缩醛产物与水的共沸物S37进入缩醛水解精馏塔(T31)，塔釜采出为反应的二醇流股S38循环至反应器(R1)。缩醛水解精馏塔(T21)总理论板数19，7-14设置反应段，添加的水S18与流股S16均从第9块理论板加入，控制总进料的水与缩醛产物的摩尔配比为1.6：1，操作压力以绝压计为1atm，回流比0.5，塔顶控制采出乙醛的纯度以循环使用，塔釜控制缩醛产物的含量，塔釜采出乙二醇与水的混合流股S24进入乙二醇精制塔(T22)。缩醛水解精馏塔(T31)总理论板数19，7-14设置反应段，流股S37从第9块理论板加入，不需要额外补充水S40，操作压力以绝压计为1atm，回流比0.4，塔顶控制采出乙醛的纯度以循环使用，塔釜控制缩醛产物的含量，塔釜采出1,2-丁二醇与水的混合流股S46进入丁二醇精制塔(T32)。乙二醇精制塔(T22)总理论板数8，操作压力以绝压计为0.1atm，回流比0.2，塔釜控制乙二醇产品的纯度。1,2-丁二醇精制塔(T32)总理论板数8，操作压力以绝压计为0.1atm，回流比0.1，塔釜控制1,2-丁二醇产品的纯度。
经上述过程后，乙二醇和1,2-丁二醇可以很好的实现分离，最终产物乙二醇的质量含量为99.91％，1,2-丁二醇的质量含量为99.0％，总回收率均超过99.5％。
实施例2
将本发明方法用于乙二醇和1,2-丁二醇混合物反应精馏分离过程，如图2所示，为实施例1的改进流程，考虑将1,2-丁二醇的缩醛产物分离出来作为产品，而不水解，包括反应器(R1)、醛/酮回收塔(T11)、乙二醇缩醛/酮产物分离塔(T12)、丁二醇缩醛/酮产物分离塔(T13)，乙二醇缩醛/酮水解精馏塔(T21)、乙二醇精制塔(T22)、1,2-丁二醇缩醛产品精制塔(T41)、 分相器(D1)，冷凝器、再沸器、泵以及相关的进料管线和连接以上设备的管线。该实施例2其他操作与实施例1相同。丁二醇缩醛产物分离塔(T13)塔顶采出流股S37从分相器(D1)进入1,2-丁二醇缩醛产品精制塔(T41)。分相器(D1)在1atm，65℃下分相，水相采出，有机相从塔顶进入1,2-丁二醇缩醛产品精制塔(T41)。1,2-丁二醇缩醛产品精制塔总理论板数为9，操作绝对压力1atm，塔顶采出1,2-丁二醇缩醛产物与水的共沸物，塔釜采出1,2-丁二醇缩醛产品，控制塔釜1,2-丁二醇缩醛产品的纯度。
经上述过程后，乙二醇和1,2-丁二醇可以很好的实现分离，并实现1,2-丁二醇产品的生产，最终产物乙二醇的质量含量为99.91％，1,2-丁二醇缩醛产品(4-乙基-2-甲基-1,3-二氧戊环)的质量含量为99.94％。
实施例3
将本发明方法用于乙二醇和1,2-丁二醇混合物反应精馏分离过程，如图3所示，为实施例1的改进流程，缩醛/酮产物考虑先水解后分离方法，该实施例3其他操作与实施例1相同。反应器R2、R3均在75℃，2atm绝对压力进行，反应停留时间控制为40分钟，采用碱性离子树脂催化剂；水解后醛回收塔(T51)总理论板数13，操作绝对压力为1.3atm，回流比5；水解后醛回收塔(T61)总理论板数13，操作绝对压力为1.3atm，回流比5；未水解缩醛回收塔(T52)总理论板数12，操作绝对压力为1atm，回流比5.7；解缩醛回收塔(T62)总理论板数12，操作绝对压力为1atm，回流比4；乙二醇精制塔(T53)总理论板数8，操作绝对压力为0.2atm，回流比0.4；1,2-丁二醇缩醛产品精制塔(T63)总理论板数8，操作绝对压力为0.2atm，回流比0.4。
经上述过程后，乙二醇和1,2-丁二醇可以很好的实现分离，最终产物乙二醇的质量含量为99.91％，1,2-丁二醇的质量含量99.52％，乙二醇和丁二醇的回收率均超过94％。
由以上实施例可见，利用本发明的乙二醇和1,2-丁二醇混合物反应分离方法与装置可行，可以很好的实现沸点接近，相对挥发度小，且形成共沸的乙二醇和1,2-丁二醇混合物的分离，整个分离工艺过程乙二醇，丁二醇(或者丁二醇缩醛/酮)产品纯度高，乙二醇回收率高。
本发明提出的乙二醇和1,2-丁二醇混合物反应分离方法与装置，已通过较佳实施例进行了描述，相关技术人员明显能在不脱离本发明内容、精神和范围内对本文所述的设备和工艺流程进行改动或适当变更与组合，来实现本发明技术。特别需要指出的是，所有相类似的替换和改动对本领域技术人员来说是显而易见的，他们都被视为包括在本发明精神、范围和内容中。

