Process system for producing succinate from maleic anhydride airflow and application of process system

Abstract

The invention provides a process system for producing succinate from maleic anhydride gas flow. The process system comprises an alcohol solvent absorption esterification system and an absorption liquid hydrogenation or secondary esterification hydrogenation coupling system which are sequentially connected through a pipeline. The alcohol solvent absorption esterification system comprises one or more stages of absorption esterification towers and a dissolved oxygen removal tower which are connected in series, the absorption liquid hydrogenation or secondary esterification hydrogenation coupling system comprises an absorption liquid hydrogenation reactor and a gas-liquid separator which are connected in sequence, and an outlet of the dissolved oxygen removal tower is connected to the absorption liquid hydrogenation reactor; the absorption esterification tower is a bubble tower, and the absorption liquid hydrogenation reactor is a fixed bed reactor. A highly integrated process system is adopted, maleic anhydride absorption-esterification coupling and absorption liquid hydrogenation are connected in series or maleic anhydride absorption-esterification coupling and secondary esterification-hydrogenation coupling are connected in series to prepare succinate, the process flow is shortened, and equipment investment and energy consumption are reduced.

Description

Technical Field

[0001] The present invention relates to the field of integrated reaction processes and apparatuses for producing succinate esters, and particularly to a process system for producing succinate esters from maleic anhydride gas stream and its applications. Background Art

[0002] Succinate esters are widely used in industries such as perfumes, foods, pharmaceuticals, pigments, coatings, plastics, and rubbers. Succinate esters are also important raw materials for synthesizing succinic acid polyester diol, a raw material for polyurethane elastomers, and biodegradable polyesters.

[0003] Currently, the mainstream synthesis methods for succinate esters are succinic acid or succinic anhydride esterification methods and hydrogenation of maleic anhydride esterification products. For example, CN101323566B discloses a method for preparing diethyl succinate. The material ratio is succinic acid: industrial ethanol: absolute ethanol: catalyst = 1: (0.8 - 1.1): 0.1: (0.01 - 0.02). Succinic acid and ethanol are subjected to esterification, filtration, alkali washing, rectification and other batch operations under the catalysis of phosphotungstic acid to obtain diethyl succinate with a yield greater than 90%. CN101092358B adds succinic acid, isopropanol, a catalyst and a dehydrating agent to an esterification reaction kettle equipped with a fractionating column to carry out batch esterification reaction to prepare diisopropyl succinate.

[0004] CN102001939B discloses a method for preparing succinic acid dicarboxylate ester. Hydrogen is mixed with maleic acid dicarboxylate ester and passed through a fixed bed reactor equipped with a hydrogenation catalyst. Part of the generated succinic acid dicarboxylate ester is used as a circulating dilution heat transfer material, and the purity of the product succinic acid dicarboxylate ester exceeds 99.6%. CN101979139B provides a catalyst for hydrogenating maleic acid dicarboxylate ester to prepare succinic acid dicarboxylate ester and its preparation method. The catalyst is prepared by an impregnation method. The content of the active component nickel is 5 - 25 wt%, and the content of the promoters Na, K, Ca, Mg, Mn, Ba or Cu is 0.5 - 8 wt%. The rest is the carrier alumina, silica, silica-alumina composite oxide or activated carbon. The maleic acid dicarboxylate ester suitable for hydrogenation by the catalyst is a carboxylic acid ester of C1 - C5. Using a fixed bed reaction process, the conversion rate of maleic acid dicarboxylate ester is 100%, and the selectivity of succinic acid dicarboxylate ester is above 99.6%. CN101745396B discloses a catalyst for hydrogenating dialkyl maleate to prepare dialkyl succinate, which is prepared by a co-precipitation method. The main active component copper oxide accounts for 40 - 60%, the promoter zinc oxide accounts for 20 - 50% and alumina accounts for 10 - 20%. Before use, the catalyst is reduced in a hydrogen stream at 300 °C for 4 hours, and then at a reaction temperature of 100 °C, a hydrogen pressure of 5 MPa, a hydrogen / ester molar ratio of 100 and a liquid hourly space velocity of 0.2 h -1Hydrogenation reaction is carried out below. CN101747189B provides a method for hydrogenating dialkyl maleate (C1-C5 alkyl ester) to dialkyl succinate. The CuZnAl catalyst is loaded into a fixed-bed reactor and reduced with hydrogen at a flow rate of 500 ml / min at 230 °C and atmospheric pressure for 12 hours, and then fed with a raw material containing dialkyl maleate for hydrogenation reaction. The liquid product is cooled and separated to obtain the target product; the hydrogenation reaction temperature is 80-120 °C, the pressure is 0.1-7 MPa, and the space velocity of dialkyl maleate is 0.1-10 h -1 , and the hydrogen-to-ester molar ratio is 5-250:1.

[0005] For the preparation of succinate by the hydrogenation of maleic anhydride esterification or the electrolytic reduction of esterification, CN102070448B discloses a method for preparing dimethyl succinate. Using maleic anhydride and methanol as raw materials, dimethyl maleate is synthesized by using an acidic cation exchange resin catalyst and a fixed-bed catalytic distillation process, and then dimethyl maleate is hydrogenated to prepare dimethyl succinate by using a Pd catalyst supported on Al2O3, SiO2, TiO2, ZrO2 or activated carbon and a fixed-bed reaction process. The esterification conversion rate of maleic anhydride is 100% and the selectivity of dimethyl maleate is greater than 99%. The hydrogenation conversion rate of dimethyl maleate and the selectivity of dimethyl succinate are both greater than 99.8%.

[0006] In summary, the existing preparation methods of succinic monoester or diester mainly use succinic acid or succinic anhydride as raw materials and directly esterify with C1-C5 alcohols; or use maleate as raw material and obtain it through catalytic hydrogenation reaction; or use commercial maleic anhydride and C1-C5 alcohols as raw materials and prepare it through two-step reaction or one-step coupling reaction of esterification hydrogenation or hydrogenation esterification. There is no report on the existing succinate synthesis technology using industrial gaseous crude maleic anhydride as raw material, and there is no attention to the overall process and system integration starting from gaseous crude maleic anhydride, which leads to limited reduction in the production cost of succinate. At present, the investment and production cost of maleic anhydride industrial plants are mainly consumed in the solvent absorption and analysis of the gas-phase maleic anhydride logistics of the oxidation reaction product, the separation and purification of maleic anhydride, and the purification and recovery of the solvent. If industrial gaseous crude maleic anhydride, the oxidation reaction product of maleic anhydride industrial plants, is used as the starting raw material, and succinic monoester or / and diester is prepared through the absorption-esterification coupling reaction with alcohol as the absorbent and the hydrogenation reaction of the absorbent solution or the esterification-hydrogenation coupling reaction in series, the process flow will be greatly shortened and the production cost will be reduced. Summary of the Invention

[0007] Aiming at the above-mentioned defects of the prior art, the purpose of the present invention is to provide a reaction system for preparing succinate by absorption-esterification-hydrogenation starting from industrial gaseous crude maleic anhydride logistics, and applying it to the low-cost production of succinate products starting from the rich and cheap gaseous crude maleic anhydride as the starting raw material.

[0008] To achieve the above object, the technical solution of the present invention is specifically as follows:

[0009] A process system for producing succinic acid ester from maleic anhydride gas stream includes an alcohol solvent absorption esterification system and an absorption liquid hydrogenation or two-esterification hydrogenation coupling system connected in sequence through pipelines, wherein:

[0010] The alcohol solvent absorption esterification system includes 1-3 stage absorption esterification towers for absorbing maleic anhydride gas stream and an absorption liquid deoxygenation tower for removing dissolved oxygen. The absorption esterification tower is provided with an alcohol absorbent inlet, a maleic anhydride gas stream inlet and an absorption liquid outlet, and the absorption liquid outlet is connected to the deoxygenation tower;

[0011] The absorption liquid hydrogenation or two-esterification hydrogenation coupling system includes an absorption liquid hydrogenation reactor and a gas-liquid separator connected in sequence. The outlet of the deoxygenation tower is connected to the absorption liquid hydrogenation reactor, and the gas phase outlet of the gas-liquid separator is connected to the hydrogen inlet of the absorption liquid hydrogenation reactor through a pipeline; The absorption esterification tower is a bubble column, and the absorption liquid hydrogenation reactor is a fixed bed reactor;

[0012] The absorption esterification tower is a bubble column, and the absorption liquid hydrogenation reactor is a fixed bed reactor.

[0013] In the present invention, the maleic anhydride gas stream comes from the maleic anhydride gas stream generated by an industrial maleic anhydride oxidation reactor and cooled, such as the process of producing maleic anhydride by benzene oxidation or n-butane oxidation.

[0014] The present invention is further configured such that the absorption esterification tower is a countercurrent spray bubble absorption esterification tower or a co-current bubble absorption esterification tower, wherein:

[0015] The countercurrent spray bubble absorption esterification tower includes an upper head, a spray absorption zone, a bubble absorption esterification zone and a lower head of the absorption tower from top to bottom. The upper head is provided with an alcohol absorbent inlet and a tail gas discharge outlet, and the lower head is provided with a maleic anhydride gas stream input port and an absorption liquid output port; A sprayer is provided above the spray absorption zone, and the sprayer evenly sprays the absorbent and sprays the maleic anhydride gas stream. The bubble absorption esterification zone is filled with particulate packing or / and solid acid catalyst. The sprayed alcohol absorbent passes through the bed from top to bottom, and the maleic anhydride gas stream passes through the bed countercurrently from bottom to top;

[0016] The co-current bubble absorption esterification tower includes an upper head, a bubble absorption esterification zone and a lower head of the absorption tower from top to bottom. The upper head is provided with an absorption liquid output port and a tail gas discharge outlet, and the lower head is provided with an alcohol absorbent input port and a maleic anhydride gas stream input port; The bubble absorption esterification zone is filled with particulate packing or / and solid acid catalyst. The alcohol absorbent and the maleic anhydride gas stream entering the co-current bubble absorption esterification tower both pass through the bed in a co-current manner from bottom to top.

[0017] The present invention is further configured such that the 1st - 3rd stage absorption - esterification tower includes a 1st - stage absorption tower with only 1 absorption tower, or a 2nd - stage absorption tower formed by 2 absorption towers in series, or a 3rd - stage absorption tower formed by 3 absorption towers in series, and the maleic anhydride gas stream and the alcohol absorbent are fed into the same tower or different towers.

[0018] The present invention is further configured such that the absorption liquid hydrogenation reactor is an adiabatic fixed - bed reactor or an isothermal fixed - bed reactor; wherein:

[0019] The adiabatic fixed - bed hydrogenation reactor includes a reactor upper head, a first - stage reactor, a second - stage reactor, and a reactor lower head from top to bottom. Among them, the upper head is provided with a liquid - phase hydrogenation raw material inlet and a hydrogen inlet, a distributor is provided inside the reactor upper head, and inert ceramic balls are filled at the top of the beds of the first - stage reactor and the second - stage reactor, and the bed bodies are filled with hydrogenation catalyst or di - esterification - hydrogenation bifunctional catalyst; a heat exchanger is provided between the first - stage reactor and the second - stage reactor, the reaction materials of the first - stage reactor are cooled by heat exchange through the heat exchanger and then enter the second - stage reactor to continue the reaction, and the lower head is provided with a reaction product outlet; the liquid - phase hydrogenation materials and hydrogen are mixed in the distributor and sequentially enter the first - stage reactor and the second - stage reactor through the distributor;

[0020] The isothermal fixed - bed reactor includes a reactor upper head, an isothermal tube - fixed - bed reactor, and a reactor lower head from top to bottom. Among them, the upper head is provided with a liquid - phase hydrogenation raw material inlet and a hydrogen inlet, a distributor is provided inside the reactor upper head, and the lower head is provided with a reaction product outlet; inert ceramic balls are filled at the top and bottom of the tubes of the isothermal tube - fixed - bed reactor, the tube beds are filled with hydrogenation catalyst or di - esterification - hydrogenation bifunctional catalyst, and a heat medium is filled between the tubes, and the heat medium enters from the bottom and exits from the top for counter - current heat exchange with the reaction materials entering from the top and exiting from the bottom.

[0021] The present invention is further configured such that the process system further includes an absorbent recovery tower, a succinic acid ester refining tower, and a tail - gas treatment system. The gas - phase outlet and the liquid - phase outlet of the gas - liquid separator are respectively connected to the hydrogen inlet of the absorption liquid hydrogenation reactor and the absorbent recovery tower, the product outlet of the absorbent recovery tower is connected to the succinic acid ester refining tower, and the tail - gas outlet of the absorption - esterification tower and the dissolved - oxygen outlet of the de - dissolved - oxygen tower are both connected to the tail - gas treatment system.

[0022] The present invention is further configured such that a vacuum system is provided at the top of the de - dissolved - oxygen tower for removing the dissolved oxygen in the absorption liquid under negative pressure, and the removed dissolved oxygen is transported to the tail - gas treatment system.

[0023] The present invention is further configured such that a connecting pipeline is further provided between the liquid - phase outlet of the absorption liquid hydrogenation reactor and the gas - liquid separator for circulating a part of the hydrogenation reaction liquid as a circulating dilution heat - transfer agent.

[0024] The present invention also provides an application of the process system for producing succinic acid esters from maleic anhydride gas stream, which comprises the following steps:

[0025] S1. Absorption-esterification of maleic anhydride gas stream: The maleic anhydride gas stream and an alcohol absorbent with or without a liquid acid catalyst are respectively fed into the absorption-esterification tower to carry out solvent absorption of maleic anhydride and esterification reaction to prepare a maleic acid ester absorbent solution, and the maleic acid ester includes maleic acid monoester and / or maleic acid diester:

[0026] S2. Removal of dissolved oxygen from the absorbent solution: The absorbent solution after absorption-esterification of the maleic anhydride gas stream in the absorption-esterification tower is fed into a de-dissolved oxygen tower, and after being de-dissolved oxygen under negative pressure, it is fed into an absorbent solution hydrogenation or two-esterification hydrogenation coupling system; Operating conditions: absorbent solution temperature 52-150 °C and tower top pressure 50.0-90.0 kPa;

[0027] S3. Absorbent solution hydrogenation or two-esterification-hydrogenation: The absorbent solution after removing dissolved oxygen and hydrogen are fed into an absorbent solution hydrogenation reactor filled with a supported metal catalyst pre-reduced and activated in advance to carry out double bond hydrogenation of maleic acid ester in the absorbent solution or two-esterification-hydrogenation of maleic acid monoester to prepare succinic acid esters, and the succinic acid esters include succinic acid monoester and / or succinic acid diester;

[0028] S4. Gas-liquid separation of the reaction product: The material after absorbent solution hydrogenation or two-esterification-hydrogenation reaction is fed into a gas-liquid separator for gas-liquid separation, and the liquid material is subjected to subsequent refining treatment of succinic acid esters.

[0029] The present invention is further configured that in step S1, the operating conditions are: the concentration of the liquid acid catalyst in the alcohol absorbent is 0-5.0 wt%, the temperature of the maleic anhydride gas stream is 55-180 °C and the volume space velocity is 200-20000 h -1 , the temperature of the alcohol absorbent is 25-85 °C and the weight space velocity is 0.1-10.0 h -1 , and through the absorption of one or more stages of serially-connected absorption-esterification towers, a maleic acid ester absorbent solution is obtained.

[0030] Preferably, in step S1, the operating conditions are: the concentration of the liquid acid catalyst in the alcohol absorbent is 0-3.0 wt%, the temperature of the maleic anhydride gas stream is 60-150 °C and the volume space velocity is 500-10000 h -1 , the temperature of the alcohol absorbent is 30-80 °C and the weight space velocity is 0.2-5.0 h -1 , and through the absorption of 1-3 stages of serially-connected absorption towers, a maleic acid ester absorbent solution is obtained;

[0031] More preferably, in step S1, the operating conditions are: the concentration of the liquid acid catalyst in the alcohol absorbent is 0-2.5 wt%, the temperature of the maleic anhydride gas stream is 65-120 °C and the volume space velocity is 1000-5000 h-1 , the temperature of the alcohol absorbent is 35 - 75 °C and the weight hourly space velocity is 0.5 - 3.0 h -1 , and the maleate absorbent solution is obtained through series absorption in 1 - 3 absorption towers

[0032] The present invention is further configured such that, in step S1, when preparing maleic monoester, the maleic anhydride gas stream and the absorbent flow in parallel or countercurrent through 1 - 3 inert packing beds, and the absorption of maleic anhydride and the monoesterification reaction occur to generate a maleic monoester absorbent solution; operating conditions: the temperature of the maleic anhydride gas stream is 60 - 150 °C and the volumetric hourly space velocity is 500 - 10000 h -1 , the temperature of the absorbent is 25 - 85 °C and the weight hourly space velocity is 0.5 - 5.0 h -1 , preferably, the temperature of the maleic anhydride gas stream is 65 - 120 °C and the volumetric hourly space velocity is 1000 - 7500 h -1 , the temperature of the absorbent is 35 - 80 °C and the weight hourly space velocity is 0.75 - 3.0 h -1 , more preferably, the temperature of the maleic anhydride gas stream is 70 - 95 °C and the volumetric hourly space velocity is 1500 - 5000 h -1 , the temperature of the absorbent is 45 - 75 °C and the weight hourly space velocity is 1.0 - 2.5 h -1 .

