Method and process system for producing poly (butylene succinate) by using gaseous crude maleic anhydride

Abstract

The invention provides a method and a process system for producing poly (butylene succinate) by using gaseous crude maleic anhydride, and the method comprises the following steps: esterifying 1, 4-butanediol to absorb maleic anhydride gas flow generated by oxidizing benzene or butane to prepare butylene succinate, hydrogenating the butylene succinate to prepare butylene succinate, hydrogenating the butylene succinate to prepare 1, 4-butanediol, and hydrogenating the 1, 4-butanediol to prepare 1, 4-butanediol. The invention relates to a catalyst for preparing poly (butylene succinate) through polycondensation of 1, 4-butanediol and butylene succinate and a process flow. The two monomers of the full-biodegradable polyester PBS and the polyester plastic can be prepared only through four steps from the gaseous crude maleic anhydride generated by oxidizing benzene or butane, the technological process is simple and convenient, the utilization rate of raw materials is high, the process is clean and environmentally friendly, the raw materials are low in price and easy to obtain, material consumption is low, energy consumption is low, investment is saved, and production cost is low.

Description

Technical Field

[0001] The present invention belongs to the technical field of polyester production, and particularly relates to a method and a process system for producing polybutylene succinate with gaseous crude maleic anhydride. Background Art

[0002] At present, the synthesis methods of the fully biodegradable plastic polybutylene succinate (PBS) mainly include ring-opening esterification polycondensation method, direct esterification polycondensation method or transesterification polycondensation method of acid monomers such as succinic anhydride, succinic acid or succinic acid ester and alcohol monomer 1,4-butanediol (BDO). Due to the high raw material prices and consumption, long process flow, high investment and energy consumption for manufacturing these monomers, the production cost of PBS remains high, lacking market competitiveness with traditional plastics such as PE and PP. Furthermore, the market acceptance of PBS plastics with excellent degradation performance is poor, increasing the difficulty of popularization and application. Therefore, reducing the manufacturing cost of PBS is the key to improving its market acceptance and achieving large-scale popularization and application.

[0003] Currently, the mainstream production routes and processes of PBS start with maleic anhydride. Through chemical synthesis, acid monomers such as succinic anhydride, succinic acid, or diethyl succinate and alcohol monomer BDO are first prepared, and then these two monomers are synthesized through ring-opening esterification, direct esterification, or transesterification polycondensation. There are many patents on the preparation of PBS acid monomers starting from maleic anhydride raw materials, covering various aspects such as preparation methods, catalysts, reaction processes, and technical routes. For example, CN105801536B discloses a method for the liquid-phase selective hydrogenation of maleic anhydride to prepare succinic anhydride, using the same catalyst of 0.01 - 1.0% Pd and a two-stage low-temperature and low-pressure reaction process, with a maleic anhydride conversion rate ≥ 99.9% and a succinic anhydride selectivity of 99.0 - 99.5%; CN114011421B discloses a method for the hydrogenation of maleic anhydride to succinic anhydride, using a basic salt-modified supported nickel or cobalt catalyst. Under the preferred reaction pressure of 1 - 2 MPa and temperature of 60 - 100 °C, the maleic anhydride conversion rate is greater than 99.9%, and the succinic anhydride selectivity exceeds 99.5%; CN107253938B discloses a production process for the direct hydrogenation of maleic anhydride to prepare high-purity succinic anhydride. A fixed-bed reactor is used, and hydrogen is added through four-stage cold hydrogen to completely react maleic anhydride. At the same time, the hydrogenation product is subjected to low-vacuum distillation in three towers (light component removal tower, γ-butyrolactone removal tower, and heavy component removal tower), and the purity of succinic anhydride reaches more than 99.9%; CN102311332B discloses a method for the hydrogenation and hydrolysis of maleic anhydride to produce succinic acid. Using a γ-butyrolactone solution of maleic anhydride with a concentration of 5 - 90 wt% as the raw material, succinic anhydride is produced by hydrogenation at low temperature and low pressure, and then hydrolyzed and crystallized to obtain high-purity succinic acid. The maleic anhydride conversion rate reaches 100%, and the succinic acid selectivity is greater than 95%; CN107473954A provides a green production method for succinic acid. In a fixed-bed reactor, one or several of maleic anhydride aqueous solution, maleic acid aqueous solution, and fumaric acid aqueous solution are subjected to hydrogenation reaction to produce succinic acid, and an ordered mesoporous material is used as the carrier to support noble metals as the hydrogenation catalyst, effectively avoiding the formation of DL-malic acid by the hydration reaction. The catalytic hydrogenation selectivity reaches 100%, the raw materials are completely converted into succinic acid, and the product purity reaches more than 99.5%; CN115745772A discloses a method for the esterification hydrogenation of maleic anhydride to produce succinic acid, including three steps: maleic anhydride esterification, hydrogenation of dimethyl maleate, and hydrolysis of diethyl succinate. The maleic anhydride esterification conversion rate is greater than 99.5%, the hydrogenation conversion rate of dimethyl maleate is greater than 99.5%, the selectivity of diethyl succinate exceeds 99.7%, and the high-pressure hydrolysis rate of diethyl succinate is greater than 98.5%. CN102863335B provides a method for the preparation of diethyl succinate, using maleic anhydride, alcohol, and hydrogen as raw materials, CO 2 as the solvent and acid catalyst, and carrying out esterification hydrogenation reaction under the action of a hydrogenation catalyst, with a diethyl succinate yield greater than 99%.

[0004] There are also numerous patent reports on the preparation of PBS alcohol monomer BDO from maleic anhydride. For example, CN114656331A uses a supported metal catalyst, with γ-butyrolactone as the solvent, and directly hydrogenates maleic anhydride to prepare BDO at a temperature of 50-80 °C and a pressure of 2-3 MPa, including three steps: hydrogenating maleic anhydride to succinic anhydride, hydrogenating succinic anhydride to γ-butyrolactone, and hydrogenating γ-butyrolactone to BDO; CN103946201B provides a method for preparing BDO by esterification hydrogenation of maleic anhydride. First, dimethyl maleate is prepared by esterifying maleic anhydride and methanol, then it is hydrogenated to dimethyl succinate under the action of a palladium catalyst, and finally dimethyl succinate is further hydrogenated to BDO in the presence of a copper catalyst, with γ-butyrolactone and tetrahydrofuran as by-products; WO8800937A1 subjects molten maleic anhydride and methanol to non-catalytic mono-esterification and acid ion exchange resin-catalyzed di-esterification, then gasifies the dimethyl maleate obtained by rectification and mixes it with hydrogen, and obtains dimethyl succinate through one-stage gas-phase hydrogenation, and then conducts two-stage gas-phase hydrogenation to obtain BDO.

[0005] There are also a large number of reports on the synthesis of PBS by esterification or transesterification polycondensation of acid monomers such as succinic anhydride, succinic acid or dimethyl succinate and alcohol monomer BDO. CN101328261B discloses a method for preparing a high molecular weight PBS degradable plastic, using succinic acid and BDO as polymerization monomers, and adopting a cerium salt-ethylene glycol antimony catalytic system to prepare PBS with a weight average molecular weight (M w ) of 56,000-125,000; CN102007159B discloses a method for continuously producing biodegradable polyesters. Using aliphatic dicarboxylic acids and aliphatic dihydroxy compounds as raw materials, through continuous esterification, condensation, polycondensation and addition polymerization reactions, a polyester product with M w of 95,000-165,000 is prepared. CN104130381B and CN104130382B report the preparation of aliphatic polyesters by three-step reactions of esterification, prepolymerization and final polymerization of succinic anhydride and aliphatic diols, and the resulting polyester M w reaches 190,000-320,000; CN114015026B discloses a method for synthesizing poly(butylene succinate) by ring-opening polymerization, using a mixture of succinic anhydride or succinic anhydride and succinic acid, and a mixture of tetrahydrofuran and 1,4-butanediol as monomers, and preparing PBS with M w of 15.7-22.6 through three steps of ring-opening esterification, negative pressure prepolymerization and vacuum final polymerization. CN102718950B uses succinic acid esters and 1,4-butanediol as raw materials, and adopts two-step polycondensation and stepwise addition of a composite catalyst to synthesize PBS with M w of 100,000-145,000. CN110563933B uses dimethyl succinate and BDO as polymerization monomers, and realizes transesterification polymerization through a high-efficiency catalyst. The product M w exceeds 250,000 and has good heat resistance and mechanical properties.

[0006] In order to reduce the manufacturing cost of PBS and improve the market competitiveness of polyester PBS, in recent years, a large number of studies have been carried out in the academic and industrial circles, mainly focusing on simplifying the synthesis process flow of monomers or polyester PBS, improving the yield and efficiency of monomer or PBS preparation, reducing the energy consumption of monomer production, optimizing the polymerization process and other single aspects of work, and great progress has been made. Moreover, attention has also begun to be paid to the system integration from raw material maleic anhydride to polyester PBS.

[0007] By co-producing two monomers of PBS through maleic anhydride hydrogenation, the preparation process flow can be shortened and the production cost of monomers can be reduced. For example, CN114181038B discloses a method for directly hydrogenating maleic anhydride to produce BDO and co-producing succinic anhydride. The first step is to hydrogenate maleic anhydride and separate γ-butyrolactone and succinic anhydride by multi-column distillation. The second step is to hydrogenate γ-butyrolactone and obtain BDO by distillation separation. The conversion rate of maleic anhydride is greater than 99%, and the total selectivity of BDO and succinic anhydride exceeds 90%. Moreover, the proportion of each product can be adjusted by adjusting the reaction conditions of the first-step maleic anhydride hydrogenation reaction. CN113332999B discloses a novel multi-component metal catalyst for hydrogenating maleic anhydride to prepare succinic anhydride and BDO. The conversion rate of maleic anhydride is 100%. When the reaction is carried out at low temperature and low pressure, the selectivity of succinic anhydride is greater than 99%. When the reaction is carried out at higher temperature and pressure, the selectivity of BDO exceeds 95%. By adjusting the reaction conditions, the purpose of producing two monomer products with the same catalyst can be achieved. CN217248853U discloses a co-production system of succinic anhydride, γ-butyrolactone, BDO and tetrahydrofuran. Its primary hydrogenation reaction unit is used for hydrogenating maleic anhydride by the solvent method to prepare succinic anhydride. The crystallization unit is used for separating crude succinic anhydride and the solvent. The secondary hydrogenation reaction unit is used for further hydrogenating succinic anhydride to generate a mixed crude product of BDO, γ-butyrolactone and tetrahydrofuran. The refining unit is used for separating the crude products.

[0008] Simplify and integrate the process flow for preparing monomers from maleic anhydride and the esterification or transesterification polycondensation synthesis of PBS from the monomers, and reduce the manufacturing cost of PBS. For example, CN114920913B provides a method for catalytic conversion of maleic anhydride to prepare PBS. Maleic anhydride is dissolved in tetrahydrofuran to prepare a maleic anhydride solution, and then hydrogenated to generate succinic anhydride and γ-butyrolactone. The γ-butyrolactone solution in tetrahydrofuran and succinic anhydride are separated; the γ-butyrolactone solution in tetrahydrofuran is further hydrogenated to obtain a BDO solution in tetrahydrofuran, and BDO and tetrahydrofuran are separated; then succinic anhydride and BDO are subjected to an esterification polycondensation reaction to produce the product PBS. CN112694602B and CN110563933B propose methods for producing PBS using maleic anhydride as a raw material. Maleic anhydride is hydrogenated to prepare succinic anhydride, and succinic anhydride is esterified with a lower aliphatic monohydric alcohol to prepare diethyl succinate, or maleic anhydride is esterified with a lower aliphatic monohydric alcohol to prepare diethyl maleate, and then hydrogenated to prepare diethyl succinate. Then, diethyl succinate and BDO are synthesized into PBS by transesterification polycondensation. CN106366297B discloses a preparation method for the overall industrial chain of polybutylene succinate. The overall reaction route is to rectify n-butane from a C4 fraction, produce liquid-phase maleic anhydride from n-butane, produce succinic anhydride from the liquid-phase maleic anhydride, subject succinic anhydride, BDO and a catalyst to ring-opening esterification in a nitrogen atmosphere, and then carry out a polycondensation reaction under vacuum conditions to obtain the target product PBS. CN 112920385B discloses a preparation method for PBS and its copolymers. First, an esterification hydrogenation product is prepared from BDO, maleic anhydride and a copolymeric dibasic acid under the action of an esterification catalyst and a hydrogenation catalyst. Then, after removing water and tetrahydrofuran from the esterification hydrogenation product, a polycondensation catalyst is added, and a reduced-pressure polycondensation reaction is carried out to prepare PBS and its copolymers.

[0009] In summary, for the existing technologies of chemically synthesizing PBS using maleic anhydride as the starting material, whether it is the process of step-by-step synthesizing acid monomers, alcohol monomers and PBS, or the process of co-producing acid monomers and alcohol monomers and then polycondensing to synthesize PBS, and the integrated process of monomers and polyesters, all use high-purity maleic anhydride products (liquid maleic anhydride) or commercial products (solid maleic anhydride) as the source raw materials. After complex and cumbersome reaction and separation processes, high-purity polymer-grade acid monomers such as succinic acid, succinic anhydride or dimethyl succinate and alcohol monomer BDO are respectively prepared first, and then these two monomers are used to synthesize polyester PBS through esterification or transesterification and polycondensation reactions. Although in recent years, through technological innovation, the co-production process and integrated process developed have shortened the process flow of synthesizing PBS, reduced the equipment investment and production cost, but it is still limited to the integration of existing single technologies of "starting from maleic anhydride products or commercial products, first synthesizing polyester monomers, and then synthesizing polyester PBS". There is still no focus on reducing the manufacturing cost of the starting material maleic anhydride and the overall process route integration from a global perspective. It still has not jumped out of the traditional thinking limitation of "using high-purity maleic anhydride products or commercial products as the starting material, and first synthesizing monomers and then synthesizing polyester", resulting in a still relatively long process flow for producing PBS and limited reduction in the manufacturing cost of PBS.