Claims (10)

  1. 一种乙二醇和1,2-丁二醇混合物的分离精制方法，其特征在于：首先将乙二醇和1,2-丁二醇与反应剂醛/酮通过缩醛或者缩酮反应形成相应的缩醛/酮产物液体混合物，回收乙二醇和1,2-丁二醇的缩醛/酮产物液体混合物中的醛/酮，液体混合物再通过分离塔，通过精馏实现不同缩醛/酮产物的分离，然后水解缩醛/酮产物获得乙二醇和1,2-丁二醇初品，最后分别经过精馏对各自初品进行提纯，得到乙二醇和1,2-丁二醇产品。
  2. 根据权利要求1所述的乙二醇和1,2-丁二醇混合物的分离精制方法，其特征在于：所述反应剂醛/酮为含1～8个碳原子的醛的一种，或者含3～8个碳原子的酮的一种。
  3. 根据权利要求1所述的乙二醇和1,2-丁二醇混合物的分离精制方法，其特征在于：所述反应剂为乙醛或丙酮。
  4. 根据权利要求1所述的乙二醇和1,2-丁二醇混合物的分离精制方法，其特征在于：水解后的乙二醇和1,2-丁二醇初品再进行乙二醇和1,2-丁二醇精制，进一步提纯乙二醇和1,2-丁二醇。
  5. 一种乙二醇和1,2-丁二醇混合物的分离精制工艺，其特征在于：方法包含以下步骤：
    ⑴含乙二醇和1,2-丁二醇的物流(S01)与反应剂醛或者酮(S02)进入到反应器(R1)中反应，采出液体混合反应产物(S03)；
    ⑵液体混合反应产物(S03)进入醛/酮回收塔(T11)，塔顶排出主要包含未反应的醛/酮的物流(S09)，循环使用，塔釜采出缩醛/酮反应的生成物和未反应的二醇混合物流S10；
    ⑶塔釜物流(S10)进入乙二醇缩醛/酮产物分离塔(T12)，塔顶采出乙二醇缩醛/酮产物与水的共沸物(S16)，塔釜采出含1,2-丁二醇缩醛/酮产物的流股(S17)；
    ⑷塔釜采出物流(S17)进入1,2-丁二醇缩醛/酮产物分离塔(T13)，塔顶采出丁二醇缩醛/酮产物与水的混合物流股(S37)，塔釜采出为反应的二醇混合物流股(S38)，循环至反应器(R1)；
    ⑸步骤⑶塔顶采出物流(S16)与补充的流股水(S18)进入缩醛/酮水解精馏塔(T21)，塔中部设置反应段，塔顶排出水解产物醛/酮(S23)，循环使用，塔釜采出含水的乙二醇流股(S24)；
    ⑹步骤⑸塔釜采出物流(S24)进入乙二醇精制塔(T22)，塔顶排除水(S30)，循环使用，塔釜采出乙二醇产品(S31)；
    ⑺步骤⑷塔顶采出物流(S37)与补充的流股水(S40)进入另外一个缩醛/酮水解精馏塔(T31)，塔中部设置反应段，塔顶排出水解产物醛/酮(S45)，循环使用，塔釜采出含水的1,2-丁二醇流股(S46)；
    ⑻步骤⑺塔釜采出物流(S46)进入1,2-丁二醇精制塔(T32)，塔顶排除水(S52)，循环使用，塔釜采出1,2-丁二醇产品(S53)。
  6. 根据权利要求1所述的乙二醇和1,2-丁二醇混合物的分离精制工艺，其特征在于：上述步骤的反应条件如下：
    步骤⑴中的反应器R1采用釜式反应器、固定床反应器的一种，操作压力以绝对压力计为0.5～10atm，反应温度40～200℃，催化采用固体酸或固体碱树脂催化剂的一种或者混合， 进入反应器中的醛/酮进料量与二醇进料总量的摩尔比为1～10；
    步骤⑵的醛/酮回收塔T11以绝对压力计的操作压力为0.1～10atm，回流比0.01～15，或者直接气相采出，不回流；
    步骤⑶的乙二醇缩醛/酮产物分离塔T12以绝对压力计的操作压力为0.5～10atm，回流比0.1～15；
    步骤⑷的丁二醇缩醛/酮产物分离塔T13以绝对压力计的操作压力为0.01～5atm，回流比0.1～15；
    步骤⑸的缩醛/酮水解精馏塔T21以绝对压力计的操作压力为0.1～10atm，回流比0.01～20，或者塔顶直接气相采出，不回流，补充的水S18与S16中的水总进料量与缩醛/酮的总进料量的摩尔比为1～10；
    步骤⑹的乙二醇精制塔T22以绝对压力计的操作压力为0.01～2atm，回流比0.01～10；
