WO2023071865A1 - 一种梯级控制连续制备丙交酯的方法及系统 - Google Patents

一种梯级控制连续制备丙交酯的方法及系统 Download PDF

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WO2023071865A1
WO2023071865A1 PCT/CN2022/125901 CN2022125901W WO2023071865A1 WO 2023071865 A1 WO2023071865 A1 WO 2023071865A1 CN 2022125901 W CN2022125901 W CN 2022125901W WO 2023071865 A1 WO2023071865 A1 WO 2023071865A1
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depolymerization
reaction
phase material
liquid phase
lactic acid
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PCT/CN2022/125901
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English (en)
French (fr)
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孙启梅
周峰
李澜鹏
刘来伍
白富栋
张雷
白毓黎
王鹏翔
李秀峥
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中国石油化工股份有限公司
中石化（大连）石油化工研究院有限公司
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Priority to EP22885743.9A priority Critical patent/EP4410781A1/en
Publication of WO2023071865A1 publication Critical patent/WO2023071865A1/zh

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    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07D—HETEROCYCLIC COMPOUNDS
    • C07D319/00—Heterocyclic compounds containing six-membered rings having two oxygen atoms as the only ring hetero atoms
    • C07D319/10—1,4-Dioxanes; Hydrogenated 1,4-dioxanes
    • C07D319/12—1,4-Dioxanes; Hydrogenated 1,4-dioxanes not condensed with other rings

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  • the invention belongs to the technical field of biodegradable materials, and in particular relates to a method and system for continuously preparing lactide with step control.
  • PLA polylactic acid
  • PBS polyhydroxyalkanoate
  • PBS polybutylene succinate
  • PBSA polybutylene succinate-adipate
  • PBAT polybutylene terephthalate-adipate
  • polylactic acid is synthesized using a two-step method.
  • the specific two-step synthesis process of polylactic acid is as follows: the first step is to prepare lactide from lactic acid; the second step is to obtain polylactic acid by ring-opening polymerization of lactide, and the molecular weight of PLA obtained in this process can reach 100,000 to one million.
  • lactide is the key to the whole synthesis process, and the process barrier is relatively high. Usually, it needs to be prepared through polycondensation and depolymerization under a catalyst, high temperature, and high vacuum system, and this process is likely to cause lactide racemization.
  • m-lactide on the one hand, will affect the optical purity of lactide, and then affect the lactide ring-opening polymerization process, so that the obtained PLA has a low molecular weight; on the other hand, it will destroy the regularity of the PLA structure, It reduces the crystallinity and reduces the mechanical properties.
  • the crude lactide obtained by the depolymerization reaction needs to be purified and refined through processes such as solvent recrystallization, water extraction, rectification, and melt crystallization to reduce m-lactide in the product; and due to L-lactide and The physical and chemical properties of m-lactide are similar, and lactide itself has the characteristics of high freezing point, boiling point and heat sensitivity, which makes separation difficult, and the overall yield is low, only about 40%-60%, and the overall economic efficiency is also biased. Low. Therefore, the racemization of the lactide synthesis process is a key factor affecting the quality and yield of lactide, and it is also the focus and difficulty of the current domestic and foreign lactide technology research.
  • US5502215A discloses a method for refining and purifying lactide.
  • SnO is used as a catalyst, and lactic acid oligomers are added to a kettle-type three-necked flask for depolymerization to prepare crude D,L-lactide.
  • This process requires strong stirring And the reaction temperature is high (reaction temperature 220 DEG C), the purity of the obtained product is low, and the coking and carbonization of the bottom material is serious, and the high temperature reaction also aggravates the racemization of lactide simultaneously.
  • US6326458B discloses a continuous process for preparing lactide and lactide polymers.
  • the depolymerization reactor in the lactide preparation and depolymerization section of the process adopts a falling film tube evaporator, and the lactic acid oligomers evaporate from The lactide vapor is extracted from the bottom of the tube reactor, and the unreacted lactic acid oligomers are discharged from the lower outlet.
  • the reaction temperature required for the falling film reaction process in this process is relatively low, which can effectively reduce the probability of lactide racemization in the depolymerization process, but the lactide yield is low.
  • CN111153886A discloses a rapid and high-yield lactide synthesis method and device, using lactic acid single component or lactic acid to add catalyst two components through a mixer to enter the oligomer preparation system, increase residence time through bottom circulation, and synthesize oligomeric lactic acid,
  • the gas phase component passes through the rectification system to increase the yield of oligomeric lactic acid;
  • the oligomeric lactic acid passes through a purification device to remove unreacted lactic acid and water; Lactide, the heavy component enters the depolymerization reactor again through reflux, and the light component is purified and recovered to obtain the lactide product.
  • Lactide can be efficiently synthesized by using this device, and crude lactide with a yield of 94%-98% can be obtained within a short residence time of 0.5-5 minutes.
  • the heavy components after depolymerization in this invention are directly refluxed into the depolymerization reactor, which not only affects the stability of the depolymerization reaction, but also increases the coking of the reaction substrate on the surface of the reactor as the molecular weight of the heavy components increases and the catalyst accumulates.
  • the probability of carbonization increases the degree of racemization of lactide and affects the continuous and stable operation of the reaction.
  • the present invention provides a method and system for continuous preparation of lactide with cascade control.
  • the present invention realizes efficient depolymerization of lactic acid oligomers through stepwise control of multi-stage series depolymerization, reduces the degree of racemization of lactide and the probability of substrate coking and carbonization, and ensures continuous and stable operation of the depolymerization process and crude
  • the stability of the lactide product composition improves the overall depolymerization reaction rate, production efficiency and lactide yield.
  • the invention provides a method for continuously preparing lactide with step control, the method comprising the following steps:
  • the first depolymerization reaction unit includes a first depolymerization reactor and a first circulation tank, the lactic acid oligomer and the depolymerization catalyst react in the first depolymerization reactor, and after the reaction The obtained first liquid-phase material enters the first circulation tank.
  • reaction in step (2) is carried out in the presence of a protonated solvent.
  • the second depolymerization reaction unit includes at least one second depolymerization reactor and at least one second circulation tank, and the first liquid phase material and optional protonated solvent are used in the second depolymerization reaction reaction in the reactor, the liquid phase material after the reaction enters the second circulation tank, and when the molecular weight of the liquid phase material is below 6000, the liquid phase material in the second circulation tank is recycled back to the second solution
  • the polymerization reactor further reacts; when the molecular weight of the liquid phase material is greater than 6000, the liquid phase material in the second circulation tank is transported to the third depolymerization reaction unit for reaction.
  • the third depolymerization reaction unit includes at least one third depolymerization reactor and at least one third circulation tank, the second liquid phase material is reacted in the third depolymerization reactor, and the reacted The liquid phase material enters the third circulation tank, and when the molecular weight of the liquid phase material is below 10000, the liquid phase material in the third circulation tank is recycled back to the third depolymerization reactor for further reaction; When the molecular weight of the liquid phase material is greater than 10000, the liquid phase material in the third circulation tank is discharged.
  • the present invention also provides a cascade control system for continuously preparing lactide, the system comprising:
  • the first depolymerization reaction unit, the lactic acid oligomer and the depolymerization catalyst react in the first depolymerization reaction unit;
  • the liquid phase material from the first depolymerization reaction unit and the optional protonated solvent circulate and react in the second depolymerization reaction unit until the molecular weight of the liquid phase material is higher than 6000;
  • the liquid phase material with a molecular weight higher than 6000 from the second depolymerization reaction unit circulates and reacts in the third depolymerization reaction unit until the molecular weight of the liquid phase material is higher than 10000;
  • the depolymerization reaction unit of each stage mainly includes depolymerization reactor and circulation tank. According to the molecular weight of lactic acid oligomers in the circulation tank, the molecular weight of lactic acid oligomers is regulated to realize the production of lactic acid oligomers.
  • the high-efficiency depolymerization reduces the degree of lactide racemization and the probability of coking and carbonization of the substrate, and ensures the conversion rate of lactic acid oligomers and system stability during the entire reaction process.
  • the conversion rate of lactic acid oligomers in the whole process can reach more than 97.0%.
  • the degree of racemization and the probability of substrate coking and carbonization in the lactide synthesis process can be further reduced , Improve product quality.
  • the content of m-lactide in the crude lactide obtained from the three depolymerization reaction units can be controlled within 6.0%.
  • the degree of lactide racemization is reduced by more than 50%, and the conversion rate of lactic acid oligomers is increased by more than 10% compared with a single-pass wiped film evaporator or multiple depolymerization reactors connected in series without cascade control .
  • Fig. 1 is a structural schematic diagram of a specific embodiment of the system for cascade control continuous preparation of lactide according to the present invention.
  • the method for continuously preparing lactide by cascade control of the present invention comprises the following steps:
  • step (1) is to carry out preliminary reaction, realizes the continuous stable reaction of material by controlling condition, and realizes the stability of discharging;
  • Step (2) is to make material react further, through regulation and control
  • the reaction time is used to control the racemization of the product, further improving the reaction efficiency and conversion rate;
  • step (3) is a deep reaction of the material, by adjusting the temperature and vacuum degree, the material with a larger molecular weight can further participate in the depolymerization reaction, and at the same time, by regulating the material
  • the one-way reaction time is used to control the racemization of the product, so that the product can meet the product requirements and ensure the reaction conversion rate.
  • Steps (1) to (3) multi-stage reactions cooperate with each other and are interrelated to ensure the stability of crude lactide output quality and yield, and realize continuous and stable operation of the entire depolymerization process .
  • the molecular weight of the lactic acid oligomer in step (1) may be 800-3000, preferably 1200-2800.
  • the molecular weights of lactic acid oligomers all refer to weight average molecular weights.
  • the method may also include preparing lactic acid oligomers according to the following procedure: sequentially dehydrating and polycondensing L-lactic acid and/or D-lactic acid.
  • the dehydration process is mainly to remove free water in lactic acid, which can be in the form of normal pressure or reduced pressure.
  • the polycondensation conditions may include: a reaction temperature of 140-170° C., an absolute pressure of 1000-2000 Pa, and a reaction time of 0.5-4 hours.
  • the amount of the depolymerization catalyst in step (1) is 0.4%-3% of the mass of the lactic acid oligomer, more preferably 0.8%-2%.
  • the depolymerization catalyst described in step (1) is a tin catalyst, more preferably at least one of stannous octoate, SnCl 2 and SnO.
  • the reaction conditions in step (1) include: a reaction temperature of 180-200° C., an absolute pressure of 500-1500 Pa, and a reaction time of 3-8 minutes.
  • the first depolymerization reaction unit includes a first depolymerization reactor and a first circulation tank, and the lactic acid oligomer and the depolymerization catalyst react in the reactor, and the first liquid-phase material obtained after the reaction enters the first circulation tank.
  • the liquid level is maintained at 50%-70%, the pressure is maintained at 10kPa-atmospheric pressure, and the temperature is maintained at 160-200°C.
  • the probability of the lactic acid oligomers continuing to undergo intermolecular polymerization can be reduced, and coking and carbonization can be reduced.
  • the conversion rate of lactic acid oligomers in the reaction described in step (1) is controlled at Between 50%-60%. In this preferred situation, the probability of the lactic acid oligomers continuing to undergo intermolecular polymerization can be reduced, and coking and carbonization can be reduced.
  • the reaction in step (2) is carried out in the presence of a protonated solvent.
  • the first liquid phase material from the first depolymerization reaction unit is mixed with the protonated solvent and then sent to the second depolymerization reaction unit.
  • the degree of racemization in the lactide synthesis process and the probability of coking and carbonization of the substrate can be further reduced, and the product quality can be improved.
  • the protonated solvent may be at least one of diamines with not less than 12 carbon atoms and glycols with not less than 12 carbon atoms.
  • the melting temperature of the protonated solvent is 80-160°C, more preferably 100-160°C.
  • the protonated solvent is at least one of C12-C18 diamines and C12-C18 diols.
  • the protonated solvent is dodecanediamine, tetradecanediamine (also known as tetradecanediamine), hexadecanediamine (also known as hexadecanediamine) , tetradecanediol (also known as tetradecanediol) and hexadecanediol (also known as hexadecanediol) at least one.
  • the amount of the protonated solvent in the reaction of step (2), may be 0.1%-6% of the mass of the lactic acid oligomer, preferably 1%-3%.