[0010] Currently, the investment and production cost of maleic anhydride industrial plants are mainly consumed in the solvent absorption and desorption of the maleic anhydride gas stream of the oxidation reaction product, the separation and purification of crude maleic anhydride, and the purification and recovery of the solvent; in addition, during the operation processes of maleic anhydride gas stream solvent absorption and desorption, maleic anhydride separation and purification, and solvent purification and recovery, due to the polymerization of acrylic acid, a by-product prone to polymerization, and the precipitation of by-products maleic acid and fumaric acid prone to crystallization in the maleic anhydride logistics, the pipelines and equipment are blocked, resulting in the inability of the plant to operate in a long cycle and requiring frequent cleaning. This is also the direction that the maleic anhydride industry strives to solve and the technological progress. If these post-treatment processes of the maleic anhydride gas stream can be omitted or downstream products can be directly produced using the maleic anhydride gas stream as the raw material, the raw material cost of maleic anhydride downstream products will be greatly reduced. Summary of the Invention

[0011] In view of this, the present invention provides an overall process route and production method for efficiently, with high quality and low cost preparing polybutylene succinate biodegradable plastics.

[0012] To achieve the above object, the technical solution of the present invention is specifically as follows:

[0013] A method for producing polybutylene succinate using gaseous crude maleic anhydride, comprising the following steps:

[0014] S1. Using 1,4-butanediol (BDO) as an absorbent, in the presence of an acid catalyst, esterifying and absorbing the crude maleic anhydride gas stream to prepare butylene glycol maleate (BM);

[0015] S2. The butanediol maleate (BM) prepared in step S1 is subjected to selective hydrogenation of the double bond in the presence of a supported noble metal catalyst to prepare butanediol succinate (BS).

[0016] S3. Part or all of the butanediol succinate (BS) prepared in step S2 is subjected to a polycondensation reaction in the presence of a polymerization catalyst to prepare polybutylene succinate (PBS).

[0017] The present invention is further configured such that the crude maleic anhydride gas stream is from the crude maleic anhydride gas stream of the product of the oxidation reaction of benzene or butane to produce maleic anhydride.

[0018] The present invention is further configured such that in step S1, the BM is a monoester or / and diester of butanediol maleate, and the reaction processes of preparing BM from BDO and maleic anhydride are shown in the following formulas (1) and (2) respectively:

[0019]

[0020] The present invention is further configured such that in step S1, the preparation of BM includes the following process: the crude maleic anhydride gas stream is quenched to 55 - 180 °C and then fed into the bottom of the esterification absorption tower, and a BDO absorbent containing 0 - 10.0 wt% of a liquid acid catalyst at a temperature of 25 - 65 °C is sprayed from the top of the tower, and the maleic anhydride gas stream and the absorbent operate in countercurrent; after 1 - 4 stages of spray esterification absorption, an esterification absorption liquid with a BM concentration of 10 - 50 wt% and a temperature of 52 - 120 °C is obtained.

[0021] Preferably, the crude maleic anhydride gas stream is quenched to 60 - 150 °C, the absorbent temperature is 25 - 60 °C, the concentration of the liquid acid catalyst is 0.1 - 5.0 wt%, the spray esterification absorption is carried out in 2 - 3 stages, the BM concentration of the esterification absorption liquid is controlled at 15 - 45 wt%, and the temperature of the esterification absorption liquid is 55 - 90 °C;

[0022] Preferably, the crude maleic anhydride gas stream is quenched to 65 - 120 °C, the absorbent temperature is 25 - 55 °C, the concentration of the liquid acid catalyst is 0.5 - 2.0 wt%, the spray esterification absorption is carried out in 2 - 3 stages, the BM concentration of the esterification absorption liquid is controlled to be 20 - 40 wt%, and the temperature of the esterification absorption liquid is controlled to be 58 - 75 °C.

[0023] The present invention is further configured such that in step S1, the acid catalyst is selected from at least one of alkyl sulfonic acid, aryl sulfonic acid, halo - carboxylic acid, halo - sulfonic acid, amides, metal salts or imide metal salts of alkyl sulfonic acid, aryl sulfonic acid or halo - carboxylic acid, and the metal element in the metal salt or its imide metal salt is selected from at least one of Li, Mg, Ba, Al, Ga, In, Fe, Cu, Ag, Zn, Sb, Bi, Ti, Zr, Sn or rare earth elements;

[0024] Preferably, the acid catalyst is selected from at least one of methanesulfonic acid, ethanesulfonic acid, propanesulfonic acid, benzenesulfonic acid, p-toluenesulfonic acid, β-naphthalenesulfonic acid, trifluoromethanesulfonic acid, pentafluoroethanesulfonic acid, trichloroacetic acid, trifluoroacetic acid, pentafluoropropionic acid, trifluoromethanesulfonamide, bis(trifluoromethanesulfonyl)imide, metal salts of fluorosulfonic acid, metal salts of methanesulfonic acid, metal salts of p-toluenesulfonic acid, metal salts of trifluoroacetic acid, metal salts of trifluoromethanesulfonic acid or metal salts of bis(trifluoromethanesulfonyl)imide;

[0025] More preferably, the metal salt of p-toluenesulfonic acid is selected from at least one of p-toluenesulfonates of Zn, Sc, Y, La, Nd, Sm or Yb; the metal salt of trifluoroacetic acid is selected from at least one of trifluoroacetates of Ba, Fe, Cu, Sc, Y, La, Nd or Yb; the metal salt of trifluoromethanesulfonic acid is selected from at least one of trifluoromethanesulfonates of Li, Mg, Al, Ga, In, Fe, Cu, Ag, Zn, Sb, Bi, Ti, Zr, Sn, Sc, Y, La, Nd, Sm or Yb; and the metal salt of bis(trifluoromethanesulfonyl)imide is selected from at least one of bis(trifluoromethanesulfonyl)imide salts of Li, Mg, Al, Fe, Cu, Zn, Bi, Zr, Sn, Sc, Y, La, Nd, Sm or Yb.

[0026] The present invention is further configured such that in step S2, the BS is butylene succinate monoester or / and diester, and the reaction processes for the selective hydrogenation of the BM double bond to prepare butylene succinate monoester or diester are respectively shown by the following formulas (3) and (4):

[0027] HOOCCH=CHCOO(CH 2 ) 4 OH+H 2 →HOOCCH 2 CH 2 COO(CH 2 ) 4 OH (3)

[0028] HO(CH 2 ) 4 OOCCH=CHCOO(CH 2 ) 4 OH+H 2 →HO(CH 2 ) 4 OOCCH 2 CH 2 COO(CH 2 ) 4 OH (4).

[0029] The present invention is further configured that in step S2, the preparation of BS includes the following process: using a slurry bed reactor, under the action of a supported noble metal catalyst, feeding the esterification absorption liquid after esterification absorption in step S1 into the slurry bed reactor, and operating in parallel flow or countercurrent flow with hydrogen to carry out selective hydrogenation reaction to prepare BS; wherein, the process conditions are: the concentration of BM in the fed esterification absorption liquid is 10-50 wt%, the feeding temperature is 52-150 °C, the weight hourly space velocity of the feed is 0.1-10.0 h -1 , the hydrogen pressure is 0.2-5.0 MPa, and the hydrogen-ester molar ratio is 2-50;

[0030] Preferably, the concentration of BM in the fed esterification absorption liquid is 15-45 wt%, the feeding temperature is 55-120 °C, the weight hourly space velocity of the feed is 0.5-5.0 h -1 , the hydrogen pressure is 0.5-3.0 MPa, and the hydrogen-ester molar ratio is 5-30;

[0031] More preferably, the concentration of BM in the fed esterification absorption liquid is 20-40 wt%, the feeding temperature is 60-90 °C, the weight hourly space velocity of the feed is 1.0-2.0 h -1 , the hydrogen pressure is 1.0-2.0 MPa, and the hydrogen-ester molar ratio is 10-20.

[0032] The present invention is further configured that in step S2, both the esterification absorption liquid and hydrogen are fed into the slurry bed reactor in parallel flow from the bottom, or the esterification absorption liquid is fed into the slurry bed reactor from the top and hydrogen is fed into the slurry bed reactor from the bottom in countercurrent flow.

[0033] The present invention is further configured that in step S2, the carrier of the supported noble metal catalyst is selected from activated carbon (AC), mesoporous carbon, carbon nanotubes, graphene, SiO 2 , Al 2 O 3 , TiO 2 or ZrO 2 , or at least one of hydrogen-type zeolite molecular sieves HM, HHEU (clinoptilolite), Hβ, HZSM-5, HZSM-22, HMCM-22, HMCM-41, HMCM-48, HMCM-49, HMCM-56, SBA-15, ZEO-1, ZEO-3 or KIT-6, and the noble metal is selected from one or more of Ru, Pd or Pt;

[0034] Preferably, the carrier of the supported noble metal catalyst is selected from AC, SiO 2 , Al 2 O 3 , SiO 2 -Al 2 O 3 , TiO 2 , ZrO2 , or HM, Hββ, HZSM-5, HZSM-22, HMCM-22, HMCM-41, HMCM-48, SBA-15, ZEO-3, KIT-6, and the noble metal is selected from one or two of Ru, Pd or Pt;

[0035] More preferably, the carrier of the supported noble metal catalyst is selected from AC, SiO 2 , Al 2 O 3 , HM, HZSM-5, HMCM-22, HMCM-48 or SBA-15, and the noble metal is Pd, Pt or Pd-Ru.

[0036] The present invention is further configured such that in step S2, the dosage of the supported noble metal catalyst is 0.2-5.0 wt% of the weight of the feed esterification absorption liquid; preferably 0.5-2.0 wt%; more preferably 0.5-1.0 wt%.

[0037] The present invention is further configured such that in step S2, in the supported noble metal catalyst, the noble metal is selected from one or more of Ru, Pd or Pt, the loading amount of Pd or Pt is 0.1-1.0 wt%, and the loading amount of Ru is 0.2-3.0 wt%.

[0038] Furthermore, in step S2, in the supported noble metal catalyst, the noble metal is selected from one or two of Ru, Pd or Pt, the loading amount of Pd or Pt is 0.2-0.7 wt%, and the loading amount of Ru is 0.5-3.0 wt%; more preferably, the noble metal is Pd, Pt or Pd-Ru, the loading amount of Pd or Pt is 0.3-0.5 wt%, and the loading amount of Ru is 1.0-2.0 wt%.

[0039] The present invention is further configured such that in step S2, the supported noble metal catalyst is a powder catalyst.

[0040] The present invention is further configured such that in step S3, the reaction process for preparing PBS by BS polycondensation is shown in the following formulas (5) and (6):

[0041] mHOOCCH 2 CH 2 COO(CH 2 ) 4 OH+HOCH 2 (CH 2 ) 2 CH 2 OH→HO(CH 2 ) 4 O[-OCCH 2 CH 2COO(CH 2 ) 4 O] m -H+mH 2 O (5)

[0042] nHO(CH 2 ) 4 OOCCH 2 CH 2 COO(CH 2 ) 4 OH→HO(CH 2 ) 4 O[-OCCH 2 CH 2 COO(CH 2 ) 4 O] n -H+(n-1)HOCH 2 (CH 2 ) 2 CH 2 OH (6).

[0043] The present invention is further configured such that in step S3, the process flow for preparing PBS by polycondensation of BS specifically includes the following steps:

[0044] S31. The liquid material after gas-liquid separation of the reaction product in step S2 is first sent to a flash distillation column for vacuum distillation after filtration and separation. A small amount of water is removed from the top of the column, and a mixture containing BS and BDO is obtained at the bottom of the column;

[0045] S32. After mixing the obtained material at the bottom of the flash distillation column and 20% - 30% of the total amount of the polymerization catalyst evenly, it is sent to an oligomerization reactor for vacuum oligomerization reaction;

[0046] S33. The oligomerization reaction product and the remaining 70 - 80% of the polymerization catalyst are mixed and then sent to a prepolymerization reactor for negative pressure prepolymerization reaction;

[0047] S34. The prepolymerization product is sent to a final polymerization reactor for vacuum final polymerization reaction to obtain PBS.

[0048] The present invention is further configured such that the polymerization catalyst is selected from at least one of titanium chloride, titanium butoxy chloride, tetraalkyl titanates, tetra-hydroxy titanates, dialkylene glycol titanates, tin chloride, stannous organic acids, tetraalkyl tins, dialkyl tin fatty acids, antimony chloride, antimony fatty diols, germanium chloride, tetrabutoxy germanium, tetraalkyl germanium; preferably, the polymerization catalyst is selected from at least one of tetrabutyl titanate, tetra-hydroxybutyl titanate, diethylene glycol titanate, dibutylene glycol titanate and antimony glycolate; more preferably, the polymerization catalyst is tetrabutyl titanate or diethylene glycol titanate;

[0049] The total amount of the polymerization catalyst used is 0.01-1.00% of the mass of BS, preferably 0.02-0.50%, and more preferably 0.05-0.10%;

[0050] The operating conditions of the flash tower are as follows: pressure 50-101 kPa, top temperature 50-100 °C, bottom temperature 60-130 °C; preferably, pressure 75-101 kPa, top temperature 60-100 °C, bottom temperature 80-110 °C;

[0051] The operating conditions of the oligomerization reaction are: temperature 150-200 °C, pressure 20-70 kPa, and material residence time 20-60 min; preferably, temperature 160-180 °C, pressure 30-50 kPa, and material residence time 30-45 min;

[0052] The operating conditions of the prepolymerization reaction are: temperature 180-230 °C, pressure 1-50 kPa, and material residence time 20-60 min; preferably, temperature 190-210 °C, pressure 5-30 kPa, and material residence time 30-45 min;

[0053] The operating conditions of the final polymerization reaction are: temperature 200-260 °C, pressure 10-200 Pa, and material residence time 60-180 min; preferably, temperature 220-240 °C, pressure 50-90 Pa, and material residence time 90-120 min.