[0033] The present invention is further configured such that, in step S1, when preparing maleic diester, the maleic anhydride gas stream and the alcohol absorbent containing a liquid acid catalyst flow in parallel or countercurrent through 1 - 3 inert packing beds or solid acid catalyst beds, and the absorption of maleic anhydride and the esterification reaction occur to generate a maleic diester absorbent solution; operating conditions: the concentration of the liquid acid catalyst is 0.05 - 5.0 wt%, the temperature of the maleic anhydride gas stream is 65 - 150 °C and the volumetric hourly space velocity is 1000 - 10000 h -1 , the temperature of the absorbent is 45 - 80 °C and the weight hourly space velocity is 0.5 - 5.0 h -1 , preferably, the concentration of the liquid acid catalyst is 0.1 - 3.0 wt%, the temperature of the maleic anhydride gas stream is 70 - 135 °C and the volumetric hourly space velocity is 1500 - 5000 h -1 , the temperature of the absorbent is 50 - 75 °C and the weight hourly space velocity is 1.0 - 4.0 h -1 , more preferably, the concentration of the liquid acid catalyst is 0.2 - 2.0 wt%, the temperature of the maleic anhydride gas stream is 75 - 120 °C and the volumetric hourly space velocity is 2000 - 3000 h -1 , the temperature of the absorbent is 55 - 70 °C and the weight hourly space velocity is 1.5 - 3.0 h -1 .

[0034] The present invention is further configured such that, when preparing maleic mono / di mixed ester in step S1, the maleic anhydride gas stream and the alcohol absorbent flow in parallel or countercurrent through 1 - 3 solid acid catalyst beds, and the absorption of maleic anhydride and the monoesterification / diesterification reaction occur to generate a maleic mixed ester absorbent solution; operating conditions: the temperature of the maleic anhydride gas stream is 80 - 180 °C and the volumetric hourly space velocity is 500 - 5000 h-1 、 The absorbent temperature is 50 - 90°C and the weight hourly space velocity is 0.5 - 5.0 h -1 , preferably the maleic anhydride gas stream temperature is 85 - 150°C and the volume hourly space velocity is 1000 - 3000 h -1 、 The absorbent temperature is 55 - 85°C and the weight hourly space velocity is 1.0 - 3.0 h -1 , more preferably the maleic anhydride gas stream temperature is 90 - 120°C and the volume hourly space velocity is 1500 - 2500 h -1 、 The absorbent temperature is 60 - 80°C and the weight hourly space velocity is 1.5 - 2.5 h -1 .

[0035] The present invention is further configured such that in the bubbling absorption and esterification zone of the absorption and esterification tower, inert packing is filled, or solid acid particle catalysts are filled in the lower layer and inert packing is filled in the upper layer, and the inert packing is selected from at least one of stainless steel, quartz, corundum or silicon carbide; preferably, in the bubbling absorption and esterification zone, inert spherical packing with a diameter of 2.0 - 5.0 mm is filled, or particle catalysts with a diameter of 2.0 - 5.0 mm are filled in the lower layer and inert spherical packing with a diameter of 2.0 - 5.0 mm is filled in the upper layer; more preferably, in the bubbling absorption and esterification zone, quartz spheres with a diameter of 3.0 - 4.0 mm are filled, or particle catalysts with a diameter of 3.0 - 4.0 mm are filled in the lower layer and quartz spheres with a diameter of 3.0 - 4.0 mm are filled in the upper layer.

[0036] The present invention is further configured such that in step S2, the operating conditions are: the absorbent liquid temperature is 55 - 120°C and the tower top pressure is 60.0 - 85.0 kPa, preferably, the absorbent liquid temperature is 60 - 90°C and the tower top pressure is 70.0 - 80.0 kPa.

[0037] The present invention is further configured such that in step S3, the absorbent liquid is mixed with hydrogen and after preheating, it is fed into the absorbent liquid hydrogenation reactor.

[0038] The present invention is further configured such that in step S3, a fixed bed reactor and a non-acidic carrier supported metal particle catalyst or a solid acid supported metal particle bifunctional catalyst are used, and the absorbent liquid is fed into the absorbent liquid hydrogenation reactor, and is operated in parallel or countercurrent with hydrogen for the hydrogenation of maleate or the two-esterification - hydrogenation of maleic monoester to prepare succinate;

[0039] The operating conditions are: the absorbent liquid temperature is 50 - 150°C and the weight hourly space velocity is 0.1 - 3.0 h -1 、 The hydrogen pressure is 0.5 - 5.0 MPa and the hydrogen - ester molar ratio is 5 - 50; preferably, the absorbent liquid temperature is 55 - 120°C and the weight hourly space velocity is 0.25 - 2.5 h -1 、 The hydrogen pressure is 0.75 - 3.0 MPa and the hydrogen - ester molar ratio is 7.5 - 40; more preferably, the absorbent liquid temperature is 60 - 90°C and the weight hourly space velocity is 0.5 - 2.0 h -1, hydrogen pressure of 1.0 to 2.0 MPa and a hydrogen to ester molar ratio of 10 to 30.

[0040] The present invention is further configured such that in step S3, when preparing monosuccinate, the maleate monoester absorption liquid from which dissolved oxygen has been removed in S2 is fed into a fixed bed reactor filled with non-acidic carrier-supported metal particle catalysts that have been pre-reduced and activated, and is operated in co-current or counter-current with hydrogen to perform double bond hydrogenation of maleate monoester to prepare monosuccinate; operating conditions: absorption liquid temperature of 50 to 120 °C and weight hourly space velocity of 0.2 to 3.0 h -1 , hydrogen pressure of 0.5 to 5.0 MPa and a hydrogen to ester molar ratio of 5 to 50; preferably, the absorption liquid temperature is 55 to 95 °C and the weight hourly space velocity is 0.25 to 2.0 h -1 , hydrogen pressure of 0.75 to 3.0 MPa and a hydrogen to ester molar ratio of 7.5 to 40; more preferably, the absorption liquid temperature is 60 to 90 °C and the weight hourly space velocity is 0.5 to 1.5 h -1 , hydrogen pressure of 1.0 to 2.0 MPa and a hydrogen to ester molar ratio of 10 to 30.

[0041] The present invention is further configured such that in step S3, when preparing disuccinate: a liquid acid catalyst is added to the maleate monoester absorption liquid or maleate mono / di mixed ester absorption liquid from which dissolved oxygen has been removed in S2, and it is fed into a fixed bed reactor filled with non-acidic carrier or solid acid carrier-supported metal particle catalysts that have been pre-reduced and activated, and is operated in co-current or counter-current with hydrogen to perform the coupling reaction of diesterification-double bond hydrogenation of maleate monoester to prepare disuccinate; or the maleate diester absorption liquid from which dissolved oxygen has been removed in S2 is fed into a fixed bed reactor filled with non-acidic carrier or solid acid carrier-supported metal particle catalysts that have been pre-reduced and activated, and is operated in co-current or counter-current with hydrogen to perform double bond hydrogenation of maleate diester to prepare disuccinate; operating conditions: absorption liquid temperature of 70 to 120 °C and weight hourly space velocity of 0.2 to 2.0 h -1 , hydrogen pressure of 0.5 to 5.0 MPa and a hydrogen to ester molar ratio of 5 to 50; preferably, the absorption liquid temperature is 75 to 105 °C and the weight hourly space velocity is 0.3 to 1.5 h -1 , hydrogen pressure of 0.75 to 3.0 MPa and a hydrogen to ester molar ratio of 7.5 to 40; more preferably, the absorption liquid temperature is 80 to 95 °C and the weight hourly space velocity is 0.5 to 1.0 h -1 , hydrogen pressure of 1.0 to 2.0 MPa and a hydrogen to ester molar ratio of 10 to 30.

[0042] The present invention is further configured that in step S3, when preparing the succinic acid mono / di mixed ester: the maleic acid monoester absorption liquid or the maleic acid mono / di mixed ester absorption liquid from which dissolved oxygen has been removed in S2 is fed into a fixed-bed reactor filled with a bifunctional catalyst of solid acid-supported metal particles that has been pre-reduced and activated, and is operated in co-current or counter-current with hydrogen gas to carry out the coupling reaction of diesterification-double bond hydrogenation of maleic acid monoester to prepare the succinic acid mono / di mixed ester; the operating conditions are: the temperature of the absorption liquid is 60-110°C and the weight hourly space velocity is 0.2-2.0 h -1 , the hydrogen pressure is 0.5-3.0 MPa and the hydrogen-to-ester molar ratio is 5-40; preferably, the temperature of the absorption liquid is 65-105°C and the weight hourly space velocity is 0.3-1.5 h -1 , the hydrogen pressure is 0.75-2.0 MPa and the hydrogen-to-ester molar ratio is 7.5-30; more preferably, the temperature of the absorption liquid is 70-95°C and the weight hourly space velocity is 0.5-1.0 h -1 , the hydrogen pressure is 1.0-1.5 MPa and the hydrogen-to-ester molar ratio is 10-20.

[0043] The present invention is further configured that the alcohol absorbent is selected from at least one of fatty alcohols, aromatic alcohols, heterocyclic alcohols, polyether polyols or polyester polyols; preferably, it is selected from n-butanol, isobutanol, n-pentanol, isopentanol, neopentanol, n-hexanol, n-heptanol, n-octanol, isooctanol, 2-ethylhexanol, isononanol, lauryl alcohol, stearyl alcohol, cyclohexanol, methylcyclohexanol, cyclohexyl propanol, allyl alcohol, butenol, methallyl alcohol, isopentenol, cinnamyl alcohol, oleyl alcohol, propargyl alcohol, butynol, benzyl alcohol, phenethyl alcohol, hydrocinnamyl alcohol, tetrahydrofurfuryl alcohol, furfuryl alcohol, ethylene glycol, polyethylene glycol, 1,2-propanediol, polypropylene glycol, 1,3-propanediol, 1,4-butanediol, 1,5-pentanediol, 1,6-hexanediol, 1,7-heptanediol, 1,8-octanediol, 1,9-nonanediol, 1,10-decanediol, 2-methyl-1,3-propanediol, neopentyl glycol, 2,2-diethyl-1,3-propanediol, cyclohexanediol, cyclohexanedimethanol, phenyl-1,3-propanediol, benzenedimethanol, 2,5-furandimethanol, butenediol, butynediol, diethylaminoethanol, diethylaminopropanol, or at least one of 2,2,2-trifluoroethanol, pentafluoroethanol, heptafluoropropanol, 2,2,2-trichloroethanol, chloropropanol or 1,3-dichloropropanol.

[0044] The present invention is further configured that the acid catalyst includes a liquid acid catalyst and a solid acid catalyst, wherein,

[0045] The liquid acid catalyst is selected from alkylsulfonic acid, arylsulfonic acid, halogenated sulfonic acid, halogenated carboxylic acid or amide or imide of halogenated sulfonic acid; preferably methanesulfonic acid, benzenesulfonic acid, p-toluenesulfonic acid, naphthalenesulfonic acid, fluorosulfonic acid, trifluoromethanesulfonic acid, pentafluoroethanesulfonic acid, trichloroacetic acid, trifluoroacetic acid or pentafluoropropionic acid, or trifluoromethanesulfonamide, bisfluorosulfonylimide, bis(trifluoromethanesulfonyl)imide; more preferably p-toluenesulfonic acid, trifluoromethanesulfonic acid or bis(trifluoromethanesulfonyl)imide;

[0046] The solid acid catalyst is selected from spheres with a diameter of 1.6 - 4.2 mm, cylindrical particles with a diameter of 1.6 - 4.2 mm and a length of 2.0 - 5.0 mm or cloverleaf particles, and the solid acid is HM, Hβ, HZSM-5, HZSM-22, HMCM-22, HMCM-49, WO3 / ZrO2 or activated carbon (AC), SiO2, HMCM-41, HMCM-48 or SBA-15 loaded with WO3 / ZrO2, H3PW 12 O 40 or Cs 2.5 H 0.5 PW 12 O 40 ; preferably HM, HZSM-5, HMCM-22, WO3 / ZrO2 or activated carbon (AC), SiO2, HMCM-41, HMCM-48 or SBA-15 loaded with WO3 / ZrO2, H3PW 12 O 40 or Cs 2.5 H 0.5 PW 12 O 40 ; more preferably HM, HZSM-5, HMCM-22, (WO3 / ZrO2) / HMCM-41, H3PW 12 O 40 / HMCM-48 or Cs 2.5 H 0.5 PW 12 O 40 / SBA-15.

[0047] The present invention is further configured such that the hydrogenation or esterification-hydrogenation catalyst comprises a metal bifunctional catalyst supported on a non-acidic carrier or a solid acid carrier. Among them, the metal particle catalyst supported on the non-acidic carrier is selected from activated carbon (AC), mesoporous carbon, carbon nanotubes, graphene, SiO2, mesoporous SiO2, TiO2, ZrO2, SBA-15, ZEO-1, ZEO-3 or TS-1 supported noble metal catalysts, preferably single-metal or bimetallic catalysts of Ru, Pd or Pt supported on SiO2, TiO2, ZrO2, SBA-15 or TS-1; more preferably Pd / SiO2, Pt / SiO2, Pd-Ru / SiO2 or Pd / SBA-15 catalysts;

[0048] The solid acid support-supported metal bifunctional catalysts use HM, Hβ, HZSM-5, HZSM-22, HMCM-22, HMCM-49, or HMCM-41, HMCM-48 or SBA-15 to support WO3 / ZrO2, H3PW 12 O 40 or Cs 2.5 H 0.5 PW 12 O 40 The solid acid supported noble metal catalyst uses HM, HZSM-5, HMCM-22, H3PW 12 O 40 / HMCM-48, Cs 2.5 H 0.5 PW 12 O 40 / SBA-15 or (WO3 / ZrO2) / HMCM-41 as the support for single or bimetallic catalysts of Ru, Pd or Pt; more preferably Pd / HM, Pt / HZSM-5, Pd-Ru / HMCM-22, Pd / (WO3 / ZrO2) / HMCM-41 or Pd / Cs 2.5 H 0.5 PW 12 O 40 / SBA-15 catalyst.

[0049] The present invention is further configured such that the hydrogenation or esterification-hydrogenation catalyst includes a non-acidic support or a solid acid support-supported metal bifunctional catalyst, and the particle size is spherical with a diameter of 1.6 to 4.2 mm, cylindrical particles with a diameter of 1.6 to 4.2 mm and a length of 2.0 to 5.0 mm, or clover particles.

[0050] Compared with the prior art, the present invention has the following beneficial effects:

[0051] (1) The process system provided by the present invention uses inexpensive and easily available industrial gaseous crude maleic anhydride as the raw material, and prepares succinic acid ester products through two-step coupling reactions, greatly reducing the raw material cost, material consumption, energy consumption, and equipment investment, and thus significantly reducing the manufacturing cost of succinic acid esters;

[0052] (2) The process system of the present invention has a very wide range of adaptability to alcohol absorbents and esterifying agents. By changing the variety of alcohol absorbent solvents and process conditions, a series of succinic acid ester products can be produced using the same reaction system and device. Moreover, by adjusting the process scheme, succinic acid monoester, diester, or mono / di mixed ester products of the same alcohol can also be produced. The product variety is rich and the scheme is flexible, which can meet various market demands and has a wide range of technical applications. Description of the Drawings

[0053] Figure 1Schematic flow diagram of a process system for producing succinate from maleic anhydride gas stream.

[0054] Figure 2 Schematic structural diagram of an absorption esterification tower, where (a) and (b) are countercurrent spray bubble absorption towers, and (c) and (d) are co-current bubble absorption towers.

[0055] Figure 3 Schematic flow diagram of the series connection mode and material flow mode of 1 - 3 stage absorption esterification towers.

[0056] Figure 4 Schematic structural diagram of an inter-stage heat exchange adiabatic fixed bed reactor for hydrogenation of absorption liquid or two-esterification - hydrogenation.

[0057] Figure 5 Schematic structural diagram of a shell-and-tube isothermal fixed bed reactor for hydrogenation of absorption liquid or two-esterification - hydrogenation.