  • the second depolymerization reaction unit includes at least one second depolymerization reactor and at least one second circulation tank, and the first liquid phase material and optional protonated solvent are Reaction in the second depolymerization reactor, the liquid phase material after the reaction enters the second circulation tank, when the molecular weight of the liquid phase material is below 6000, the liquid phase material in the second circulation tank Circulating back to the second depolymerization reactor for further reaction; when the molecular weight of the liquid phase material is greater than 6000, the liquid phase material in the second circulation tank is transported to the third depolymerization reaction unit for reaction.
  • the liquid phase material in the second circulation tank is recycled back to the second depolymerization reactor for further reaction; when the liquid phase material When the molecular weight is greater than 6000 and less than 10000, the liquid phase material in the second circulation tank is transported to the third depolymerization reaction unit for reaction.
  • the second depolymerization reaction unit may include a second depolymerization reactor and a second circulation tank, or may include two or more second depolymerization reactors and two or more The second circulation tank.
  • each second depolymerization reactor is respectively equipped with a second circulation tank, and each second depolymerization reactor and its correspondingly configured second circulation tank form a circulation reaction unit.
  • the second depolymerization reaction unit includes two second depolymerization reactors and two second circulation tanks, wherein the second depolymerization reactor A and the second circulation tank a constitute a circulation reaction Unit 2-1, the second depolymerization reactor B and the second circulation tank b constitute another circulation reaction unit 2-2, the liquid phase material (that is the first liquid phase material) from the first depolymerization reaction unit First enter the second depolymerization reactor A of the circulation reaction unit 2-1 for reaction, and the liquid phase material after the reaction enters the second circulation tank a, when the molecular weight of the liquid phase material is below 4500 (such as 3000-45000) When the liquid phase material in the second circulation tank a is recycled back to the second depolymerization reactor A for further reaction; when the molecular weight of the liquid phase material is greater than 4500 and less than 6000, the liquid phase material in the second circulation tank a is transported To react in the second depolymerization reactor B of the circulation reaction unit 2-2, the liquid phase material after
  • the second depolymerization reaction unit preferably includes a second depolymerization reactor and a second circulation tank.
  • the reaction conditions in the second depolymerization reactor may include: a reaction temperature of 200-220° C., an absolute pressure of 400-1000 Pa, and a one-way reaction time of 2-5 minutes.
  • the feeding amount of the lactic acid oligomer is 3-5 times of the actual reaction amount.
  • the residence time of the lactic acid oligomers on the surface of the second depolymerization reactor can be reduced, the occurrence of polymerization can be suppressed, the yield can be improved, and the product quality can be guaranteed.
  • the liquid level is maintained at 50%-70%, the pressure is maintained at 10kPa-atmospheric pressure, and the temperature is maintained at 160-200°C.
  • the conversion rate of lactic acid oligomers in the depolymerization reaction process of step (2) can reach more than 70%.
  • the third depolymerization reaction unit includes at least one third depolymerization reactor and at least one third circulation tank, and the second liquid phase material is reaction in the tank, the liquid phase material after the reaction enters the third circulation tank, and when the molecular weight of the liquid phase material is below 10000, the liquid phase material in the third circulation tank is recycled back to the third solution
  • the polymerization reactor further reacts; when the molecular weight of the liquid phase material is greater than 10000, the liquid phase material in the third circulation tank is discharged.
  • the third depolymerization reaction unit may include a third depolymerization reactor and a third circulation tank, or may include two or more third depolymerization reactors and two or more The third circulation tank.
  • the third depolymerization reaction unit includes two or more third depolymerization reactors and two or more third circulation tanks, each third depolymerization reactor is respectively equipped with a third circulation tank, and each third depolymerization reactor and its correspondingly configured third circulation tank form a circulation reaction unit.
  • the third depolymerization reaction unit includes two third depolymerization reactors and two third circulation tanks, wherein the third depolymerization reactor C and the third circulation tank c constitute a circulation reaction Unit 3-1, the third depolymerization reactor D and the third circulation tank d form another circulation reaction unit 3-2, and the liquid phase material (that is, the second liquid phase material) from the second depolymerization reaction unit First enter the third depolymerization reactor C of the circulation reaction unit 3-1 to react, and the liquid phase material after the reaction enters the third circulation tank c, when the molecular weight of the liquid phase material is below 8000 (such as greater than 6000 and less than equal to 8000), the liquid phase material in the third circulation tank c is recycled back to the third depolymerization reactor C for further reaction; when the molecular weight of the liquid phase material is greater than 8000 and less than 10000, the liquid phase material in the third circulation tank c is The phase material is transported to the third depolymerization reactor D of the circulation
  • the third circulation tank d When the molecular weight of the liquid phase material is below 10000, the third circulation tank d The liquid phase material in the tank d is recycled to the third depolymerization reactor D for further reaction; when the molecular weight of the liquid phase material is greater than 10000, the liquid phase material in the third circulation tank d is discharged from the system.
  • the third depolymerization reaction unit preferably includes a third depolymerization reactor and a third circulation tank.
  • the reaction conditions in the third depolymerization reactor may include: a reaction temperature of 220-240° C., an absolute pressure of 200-800 Pa, and a one-way reaction time of 1-4 minutes.
  • the feeding amount of the lactic acid oligomer is 4-6 times of the actual reaction amount.
  • the residence time of the lactic acid oligomers on the surface of the third depolymerization reactor can be reduced, the polymerization reaction can be suppressed, the yield can be increased, and the product quality can be guaranteed.
  • the liquid level is maintained at 10%-30%, the pressure is maintained at 10kPa-atmospheric pressure, and the temperature is maintained at 160-200°C.
  • the liquid phase material with a molecular weight greater than 10,000 discharged from the third depolymerization reaction unit is a lactic acid high polymer, which can be hydrolyzed to return to lactic acid.
  • the conversion rate of lactic acid oligomers in the depolymerization reaction process of step (3) can reach more than 70%.
  • the conversion rate of lactic acid oligomers in the whole cascade controlled depolymerization reaction process can reach more than 97.0%.
  • the reactors in the first depolymerization reaction unit, the second depolymerization reaction unit and the third depolymerization reaction unit are each wiped film depolymerization reactors, more preferably thin film evaporators, molecular distillation evaporators or other stirred film evaporators.
  • the gas-phase crude lactide generated by the three depolymerization reaction units is discharged from the top of each depolymerization reactor, and the mass composition of the obtained crude lactide is: L-lactide content 82%-92% %, the content of m-lactide is 1.0%-6%, the content of L-lactic acid is 0.5%-6%, and the content of dimer and trimer is 1.5%-6%.
  • the gas-phase crude lactide produced by the three depolymerization reaction units enters the separation and purification process, and can be directly refined through rectification or other purification and purification processes, and the quality of the obtained product meets the requirements of polymerization-grade lactide monomer need.
  • the cascade control system of the present invention for continuously preparing lactide comprises:
  • the first depolymerization reaction unit, the lactic acid oligomer and the depolymerization catalyst react in the first depolymerization reaction unit;
  • the liquid phase material from the first depolymerization reaction unit and the optional protonated solvent circulate and react in the second depolymerization reaction unit until the molecular weight of the liquid phase material is higher than 6000;
  • the liquid phase material with a molecular weight higher than 6000 from the second depolymerization reaction unit circulates and reacts in the third depolymerization reaction unit until the molecular weight of the liquid phase material is higher than 10000;
  • the first depolymerization reaction unit includes a first depolymerization reactor and a first circulation tank, and the lactic acid oligomer and the depolymerization catalyst react in the first depolymerization reactor, The liquid phase material obtained after the reaction enters the first circulation tank.
  • the second depolymerization reaction unit includes at least one second depolymerization reactor and at least one second circulation tank, the liquid phase material from the first depolymerization reaction unit and optional protonated solvent React in the second depolymerization reactor, the liquid phase material after the reaction enters the described second circulation tank, when the molecular weight of the liquid phase material is below 6000, the liquid phase in the second circulation tank The material is recycled back to the second depolymerization reactor for further reaction; when the molecular weight of the liquid phase material is greater than 6000, the liquid phase material in the second circulation tank is transported to the third depolymerization reaction unit for reaction .
  • the third depolymerization reaction unit includes at least one third depolymerization reactor and at least one third circulation tank, and the liquid phase material with a molecular weight higher than 6000 from the second depolymerization reaction unit is in the The reaction in the third depolymerization reactor, the liquid phase material after the reaction enters the third circulation tank, when the molecular weight of the liquid phase material is below 10000, the liquid phase material in the third circulation tank is circulated return to the third depolymerization reactor for further reaction; when the molecular weight of the liquid phase material is greater than 10,000, the liquid phase material in the third circulation tank is discharged from the system.
  • the reactors in the first depolymerization reaction unit, the second depolymerization reaction unit, and the third depolymerization reaction unit may each be a wiped film depolymerization reactor, preferably a thin film evaporator, molecular distillation evaporator or other stirred film evaporator.
  • the system for continuously preparing lactide with step control includes a first depolymerization reaction unit, a second depolymerization reaction unit and a third depolymerization reaction unit, wherein , the first depolymerization reaction unit includes a first depolymerization reactor I and a first circulation tank IV, the second depolymerization reaction unit includes a second depolymerization reactor II and a second circulation tank V, and the second depolymerization reaction unit includes a second depolymerization reactor II and a second circulation tank V.
  • the second circulation tank V is equipped with a first molecular weight detection and control component VII
  • the third depolymerization reaction unit includes a third depolymerization reactor III and a third circulation tank VI
  • the third circulation tank VI is equipped with a second molecular weight Detection and Control Module VII.
  • the specific operation process of the system is: firstly, the lactic acid oligomer 01 is transported to the first depolymerization reactor I for reaction, and the unreacted lactic acid oligomer is discharged into the first circulation tank IV, and the pressure is controlled at 10kPa-normal pressure, the temperature is 160-200°C, and when the liquid level in the tank reaches 50%-70%, the discharge material 05 of the first circulation tank is transported to the second depolymerization reactor II alone or further mixed with the protonation solvent 10 Reaction, unreacted lactic acid oligomers enter the second circulation tank V, control the pressure at 10kPa-atmospheric pressure, and the temperature at 160-200°C, wait until the liquid level in the tank reaches 50%-70%, and detect the second circulation tank V
  • the molecular weight of the middle lactic acid oligomer if the molecular weight is not higher than 6000 (preferably 3000-6000), it will be returned to the second depolymerization reactor II as the second circulating tank recycle material 06
  • the experimental methods in the following examples are conventional methods in the art unless otherwise specified.
  • the experimental materials used in the following examples can be purchased from biochemical reagent stores unless otherwise specified.
  • the lactic acid used in the embodiment of the present invention is heat-resistant grade L-lactic acid with a lactic acid content of 88% or more, and its optical purity is not lower than 99.0%.
  • the present invention uses a Malvern Viscotek OMNISEC GPC/SEC gel chromatograph to analyze the molecular weight of lactic acid oligomers.
  • PS polystyrene
  • the chromatographic column model is T3000
  • the size is 300mmL ⁇ 8.0mm
  • the column temperature is 40°C
  • the flow rate is 1.0mL/min
  • the sample concentration is 2-5mg/mL
  • single time The injection volume is 500 ⁇ L.
  • the present invention uses Agilent high performance liquid chromatography to analyze the chemical purity of lactide, L-lactic acid and dimer, trimer content, ultraviolet detector, adopts phosphoric acid and acetonitrile as mobile phase, and the chromatographic column model is ZORBAX SB-Aq , the length of the column is 250mm, the inner diameter of the column is 4.6mm, and the particle size of the inner filler is 5 ⁇ m. Detection wavelength: 200nm, column temperature: 40°C, flow rate: 1mL/min, injection volume: 5 ⁇ L.
  • the Agilent gas chromatograph of the present invention analyzes the lactide content of different optical isomers, selects the CYCLOSIL-B model chromatographic column, the temperature of the gasification chamber is 250°C, the temperature of the detector is 280°C, a hydrogen flame ion detector, and a column temperature program Raise the initial temperature to 100°C, keep it for 5min, raise it to 140°C at a rate of 4°C/min, keep it for 7min, raise it to 200°C at a rate of 8°C/min, keep it for 20min, the flow rate of carrier gas N2 is 1.4mL/min, hydrogen The flow rate is 30mL/min, the air flow rate is 400mL/min, and the injection volume is 0.5 ⁇ L.
