CN107987265B

CN107987265B - Continuous preparation method of polyester polyol

Info

Publication number: CN107987265B
Application number: CN201610956110.4A
Authority: CN
Inventors: 孙双翼; 鞠昌迅; 黎源; 华卫琦
Original assignee: Wanhua Chemical Group Co Ltd
Current assignee: Wanhua Chemical Group Co Ltd
Priority date: 2016-10-27
Filing date: 2016-10-27
Publication date: 2020-04-10
Anticipated expiration: 2036-10-27
Also published as: CN107987265A

Abstract

The invention relates to a continuous preparation method of polyester polyol. Comprises three reaction stages of esterification, flash evaporation and polycondensation. The esterification reaction is carried out at 130-260 ℃, preferably 180-240 ℃, more preferably 200-230 ℃, the esterification rate is about 50-95%, and the loss of the raw material polyol can be reduced by carrying out the esterification reaction at a higher temperature. Through the high-efficient quick micromolecules such as desorption water of flash distillation, promote reaction rate, practice thrift the production time, reduce the energy consumption. The mixture of the catalyst and the polyester polyol is used for feeding, and the hydrolysis of the catalyst can be effectively reduced from the aspects of thermodynamics and kinetics by changing the ligand, so that the using amount of the catalyst is reduced. The polyester polyol produced by the process has the advantages of low catalyst residue, low color number and stable product quality.

Description

Continuous preparation method of polyester polyol

Technical Field

The invention belongs to the field of organic high molecular compounds and preparation thereof, and particularly relates to a continuous preparation method of polyester polyol, and more particularly relates to a synthesis process of polyester polyol with low catalyst residue.

Background

Polyester polyols are one of the important raw materials for the production of polyurethanes. Black and white materials are required for the production of polyurethanes. The main component of the white material is polyol, and can be divided into polyester polyol and polyether polyol. The polyester polyurethane has the advantages of high mechanical strength, oil resistance, heat resistance, wear resistance, hydrolysis resistance, low temperature resistance, oxidation resistance and poor stability to acid and alkali compared with polyether polyurethane. Currently, polyester polyurethane is mainly used for producing polyurethane resin, polyurethane shoe sole stock solution, polyurethane rigid foam, polyurethane rubber, polyurethane screen, polyurethane casting products, polyurethane adhesives, thermoplastic polyester elastomers and the like.

Generally, polyester polyols are linear polymers prepared by the polycondensation of dibasic acids and glycols. The general division into two stages of reaction: an esterification section and a polycondensation section. The acid and alcohol in the esterification section react to generate a large amount of water, and the water is separated from a small amount of alcohol carried in the water by a rectifying column and then is removed out of the system to push the reaction to be carried out towards the esterification direction. And in the polycondensation section, residual water and a part of small molecular alcohol are removed by the system under high-temperature vacuum, so that the molecular weight is increased, and a target product is generated. The traditional polyester polyol production process is a batch process. DuPont, Bayer, Basff, etc. have applied for related patents on batch production processes of different polyester polyols, such as US2006069174, CN 1668668, DE 102006048288, CN 101168592, CN 201343512, etc. Although the intermittent production has the advantages of simple device and flexible production, and can be conveniently switched among various brands, the product quality fluctuation is large, the quality stability of subsequent polyurethane products is seriously influenced, and the intermittent production has the defects of long reaction period, low device utilization rate, high energy consumption loss caused by repeated temperature rise and temperature reduction, high cost and the like compared with the continuous production.

The dow company developed a semi-continuous process for the production of polyester polyols, as described in WO 02008037400, CN101516965, by first subjecting a polybasic acid and a polyhydric alcohol to esterification in a batch stirred tank, and then continuously passing the resulting oligomer through a fixed bed reactor for the subsequent continuous polycondensation.