[0058] Figure 6 Schematic structural diagram of a fixed bed reactor for double bond hydrogenation of absorption liquid or two-esterification - double bond hydrogenation, where (a) is co-current isothermal trickle fixed bed, (b) is co-current adiabatic trickle fixed bed, (c) is countercurrent adiabatic bubble fixed bed, (d) is co-current adiabatic bubble fixed bed).

[0059] Among them, Figure 2 In (a) and (b), 11. Upper head of the absorption tower, 12. Spray absorption area, 13. Bubble absorption esterification area, 14. Lower head of the absorption tower, 15. Alcohol absorbent inlet, 16. Tail gas discharge port, 17. Maleic anhydride gas stream input port, 18. Absorption liquid output port, 19. Sprayer. Figure 2 In (c) and (d), 21. Upper head of the absorption tower, 22. Bubble absorption esterification area, 23. Lower head of the absorption tower, 24. Absorption liquid output port, 25. Tail gas discharge port, 26. Alcohol absorbent input port, 27. Maleic anhydride gas stream input port. Figure 4 In it, 31. Upper head of the reactor, 32. First stage reactor, 33. Second stage reactor, 34. Lower head of the reactor, 35. Liquid phase hydrogenation raw material inlet, 36. Hydrogen inlet, 37. Distributor, 38. Heat exchanger, 39. Reaction product outlet. Figure 5 In it, 41. Upper head of the reactor, 42. Isothermal shell-and-tube fixed bed reactor, 43. Lower head of the reactor, 44. Liquid phase hydrogenation raw material inlet, 45. Hydrogen inlet, 46. Reaction product outlet, 47. Distributor. Specific implementation mode

[0060] The present invention will be further described below in conjunction with specific embodiments. It should be noted that the embodiments described in this part are only a part of the present invention, rather than all embodiments. Based on the embodiments of the present invention, all other embodiments obtained by those skilled in the art without creative work fall within the scope of protection of the present invention.

[0061] As Figure 1 shown, a process system for producing succinate from maleic anhydride gas stream provided by the present invention includes an alcohol solvent absorption esterification system and an absorption liquid hydrogenation or two-esterification hydrogenation coupling system connected in sequence through pipelines. Among them,

[0062] The alcohol solvent absorption esterification system includes a 1-3 stage absorption esterification tower for absorbing maleic anhydride gas stream and a deoxygenation tower for removing dissolved oxygen in the absorption liquid. The absorption esterification tower is provided with an alcohol absorbent inlet, a maleic anhydride gas stream inlet and an absorption liquid outlet, and the absorption liquid outlet is connected to the deoxygenation tower;

[0063] The absorption liquid hydrogenation or two-esterification hydrogenation coupling system includes an absorption liquid hydrogenation reactor and a gas-liquid separator connected in sequence. The absorption liquid outlet of the deoxygenation tower is connected to the absorption liquid hydrogenation reactor, and the gas-phase outlet pipeline of the gas-liquid separator is connected to the absorption liquid hydrogenation reactor;

[0064] The absorption esterification tower is a bubble column, and the absorption liquid hydrogenation reactor is a fixed bed reactor.

[0065] In the present invention, the maleic anhydride gas stream comes from the cooled maleic anhydride gas stream generated by an industrial maleic anhydride oxidation reactor, such as the maleic anhydride gas stream produced by the benzene oxidation or n-butane oxidation process for producing maleic anhydride.

[0066] In the embodiments of the present invention, the process system further includes a tail gas treatment system, and an absorbent recovery tower and a succinate refining tower connected in sequence. The absorbent recovery tower is used to recover the alcohol absorbent to realize the recycling of the alcohol absorbent. The liquid-phase outlet of the gas-liquid separator is connected to the absorbent recovery tower, and the product outlet of the absorbent recovery tower is connected to the succinate refining tower; the gas-phase outlet of the gas-liquid separator is connected to the hydrogen inlet of the absorption liquid hydrogenation reactor to recycle the separated hydrogen; the tail gas outlet of the absorption esterification tower and the dissolved oxygen outlet of the deoxygenation tower are both connected to the tail gas treatment system.

[0067] A vacuum system (not shown in the figure) is provided at the top of the deoxygenation tower to remove dissolved oxygen in the absorption liquid under negative pressure, and the removed dissolved oxygen is transported to the tail gas treatment system.

[0068] In an embodiment of the present invention, a connecting pipeline is further provided between the liquid phase outlet of the absorption liquid hydrogenation reactor and the liquid phase outlet of the gas-liquid separator, for circulating a part of the hydrogenation reaction liquid to circulate and dilute the heat removal.

[0069] Reference Figure 2 , the absorption esterification tower is selected as a countercurrent spray bubbling absorption esterification tower or a co-current bubbling absorption esterification tower. Among them, as Figure 2 a and Figure 2 shown in b, the countercurrent spray bubbling absorption esterification tower includes an upper head 11 of the absorption tower, a spray absorption zone 12, a bubbling absorption esterification zone 13 and a lower head 14 of the absorption tower from top to bottom. Among them, the upper head 11 is provided with an alcohol absorbent inlet 15 and a tail gas discharge port 16, and the lower head 14 is provided with a maleic anhydride gas flow input port 17 and an absorption liquid output port 18; a sprayer 19 is arranged above the spray absorption zone, and the sprayer 19 uniformly sprays the alcohol absorbent and sprays the maleic anhydride gas flow. The bubbling absorption esterification zone 13 is filled with granular packing or / and solid acid catalyst. The sprayed alcohol absorbent passes through the bed from top to bottom, and the maleic anhydride gas flow passes through the bed countercurrently from bottom to top. Figure 2 In a, the bubbling absorption esterification zone 13 is filled with granular packing, Figure 2 In b, the bubbling absorption esterification zone 13 is filled with granular packing and solid acid catalyst.

[0070] Using Figure 2 a and Figure 2 shown in b, in the case of the countercurrent of the maleic anhydride gas flow and the absorbent, the maleic anhydride gas flow is fed into the absorption esterification tower from the maleic anhydride gas flow input port 17 on the lower head 14 of the absorption tower, passes through the packing or solid acid catalyst bed layer and the spray absorption zone 12 from the bottom of the absorption esterification tower to the top, and the tail gas after absorption and esterification is discharged from the tail gas discharge port 16 on the upper head 11 of the absorption tower; at the same time, the alcohol absorbent or the absorbent containing liquid acid catalyst is sprayed into the spray absorption zone 12 of the absorption tower from the alcohol absorbent inlet 15 on the upper head 11 of the absorption tower through the sprayer 19, passes through the packing or solid acid catalyst bed layer from the top of the absorption tower through the spray absorption zone 12 and the bubbling absorption esterification zone 13, and flows towards the bottom of the absorption tower; the maleic anhydride gas flow and the absorbent are countercurrent, and an esterification reaction occurs while the absorbent absorbs the maleic anhydride gas flow, and the obtained absorption liquid is taken out from the absorption liquid output port 18 on the lower head 14 of the absorption tower.

[0071] As Figure 2 c and Figure 2As shown in Figure d, the co-current bubble absorption esterification tower includes an upper head 21 of the absorption tower, a bubble absorption esterification zone 22, and a lower head 23 of the absorption tower from top to bottom. An absorption liquid outlet 24 and a tail gas outlet 25 are provided on the upper head 21 of the absorption tower. An alcohol absorbent inlet 26 and a maleic anhydride gas stream inlet 27 are provided on the lower head 23 of the absorption tower. The bubble absorption esterification zone 22 is filled with particulate packing or / and solid acid catalyst. The alcohol absorbent and the maleic anhydride gas stream entering the co-current bubble absorption esterification tower both flow upward through the bed from bottom to top in parallel. Figure 2 In Figure c, the bubble absorption esterification zone 13 is filled with particulate packing. Figure 2 In Figure d, the bubble absorption esterification zone 13 is filled with particulate packing and solid acid catalyst.

[0072] Adopt Figure 2 Figure c and Figure 2 In the device shown in Figure d, in the case of co-current flow of the maleic anhydride gas stream and the absorbent, the maleic anhydride gas stream and the alcohol absorbent or the absorbent containing liquid acid catalyst are respectively fed into the absorption tower from the alcohol absorbent inlet 26 and the maleic anhydride gas stream inlet 27 on the lower head 23 of the absorption tower, and flow upward in parallel from the bottom of the absorption tower through the packing or the solid acid catalyst bed layer in the bubble absorption esterification zone 22 to the top. The tail gas after absorption-esterification is discharged from the tail gas outlet 25 on the upper head 11 of the absorption tower, and the absorption liquid is taken out from the absorption liquid outlet 24 on the upper head 21 of the absorption tower.

[0073] Refer to Figure 3 , the 1-3 stage absorption esterification tower includes a 1-stage absorption tower with only 1 absorption tower, or a 2-stage absorption tower formed by two absorption towers in series, or a 3-stage absorption tower formed by three absorption towers in series. The maleic anhydride gas stream and the alcohol absorbent are fed into the same tower or different towers; the feeding conditions of the multi-stage absorption esterification tower include the following four cases:

[0074] (1) Refer to Figure 3 Figure a, the maleic anhydride gas stream and the alcohol absorbent are fed into different towers in countercurrent: the maleic anhydride gas stream is fed into the bottom of the first-stage absorption tower, flows upward and undergoes a first-stage absorption-esterification reaction, flows out from the upper end of the tower and then is fed into the bottom of the second-stage absorption tower, flows upward and undergoes a second-stage absorption-esterification reaction, flows out from the upper end of the tower and then is fed into the bottom of the third-stage absorption tower, flows upward and undergoes a third-stage absorption-esterification reaction, and flows out from the upper end of the tower and is fed into the tail gas treatment system; while the fresh absorbent is fed into the top of the third-stage absorption tower, flows downward and absorbs maleic anhydride. After the absorption liquid flows out from the lower end of the third-stage absorption tower, it is fed into the top of the second-stage absorption tower, flows downward and continues to absorb maleic anhydride. After the absorption liquid flows out from the lower end of the second-stage absorption tower, it is fed into the top of the first-stage absorption tower, flows downward and further absorbs maleic anhydride, and the high-concentration absorption liquid is taken out from the lower end of the first-stage absorption tower;

[0075] (2) Refer to Figure 3b. The maleic anhydride gas stream and the absorbent are fed into the same tower in countercurrent: The maleic anhydride gas stream and the fresh absorbent are respectively fed into the bottom and the top of the first-stage absorption tower. The maleic anhydride gas stream flows upward from the bottom, while the absorbent flows downward from the top, and a first-stage absorption-esterification reaction occurs. After the gas stream flows out from the upper end of the tower and the absorbent solution flows out from the lower end, they are respectively fed into the bottom and the top of the second-stage absorption tower, flowing in countercurrent and undergoing a second-stage absorption-esterification reaction. After the gas stream flows out from the upper end of the tower and the absorbent solution flows out from the lower end, they are respectively fed into the bottom and the top of the third-stage absorption tower, flowing in countercurrent and undergoing a third-stage absorption-esterification reaction. After the gas stream flows out from the upper end of the tower, it is fed into the tail gas treatment system, and the high-concentration absorbent solution is withdrawn from the lower end of the third-stage absorption tower;

[0076] (3) See Figure 3 c. The maleic anhydride gas stream and the absorbent are fed into the same tower in cocurrent: The maleic anhydride gas stream and the fresh absorbent are respectively fed into the bottom and the lower end of the first-stage absorption tower, flowing in cocurrent upward from the bottom and undergoing a first-stage absorption-esterification reaction. After the gas stream flows out from the top of the tower and the absorbent solution flows out from the upper end, they are respectively fed into the bottom and the lower end of the second-stage absorption tower, flowing in cocurrent upward from the bottom and undergoing a second-stage absorption-esterification reaction. After the gas stream flows out from the top of the tower and the absorbent solution flows out from the upper end, they are respectively fed into the bottom and the lower end of the third-stage absorption tower, flowing in cocurrent upward from the bottom and undergoing a third-stage absorption-esterification reaction. After the gas stream flows out from the top of the tower, it is fed into the tail gas treatment system, and the high-concentration absorbent solution is withdrawn from the upper end of the third-stage absorption tower;

[0077] (4) See Figure 3 d. The maleic anhydride gas stream and the absorbent are fed into different towers in cocurrent: The maleic anhydride gas stream is fed into the bottom of the first-stage absorption tower, flows upward from the bottom and undergoes a first-stage absorption-esterification reaction, flows out from the top of the tower and then is fed into the bottom of the second-stage absorption tower, flows upward from the bottom and undergoes a second-stage absorption-esterification reaction, flows out from the top of the tower and then is fed into the bottom of the third-stage absorption tower, flows upward from the bottom and undergoes a third-stage absorption-esterification reaction, and flows out from the top of the tower and is fed into the tail gas treatment system; while the fresh absorbent is fed into the lower end of the third-stage absorption tower, flows upward from the bottom and absorbs maleic anhydride. After the absorbent solution flows out from the upper end of the third-stage absorption tower, it is fed into the lower end of the second-stage absorption tower, flows upward from the bottom and continues to absorb maleic anhydride. After the absorbent solution flows out from the upper end of the second-stage absorption tower, it is fed into the lower end of the first-stage absorption tower, flows upward from the bottom and further absorbs maleic anhydride, and the high-concentration absorbent solution is finally withdrawn from the upper end of the first-stage absorption tower.

[0078] In the embodiments of the present invention, the absorbent hydrogenation reactor is an adiabatic fixed-bed reactor or an isothermal fixed-bed reactor;

[0079] Such as Figure 4As shown in the figure, when an adiabatic fixed-bed reactor is selected, the adiabatic fixed-bed hydrogenation reactor includes a reactor upper head 31, a first-stage reactor 32, a second-stage reactor 33, and a reactor lower head 34 from top to bottom. Among them, a liquid-phase hydrogenation raw material inlet 35 and a hydrogen inlet 36 are provided on the reactor upper head 31, a distributor 37 is provided inside the reactor upper head 31, inert porcelain balls are filled at the top of the beds of the first-stage reactor 32 and the second-stage reactor 33, and a hydrogenation catalyst or a diesterification-hydrogenation bifunctional catalyst is filled in the reactor body; a heat exchanger 38 is provided between the first-stage reactor 32 and the second-stage reactor 33. The reaction materials of the first-stage reactor 32 enter the second-stage reactor 33 through heat exchange and cooling by the heat exchanger 38 for continuous reaction, and a reaction product outlet 39 is provided on the lower head; the liquid-phase hydrogenation material and hydrogen are mixed in the distributor 37 and enter the first-stage reactor 31 and the second-stage reactor 32 in sequence through the distributor 37.

[0080] As Figure 5 shown in the figure, when an isothermal fixed-bed reactor is selected, the isothermal fixed-bed hydrogenation reactor includes a reactor upper head 41, an isothermal tube fixed-bed reactor 42, and a reactor lower head 43 from top to bottom. Among them, a liquid-phase hydrogenation raw material inlet 44 and a hydrogen inlet 45 are provided on the reactor upper head 41, a reaction product outlet 46 is provided on the reactor lower head 43, and a distributor 47 is provided inside the reactor upper head 41; inert porcelain balls are filled at both the top and bottom of the tubes of the isothermal tube fixed-bed reactor 42, a hydrogenation catalyst or a diesterification-hydrogenation bifunctional catalyst is filled in the tube reactor body, and a heat medium M is filled between the tubes. The heat medium M enters from the bottom and exits from the top for countercurrent heat exchange with the reaction materials entering from the top and exiting from the bottom.

[0081] In an embodiment of the present invention, the absorption liquid hydrogenation reactor is selected from a cocurrent isothermal trickle fixed-bed reactor, a cocurrent adiabatic trickle fixed-bed reactor, a countercurrent adiabatic bubble fixed-bed reactor, or a cocurrent adiabatic bubble fixed-bed reactor. As Figure 6 shown in the figure, it includes the following four operation modes:

[0082] (1) Using a cocurrent isothermal trickle fixed-bed reactor, as Figure 6 shown in a, both the absorption liquid and hydrogen are sent into the reactor from the top of the trickle bed reactor through the distributor, flow cocurrently from top to bottom through the hydrogenation catalyst or diesterification-hydrogenation bifunctional catalyst bed layer, and the reacted materials are taken out from the bottom of the reactor;

[0083] (2) Using a cocurrent adiabatic trickle fixed-bed reactor, as Figure 6 shown in b: both the absorption liquid and hydrogen are sent into the reactor from the top of the trickle bed reactor through the distributor, flow cocurrently from top to bottom through the hydrogenation catalyst or diesterification-hydrogenation bifunctional catalyst bed layer, and the reacted materials are taken out from the bottom of the reactor;

[0084] (3) An adiabatic countercurrent bubble fixed-bed reactor is adopted, as shown in Figure 6 c: The absorbent liquid is fed into the reactor from the top of the bubble bed reactor through a distributor and flows downward, while hydrogen is fed into the reactor from the bottom and flows upward. The absorbent liquid and hydrogen flow countercurrently through the hydrogenation catalyst or the bed layer of the diesterification-hydrogenation bifunctional catalyst. The recycled hydrogen after the reaction is discharged from the top of the reactor, and the liquid product is withdrawn from the bottom of the reactor;

[0085] (4) A cocurrent adiabatic bubble fixed-bed reactor is adopted, as shown in Figure 6 d: The absorbent liquid and hydrogen are respectively fed into the reactor from the bottom of the bubble bed reactor through a distributor and flow cocurrently upward through the hydrogenation catalyst or the bed layer of the diesterification-hydrogenation bifunctional catalyst. The recycled hydrogen after the reaction is discharged from the top of the reactor, and the liquid product is withdrawn from the upper end of the reactor.