    步骤⑺的缩醛/酮水解精馏塔T31以绝对压力计的操作压力为0.1～10atm，回流比0.01～20，或者塔顶直接气相采出，补充的水S40与S37中的水总进料量与缩醛/酮的总进料量的摩尔比为1～10；
    步骤⑻的1,2-丁二醇精制塔T32以绝对压力计的操作压力为0.01～2atm，回流比0.01～10。
  7. 根据权利要求1所述的乙二醇和1,2-丁二醇混合物的分离精制工艺，其特征在于：所述步骤⑸和步骤⑺的缩醛/酮产物的水解采用先反应后分离的方式进行，物流(S16)与(S46)分别进入各自的水解单元，水解单元包含一个反应器(R2或R3)，一个醛/酮分离塔(T51或T61)，一个缩醛/酮产物回收塔(T52或T62)，产品精制塔(T53或T63)。
  8. 根据权利要求7所述的乙二醇和1,2-丁二醇混合物的分离精制工艺，其特征在于：所述反应器(R2或R3)操作压力以绝对压力计为0.5～10atm，反应温度40～200℃，催化采用固体酸或固体碱树脂催化剂的一种或者混合；醛/酮分离塔(T51，T61)操作压力以绝压计为0.5～10atm，回流比0.01～15，或者塔顶直接气相采出；缩醛/酮产物回收塔(T52，T62)操作压力以绝压计为0.1～10atm，回流比0.1～15；产品精制塔(T53，T63)操作压力以绝压计为0.01～2atm，回流比0.01～10。
  9. 根据权利要求5所述的乙二醇和1,2-丁二醇混合物的分离精制工艺，其特征在于：1,2-丁二醇的缩醛/酮产物不进行水解，直接分离获得缩醛/酮产品，如附图2所示的塔T41，塔顶设置分相器，操作温度20～150℃，水相排出，有机相回流，塔T41以绝对压力计为0.05～5atm。
  10. 一套乙二醇和1,2-丁二醇混合物的分离精制装置，其特征在于：装置包含反应器(R1)、醛/酮回收塔(T11)、乙二醇缩醛/酮产物分离塔(T12)、丁二醇缩醛/酮产物分离塔(T13)，缩醛/酮水解精馏塔(T21、T31)、乙二醇精制塔(T22)、1,2-丁二醇精制塔(T32)、冷凝器、再沸器、泵以及相关的进料管线和连接以上设备的管线；所有精馏塔塔顶和塔釜均设置冷凝器和再沸器。反应器(R1)出料进入、醛/酮回收塔(T11)，塔顶采出未反应的醛/酮泵送回反应器(R1)，乙二醇缩醛/酮产物分离塔(T12)；乙二醇缩醛/酮产物分离塔(T12)塔顶采出乙二醇缩醛/酮产物与水的混合物进入乙二醇缩醛/酮产物水解精馏塔(T21)，塔釜采出流股进入丁二醇缩醛/酮产物分离塔(T13)；丁二醇缩醛/酮产物分离塔(T13)塔顶采出丁二醇缩醛/酮产 物与水的混合物进入丁二醇缩醛/酮产物水解精馏塔(T31)，塔釜采出未反应的二醇混合物循环使用；乙二醇缩醛/酮产物水解精馏塔(T21)中部设置反应段，塔顶采出醛/酮产物，用于循环使用，塔釜采出乙二醇与水的混合物进入乙二醇精制塔(T22)；乙二醇精制塔(T22)塔顶脱除水，塔釜为乙二醇产品；丁二醇缩醛/酮产物水解精馏塔(T31)中部设置反应段，塔顶采出醛/酮产物，用于循环使用，塔釜采出丁二醇与水的混合物进入丁二醇精制塔(T32)；丁二醇精制塔(T32)塔顶脱除水，塔釜为1,2-丁二醇产品。
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Cited By (5)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN114470836A (zh) * 2021-12-28 2022-05-13 天津大学 基于切割乙酸乙酯的煤制乙醇液相产物的分离装置和分离方法
CN114656332A (zh) * 2020-12-22 2022-06-24 中国石油化工股份有限公司 组合物及其制备方法和应用
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CN115304454A (zh) * 2021-05-07 2022-11-08 中国石油化工股份有限公司 一种聚酯生产过程中回收乙二醇的分离方法及系统