The dupont company designs an upflow polyester prepolycondensation reactor (UFFP), the tray in the reactor has more dead zones and must rely on vigorous bubbling for self-cleaning, and the dupont company applies a plate-type rectifying tower to polyester prepolycondensation later, as described in patents US 5786443, US849849, CN1137278, etc. The reactor has no mechanical stirring and simple structure, but has small operation elasticity and relatively short residence time in the reactor, and is not suitable for the continuous production of polyester polyol. The university of eastern China developed up to two-stage continuous production processes based on column reactors, as described in CN 102432846 and CN 102585191. The polyester polyol synthesized by the continuous method has the advantages of stable product quality, high automation degree, short reaction time, low energy and labor cost consumption in producing products with unit mass, and the like. But the temperature control and residence time requirements of the various sections of the reactor are high.

The esterification reaction of polyester polyol is that acid and alcohol react to generate ester and water, the generated water and the entrained small molecular alcohol are separated by a rectifying tower, water is extracted from the top of the tower, and the alcohol returns to the reaction kettle from the bottom of the tower to continue the reaction. The temperature of the reaction kettle has important influence on production, and the higher the temperature is in a certain range, the more favorable the reaction and the water evaporation are. However, the traditional batch reaction process and the existing continuous reaction process are limited by the process principle, and the initial reaction temperature cannot be increased very high. In the initial stage of the existing reaction process, main materials in a reaction kettle comprise a large amount of raw material micromolecule alcohol, on one hand, the boiling point of initial reaction liquid containing a large amount of micromolecule alcohol and water produced by reaction is relatively lower than the later-stage reaction temperature (about 200 ℃), and the safety hazard of bumping is caused when the temperature is increased, on the other hand, if the temperature is too high, the gasification of the alcohol is intensified, and the excessive alcohol is extracted from the tower top along with water, so that the raw material proportion is unbalanced, and finally, the product index fluctuation and the raw material waste are caused. Therefore, the reaction temperature must be controlled at 130 ℃ to 160 ℃ in the initial stage of the batch process, and then the temperature is gradually increased after the conversion rate is more than 25% as the reaction proceeds. For example, CN201310673843.3, the reaction is carried out at the temperature of 150 ℃ and 160 ℃, then the temperature is increased by 200 ℃ and 210 ℃, and the temperature is kept for 1-2h every time the temperature is increased by 15-20 ℃; the CN201510941271.1 patent is first heated to 140 ℃ and 145 ℃ for reaction, then heated to 170 ℃ and 175 ℃ for reaction for 2-3 hours, and the final reaction temperature is 210 ℃ and 230 ℃. The process of heating up for many times has three problems, firstly, the reaction temperature in the initial stage is low, so that the time consumption is long, secondly, the process fluctuation caused by frequent heating up operation in the later stage has obvious influence on the product quality, and thirdly, the operation is complicated and the manpower is wasted. In addition, as an intermittent reaction, the temperature must be reduced after production is completed to carry out the next feeding and reaction, which causes energy waste.

The esterification reaction of the polyester polyol is a reversible reaction, and in order to promote the reaction to proceed towards the esterification direction, the key is to rapidly, efficiently and low-efficiently remove the generated byproduct, namely small molecular water or small molecular alcohol. The traditional reaction kettle accelerates the removal of small molecules by enhancing stirring or bubbling nitrogen, and has high energy consumption and limited effect.

The late stage of esterification reaction of polyester polyol usually needs to add a catalyst, and the common catalyst comprises metal organic compounds such as titanium, tin, antimony, bismuth, zirconium and the like. The metal organic compounds of titanium and tin are most widely used. The relevant documents and patents CN201510941271.1 indicate that tin-containing compounds cause blackening problems and other disadvantages when exposed to oxidation, and that tin-containing compounds have a very negative environmental impact and are extremely biocides. Titanium alkoxides are susceptible to hydrolysis and the solution to this is industrially to add large amounts of titanium alkoxide so that even after partial hydrolysis of the catalyst, the remaining catalyst is still sufficient to catalyze the reaction in an economical time, but this method brings too high a catalyst residue to the detriment of the properties of the final polyurethane product. In addition, titanium catalysts are known to easily increase the color number of products. The greater the catalyst concentration, the longer the contact time with the product, and the more marked the increase in the color number of the product.