[0054] The present invention is further configured such that in step S3, when a part of BS is taken for polycondensation reaction to prepare PBS, it further includes step S4: in the presence of a hydrogenation catalyst, the remaining BS is subjected to a hydrogenation reaction to prepare BDO as the absorbent in step (1), and the reaction route is shown in the following formulas (7) and (8):

[0055] HOOCCH 2 CH 2 COO(CH 2 ) 4 OH + 4H 2 → 2HOCH 2 (CH 2 ) 2 CH 2 OH + H 2 O (7)

[0056] HO(CH 2 ) 4 OOCCH 2 CH 2 COO(CH 2 ) 4 OH + 4H 2 → 3HOCH2 (CH 2 ) 2 CH 2 OH (8).

[0057] The present invention is further configured such that in step S4, the hydrogenation catalyst is selected from CuZnXY or CuZnXY / Z catalyst, wherein X is selected from at least one element of Al, Zr, Mo, W, Cr, Mn or Ni, Y is selected from one of alkaline earth or / and rare earth elements, the carrier Z is selected from a high specific surface area carrier or a composite carrier of a high specific surface area carrier and a high thermal conductivity carrier, the high specific surface area carrier is selected from at least one of activated carbon, SiO 2 , γ-Al 2 O 3 , θ-Al 2 O 3 , SiO 2 -Al 2 O 3 , TiO 2 , molecular sieve, diatomite or bentonite, and the high thermal conductivity carrier is selected from at least one of graphene, Si, α-SiO 2 , α-Al 2 O 3 , BeO, SiC, Mo 2 C, BN or AlN;

[0058] Preferably, in the CuZnXY or CuZnXY / Z catalyst, X is selected from at least one of Al, Zr, Cr or Mn, Y is selected from at least one of Mg, Ca, Ba, La, Ce, Sm or Y, the carrier Z is selected from a high specific surface area carrier or a composite carrier of a high specific surface area carrier and a high thermal conductivity carrier, the high specific surface area carrier is selected from SiO 2 , γ-Al 2 O 3 , θ-Al 2 O 3 , SiO 2 -Al 2 O 3 or molecular sieve β, ZSM-5, MCM-22, SBA-15, and the high thermal conductivity carrier is selected from Si, α-Al 2 O 3 or SiC;

[0059] More preferably, the hydrogenation catalyst is selected from at least one of the particulate catalysts of CuZnAlCe, CuZnZrCe, CuZnAlCe / SiO 2 or CuZnZrCe / SiO 2 -Si.

[0060] The present invention is further configured such that in step S4, the remaining BS is prepared into BDO through a hydrogenation reaction, which specifically includes the following process:

[0061] S41. Feed the remaining BS into a hydrogenation reactor for hydrogenation reaction;

[0062] S42. After the hydrogenation reaction product is condensed and cooled, it is sent to a gas-liquid separator for separation, and the liquid is sent to a light component removal column;

[0063] S43. Feed the bottom material of the light component removal column described in step S42 into a BDO refining column to obtain high-purity BDO, which is used as the absorbent in step S1;

[0064] Among them, the hydrogenation reactor is a fixed-bed reactor, preferably an isothermal fixed-bed or adiabatic fixed-bed reactor, and more preferably a tube bundle isothermal fixed-bed or inter-stage heat exchange adiabatic fixed-bed reactor;

[0065] The hydrogenation reaction conditions are as follows: the mass concentration of the BDO solution of the fed BS is 8-40 wt%, the feed weight hourly space velocity is 0.1-5.0 h -1 , the feed temperature is 200-300 °C, the reaction pressure is 2.0-8.0 MPa, and the hydrogen-ester molar ratio is 20-100; preferably, the concentration of the fed BS is 12-35 wt%, the feed weight hourly space velocity is 0.3-3.0 h -1 , the feed temperature is 220-280 °C, the reaction pressure is 2.5-6.5 MPa, and the hydrogen-ester molar ratio is 30-80; more preferably, the concentration of the fed BS is 15-30 wt%, the feed weight hourly space velocity is 0.5-1.5 h -1 , the feed temperature is 240-260 °C, the reaction pressure is 3.0-5.0 MPa, and the hydrogen-ester molar ratio is 40-70;

[0066] The operating conditions of the light component removal column are as follows: the top pressure is 10-101 kPa, the top temperature is 75-130 °C, and the bottom temperature is 120-180 °C. Preferably, the top pressure is 15-75 kPa, the top temperature is 85-120 °C, and the bottom temperature is 125-150 °C. More preferably, the top pressure is 20-50 kPa, the top temperature is 95-110 °C, and the bottom temperature is 130-140 °C;

[0067] The operating conditions of the BDO refining column are as follows: the top pressure is 2-50 kPa, the number of theoretical plates is 20-50, the reflux ratio is 0.5-5, the top temperature is 120-150 °C, and the bottom temperature is 160-200 °C. Preferably, the top pressure is 5-40 kPa, the number of theoretical plates is 25-45, the reflux ratio is 0.7-4.0, the top temperature is 125-145 °C, and the bottom temperature is 165-195 °C. More preferably, the top pressure is 10-30 kPa, the number of theoretical plates is 30-40, the reflux ratio is 1.0-3.0, the top temperature is 130-140 °C, and the bottom temperature is 170-190 °C.

[0068] The present invention is further configured such that in step S2, after the reaction, the material is subjected to gas-liquid separation. The separated hydrogen gas is sent to the slurry bed reactor for BM hydrogenation for recycling, or is transported to step S4 for recycling in the hydrogenation reactor for BS hydrogenation; after the separated liquid-phase material is subjected to solid-liquid separation, the supported noble metal catalyst for BM hydrogenation is recovered. The method of solid-liquid separation is selected from filtration, pressure filtration, centrifugation or sedimentation separation; the liquid part after separating the catalyst is divided into two paths, one path is sent to the flash distillation tower in step S3, and the other path is sent to the BS hydrogenation reactor in step S4.

[0069] The present invention is further configured such that in step S3, the gaseous materials from the oligomerization reaction, prepolymerization reaction and final polymerization reaction are collected and condensed and then sent to step S4 for use in the hydrogenation reactor for BS hydrogenation; the obtained final polymerization product PBS is cooled and pelletized underwater to obtain PBS pellet products.

[0070] The present invention is further configured such that in step S1, the tail gas part after the crude maleic anhydride gas stream is subjected to esterification absorption is recycled back to the maleic anhydride oxidation reactor, and the remaining tail gas is sent to the incinerator.

[0071] Further, 45-85 vol% of the tail gas after the crude maleic anhydride gas stream is subjected to esterification absorption is recycled back to the maleic anhydride oxidation reactor; preferably, the tail gas recycle amount is controlled to be 50-80 vol%; more preferably, the tail gas recycle amount is controlled to be 55-75 vol%.

[0072] The present invention is further configured such that in step S4, 90-95 wt% of the material drawn from the bottom of the BDO refining tower is sent back to the hydrogenation reactor for BS hydrogenation, and the remaining 5-10 wt% is discharged as a heavy component and sent to the incinerator.

[0073] The present invention also provides a process system for producing polybutylene succinate from gaseous crude maleic anhydride, which includes a maleic anhydride oxidation reactor, a maleic anhydride esterification absorption tower, a BM hydrogenation reaction system and a BS hydrogenation reaction system connected in sequence. The BM hydrogenation reaction system includes a BM hydrogenation reactor and a gas-liquid separator connected in sequence. The BS polymerization reaction system includes a flash distillation tower, an oligomerization and polycondensation kettle, a prepolymerization kettle and a final polycondensation kettle connected in sequence. The BS hydrogenation reaction system includes a BS hydrogenation reactor, a second gas-liquid separator, a light component removal tower and a BDO refining tower connected in sequence;

[0074] wherein, the liquid-phase outlet of the gas-liquid separator of the BM hydrogenation reaction system is connected to the BS hydrogenation reactor and the flash distillation tower; the top of the BDO refining tower is connected to the maleic anhydride esterification absorption tower; control valves are installed on each connecting pipeline in the process system.

[0075] The present invention is further configured such that the tail gas pipeline of the maleic anhydride esterification absorption tower is connected to the maleic anhydride oxidation reactor and the incinerator, the gas phase outlet of the gas-liquid separator is connected to the BM hydrogenation reactor and the BS hydrogenation reactor, the gas phase outlet of the second gas-liquid separator is connected to the BS hydrogenation reactor, and the bottom of the BDO refining tower is connected to the BS hydrogenation reactor and the post-treatment device.

[0076] Compared with the prior art, the present invention has the following beneficial effects:

[0077] (1) The principle and process of the present invention are as follows: BDO is used as the absorbent and esterifying agent for the maleic anhydride gas stream generated by the oxidation of benzene or butane in the industrial maleic anhydride plant. While absorbing maleic anhydride, an esterification reaction occurs to generate maleic acid-1,4-butanediol monoester and diester (collectively referred to as BM); BM then undergoes double bond hydrogenation at low temperature and low pressure with high activity and high selectivity to generate polyester precursors succinic acid-1,4-butanediol monoester and diester (collectively referred to as BS); then part of the BS undergoes three-stage condensation polymerization reactions of oligomerization, prepolymerization and final polymerization to synthesize polyester PBS, and the remaining part of the BS undergoes efficient ester group hydrogenation at higher temperature and pressure to generate BDO, which is recycled as the maleic anhydride absorbent. The present invention adopts a brand-new process route with different principles, fundamentally changing the overall process flow for synthesizing PBS;

[0078] (2) The present invention greatly simplifies and shortens the overall process flow, can significantly reduce the overall investment, and makes the operation more convenient. The present invention realizes the production of polyester PBS starting from gaseous crude maleic anhydride, fundamentally changing and shortening the overall process route and flow for producing PBS, specifically manifested as: skipping and eliminating the production steps of maleic anhydride, polyester acid monomers and alcohol monomers in the prior art, as well as the process of esterifying acid monomers and alcohol monomers to prepare polyester precursor BS, and directly obtaining polyester precursor BS from the maleic anhydride gas stream through two steps of esterification absorption and double bond hydrogenation; because gaseous crude maleic anhydride is used as the raw material, the processes of solvent absorption and desorption, crude maleic anhydride separation and refining, solvent purification and recovery, etc. that must be carried out in the prior art for producing maleic anhydride products are eliminated, thereby eliminating a large amount of investment in the maleic anhydride absorption and separation device, and also avoiding the cumbersome operation of frequent shutdown and cleaning due to pipeline and equipment blockage; because the same solvent BDO is used in each reaction step and the target product has high selectivity, the esterification absorption liquid and the double bond hydrogenation product do not need to be separated and refined. Only the gaseous material of negative pressure polymerization and the ester group hydrogenation product need to be separated and refined together once to obtain the precursor for producing PBS and the recycled absorption solvent BDO, eliminating the process of separating and refining the target product in each reaction step of the prior art.

[0079] (3) The raw materials of the present invention are inexpensive, with low material and energy consumption, which can significantly reduce the production cost of PBS and improve the market competitiveness of products. In the whole process of the overall process of producing PBS in the present invention, a single solvent BDO is used. At the same time, BDO is also the esterifying agent of maleic anhydride and the structural unit component of PBS. Therefore, in the overall process from the esterification absorption of the maleic anhydride gas stream to the final obtaining of PBS, only one separation and purification of materials is required. Compared with the prior art, the raw materials are inexpensive, and the material and energy consumption of the overall process are low, greatly reducing the raw material cost and operation cost of producing PBS, and thus greatly reducing the comprehensive production cost of PBS.

[0080] (4) The present invention uses BDO as the absorbent and esterifying agent, and hydrogenates and saturates the unsaturated oxygen-containing compounds, which not only avoids the blockage of pipelines and equipment, but also improves the yield and output of the target product. Compared with the prior art that uses high-boiling heavy solvents DBP or DIBE, BDO not only has good solubility for maleic anhydride and its esterification products, but also has good solubility for by-products such as acrylic acid, maleic acid, and fumaric acid in the maleic anhydride gas stream, and reacts with them to form 1,4-butanediol esters of acrylic acid and BM, so that maleic acid, fumaric acid and their esterification products with high freezing points will not crystallize out; in addition, the esterification absorption liquid containing unsaturated oxygen-containing compounds is hydrogenated and saturated in a mild condition of low temperature and low pressure in a timely manner to prevent the occurrence of double bond polymerization reaction to form polymers. The present invention avoids the crystallization precipitation of by-products and the formation of double bond polymers in the absorption and subsequent separation and purification processes of the maleic anhydride gas stream, thereby avoiding the blockage of pipelines and equipment in the maleic anhydride device and greatly extending its operation cycle; because the by-products maleic acid, fumaric acid and the esterification product BM of BDO in the maleic anhydride gas stream are the same as the esterification products of maleic anhydride and BDO, and all become the precursors of the polyester PBS after hydrogenation, the present invention can also convert the by-products of the maleic anhydride device into the target product, thereby improving the yield and output of the final product PBS. Brief Description of the Drawings

[0081] Figure 1 Process flow diagram for producing PBS from maleic anhydride gas stream generated by benzene or butane oxidation;

[0082] Figure 2 Material flow diagram for producing PBS from maleic anhydride gas stream generated by benzene or butane oxidation. Detailed Description of the Invention

[0083] The following combines specific embodiments and refers to the attached Figure 1 and 2 to further elaborate on the present invention. It should be noted that the embodiments described in this part are only a part of the embodiments of the present invention, rather than all embodiments. Based on the embodiments of the present invention, all other embodiments obtained by those of ordinary skill in the art without creative labor fall within the scope of protection of the present invention.

[0084] Symbol Explanation:

[0085] SAA is succinic anhydride, BDO is 1,4 - butanediol, GBL is γ - butyrolactone, DMS is dimethyl succinate, BM is a general term for butylene glycol mono - ester or / and diester of maleic acid, BS is a general term for butylene glycol mono - ester or / and diester of succinic acid, and PBS is poly(butylene succinate).

[0086] In the following process, the total yield of PBS is the ratio of the weight of the PBS pellet product to the weight of maleic anhydride in the crude maleic anhydride gas stream fed into the absorption tower × 100%.

[0087] The present invention provides a method for producing poly(butylene succinate) from gaseous crude maleic anhydride, comprising the following steps:

[0088] S1. Using BDO as an absorbent, in the presence of an acid catalyst, esterifying and absorbing the crude maleic anhydride gas stream to prepare BM, and the crude maleic anhydride gas stream comes from the product crude maleic anhydride gas stream of the oxidation reaction of benzene or butane to maleic anhydride.