[0086] The present invention also provides an application of the above process system for producing succinic acid esters from maleic anhydride gas stream, which is used for producing succinic acid esters from maleic anhydride gas stream and includes the following steps:

[0087] S1. Absorption-esterification of maleic anhydride gas stream: The maleic anhydride gas stream and an alcohol absorbent containing or not containing a liquid acid catalyst are respectively fed into the absorption-esterification tower to carry out solvent absorption of maleic anhydride and an esterification reaction to prepare a maleate absorbent solution:

[0088] S2. Removal of dissolved oxygen from the absorbent liquid: The absorbent liquid after absorption-esterification of the maleic anhydride gas stream in the absorption-esterification tower is fed into a de-dissolved oxygen tower, and after being de-dissolved oxygen under negative pressure, it is fed into the absorbent liquid hydrogenation or diesterification-hydrogenation coupling system; Operating conditions: absorbent liquid temperature 52 - 150 °C and top pressure of the tower 50.0 - 90.0 kPa;

[0089] S3. Absorbent liquid hydrogenation or diesterification-hydrogenation: The absorbent liquid after removing dissolved oxygen and hydrogen are fed into an absorbent liquid hydrogenation reactor filled with a supported metal catalyst that has been pre-reduced and activated to carry out double-bond hydrogenation of maleate in the absorbent liquid or diesterification-hydrogenation of maleic monoester to prepare succinic acid esters:

[0090] S4. Gas-liquid separation of the reaction product: The material after absorbent liquid hydrogenation or diesterification-hydrogenation reaction is fed into a gas-liquid separator for gas-liquid separation, and the liquid material is subjected to subsequent succinic acid ester refining treatment after absorbent recovery.

[0091] In an embodiment of the present invention, in step S1, the operating conditions are: the concentration of the liquid acid catalyst in the alcohol absorbent is 0 - 5.0 wt%, the temperature of the maleic anhydride gas stream is 55 - 180 °C and the volume space velocity is 200 - 20000 h -1 , the temperature of the alcohol absorbent is 25 - 85 °C and the weight space velocity is 0.1 - 10.0 h -1, after absorption in an absorption esterification tower in series of one or more stages, a maleate absorption solution is obtained; preferably, the operating conditions are: the concentration of the liquid acid catalyst in the alcohol absorbent is 0-3.0 wt%, the temperature of the maleic anhydride gas stream is 60-150 °C, and the volume space velocity is 500-10,000 h -1 , the temperature of the alcohol absorbent is 30-80 °C, and the weight space velocity is 0.2-5.0 h -1 , a maleate absorption solution is obtained through series absorption in 1-3 absorption towers; more preferably, the operating conditions are: the concentration of the liquid acid catalyst in the alcohol absorbent is 0-2.5 wt%, the temperature of the maleic anhydride gas stream is 65-120 °C, and the volume space velocity is 1000-5000 h -1 , the temperature of the alcohol absorbent is 35-75 °C, and the weight space velocity is 0.5-3.0 h -1 , a maleate absorption solution is obtained through series absorption in 1-3 absorption towers.

[0092] In an embodiment of the present invention, in step S2, the temperature of the absorption solution is 55-120 °C, and the tower top pressure is 60.0-85.0 kPa. More preferably, the temperature of the absorption solution is 60-90 °C, and the tower top pressure is 70.0-80.0 kPa;

[0093] In an embodiment of the present invention, in step S3, a fixed bed reactor and a metal particle catalyst supported on a non-acidic carrier or a bifunctional catalyst of a solid acid supported on metal particles are used. The absorption solution is fed into an absorption solution hydrogenation reactor and operated in parallel or countercurrent with hydrogen for maleate hydrogenation or maleic monoester diesterification-hydrogenation to prepare succinate;

[0094] The operating conditions are: the temperature of the absorption solution is 50-150 °C, and the weight hourly space velocity is 0.1-3.0 h -1 , the hydrogen pressure is 0.5-5.0 MPa, and the hydrogen to ester molar ratio is 5-50; preferably, the temperature of the absorption solution is 55-120 °C, and the weight hourly space velocity is 0.25-2.5 h -1 , the hydrogen pressure is 0.75-3.0 MPa, and the hydrogen to ester molar ratio is 7.5-40; more preferably, the temperature of the absorption solution is 60-90 °C, and the weight hourly space velocity is 0.5-2.0 h -1 , the hydrogen pressure is 1.0-2.0 MPa, and the hydrogen to ester molar ratio is 10-30.

[0095] The technical solutions of the present invention will be described in detail below with specific embodiments

[0096] Examples 1-4: Synthesis of succinic monoester by maleic anhydride gas stream absorption-monoesterification and hydrogenation in series

[0097] Example 1

[0098] The process system for producing succinate from maleic anhydride gas stream includes the following processes:

[0099] S1. Absorption - Mono - esterification of Maleic Anhydride Gas Stream

[0100] Adopt a counter - current spray bubble absorption - esterification tower as shown in Figure 2 Figure a and a 3 - stage series absorption method with different tower feeding as shown in Figure 3 Figure a. Feed the maleic anhydride gas stream and the alcohol absorbent butanol into the absorption - esterification tower filled with corundum small ball packing with a diameter of 3.2 mm respectively. Among them,

[0101] The alcohol absorbent butanol enters from the top of the counter - current spray bubble absorption - esterification tower. The temperature of the first - stage feeding butanol is 40°C and the weight hourly space velocity is 2.0 h -1 , the temperature of the maleic anhydride gas stream is 90°C and the volume hourly space velocity is 5000 h -1 , and the maleic anhydride gas stream enters from the bottom of the tower; The absorption liquid obtained from the bottom of the counter - current spray bubble absorption - esterification tower has a temperature of 65 - 70°C, a maleic acid monobutyl ester concentration of 46 - 50 wt%, a maleic anhydride esterification rate of 100%, and a maleic acid monobutyl ester selectivity of 100%.

[0102] S2. Removal of Dissolved Oxygen from Maleic Acid Mono - ester Absorption Liquid

[0103] Feed the absorption liquid discharged from the bottom of the counter - current spray bubble absorption tower in step S1 into the de - dissolved oxygen tower. After removing the dissolved oxygen under negative pressure, it is sent to the hydrogenation reactor. Operating conditions: the temperature is the temperature of the absorption liquid obtained in step S1, and the top pressure of the de - dissolved oxygen tower is 85 kPa.

[0104] S3. Double - bond Hydrogenation of Maleic Acid Mono - ester Absorption Liquid

[0105] Feed the maleic acid mono - ester absorption liquid from S2 into a fixed - bed reactor filled with 0.5 wt% Pd / SBA - 15 catalyst that has been pre - reduced and activated as shown in Figure 4 Figure. Operate in co - current or counter - current with hydrogen to carry out the double - bond hydrogenation of the maleic acid mono - ester absorption liquid.

[0106] In this example, for the double - bond hydrogenation of maleic acid monobutyl ester absorption liquid: adopt an adiabatic fixed - bed reactor with inter - stage heat exchange as shown in Figure 4 Figure. The material inlet mode and structure of the fixed - bed reactor adopted are like a co - current adiabatic trickle fixed - bed reactor as shown in Figure 6 Figure b. The absorption liquid and hydrogen enter the fixed - bed reactor from the top in co - current, and the hydrogenation product is output from the bottom. Among them, the temperature of the maleic acid monobutyl ester absorption liquid is 60°C and the weight hourly space velocity is 0.5 h -1 , the hydrogen pressure is 2.0 MPa and the hydrogen - to - ester molar ratio is 10. The results are: the double - bond hydrogenation rate of maleic acid monobutyl ester is 100%, and the selectivity of monobutyl succinate is 100%.

[0107] S4. Gas - Liquid Separation of Hydrogenation Reaction Product of Absorption Liquid

[0108] The hydrogenation reaction material from S3 is fed into a gas-liquid separator. The hydrogen gas separated by gas-liquid separation is pressurized by a compressor and then fed into the hydrogenation reactor for recycling use, and the liquid material is fed into the succinic acid monoester refining system after being recovered by an absorbent.

[0109] Example 2

[0110] S1. Absorption - monoesterification of maleic anhydride gas stream:

[0111] Adopt the Figure 2 countercurrent spray bubble absorption - esterification tower shown in a and the Figure 3 same - tower feed 1 - stage absorption esterification tower shown in b (only use the Figure 3 first - stage absorption esterification tower of b) for the absorption method, and feed the maleic anhydride gas stream and the alcohol absorbent butanol into the absorption esterification tower filled with corundum small balls with a diameter of 3.2 mm respectively. Among them,

[0112] the alcohol absorbent enters from the top of the countercurrent spray bubble absorption - esterification tower, the feed temperature of 1,4 - butanediol is 60 °C and the weight hourly space velocity is 1.0 h -1 , the maleic anhydride gas stream temperature is 90 °C and the volume hourly space velocity is 1000 h -1 , the maleic anhydride gas stream enters from the bottom of the tower, and the obtained absorption liquid has a temperature of ~80 °C and a concentration of maleic acid monohydroxybutyl ester of ~35 wt%, the maleic anhydride esterification rate is 100%, and the selectivity of maleic acid monohydroxybutyl ester is 100%.

[0113] S2. Removal of dissolved oxygen from the maleic acid monoester absorption liquid:

[0114] Feed the absorption liquid discharged from the bottom of the countercurrent spray bubble absorption tower in step S1 into a de - dissolved oxygen tower, and after removing the dissolved oxygen under negative pressure, feed it into the hydrogenation reactor. Operating conditions: the temperature is the temperature of the absorption liquid in step S1, and the pressure at the top of the de - dissolved oxygen tower is 70 kPa.

[0115] S3. Double - bond hydrogenation of the maleic acid monoester absorption liquid:

[0116] Feed the maleic acid monoester absorption liquid from S2 into the Figure 4 fixed - bed reactor filled with 0.5 wt% Pd / SBA - 15 catalyst that has been pre - reduced and activated as shown in Figure 4 , and operate in co - current or counter - current with hydrogen gas to carry out the double - bond hydrogenation of the maleic acid monoester absorption liquid. In this example, for the double - bond hydrogenation of the maleic acid monohydroxybutyl ester absorption liquid, use the Figure 6 inter - stage heat - exchange adiabatic fixed - bed reactor shown in b and the -1 co - current adiabatic trickle - bed fixed - bed reactor with the structure shown in b. The temperature of the maleic acid monohydroxybutyl ester absorption liquid is 80 °C and the weight hourly space velocity is 0.75 h -1 , the hydrogen pressure is 1.5 MPa and the hydrogen - ester molar ratio is 20. The results are: the double - bond hydrogenation rate of maleic acid monohydroxybutyl ester is 100%, and the selectivity of monohydroxybutyl succinate is 99.3%.

[0117] S4. Gas-liquid separation of the hydrogenation reaction product of the absorbent:

[0118] Feed the hydrogenation reaction material from S3 into a gas-liquid separator. The hydrogen gas separated by gas-liquid separation is pressurized by a compressor and then sent back to the hydrogenation reactor for recycling, and the liquid material is sent to the succinic acid monoester refining system after being recovered by the absorbent.

[0119] Example 3

[0120] S1. Absorption-monoesterification of maleic anhydride gas stream:

[0121] Adopt a co-current bubbling absorption esterification tower as shown in Figure 2 c and a counter-current feeding two-stage series absorption method as shown in Figure 3 c (only use the first and second stage absorption esterification towers of Figure 3 c), and feed the maleic anhydride gas stream and the alcohol absorbent lauryl alcohol into the absorption esterification tower filled with corundum balls with a diameter of 3.2 mm respectively. Among them, both the alcohol absorbent and the maleic anhydride gas stream enter from the bottom of the co-current bubbling absorption esterification tower. The temperature of the maleic anhydride gas stream is 120 °C and the volumetric space velocity is 3000 h -1 , the temperature of the lauryl alcohol fed in the first stage is 65 °C and the weight space velocity is 2.0 h -1 ; the obtained absorbent has a temperature of ~95 °C and a concentration of monolauryl maleate of ~42 wt%, the esterification rate of maleic anhydride is 100%, and the selectivity of monolauryl maleate is 100%.

[0122] S2. Removal of dissolved oxygen from the monolauryl maleate absorbent:

[0123] Feed the absorbent discharged from the bottom of the counter-current spray bubbling absorption tower in step S1 into a de-dissolved oxygen tower, and after removing dissolved oxygen under negative pressure, send it to the hydrogenation reactor. Operating conditions: the temperature is the temperature of the absorbent in step S1, and the pressure at the top of the de-dissolved oxygen tower is 75 kPa.

[0124] S3. Double bond hydrogenation of the monolauryl maleate absorbent:

[0125] Feed the monolauryl maleate absorbent from S2 into a fixed bed reactor filled with 0.5 wt% Pd / SBA-15 catalyst that has been pre-reduced and activated, and operate in co-current or counter-current with hydrogen gas to carry out double bond hydrogenation of the monolauryl maleate absorbent. In this example, for the double bond hydrogenation of the monolauryl maleate absorbent, a tubular isothermal fixed bed reactor as shown in Figure 5 is used. The material inlet mode and structure of the fixed bed reactor adopted are those of a co-current isothermal trickle fixed bed reactor with the structure shown in Figure 5 a. The temperature of the monolauryl maleate absorbent is 85 °C and the weight hourly space velocity is 1.0 h Figure 6 a, and the structure is a co-current isothermal trickle fixed bed reactor as shown in -1, hydrogen pressure of 1.0 MPa and a hydrogen to ester molar ratio of 25. The results were as follows: the double bond hydrogenation rate of lauryl maleate was 100%, and the selectivity of monolauryl succinate was 99.1%.

[0126] S4. Gas-liquid separation of the hydrogenation reaction product of the absorbent:

[0127] The hydrogenation reaction material from S3 is sent to a gas-liquid separator. The hydrogen gas separated by gas-liquid separation is pressurized by a compressor and then sent back to the hydrogenation reactor for recycling. The liquid material is sent to the monolauryl succinate refining system after being recovered by the absorbent.

[0128] Example 4

[0129] S1. Absorption - monoesterification of maleic anhydride gas stream:

[0130] Adopt the co-current bubble absorption esterification tower as shown in Figure 2 c and the same tower feed two-stage series absorption method as shown in Figure 3 d (only use the first and second stage absorption esterification towers of Figure 3 d). The maleic anhydride gas stream and the alcohol absorbent polyethylene glycol 400 (PEG - 400) are respectively sent into the absorption esterification tower filled with corundum balls with a diameter of 3.2 mm. Among them, both the alcohol absorbent and the maleic anhydride gas stream enter from the bottom of the co-current bubble absorption esterification tower. The temperature of the maleic anhydride gas stream is 120 °C and the volume space velocity is 2500 h -1 , the temperature of PEG - 400 is 70 °C and the weight space velocity is 2.5 h -1 ; the temperature of the obtained absorbent is ~98 °C and the concentration of monolauryl maleate in PEG - 400 is ~45 wt%. The esterification rate of maleic anhydride is 100%, and the selectivity of monolauryl maleate in PEG - 400 is 100%.

[0131] S2. Removal of dissolved oxygen from the monolauryl maleate absorbent:

[0132] The absorbent discharged from the bottom of the countercurrent spray bubble absorption tower in step S1 is sent into a de-dissolved oxygen tower, and after the dissolved oxygen is removed under negative pressure, it is sent into the hydrogenation reactor. The operating conditions are: the temperature is the temperature of the absorbent, and the pressure at the top of the de-dissolved oxygen tower is 70 kPa.