The catalyst residue is reduced, which is beneficial to improving the product quality of the polyester polyol and improving the color number of the product. However, in the prior art, the reaction time is usually required to be greatly prolonged after the catalyst dosage is reduced, so that the production cost is increased, and the color number is increased due to the overlong reaction time. At present, the batch process for producing polyester polyol comprises two processes of adding a catalyst at the beginning of the reaction and adding the catalyst after the reaction is finished in a normal pressure section and then vacuumizing. Patent cn201210027942.x uses 0.03 wt% of catalyst and has to be added to the reactor from the beginning. CN201210484831.1 used 441ppm of tetrabutyl titanate and was added to the reaction at the beginning. Patent CN201310655647.3 uses a method of adding catalyst in the middle of the atmospheric pressure stage, and the amount of the catalyst is more than 200 ppm. The document Ind.Eng.chem.Res.2012,51,12258-12265 for the synthesis of polyester resins uses a process in which a catalyst is added after the end of the atmospheric section and then the vacuum is switched on, the catalyst being 1000ppm of tetrabutyltitanate. A two-stage continuous process such as CN 102432846 is carried out by adding the catalyst after the atmospheric pressure stage is finished.

The disadvantages of adding the catalyst immediately after the end of the atmospheric section are significant. The reaction solution after the atmospheric pressure stage still contains water, and the acid value of the reaction solution is still high, as mentioned in the CN 102432846 example, the acid value is 35-45 mgKOH/g. These residual acids continue to react with the still-formed water, and the catalyst rapidly decomposes upon contact with water. The residual moisture in the system can not be sufficiently removed by the existing process, and if only a small amount of catalyst is added, the residual catalyst after hydrolysis is too little, so that the effective catalytic reaction can not be carried out. On the other hand, the reduction in the amount of catalyst causes great difficulty in accurately weighing and adding the catalyst, which deteriorates the stability of the production process and the product index. This results in the amount of catalyst actually used in industrial production still being high.

Disclosure of Invention

The invention aims to overcome the defects of the prior art and provide a continuous preparation method of polyester polyol. The method has the advantages of low cost, simple operation, stable product quality, low residue of the obtained polyester polyol catalyst and low color number.

In order to achieve the technical effects, the invention adopts the following technical scheme:

a continuous process for the preparation of polyester polyols comprising the steps of: the method comprises three reaction stages of esterification, flash evaporation and polycondensation,

(1) under the protection of inert gas, continuously feeding polybasic acid (anhydride) and polyalcohol into an esterification kettle for esterification reaction, and evaporating water generated in the reaction; the reaction temperature is 130-260 ℃, preferably 180-240 ℃, and more preferably 200-230 ℃; the esterification rate is 50-95%, preferably 75-90%;

(2) continuously feeding the material obtained in the step (1) into a flash evaporation kettle for flash evaporation, wherein the flash evaporation temperature is 200-; the esterification rate is 90-99%, preferably 95-99%;

(3) continuously feeding the material obtained in the step (2) and the catalyst into a polycondensation kettle, and carrying out polycondensation reaction at the temperature of 200-.

The residence time of the step (1) is 2-4 h.

The reaction pressure of the step (1) is normal pressure-0.4 Mpa.

The residence time of the step (2) is 0.5 to 2 hours.

The residence time of the step (3) is 2-6 h.

In the esterification reaction stage of the present invention, the esterification reactor is in a continuous feeding and withdrawing state, and the material inventory in the esterification reactor is about 60-360 times of the feeding amount per minute, preferably about 120-240 times. The main material in the kettle is esterification product with esterification rate of about 50-95%. On the one hand, the majority of the starting polyol is consumed and is present in a much lower proportion than when the batch process is started. On the other hand, the polyester intermediate product generated by the reaction of the polyhydric alcohol and the polybasic acid (anhydride) has a very high boiling point, and the boiling point of the reaction liquid is obviously improved. Therefore, the invention can carry out esterification reaction at higher temperature without the danger of bumping and the influence of gasification of a large amount of raw material polyol, thereby shortening the reaction time and reducing the production cost.