[0089] S2. Subjecting the BM obtained in step S1 to selective hydrogenation of double bonds in the presence of a supported noble metal catalyst to prepare BS;

[0090] S3. Subjecting part or all of the BS obtained in step S2 to polycondensation reaction in the presence of a polymerization catalyst to obtain PBS;

[0091] The material flow chart for the above production of PBS is as Figure 1 shown.

[0092] In an embodiment of the present invention, in step S1, the preparation of butylene glycol maleate includes the following process: After quickly cooling the crude maleic anhydride gas stream to 55 - 180°C, it is sent to the bottom of the esterification absorption tower, and the BDO absorbent containing 0 - 10.0 wt% of liquid acid catalyst at a temperature of 25 - 65°C is sprayed from the top of the tower. The maleic anhydride gas stream and the absorbent operate counter - currently; after 1 - 4 - stage spray esterification absorption, an esterification absorption liquid with a butylene glycol maleate concentration of 10 - 50 wt% and a temperature of 52 - 120°C is obtained; preferably, the crude maleic anhydride gas stream is quickly cooled to 60 - 150°C, the absorbent temperature is 25 - 60°C, the liquid acid catalyst concentration is 0.1 - 5.0 wt%, the spray esterification absorption is carried out in 2 - 3 stages, the butylene glycol maleate concentration of the esterification absorption liquid is controlled at 15 - 45 wt%, and the temperature of the esterification absorption liquid is 55 - 90°C; more preferably, the crude maleic anhydride gas stream is quickly cooled to 65 - 120°C, the absorbent temperature is 25 - 55°C, the liquid acid catalyst concentration is 0.5 - 2.0 wt%, the spray esterification absorption is carried out in 2 - 3 stages, the butylene glycol maleate concentration of the esterification absorption liquid is controlled at 20 - 40 wt%, and the temperature of the esterification absorption liquid is controlled at 58 - 75°C.

[0093] In an embodiment of the present invention, in step S1, the acid catalyst is selected from at least one of alkylsulfonic acid, arylsulfonic acid, halogenated carboxylic acid, halogenated sulfonic acid, alkylsulfonic acid, arylsulfonic acid or amide, metal salt or imide metal salt of halogenated carboxylic acid, wherein the metal element in the metal salt or imide metal salt is selected from at least one of Li, Mg, Ba, Al, Ga, In, Fe, Cu, Ag, Zn, Sb, Bi, Ti, Zr, Sn or rare earth elements; preferably, the acid catalyst is selected from at least one of methanesulfonic acid, ethanesulfonic acid, propanesulfonic acid, benzenesulfonic acid, p-toluenesulfonic acid, β-naphthalenesulfonic acid, trifluoromethanesulfonic acid, pentafluoroethanesulfonic acid, trichloroacetic acid, trifluoroacetic acid, pentafluoropropionic acid, trifluoromethanesulfonamide, bis(trifluoromethanesulfonyl)imide, metal salt of fluorosulfonic acid, metal salt of methanesulfonic acid, metal salt of p-toluenesulfonic acid, metal salt of trifluoroacetic acid, metal salt of trifluoromethanesulfonic acid or bis(trifluoromethanesulfonyl)imide metal salt; more preferably, the metal salt of p-toluenesulfonic acid is selected from at least one of p-toluenesulfonates of Zn, Sc, Y, La, Nd, Sm or Yb; the metal salt of trifluoroacetic acid is selected from at least one of trifluoroacetates of Ba, Fe, Cu, Sc, Y, La, Nd or Yb; the metal salt of trifluoromethanesulfonic acid is selected from at least one of trifluoromethanesulfonates of Li, Mg, Al, Ga, In, Fe, Cu, Ag, Zn, Sb, Bi, Ti, Zr, Sn, Sc, Y, La, Nd, Sm or Yb; the bis(trifluoromethanesulfonyl)imide metal salt is selected from at least one of bis(trifluoromethanesulfonyl)imide salts of Li, Mg, Al, Fe, Cu, Zn, Bi, Zr, Sn, Sc, Y, La, Nd, Sm or Yb.

[0094] In an embodiment of the present invention, in step S2, the preparation of butylene succinate includes the following process: using a slurry bed reactor, under the action of a supported noble metal catalyst, feeding the esterification absorption liquid after esterification absorption in step S1 into the slurry bed reactor, and operating in parallel or countercurrent with hydrogen to carry out a selective hydrogenation reaction to prepare butylene succinate; wherein, the process conditions are: the concentration of butylene glycol maleate in the fed esterification absorption liquid is 10-50 wt%, the feeding temperature is 52-150 °C, the weight hourly space velocity of the feed is 0.1-10.0 h -1 -1, the hydrogen pressure is 0.2-5.0 MPa and the hydrogen-to-ester molar ratio is 2-50; preferably, the concentration of butylene glycol maleate in the fed esterification absorption liquid is 15-45 wt%, the feeding temperature is 55-120 °C, the weight hourly space velocity of the feed is 0.5-5.0 h -1 -1, the hydrogen pressure is 0.5-3.0 MPa and the hydrogen-to-ester molar ratio is 5-30; more preferably, the concentration of butylene glycol maleate in the fed esterification absorption liquid is 20-40 wt%, the feeding temperature is 60-90 °C, the weight hourly space velocity of the feed is 1.0-2.0 h -1 -1, the hydrogen pressure is 1.0-2.0 MPa and the hydrogen-to-ester molar ratio is 10-20.

[0095] In an embodiment of the present invention, in step S2, both the esterification absorption liquid and hydrogen are fed into the slurry bed reactor in parallel from the bottom, or the esterification absorption liquid is fed into the slurry bed reactor from the top and hydrogen is fed into the slurry bed reactor from the bottom in countercurrent.

[0096] In an embodiment of the present invention, in step S2, the dosage of the supported noble metal catalyst is 0.2 - 5.0 wt% of the weight of the fed esterification absorption liquid, and the carrier is selected from activated carbon (AC), mesoporous carbon, carbon nanotubes, graphene, SiO 2 , Al 2 O 3 , TiO 2 or ZrO 2 , or at least one of the hydrogen-type zeolite molecular sieves HM, HHEU (clinoptilolite), Hβ, HZSM-5, HZSM-22, HMCM-22, HMCM-41, HMCM-48, HMCM-49, HMCM-56, SBA-15, ZEO-1, ZEO-3 or KIT-6, the noble metal is selected from one or more of Ru, Pd or Pt, the loading amount of Pd or Pt is 0.1 - 1.0 wt%, and the loading amount of Ru is 0.2 - 3.0 wt%; preferably, the dosage of the supported noble metal catalyst is 0.5 - 2.0 wt% of the weight of the fed esterification absorption liquid, and the carrier is selected from AC, SiO 2 , Al 2 O 3 , SiO 2 -Al 2 O 3 , TiO 2 or ZrO 2 , or HM, Hβ, HZSM-5, HZSM-22, HMCM-22, HMCM-41, HMCM-48, SBA-15, ZEO-3 or KIT-6, the noble metal is selected from one or two of Ru, Pd or Pt, the loading amount of Pd or Pt is 0.2 - 0.7 wt%, and the loading amount of Ru is 0.5 - 3.0 wt%; more preferably, the dosage of the supported noble metal catalyst is 0.5 - 1.0 wt% of the weight of the fed esterification absorption liquid, and the carrier is selected from AC, SiO 2 , Al 2 O 3 , HM, HZSM-5, HMCM-22, HMCM-48 or SBA-15, the noble metal is Pd, Pt or Pd-Ru, the loading amount of Pd or Pt is 0.3 - 0.5 wt%, and the loading amount of Ru is 1.0 - 2.0 wt%.

[0097] In an embodiment of the present invention, in step S2, the supported noble metal catalyst is a powder catalyst.

[0098] In an embodiment of the present invention, in step S3, the process flow of preparing polybutylene succinate by polycondensation of butylene succinate specifically includes the following steps:

[0099] S31. The liquid material after gas-liquid separation of the reaction product in step S2 is first sent to a flash distillation column for vacuum distillation after filtration and separation of part or all of the liquid. A small amount of water is removed from the top of the column, and a mixture containing butylene succinate and BDO is obtained at the bottom of the column.

[0100] S32. After mixing the obtained material at the bottom of the flash distillation column with 20% - 30% of the total amount of the polymerization catalyst evenly, it is sent to an oligomerization reactor for vacuum oligomerization reaction.

[0101] S33. The oligomerization reaction product is mixed with the remaining 70 - 80% of the polymerization catalyst and then sent to a prepolymerization reactor for negative pressure prepolymerization reaction.

[0102] S34. The prepolymer product is sent to a final polymerization reactor for vacuum final polymerization reaction to obtain polybutylene succinate.

[0103] In an embodiment of the present invention, the polymerization catalyst is selected from at least one of titanium chloride, titanium butoxy chloride, tetraalkyl titanate, tetrahydroxy titanate, dialkylene glycol titanate, tin chloride, stannous organic acid, tetraalkyl tin, dialkyl tin fatty acid, antimony chloride, antimony fatty diol, germanium chloride, tetrabutoxy germanium, tetraalkyl germanium; preferably, the polymerization catalyst is selected from at least one of tetrabutyl titanate, tetra(2-hydroxybutyl) titanate, diethylene glycol titanate, dibutylene glycol titanate and antimony glycolate; more preferably, the polymerization catalyst is tetrabutyl titanate or diethylene glycol titanate;

[0104] The total amount of the polymerization catalyst used is 0.01 - 1.00% of the mass of butylene succinate, preferably 0.02 - 0.50%, more preferably 0.05 - 0.10%;

[0105] The operating conditions of the flash distillation column are: pressure 50 - 101 kPa, top temperature 50 - 100 °C, bottom temperature 60 - 130 °C; preferably, pressure 75 - 101 kPa, top temperature 60 - 100 °C, bottom temperature 80 - 110 °C;

[0106] The operating conditions of the oligomerization reaction are: temperature 150 - 200 °C, pressure 20 - 70 kPa and material residence time 20 - 60 min; preferably, temperature 160 - 180 °C, pressure 30 - 50 kPa and material residence time 30 - 45 min;

[0107] The operating conditions of the prepolymerization reaction are as follows: temperature 180 - 230 °C, pressure 1 - 50 kPa, and residence time of the material 20 - 60 min; preferably, temperature 190 - 210 °C, pressure 5 - 30 kPa, and residence time of the material 30 - 45 min;

[0108] The operating conditions of the post-polymerization reaction are as follows: temperature 200 - 260 °C, pressure 10 - 200 Pa, and residence time of the material 60 - 180 min; preferably, temperature 220 - 240 °C, pressure 50 - 90 Pa, and residence time of the material 90 - 120 min.

[0109] In the embodiment of the present invention, when preparing polybutylene succinate by polycondensation of a part of butylene succinate prepared in step S2, it further includes step S4: subjecting the remaining part of butylene succinate to a hydrogenation reaction in the presence of a hydrogenation catalyst to prepare 1,4-butanediol, which is used as the absorbent in step (1).

[0110] In the embodiment of the present invention, in step S4, the hydrogenation catalyst is selected from CuZnXY or CuZnXY / Z catalyst, wherein X is selected from at least one element of Al, Zr, Mo, W, Cr, Mn or Ni, Y is selected from one of alkaline earth or / and rare earth elements, and the carrier Z is selected from a high specific surface area carrier or a composite carrier of a high specific surface area carrier and a high thermal conductivity carrier. The high specific surface area carrier is selected from at least one of activated carbon, SiO 2 、γ-Al 2 O 3 、θ-Al 2 O 3 、SiO 2 -Al 2 O 3 、TiO 2 、molecular sieve, diatomaceous earth or bentonite, and the high thermal conductivity carrier is selected from at least one of graphene, Si, α-SiO 2 、α-Al 2 O 3 、BeO、SiC、Mo 2 C、BN or AlN; preferably, in the CuZnXY or CuZnXY / Z catalyst, X is selected from at least one of Al, Zr, Cr or Mn, Y is selected from at least one of Mg, Ca, Ba, La, Ce, Sm or Y, and the carrier Z is selected from a high specific surface area carrier or a composite carrier of a high specific surface area carrier and a high thermal conductivity carrier. The high specific surface area carrier is selected from SiO 2 、γ-Al 2 O 3 、θ-Al 2 O 3 、SiO 2 -Al 2 O3 or molecular sieve β, ZSM-5, MCM-22, SBA-15, and the high thermal conductivity carrier is selected from Si, α-Al 2 O 3 or SiC; more preferably, the hydrogenation catalyst is selected from at least one of CuZnAlCe, CuZnZrCe, CuZnAlCe / SiO 2 or CuZnZrCe / SiO 2 -Si particulate catalysts.