  • m 0 is the quality of crude lactide
  • y 0 is the purity of L-lactide in crude lactide
  • m is the quality of lactide product
  • M is the lactide that a certain amount of lactic acid oligomers can theoretically convert into The quality of the ester, that is, the quality of the lactic acid oligomers.
  • ⁇ pure substance represents the specific rotation of pure lactide
  • ⁇ measured sample represents the specific rotation of the measured substance
  • the embodiment of the present invention is carried out according to the device and process shown in Figure 1.
  • the lactic acid oligomer 01 is transported to the first scraped film depolymerization reactor I for reaction, and the unreacted lactic acid oligomer is discharged into the circulation tank IV, and the control
  • the pressure is 10kPa-atmospheric pressure
  • the temperature is 160-200°C
  • the first circulating tank discharge material 05 is transported to the second scraping film solution alone or further mixed with the protonated solvent 10
  • the reaction is carried out in the polymerization reactor II, the unreacted lactic acid oligomer enters the circulation tank V, the pressure is controlled at 10kPa-atmospheric pressure, the temperature is 160-200°C, and the liquid level in the tank reaches 50%-70%, the detection cycle
  • the molecular weight of the oligomer in tank V if the molecular weight is lower than 6000, preferably 3000-6000, it will be returned to the second depol
  • Material 07 is transported to the third scraped film depolymerization reactor III for reaction, and the unreacted lactic acid oligomer enters the circulation tank VI, the pressure is adjusted at 10kPa-atmospheric pressure, the temperature is 160-200°C, and the liquid level in it reaches 10 %-30%, detect the molecular weight of oligomers in the circulation tank VI, if the molecular weight is lower than 10,000, it will be returned to the third depolymerization reactor III as the circulating material 08 of the third circulation tank, if the molecular weight is greater than 10,000, it will be used as lactic acid Polymer 09 is discharged from the reaction system, and the discharged material is lactic acid polymer 10, which can be hydrolyzed to form lactic acid for reuse.
  • the gas-phase crude lactide generated by the three depolymerization reactors is discharged from the top of each depolymerization reactor to obtain the crude lactide product 04, which enters the separation and purification section.
  • Lactic acid to remove free water take 4000g L-lactic acid (the content of lactic acid is about 88.0%, and the optical purity is 99.2%), add it to the reactor with a stirring system, and use a vacuum circulating pump to maintain the pressure of the system at 50kPa Start heating under vacuum, gradually heat to 110-120°C, and dehydrate for 2 hours. At this time, the free water in the reaction system is slowly distilled out of the reaction system.
  • lactic acid oligomer Take 3000g of lactic acid oligomer, add 30g of stannous octoate catalyst, mix well and transport to the first scraped film depolymerization reactor, control the depolymerization reaction conditions as follows: vacuum degree 600Pa, reaction temperature 190 °C, one-way reaction time 4min , the unreacted lactic acid oligomers are discharged into the first circulation tank, the temperature is controlled at 180° C., and the pressure is controlled at 20 kPa. After the liquid level increases to 60%, the molecular weight of the lactic acid oligomers is measured to be 3216. The conversion rate of lactic acid oligomers in this process was 54.6%.
  • the lactic acid oligomers in the first circulation tank are transported to the second wiped-film depolymerization reactor, the reaction temperature is controlled at 210°C, the vacuum degree is 400Pa, the single-pass reaction time is 3min, and the feed amount of lactic acid oligomers is 1% of the actual reaction amount 4 times, the unreacted lactic acid oligomers are discharged into the second circulation tank, the liquid level is maintained at 50%, the pressure is maintained at 10kPa, and the temperature is maintained at 180°C. After testing, the molecular weight of the lactic acid oligomers in the second circulation tank is 4915.
  • the lactic acid oligomer in the second circulation tank was transported to the third depolymerization reactor, the reaction temperature was 230°C, the vacuum degree was 300Pa, the single-pass reaction time was 2min, and the feeding amount of the lactic acid oligomer was 5 times of the actual reaction amount.
  • the reacted lactic acid oligomers are discharged into the third circulation tank, and the temperature is controlled at 180°C and the pressure is controlled at 20kPa.
  • the molecular weight of the lactic acid oligomers is detected to be 7864, and the lactic acid oligomers are recycled to the third depolymerization reactor Continue to participate in the reaction, and discharge the system after the molecular weight of the lactic acid oligomer increases to 10,000.
  • the conversion rate of lactic acid oligomers in this process was 71.1%.
  • composition of the crude lactide obtained in the whole process is: 89.7% of L-lactide, 3.2% of m-lactide, 2.5% of L-lactic acid, and 3.4% of lactic acid dimer and trimer.
  • the conversion rate of lactic acid oligomers in the whole step cycle depolymerization process can reach 98.3%.
  • the above-mentioned crude lactide product is purified by a two-stage rectification system, and the chemical purity and optical purity of the refined product can meet the requirements of the polymerization-grade lactide monomer.
  • lactic acid oligomer Take 3000g of lactic acid oligomer, add 30g of stannous octoate catalyst, mix well and transport to the first scraped film depolymerization reactor, control the depolymerization reaction conditions: vacuum degree 1500Pa, reaction temperature 230 °C, one-way reaction time 8min , the unreacted lactic acid oligomers are discharged into the first circulation tank, the temperature is controlled at 200° C., and the pressure is controlled at 20 kPa. After the liquid level increases to 60%, the measured molecular weight of the lactic acid oligomers is 3713. The conversion rate of lactic acid oligomers in this process was 50.8%.
  • the lactic acid oligomers in the first circulation tank are transported to the second scraped film depolymerization reactor, the reaction temperature is controlled at 220°C, the vacuum degree is 1000Pa, the single-pass reaction time is 5min, and the feed amount of lactic acid oligomers is the actual reaction amount 3 times, the unreacted lactic acid oligomers are discharged into the second circulation tank, the liquid level is maintained at 50%, the pressure is maintained at 20kPa, and the temperature is maintained at 180°C. After testing, the molecular weight of the lactic acid oligomers in the second circulation tank is 6032. Transported to the third depolymerization reactor. The conversion rate of lactic acid oligomers in this process was 71.2%.
  • the lactic acid oligomer in the second circulation tank was transported to the third depolymerization reactor, the reaction temperature was 240°C, the vacuum degree was 800Pa, the single-pass reaction time was 4min, and the feeding amount of the lactic acid oligomer was 4 times of the actual reaction amount.
  • the reacted lactic acid oligomers are discharged into the third circulation tank, the temperature is controlled at 180°C, and the pressure is controlled at 20kPa.
  • the molecular weight of the lactic acid oligomers is detected to be 7357, and the lactic acid oligomers are recycled to the third depolymerization reactor Continue to participate in the reaction, and discharge the system after the molecular weight of the lactic acid oligomer increases to 10,000.
  • the conversion rate of lactic acid oligomers in this process was 70.3%.
  • composition of the crude lactide obtained in the whole process is: 84.6% of L-lactide, 5.9% of m-lactide, 1.5% of L-lactic acid, and 4.0% of lactic acid dimer and trimer.
  • the conversion rate of lactic acid oligomers in the whole step cycle depolymerization process can reach 97.1%.
  • the above-mentioned crude lactide product is purified by a two-stage rectification system, and the chemical purity and optical purity of the refined product can meet the requirements of the polymerization-grade lactide monomer.
  • lactic acid oligomer Take 3000g of lactic acid oligomer, add 30g of stannous octoate catalyst, mix well and transport to the first scraped film depolymerization reactor, control the depolymerization reaction conditions as follows: vacuum degree 500Pa, reaction temperature 180 °C, one-way reaction time 3min , the feeding amount of lactic acid oligomers is 4 times of the actual reaction amount; the unreacted lactic acid oligomers are discharged into the first circulation tank, the temperature is controlled at 180°C, and the pressure is controlled at 20kPa. After the liquid level increases to 60%, The measured molecular weight of the lactic acid oligomer was 3461. The conversion rate of lactic acid oligomers in this process was 51.8%.
  • the lactic acid oligomers in the first circulation tank are transported to the second wiped-film depolymerization reactor, the reaction temperature is controlled at 200°C, the vacuum degree is 400Pa, the single-pass reaction time is 2min, and the feed amount of lactic acid oligomers is the actual reaction amount 5 times, the unreacted lactic acid oligomers are discharged into the second circulation tank, the liquid level is maintained at 50%, the pressure is maintained at 20kPa, and the temperature is maintained at 180°C.
  • the molecular weight of the lactic acid oligomers in the second circulation tank is 4456.
  • the lactic acid oligomer in the second circulation tank was transported to the third depolymerization reactor, the reaction temperature was 220°C, the vacuum degree was 200Pa, the single-pass reaction time was 1min, and the feeding amount of the lactic acid oligomer was 6 times of the actual reaction amount.
  • the reacted lactic acid oligomers are discharged into the third circulation tank, and the temperature is controlled at 180°C and the pressure is controlled at 20kPa.
  • the molecular weight of the lactic acid oligomers is detected to be 7071, and the lactic acid oligomers are recycled to the third depolymerization reactor Continue to participate in the reaction, and discharge the system after the molecular weight of the lactic acid oligomer increases to 10,000.
  • the conversion rate of lactic acid oligomers in this process was 72.4%.
  • composition of crude lactide obtained in the whole process is: L-lactide 89.2%, m-lactide 3.2%, L-lactic acid 2.9%, lactic acid dimer and trimer content 4.7%, the whole cascade cycle
  • the conversion rate of lactic acid oligomers in the depolymerization process can reach 97.6%.
  • the above-mentioned crude lactide product is purified by a two-stage rectification system, and the chemical purity and optical purity of the refined product can meet the requirements of the polymerization-grade lactide monomer.
  • Lactide was prepared according to the method of Example 1, the difference being that the stannous chloride catalyst of the same quality was used.
  • the molecular weight of the lactic acid oligomer at the outlet of the first circulation tank was detected to be 3527, and the conversion rate of the lactic acid oligomer was 53.8%.
  • the molecular weight of the lactic acid oligomer in the second circulation tank is detected to be 5319, and it is directly circulated back to the second depolymerization reactor to continue to participate in the reaction.
  • the molecular weight of the exported lactic acid oligomer is greater than 6000, it is transported to the third depolymerization reactor, and the conversion rate of the lactic acid oligomer in this process is 71.7%.
  • the molecular weight of the lactic acid oligomer at the outlet of the third circulation tank was measured to be 8213, and it was circulated back to the third depolymerization reactor to continue to participate in the reaction.
  • the molecular weight of the lactic acid oligomer increased to 10000, it was discharged system.
  • the conversion rate of lactic acid oligomers in this process was 70.8%.
  • the composition of crude lactide obtained in the whole process is: L-lactide 88.2%, m-lactide 3.9%, L-lactic acid 2.7%, lactic acid dimer and trimer content 4.4%.
  • the conversion rate of lactic acid oligomers in the whole cascade cycle depolymerization process can reach 97.8%.
  • the above-mentioned crude lactide product is purified by a two-stage rectification system, and the chemical purity and optical purity of the refined product can meet the requirements of the polymerization-grade lactide monomer.
  • Lactide was prepared according to the method of Example 1, except that a depolymerization evaporator in the form of molecular distillation was used. After the reaction in the first depolymerization reactor, the molecular weight of the lactic acid oligomer at the outlet of the first circulating tank was detected to be 3009, and the conversion rate of the lactic acid oligomer was 55.3%. After the lactic acid oligomer enters the second depolymerization reactor for reaction, the conversion rate of the lactic acid oligomer in this process is 73.5%. After the reaction in the third depolymerization reactor, the conversion rate of lactic acid oligomers in this process was 72.7%.
  • composition of crude lactide obtained in the whole process is: 91.3% of L-lactide, 2.2% of m-lactide, 2.9% of L-lactic acid, and 3.0% of lactic acid dimer and trimer.
  • the conversion rate of lactic acid oligomers in the whole step cycle depolymerization process can reach 98.7%.
  • the above-mentioned crude lactide product is purified by a two-stage rectification system, and the chemical purity and optical purity of the refined product can meet the requirements of the polymerization-grade lactide monomer.