A flash evaporation process is added between the esterification reaction and the polycondensation reaction, so that small molecules are removed more efficiently and rapidly, the reaction speed is increased, the production time is saved, and the energy consumption is reduced.

The water in the reaction liquid after the reaction in the normal pressure section can be removed efficiently and quickly by using the flash evaporation technology, the reaction liquid is promoted to continue to react in a high-temperature vacuum environment, the acid value of the system is reduced to be less than or equal to 20mgKOH/g, and therefore the amount of water generated in the subsequent reaction is reduced. And the possibility of catalyst hydrolysis is effectively reduced in subsequent reactions, so that the low-concentration catalyst can be used for achieving the reaction speed which can be achieved by the prior process by using the high-concentration catalyst.

The water generated by the reaction in the step (1) is distilled and extracted, and the extraction temperature is controlled to be 96-104 ℃, preferably 98-102 ℃.

The material obtained in the step (1) can be preheated and pressurized by nitrogen before being subjected to flash evaporation, so that the flash evaporation effect is improved. The preheating temperature is 200-260 ℃, preferably 200-230 ℃. Nitrogen flow (in standard/hour) as feed flow (in m)³Calculated as a/h) of 0.1 to 1000 times, preferably 1 to 30 times.

The polybasic acid (anhydride) is selected from one or more of C2-C12 dibasic, tribasic and tetrabasic acids and corresponding acid anhydrides, preferably selected from one or more of succinic acid, glutaric acid, adipic acid, sebacic acid, dodecanedioic acid, phthalic acid, isophthalic acid, terephthalic acid, trimellitic acid and phthalic anhydride.

The polyhydric alcohol is selected from one or more of C2-C12 dihydric, trihydric and tetrahydric alcohols, preferably one or more of ethylene glycol, diethylene glycol, 1, 2-propanediol, 1, 3-butanediol, 1, 4-butanediol, 1, 5-pentanediol, 1, 6-hexanediol, neopentyl glycol, 3-methyl-1, 5-pentanediol, glycerol, trimethylolpropane and pentaerythritol.

The molar ratio of the polycarboxylic acid (anhydride) to the polyol in the present invention is 1:1.05 to 1:1.5, which varies depending on the molecular weight (or hydroxyl value) of the objective product. And (4) if the product index is adjusted in the step (3), strict requirements on the feeding quality fluctuation in the step (1) are not made. If no adjustment of the product index is carried out in step (3), the fluctuation in the feed quality in step (1) is required to be controlled to within 0.3 wt%, preferably within 0.1 wt%, from the viewpoint of production stability.

The catalyst of the invention is selected from one or more of titanium, tin, antimony, zirconium and bismuth based catalysts, and preferably comprises one or more of isopropyl titanate, tetrabutyl titanate, dibutyltin dilaurate, stannous octoate and bismuth laurate.

The catalyst according to the invention is used in an amount of 2 to 20ppm, preferably 9 to 15ppm, relative to the total weight of the charge. The total weight of the fed materials is the total weight of the polyalcohol and the polybasic acid (anhydride).

The catalyst can be directly added into a polycondensation kettle or added into a connecting pipeline between a flash evaporation kettle and the polycondensation kettle.

The timing of the addition of the catalyst is an important factor for the present invention, and the system must have sufficiently reacted when the catalyst is added and the water produced by the reaction has sufficiently been removed from the reaction system. It is recommended that when the catalyst is added, the acid value of the material obtained in step (2) is not more than 20mgKOH/g, preferably not more than 15mgKOH/g, more preferably not more than 10 mgKOH/g.

The catalyst in the step (3) of the present invention may be added in the form of a pure substance or in the form of a mixture of the catalyst and the polyester polyol. Preferably, the catalyst of the present invention is mixed with a polyester polyol to prepare a mixture, and the catalyst is added to the reactor in the form of the mixture.