[0111] In an embodiment of the present invention, in step S4, the remaining butylene succinate is subjected to a hydrogenation reaction to prepare 1,4-butanediol, which specifically includes the following process:

[0112] S41. Feed the remaining butylene succinate into a hydrogenation reactor for hydrogenation reaction;

[0113] S42. After the hydrogenation reaction product is condensed and cooled, it is sent to a gas-liquid separator for separation, and the liquid is sent to a light component removal tower;

[0114] S43. Feed the bottom material of the light component removal tower in step S42 into a BDO refining tower to obtain high-purity 1,4-butanediol, which is used as the absorbent in step S1;

[0115] Among them, the hydrogenation reactor is a fixed-bed reactor, preferably an isothermal fixed-bed or adiabatic fixed-bed reactor, more preferably a tube-type isothermal fixed-bed or inter-stage heat exchange adiabatic fixed-bed reactor;

[0116] The hydrogenation reaction conditions are as follows: the mass concentration of the 1,4-butanediol solution of the fed butylene succinate is 8-40 wt%, the feed weight hourly space velocity is 0.1-5.0 h -1 , the feed temperature is 200-300 °C, the reaction pressure is 2.0-8.0 MPa, and the hydrogen-to-ester molar ratio is 20-100; preferably, the concentration of the fed butylene succinate is 12-35 wt%, the feed weight hourly space velocity is 0.3-3.0 h -1 , the feed temperature is 220-280 °C, the reaction pressure is 2.5-6.5 MPa, and the hydrogen-to-ester molar ratio is 30-80; more preferably, the concentration of the fed butylene succinate is 15-30 wt%, the feed weight hourly space velocity is 0.5-1.5 h -1 , the feed temperature is 240-260 °C, the reaction pressure is 3.0-5.0 MPa, and the hydrogen-to-ester molar ratio is 40-70;

[0117] The operating conditions of the light component removal column are as follows: the top pressure is 10 - 101 kPa, the top temperature is 75 - 130 °C, and the bottom temperature is 120 - 180 °C. Preferably, the top pressure is 15 - 75 kPa, the top temperature is 85 - 120 °C, and the bottom temperature is 125 - 150 °C. More preferably, the top pressure is 20 - 50 kPa, the top temperature is 95 - 110 °C, and the bottom temperature is 130 - 140 °C;

[0118] The operating conditions of the BDO refining column are as follows: the top pressure is 2 - 50 kPa, the number of theoretical plates is 20 - 50, the reflux ratio is 0.5 - 5, the top temperature is 120 - 150 °C, and the bottom temperature of the column kettle is 160 - 200 °C. Preferably, the top pressure is 5 - 40 kPa, the number of theoretical plates is 25 - 45, the reflux ratio is 0.7 - 4.0, the top temperature is 125 - 145 °C, and the bottom temperature of the column kettle is 165 - 195 °C. More preferably, the top pressure is 10 - 30 kPa, the number of theoretical plates is 30 - 40, the reflux ratio is 1.0 - 3.0, the top temperature is 130 - 140 °C, and the bottom temperature of the column kettle is 170 - 190 °C.

[0119] In the embodiment of the present invention, in step S2, after the reaction material is subjected to gas-liquid separation, the separated hydrogen is sent to the slurry bed reactor for BM hydrogenation for recycling, or transported to step S4 for recycling in the hydrogenation reactor for BS hydrogenation; after the separated liquid-phase material is subjected to solid-liquid separation, the supported noble metal catalyst for BM hydrogenation is recovered. The method of solid-liquid separation is selected from filtration, pressure filtration, centrifugation or sedimentation separation; the liquid part after separating the catalyst is divided into two paths, one path is sent to the flash tower in step S3, and the other path is sent to the BS hydrogenation reactor in step S4.

[0120] In the embodiment of the present invention, in step S3, the gas-phase materials of the oligomerization reaction, prepolymerization reaction and final polymerization reaction are collected and condensed and then sent to step S4 for use in the hydrogenation reactor for BS hydrogenation; the obtained final polymerization product PBS is cooled and pelletized underwater to obtain PBS pellet products.

[0121] In the embodiment of the present invention, in step S1, a part of the tail gas after the crude maleic anhydride gas stream is esterified and absorbed is recycled back to the maleic anhydride oxidation reactor, and the remaining tail gas is sent to the incinerator; further, 45 - 85 vol% of the tail gas after the crude maleic anhydride gas stream is esterified and absorbed is recycled back to the maleic anhydride oxidation reactor; preferably, the tail gas circulation amount is controlled to be 50 - 80 vol%; more preferably, the tail gas circulation amount is controlled to be 55 - 75 vol%.

[0122] In the embodiment of the present invention, in step S4, 90 - 95 wt% of the material drawn from the bottom of the BDO refining column is sent back to the hydrogenation reactor for BS hydrogenation, and the remaining 5 - 10 wt% is discharged as a heavy component and sent to the incinerator.

[0123] The present invention also provides a process system for producing polybutylene succinate from gaseous crude maleic anhydride. As Figure 2 shown, it includes a maleic anhydride oxidation reactor, a maleic anhydride esterification absorption tower, a BM hydrogenation reaction system, and a BS hydrogenation reaction system connected in sequence. The BM hydrogenation reaction system includes a BM hydrogenation reactor and a gas-liquid separator connected in sequence. The BS polymerization reaction system includes a flash tower, a homogeneous polycondensation kettle, a pre-polycondensation kettle, and a final polycondensation kettle connected in sequence. The BS hydrogenation reaction system includes a BS hydrogenation reactor, a second gas-liquid separator, a light component removal tower, and a BDO refining tower connected in sequence;

[0124] Among them, the liquid phase outlet of the gas-liquid separator of the BM hydrogenation reaction system is connected to the BS hydrogenation reactor and the flash tower; the top of the BDO refining tower is connected to the maleic anhydride esterification absorption tower; control valves are installed on each connecting pipeline in the process system.

[0125] In the embodiment of the present invention, the tail gas pipeline of the maleic anhydride esterification absorption tower is connected to the maleic anhydride oxidation reactor and the incinerator. The gas phase outlet of the gas-liquid separator is connected to the BM hydrogenation reactor and the BS hydrogenation reactor. The gas phase outlet of the second gas-liquid separator is connected to the BS hydrogenation reactor. The bottom of the BDO refining tower is connected to the BS hydrogenation reactor and the post-treatment device.

[0126] Example 1

[0127] Referring to the attached Figure 1 and 2 , the maleic anhydride gas stream is esterified, absorbed, hydrogenated, and polycondensed to synthesize PBS. The specific process is as follows:

[0128] (1) Preparation of butanediol maleate (BM) by maleic anhydride esterification absorption:

[0129] The crude maleic anhydride gas stream at the outlet of the reactor for producing maleic anhydride by oxidizing benzene or butane is quenched to 120 °C and sent to the bottom of the esterification absorption tower. A BDO absorbent containing 1.0 wt% p-toluenesulfonic acid at a temperature of 45 °C is sprayed from the top of the tower. The maleic anhydride gas stream and the absorbent operate countercurrently. After 3-stage spray esterification absorption, an esterification absorption liquid (BDO solution with a BM concentration of 35.02 wt%) at a temperature of 75 °C is obtained. The tail gas after absorption is recycled back to the maleic anhydride oxidation reactor at 65 vol%, and the remaining tail gas is sent to the tail gas incinerator.

[0130] After detection, based on maleic anhydride in the maleic anhydride gas stream, in this process, the maleic anhydride esterification conversion rate is 100 mol%, and the BM selectivity is 99.28 mol%. The mass percentage content of each main substance in the esterification absorption liquid: BM is 35.02 wt%, BDO is 56.72 wt%, butanediol acrylate is 1.52 wt%, butanediol acetate is 2.26 wt%, and water is 3.56 wt%.

[0131] (2) Preparation of butylene succinate (BS) by hydrogenation of esterified absorption liquid:

[0132] The esterified absorption liquid with a BM concentration of 35.02 wt% obtained in step (1) was fed into the top of the slurry bed reactor, and hydrogen was fed countercurrently from the bottom. Under the action of Pd / AC powder catalyst in the slurry bed reactor, the absorption liquid temperature was 75°C and the weight hourly space velocity was 1.2 h -1 , hydrogen pressure of 1.0 MPa and hydrogen to ester molar ratio of 15, double bond hydrogenation of BM to prepare BS; Pd loading in Pd / AC catalyst is 1.0 wt%, and the amount of Pd / AC powder catalyst is 2.0 wt% of the weight of esterification absorption liquid;,

[0133] After the reaction material is separated into gas and liquid, the hydrogen is pressurized by a compressor and sent to the subsequent BS hydrogenation reactor. The liquid material is filtered and the BM hydrogenation catalyst is recovered and recycled. The filtrate after recovering the BM hydrogenation catalyst is divided into two paths, one is sent to the flash tower, and the other is sent to the subsequent BS hydrogenation reactor.

[0134] After testing, in step (2), based on BM in the absorption liquid, the BM double bond hydrogenation conversion rate was 100 mol%, and the BS selectivity was 99.95 mol%. The mass percentage of each main substance in the liquid product was: BS was 35.55 wt%, BDO was 56.05 wt%, butylene glycol propionate was 1.55 wt%, butylene glycol acetate was 2.30 wt%, and water was 3.60 wt%.

[0135] (3) BS polycondensation to synthesize polybutylene succinate (PBS):

[0136] After part of the filtrate in step (2) is sent to a flash tower, it is distilled under reduced pressure at a pressure of 90 kPa, a tower top temperature of 70° C. and a tower bottom temperature of 95° C., a small amount of water and other components are removed from the tower top, and a mixture containing BS and BDO is obtained at the bottom of the tower; a polymerization catalyst tetrabutyl titanate is added to the obtained flash tower bottom material, the mass of the tetrabutyl titanate is 0.01% of the mass of BS, and after stirring evenly, it is sent to a polymerization reactor, and a polymerization reaction is carried out at a temperature of 160° C., a pressure of 50 kPa and a residence time of 45 min; then, tetrabutyl titanate in an amount of 0.04% of the mass of BS is added to the polymerization product, and the product is sent to a prepolymerization reactor, and a prepolymerization reaction is carried out at a temperature of 200° C., a pressure of 10 kPa and a residence time of 30 min; finally, the prepolymerization product is sent to a final polymerization reactor, and a final polymerization reaction is carried out at a temperature of 240° C., a pressure of 50 Pa and a residence time of 120 min to obtain PBS; the final polymerization product PBS is cooled underwater and pelletized to obtain a PBS pellet product. In this process, the gaseous materials in the polymerization, prepolymerization and final polymerization processes are collected, condensed and sent to the subsequent BS hydrogenation reactor.

[0137] After detection, in this step, based on BS in the feed of the flash tower, the PBS yield is 70.65 wt%; the composition of the gas-phase condensate: BDO is 85.92 wt%, tetrahydrofuran is 4.32 wt%, water is 2.50 wt%, butylene succinate is 0.22 wt%, butylene propionate is 2.36 wt%, and butylene acetate is 3.58 wt%.

[0138] The process for producing PBS in this example further includes step (4):

[0139] (4) Hydrogenation of the BS ester group to prepare 1,4-butanediol (BDO):

[0140] Mix the hydrogen gas separated by gas-liquid separation in step (2) with fresh hydrogen gas and send it into the BS hydrogenation fixed-bed reactor filled with CuZnAlCe granular catalyst. At the same time, mix a part of the remaining filtrate in step (2) with the gas-phase condensate product collected in step (3) and send it into the BS hydrogenation fixed-bed reactor; at a feed BS concentration of 25.26 wt%, a weight hourly space velocity of 0.5 h -1 , a temperature of 260 °C, a pressure of 3.5 MPa, and a hydrogen-to-ester molar ratio of 50, carry out the hydrogenation reaction.

[0141] After the BS hydrogenation product is condensed and cooled, it is sent to the second gas-liquid separator for separation. The hydrogen gas is pressurized by a compressor and sent back to the BS hydrogenation reactor for recycling. The liquid is sent to the light-component removal tower. At a pressure of 30 kPa, a top temperature of 95 °C, and a bottom temperature of 130 °C, a small amount of light components containing tetrahydrofuran, ethanol, propanol, and water are removed from the top of the light-component removal tower. The bottom material is sent to the BDO refining tower; at a pressure of 10 kPa, a reflux ratio of 1.0, a top temperature of 130 °C, and a bottom temperature of 180 °C, BDO with a purity greater than 99.52 wt% is obtained at the top of the BDO refining tower and sent to the maleic anhydride esterification absorption tower as an absorbent. 95 wt% of the bottom product is sent back to the BS hydrogenation reactor, and the remaining 5 wt% is discharged as a heavy component and sent to the incinerator.

[0142] After detection, in step (4), based on BS in the hydrogenation feed, the BS hydrogenation conversion rate is 98.52 mol%, and the BDO selectivity is 95.86 mol%; the mass percentages of the main substances in the hydrogenation product are: BDO is 91.25 wt%, tetrahydrofuran is 2.25 wt%, ethanol is 0.90 wt%, propanol is 0.71 wt%, BS is 0.37 wt%, and water is 3.78 wt%; the BDO separation and refining yield is 96.50 wt%.

[0143] In this example, based on maleic anhydride in the maleic anhydride gas stream (maleic anhydride is the only raw material), from gaseous maleic anhydride to PBS product, the total PBS yield in the whole process is 76.60 wt%.

[0144] Example 2

[0145] Refer to the attached Figure 1 and 2 Maleic anhydride gas stream is esterified, absorbed and hydrogenated and polycondensed to synthesize polyester PBS, specifically including the following steps:

[0146] (1) Preparation of butanediol maleate (BM) by esterification and absorption of maleic anhydride:

[0147] The maleic anhydride gas stream at the outlet of the reactor for producing maleic anhydride by oxidizing benzene or butane is quenched to 90 °C and sent to the bottom of the esterification absorption tower. The absorbent, 1,4-butanediol (BDO) containing 0.5 wt% trifluoromethanesulfonic acid at a temperature of 30 °C, is sprayed from the top of the tower. The maleic anhydride gas stream and the absorbent operate in countercurrent. After two-stage spray esterification absorption, an absorption liquid (BDO solution with a BM concentration of 27.50 wt%) at a temperature of 65 °C is obtained. The tail gas after absorption is recycled back to the maleic anhydride oxidation reactor at 75 vol%, and the remaining tail gas is sent to the tail gas incinerator.

[0148] Based on maleic anhydride in the maleic anhydride gas stream, the esterification conversion rate of maleic anhydride is 100 mol%, and the selectivity of BM is 99.52 mol%. The mass percentage content of each main substance in the absorption liquid: BM is 27.50 wt%, BDO is 64.95 wt%, butyl acrylate is 1.50 wt%, butyl acetate is 2.21 wt%, and water is 3.72 wt%.