  • Lactide is prepared according to the method of Example 1, the difference is that: before the unreacted lactic acid oligomers in the first depolymerization reactor are sent into the second depolymerization reactor, 1.0% tetradecane binary alcohol.
  • the conversion rate of lactic acid oligomers in the second depolymerization reactor was 74.4%.
  • the conversion rate of lactic acid oligomers in the third depolymerization reactor was 72.4%.
  • the composition of crude lactide obtained in the whole process is: 91.3% of L-lactide, 1.8% of m-lactide, 2.1% of L-lactic acid, and 3.3% of lactic acid dimer and trimer.
  • the conversion rate of lactic acid oligomers in the whole step cycle depolymerization process can reach 98.6%.
  • the degree of racemization of the crude lactide product obtained in each step is low, and the material handling capacity in the whole process is high.
  • the above-mentioned crude lactide product is purified by a two-stage rectification system, and the chemical purity and optical purity of the refined product can meet the requirements of the polymerization-grade lactide monomer.
  • Lactide is prepared according to the method of Example 1, the difference is that: before the unreacted lactic acid oligomers in the first depolymerization reactor are sent into the second depolymerization reactor, 1.0% dodecane binary amine.
  • the conversion rate of lactic acid oligomers in the second depolymerization reactor was 72.9%.
  • the conversion rate of lactic acid oligomers in the third depolymerization reactor was 71.9%.
  • composition of crude lactide obtained in the whole process is: 91.5% of L-lactide, 1.6% of m-lactide, 2.2% of L-lactic acid, and 3.1% of lactic acid dimer and trimer.
  • the conversion rate of lactic acid oligomers in the whole step cycle depolymerization process can reach 98.7%.
  • the above-mentioned crude lactide product is purified and refined through a two-stage rectification system.
  • the chemical purity and optical purity of the refined product can meet the requirements of the polymerization-grade lactide monomer.
  • Lactide was prepared according to the method of Example 1, except that the molecular weight of the circulation tank was not detected and adjusted according to the detection results.
  • the composition of crude lactide obtained in the whole process is: L-lactide 85.4%, m-lactide 6.5%, L-lactic acid content 4.7%, dimer and trimer content 5.8%.
  • the conversion rate of lactic acid oligomers in the whole step cycle depolymerization process can reach 95.3%.
  • the quality of the obtained crude lactide product is equivalent to that of step control, the conversion rate of lactic acid oligomers is on the low side, and the reaction treatment capacity of the whole process is only about 200g/h (while the treatment capacity in Example 1 can reach more than 300g/h), And with the continuous increase of the molecular weight of the circulating oligomers, the composition of the inlet of the depolymerization reactor is constantly changing, the depolymerization reaction rate is constantly changing, the degree of racemization is constantly increasing, and the composition of the export product is in a dynamic change. The stability is poor, and the device needs to be stopped regularly for slag discharge operation.
  • the lactic acid oligomer depolymerization technical scheme in the patent application of Nanjing University CN111153886A is used to prepare lactide, that is, the first-stage cycle depolymerization process is adopted.
  • the depolymerization reactor form in this comparative example is the same as that of Example 1, and the depolymerization reaction conditions are controlled. It is: the degree of vacuum is 300Pa, the reaction temperature is 210°C, and the one-way residence time is about 2min. After the reaction, the heavy components are discharged into the circulation tank, mixed with fresh lactic acid oligomers, and then transported to the depolymerization reactor to control the reaction process.
  • Fresh materials The mass ratio to the circulating material is 1:3; as the reaction progresses, the liquid level of the circulating tank is controlled to maintain at 60%, the pressure at 50kPa, and the temperature at 180°C. Use the molecular weight of lactic acid oligomers at the outlet of the circulation tank or the cumulative amount of catalyst circulation as a reference to carry out regular slagging, and carry out slagging when the molecular weight of the oligomers is greater than 10,000.
  • the content of L-lactide is 85.2%
  • the content of m-lactide is 8.6%
  • the content of L-lactic acid is 2.4%
  • the content of dimer and trimer is 3.2%.
  • the conversion rate of lactic acid oligomers in the crude lactide synthesis process reaches 95.2%.
  • the product yield can reach 95%
  • the racemization of the product is serious, and as the reaction progresses, the composition of the material at the inlet of the depolymerization reactor changes greatly, which makes the depolymerization reaction rate change greatly, and the depolymerization reaction time of the system is long ,
  • the material composition of the product export is unstable.
  • the test of depolymerization of lactic acid oligomers to prepare lactide is carried out, that is, the depolymerization is carried out in the form of three reactors connected in series, and the material of the third reactor is recycled back to the first reactor for reaction , and regularly discharge slag from the bottom of the third reactor, the form of the depolymerization reactor adopted is the same as in Example 1.