Families Citing this family (6)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
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Citations (10)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US4966658A (en) * 1989-12-27 1990-10-30 Lloyd Berg Recovery of ethylene glycol from butanediol isomers by azeotropic distillation
CN103193594A (zh) * 2012-01-10 2013-07-10 中国石油化工股份有限公司 用于分离乙二醇和1，2-丁二醇的方法
CN103772146A (zh) * 2012-10-25 2014-05-07 中国石油化工股份有限公司 分离乙二醇和1,2-丁二醇的方法
CN105541551A (zh) * 2016-02-04 2016-05-04 天津大学 乙二醇和1,2-丁二醇的反应精馏分离精制新方法、工艺及装置
CN105541555A (zh) * 2016-02-04 2016-05-04 天津大学 用于分离乙二醇、丙二醇和丁二醇的反应精馏方法及装置
CN105622338A (zh) * 2016-02-04 2016-06-01 天津大学 乙二醇和1,2-丁二醇的分离方法、工艺及装置
CN105622337A (zh) * 2016-02-04 2016-06-01 天津大学 煤制乙二醇液相产物分离的反应精馏耦合新工艺及装置
CN105622343A (zh) * 2016-02-04 2016-06-01 天津大学 生物质基制乙二醇液相产物的反应精馏分离新工艺及装置
CN105693466A (zh) * 2016-03-10 2016-06-22 天津大学 一种二醇缩醛/酮产物高效水解的反应精馏方法及装置
CN105693687A (zh) * 2016-03-10 2016-06-22 天津大学 一种二醇缩醛/酮反应的高效反应精馏方法及装置

Family Cites Families (2)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US3929915A (en) * 1974-07-31 1975-12-30 Du Pont Process for the production of butanediol
US6548681B1 (en) * 2001-06-26 2003-04-15 Board Of Trustees Of Michigan State University Process for the recovery of a polyol from an aqueous solution

Patent Citations (10)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US4966658A (en) * 1989-12-27 1990-10-30 Lloyd Berg Recovery of ethylene glycol from butanediol isomers by azeotropic distillation
CN103193594A (zh) * 2012-01-10 2013-07-10 中国石油化工股份有限公司 用于分离乙二醇和1，2-丁二醇的方法
CN103772146A (zh) * 2012-10-25 2014-05-07 中国石油化工股份有限公司 分离乙二醇和1,2-丁二醇的方法
CN105541551A (zh) * 2016-02-04 2016-05-04 天津大学 乙二醇和1,2-丁二醇的反应精馏分离精制新方法、工艺及装置
CN105541555A (zh) * 2016-02-04 2016-05-04 天津大学 用于分离乙二醇、丙二醇和丁二醇的反应精馏方法及装置
CN105622338A (zh) * 2016-02-04 2016-06-01 天津大学 乙二醇和1,2-丁二醇的分离方法、工艺及装置
CN105622337A (zh) * 2016-02-04 2016-06-01 天津大学 煤制乙二醇液相产物分离的反应精馏耦合新工艺及装置
CN105622343A (zh) * 2016-02-04 2016-06-01 天津大学 生物质基制乙二醇液相产物的反应精馏分离新工艺及装置
CN105693466A (zh) * 2016-03-10 2016-06-22 天津大学 一种二醇缩醛/酮产物高效水解的反应精馏方法及装置
CN105693687A (zh) * 2016-03-10 2016-06-22 天津大学 一种二醇缩醛/酮反应的高效反应精馏方法及装置

Cited By (6)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN114656332A (zh) * 2020-12-22 2022-06-24 中国石油化工股份有限公司 组合物及其制备方法和应用
CN115215729A (zh) * 2021-04-14 2022-10-21 中国石油化工股份有限公司 一种聚酯级煤基乙二醇的制法
CN115304454A (zh) * 2021-05-07 2022-11-08 中国石油化工股份有限公司 一种聚酯生产过程中回收乙二醇的分离方法及系统
CN114470836A (zh) * 2021-12-28 2022-05-13 天津大学 基于切割乙酸乙酯的煤制乙醇液相产物的分离装置和分离方法
CN114470836B (zh) * 2021-12-28 2023-07-18 天津大学 基于切割乙酸乙酯的煤制乙醇液相产物的分离装置和分离方法
CN115212600A (zh) * 2022-08-22 2022-10-21 福建永荣科技有限公司 一种节能高效型环己酮精制系统

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