[0133] S3. Double bond hydrogenation of the monolauryl maleate absorbent:

[0134] The monolauryl maleate absorbent from S2 is sent into a fixed bed reactor filled with 0.5 wt% Pd / SBA - 15 catalyst that has been pre-reduced and activated as shown in Figure 5 and operated in co-current or counter-current with hydrogen gas for the double bond hydrogenation of the monolauryl maleate absorbent. In this example, for the double bond hydrogenation of monolauryl maleate in PEG - 400, a tubular isothermal fixed bed reactor as shown in Figure 5 and an attached Figure 6Co-current isothermal trickle-bed reactor with the structure shown in a, temperature of maleic acid PEG-400 monoester absorbent is 90 °C and weight hourly space velocity is 1.5 h -1 , hydrogen pressure is 1.0 MPa and molar ratio of hydrogen to ester is 30. The results are as follows: hydrogenation rate of double bond of maleic acid PEG-400 ester is 100%, and selectivity of succinic acid PEG-400 monoester is 98.9%.

[0135] S4. Gas-liquid separation of hydrogenation reaction product of absorbent:

[0136] Feed the hydrogenation reaction material from S3 into the gas-liquid separator. The hydrogen separated by gas-liquid separation is pressurized by a compressor and then sent back to the hydrogenation reactor for recycling, and the liquid material is sent to the succinic acid monoester refining system after being recovered by the absorbent.

[0137] Examples 5-8: Synthesis of diethyl succinate by tandem reaction of maleic anhydride gas absorption - di-esterification and hydrogenation

[0138] Example 5

[0139] S1. Maleic anhydride gas absorption - di-esterification:

[0140] Feed the maleic anhydride gas stream and the absorbent containing liquid acid catalyst into the solvent absorption - esterification coupling reaction tower filled with quartz balls with a diameter of 2.6 mm or solid acid catalyst shown in a-2d respectively. The maleic anhydride gas stream and the absorbent containing liquid acid catalyst flow through the packing or catalyst bed of the 1-3 stage series absorption - esterification tower shown in parallel or countercurrent, and the absorption - di-esterification coupling reaction of maleic anhydride occurs. Figure 2 shown, and the maleic anhydride gas stream and the absorbent containing liquid acid catalyst flow through the packing or catalyst bed of the 1-3 stage series absorption - esterification tower shown in parallel or countercurrent, and the absorption - di-esterification coupling reaction of maleic anhydride occurs. Figure 3 In this example, for the preparation of dibutyl maleate absorbent, a countercurrent spray bubbling absorption esterification tower filled with quartz balls shown in a and a co-tower feeding 2-stage series absorption method shown in b (only using the first and second stage absorption esterification towers of b) are adopted. The temperature of the maleic anhydride gas stream is 90 °C and the volume hourly space velocity is 2500 h

[0141] In this example, for the preparation of dibutyl maleate absorbent, a countercurrent spray bubbling absorption esterification tower filled with quartz balls shown in a and a co-tower feeding 2-stage series absorption method shown in b (only using the first and second stage absorption esterification towers of b) are adopted. The temperature of the maleic anhydride gas stream is 90 °C and the volume hourly space velocity is 2500 h Figure 2 shown, the temperature of the first stage feed butanol is 60 °C and the weight hourly space velocity is 2.0 h Figure 3 shown, the concentration of trifluoromethanesulfonic acid in the absorbent butanol is 2.0 wt%; the obtained absorbent temperature is ~80 °C and the dibutyl maleate concentration is ~43 wt%, the esterification rate of maleic anhydride is 100%, and the selectivity of dibutyl maleate is 98.5%. Figure 3 shown, the temperature of the first stage feed butanol is 60 °C and the weight hourly space velocity is 2.0 h -1 , the concentration of trifluoromethanesulfonic acid in the absorbent butanol is 2.0 wt%; the obtained absorbent temperature is ~80 °C and the dibutyl maleate concentration is ~43 wt%, the esterification rate of maleic anhydride is 100%, and the selectivity of dibutyl maleate is 98.5%. -1 , the concentration of trifluoromethanesulfonic acid in the absorbent butanol is 2.0 wt%; the obtained absorbent temperature is ~80 °C and the dibutyl maleate concentration is ~43 wt%, the esterification rate of maleic anhydride is 100%, and the selectivity of dibutyl maleate is 98.5%.

[0142] S2. Removal of dissolved oxygen from maleic acid diester absorbent:

[0143] Feed the absorbent from S1 into the dissolved oxygen removal tower, and after removing the dissolved oxygen under negative pressure, send it to the hydrogenation reactor. Operating conditions: temperature of the absorbent, pressure at the top of the dissolved oxygen removal tower is 75 kPa.

[0144] S3. Double bond hydrogenation of maleic acid diester absorbent:

[0145] The maleic acid diester absorbent from S2 is fed into the fixed-bed reactor filled with 0.5 wt% Pd / HM catalyst pre-reduced and activated, and operated in co-current or counter-current with hydrogen for double bond hydrogenation of the maleic acid diester absorbent. In this example, for the double bond hydrogenation of dibutyl maleate absorbent, a tubular isothermal fixed-bed reactor shown in Figure 5 and a co-current isothermal trickle fixed-bed reactor with the structure shown in Figure 5 are used. The temperature of the dibutyl maleate absorbent is 80 °C, the weight hourly space velocity is 0.5 h Figure 6 -1, the hydrogen pressure is 1.5 MPa, and the hydrogen to ester molar ratio is 15. The results are as follows: the double bond hydrogenation rate of dibutyl maleate is 100%, and the selectivity of dibutyl succinate is 100%. -1

[0146] S4. Gas-liquid separation of hydrogenation reaction products:

[0147] The hydrogenation reaction material from S3 is fed into the gas-liquid separator. The separated hydrogen is pressurized by a compressor and then sent back to the hydrogenation reactor for recycling, and the liquid material is sent to the succinic acid ester refining system after being recovered by the absorbent.

[0148] Example 6

[0149] S1. Maleic anhydride gas absorption - di-esterification:

[0150] The maleic anhydride gas stream and the absorbent containing liquid acid catalyst are respectively fed into the solvent absorption - esterification coupling reaction tower filled with quartz balls with a diameter of 2.6 mm or solid acid catalyst shown in Figure 2 a - 2d. The maleic anhydride gas stream and the absorbent containing liquid acid catalyst flow through the packing or catalyst bed of the 1 - 3 stage series absorption - esterification tower shown in Figure 3 in co-current or counter-current to carry out the absorption - di-esterification coupling reaction of maleic anhydride.

[0151] In this example, for the preparation of diisooctyl maleate absorbent, a counter-current spray bubble absorption esterification tower filled with HM zeolite catalyst pellets shown in Figure 2 b and a 3-stage absorption method with different tower feeds shown in Figure 3 a are used. The temperature of the maleic anhydride gas stream is 120 °C and the volume hourly space velocity is 10000 h -1 , the temperature of isooctanol in the first-stage feed is 70 °C and the weight hourly space velocity is 3.0 h -1 , the concentration of p-toluenesulfonic acid in the absorbent isooctanol is 3.0 wt%; the obtained absorbent temperature is ~98 °C and the concentration of diisooctyl maleate is ~40 wt%, the esterification rate of maleic anhydride is 100%, and the selectivity of diisooctyl maleate is 98.2%.

[0152] S2. Maleic acid diester absorption liquid deoxygenation:

[0153] The absorption liquid from S1 is sent to the deoxygenation tower, and after the dissolved oxygen is removed by negative pressure, it is sent to the hydrogenation reactor. The operating conditions are: the temperature of the absorption liquid and the top pressure of the deoxygenation tower is 60kPa.

[0154] S3. Double bond hydrogenation of maleic acid diester absorption liquid:

[0155] The maleic acid diester absorption liquid from S2 is sent to the attached Figure 5 In the fixed bed reactor filled with 0.5 wt% Pd / HM catalyst which has been previously reduced and activated, double bond hydrogenation of maleic acid diester absorption liquid is carried out in co-current or counter-current operation with hydrogen. Figure 5 The isothermal fixed bed reactor and the attached Figure 6 The co-current isothermal trickle fixed bed reactor with the structure shown in a, the temperature of the diisooctyl maleate absorption liquid is 98°C and the weight hourly space velocity is 1.0h -1 , hydrogen pressure 1.0 MPa and hydrogen ester molar ratio 20. The results are: double bond hydrogenation rate of diisooctyl maleate is 100%, and selectivity of diisooctyl succinate is 99.2%.

[0156] S4. Gas-liquid separation of hydrogenation reaction products:

[0157] The hydrogenation reaction materials from S3 are sent to the gas-liquid separator, the hydrogen separated from the gas and liquid is pressurized by the compressor and sent to the hydrogenation reactor for recycling, and the liquid materials are recovered by the absorbent and sent to the succinate refining system.

[0158] Example 7

[0159] S1. Maleic anhydride gas flow absorption-diesterification:

[0160] The maleic anhydride gas stream and the absorbent containing the liquid acid catalyst are respectively sent to the adjacent Figure 2 In the solvent absorption-esterification coupling reaction tower filled with quartz ball filler or solid acid catalyst shown in a to 2d, the maleic anhydride gas flow and the absorbent containing liquid acid catalyst flow in parallel or countercurrent through the attached Figure 3 The packing or catalyst bed of the 1 to 3-stage absorption-esterification tower connected in series causes the absorption-diesterification coupling reaction of maleic anhydride.

[0161] In this embodiment, the preparation of difurfuryl maleate absorption liquid: using the attached Figure 2 The parallel flow bubbling absorption esterification tower filled with quartz balls shown in c and the attached Figure 3 c shows a two-stage series absorption method with different tower feed (only using Figure 3 c) with a maleic anhydride gas flow temperature of 120 °C and a volume space velocity of 3000 h-1 , the temperature of the first-stage feed furfuryl alcohol is 65 °C and the weight hourly space velocity is 1.5 h -1 , the concentration of trifluoromethanesulfonic acid in the absorbent furfuryl alcohol is 2.5 wt%; the temperature of the obtained absorption liquid is ~95 °C and the concentration of di-furfuryl maleate is ~38 wt%, the esterification rate of maleic anhydride is 100%, and the selectivity of di-furfuryl maleate is 98.6%.

[0162] S2. Removal of dissolved oxygen from the di-furfuryl maleate absorption liquid:

[0163] The absorption liquid from S1 is sent to a deaeration tower, and after removing dissolved oxygen under negative pressure, it is sent to a hydrogenation reactor. Operating conditions: the temperature of the absorption liquid, and the pressure at the top of the deaeration tower is 80 kPa.

[0164] S3. Double-bond hydrogenation of the di-furfuryl maleate absorption liquid:

[0165] The di-furfuryl maleate absorption liquid from S2 is sent to the fixed-bed reactor filled with 0.5 wt% Pd / HM catalyst that has been pre-reduced and activated as shown in Figure 4 . It is operated in parallel or countercurrent with hydrogen to carry out the double-bond hydrogenation of the di-furfuryl maleate absorption liquid. In this example, for the double-bond hydrogenation of the di-furfuryl maleate absorption liquid: a staged-intercooled adiabatic fixed-bed reactor as shown in Figure 4 and a co-current adiabatic trickle fixed-bed reactor with the structure shown in Figure 6 b are used. The temperature of the di-furfuryl maleate absorption liquid is 95 °C and the weight hourly space velocity is 1.0 h -1 , the hydrogen pressure is 2.0 MPa, and the hydrogen-to-ester molar ratio is 10. The results are: the double-bond hydrogenation rate of di-furfuryl maleate is 100%, and the selectivity of di-furfuryl succinate is 99.3%.

[0166] S4. Gas-liquid separation of the hydrogenation reaction product:

[0167] The hydrogenation reaction material from S3 is sent to a gas-liquid separator. The hydrogen separated from the gas-liquid separation is pressurized by a compressor and sent back to the hydrogenation reactor for recycling, and the liquid material is sent to the di-furfuryl succinate refining system after the absorbent is recovered.

[0168] Example 8

[0169] S1. Absorption-esterification of maleic anhydride gas stream:

[0170] The maleic anhydride gas stream and the absorbent containing a liquid acid catalyst are respectively sent into the solvent absorption-esterification coupling reaction tower filled with quartz balls with a diameter of 2.6 mm or a solid acid catalyst as shown in Figure 2 a~2d. The maleic anhydride gas stream and the absorbent containing a liquid acid catalyst flow through the packing or catalyst bed of the 1~3-stage series absorption-esterification tower as shown in Figure 3 in parallel or countercurrent, and the absorption-esterification coupling reaction of maleic anhydride occurs.

[0171] In this example, the preparation of the dihydroxyethyl maleate absorbent: Using the Figure 2 cocurrent bubble absorption esterification tower filled with HZSM-5 zeolite catalyst pellets shown in Figure 3 Figure d and the same tower feed 2-stage series absorption method shown in Figure 3 Figure d (only using the first and second stage absorption esterification towers of -1 Figure d), the maleic anhydride gas stream temperature is 120 °C and the volume space velocity is 1500 h -1 , the ethylene glycol temperature is 60 °C and the weight space velocity is 1.5 h

[0172] S2. Removal of dissolved oxygen from the dihydroxyethyl maleate absorbent:

[0173] Send the absorbent from S1 to the deaeration tower, and after removing dissolved oxygen under negative pressure, send it to the hydrogenation reactor. Operating conditions: the temperature of the absorbent, and the top pressure of the deaeration tower is 65 kPa.

[0174] S3. Double bond hydrogenation of the dihydroxyethyl maleate absorbent:

[0175] Send the dihydroxyethyl maleate absorbent from S2 into the Figure 4 fixed bed reactor filled with 0.5 wt% Pd / HM catalyst that has been pre-reduced and activated shown in

[0176] Figure, and operate in cocurrent or countercurrent with hydrogen to carry out double bond hydrogenation of the dihydroxyethyl maleate absorbent. Figure 4 In this example, the double bond hydrogenation of the dihydroxyethyl maleate absorbent: Using the Figure 6 interstage heat exchange adiabatic fixed bed reactor shown in Figure b and the cocurrent adiabatic trickle fixed bed reactor with the structure shown in -1 Figure b, the dihydroxyethyl maleate absorbent temperature is 92 °C and the weight hourly space velocity is 0.75 h

[0177] S4. Gas-liquid separation of the hydrogenation reaction product:

[0178] Send the hydrogenation reaction material from S3 to the gas-liquid separator. The separated hydrogen is pressurized by a compressor and sent back to the hydrogenation reactor for recycling, and the liquid material is sent to the succinic acid ester refining system after absorbent recovery.

[0179] Examples 9-12 Maleic anhydride gas stream absorption - monoesterification and diesterification - hydrogenation tandem synthesis of succinic acid diester

[0180] Example 9

[0181] S1. Maleic anhydride gas flow absorption-monoesterification:

[0182] The maleic anhydride gas stream and the alcohol absorbent are respectively sent to the Figure 2 In the solvent absorption-monoesterification coupling reaction tower filled with 2.8 mm diameter silicon carbide beads shown in a to 2d, the maleic anhydride gas flow and the absorbent flow in parallel or countercurrent through the attached Figure 3 In the packing bed of the 1 to 3-stage series absorption and esterification tower shown, maleic anhydride absorption-monoesterification coupling reaction occurs to generate absorption liquids of monocyclohexyl maleate, mono(diethylamino)ethyl ester, monohydroxybutyl ester or monohydroxybutene ester respectively.

[0183] In this embodiment, cyclohexanol is used as an alcohol absorbent to prepare the monocyclohexyl maleate absorption liquid: Figure 2 a) a countercurrent spray bubbling absorption esterification tower filled with silicon carbide beads and Figure 3 b shows the same tower feed two-stage series absorption method (only using Figure 3 b) with a maleic anhydride gas flow temperature of 80 °C and a volume space velocity of 3000 h -1 The first stage feed cyclohexanol temperature is 45°C and the weight space velocity is 2.0h -1 The obtained absorption liquid temperature is ~65℃ and the concentration of monocyclohexyl maleate is ~45wt%, the maleic anhydride esterification rate is 100%, and the monocyclohexyl maleate selectivity is 100%.

[0184] S2. Maleic acid monoester absorption liquid deoxygenation:

[0185] The absorption liquid from S1 is sent to the deoxygenation tower, and after the dissolved oxygen is removed by negative pressure, it is sent to the hydrogenation reactor. The operating conditions are: the absorption liquid temperature and the deoxygenation tower top pressure are 75kPa.

[0186] S3. Diesterification of maleic acid monoester absorption liquid - double bond hydrogenation:

[0187] 2.0 wt% of bis(trifluoromethanesulfonyl)imide catalyst was added to the maleic acid monoester absorption liquid from S2, and then fed to the attached Figure 5 The fixed bed reactor filled with a pre-reduced activated supported metal catalyst is operated in co-current or counter-current with hydrogen to carry out diesterification and double bond hydrogenation of the maleic acid monoester absorption liquid.