  • the same temperature of the three reactors in series is 200°C, and the vacuum degree increases successively, being respectively 600Pa, 400Pa and 300Pa (same as Example 1), the feed rate of the first depolymerization reactor is the same as that of Example 1, and the lactic acid
  • the conversion rate of oligomer in the first depolymerization reactor is 54.9%, and the unreacted material in the first depolymerization reactor is continuously transported to the second depolymerization reactor, and the conversion rate is 49.3%, while the second depolymerization reaction
  • the unreacted material in the tank is transported to the third depolymerization reactor for reaction, and the conversion rate is 40.1%, and the unreacted components are treated as slagging at this time, and the conversion rate of the whole reaction lactic acid oligomer is 86%, and the crude lactide
  • the content of m-lactide in the product is 5.8%, but the composition of crude lactide is relatively stable; but if the unreacted components are recycled
  • the present invention adopts multi-stage series depolymerization reaction and performs cascade control, regulates and controls according to the molecular weight of lactic acid oligomers in the circulation tank, realizes efficient depolymerization of lactic acid oligomers, reduces The degree of lactide racemization and the probability of coking and carbonization of the substrate are determined, and the conversion rate of lactic acid oligomers and system stability during the entire reaction process are guaranteed.
  • the conversion rate of lactic acid oligomers in the whole process can reach more than 97.0%.

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Abstract

本发明属于可生物降解材料技术领域，公开了一种梯级控制连续制备丙交酯的方法及系统。该方法包括：(1)使乳酸低聚物和解聚催化剂在第一解聚反应单元中进行反应，得到第一液相物料；(2)使第一液相物料在第二解聚反应单元中循环进行反应，直至液相物料的分子量高于6000，得到第二液相物料；(3)使第二液相物料在第三解聚反应单元中循环进行反应，直至液相物料的分子量高于10000；(4)从第一解聚反应单元、第二解聚反应单元和第三解聚反应单元中收集气相粗丙交酯并提纯。本发明实现了乳酸低聚物的高效解聚，降低了丙交酯消旋化程度和底物结焦碳化几率，保障了整个反应过程乳酸低聚物的转化率及系统稳定性。

Description

一种梯级控制连续制备丙交酯的方法及系统
相关申请的交叉引用
本申请要求2021年10月31日提交的中国专利申请202111279100.9的权益，该申请的内容通过引用被合并于本文。
技术领域
本发明属于可生物降解材料技术领域，具体涉及一种梯级控制连续制备丙交酯的方法及系统。
背景技术
目前，全球每年一次性塑料制品消费达1.2亿吨，只有10％被回收利用，另外约12％被焚烧，超过70％被丢弃到土壤、空气和海洋中。投放至海洋的塑料垃圾每年超过800万吨，并且这一数字还在不断上升，预计到2025年，全球海洋塑料垃圾量将达2.5亿吨。传统的一次性塑料制品，使用寿命较短，但物化性质稳定，自然降解难，大量一次性塑料制品废弃物导致的各类环境问题频出，已经严重危害到土地、水体及动物、人类的健康安全。全球已有近90个国家和地区出台了控制或者禁止一次性不可降解塑料制品的相关政策或规定。
目前商业化的生物降解塑料有聚乳酸(PLA)、聚羟基脂肪酸酯(PHA)、聚丁二酸丁二醇酯(PBS)、聚丁二酸-己二酸丁二醇酯(PBSA)、聚对苯二甲酸-己二酸丁二醇酯(PBAT)等，其中PLA目前应用最为广泛，应用前景最为突出。它不仅具有一般高分子材料的基本性能，而且在加工性能、物理机械性能和生物降解性能等方面更为优良，可广泛应用到包装行业、纺织行业、农用行业、消费品市场等，正逐步发展成为国民经济和社会发展所必需的基础性大宗原材料。
通常，聚乳酸采用两步法合成。具体的两步法合成聚乳酸的过程如下：第一步乳酸制备丙交酯；第二步丙交酯开环聚合制得聚乳酸，该过程所得PLA分子量可达到十万至百万。其中，丙交酯是整个合成过程的关键，工艺壁垒相对较高，通常需在催化剂、高温、高真空体系下通过缩聚、解聚过程制备得到，而该过程易造成丙交酯消旋化。这主要是由于生成丙交酯的解聚过程主要是发生在乳酸低 聚物分子链上的“背咬合”酯催化反应，具体过程为：在催化剂、高温、高真空作用下，乳酸低聚物链上的羰基被激活，链首段的羟基攻击带正电荷的羰基，使得酯键断裂(“正咬”过程)，形成L/D-丙交酯；但在碱性氧化物存在、催化剂过量或过高温度下，乳酸低聚物末端的羧酸阴离子攻击与该乳酸单元相邻的单元上的手性碳原子，从而使得次甲基碳与酯氧键之间的键断裂(“反咬”过程)，构型发生反转，得到内消旋丙交酯(m-丙交酯)，如式(1)所示。而m-丙交酯的存在，一方面会影响丙交酯的光学纯度，进而影响丙交酯开环聚合过程，使制得的PLA分子量偏低；另一方面会破坏PLA结构的规整性，使其结晶度降低，力学性能下降。
因此，解聚反应制得的粗丙交酯需要通过溶剂重结晶、水萃、精馏、熔融结晶等工艺进行提纯精制，以减少产品中m-丙交酯；而由于L-丙交酯和m-丙交酯物化性质相近，且丙交酯自身具有高凝点、沸点及热敏性等特性，使得分离困难，且整体收率较低，仅在40％-60％左右，整体经济性也偏低。因此，丙交酯合成过程的消旋化，是影响丙交酯品质和收率的关键因素，也是当前国内外丙交酯技术研究的重点和难点。
Figure PCTCN2022125901-appb-000001
Morteza Ehsani等在“Lactide synthesis optimization:investigation of the temperature,catalyst and pressure effects”一文中详细探究了温度、催化剂等因素对于解聚过程的影响，在反应温度考察时发现，当反应温度低时，m-丙交酯生成较少，随着温度的升高，m-丙交酯含量显著增加，反应温度为230℃时，m-丙交酯含量达到25.52％；而考察氧化亚锡、氯化亚锡、辛酸亚锡、三氧化锑及硫酸对 丙交酯合成过程的影响时，发现采用SnCl 2和硫酸做催化剂所得丙交酯产品纯度最高，且m-丙交酯含量最少；但催化剂浓度不能太高，以SnCl 2为例，随着SnCl 2浓度的增加，丙交酯生成产率增大，但过多的催化剂会加剧消旋化反应速率。
US5502215A公开了一种丙交酯精制提纯方法，该专利以SnO为催化剂，将乳酸低聚物加入釜式三口瓶中进行解聚反应制备得到粗D,L-丙交酯，该过程需强搅拌且反应温度高(反应温度220℃)，所得产品纯度低、且釜底底物结焦碳化严重，同时高温反应也加剧了丙交酯的消旋化。
US6326458B公开了一种制备丙交酯和丙交酯聚合物的连续工艺，该工艺在丙交酯制备解聚工段解聚反应器采用的是降膜式列管蒸发器，乳酸低聚物从蒸发器顶端加入，丙交酯蒸汽从列管反应器底部采出，未反应的乳酸低聚物从下出料口排出。该过程降膜反应过程所需反应温度相对较低，可有效降低解聚过程丙交酯消旋化几率，但丙交酯收率较低，为保持高的丙交酯收率，一般需降低进料速度，这又会造成低聚物在降膜反应器表面停留时间增加，未解聚的乳酸低聚物在高温高真空体系下快速聚合，造成低聚物分子量偏大，进一步影响解聚速率，而且易造成低聚物在降膜列管反应器表面的结焦、碳化。
CN111153886A公开了一种快速高产丙交酯的合成方法及装置，采用乳酸单组分或乳酸加入催化剂双组分经过混合器进入寡聚物制备系统，经过底部循环增加停留时间，合成寡聚乳酸，气相组分经过精馏系统提高寡聚乳酸产率；寡聚乳酸经过提纯装置去除未反应乳酸和水；脱轻后的寡聚乳酸加入催化剂后经过混合器，进入解聚反应器解聚为丙交酯，重组分经过回流再次进入解聚反应器，轻组分经过提纯回收系统后得到丙交酯产物。采用此装置可以高效合成丙交酯，在短停留时间0.5-5分钟内可获得产率为94％-98％的粗丙交酯，利用轻组分经过简单纯化系统后丙交酯产品中L-丙交酯、D-丙交酯或DL-丙交酯含量为94％-98％，m-丙交酯含量0.5％-5.5％。但是，该发明解聚后的重组分直接回流进入解聚反应器，不仅对解聚反应稳定性会造成影响，而且随着重组分分子量增加及催化剂累积，易增加反应底物在反应器表面结焦、碳化几率，增加了丙交酯的消旋化程度，影响反应的连续稳定运行。
发明内容
针对现有技术的不足，本发明提供了一种梯级控制连续制备丙交酯的方法及 系统。本发明通过对多级串联解聚进行梯级控制，实现了乳酸低聚物的高效解聚，降低了丙交酯消旋化程度和底物结焦碳化几率，保证了解聚过程的连续稳定操作及粗丙交酯产品组成的稳定性，提高了整体解聚反应速率、生产效率和丙交酯收率。
本发明提供一种梯级控制连续制备丙交酯的方法，该方法包括以下步骤：
(1)使乳酸低聚物和解聚催化剂在第一解聚反应单元中进行反应，得到第一液相物料；
(2)使所述第一液相物料在第二解聚反应单元中循环进行反应，直至液相物料的分子量高于6000，得到第二液相物料；
(3)使所述第二液相物料在第三解聚反应单元中循环进行反应，直至液相物料的分子量高于10000；
(4)从所述第一解聚反应单元、所述第二解聚反应单元和所述第三解聚反应单元中收集气相粗丙交酯并提纯。
优选的，所述第一解聚反应单元包括第一解聚反应器和第一循环罐，所述乳酸低聚物和所述解聚催化剂在所述第一解聚反应器中反应，反应后得到的所述第一液相物料进入所述第一循环罐。
优选的，步骤(2)所述反应在质子化溶剂的存在下进行。
优选的，所述第二解聚反应单元包括至少一个第二解聚反应器和至少一个第二循环罐，所述第一液相物料以及可选的质子化溶剂在所述第二解聚反应器中反应，反应后的液相物料进入所述第二循环罐，当所述液相物料的分子量为6000以下时，将所述第二循环罐中的液相物料循环回所述第二解聚反应器进一步反应；当所述液相物料的分子量大于6000时，将所述第二循环罐中的液相物料输送到所述第三解聚反应单元进行反应。
优选的，所述第三解聚反应单元包括至少一个第三解聚反应器和至少一个第三循环罐，所述第二液相物料在所述第三解聚反应器中反应，反应后的液相物料进入所述第三循环罐，当所述液相物料的分子量为10000以下时，将所述第三循环罐中的液相物料循环回所述第三解聚反应器进一步反应；当所述液相物料的分子量大于10000时，将所述第三循环罐中的液相物料排出。
本发明还提供一种梯级控制连续制备丙交酯的系统，该系统包括：
第一解聚反应单元，乳酸低聚物和解聚催化剂在所述第一解聚反应单元中进行反应；
第二解聚反应单元，来自所述第一解聚反应单元的液相物料以及可选的质子化溶剂在所述第二解聚反应单元循环进行反应，直至液相物料的分子量高于6000；
第三解聚反应单元，来自所述第二解聚反应单元的分子量高于6000的液相物料在所述第三解聚反应单元循环进行反应，直至液相物料的分子量高于10000；
以及用于从所述第一解聚反应单元、所述第二解聚反应单元和所述第三解聚反应单元中收集气相粗丙交酯并提纯的装置。
与现有技术相比，本发明的有益效果体现在：
(1)采用多级串联解聚反应并进行梯级控制，每级解聚反应单元主要包括解聚反应器与循环罐，根据循环罐中乳酸低聚物分子量情况进行调控，实现了乳酸低聚物的高效解聚，降低了丙交酯消旋化程度和底物结焦碳化几率，保障了整个反应过程乳酸低聚物的转化率及系统稳定性。整个过程乳酸低聚物转化率可达97.0％以上。