[0149] (2) Hydrogenation of the esterification absorption liquid to prepare butanediol succinate (BS):

[0150] The absorption liquid with a BM concentration of 27.50 wt% obtained in step (1) is sent into the reactor from the top, and hydrogen is sent in countercurrently from the bottom. Under the action of Pd / SiO 2 powder catalyst in a slurry bed reactor, at an absorption liquid temperature of 65 °C, a weight hourly space velocity of 2.0 h -1 , a hydrogen pressure of 1.5 MPa and a hydrogen to ester molar ratio of 10, the double bond of BM is hydrogenated to prepare BS; the Pd loading in the Pd / SiO 2 powder catalyst is 0.5 wt%, and the dosage of the Pd / SiO 2 powder catalyst is 4.0 wt% of the weight of the esterification absorption liquid;

[0151] The reaction product stream is subjected to gas-liquid separation. The hydrogen is pressurized by a compressor and sent to the subsequent BS hydrogenation reactor. The liquid material is filtered to recover the BM hydrogenation catalyst and recycled. The filtrate is divided into two paths, one path is sent to the flash tower, and the other path is sent to the subsequent BS hydrogenation reactor.

[0152] Based on BM in the absorbent liquid, the hydrogenation conversion rate of the BM double bond is 100 mol%, and the selectivity of BS is 100 mol%; the mass percentage content of each main substance in the liquid-phase product: BS is 27.78 wt%, BDO is 64.35 wt%, butylene glycol propionate is 1.53 wt%, butylene glycol acetate is 2.25 wt%, and water is 3.75 wt%.

[0153] (3) Polycondensation of BS to synthesize poly(butylene succinate) (PBS):

[0154] Part of the filtrate from step (2) is sent to a flash tower and subjected to vacuum distillation at a pressure of 70 kPa, a top temperature of 80 °C, and a bottom temperature of 100 °C. A small amount of components such as water are removed from the top of the tower, and the bottom of the tower is a mixture containing BS and BDO; tetrabutyl titanate equivalent to 0.02% of the mass of BS is added to the bottom material, and after stirring evenly, it is sent to an oligomerization reactor and subjected to oligomerization reaction at a temperature of 170 °C, a pressure of 30 kPa, and a residence time of 30 min; tetrabutyl titanate equivalent to 0.05% of the mass of BS is added to the oligomerization product and then sent to a prepolymerization reactor and subjected to prepolymerization reaction at a temperature of 190 °C, a pressure of 5 kPa, and a residence time of 45 min; then the prepolymerization product is sent to a final polymerization reactor and subjected to final polymerization reaction at a temperature of 230 °C, a pressure of 30 Pa, and a residence time of 90 min to obtain PBS; finally, the final polymerization product PBS is cooled and pelletized underwater to obtain PBS pellet products; the gas-phase products of the oligomerization, prepolymerization, and final polymerization reactions are collected, condensed, and sent to the subsequent BS hydrogenation reactor.

[0155] Based on BS in the flash tower feed, the yield of PBS is 69.82 wt%, and the yield of BDO is 16.56 wt%; the composition of the gas-phase condensate: BDO is 87.20 wt%, tetrahydrofuran is 3.95 wt%, water is 2.28 wt%, butylene glycol succinate is 0.45 wt%, butylene glycol propionate is 2.18 wt%, and butylene glycol acetate is 3.35 wt%; the performance indicators of the obtained PBS pellet products are: weight-average molecular weight of 176,000, melt index of 23.8 g / 10 min, acid value of 21.2 mol / t, melting point of 115 °C, and tensile strength of 46 MPa.

[0156] (4) Hydrogenation of the ester group of BS to prepare 1,4-butanediol (BDO):

[0157] The hydrogen gas separated from the gas-liquid separation in step (2) is mixed with fresh hydrogen gas and sent to a BS hydrogenation fixed-bed reactor filled with CuZnZrCe / SiO 2 particle catalyst. At the same time, the remaining part of the filtrate from step (2) is mixed with the gas-phase condensate of the polycondensation reaction collected in step (3) and sent to the BS hydrogenation reactor. At a feed BS concentration of 19.50 wt% and a weight hourly space velocity of 1.0 h -1、The hydrogenation reaction is carried out at a temperature of 240 °C, a pressure of 5.0 MPa and a molar ratio of hydrogen to ester of 35; after the hydrogenation product is condensed and cooled, it is sent to a gas-liquid separator for separation. The hydrogen is pressurized by a compressor and sent back to the BS hydrogenation reactor for recycling, and the liquid is sent to a light component removal column; at a pressure of 50 kPa, a top temperature of 110 °C and a bottom temperature of 140 °C, a small amount of light components containing tetrahydrofuran, ethanol, propanol and water are removed from the top of the light component removal column, and the bottom material is sent to a BDO purification column; at a pressure of 20 kPa, a reflux ratio of 2.0, a top temperature of 140 °C and a bottom temperature of 190 °C, BDO with a purity greater than 99.55 wt% is obtained at the top of the BDO purification column and sent to the maleic anhydride esterification absorption column as an absorbent, and 95 wt% of the bottom product is sent back to the BS hydrogenation reactor, and the remaining 5 wt% is discharged as a heavy component and sent to an incinerator.

[0158] Based on BS in the hydrogenation feed, the hydrogenation conversion rate of BS is 96.68 mol%, and the selectivity of BDO is 96.65 mol%; the mass percentages of the main substances in the hydrogenation product are: BDO is 91.75 wt%, tetrahydrofuran is 2.08 wt%, ethanol is 0.82 wt%, propanol is 0.68 wt%, BS is 0.96 wt%, and water is 3.55 wt%; the yield of BDO separation and purification is 96.32 wt%.

[0159] Based on maleic anhydride in the maleic anhydride gas stream (maleic anhydride is the only raw material), the total yield of PBS in the whole process from gaseous maleic anhydride to PBS product is 77.25 wt%.

[0160] Example 3

[0161] Refer to Appendix Figure 1 and 2 , esterify, absorb and hydrogenate and polycondense maleic anhydride gas stream to synthesize polyester PBS, which specifically includes the following steps:

[0162] (1) Prepare butanediol maleate (BM) by esterifying and absorbing maleic anhydride:

[0163] The maleic anhydride gas stream at the outlet of the reactor for oxidizing benzene or butane to produce maleic anhydride is quenched to 150 °C and sent to the bottom of the esterification absorption column. The absorbent, 1,4-butanediol (BDO) containing 1.5 wt% samarium trifluoromethanesulfonate at a temperature of 55 °C, is sprayed from the top of the column. The maleic anhydride gas stream and the absorbent operate countercurrently. After 4-stage spray esterification absorption, an absorption liquid (BDO solution with a BM concentration of 42.50 wt%) at a temperature of 95 °C is obtained. 50 vol% of the tail gas after absorption is recycled to the maleic anhydride oxidation reactor, and the remaining tail gas is sent to the tail gas incinerator.

[0164] Based on maleic anhydride in the maleic anhydride gas stream, the esterification conversion rate of maleic anhydride is 100 mol%, and the selectivity of BM is 99.15 mol%; the mass percentage content of each main substance in the absorption liquid: BM is 42.50 wt%, BDO is 49.75 wt%, butylene glycol acrylate is 1.60 wt%, butylene glycol acetate is 2.52 wt%, and water is 3.45 wt%.

[0165] (2) Preparation of butylene succinate (BS) by hydrogenation of the esterification absorption liquid:

[0166] The absorption liquid with a BM concentration of 42.50 wt% obtained in step (1) is fed into the reactor from the top, while hydrogen is fed in countercurrent from the bottom. Under the action of the Pd-Ru / HM powder catalyst in the slurry bed reactor, at an absorption liquid temperature of 95 °C, a weight hourly space velocity of 1.0 h -1 , a hydrogen pressure of 2.0 MPa and a hydrogen-to-ester molar ratio of 20, the double bond hydrogenation of BM is carried out to prepare BS; the Pd loading in the Pd-Ru / HM powder catalyst is 0.3 wt%, the Ru loading is 2.0 wt%, and the dosage of the Pd-Ru / HM powder catalyst is 2.0 wt% of the weight of the esterification absorption liquid;

[0167] The reaction product stream is subjected to gas-liquid separation. The hydrogen is pressurized by a compressor and then fed into the subsequent BS hydrogenation reactor. The liquid material is filtered to recover the BM hydrogenation catalyst and recycled. The filtrate is divided into two paths, one path is fed into the flash tower, and the other path is fed into the subsequent BS hydrogenation reactor.

[0168] Based on BM in the absorption liquid, the double bond hydrogenation conversion rate of BM is 100 mol%, and the selectivity of BS is 99.12 mol%; the mass percentage content of each main substance in the liquid phase product: BS is 42.62 wt%, BDO is 49.30 wt%, butylene glycol propionate is 1.65 wt%, butylene glycol acetate is 2.55 wt%, and water is 3.58 wt%.

[0169] (3) Polycondensation of BS to synthesize polybutylene succinate (PBS):

[0170] Part of the filtrate from step (2) is fed into a flash tower and subjected to vacuum distillation at a pressure of 90 kPa, a top temperature of 90 °C, and a bottom temperature of 110 °C. A small amount of components such as water are removed from the top of the tower, and the bottom is a mixture containing BS and BDO. Titanium dibutoxide equivalent to 0.01% of the mass of BS is added to the bottom material, and after stirring evenly, it is fed into an oligomerization reactor and subjected to oligomerization reaction at a temperature of 180 °C, a pressure of 50 kPa, and a residence time of 30 min. Titanium dibutoxide equivalent to 0.03% of the mass of BS is added to the oligomerization product and then fed into a prepolymerization reactor and subjected to prepolymerization reaction at a temperature of 210 °C, a pressure of 20 kPa, and a residence time of 45 min. Then the prepolymerization product is fed into a final polymerization reactor and subjected to final polymerization reaction at a temperature of 240 °C, a pressure of 70 Pa, and a residence time of 135 min to obtain PBS. Finally, the final polymerization product PBS is cooled and pelletized underwater to obtain PBS pellet products; the gas-phase products of the oligomerization, prepolymerization, and final polymerization reactions are collected, condensed, and fed into the subsequent BS hydrogenation reactor.

[0171] Based on BS in the flash tower feed, the PBS yield is 71.26 wt%, and the BDO yield is 12.85 wt%; the composition of the gas-phase condensate: BDO is 81.85 wt%, tetrahydrofuran is 6.90 wt%, water is 4.35 wt%, butylene succinate is 0.16 wt%, butylene propionate is 2.58 wt%, and butylene acetate is 3.88 wt%; the performance indicators of the obtained PBS pellet products are: weight-average molecular weight of 152,000, melt index of 30.6 g / 10 min, acid value of 25.7 mol / t, melting point of 113 °C, and tensile strength of 43 MPa.

[0172] (4) Hydrogenation of the BS ester group to prepare 1,4-butanediol (BDO):

[0173] The hydrogen gas separated from the gas-liquid separation in step (2) is mixed with fresh hydrogen gas and fed into a BS hydrogenation fixed-bed reactor filled with CuZnZrCe / SiO 2 -Si particle catalyst. At the same time, the remaining part of the filtrate from step (2) is mixed with the gas-phase condensate of the polycondensation reaction collected in step (3) and fed into the BS hydrogenation reactor. At a feed BS concentration of 30.25 wt% and a weight hourly space velocity of 0.5 h -1, the hydrogenation reaction is carried out at a temperature of 250 °C, a pressure of 4.0 MPa and a molar ratio of hydrogen to ester of 80; after the hydrogenation product is condensed and cooled, it is sent to a gas-liquid separator for separation. The hydrogen is pressurized by a compressor and sent to the BS hydrogenation reactor for recycling, and the liquid is sent to the light component removal column; at a pressure of 70 kPa, a top temperature of 115 °C and a bottom temperature of 150 °C, a small amount of light components containing tetrahydrofuran, ethanol, propanol and water are removed from the top of the light component removal column, and the bottom material is sent to the BDO purification column; at a pressure of 30 kPa, a reflux ratio of 3.0, a top temperature of 140 °C and a bottom temperature of 195 °C, BDO with a purity greater than 99.28 wt% is obtained from the top of the BDO purification column and sent to the maleic anhydride esterification absorption column as an absorbent, 95 wt% of the bottom product is sent back to the BS hydrogenation reactor, and the remaining 5 wt% is discharged as a heavy component and sent to the incinerator.

[0174] Based on BS in the hydrogenation feed, the hydrogenation conversion rate of BS is 97.88 mol%, and the selectivity of BDO is 96.36 mol%; the mass percentages of the main substances in the hydrogenation product are: BDO is 88.38 wt%, tetrahydrofuran is 4.82 wt%, ethanol is 1.02 wt%, propanol is 0.85 wt%, BS is 0.18 wt%, and water is 4.65 wt%; the yield of BDO separation and purification is 96.25 wt%.

[0175] Based on maleic anhydride in the maleic anhydride gas stream (maleic anhydride is the only raw material), the total yield of PBS in the whole process from gaseous maleic anhydride to PBS product is 75.92 wt%.

[0176] Comparative Example 1

[0177] The absorption and separation of maleic anhydride gas stream to produce maleic anhydride and the synthesis of PBS by ring-opening esterification and polycondensation of maleic anhydride via succinic anhydride specifically include the following steps:

[0178] (1) The absorption, desorption and separation and purification of maleic anhydride gas stream to produce maleic anhydride:

[0179] The maleic anhydride gas stream at the outlet of the maleic anhydride oxidation reactor is quenched to about 150 °C and sent to the absorption column to be sprayed with a 55 °C dibutyl phthalate (DBP) absorption solvent. After 3-stage countercurrent absorption, an absorption liquid containing about 32 wt% maleic anhydride at about 60 °C is obtained at the bottom of the absorption column; the gas at the top of the absorption column is washed with demineralized water, 70 vol% is sent back to the oxidation reactor to be mixed with fresh air and recycled, and the remaining tail gas is sent to the incinerator for incineration.

[0180] The absorption liquid is directly fed into the solvent desorption column, and the crude maleic anhydride containing light components is separated from the solvent containing heavy components through negative pressure flashing; the crude maleic anhydride is fed into the product refining system, and by-products such as acetic acid and acrylic acid are removed in the light component removal column at a pressure of 10 kPa, a top temperature of 130 °C, and a bottom temperature of 170 °C. It is rectified in the product column at a pressure of 5 kPa, a top temperature of 160 °C, a bottom temperature of 220 °C, and a reflux ratio of 0.5 to obtain a maleic anhydride product with a purity greater than 99.50%; the solvent containing heavy components enters the demineralized water washing column. After washing and extracting the acidic components, it is sent to a centrifuge to separate the aqueous phase and the solvent phase. The solvent DBP is sent back to the absorption column for recycling, and the aqueous phase is sent to the wastewater treatment system. The maleic anhydride yield in the maleic anhydride gas stream is 96.60 wt%, and the recovery rate of the absorption solvent DBP is 99.50%.