[0188] In this embodiment, the maleic acid monocyclohexyl ester is diesterified and double bond hydrogenated by using Figure 5 The isothermal fixed bed reactor shown in FIG. Figure 6 a co-current isothermal trickle fixed bed reactor with the structure shown in FIG. 1, the hydrogenation catalyst is 0.5 wt% Pd / HM, the temperature of the monocyclohexyl maleate absorption liquid is 85°C and the weight hourly space velocity is 0.5 h -1, hydrogen pressure of 2.5 MPa and a hydrogen to ester molar ratio of 20. The results were: the diesterification rate of monocyclohexyl maleate was 98.2%, the double bond hydrogenation rate was 100%, and the selectivity of dicyclohexyl succinate was 98.5%.

[0189] S4. Gas-liquid separation of the diesterification-hydrogenation product:

[0190] Feed the diesterification-hydrogenation coupling reaction material from S3 into a gas-liquid separator. The separated hydrogen is pressurized by a compressor and then sent back to the diesterification-hydrogenation reactor for recycling, and the liquid material is sent to the succinate refining system after being recovered by an absorbent.

[0191] Example 10

[0192] S1. Maleic anhydride gas absorption - monoesterification:

[0193] Preparation of the absorption solution of mono(diethylamino)ethyl maleate: Use the Figure 2 countercurrent spray bubble absorption esterification tower filled with silicon carbide balls with a diameter of 2.8 mm shown in Figure 3 a and the -1 hetero-tower feeding three-stage absorption method shown in -1 a. Feed the maleic anhydride gas stream and the alcohol absorbent diethylaminoethanol into the countercurrent spray bubble absorption esterification tower respectively. The temperature of the maleic anhydride gas stream is 90 °C and the volume space velocity is 5000 h

[0194] S2. Removal of dissolved oxygen from the monoester absorption solution:

[0195] Feed the absorption solution from S1 into a deaeration tower, and after removing dissolved oxygen under negative pressure, send it to a hydrogenation reactor. The operating conditions are: the temperature of the absorption solution, and the pressure at the top of the deaeration tower is 70 kPa.

[0196] S3. Diesterification - double bond hydrogenation of the monoester absorption solution:

[0197] Add 2.0 wt% of the bis(trifluoromethanesulfonyl)imide catalyst to the monoester absorption solution from S2, and then feed it into a fixed bed reactor filled with a pre-reduced and activated supported metal catalyst, and operate in co-current or counter-current with hydrogen to carry out the diesterification - double bond hydrogenation of the monoester absorption solution.

[0198] In this example, for the diesterification - double bond hydrogenation of mono(diethylamino)ethyl maleate, use the Figure 4 inter-stage heat exchange adiabatic fixed bed reactor shown in Figure 6The co-current adiabatic trickle-bed fixed-bed reactor with the structure shown in b, the hydrogenation catalyst is 0.5 wt% Pd / SiO2, the temperature of the maleic acid mono(diethylamino)ethyl ester absorbent is 80 °C and the weight hourly space velocity is 0.75 h -1 , the hydrogen pressure is 1.0 MPa and the hydrogen-ester molar ratio is 10. The results are as follows: the diesterification rate of maleic acid mono(diethylamino)ethyl ester is 98.5%, the double bond hydrogenation rate is 100%, and the selectivity of di(diethylamino)ethyl succinate is 99.2%.

[0199] S4. Gas-liquid separation of the diesterification-hydrogenation product:

[0200] The diesterification-hydrogenation coupling reaction material from S3 is sent into a gas-liquid separator. The separated hydrogen is pressurized by a compressor and then sent into the diesterification-hydrogenation reactor for recycling, and the liquid material is sent into the succinate refining system after being recovered by the absorbent.

[0201] Example 11

[0202] S1. Maleic anhydride gas absorption-monoesterification:

[0203] The maleic anhydride gas stream and the alcohol absorbents cyclohexanol, diethylaminoethanol, 1,4-butanediol or butenediol are respectively sent into the solvent absorption-monoesterification coupling reaction tower filled with silicon carbide spheres with a diameter of 2.8 mm shown in Attachment Figure 2 a to 2d. The maleic anhydride gas stream and the absorbent flow co-currently or counter-currently through the packing bed of the 1-3 stage series absorption-esterification tower shown in Attachment Figure 3 to generate absorbents of maleic acid monocylcohexyl ester, mono(diethylamino)ethyl ester, monohydroxybutyl ester or monohydroxybutenyl ester through the absorption-monoesterification coupling reaction of maleic anhydride. Among them:

[0204] For the preparation of the maleic acid monohydroxybutyl ester absorbent, a co-current bubbling absorption-esterification tower filled with silicon carbide spheres with a diameter of 2.8 mm shown in Figure 2 c and the different tower feeding 2-stage series absorption method shown in Figure 3 c (only using the first and second stage absorption-esterification towers of Figure 3 c) are used. The maleic anhydride gas stream and the alcohol absorbent 1,4-butanediol are respectively sent into the co-current bubbling absorption-esterification tower. The temperature of the maleic anhydride gas stream is 95 °C and the volume hourly space velocity is 5000 h -1 , the temperature of the 1,4-butanediol fed in the first stage is 50 °C and the weight hourly space velocity is 3.0 h -1 ; the obtained absorbent has a temperature of ~78 °C and a maleic acid monohydroxybutyl ester concentration of ~39 wt%, the maleic anhydride esterification rate is 100%, and the selectivity of maleic acid monohydroxybutyl ester is 98.8%.

[0205] S2. Removal of dissolved oxygen from the maleic acid monoester absorbent:

[0206] The absorbent liquid from S1 is fed into the deaeration tower, and after the dissolved oxygen is removed under negative pressure, it is fed into the hydrogenation reactor. The operating conditions are as follows: the temperature of the absorbent liquid, and the pressure at the top of the deaeration tower is 80 kPa.

[0207] S3. Diesterification - double bond hydrogenation of maleic acid monoester absorbent liquid:

[0208] Add 2.0 wt% of bis(trifluoromethanesulfonyl)imide catalyst to the maleic acid monoester absorbent liquid from S2, and then feed it into the Figure 4 fixed - bed reactor filled with the pre - reduced and activated supported metal catalyst as shown, and operate in co - current or counter - current with hydrogen to carry out the diesterification - double bond hydrogenation of maleic acid monoester absorbent liquid.

[0209] In this example, for the diesterification - double bond hydrogenation of monohydroxybutyl maleate: use Figure 4 the inter - stage heat - exchange adiabatic fixed - bed reactor as shown and Figure 6 the co - current adiabatic trickle - bed fixed - bed reactor with the structure shown in b. The hydrogenation catalyst is 0.5 wt% Pd / [(WO3 / ZrO2) / HMCM - 41]. The temperature of the monohydroxybutyl maleate absorbent liquid is 90 °C and the weight hourly space velocity is 0.5 h -1 、the hydrogen pressure is 1.0 MPa and the hydrogen - to - ester molar ratio is 20. The results are as follows: the diesterification rate of monohydroxybutyl maleate is 99.3%, the double - bond hydrogenation rate is 100%, and the selectivity of dihydroxybutyl succinate is 99.0%.

[0210] S4. Gas - liquid separation of the diesterification - hydrogenation product:

[0211] Feed the diesterification - hydrogenation coupling reaction material from S3 into the gas - liquid separator. The separated hydrogen is pressurized by a compressor and then fed into the diesterification - hydrogenation reactor for recycling. The liquid material is sent to the succinic acid ester refining system after the absorbent is recovered.

[0212] Example 12

[0213] S1. Maleic anhydride gas absorption - monoesterification:

[0214] Feed the maleic anhydride gas stream and the alcohol absorbents cyclohexanol, diethylaminoethanol, 1,4 - butanediol or butenediol into the Figure 2 solvent absorption - monoesterification coupling reaction tower filled with 2.8 - mm - diameter silicon carbide ball packing as shown in a~2d respectively. The maleic anhydride gas stream and the absorbent flow through the Figure 3 packing bed of the 1 - 3 - stage series - connected absorption - esterification tower as shown in parallel or counter - current, and the absorption - monoesterification coupling reaction of maleic anhydride occurs to generate the absorbent liquids of maleic acid monocylclohexyl ester, mono(diethylamino)ethyl ester, monohydroxybutyl ester or monohydroxybutenyl ester respectively. Among them:

[0215] Preparation of monohydroxybutenyl maleate absorbent liquid: Use Figure 2The co-current bubble absorption esterification tower filled with silicon carbide balls with a loading diameter of 2.8 mm as shown in d and Figure 3 The same tower feed 2-stage series absorption method as shown in d (only using Figure 3 The first and second stage absorption esterification towers of d), the maleic anhydride gas stream and the alcohol absorbent butylene glycol are respectively sent into the co-current bubble absorption esterification tower. The temperature of the maleic anhydride gas stream is 90 °C and the volumetric space velocity is 2500 h -1 , the temperature of butylene glycol is 45 °C and the weight space velocity is 2.0 h -1 ; The obtained absorption liquid has a temperature of ~72 °C and a concentration of monohydroxybutenyl maleate of ~42 wt%, the esterification rate of maleic anhydride is 100%, and the selectivity of monohydroxybutenyl maleate is 99.2%.

[0216] S2. Removal of dissolved oxygen from the monoesters of maleic acid absorption liquid:

[0217] The absorption liquid from S1 is sent into a deaeration tower, and after removing dissolved oxygen under negative pressure, it is sent into a hydrogenation reactor. The operating conditions are: the temperature of the absorption liquid, and the pressure at the top of the deaeration tower is 85 kPa.

[0218] S3. Diesterification-double bond hydrogenation of the monoesters of maleic acid absorption liquid:

[0219] Add 2.0 wt% of bis(trifluoromethanesulfonyl)imide catalyst to the monoesters of maleic acid absorption liquid from S2, and then send it into the Figure 5 Fixed bed reactor filled with a pre-reduced and activated supported metal catalyst as shown, and operate in co-current or counter-current with hydrogen to carry out the diesterification-double bond hydrogenation of the monoesters of maleic acid absorption liquid.

[0220] In this example, the diesterification-double bond hydrogenation of monohydroxybutenyl maleate: use Figure 5 The tubular isothermal fixed bed reactor as shown and Figure 6 The co-current isothermal trickle fixed bed reactor with the structure shown in a. The hydrogenation catalyst is 0.5 wt% Pd / (Cs 2.5 H 0.5 PW 12 O 40 / SBA-15), the temperature of the monohydroxybutenyl maleate absorption liquid is 95 °C and the weight hourly space velocity is 1.0 h -1 、The hydrogen pressure is 1.5 MPa and the hydrogen / ester molar ratio is 15. The results are: the diesterification rate of monohydroxybutenyl maleate is 99.6%, the hydrogenation rate of the maleoyl double bond is 100%, and the selectivity of dihydroxybutenyl succinate is 98.6%.

[0221] S4. Gas-liquid separation of the diesterification-hydrogenation product:

[0222] The reaction materials for the two-esterification and hydrogenation coupling from S3 are sent to a gas-liquid separator. The separated hydrogen gas is pressurized by a compressor and then sent to the two-esterification and hydrogenation reactor for recycling. The liquid materials are sent to the succinic acid ester refining system after being recovered by an absorbent.

[0223] Examples 13 to 16: Synthesis of diethyl succinate by the tandem combination of maleic anhydride gas absorption-esterification and two-esterification-hydrogenation

[0224] Example 13

[0225] S1. Maleic anhydride gas absorption-esterification:

[0226] The maleic anhydride gas stream and the absorbent are respectively sent into the solvent absorption-esterification coupling reaction tower filled with solid acid catalysts with a loading diameter of 2.8 mm as shown in FIGS. a to 2d. The maleic anhydride gas stream and the absorbent flow through the catalyst beds of the 1st to 3rd stage series absorption-esterification towers shown in FIGS. in parallel or countercurrent, and the absorption-esterification coupling reaction of maleic anhydride occurs. Figure 2 In this example, the preparation of the maleic acid mono / diallyl ester absorbent: Use the countercurrent spray bubble absorption-esterification tower filled with HM pellets as shown in FIG. a and the same tower feeding 2-stage series absorption method as shown in FIG. b (only use the first and second stage absorption-esterification towers of FIG. b). The maleic anhydride gas stream and the absorbent allyl alcohol are respectively sent into the countercurrent spray bubble absorption-esterification tower. The temperature of the maleic anhydride gas stream is 90 °C and the volume space velocity is 2500 h Figure 3 ; The temperature of the first-stage feeding allyl alcohol is 60 °C and the weight space velocity is 2.0 h

[0227] ; The obtained absorbent has a temperature of ~78 °C and a maleic acid allyl ester concentration of ~40 wt%, a maleic anhydride esterification rate of 100%, and a maleic acid allyl ester selectivity of 100%. Figure 2 Figure 3 Figure 3

[0228] b. The first and second stage absorption-esterification towers), the maleic anhydride gas stream and the absorbent allyl alcohol are respectively sent into the countercurrent spray bubble absorption-esterification tower. The temperature of the maleic anhydride gas stream is 90 °C and the volume space velocity is 2500 h -1 ; The temperature of the first-stage feeding allyl alcohol is 60 °C and the weight space velocity is 2.0 h -1 ; The obtained absorbent has a temperature of ~78 °C and a maleic acid allyl ester concentration of ~40 wt%, a maleic anhydride esterification rate of 100%, and a maleic acid allyl ester selectivity of 100%.

[0228] S2. Removal of dissolved oxygen from the maleic acid ester absorbent:

[0229] The absorbent from S1 is sent into a dissolved oxygen removal tower, and after the dissolved oxygen is removed under negative pressure, it is sent into a hydrogenation reactor. Operating conditions: the temperature of the absorbent, the pressure at the top of the dissolved oxygen removal tower is 80 kPa.

[0230] S3. Two-esterification and double bond hydrogenation of the maleic acid ester absorbent:

[0231] 2.5 wt% of p-toluenesulfonic acid catalyst is added to the maleic acid mixed ester absorbent from S2, and it is sent into the fixed bed reactor filled with 0.5 wt% Pd / SiO2 catalyst that has been pre-reduced and activated, and operates in parallel or countercurrent with hydrogen gas to carry out the two-esterification and double bond hydrogenation of the maleic acid ester absorbent. Figure 5 Shown in FIG., and operate in parallel or countercurrent with hydrogen gas to carry out the two-esterification and double bond hydrogenation of the maleic acid ester absorbent.

[0232] In this embodiment, the diallylation of maleic acid and the hydrogenation of double bonds are as follows: Figure 5 The tubular isothermal fixed-bed reactor shown and Figure 6 the co-current isothermal trickle-bed fixed-bed reactor with the structure shown in Fig. a. The temperature of the maleic acid allyl ester absorption liquid is 78 °C, and the weight hourly space velocity is 0.5 h -1 . The hydrogen pressure is 1.0 MPa, and the molar ratio of hydrogen to ester is 10. The results are as follows: the diallylation rate of maleic acid allyl ester is 99.8%, the double bond hydrogenation rate is 100%, and the selectivity of diallyl succinate is 99.3%.

[0233] S4. Gas-liquid separation of the diallylation-hydrogenation product:

[0234] The diallylation-hydrogenation coupling reaction material from S3 is sent into the gas-liquid separator. The separated hydrogen is pressurized by a compressor and then sent back to the diallylation-hydrogenation reactor for recycling. The liquid material is sent into the diallyl succinate refining system after being recovered by the absorbent.

[0235] Example 14

[0236] S1. Absorption-esterification of maleic anhydride gas stream:

[0237] The maleic anhydride gas stream and the absorbent are respectively sent into the Figure 2 solvent absorption-esterification coupling reaction tower filled with solid acid catalyst with a diameter of 2.8 mm shown in Fig. a-2d. The maleic anhydride gas stream and the absorbent flow through the catalyst bed of the 1-3 stage series absorption-esterification tower shown in Figure 3 Fig. in parallel or countercurrent, and the absorption-esterification coupling reaction of maleic anhydride occurs.

[0238] In this embodiment, the preparation of the maleic acid mono / dihydrocinnamyl ester absorption liquid is as follows: Figure 2 The co-current bubble absorption esterification tower filled with (WO3 / ZrO2) / HMCM-41 pellets shown in Fig. d and the co-tower feeding 2-stage series absorption method shown in Figure 3 Fig. d (only using the first and second stage absorption-esterification towers of Figure 3 Fig. d) are adopted. The maleic anhydride gas stream and the absorbent hydrocinnamyl alcohol are respectively sent into the absorption-esterification tower. The temperature of the maleic anhydride gas stream is 120 °C, and the volume hourly space velocity is 1500 h -1 , the temperature of hydrocinnamyl alcohol is 65 °C, and the weight hourly space velocity is 1.5 h -1 ; the obtained absorption liquid has a temperature of ~95 °C and a maleic acid hydrocinnamyl ester concentration of ~43 wt%. The esterification rate of maleic anhydride is 100%, and the selectivity of maleic acid hydrocinnamyl ester is 100%.