(2)通过不同梯级解聚反应器设置不同操作参数(包括反应温度、绝对压力(也称为真空度)以及反应时间或单程反应时间等)，保证了每一级解聚反应系统乳酸低聚物进料的稳定性、反应操作参数的稳定控制，提高每一梯级解聚反应器的反应速率，从而提高整体反应效率，保障了粗丙交酯出料品质和收率的稳定性，实现了整个解聚过程的连续稳定操作。与单一循环解聚系统过程相比，梯级解聚过程解聚效率提升30％。
(3)在较优选的实施方式中，在采用多级串联解聚反应单元进行梯级控制基础上，结合使用质子化溶剂，可以进一步降低丙交酯合成过程消旋化程度和底物结焦碳化几率，提高产品品质。三个解聚反应单元所得粗丙交酯中m-丙交酯含量均可控制在6.0％以内。相较釜式循环解聚反应器，丙交酯消旋化程度降低50％以上，相较单程刮膜蒸发器或多解聚反应器串联非梯级控制，乳酸低聚物转化率提高10％以上。
附图说明
图1为本发明所述的梯级控制连续制备丙交酯的系统的一种具体实施方式 的结构示意图。
附图标记说明
Ⅰ-第一解聚反应器、Ⅱ-第二解聚反应器、Ⅲ-第三解聚反应器、Ⅳ-第一循环罐、Ⅴ-第二循环罐、Ⅵ-第三循环罐、Ⅶ-第一分子量检测与控制组件、Ⅷ-第二分子量检测与控制组件；01-乳酸低聚物、02-第二循环罐进料、03-第三循环罐进料、04-粗丙交酯、05-第一循环罐出料、06-第二循环罐循环料、07-第二循环罐出料、08-第三循环罐循环料、09-乳酸高聚物，10-质子化溶剂。
具体实施方式
以下结合附图对本发明的具体实施方式进行详细说明。应当理解的是，此处所描述的具体实施方式仅用于说明和解释本发明，并不用于限制本发明。
本发明所述的梯级控制连续制备丙交酯的方法包括以下步骤：
(1)使乳酸低聚物和解聚催化剂在第一解聚反应单元中进行反应，得到第一液相物料；
(2)使所述第一液相物料在第二解聚反应单元中循环进行反应，直至液相物料的分子量高于6000，得到第二液相物料；
(3)使所述第二液相物料在第三解聚反应单元中循环进行反应，直至液相物料的分子量高于10000；
(4)从所述第一解聚反应单元、所述第二解聚反应单元和所述第三解聚反应单元中收集气相粗丙交酯并提纯。
在本发明所述的方法中，步骤(1)是进行初步的反应，通过控制条件实现物料的连续稳定反应，并实现出料的稳定性；步骤(2)是使物料进一步反应，通过调控单程反应时间来控制产物的消旋化，进一步提高反应效率和转化率；步骤(3)是物料深度反应，通过调节温度和真空度，使更大分子量的物料进一步参与解聚反应，同时通过调控物料单程反应时间来控制产物的消旋化，使产物达到产品要求，保证反应转化率。步骤(1)至(3)多级反应(优选为三级反应)相互配合并相互关联，保障了粗丙交酯出料品质和收率的稳定性，实现了整个解聚过程的连续稳定操作。
在本发明所述的方法中，步骤(1)所述乳酸低聚物的分子量可以为800-3000， 优选为1200-2800。在本文中，乳酸低聚物的分子量均是指重均分子量。
在本发明中，所述方法还可以包括按照以下工序制备乳酸低聚物：将L-乳酸和/或D-乳酸依次进行脱水、缩聚。所述脱水工序主要是脱除乳酸中的游离水，可以采用常压或减压形式。所述缩聚的条件可以包括：反应温度为140-170℃，绝对压力为1000-2000Pa，反应时间0.5-4h。
在本发明所述的方法中，优选的，步骤(1)所述解聚催化剂的用量为所述乳酸低聚物质量的0.4％-3％，进一步优选为0.8％-2％。
在本发明所述的方法中，优选的，步骤(1)所述解聚催化剂为锡类催化剂，进一步优选为辛酸亚锡、SnCl 2和SnO中的至少一种。
在本发明所述的方法中，优选的，步骤(1)所述反应的条件包括：反应温度为180-200℃，绝对压力为500-1500Pa，反应时间为3-8min。
按照本发明的一种具体实施方式，所述第一解聚反应单元包括第一解聚反应器和第一循环罐，所述乳酸低聚物和所述解聚催化剂在所述第一解聚反应器中反应，反应后得到的所述第一液相物料进入所述第一循环罐。
在优选情况下，在所述第一循环罐中，液位维持在50％-70％，压力维持在10kPa-常压，温度维持在160-200℃。在该优选情况下，可以减少乳酸低聚物继续发生分子间聚合反应的几率，减少结焦碳化。
在本发明所述的方法中，优选情况下，通过控制输送至第一解聚反应器中的乳酸低聚物的量，将步骤(1)所述反应中乳酸低聚物的转化率控制在50％-60％之间。在该优选情况下，可以减少乳酸低聚物继续发生分子间聚合反应的几率，减少结焦碳化。
在本发明所述的方法中，优选情况下，步骤(2)所述反应在质子化溶剂的存在下进行。在具体的实施方式中，将来自所述第一解聚反应单元的第一液相物料与质子化溶剂混合后输送至所述第二解聚反应单元。在该优选情况下，通过使用质子化溶剂，可进一步降低丙交酯合成过程消旋化程度和底物结焦碳化几率，提高产品品质。
在本发明中，所述质子化溶剂可以为碳原子数不小于12的二元胺和碳原子数不小于12的二元醇中的至少一种。优选地，所述质子化溶剂的熔融温度为80-160℃，更优选为100-160℃。进一步优选地，所述质子化溶剂为C12-C18的 二元胺和C12-C18的二元醇中的至少一种。更优选地，所述质子化溶剂为十二烷二元胺、十四烷二元胺(也称为十四烷二胺)、十六烷二元胺(也称为十六烷二胺)、十四烷二元醇(也称为十四烷二醇)和十六烷二元醇(也称为十六烷二醇)中的至少一种。
在本发明所述的方法中，在步骤(2)所述反应中，所述质子化溶剂的用量可以为乳酸低聚物质量的0.1％-6％，优选为1％-3％。
按照本发明的一种具体实施方式，所述第二解聚反应单元包括至少一个第二解聚反应器和至少一个第二循环罐，所述第一液相物料以及可选的质子化溶剂在所述第二解聚反应器中反应，反应后的液相物料进入所述第二循环罐，当所述液相物料的分子量为6000以下时，将所述第二循环罐中的液相物料循环回所述第二解聚反应器进一步反应；当所述液相物料的分子量大于6000时，将所述第二循环罐中的液相物料输送到所述第三解聚反应单元进行反应。在优选情况下，当所述液相物料的分子量为3000-6000时，将所述第二循环罐中的液相物料循环回所述第二解聚反应器进一步反应；当所述液相物料的分子量大于6000且小于10000时，将所述第二循环罐中的液相物料输送到所述第三解聚反应单元进行反应。
在上述具体实施方式中，所述第二解聚反应单元可以包括一个第二解聚反应器和一个第二循环罐，也可以包括两个或以上的第二解聚反应器和两个或以上的第二循环罐。优选地，当所述第二解聚反应单元包括两个或以上的第二解聚反应器和两个或以上的第二循环罐时，每个第二解聚反应器各自对应配置一个第二循环罐，且每个第二解聚反应器及其对应配置的第二循环罐构成一个循环反应单元。在一个具体例子中，所述第二解聚反应单元包括两个第二解聚反应器和两个第二循环罐，其中，第二解聚反应器A和第二循环罐a构成一个循环反应单元2-1，第二解聚反应器B和第二循环罐b构成另一个循环反应单元2-2，来自所述第一解聚反应单元的液相物料(也即第一液相物料)先进入所述循环反应单元2-1的第二解聚反应器A中进行反应，反应后的液相物料进入第二循环罐a，当液相物料的分子量为4500以下(如3000-45000)时，将第二循环罐a中的液相物料循环回第二解聚反应器A进一步反应；当液相物料的分子量大于4500且小于6000时，将第二循环罐a中的液相物料输送至循环反应单元2-2的第二解聚反应器B 中进行反应，反应后的液相物料进入第二循环罐b，当液相物料的分子量为6000以下(如4500-6000)时，将第二循环罐b中的液相物料循环回第二解聚反应器B进一步反应；当液相物料的分子量大于6000且小于10000时，将第二循环罐b中的液相物料输送到所述第三解聚反应单元进行反应。在实际操作过程中，所述第二解聚反应单元中包含的由一个解聚反应器和一个循环罐构成的循环反应单元的数量越多，可实现更精细地控制不同阶段乳酸低聚物的分子量，从而获得更好的反应效果，然而，综合考虑成本和效果提升的幅度，所述第二解聚反应单元优选包括一个第二解聚反应器和一个第二循环罐。
在本发明所述的方法中，所述第二解聚反应器中的反应条件可以包括：反应温度为200-220℃，绝对压力为400-1000Pa，单程反应时间为2-5min。
在本发明所述的方法中，在优选情况下，在所述第二解聚反应器的单程反应中，乳酸低聚物的进料量为实际反应量的3-5倍。在该优选情况下，可以降低乳酸低聚物在第二解聚反应器表面的停留时间，抑制聚合反应发生，提高收率，保证产品品质。
在本发明所述的方法中，在优选情况下，在所述第二循环罐中，液位维持在50％-70％，压力维持在10kPa-常压，温度维持在160-200℃。
在本发明所述的方法中，步骤(2)的解聚反应过程乳酸低聚物的转化率可达到70％以上。
按照本发明的一种具体实施方式，所述第三解聚反应单元包括至少一个第三解聚反应器和至少一个第三循环罐，所述第二液相物料在所述第三解聚反应器中反应，反应后的液相物料进入所述第三循环罐，当所述液相物料的分子量为10000以下时，将所述第三循环罐中的液相物料循环回所述第三解聚反应器进一步反应；当所述液相物料的分子量大于10000时，将所述第三循环罐中的液相物料排出。
在上述具体实施方式中，所述第三解聚反应单元可以包括一个第三解聚反应器和一个第三循环罐，也可以包括两个或以上的第三解聚反应器和两个或以上的第三循环罐。优选地，当所述第三解聚反应单元包括两个或以上的第三解聚反应器和两个或以上的第三循环罐时，每个第三解聚反应器各自对应配置一个第三循环罐，且每个第三解聚反应器及其对应配置的第三循环罐构成一个循环反应单元。 在一个具体例子中，所述第三解聚反应单元包括两个第三解聚反应器和两个第三循环罐，其中，第三解聚反应器C和第三循环罐c构成一个循环反应单元3-1，第三解聚反应器D和第三循环罐d构成另一个循环反应单元3-2，来自所述第二解聚反应单元的液相物料(也即第二液相物料)先进入所述循环反应单元3-1的第三解聚反应器C中进行反应，反应后的液相物料进入第三循环罐c，当液相物料的分子量为8000以下(如大于6000且小于等于8000)时，将第三循环罐c中的液相物料循环回第三解聚反应器C进一步反应；当液相物料的分子量大于8000且小于10000时，将第三循环罐c中的液相物料输送至循环反应单元3-2的第三解聚反应器D中进行反应，反应后的液相物料进入第三循环罐d，当液相物料的分子量为10000以下时，将第三循环罐d中的液相物料循环回第三解聚反应器D进一步反应；当液相物料的分子量大于10000时，将第三循环罐d中的液相物料排出体系。在实际操作过程中，所述第三解聚反应单元中包含的由一个解聚反应器和一个循环罐构成的循环反应单元的数量越多，可实现更精细地控制不同阶段乳酸低聚物的分子量，从而获得更好的反应效果，然而，综合考虑成本和效果提升的幅度，所述第三解聚反应单元优选包括一个第三解聚反应器和一个第三循环罐。
在本发明所述的方法中，所述第三解聚反应器中的反应条件可以包括：反应温度为220-240℃，绝对压力为200-800Pa，单程反应时间1-4min。
在本发明所述的方法中，在优选情况下，在所述第三解聚反应器的单程反应中，乳酸低聚物的进料量为实际反应量的4-6倍。在该优选情况下，可以降低乳酸低聚物在第三解聚反应器表面停留时间，抑制聚合反应发生，提高收率，保证产品品质。
在本发明所述的方法中，在优选情况下，在所述第三循环罐中，液位维持在10％-30％，压力维持在10kPa-常压，温度维持在160-200℃。
在本发明所述的方法中，从所述第三解聚反应单元中排出的分子量大于10000的液相物料是乳酸高聚物，可以通过水解回相乳酸。
在本发明所述的方法中，步骤(3)的解聚反应过程乳酸低聚物的转化率可达到70％以上。
按照本发明所述的方法，整个梯级控制解聚反应过程乳酸低聚物的转化率可 达到97.0％以上。
在本发明所述的方法中，在优选情况下，所述第一解聚反应单元、所述第二解聚反应单元和所述第三解聚反应单元中的反应器(即第一解聚反应器、第二解聚反应器和第三解聚反应器)各自为刮膜解聚反应器，更优选为薄膜蒸发器、分子蒸馏蒸发器或其他搅拌膜式蒸发器。
按照本发明所述的方法，三个解聚反应单元生成的气相粗丙交酯由各解聚反应器顶部排出，所得的粗丙交酯质量组成为：L-丙交酯含量82％-92％，m-丙交酯含量1.0％-6％，L-乳酸的含量为0.5％-6％，二聚体及三聚体含量为1.5％-6％。
按照本发明所述的方法，三个解聚反应单元生成的气相粗丙交酯进入分离提纯工序，可以直接通过精馏或其他提纯精制工艺进行精制，所得产品品质满足聚合级丙交酯单体需求。
本发明所述的梯级控制连续制备丙交酯的系统包括：
第一解聚反应单元，乳酸低聚物和解聚催化剂在所述第一解聚反应单元中进行反应；
第二解聚反应单元，来自所述第一解聚反应单元的液相物料以及可选的质子化溶剂在所述第二解聚反应单元循环进行反应，直至液相物料的分子量高于6000；
第三解聚反应单元，来自所述第二解聚反应单元的分子量高于6000的液相物料在所述第三解聚反应单元循环进行反应，直至液相物料的分子量高于10000；
以及用于从所述第一解聚反应单元、所述第二解聚反应单元和所述第三解聚反应单元中收集气相粗丙交酯并提纯的装置。
在优选情况下，所述第一解聚反应单元包括第一解聚反应器和第一循环罐，所述乳酸低聚物和所述解聚催化剂在所述第一解聚反应器中反应，反应后得到的液相物料进入所述第一循环罐。