[0181] (2) Hydrogenation of maleic anhydride to produce succinic anhydride with co-production of γ-butyrolactone and separation and purification:

[0182] Using the maleic anhydride product obtained in step (1), a maleic anhydride solution with a concentration of about 25 wt% is prepared with γ-butyrolactone as the solvent. The maleic anhydride solution and hydrogen are fed into a fixed-bed hydrogenation reactor filled with 0.3 wt% Pd - 2.0 wt% Ni / Al 2 O 3 catalyst in a co-current manner from the top. The maleic anhydride hydrogenation reaction is carried out at a feed weight hourly space velocity of 1.0 h -1 , a reaction temperature of 130 °C, a hydrogen pressure of 3.0 MPa, and a hydrogen-to-ester molar ratio of 50 to produce a mixture of succinic anhydride and γ-butyrolactone. The reaction product is subjected to gas-liquid separation. The hydrogen is pressurized by a compressor and sent back to the maleic anhydride hydrogenation reactor for recycling. The liquid-phase material is fed into the γ-butyrolactone removal column. At a pressure of 30 kPa, a top temperature of 95 °C, and a bottom temperature of 165 °C, a γ-butyrolactone material containing tetrahydrofuran is obtained at the top of the column, and the crude succinic anhydride at the bottom is fed into the product column; it is rectified at a pressure of 5 kPa, a top temperature of 156 °C, and a bottom temperature of 205 °C. A succinic anhydride with a purity greater than 99.50 wt% is obtained at the top of the product column.

[0183] The maleic anhydride hydrogenation conversion rate is 100 mol%, the succinic anhydride selectivity is 45.82 mol%, and the γ-butyrolactone selectivity is 48.56 mol%; the separation and purification succinic anhydride yield is 98.62 wt%; based on the feed maleic anhydride, the weight yield of succinic anhydride is 46.11 wt%, and that of γ-butyrolactone is 42.61 wt%.

[0184] (3) Hydrogenation of γ-butyrolactone to produce 1,4-butanediol and separation and purification:

[0185] The γ-butyrolactone material containing tetrahydrofuran obtained in step (2) is added with tetrahydrofuran to prepare a solution with a γ-butyrolactone concentration of 40 wt%, and is fed into a fixed-bed reactor filled with CuZnAlMoO particle catalyst in a co-current manner with hydrogen. At a γ-butyrolactone weight hourly space velocity of 0.20 h-1 Perform a hydrogenation reaction at a temperature of 190 °C, a pressure of 5.5 MPa, and a hydrogen-to-ester molar ratio of 200; after the hydrogenation product is condensed and cooled, it is sent to a gas-liquid separator for separation. The hydrogen is sent to the γ-butyrolactone hydrogenation reactor for recycling after being pressurized by a compressor, and the liquid is sent to a light component removal column; at a pressure of 70 kPa, a top temperature of 90 °C, and a bottom temperature of 130 °C, the light components containing tetrahydrofuran and water are removed from the top of the light component removal column, and the bottom material is sent to a γ-butyrolactone removal column; at a pressure of 35 kPa, a top temperature of 105 °C, a bottom temperature of 175 °C, and a reflux ratio of 2.0, γ-butyrolactone is obtained from the top of the γ-butyrolactone removal column, and the bottom material is sent to a BDO purification column; at a pressure of 20 kPa, a reflux ratio of 3.0, a top temperature of 140 °C, and a kettle temperature of 195 °C, BDO with a purity greater than 99.28 wt% is obtained from the top of the BDO purification column.

[0186] The conversion rate of γ-butyrolactone is 96.80 mol%, the selectivity of BDO is 95.36 mol%, and the selectivity of tetrahydrofuran is 3.95%; the yield of separated and purified BDO is 97.55 mol%; the yield of BDO based on the feed γ-butyrolactone is 94.22 wt%.

[0187] (4) Esterification of succinic anhydride and 1,4-butanediol to synthesize butanediol succinate (BS):

[0188] Mix the succinic anhydride prepared in step (2) and the 1,4-butanediol prepared in step (3) at a molar ratio of 1.00:1.25, and add p-toluenesulfonic acid at 1.00% of the mass of succinic anhydride, and send it to a reaction kettle for esterification reaction at a temperature of 180 °C, a pressure of 0.2 MPa, and a residence time of 30 min; the reacted material is sent to a flash tank to remove water and tetrahydrofuran at 90 kPa to obtain an esterification product of BS and BDO.

[0189] The conversion rate of succinic anhydride is 100 mol%, the selectivity of BS based on succinic anhydride is 100 mol%; the by-product yield of tetrahydrofuran based on the feed BDO is 3.56 mol%; the yield of BS based on the total feed amount is 96.37 wt%.

[0190] (5) Polycondensation of BS to synthesize polybutylene succinate (PBS):

[0191] Add tetrabutyl titanate equivalent to 0.02% of the mass of BS to the esterification product obtained in step (5), stir evenly and then send it into the oligomerization reactor, and carry out oligomerization reaction at a temperature of 180 °C, a pressure of 50 kPa and a residence time of 30 min; add tetrabutyl titanate equivalent to 0.04% of the mass of BS to the oligomerization product and then send it into the prepolymerization reactor, and carry out prepolymerization reaction at a temperature of 200 °C, a pressure of 10 kPa and a residence time of 30 min; then send the prepolymerization product into the final polymerization reactor, and carry out final polymerization reaction at a temperature of 240 °C, a pressure of 50 Pa and a residence time of 120 min to obtain PBS; finally, the final polymerization product PBS is cooled and pelletized underwater to obtain PBS pellet products.

[0192] Based on BS in the feedstock, the PBS yield is 83.99 wt%, and the BDO yield is 4.50 wt%; the composition of the gas-phase condensate: BDO is 28.12 wt%, tetrahydrofuran is 22.52 wt%, water is 49.04 wt%, and butylene succinate is 0.32 wt%; the performance indicators of the obtained PBS pellet products are: weight-average molecular weight of 155,000, melt index of 25.2 g / 10 min, acid value of 32.6 mol / t, melting point of 114 °C, and tensile strength of 45 MPa.

[0193] Based on maleic anhydride in the maleic anhydride gas stream (maleic anhydride is the only raw material), the total yield of the whole process of producing PBS from the maleic anhydride gas stream is 71.25 wt%.

[0194] Comparative Example 2

[0195] Maleic anhydride is produced by absorption separation of the maleic anhydride gas stream, and PBS is synthesized by direct esterification polycondensation of maleic anhydride via succinic acid, which specifically includes the following steps:

[0196] (1) Absorption, analysis and separation and purification of maleic anhydride in the maleic anhydride gas stream to produce maleic anhydride: the same as step (1) of Comparative Example 1.

[0197] (2) Hydrogenation of maleic anhydride to produce succinic anhydride with co-production of γ-butyrolactone and separation and purification: the same as step (2) of Comparative Example 1.

[0198] (3) Hydrogenation of γ-butyrolactone to prepare 1,4-butanediol and separation and purification: the same as step (3) of Comparative Example 1.

[0199] (4) Hydrolysis and crystallization of succinic anhydride to prepare succinic acid:

[0200] According to the molar ratio of succinic anhydride: water = 1:10, send the succinic anhydride and softened water refined in step (2) into the hydrolysis reactor respectively, and carry out hydrolysis reaction at a reaction temperature of 80 °C and a residence time of 30 min; the hydrolysis reaction liquid is then concentrated, crystallized and recrystallized to obtain succinic acid with a purity greater than 98.5 wt%. Based on succinic anhydride, the total yield of succinic acid hydrolysis and crystallization is 114.50 wt%.

[0201] (5) Esterification synthesis of BS from succinic acid and 1,4-butanediol and dehydration:

[0202] Mix the succinic acid prepared in step (4) and the 1,4-butanediol prepared in step (3) according to a molar ratio of 1.00:1.25, add p-toluenesulfonic acid at 1.00% of the mass of succinic acid, and send it into a reaction kettle for esterification reaction at a temperature of 180 °C, a pressure of 0.2 MPa, and a residence time of 30 min; the reacted material is sent into a flash tank to remove water and tetrahydrofuran at 90 kPa to obtain an esterification product of BS and BDO. The conversion rate of succinic acid is 85.62 mol%, the selectivity of BS based on succinic acid is 100 mol%; the yield of BS based on the total feed amount is 70.58 wt%; the residual rate of succinic acid based on the fed succinic acid is 14.38 mol%; the by-product yield of tetrahydrofuran based on the fed BDO is 3.66 mol%, and the residual rate of BDO is 16.34 mol%.

[0203] (5) Polycondensation synthesis of polybutylene succinate (PBS) from BS: The same as step (5) of Comparative Example 1.

[0204] Based on the total amount of BS and succinic acid in the feed, the yield of PBS is 74.68 wt%; the composition of the gas-phase condensate: BDO is 14.52 wt%, tetrahydrofuran is 21.80 wt%, water is 62.48 wt%, and butylene succinate is 1.20 wt%; the performance indicators of the obtained PBS pellet product are: weight-average molecular weight of 136,000, melt index of 32.8 g / 10 min, acid value of 43.5 mol / t, melting point of 113 °C, and tensile strength of 42 MPa.

[0205] Based on maleic anhydride in the maleic anhydride gas stream (maleic anhydride is the only raw material), the total yield of the whole process of producing PBS from the maleic anhydride gas stream is 70.55 wt%.

[0206] Comparative Example 3

[0207] Absorption separation of maleic anhydride gas stream to produce maleic anhydride and synthesis of PBS by transesterification polycondensation of maleic anhydride via dimethyl succinate, specifically including the following steps:

[0208] (1) Absorption, analysis, and separation and purification of maleic anhydride gas stream to produce maleic anhydride: The same as step (1) of Comparative Example 1.

[0209] (2) Hydrogenation of maleic anhydride to produce succinic anhydride and co-production of γ-butyrolactone and separation and purification: The same as step (2) of Comparative Example 1.

[0210] (3) Hydrogenation of γ-butyrolactone to prepare 1,4-butanediol and separation and purification: The same as step (3) of Comparative Example 1.

[0211] (4) Esterification of succinic anhydride with methanol to prepare dimethyl succinate and separation and purification:

[0212] According to the molar ratio of succinic anhydride:methanol = 1.0:2.1, the succinic anhydride refined in step (2) and methanol are fed into an esterification reactor, and a pre-esterification reaction is carried out under normal pressure, at a temperature of 60 °C and a residence time of 30 min; then the monomethyl succinate material generated by the pre-esterification is fed into a catalytic distillation column filled with macroporous cation exchange resin Tulsimer T-63MP, and a deep esterification reaction is carried out under normal pressure, at 120 °C and a residence time of 60 min. The water and methanol generated by the deep esterification are distilled out from the top of the catalytic distillation column and fed into an alcohol-water separation column to recover methanol, and dimethyl succinate is taken out from the bottom of the catalytic distillation column. The conversion rate of succinic anhydride is 100 mol%, and the selectivity of dimethyl succinate is 98.05 mol%; based on succinic anhydride, the yield of dimethyl succinate is 140.20 wt%.

[0213] (5) Transesterification of dimethyl succinate and 1,4-butanediol to synthesize BS and methanol removal:

[0214] Mix the dimethyl succinate prepared in step (4) and the 1,4-butanediol prepared in step (3) according to the molar ratio of 1.00:1.25, add lithium acetate at 0.1% of the mass of dimethyl succinate, and feed it into a reactive distillation kettle to carry out a continuous transesterification reaction at 160 °C, normal pressure and a residence time of 45 min. Methanol and a small amount of tetrahydrofuran are distilled out from the top of the column, and a mixture of BS and BDO is continuously taken out from the bottom of the column. The conversion rate of dimethyl succinate is 99.5 mol%, and the selectivity of BS based on dimethyl succinate is 100 mol%; based on the total feed amount, the yield of BS is 84.30 wt%; based on the feed BDO, the by-product yield of tetrahydrofuran is 0.68 mol%.

[0215] (6) Polycondensation of BS to synthesize polybutylene succinate (PBS): The same as step (5) of Comparative Example 1.

[0216] Based on BS in the esterification product, the yield of PBS is 78.90 wt%, and the yield of BDO is 5.03 wt%; the composition of the gas-phase condensate: BDO is 23.76 wt%, tetrahydrofuran is 18.78 wt%, water is 4.71 wt%, methanol is 52.42 wt%, and butylene succinate is 0.33 wt%; the performance indexes of the obtained PBS pellet product are: weight-average molecular weight of 151,000, melt index of 26.4 g / 10 min, acid value of 21.5 mol / t, melting point of 114 °C, and tensile strength of 44 MPa.

[0217] Based on maleic anhydride in the maleic anhydride gas stream (maleic anhydride is the only raw material), the total yield of the whole process of producing PBS from the maleic anhydride gas stream is 70.02 wt%.

[0218] According to the process flow and yield results of producing PBS from maleic anhydride gas stream in the above Examples 1-3 and Comparative Examples 1-3, it can be seen that by using the method of the present invention, polyester PBS can be produced from maleic anhydride gas stream only through four steps of esterification absorption, BM hydrogenation, BS hydrogenation and BS polycondensation, and only one separation and purification of materials is required in the whole process. The total yield of PBS in the whole process is 75.92-77.25 wt%. When using the most simplified process flow in the prior art to produce polyester PBS from maleic anhydride gas stream, the succinic anhydride ring-opening esterification method requires five steps, and both the succinic acid direct esterification method and the succinic acid diester transesterification method require six steps. Moreover, separation and purification of materials are required for each step of these three methods. The total yield of PBS in the whole process is 70.02-71.25 wt%. Therefore, the present invention has developed a new process route for directly synthesizing fully biodegradable plastic PBS from gaseous crude maleic anhydride, significantly simplifying and shortening the process flow, and essentially and fundamentally eliminating the problems of by-products such as acrylic acid, maleic acid and fumaric acid in maleic anhydride production from blocking pipelines and equipment, making the production operation extremely simple.