The polyester polyol mixed with the catalyst can be any polyester polyol available, preferably the starting material for its preparation is selected from the group consisting of the acid and alcohol used in the catalyst to prepare the corresponding product. The hydroxyl number of the polyester polyols used is preferably close to or slightly higher than that of the product, although polyester polyols having a lower hydroxyl number than that of the product may also be used. On the premise of not influencing the quality and the service performance of subsequent products, polyester polyol with a structure different from that of the products can also be used. For example, a product of dodecanedioic acid/hexanediol, a polyester polyol (less expensive and readily available) using adipic acid/butanediol is mixed with a catalyst.

More preferably, the polyester polyol mixed with the catalyst is a polyester polyol having the same structure as the product to be prepared.

The weight ratio of the catalyst to the polyester polyol in the mixture of the catalyst and the polyester polyol in the step (3) of the present invention is 1:2 to 1:50, preferably 1:5 to 1: 20. The preparation method comprises the steps of heating and melting the polyester polyol under the protection of inert gas according to the proportion, adding the catalyst, and stirring and mixing at high speed. The mixture can be used directly in liquid state, or cooled to solid state and cut for use. The mixture needs to be preserved without exposure to moisture. Under sealed preservation, the catalytic activity of the mixture solid is not obviously reduced after being preserved for 1 year.

In step (1), water produced by the reaction and the polyol as a raw material need to be separated by a rectifying column. The alcohol content in the fraction extracted from the tower top can be effectively reduced by controlling the temperature of the tower top at 98-102 ℃. In the case where the hydroxyl value of the product is not critical or the hydroxyl value is adjusted in step (3), the overhead temperature may be controlled to 96 to 104 ℃.

In each step of the reaction, inert gas can be continuously introduced to take away water generated by the reaction, which is beneficial to quickly finishing the reaction, but the problem of increasing the loss of raw material polyol is also solved. The amount of inert gas introduced is measured from the viewpoint of economy.

Can set up the sampling point in flash tank discharging tube and carry out the analysis of taking a sample, carry out product index's adjustment under necessary condition: when the hydroxyl value is too small, the calculated amount of alcohol is added in the polycondensation kettle, when the hydroxyl value is too large, the retention time of the product in the polycondensation kettle is prolonged, and vacuum is enhanced or nitrogen bubbling is carried out to take away additional alcohol.

And setting sampling points in a discharge pipeline of the polycondensation kettle for sampling analysis, and selecting the flow direction of a product according to an analysis result. The product can be beaten to the packaging process and carry out ejection of compact packing, can beat the polycondensation cauldron back again and carry out appropriate adjustment: when the hydroxyl value is too small, supplementing calculated amount of alcohol into the polycondensation kettle, when the hydroxyl value is too large, prolonging the retention time of the product in the polycondensation kettle, and optionally enhancing vacuum or carrying out nitrogen bubbling to take away additional alcohol; when the acid value is too large, the residence time of the product in the polycondensation vessel is prolonged.

The key element to avoid catalyst hydrolysis is to reduce the water content of the reaction system when the catalyst is added. In the prior art, the catalyst is added at the beginning or in the middle of the atmospheric section or before the vacuum section. In this case, the amount of the catalyst used is large because the water content of the reaction system is high or a large amount of water is produced during the reaction. Particularly, when the catalyst is added in the middle of the normal pressure section or before the vacuum section is started, the phenomenon that white mist is emitted out of the reaction kettle when the catalyst is added can be obviously observed, and the catalyst on the needle head or the bottle opening is immediately whitened after contacting the white mist. In the invention, the reaction system is firstly vacuumized, and most of water is removed. The reaction rate is reduced in the later stage of the reaction, the generated water is little, and the catalyst is added under the condition that the generated water can be effectively pumped away by a vacuum system, so that the catalyst hydrolysis is greatly reduced.