在优选情况下，所述第二解聚反应单元包括至少一个第二解聚反应器和至少一个第二循环罐，来自所述第一解聚反应单元的液相物料以及可选的质子化溶剂在所述第二解聚反应器中反应，反应后的液相物料进入所述第二循环罐，当所述液相物料的分子量为6000以下时，将所述第二循环罐中的液相物料循环回所述第二解聚反应器进一步反应；当所述液相物料的分子量大于6000时，将所述第 二循环罐中的液相物料输送到所述第三解聚反应单元进行反应。
在优选情况下，所述第三解聚反应单元包括至少一个第三解聚反应器和至少一个第三循环罐，来自所述第二解聚反应单元的分子量高于6000的液相物料在所述第三解聚反应器中反应，反应后的液相物料进入所述第三循环罐，当所述液相物料的分子量为10000以下时，将所述第三循环罐中的液相物料循环回所述第三解聚反应器进一步反应；当所述液相物料的分子量大于10000时，将所述第三循环罐中的液相物料排出系统。
在本发明所述的系统中，所述第一解聚反应单元、所述第二解聚反应单元和所述第三解聚反应单元中的反应器(即第一解聚反应器、第二解聚反应器和第三解聚反应器)各自可以为刮膜解聚反应器，优选为薄膜蒸发器、分子蒸馏蒸发器或其他搅拌膜式蒸发器。
根据本发明的一种具体实施方式，如图1所示，所述梯级控制连续制备丙交酯的系统包括第一解聚反应单元、第二解聚反应单元和第三解聚反应单元，其中，所述第一解聚反应单元包括第一解聚反应器I和第一循环罐Ⅳ，所述第二解聚反应单元包括第二解聚反应器Ⅱ和第二循环罐Ⅴ，所述第二循环罐Ⅴ配置有第一分子量检测与控制组件Ⅶ，所述第三解聚反应单元包括第三解聚反应器Ⅲ和第三循环罐Ⅵ，所述第三循环罐Ⅵ配置有第二分子量检测与控制组件ⅦI。该系统的具体运行过程是：首先将乳酸低聚物01输送至第一解聚反应器Ⅰ中进行反应，未反应的乳酸低聚物排入第一循环罐Ⅳ中，控制压力在10kPa-常压，温度在160-200℃，待罐中液位达到50％-70％，将第一循环罐出料05单独或者进一步与质子化溶剂10混合后输送至第二解聚反应器Ⅱ中进行反应，未反应的乳酸低聚物进入第二循环罐Ⅴ中，控制压力在10kPa-常压，温度在160-200℃，待罐中液位达到50％-70％，检测第二循环罐Ⅴ中乳酸低聚物分子量，若分子量不高于6000(优选3000-6000)，则作为第二循环罐循环料06返回第二解聚反应器Ⅱ中，若分子量高于6000，则作为第二循环罐出料07输送至第三解聚反应器Ⅲ进行反应，未反应的乳酸低聚物进入第三循环罐Ⅵ中，调控压力在10kPa-常压，温度在160-200℃，待其中液位达到10％-30％，检测第三循环罐Ⅵ中低聚物分子量，若分子量低于1万，则作为第三循环罐循环料08返回第三解聚反应器Ⅲ中，若分子量大于1万，则作为乳酸高聚物09排出反应体系，排出的物料为乳酸高聚物， 可通过水解形成乳酸进行回用。三个解聚反应器生成的气相粗丙交酯由各解聚反应器顶部排出，得到粗丙交酯产品04，进入分离提纯工序。
下面通过实施例来进一步说明本发明制备丙交酯的方法及系统。实施例在以本发明技术方案为前提下进行实施，给出了详细的实施方式和具体操作过程，但本发明的保护范围不限于下述实施例。
以下实施例中的实验方法，如无特殊说明，均为本领域常规方法。下述实施例中所用的实验材料，如无特殊说明，均可从生化试剂商店购买得到。
本发明实施例采用的乳酸为乳酸含量在88％及以上的耐热级L-乳酸，其光学纯度不低于99.0％。
本发明采用马尔文Viscotek OMNISEC GPC/SEC凝胶色谱仪分析乳酸低聚物的分子量。采用传统校正方法，以聚苯乙烯(PS)为内标，色谱柱型号为T3000，尺寸为300mmL×8.0mm，柱温40℃，流速1.0mL/min，样品浓度2-5mg/mL，单次进样量500μL。
本发明采用安捷伦高效液相色谱仪来分析丙交酯化学纯度、L-乳酸及二聚体、三聚体含量，紫外检测器，采用磷酸和乙腈做流动相，色谱柱型号为ZORBAX SB-Aq，柱长250mm，柱内径4.6mm，内装填料粒径5μm。检测波长：200nm，柱温：40℃，流速：1mL/min，进样量：5μL。
本发明采相安捷伦气相色谱仪来分析不同光学异构体丙交酯含量，选用CYCLOSIL-B型号色谱柱、气化室温度250℃，检测器温度280℃，氢火焰离子检测器，柱温程序升温初温100℃，保持5min，以4℃/min的速率升温至140℃，保持7min，以8℃/min的速率升温至200℃，保持20min，载气N 2流量1.4mL/min，氢气流量30mL/min，空气流量400mL/min，进样量0.5μL。
丙交酯提纯过程的收率Y和整个制备提纯过程产品收率Y 总的计算公式如下：
Figure PCTCN2022125901-appb-000002
Figure PCTCN2022125901-appb-000003
其中，m 0为粗丙交酯的质量，y 0为粗丙交酯中L-丙交酯纯度，m为丙交酯产品质量，M为一定量乳酸低聚物理论能转化成的丙交酯的质量，即乳酸低聚物的质量。
采用WZZ-2S自动旋光仪分析样品的比旋光，从而表征样品的光学纯度，纯 L-丙交酯的比旋光为-278，纯D-丙交酯的比旋光为+278，m-丙交酯的比旋光度为0，样品光学纯度X计算公式如下；
Figure PCTCN2022125901-appb-000004
其中，α 纯物质表示纯丙交酯的比旋光，α 被测样品表示被测物质比旋光。
本发明实施例按照图1的装置和流程进行，首先将乳酸低聚物01输送至第一刮膜解聚反应器Ⅰ中进行反应，未反应的乳酸低聚物排入循环罐Ⅳ中，控制压力在10kPa-常压，温度在160-200℃，待罐中液位50％-70％，将第一循环罐出料05单独或者进一步与质子化溶剂10混合后输送至第二刮膜解聚反应器Ⅱ中进行反应，未反应的乳酸低聚物进入循环罐Ⅴ中，控制压力在10kPa-常压，温度在160-200℃，待罐中液位达到50％-70％，检测循环罐Ⅴ中低聚物分子量，若分子量低于6000优选3000-6000，则作为第二循环罐循环料06返回第二解聚反应器Ⅱ中，若分子量高于6000，则作为第二循环罐出料07输送至第三刮膜解聚反应器Ⅲ进行反应，未反应的乳酸低聚物进入循环罐Ⅵ中，调控压力在10kPa-常压，温度在160-200℃，待其中液位达到10％-30％，检测循环罐Ⅵ中低聚物分子量，若分子量低于1万，则作为第三循环罐循环料08返回第三解聚反应器Ⅲ中，若分子量大于1万，则作为乳酸高聚物09排出反应体系，排出的物料为乳酸高聚物10，可通过水解形成乳酸进行回用。三个解聚反应器生成的气相粗丙交酯由各解聚反应器顶部排出，得到粗丙交酯产品04，进入分离提纯工段。
乳酸低聚物的制备：
(1)乳酸脱游离水：取4000g L-乳酸(其中乳酸含量在88.0％左右，光学纯度为99.2％)，加入带搅拌系统的反应釜之中，采用真空循环水泵，维持系统的压力在50kPa左右，在真空下开始加热，逐步加热至110-120℃，脱水2h，此时反应体系中的游离水被缓慢蒸出反应体系。
(2)乳酸低聚物制备：在体系内的游离水几乎被完全脱除后，提高系统的真空度，将体系的压力缓慢降至1.2kPa左右，料液的温度逐步升至160℃，反应2.5h，此时，乳酸分子间发生缩聚反应，体系内反应生成的水分被蒸出体系，得到分子量为1901的乳酸低聚物。
实施例1
取乳酸低聚物3000g，加入30g辛酸亚锡催化剂，混匀后输送至第一刮膜解聚反应器中，控制解聚反应条件为：真空度600Pa，反应温度190℃，单程反应时间为4min，未反应乳酸低聚物排入第一循环罐，温度控制在180℃、压力控制在20kPa，待液位增至60％后，测得乳酸低聚物分子量为3216。该过程乳酸低聚物转化率为54.6％。
将第一循环罐中乳酸低聚物输送至第二刮膜解聚反应器，控制反应温度210℃，真空度400Pa，单程反应时间为3min，乳酸低聚物的进料量为实际反应量的4倍，未反应乳酸低聚物排入第二循环罐中，液位维持在50％，压力维持在10kPa，温度维持在180℃，经检测第二循环罐中乳酸低聚物分子量为4915，通过调控循环回第二解聚反应器继续参与反应；当测得第二循环罐出口低聚物分子量大于6000时，输送至第三解聚反应器。该过程乳酸低聚物转化率为72.4％。
将第二循环罐中乳酸低聚物输送至第三解聚反应器，反应温度230℃，真空度300Pa，单程反应时间2min，乳酸低聚物的进料量为实际反应量的5倍，未反应乳酸低聚物排入第三循环罐，温度控制在180℃、压力控制在20kPa，待液位增至20％后，检测乳酸低聚物分子量为7864，循环回第三解聚反应器中继续参与反应，待乳酸低聚物分子量增至10000后，排出体系。该过程乳酸低聚物转化率为71.1％。
整个过程获得的粗丙交酯的组成为：L-丙交酯89.7％，m-丙交酯3.2％，L-乳酸2.5％，乳酸二聚体、三聚体含量为3.4％。整个梯级循环解聚过程乳酸低聚物的转化率可达到98.3％。
将上述粗丙交酯产品经两级精馏系统提纯精制产品化学纯度、光学纯度均可满足聚合级丙交酯单体要求。
实施例2
取乳酸低聚物3000g，加入30g辛酸亚锡催化剂，混匀后输送至第一刮膜解聚反应器中，控制解聚反应条件为：真空度1500Pa，反应温度230℃，单程反应时间为8min，未反应乳酸低聚物排入第一循环罐，温度控制在200℃、压力控制在20kPa，待液位增至60％后，测得到乳酸低聚物分子量为3713。该过程乳酸低聚物转化率为50.8％。
将第一循环罐中乳酸低聚物输送至第二刮膜解聚反应器，控制反应温度220℃，真空度1000Pa，单程反应时间为5min，乳酸低聚物的进料量为实际反应量的3倍，未反应乳酸低聚物排入第二循环罐中，液位维持在50％，压力维持在20kPa，温度维持在180℃，经检测第二循环罐中乳酸低聚物分子量为6032，输送至第三解聚反应器中。该过程乳酸低聚物转化率为71.2％。
将第二循环罐中乳酸低聚物输送至第三解聚反应器，反应温度240℃，真空度800Pa，单程反应时间4min，乳酸低聚物的进料量为实际反应量的4倍，未反应乳酸低聚物排入第三循环罐，温度控制在180℃、压力控制在20kPa，待液位增至20％后，检测乳酸低聚物分子量为7357，循环回第三解聚反应器中继续参与反应，待乳酸低聚物分子量增至10000后，排出体系。该过程乳酸低聚物转化率为70.3％。
整个过程获得的粗丙交酯的组成为：L-丙交酯84.6％，m-丙交酯5.9％，L-乳酸1.5％，乳酸二聚体、三聚体含量为4.0％。整个梯级循环解聚过程乳酸低聚物的转化率可达到97.1％。
将上述粗丙交酯产品经两级精馏系统提纯精制产品化学纯度、光学纯度均可满足聚合级丙交酯单体要求。
实施例3
取乳酸低聚物3000g，加入30g辛酸亚锡催化剂，混匀后输送至第一刮膜解聚反应器中，控制解聚反应条件为：真空度500Pa，反应温度180℃，单程反应时间为3min，乳酸低聚物的进料量为实际反应量的4倍；未反应乳酸低聚物排入第一循环罐，温度控制在180℃、压力控制在20kPa，待液位增至60％后，测得乳酸低聚物分子量为3461。该过程乳酸低聚物转化率为51.8％。
将第一循环罐中乳酸低聚物输送至第二刮膜解聚反应器，控制反应温度200℃，真空度400Pa，单程反应时间为2min，乳酸低聚物的进料量为实际反应量的5倍，未反应乳酸低聚物排入第二循环罐中，液位维持在50％，压力维持在20kPa，温度维持在180℃，经检测第二循环罐中乳酸低聚物分子量为4456，循环回第二解聚反应器继续参与反应，当测得第二循环罐出口低聚物分子量大于6000时，输送至第三解聚反应器。该过程乳酸低聚物转化率为71.9％。
将第二循环罐中乳酸低聚物输送至第三解聚反应器，反应温度220℃，真空 度200Pa，单程反应时间1min，乳酸低聚物的进料量为实际反应量的6倍，未反应乳酸低聚物排入第三循环罐，温度控制在180℃、压力控制在20kPa，待液位增至20％后，检测乳酸低聚物分子量为7071，循环回第三解聚反应器中继续参与反应，待乳酸低聚物分子量增至10000后，排出体系。该过程乳酸低聚物转化率为72.4％。
整个过程获得的粗丙交酯的组成为：L-丙交酯89.2％，m-丙交酯3.2％，L-乳酸2.9％，乳酸二聚体、三聚体含量为4.7％，整个梯级循环解聚过程乳酸低聚物的转化率可达到97.6％。
将上述粗丙交酯产品经两级精馏系统提纯精制产品化学纯度、光学纯度均可满足聚合级丙交酯单体要求。
实施例4
按照实施例1的方法制备丙交酯，不同之处在于：采用相同质量的氯化亚锡催化剂。在第一解聚反应器中反应后，经检测第一循环罐出口乳酸低聚物分子量为3527，乳酸低聚物转化率为53.8％。该乳酸低聚物进入第二解聚反应器反应后，经检测第二循环罐中乳酸低聚物分子量为5319，直接循环回第二解聚反应器继续参与反应，当测得第二循环罐出口乳酸低聚物分子量大于6000时，输送至第三解聚反应器中，该过程乳酸低聚物转化率为71.7％。在第三解聚反应器反应后，测得第三循环罐出口乳酸低聚物分子量为8213，循环回第三解聚反应器中继续参与反应，待乳酸低聚物分子量增至10000后，排出体系。该过程乳酸低聚物转化率为70.8％。
整个过程获得的粗丙交酯组成为：L-丙交酯88.2％，m-丙交酯3.9％，L-乳酸2.7％，乳酸二聚体、三聚体含量为4.4％。整个梯级循环解聚过程乳酸低聚物的转化率可达到97.8％。
将上述粗丙交酯产品经两级精馏系统提纯精制产品化学纯度、光学纯度均可满足聚合级丙交酯单体要求。
实施例5