[0239] S2. Removal of dissolved oxygen from the maleic acid ester absorption liquid:

[0240] The absorption liquid from S1 is sent into the dissolved oxygen removal tower, and after the dissolved oxygen is removed under negative pressure, it is sent into the hydrogenation reactor. The operating conditions are: the temperature of the absorption liquid, and the pressure at the top of the dissolved oxygen removal tower is 90 kPa.

[0241] S3. Diesterification - double bond hydrogenation of maleate absorbent:

[0242] Add 2.5 wt% p - toluenesulfonic acid catalyst to the maleic acid mixed ester absorbent from S2, and feed it into the fixed - bed reactor filled with 0.5 wt% Pd / SiO2 catalyst that has been pre - reduced and activated as shown in Figure 5 Perform diesterification - double bond hydrogenation of the maleate absorbent by co - current or counter - current operation with hydrogen.

[0243] In this example, for the diesterification - double bond hydrogenation of hydrocinnamyl maleate: Use the tubular isothermal fixed - bed reactor shown in Figure 5 and the co - current isothermal trickle - bed fixed - bed reactor with the structure shown in Figure 6 a. The temperature of the hydrocinnamyl maleate absorbent is 95 °C, the weight hourly space velocity is 1.0 h -1 , the hydrogen pressure is 1.0 MPa, and the hydrogen - ester molar ratio is 20. The results are as follows: the diesterification rate of hydrocinnamyl maleate is 98.5%, the conversion rate of double - bond hydrogenation is 100%, and the selectivity of dihydrocinnamyl succinate is 98.8%.

[0244] S4. Gas - liquid separation of the diesterification - hydrogenation product:

[0245] Feed the reaction material of the diesterification - hydrogenation coupling from S3 into a gas - liquid separator. The separated hydrogen is pressurized by a compressor and then fed into the diesterification - hydrogenation reactor for recycling. The liquid material is fed into the succinic acid diester refining system after being recovered by the absorbent.

[0246] Example 15

[0247] S1. Maleic anhydride gas absorption - esterification:

[0248] Feed the maleic anhydride gas stream and the absorbent into the solvent absorption - esterification coupling reaction tower filled with a solid acid catalyst with a diameter of 2.8 mm as shown in Figure 2 a - 2d. The maleic anhydride gas stream and the absorbent pass through the catalyst bed of the 1 - 3 - stage series - connected absorption - esterification tower shown in Figure 3 either co - currently or counter - currently to carry out the absorption - esterification coupling reaction of maleic anhydride.

[0249] Preparation of maleic acid mono / bis(4 - hydroxymethylcyclohexyl)methyl ester absorbent: Use the counter - current spray - bubbling absorption - esterification tower filled with HZSM - 5 pellets shown in Figure 2 a and the 3 - stage absorption method with different tower feeds shown in Figure 3 a. The temperature of the maleic anhydride gas stream is 120 °C and the volume hourly space velocity is 3000 h -1 , the temperature of the 1,4 - cyclohexanedimethanol fed into the first stage is 60 °C and the weight hourly space velocity is 3.0 h -1;The obtained absorbent has a temperature of ~95 °C and a concentration of 4-hydroxymethylcyclohexyl methyl maleate of ~45 wt%, the maleic anhydride esterification rate is 100%, and the selectivity of 4-hydroxymethylcyclohexyl methyl maleate is 100%.

[0250] S2. Dissolved oxygen removal from the maleate absorbent:

[0251] Feed the absorbent from S1 into a dissolved oxygen removal tower, remove the dissolved oxygen under negative pressure, and then feed it into a hydrogenation reactor. Operating conditions: the temperature of the absorbent in each example, and the pressure at the top of the dissolved oxygen removal tower is 90 kPa.

[0252] S3. Diesterification - double bond hydrogenation of the maleate absorbent:

[0253] Add 2.5 wt% p-toluenesulfonic acid catalyst to the maleic acid mixed ester absorbent from S2, and feed it into the fixed bed reactor filled with 0.5 wt% Pd / SiO2 catalyst that has been pre-reduced and activated as shown in Figure 4 Perform the diesterification - double bond hydrogenation of the maleate absorbent by co-current or counter-current operation with hydrogen.

[0254] In this example, for the diesterification - double bond hydrogenation of 4-hydroxymethylcyclohexyl methyl maleate: Use the Figure 4 interstage heat exchange adiabatic fixed bed reactor shown in and the Figure 6 co-current adiabatic trickle fixed bed reactor with the structure shown in b. The temperature of the 4-hydroxymethylcyclohexyl methyl maleate absorbent is 95 °C and the weight hourly space velocity is 1.0 h -1 、hydrogen pressure is 1.0 MPa and the hydrogen - ester molar ratio is 10. The results are as follows: the diesterification rate of 4-hydroxymethylcyclohexyl methyl maleate is 98.2%, the conversion rate of double bond hydrogenation is 100%, and the selectivity of di(4-hydroxymethylcyclohexyl) methyl succinate is 99.2%.

[0255] S4. Gas - liquid separation of the diesterification - hydrogenation product:

[0256] Feed the diesterification - hydrogenation coupling reaction material from S3 into a gas - liquid separator. The separated hydrogen is pressurized by a compressor and then fed into the diesterification - hydrogenation reactor for recycling. The liquid material is fed into the di - ester of succinic acid refining system after being recovered by the absorbent.

[0257] Example 16

[0258] S1. Maleic anhydride gas absorption - esterification:

[0259] Feed the maleic anhydride gas stream and the absorbent into the solvent absorption - esterification coupling reaction tower filled with a solid acid catalyst with a diameter of 2.8 mm as shown in Figure 2 a - 2d respectively. The maleic anhydride gas stream and the absorbent pass through the Figure 3In the catalyst beds of the 1st - 3rd stage series - connected absorption - esterification towers shown, the absorption - esterification coupling reaction of maleic anhydride occurs.

[0260] In this example, the preparation of the maleic acid mono / di(triethylene glycol) ester absorbent: Use Figure 2 The co - current bubble - column absorption - esterification tower filled with H3PW 12 O 40 / SBA - 15 pellets and the Figure 3 The 2 - stage series - connected absorption method with different - tower feeding shown in Figure 3 (only use the first and second - stage absorption - esterification towers in ( -1 Feed the maleic anhydride gas stream and the absorbent triethylene glycol into the counter - current spray bubble - column absorption - esterification tower respectively. The temperature of the maleic anhydride gas stream is 95°C and the volumetric space velocity is 2500 h -1 ; The temperature of the triethylene glycol fed into the first stage is 60°C and the weight space velocity is 2.5 h

[0261] S2. Removal of dissolved oxygen from the maleic acid ester absorbent:

[0262] Feed the absorbent from S1 into the de - dissolved - oxygen tower, and after removing the dissolved oxygen under negative pressure, send it into the hydrogenation reactor. Operating conditions: the temperature of the absorbent, and the pressure at the top of the de - dissolved - oxygen tower is 70 kPa.

[0263] S3. Diesterification - double - bond hydrogenation of the maleic acid ester absorbent:

[0264] Add 2.5 wt% p - toluenesulfonic acid catalyst to the maleic acid mixed - ester absorbent from S2, and send it into the Figure 4 Fixed - bed reactor filled with 0.5 wt% Pd / SiO2 catalyst that has been pre - reduced and activated, and operate in co - current or counter - current with hydrogen to carry out the diesterification - double - bond hydrogenation of the maleic acid ester absorbent.

[0265] In this example, for the diesterification - double - bond hydrogenation of triethylene glycol maleate: Use Figure 4 The inter - stage heat - exchange adiabatic fixed - bed reactor shown and Figure 6 The co - current adiabatic trickle - bed fixed - bed reactor shown in ( -1 The temperature of the triethylene glycol maleate absorbent is 80°C and the weight hourly space velocity is 0.50 h

[0266] S4. Gas - liquid separation of the diesterification - hydrogenation product:

[0267] Feed the reaction materials of the two-esterification-hydrogenation coupling from S3 into a gas-liquid separator. The separated hydrogen is pressurized by a compressor and then sent into the two-esterification-hydrogenation reactor for recycling. The liquid materials are sent into the succinic acid diester refining system after being recovered by the absorbent.

[0268] Comparative Example 1

[0269] Maleic anhydride gas absorption-desorption-refining to produce maleic anhydride and maleic anhydride esterification-hydrogenation to produce monobutyl succinate includes the following steps:

[0270] (1) Solvent absorption of maleic anhydride gas: Feed the oxidized reaction gas generated by the maleic anhydride unit after cooling into the countercurrent spray bubble absorption tower shown in Attachment Figure 2 a and the 3-stage series absorption method with different tower feeds shown in Attachment Figure 3 a. The temperature of the maleic anhydride gas stream is 120 °C and the volumetric space velocity is 2000 h -1 . The temperature of the absorption solvent dibutyl phthalate (DBP) for the first-stage feed is 40 °C and the weight space velocity is 1.5 h -1 . The obtained absorption liquid has a temperature of ~75 °C and a maleic anhydride concentration of ~25 wt%.

[0271] (2) Solvent desorption of the absorption liquid: Feed the absorption liquid from the solvent absorption tower into the solvent desorption tower, and separate the maleic anhydride containing light components from the DBP solvent containing heavy components through vacuum flashing. The crude maleic anhydride is sent into the product refining system, and the DBP solvent containing heavy components is sent into the solvent purification system.

[0272] (3) Purification of the desorbed solvent DBP: The solvent containing heavy components sent into the solvent purification system first enters the demineralized water washing tower. After washing and extracting the acidic components, it is then sent into a centrifuge to separate the aqueous phase and the solvent phase. The solvent DBP is sent back to the solvent absorption tower for recycling, and the aqueous phase is sent into the wastewater treatment system.

[0273] (4) Separation and refining of the desorbed liquid: The crude maleic anhydride sent into the product refining system is subjected to vacuum distillation in the light component removal tower. By-products such as acetic acid and acrylic acid are removed from the top of the tower, and the bottom material of the tower is sent into the product tower; through vacuum distillation, maleic anhydride products are obtained from the top of the product tower. Operating conditions: The pressure for light component removal is 10 kPa, the top temperature is 130 °C, and the bottom temperature is 170 °C; the pressure of the product tower is 5 kPa, the top temperature is 130 °C, and the bottom temperature is 140 °C.

[0274] (5) Esterification of maleic anhydride to prepare monobutyl maleate: Mix the maleic anhydride from (4) with the esterifying agent butanol, and feed it into a reaction kettle filled with 3.2 mm corundum ball packing to carry out the mono-esterification reaction of maleic anhydride to produce monobutyl maleate. Control the molar ratio of alcohol to anhydride to be 3.1:1.0, the reaction temperature to be 60 °C, and the feed weight space velocity to be 1.0 h -1 . The esterification rate of maleic anhydride is 100%, and the selectivity of monobutyl maleate is 100%.

[0275] (6) Hydrogenation of monobutyl maleate: The same as S3 in Example 1, except that the monoesters absorption liquid from S2 is changed to the monoesters esterification liquid from (5). The double bond hydrogenation rate of monobutyl maleate is 100%, and the selectivity of monobutyl succinate is 100%.

[0276] (7) Gas-liquid separation of hydrogenation product: The same as S4 in Example 1, except that the hydrogenation reaction material from S3 is changed to the hydrogenation reaction material from (6).

[0277] Comparative Example 2

[0278] Maleic anhydride gas absorption - desorption - purification to produce maleic anhydride and maleic anhydride hydrogenation - esterification to produce monobutyl succinate

[0279] Steps (1) to (4) for solvent absorption, solvent desorption, solvent purification of maleic anhydride gas stream and purification of desorbed liquid to prepare product maleic anhydride are the same as in Comparative Example 1.

[0280] (5) Hydrogenation of product maleic anhydride to produce succinic anhydride: The product maleic anhydride from (4) is mixed with the solvent γ - butyrolactone to prepare a maleic anhydride solution with a concentration of 48 wt%, and is fed into an inter - stage heat - exchange and co - current adiabatic trickle - bed fixed - bed reactor filled with 0.5 wt% Pd / SBA - 15 catalyst that has been pre - reduced and activated, for hydrogenation of maleic anhydride to produce succinic anhydride. Control the reaction temperature at 60 °C, the feed weight hourly space velocity at 0.5 h -1 , the hydrogen pressure at 2.0 MPa and the hydrogen - anhydride molar ratio at 10. The double bond hydrogenation rate of maleic anhydride is 100%, and the selectivity of succinic anhydride is 99.6%.

[0281] (6) Gas - liquid separation of maleic anhydride hydrogenation product: The hydrogenation reaction material from (5) is fed into a gas - liquid separator. The separated hydrogen is pressurized by a compressor and sent back to the (5) hydrogenation reactor for recycling, and the liquid material is fed into (7).

[0282] (7) Recovery of maleic anhydride hydrogenation solvent: The liquid material from (6) is fed into a γ - butyrolactone recovery tower for solvent. After vacuum distillation, the γ - butyrolactone distilled from the top of the tower is condensed and sent to a storage tank for recycling as the hydrogenation solvent in (5), and the bottom material succinic anhydride is fed into (8).

[0283] (8) Esterification of succinic anhydride to produce monobutyl succinate: The succinic anhydride from (7) is mixed with the esterifying agent butanol and fed into a reaction kettle filled with corundum balls with a diameter of 3.2 mm as packing for the mono - esterification reaction of succinic anhydride to produce monobutyl succinate, and then the esterification reaction liquid is fed into the monobutyl succinate refining system. Control the alcohol - anhydride molar ratio of the esterification reaction at 3.1:1.0, the reaction temperature at 60 °C, and the feed weight hourly space velocity at 1.0 h -1 , the esterification rate of succinic anhydride is 100%, and the selectivity of monobutyl succinate is 100%.

[0284] Comparative Example 3

[0285] Preparation of maleic anhydride by solvent absorption - desorption - purification of maleic anhydride in gas stream and preparation of dibutyl succinate by esterification - hydrogenation of maleic anhydride

[0286] Steps (1) - (4): Solvent absorption, solvent desorption, solvent purification of maleic anhydride in gas stream and purification of desorbed liquid to prepare the product maleic anhydride are the same as in Comparative Example 1.

[0287] (5) Esterification of the product maleic anhydride with butanol to prepare dibutyl maleate: Mix the product maleic anhydride from (4) with butanol containing 2.0 wt% trifluoromethanesulfonic acid, and feed it into an esterification reactor filled with quartz ball packing with a diameter of 2.6 mm to carry out the two - esterification reaction of maleic anhydride and butanol. Control the molar ratio of alcohol to anhydride at 3.3:1.0, the reaction temperature at 80 °C, and the weight hourly space velocity of butanol at 1.0 h -1 , the esterification rate of maleic anhydride is 100%, and the selectivity of dibutyl maleate is 98.0%.

[0288] (6) Hydrogenation of dibutyl maleate to prepare dibutyl succinate: Feed the esterification liquid from (5) into a co - current isothermal tubular trickle - bed fixed - bed reactor filled with 0.5 wt% Pd / HM catalyst that has been pre - reduced and activated to carry out the double - bond hydrogenation of dibutyl maleate to prepare dibutyl succinate. Control the reaction temperature at 80 °C, the weight hourly space velocity of the esterification liquid at 0.5 h -1 , the hydrogen pressure at 1.5 MPa, and the molar ratio of hydrogen to ester at 15. The double - bond hydrogenation rate of dibutyl maleate is 100%, and the selectivity of dibutyl succinate is 100%.

[0289] (7) Gas - liquid separation of the hydrogenation product: Feed the hydrogenation material from (6) into a gas - liquid separator. The separated hydrogen is pressurized by a compressor and then sent back to (6) for recycling, and the liquid material is sent to the dibutyl succinate purification system.

[0290] Comparative Example 4

[0291] Preparation of maleic anhydride by solvent absorption - desorption - purification of maleic anhydride in gas stream and preparation of dibutyl succinate by hydrogenation - esterification of maleic anhydride

[0292] Steps (1) - (7): Solvent absorption, solvent desorption, solvent purification of maleic anhydride in gas stream and purification of desorbed liquid to prepare the product maleic anhydride, as well as solvent hydrogenation of maleic anhydride to prepare succinic anhydride, gas - liquid separation of the hydrogenation product, and recovery of the hydrogenation solvent are all the same as in Comparative Example 2.