In order to minimize contact and reaction between the catalyst and water, in the present invention, the catalyst is mixed with the polyester polyol and then added in the form of a solid or liquid. The occurrence of hydrolysis reaction is reduced from three aspects. Firstly, the catalyst has a coordination effect with polyester polyol or residual carboxyl, the periphery of the coordinated catalyst is surrounded by long-chain product molecules, molecular ligands with small steric hindrance are greatly increased, and the reaction is not favorable for water molecule attack from the aspect of dynamics. Secondly, the coordinated polyester polyol has a larger molecular weight than a small molecule alcohol ligand (such as a butanol-based ligand, an isopropanol-based ligand), is more difficult to leave than a small molecule, and is difficult to participate in a hydrolysis reaction interchanged with water. Finally, because the catalyst is wrapped by the polyester polyol, the polyester polyol is solid or high-viscosity liquid in a feeding state, the diffusion of residual water vapor in the reaction kettle can be greatly reduced, the residual water vapor contacts the wrapped catalyst, and the reaction does not occur without contact.

The invention has the beneficial effects that: in the esterification reaction stage, compared with the traditional process, the method greatly reduces the proportion of raw material polyol in the reaction liquid, reduces the loss of the polyol, greatly improves the reaction temperature and shortens the reaction time. The reaction condition is more stable, and the product quality stability is more guaranteed. By mixing and feeding the catalyst and the polyester polyol, the using amount of the catalyst is reduced, the contact time of the catalyst and water is reduced, the obtained product has low catalyst residue, the metal ion residue of the catalyst is less than or equal to 4ppm, and the metal ion residue of the product obtained by the conventional process is usually 5-20 ppm; the color number is less than or equal to 30, the acid value can reach below 1mgKOH/g, and the product quality is stable.

Detailed Description

The technical solution of the present invention is illustrated by the following specific examples, and specific embodiments and specific operation procedures are given, but the scope of the present invention is not limited to the following examples.

In the invention, the hydroxyl value and the acid value are measured by using a Mettler potentiometric titrator, the Ti element residue is measured by using ICP, and the color number is measured by comparing a colorimetric tube visual colorimetry with a standard colorimetric card.

Example 1

Uniformly mixing adipic acid and 1, 4-butanediol in a slurry tank according to the mass ratio of 1.45:1, adding about 0.75kg of mixture into a 5L esterification kettle per hour under the continuous and stable condition, maintaining the esterification kettle at normal pressure, controlling the reaction temperature to be about 180 ℃, controlling the temperature at the top of the tower to be 96-100 ℃, and controlling the material in the kettle to be 1.5L. The water collecting speed is 0.098kg/h, which is converted into the esterification rate of 90%. The materials at the bottom of the esterification kettle are continuously led into a flash evaporation kettle of 10L. The reaction in a 10L flash evaporation kettle is maintained at a pressure of 18kPa and a temperature of about 200 ℃, the volume of the materials in the kettle after stabilization is about 1L, and the acid value is about 12 mgKOH/g. The material at the bottom of the flash evaporation kettle is continuously introduced into a 10L polycondensation kettle, 16ppm isopropyl titanate catalyst (0.012g/h) is added, the reaction temperature is maintained at about 200-210 ℃, the reaction pressure is maintained at 200Pa, the material volume in the kettle is about 3L under a stable condition, and the polycondensation kettle is continuously discharged. Sampling to obtain the product with acid value of 0.2-0.6mgKOH/g, hydroxyl value of 53-58mgKOH/g, color number of 20 and Ti residue of 2.5-3.7 ppm.