按照实施例1的方法制备丙交酯，不同之处在于：采用分子蒸馏形式的解聚蒸发器。在第一解聚反应器中反应后，经检测第一循环罐出口乳酸低聚物分子量 为3009，乳酸低聚物转化率为55.3％。该乳酸低聚物进入第二解聚反应器反应后，该过程乳酸低聚物转化率为73.5％。在第三解聚反应器反应后，该过程乳酸低聚物转化率为72.7％。
整个过程获得的粗丙交酯组成为：L-丙交酯91.3％，m-丙交酯2.2％，L-乳酸2.9％，乳酸二聚体、三聚体含量为3.0％。整个梯级循环解聚过程乳酸低聚物的转化率可达到98.7％。
将上述粗丙交酯产品经两级精馏系统提纯精制产品化学纯度、光学纯度均可满足聚合级丙交酯单体要求。
实施例6
按照实施例1的方法制备丙交酯，不同之处在于：第一解聚反应器中未反应的乳酸低聚物在送入第二解聚反应器之前，加入1.0％的十四烷二元醇。第二解聚反应反应器中乳酸低聚物转化率为74.4％。第三解聚反应器中乳酸低聚物转化率为72.4％。
整个过程获得的粗丙交酯组成为：L-丙交酯91.3％，m-丙交酯1.8％，L-乳酸2.1％，乳酸二聚体、三聚体含量为3.3％。整个梯级循环解聚过程乳酸低聚物的转化率可达到98.6％。所得各梯级所得粗丙交酯产品消旋化程度低，整个过程物料处理量高。
将上述粗丙交酯产品经两级精馏系统提纯精制产品化学纯度、光学纯度均可满足聚合级丙交酯单体要求。
实施例7
按照实施例1的方法制备丙交酯，不同之处在于：第一解聚反应器中未反应的乳酸低聚物在送入第二解聚反应器之前，加入1.0％的十二烷二元胺。第二解聚反应反应器中乳酸低聚物转化率为72.9％。第三解聚反应器中乳酸低聚物转化率为71.9％。
整个过程获得的粗丙交酯组成为：L-丙交酯91.5％，m-丙交酯1.6％，L-乳酸2.2％，乳酸二聚体、三聚体含量为3.1％。整个梯级循环解聚过程乳酸低聚物的转化率可达到98.7％。
将上述粗丙交酯产品经两级精馏系统提纯精制产品化学纯度、光学纯度均可 满足聚合级丙交酯单体要求。
对比例1
按照实施例1的方法制备丙交酯，不同之处在于：未对循环罐分子量进行检测并根据检测结果进行调控。整个过程获得的粗丙交酯组成为：L-丙交酯85.4％，m-丙交酯6.5％，L-乳酸的含量为4.7％，二聚体、三聚体含量为5.8％。整个梯级循环解聚过程乳酸低聚物的转化率可达到95.3％。尽管所得粗丙交酯产品品质与梯级控制相当，但乳酸低聚物转化率偏低，且整个过程反应处理量只有200g/h左右(而实施例1中处理量能够达到300g/h以上)，且随着循环低聚物分子量的不断增大，解聚反应器进口组成不断变化，解聚反应速率不断变化、消旋化程度不断增大，出口产品组成处于动态变化之中，连续过程中反应稳定性差，且需要定期停装置进行排渣操作。
对比例2
采用南京大学CN111153886A专利申请中乳酸低聚物解聚技术方案进行丙交酯制备，即采用一级循环解聚工艺，该对比例中解聚反应器形式与实施例1相同，控制解聚反应条件为：真空度在300Pa，反应温度210℃，单程停留时间为2min左右，反应后重组分排入循环罐，并与新鲜乳酸低聚物混合后输送至解聚反应器中，控制反应过程新鲜物料与循环物料质量比为1:3；随着反应的进行，控制循环罐液位维持在60％，压力维持在50kPa，温度维持在180℃。以循环罐出口乳酸低聚物分子量或催化剂循环累积量为参考进行定期排渣，当低聚物分子量大于10000时进行排渣。
经分析所粗丙交酯产品中，L-丙交酯含量为85.2％，m-丙交酯含量为8.6％，L-乳酸的含量为2.4％，二聚体、三聚体含量为3.2％，粗丙交酯合成过程乳酸低聚物转化率达95.2％。尽管产品收率可达到95％，但是产品消旋化严重，且随着反应的进行，解聚反应器进口物料组成变化较大、使得解聚反应速率变化亦较大，体系解聚反应时间长，产品出口物料组成不稳定。
对比例3
采用与日本专利申请JPH08333359A相似的技术方案进行乳酸低聚物解聚制备丙交酯的试验，即采用三反应器串联的形式进行解聚，第三反应器的物料循环回第一反应器进行反应，并从第三反应器底定期排渣，所采用解聚反应器形式同实施例1。三反应器串联反应温度相同均为200℃，真空度依次升高，分别为600Pa、400Pa和300Pa(与实施例1相同)，第一解聚反应器的进料速率与实施例1相同，乳酸低聚物在第一解聚反应器中转化率为54.9％，第一解聚反应器中未反应的物料连续输送至第二解聚反应器，转化率为49.3％，而第二解聚反应器未反应的物料输送至第三解聚反应器反应，转化率为40.1％，如此时未反应的组分做排渣处理，整个反应乳酸低聚物的转化率为86％，粗丙交酯产品中m-丙交酯含量为5.8％，但粗丙交酯组成相对稳定；但如果此时将未反应的组分再次循环回第一解聚反应器，此时，各反应器进口的组成变化将较大，产品收率虽有提升，但是反应效率、反应系统的稳定性被破坏。
由上述实施例和对比例可以看出，本发明采用多级串联解聚反应并进行梯级控制，根据循环罐中乳酸低聚物分子量情况进行调控，实现了乳酸低聚物的高效解聚，降低了丙交酯消旋化程度和底物结焦碳化几率，保障了整个反应过程乳酸低聚物的转化率及系统稳定性。整个过程乳酸低聚物转化率可达97.0％以上。
以上详细描述了本发明的优选实施方式，但是，本发明并不限于此。在本发明的技术构思范围内，可以对本发明的技术方案进行多种简单变型，包括各个技术特征以任何其它的合适方式进行组合，这些简单变型和组合同样应当视为本发明所公开的内容，均属于本发明的保护范围。

Claims (29)

  1. 一种梯级控制连续制备丙交酯的方法，其特征在于，该方法包括以下步骤：
    (1)使乳酸低聚物和解聚催化剂在第一解聚反应单元中进行反应，得到第一液相物料；
    (2)使所述第一液相物料在第二解聚反应单元中循环进行反应，直至液相物料的分子量高于6000，得到第二液相物料；
    (3)使所述第二液相物料在第三解聚反应单元中循环进行反应，直至液相物料的分子量高于10000；
    (4)从所述第一解聚反应单元、所述第二解聚反应单元和所述第三解聚反应单元中收集气相粗丙交酯并提纯。
  2. 根据权利要求1所述的方法，其特征在于，在步骤(1)中，所述乳酸低聚物的分子量为800-3000，优选为1200-2800。
  3. 根据权利要求1或2所述的方法，其特征在于，所述方法还包括按照以下工序制备乳酸低聚物：将L-乳酸和/或D-乳酸依次进行脱水、缩聚，
    优选地，所述缩聚的条件包括：反应温度为140-170℃，绝对压力为1000-2000Pa，反应时间为0.5-4h。
  4. 根据权利要求1所述的方法，其特征在于，在步骤(1)中，所述解聚催化剂的用量为所述乳酸低聚物质量的0.4％-3％，优选为0.8％-2％。
  5. 根据权利要求1或4所述的方法，其特征在于，在步骤(1)中，所述解聚催化剂为锡类催化剂，优选为辛酸亚锡、SnCl 2和SnO中的至少一种。
  6. 根据权利要求1所述的方法，其特征在于，步骤(1)所述反应的条件包括：反应温度为180-200℃，绝对压力为500-1500Pa，反应时间为3-8min。
  7. 根据权利要求1所述的方法，其特征在于，所述第一解聚反应单元包括第一解聚反应器和第一循环罐，所述乳酸低聚物和所述解聚催化剂在所述第一解聚反应器中反应，反应后得到的所述第一液相物料进入所述第一循环罐。
  8. 根据权利要求7所述的方法，其特征在于，在所述第一循环罐中，液位维持在50％-70％，压力维持在10kPa-常压，温度维持在160-200℃。
  9. 根据权利要求1、6、7或8所述的方法，其特征在于，步骤(1)所述反应中，乳酸低聚物的转化率控制在50％-60％之间。
  10. 根据权利要求1所述的方法，其特征在于，步骤(2)所述反应在质子化溶剂的存在下进行。
  11. 根据权利要求10所述的方法，其特征在于，所述质子化溶剂为碳原子数不小于12的二元胺和碳原子数不小于12的二元醇中的至少一种。
  12. 根据权利要求11所述的方法，其特征在于，所述质子化溶剂的熔融温度为80-160℃，优选为100-160℃。
  13. 根据权利要求11或12所述的方法，其特征在于，所述质子化溶剂为C12-C18的二元胺和C12-C18的二元醇中的至少一种，优选为十二烷二元胺、十四烷二元胺、十六烷二元胺、十四烷二元醇和十六烷二元醇中的至少一种。
  14. 根据权利要求10、11或12所述的方法，其特征在于，在步骤(2)所述反应中，所述质子化溶剂的用量为乳酸低聚物质量的0.1％-6％，优选为1％-3％。
  15. 根据权利要求1、10、11或12所述的方法，其特征在于，所述第二解聚反应单元包括至少一个第二解聚反应器和至少一个第二循环罐，所述第一液相物料以及可选的质子化溶剂在所述第二解聚反应器中反应，反应后的液相物料进 入所述第二循环罐，当所述液相物料的分子量为6000以下时，将所述第二循环罐中的液相物料循环回所述第二解聚反应器进一步反应；当所述液相物料的分子量大于6000时，将所述第二循环罐中的液相物料输送到所述第三解聚反应单元进行反应。
  16. 根据权利要求15所述的方法，其特征在于，当所述液相物料的分子量为3000-6000时，将所述第二循环罐中的液相物料循环回所述第二解聚反应器进一步反应；当所述液相物料的分子量大于6000且小于10000时，将所述第二循环罐中的液相物料输送到所述第三解聚反应单元进行反应。
  17. 根据权利要求15所述的方法，其特征在于，所述第二解聚反应器中的反应条件包括：反应温度为200-220℃，绝对压力为400-1000Pa，单程反应时间为2-5min。
  18. 根据权利要求15所述的方法，其特征在于，在所述第二解聚反应器的单程反应中，乳酸低聚物的进料量为实际反应量的3-5倍。
  19. 根据权利要求15所述的方法，其特征在于，在所述第二循环罐中，液位维持在50％-70％，压力维持在10kPa-常压，温度维持在160-200℃。
  20. 根据权利要求1所述的方法，其特征在于，所述第三解聚反应单元包括至少一个第三解聚反应器和至少一个第三循环罐，所述第二液相物料在所述第三解聚反应器中反应，反应后的液相物料进入所述第三循环罐，当所述液相物料的分子量为10000以下时，将所述第三循环罐中的液相物料循环回所述第三解聚反应器进一步反应；当所述液相物料的分子量大于10000时，将所述第三循环罐中的液相物料排出。
  21. 根据权利要求20所述的方法，其特征在于，所述第三解聚反应器中的反应条件包括：反应温度为220-240℃，绝对压力为200-800Pa，单程反应时间 1-4min。
  22. 根据权利要求20或21所述的方法，其特征在于，在所述第三解聚反应器的单程反应中，乳酸低聚物的进料量为实际反应量的4-6倍。
  23. 根据权利要求20或21所述的方法，其特征在于，在所述第三循环罐中，液位维持在10％-30％，压力维持在10kPa-常压，温度维持在160-200℃。
  24. 根据权利要求1、7、15或20所述的方法，其特征在于，所述第一解聚反应单元、所述第二解聚反应单元和所述第三解聚反应单元中的反应器各自为刮膜解聚反应器，优选为薄膜蒸发器、分子蒸馏蒸发器或其他搅拌膜式蒸发器。
  25. 一种梯级控制连续制备丙交酯的系统，其特征在于，该系统包括：
    第一解聚反应单元，乳酸低聚物和解聚催化剂在所述第一解聚反应单元中进行反应；
    第二解聚反应单元，来自所述第一解聚反应单元的液相物料以及可选的质子化溶剂在所述第二解聚反应单元循环进行反应，直至液相物料的分子量高于6000；
    第三解聚反应单元，来自所述第二解聚反应单元的分子量高于6000的液相物料在所述第三解聚反应单元循环进行反应，直至液相物料的分子量高于10000；
    以及用于从所述第一解聚反应单元、所述第二解聚反应单元和所述第三解聚反应单元中收集气相粗丙交酯并提纯的装置。
  26. 根据权利要求25所述的系统，其特征在于，所述第一解聚反应单元包括第一解聚反应器和第一循环罐，所述乳酸低聚物和所述解聚催化剂在所述第一解聚反应器中反应，反应后得到的液相物料进入所述第一循环罐。
  27. 根据权利要求25或26所述的系统，其特征在于，所述第二解聚反应单元包括至少一个第二解聚反应器和至少一个第二循环罐，来自所述第一解聚反 应单元的液相物料以及可选的质子化溶剂在所述第二解聚反应器中反应，反应后的液相物料进入所述第二循环罐，当所述液相物料的分子量为6000以下时，将所述第二循环罐中的液相物料循环回所述第二解聚反应器进一步反应；当所述液相物料的分子量大于6000时，将所述第二循环罐中的液相物料输送到所述第三解聚反应单元进行反应。
  28. 根据权利要求25或26所述的系统，其特征在于，所述第三解聚反应单元包括至少一个第三解聚反应器和至少一个第三循环罐，来自所述第二解聚反应单元的分子量高于6000的液相物料在所述第三解聚反应器中反应，反应后的液相物料进入所述第三循环罐，当所述液相物料的分子量为10000以下时，将所述第三循环罐中的液相物料循环回所述第三解聚反应器进一步反应；当所述液相物料的分子量大于10000时，将所述第三循环罐中的液相物料排出系统。
  29. 根据权利要求25-28中任意一项所述的系统，其特征在于，所述第一解聚反应单元、所述第二解聚反应单元和所述第三解聚反应单元中的反应器各自为刮膜解聚反应器，优选为薄膜蒸发器、分子蒸馏蒸发器或其他搅拌膜式蒸发器。
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