[0219] 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 in the protection scope of the present invention.

Claims

1. A method for producing polybutylene succinate using gaseous crude maleic anhydride, characterized in that: The steps include: S1. Using 1,4-butanediol as an absorbent, in the presence of an acid catalyst, esterification and absorption of crude maleic anhydride gas stream to prepare butanediol butylene glycol ester; S2. The butanediol succinate prepared in step S1 is selectively hydrogenated by double bonds in the presence of a supported noble metal catalyst to prepare butanediol succinate; S3. subjecting part or all of the butylene succinate prepared in step S2 to a condensation reaction in the presence of a polymerization catalyst to obtain polybutylene succinate.

2. The method according to claim 1, characterized in that In step S1, the preparation of butylene glycol butene ester includes the following process: after the crude maleic anhydride gas flow is rapidly cooled to 55-180° C., it is sent to the bottom of an esterification absorption tower, and a BDO absorbent containing 0-10.0 wt% of a liquid acid catalyst at a temperature of 25-65° C. is sprayed from the top of the tower, and the maleic anhydride gas flow and the absorbent are operated in countercurrent; after 1-4 stages of spray esterification absorption, an esterification absorption liquid with a butylene glycol butene ester concentration of 10-50 wt% and a temperature of 52-120° C. is obtained; Preferably, the crude maleic anhydride gas stream is quenched to 60-150°C, the absorbent temperature is 25-60°C, the liquid acid catalyst concentration is 0.1-5.0wt%, the spray esterification absorption adopts 2-3 levels, the esterification absorption liquid butanediol ester concentration is controlled at 15-45wt%, and the esterification absorption liquid temperature is 55-90°C; More preferably, the crude maleic anhydride gas stream is quenched to 65-120°C, the absorbent temperature is 25-55°C, the liquid acid catalyst concentration is 0.5-2.0wt%, the spray esterification absorption adopts 2-3 levels, the esterification absorption liquid butanediol butylene glycol ester concentration is controlled to 20-40wt%, and the esterification absorption liquid temperature is controlled to 58-75°C.

3. The method according to claim 1 or 2, characterized in that: In step S1, the acid catalyst is selected from at least one of alkyl sulfonic acid, aryl sulfonic acid, halogenated carboxylic acid, halogenated sulfonic acid, amide, metal salt or metal imide salt thereof of alkyl sulfonic acid, aryl sulfonic acid or halogenated carboxylic acid, wherein the metal element in the metal salt or metal imide salt thereof is selected from at least one of Li, Mg, Ba, Al, Ga, In, Fe, Cu, Ag, Zn, Sb, Bi, Ti, Zr, Sn or rare earth elements; Preferably, the acid catalyst is selected from at least one of methanesulfonic acid, ethanesulfonic acid, propanesulfonic acid, benzenesulfonic acid, p-toluenesulfonic acid, β-naphthalenesulfonic acid, trifluoromethanesulfonic acid, pentafluoroethanesulfonic acid, trichloroacetic acid, trifluoroacetic acid, pentafluoropropionic acid, trifluoromethanesulfonamide, bistrifluoromethanesulfonimide, fluorosulfonic acid metal salt, methanesulfonic acid metal salt, p-toluenesulfonic acid metal salt, trifluoroacetic acid metal salt, trifluoromethanesulfonic acid metal salt or bistrifluoromethanesulfonimide metal salt.

4. The method according to claim 1, characterized in that: In step S2, the preparation of butylene succinate includes the following process: using a slurry bed reactor, under the action of a supported precious metal catalyst, the esterification absorption liquid after esterification absorption in step S1 is fed into a slurry bed reactor, and operated in parallel or countercurrent with hydrogen to perform a selective hydrogenation reaction to prepare butylene succinate; wherein the process conditions are: the concentration of butylene succinate in the feed esterification absorption liquid is 10-50wt%, the feed temperature is 52-150°C, and the feed weight hourly space velocity is 0.1-10.0h -1 , hydrogen pressure is 0.2-5.0 MPa and hydrogen-ester molar ratio is 2-50; Preferably, the concentration of butylene glycol butylene ester in the feed esterification absorption liquid is 15-45wt%, the feed temperature is 55-120°C, and the feed weight hourly space velocity is 0.5-5.0h -1 , hydrogen pressure is 0.5-3.0MPa and hydrogen-ester molar ratio is 5-30; More preferably, the concentration of butylene glycol butylene ester in the feed esterification absorption liquid is 20-40 wt%, the feed temperature is 60-90°C, and the feed weight hourly space velocity is 1.0-2.0 h -1 , the hydrogen pressure is 1.0-2.0 MPa and the hydrogen-ester molar ratio is 10-20.

5. The method according to claim 1 or 4, characterized in that: In step S2, the amount of the supported noble metal catalyst is 0.2-5.0wt% of the weight of the feed esterification absorption liquid, the carrier of the supported noble metal catalyst is selected from activated carbon, mesoporous carbon, carbon nanotubes, graphene, SiO2, Al2O3, TiO2 or ZrO2, or at least one of hydrogen-type zeolite molecular sieves HM, HHEU (oblique hair), Hβ, HZSM-5, HZSM-22, HMCM-22, HMCM-41, HMCM-48, HMCM-49, HMCM-56, SBA-15, ZEO-1, ZEO-3 or KIT-6, the noble metal is selected from one or more of Ru, Pd or Pt, the loading amount of Pd or Pt is 0.1-1.0wt%, and the loading amount of Ru is 0.2-3.0wt%; Preferably, the amount of the supported noble metal catalyst is 0.5-2.0wt% of the weight of the feed esterification absorption liquid, the carrier of the supported noble metal catalyst is selected from activated carbon, SiO2, Al2O3, SiO2-Al2O3, TiO2, ZrO2, or HM, Hβ, HZSM-5, HZSM-22, HMCM-22, HMCM-41, HMCM-48, SBA-15, ZEO-3, KIT-6, the noble metal is selected from one or two of Ru, Pd or Pt, the loading amount of Pd or Pt is 0.2-0.7wt%, and the loading amount of Ru is 0.5-3.0wt%; More preferably, the amount of the supported precious metal catalyst is 0.5-1.0 wt% of the weight of the feed esterification absorption liquid, the carrier of the supported precious metal catalyst is selected from activated carbon, SiO2, Al2O3, HM, HZSM-5, HMCM-22, HMCM-48 or SBA-15, the precious metal is selected from Pd, Pt or Pd-Ru, the Pd or Pt loading is 0.3-0.5 wt%, and the Ru loading is 1.0-2.0 wt%.

6. The method according to claim 1 or 4, characterized in that: In step S3, the process flow of preparing polybutylene succinate by polycondensation of butylene succinate specifically comprises the following steps: S31. The liquid material of the reaction product of step (2) after gas-liquid separation is filtered, and part or all of the liquid after separation is first sent to a flash tower for vacuum distillation, a small amount of water is removed from the top of the tower, and a mixture containing butylene succinate and BDO is obtained in the bottom of the tower; S32. After the obtained flash tower reactor material and 20% to 30% of the total amount of the polymerization catalyst are mixed uniformly, the mixture is fed into a polymerization reactor for reduced pressure polymerization reaction; S33. The oligomerization reaction product and the remaining 70 to 80% of the polymerization catalyst are mixed and fed into a prepolymerization reactor for negative pressure prepolymerization; S34. The prepolymerization product is sent to a final polymerization reactor to carry out a vacuum final polymerization reaction to obtain polybutylene succinate.

7. The method according to claim 1 or 6, characterized in that: In step S3, the polymerization catalyst is selected from at least one of titanium chloride, butoxytitanium chloride, tetraalkyl titanate, tetrahydroxy titanate, dialkyl titanate, tin chloride, organic acid stannous, tetraalkyl tin, fatty acid dialkyl tin, antimony chloride, fatty diol antimony, germanium chloride, tetrabutoxy germanium, and tetraalkyl germanium; preferably at least one of tetrabutyl titanate, tetrahydroxybutyl titanate, diethylene glycol titanate, dibutylene glycol titanate, and ethylene glycol antimony; The total amount of the polymerization catalyst is 0.01-1.00% of the mass of butylene succinate, preferably 0.02-0.50%, more preferably 0.05-0.10%; The operating conditions of the flash tower are: pressure 50-101 kPa, tower top temperature 50-100°C, tower bottom temperature 60-130°C; preferably, pressure 75-101 kPa, tower top temperature 60-100°C, tower bottom temperature 80-110°C; The operating conditions of the polymerization reaction are: temperature 150-200°C, pressure 20-70kPa and material residence time 20-60min; preferably, temperature 160-180°C, pressure 30-50kPa and material residence time 30-45min; The prepolymerization reaction operating conditions are: temperature 180-230°C, pressure 1-50kPa and material residence time 20-60min; preferably, temperature 190-210°C, pressure 5-30kPa and material residence time 30-45min; The final polymerization reaction operating conditions are: temperature 200-260° C., pressure 10-200 Pa and material residence time 60-180 min; preferably, temperature 220-240° C., pressure 50-90 Pa and material residence time 90-120 min.

8. The method according to claim 1, characterized in that In step S3, when a portion of butylene succinate is subjected to polycondensation reaction to prepare polybutylene succinate, the process further comprises step S4: in the presence of a hydrogenation catalyst, the remaining butylene succinate is subjected to hydrogenation reaction to prepare 1,4-butanediol, which is used as an absorbent in step (1); wherein the hydrogenation catalyst is a CuZnXY or CuZnXY / Z catalyst, wherein X is selected from at least one element selected from Al, Zr, Mo, W, Cr, Mn or Ni, and Y is selected from alkaline earth or / and One of the rare earth elements, the carrier Z is selected from a high specific surface carrier or a composite carrier of a high specific surface carrier and a high thermal conductivity carrier, the high specific surface carrier is selected from at least one of activated carbon, SiO2, γ-Al2O3, θ-Al2O3, SiO2-Al2O3, TiO2, molecular sieve, diatomaceous earth or bentonite, and the high thermal conductivity carrier is selected from at least one of graphene, Si, α-SiO2, α-Al2O3, BeO, SiC, Mo2C, BN or AlN.

9. The method according to claim 8, characterized in that In step S4, the remaining butylene succinate is subjected to a hydrogenation reaction to prepare 1,4-butanediol, which specifically includes the following process: S41. The remaining butylene succinate is fed into a hydrogenation reactor for hydrogenation reaction; S42. The hydrogenation reaction product is condensed and cooled and then sent to a gas-liquid separator for separation, and the liquid is sent to a light component removal tower; S43. Send the bottom material of the light component removal tower in step S42 to a BDO refining tower to obtain high-purity 1,4-butanediol, which is used as an absorbent in step S1; The hydrogenation reactor is a fixed bed reactor, preferably an isothermal fixed bed reactor or an adiabatic fixed bed reactor; The hydrogenation reaction conditions are: the mass concentration of 1,4-butanediol solution of butylene succinate is 8-40wt%, the feed weight hourly space velocity is 0.1-5.0h -1 , feed temperature 200-300°C, reaction pressure 2.0-8.0 MPa and hydrogen-ester molar ratio 20-100; preferably, feed butylene succinate concentration is 12-35 wt%, feed weight hourly space velocity is 0.3-3.0 h -1 , feed temperature 220-280°C, reaction pressure 2.5-6.5 MPa and hydrogen-ester molar ratio 30-80; more preferably, feed butylene succinate concentration 15-30 wt%, feed weight hourly space velocity 0.5-1.5 h -1 , feed temperature 240-260°C, reaction pressure 3.0-5.0MPa and hydrogen-ester molar ratio 40-70; The operating conditions of the light component removal tower are: tower top pressure 10-101 kPa, tower top temperature 75-130 ° C, tower bottom temperature 120-180 ° C, preferably tower top pressure 15-75 kPa, tower top temperature 85-120 ° C, tower bottom temperature 125-150 ° C, more preferably tower top pressure 20-50 kPa, tower top temperature 95-110 ° C, tower bottom temperature 130-140 ° C; The operating conditions of the BDO refining tower are: a tower top pressure of 2 to 50 kPa, a theoretical plate number of 20 to 50, a reflux ratio of 0.5 to 5, a tower top temperature of 120 to 150° C., and a tower bottom temperature of 160 to 200° C., preferably a tower top pressure of 5 to 40 kPa, a theoretical plate number of 25 to 45, a reflux ratio of 0.7 to 4.0, a tower top temperature of 125 to 145° C., and a tower bottom temperature of 165 to 195° C., and more preferably a tower top pressure of 10 to 30 kPa, a theoretical plate number of 30 to 40, a reflux ratio of 1.0 to 3.0, a tower top temperature of 130 to 140° C., and a tower bottom temperature of 170 to 190° C.

10. The method according to claim 9, characterized in that In step S3, the gaseous materials of the polymerization reaction, prepolymerization reaction and final polymerization reaction are collected, condensed and then sent to the hydrogenation reactor for hydrogenation of butanediol succinate in step (4).

11. A process system for producing polybutylene succinate using gaseous crude maleic anhydride, characterized in that: The process comprises a maleic anhydride oxidation reactor, a maleic anhydride esterification absorption tower, a BM hydrogenation reaction system and a BS hydrogenation reaction system which are connected in sequence. The BM hydrogenation reaction system comprises a BM hydrogenation reactor and a gas-liquid separator which are connected in sequence. The BS polymerization reaction system comprises a flash tower, a polycondensation kettle, a pre-polycondensation kettle and a final polycondensation kettle which are connected in sequence. The BS hydrogenation reaction system comprises a BS hydrogenation reactor, a second gas-liquid separator, a light component removal tower and a BDO refining tower which are connected in sequence. The liquid phase outlet of the gas-liquid separator of the BM hydrogenation reaction system is connected to the BS hydrogenation reactor and the flash tower. The top of the BDO refining tower is connected to the maleic anhydride esterification absorption tower. Control valves are installed on each connecting pipeline in the process system.

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