Example 2

Mixing adipic acid, neopentyl glycol and 1, 6-hexanediol according to the proportion of 1: 0.246: the mass ratio of 0.682 is mixed evenly in the slurry tank. Under the continuous and stable condition, about 0.75kg of mixture is added into a 5L esterification kettle per hour, the esterification kettle maintains normal pressure, the reaction temperature is about 230 ℃, the temperature at the top of the esterification kettle is controlled to be 98-100 ℃, and the volume of the materials in the esterification kettle is about 2.3L. The water collecting speed is 0.077kg/h, which is equivalent to 80 percent of esterification rate. The materials at the bottom of the esterification kettle are continuously led into a 10L flash evaporation kettle after being preheated to 230 ℃. The reaction in a 10L flash evaporation kettle is maintained at a pressure of 10kPa, a reaction temperature of about 230 ℃ and 245 ℃, the volume of the materials in the kettle is 1.5L after stabilization, and the acid value is 18 mgKOH/g. The material at the bottom of the flash evaporation kettle is continuously introduced into a 10L polycondensation kettle, a mixture of isopropyl titanate and poly adipic acid-1, 6-hexanediol glycol (hydroxyl value is 78mgKOH/g) in a ratio of 1:20 is continuously added, 0.012g/h equivalent to the addition of isopropyl titanate is added, the reaction temperature is maintained at about 230-. And continuously discharging materials from the polycondensation kettle. Sampling to obtain the product with acid value of 0.4-0.8mgKOH/g, hydroxyl value of 73-79mgKOH/g, color number of 20 and Ti residue of 2.5-3.7 ppm.

Example 3

Mixing dodecanedioic acid and 1, 6-hexanediol according to the proportion of 1: the mass ratio of 0.565 is mixed uniformly in the slurry tank. Under the continuous and stable condition, about 0.75kg of mixture is added into a 5L esterification kettle per hour, the esterification kettle maintains normal pressure, the reaction temperature is 200 ℃, the temperature at the top of the esterification kettle is controlled to be 98-100 ℃, and the volume of the materials in the esterification kettle is about 1.5L. The water collecting speed is 0.071kg/h, and the esterification rate is 95 percent. The materials at the bottom of the esterification kettle are continuously led into a flash evaporation kettle of 10L. The reaction in a 10L flash evaporation kettle is maintained at a pressure of 10kPa and a reaction temperature of about 200 ℃ and 215 ℃, the volume of the materials in the kettle is 1.5L after stabilization, and the acid value is 5.6 mgKOH/g. Continuously introducing the bottom material of the flash evaporation kettle into a 10L polycondensation kettle, continuously adding a mixture of isopropyl titanate and 1:10 poly (adipic acid-1, 6-hexanediol glycol) (hydroxyl value is 56mgKOH/g), wherein 0.012g/h is equivalent to adding isopropyl titanate, maintaining the reaction temperature at about 200-215 ℃, the reaction pressure at 100-150Pa and the volume of the material in the kettle at a stable state at about 1.5-3L. And continuously discharging materials from the polycondensation kettle. Sampling to obtain the product with acid value of 0.2-0.6mgKOH/g, hydroxyl value of 28-34mgKOH/g, color number of 20 and Ti residue of 2.5-3.7 ppm.

Comparative example 1

2.489kg of adipic acid, 1.718kg of 1, 4-butanediol and 200mg (47ppm) of isopropyl titanate were weighed into a 5L stainless steel reaction vessel equipped with an oil bath jacket, a stirrer, a gas inlet, and a rectification column. Vacuumizing, and introducing nitrogen for three times. The reaction temperature was 161 ℃. Timing is started when the temperature at the top of the tower rises, and the distillate is continuously extracted, wherein the temperature at the top of the tower is controlled to be 98-100 ℃. After 2 hours of reaction, the temperature was raised to 200 ℃ in half an hour and the reaction was continued at 200 ℃ for a total time of 5.5 hours. The system pressure was reduced below 1kPa by an oil pump and the reaction was continued at 200 ℃ for 12 hours. The acid value of the product is 0.33mgKOH/g, the hydroxyl value is 56.68mgKOH/g, the color number is 40, and the Ti residue is 5.8 ppm.

Comparative example 1 has a longer reaction time than the examples. Comparative example 1 has problems that the product color number is too large and Ti residue is high.