[0293] (8) Esterification of succinic anhydride to prepare dibutyl succinate: Mix the liquid phase material from (7) with butanol containing 2.0 wt% trifluoromethanesulfonic acid, feed it into an esterification reactor filled with quartz ball packing with a diameter of 2.6 mm to carry out the two - esterification reaction of succinic anhydride and butanol, and then send the esterification reaction liquid to the dibutyl succinate purification system. Control the molar ratio of alcohol to anhydride in the esterification reaction at 3.3:1.0, the reaction temperature at 80 °C, and the weight hourly space velocity of butanol at 1.0 h -1, the esterification rate of succinic anhydride is 100%, and the selectivity of dibutyl succinate is 98.6%.

[0294] It can be seen from Examples 1 to 16 that starting from the oxidation reaction product gas (gaseous crude maleic anhydride) of an industrial maleic anhydride plant and using the reaction system and key equipment reactor of the present invention, whether it is to prepare a monoalkyl or monohydroxyalkyl succinate by a series connection of an alcohol absorption - monoesterification coupling reaction and a hydrogenation reaction, or to prepare a dialkyl or dihydroxyalkyl succinate by a series connection of an alcohol absorption - monoesterification coupling reaction and a diesterification - hydrogenation coupling reaction, an alcohol absorption - diesterification coupling reaction and a hydrogenation reaction, or an alcohol absorption - mono / di mixed esterification coupling reaction and a diesterification - hydrogenation coupling reaction, good reaction effects can be achieved. The maleic anhydride esterification rate is 100%, the double bond hydrogenation rate is 100%, the selectivity of the product monoester of succinic acid is 98.9 - 100%, or the selectivity of the product diester of succinic acid is 98.5 - 100%.

[0295] Comparing Example 1 and 5 with Comparative Examples 1 to 4, it can be seen that when preparing a monoester or diester of succinic acid starting from the oxidation reaction product gas of an industrial maleic anhydride plant, using the existing technical routes and systems, that is: first, absorb gaseous crude maleic anhydride through a solvent (such as DBP), solvent desorption, solvent purification, and desorption liquid refining processes and systems to obtain the product maleic anhydride (refined maleic anhydride), then carry out monoesterification or diesterification of the product maleic anhydride with an esterifying agent alcohol to prepare maleic mono- or diester, and then hydrogenate the maleic ester to prepare a mono- or diester of succinic acid, or hydrogenate the product maleic anhydride to prepare succinic anhydride, and then carry out monoesterification or diesterification of succinic anhydride with an esterifying agent alcohol to prepare a mono- or diester of succinic acid; the entire system from gaseous crude maleic anhydride through solvent absorption to the reaction to produce succinic acid ester requires 7 - 8 steps, and the process is very cumbersome. In the present invention, for the entire reaction system from gaseous crude maleic anhydride to the reaction to produce succinic acid ester, due to the adoption of key equipment such as absorption - esterification and esterification - hydrogenation coupling reactors, as well as coupling reaction and series reaction technologies, the process and system are greatly simplified, and it only takes 4 steps to synthesize succinic acid ester, eliminating the processes such as solvent desorption, purification, and desorption liquid refining for preparing refined maleic anhydride in the existing system.

[0296] Obviously, the system of the present invention is very simple and straightforward. The key is the adoption of coupling reaction technology and reactors; the present invention not only simplifies the system and process for synthesizing succinic acid ester, greatly reduces equipment investment, makes the operation more convenient, but also significantly reduces energy consumption and material consumption (reducing the loss of maleic anhydride in the existing system), thereby greatly reducing the total production cost; in addition, the present invention can also produce a series of different succinic acid ester products through the same set of reaction system and equipment by adjusting process parameters, with a wide range of technical applications, flexible product schemes, and the ability to adjust product varieties according to market demand, thereby enhancing the competitiveness of products and production enterprises.

[0297] The above are only the preferred embodiments of the present invention and are not intended to limit the present invention. Any modifications, equivalent replacements, improvements, etc. made within the spirit and principle of the present invention shall be included within the protection scope of the present invention.

Claims

1. A process system for producing succinates from maleic anhydride gas flow, characterized in that: It includes an alcohol solvent absorption esterification system and an absorption liquid hydrogenation or diesterification hydrogenation coupling system which are sequentially connected through pipelines, wherein: The alcohol solvent absorption esterification system comprises 1 to 3-stage absorption esterification towers for absorbing maleic anhydride gas flow and an absorption liquid deoxygenation tower for removing dissolved oxygen. The absorption esterification tower is provided with an alcohol absorbent inlet, a maleic anhydride gas flow inlet and an absorption liquid outlet, and the absorption liquid outlet is connected to the deoxygenation tower. The absorption liquid hydrogenation or diesterification hydrogenation coupling system comprises an absorption liquid hydrogenation reactor and a gas-liquid separator connected in sequence, the outlet of the deoxygenation tower is connected to the absorption liquid hydrogenation reactor, and the gas phase outlet pipeline of the gas-liquid separator is connected to the hydrogen inlet of the absorption liquid hydrogenation reactor; the absorption esterification tower is a bubbling tower, and the absorption liquid hydrogenation reactor is a fixed bed reactor.

2. The process system for producing succinate esters from maleic anhydride gas flow according to claim 1, characterized in that: The absorption esterification tower is selected from a countercurrent spray bubbling absorption esterification tower or a parallel current bubbling absorption esterification tower, wherein: The countercurrent spray bubbling absorption esterification tower comprises, from top to bottom, an upper end cap of the absorption tower, a spray absorption zone, a bubbling absorption esterification zone and a lower end cap of the absorption tower, wherein the upper end cap is provided with an alcohol absorbent inlet and an exhaust gas outlet, and the lower end cap is provided with a maleic anhydride gas flow inlet and an absorption liquid outlet; a sprayer is provided above the spray absorption zone, and the sprayer evenly sprays the absorbent and the maleic anhydride gas flow; the bubbling absorption esterification zone is filled with granular fillers and / or solid acid catalysts, and the sprayed alcohol absorbent passes through the bed from top to bottom, and the maleic anhydride gas flow countercurrently passes through the bed from bottom to top; The parallel-current bubbling absorption esterification tower comprises, from top to bottom, an upper end cap of the absorption tower, a bubbling absorption esterification zone and a lower end cap of the absorption tower, wherein the upper end cap is provided with an absorption liquid outlet and an exhaust gas outlet, and the lower end cap is provided with an alcohol absorbent inlet and a maleic anhydride airflow inlet; the bubbling absorption esterification zone is filled with granular fillers and / or solid acid catalysts, and the alcohol absorbent and maleic anhydride airflow entering the parallel-current bubbling absorption esterification tower both flow through the bed from bottom to top in parallel.

3. The process system for producing succinate esters from maleic anhydride gas flow according to claim 1, characterized in that: The 1 to 3 stage absorption esterification towers include a 1 stage absorption tower with only one absorption tower, or a 2 stage absorption tower with two absorption towers connected in series, or a 3 stage absorption tower with three absorption towers connected in series. The maleic anhydride gas flow and the alcohol absorbent are fed into the same tower or different towers.

4. The process system for producing succinate esters from maleic anhydride gas flow according to claim 1, characterized in that: The absorption liquid hydrogenation reactor is an adiabatic fixed bed reactor or an isothermal fixed bed reactor; When an adiabatic fixed bed reactor is selected, the adiabatic fixed bed hydrogenation reactor comprises, from top to bottom, an upper end cap of the reactor, a first stage reactor, a second stage reactor and a lower end cap of the reactor, wherein the upper end cap of the reactor is provided with a liquid phase hydrogenation raw material inlet and a hydrogen inlet, and a distributor is provided in the upper end cap of the reactor, and the top of the bed of the first stage reactor and the second stage reactor are filled with inert porcelain balls and the bed is filled with a hydrogenation catalyst or a diesterification-hydrogenation bifunctional catalyst; a heat exchanger is provided between the first stage reactor and the second stage reactor, and the reaction material of the first stage reactor enters the second stage reactor for further reaction after heat exchange and cooling in the heat exchanger, and a reaction product outlet is provided on the lower end cap; the liquid phase hydrogenation material is mixed with hydrogen in the distributor, and enters the first stage reactor and the second stage reactor in sequence through the distributor; The isothermal fixed bed reactor comprises, from top to bottom, an upper reactor head, an isothermal tube fixed bed reactor and a lower reactor head, wherein the upper reactor head is provided with a liquid phase hydrogenation feedstock inlet and a hydrogen inlet, a distributor is provided in the upper reactor head, and the lower reactor head is provided with a reaction product outlet; the top and bottom ends of the tubes of the isothermal tube fixed bed reactor are both filled with inert porcelain balls, the tube bed is filled with a hydrogenation catalyst or a diesterification-hydrogenation dual-function catalyst, and heat medium is filled between the tubes, and the heat medium enters from the bottom and exits from the top in countercurrent heat exchange with the reaction materials that enter from the top and exit from the bottom.

5. The process system for producing succinate esters from maleic anhydride gas flow according to claim 1, characterized in that: The process system also includes an absorbent recovery tower, a succinate refining tower and a tail gas treatment system. The gas phase outlet and the liquid phase outlet of the gas-liquid separator are respectively connected to the hydrogen inlet of the absorption liquid hydrogenation reactor and the absorbent recovery tower, the product outlet of the absorbent recovery tower is connected to the succinate refining tower, and the tail gas outlet of the absorption esterification tower and the dissolved oxygen outlet of the de-dissolved oxygen are both connected to the tail gas treatment system.

6. An application of the process system according to any one of claims 1 to 5, characterized in that: The method for producing succinate esters using maleic anhydride gas stream comprises the following steps: S1. Absorption-esterification of maleic anhydride gas stream: The maleic anhydride gas stream and the alcohol absorbent containing or not containing a liquid acid catalyst are respectively fed into the absorption-esterification tower to carry out solvent absorption and esterification reaction of maleic anhydride to prepare maleic acid ester absorption liquid: the maleic acid ester includes maleic acid monoester and / or maleic acid diester; S2. Removing dissolved oxygen from the absorption liquid: sending the absorption liquid of the maleic anhydride gas flow in the absorption esterification tower to a dissolved oxygen removal tower, removing dissolved oxygen under negative pressure, and then sending it to the absorption liquid hydrogenation or diesterification hydrogenation coupling system; operating conditions: absorption liquid temperature 52-150°C and tower top pressure 50.0-90.0 kPa; preferably, the absorption liquid temperature 55-120°C and the tower top pressure 60.0-85.0 kPa; more preferably, the absorption liquid temperature 60-90°C and the tower top pressure 70.0-80.0 kPa; S3. hydrogenation or diesterification-hydrogenation of the absorption liquid: the absorption liquid and hydrogen after the removal of dissolved oxygen are fed into an absorption liquid hydrogenation reactor filled with a pre-reduced activated supported metal catalyst to hydrogenate the double bonds of maleate in the absorption liquid or diesterification-hydrogenation of maleate monoester to prepare succinate, wherein the succinate includes succinate monoester and / or succinate diester; S4. Gas-liquid separation of reaction products: The material after hydrogenation of the absorption liquid or diesterification-hydrogenation reaction is sent to a gas-liquid separator for gas-liquid separation, and the liquid material is subjected to subsequent succinate ester refining treatment.

7. The use according to claim 6, characterized in that: In step S1, the operating conditions are: the concentration of the liquid acid catalyst in the alcohol absorbent is 0-5.0 wt%, the temperature of the maleic anhydride gas flow is 55-180°C and the volume space velocity is 200-20000 h -1 , alcohol absorbent temperature 25~85℃ and weight space velocity 0.1~10.0h -1 , after being absorbed by one or more absorption and esterification towers connected in series, a maleate absorption liquid is obtained; Preferably, the operating conditions are: the concentration of liquid acid catalyst in the alcohol absorbent is 0-3.0wt%, the temperature of maleic anhydride gas flow is 60-150°C and the volume space velocity is 500-10000h -1 , alcohol absorbent temperature 30~80℃ and weight space velocity 0.2~5.0h -1 , through 1 to 3 absorption towers in series, maleate absorption liquid is obtained; More preferably, the operating conditions are: the concentration of the liquid acid catalyst in the alcohol absorbent is 0 to 2.5 wt%, the temperature of the maleic anhydride gas flow is 65 to 120°C and the volume space velocity is 1000 to 5000 h -1 , alcohol absorbent temperature 35 ~ 75 ℃ and weight space velocity 0.5 ~ 3.0h -1 , and the maleate absorption liquid is obtained through absorption in series in 1 to 3 absorption towers.

8. The use according to claim 6, characterized in that: In step S3, a fixed bed reactor and a non-acidic carrier-supported metal particle catalyst or a solid acid-supported metal particle bifunctional catalyst are used to feed the absorption liquid into an absorption liquid hydrogenation reactor, and maleate hydrogenation or maleate monoester diesterification-hydrogenation is performed in parallel or countercurrent operation with hydrogen to prepare succinate; The operating conditions are: absorption liquid temperature 50-150°C and weight hourly space velocity 0.1-3.0h -1 , hydrogen pressure 0.5-5.0 MPa and hydrogen ester molar ratio 5-50; preferably, the absorption liquid temperature 55-120°C and the weight hourly space velocity 0.25-2.5h -1 , hydrogen pressure 0.75-3.0 MPa and hydrogen ester molar ratio 7.5-40; more preferably, the absorption liquid temperature 60-90 ° C and the weight hourly space velocity 0.5-2.0h -1 , hydrogen pressure 1.0-2.0 MPa and hydrogen-ester molar ratio 10-30.

9. The use according to claim 6, characterized in that: The alcohol absorbent is selected from at least one of aliphatic alcohols, aromatic alcohols, heterocyclic alcohols, polyether polyols or polyester polyols; preferably, selected from n-butanol, isobutanol, n-pentanol, isopentanol, neopentyl alcohol, n-hexanol, n-heptanol, n-octanol, isooctyl alcohol, 2-ethylhexanol, isononanol, lauryl alcohol, stearyl alcohol, cyclohexanol, methylcyclohexanol, cyclohexylpropanol, allyl alcohol, butenol, methylallyl alcohol, isopentenol, cinnamyl alcohol, oleyl alcohol, propynyl alcohol, butynol, benzyl alcohol, phenylethyl alcohol, hydrogenated cinnamyl alcohol, tetrahydrofurfuryl alcohol, furfuryl alcohol, ethylene glycol, polyethylene glycol, 1,2-propylene glycol, polypropylene glycol, 1,3-propylene glycol, 1 ,4-butanediol, 1,5-pentanediol, 1,6-hexanediol, 1,7-heptanediol, 1,8-octanediol, 1,9-nonanediol, 1,10-decanediol, 2-methyl-1,3-propanediol, neopentyl glycol, 2,2-diethyl-1,3-propanediol, cyclohexanediol, cyclohexanedimethanol, phenyl-1,3-propanediol, benzenedimethanol, 2,5-furandimethanol, butene diol, butynediol, diethylaminoethanol, diethylaminopropanol, or at least one of 2,2,2-trifluoroethanol, pentafluoroethanol, heptafluoropropanol, 2,2,2-trichloroethanol, chloropropanol or 1,3-dichloropropanol.

10. The use according to claim 6, characterized in that: The hydrogenation or esterification-hydrogenation catalyst comprises a non-acidic carrier or a solid acid carrier-supported metal bifunctional catalyst, wherein: the non-acidic carrier-supported metal particle catalyst is selected from activated carbon (AC), mesoporous carbon, carbon nanotubes, graphene, SiO2, mesoporous SiO2, TiO2, ZrO2, SBA-15, ZEO-1, ZEO-3 or TS-1-supported precious metal catalyst; preferably, SiO2, TiO2, ZrO2, SBA-15 or TS-1-supported Ru, Pd or Pt single metal or bimetallic catalyst; more preferably, Pd / SiO2, Pt / SiO2, Pd-Ru / SiO2 or Pd / SBA-15 catalyst; Solid acid carrier supported metal bifunctional catalysts are selected from HM, Hβ, HZSM-5, HZSM-22, HMCM-22, HMCM-49, or HMCM-41, HMCM-48 or SBA-15 supported WO3 / ZrO2, H3PW 12 O 40 or Cs 2.5 H 0.5 PW 12 O 40 The solid acid is a carrier-supported precious metal catalyst; preferably HM, HZSM-5, HMCM-22, H3PW 12 O 40 / HMCM-48, Cs 2.5 H 0.5 PW 12 O 40 / SBA-15 or (WO3 / ZrO2) / HMCM-41 is a single or bimetallic catalyst with Ru, Pd or Pt as the carrier; more preferably Pd / HM, Pt / HZSM-5, Pd-Ru / HMCM-22, Pd / (WO3 / ZrO2) / HMCM-41 or Pd / Cs 2.5 H 0.5 PW 12 O 40 / SBA-15 catalyst.

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