Claims

1. A continuous process for the preparation of polyester polyols comprising the steps of: the method comprises three reaction stages of esterification, flash evaporation and polycondensation,

(1) under the protection of inert gas, continuously feeding one or more of polybasic acid and corresponding anhydride and polyalcohol into an esterification kettle for esterification reaction, and evaporating water generated in the reaction; the reaction temperature is 180-240 ℃; the esterification rate is 50-95%;

(2) continuously feeding the material obtained in the step (1) into a flash evaporation kettle for flash evaporation, wherein the flash evaporation temperature is 200 ℃ and 260 ℃, and the pressure is less than or equal to 50 kPa; the esterification rate is 90-99%;

(3) continuously feeding the material and the catalyst obtained in the step (2) into a polycondensation kettle, and carrying out polycondensation reaction at the temperature of 200-;

the dosage of the catalyst is 2-20ppm relative to the total weight of the feed;

the acid value of the material obtained in the step (2) is less than or equal to 20 mgKOH/g.

2. The method as claimed in claim 1, wherein the reaction temperature in step (1) is 200-230 ℃; the esterification rate is 75-90%;

the flash evaporation temperature in the step (2) is 200 ℃ and 230 ℃, and the pressure is less than or equal to 20 kPa; the esterification rate is 95-99%;

the step (3) is carried out at the temperature of 200 ℃ and 230 ℃ and under the reaction pressure of 5kPa or less.

3. The method according to claim 2, wherein the pressure in the step (2) is less than or equal to 10 kPa; the reaction pressure in the step (3) is less than 0.2 kPa.

4. The process according to claim 1, wherein the polyacid and the corresponding anhydride are selected from one or more of the di-, tri-, tetra-and corresponding anhydrides of C2-C12.

5. The method according to claim 4, wherein the polyacid and the corresponding anhydride are selected from one or more of succinic acid, glutaric acid, adipic acid, sebacic acid, dodecanedioic acid, phthalic acid, isophthalic acid, terephthalic acid, trimellitic acid, phthalic anhydride.

6. The method of claim 1, wherein the polyol is selected from one or more of the group consisting of C2-C12 dihydric, trihydric and tetrahydric alcohols.

7. The method according to claim 6, wherein the polyol is selected from one or more of ethylene glycol, diethylene glycol, 1, 2-propanediol, 1, 3-butanediol, 1, 4-butanediol, 1, 5-pentanediol, 1, 6-hexanediol, neopentyl glycol, 3-methyl-1, 5-pentanediol, glycerol, trimethylolpropane, pentaerythritol.

8. The process of claim 1 wherein the catalyst is selected from one or more of the group consisting of titanium, tin, antimony, zirconium and bismuth based catalysts.

9. The method of claim 8, wherein the catalyst is selected from one or more of isopropyl titanate, tetrabutyl titanate, dibutyltin dilaurate, stannous octoate, and bismuth laurate.

10. The process according to claim 1, characterized in that the catalyst is used in an amount of 9 to 15ppm relative to the total weight of the charge.

11. The method as claimed in claim 1, wherein the material obtained in step (2) has an acid value of 15mgKOH/g or less.

12. The method as claimed in claim 11, wherein the material obtained in step (2) has an acid value of 10mgKOH/g or less.

13. The method of claim 1, wherein the catalyst in step (3) is added to the reactor in the form of a mixture of the catalyst and the polyester polyol.

14. The method of claim 13, wherein the weight ratio of catalyst to polyester polyol in the mixture of catalyst and polyester polyol is from 1:2 to 1: 50.

15. The method of claim 13, wherein the weight ratio of catalyst to polyester polyol in the mixture of catalyst and polyester polyol is from 1:5 to 1: 20.

16. The method as claimed in claim 1, wherein the material obtained in step (1) is preheated at a temperature of 200-260 ℃ before flash evaporation; pressurizing the material obtained in the step (1) by using nitrogen, wherein the nitrogen flow is 0.1-1000 times of the material flow, the nitrogen flow is measured in standard square/hour, and the material flow is measured in m³And (h) calculating.

17. The method as claimed in claim 16, wherein the material obtained in step (1) is preheated at a temperature of 200 ℃ to 230 ℃ before the flash evaporation; pressurizing the material obtained in the step (1) by using nitrogen, wherein the nitrogen flow is 1-30 times of the material flow, the nitrogen flow is measured in standard square/hour, and the material flow is measured in m³And (h) calculating.