MXPA99010150A - Pressure polymerization of polyester - Google Patents

Abstract

A process is disclosed for the preparation of a polyester polymer or polyester copolymer under superatmospheric pressure conditions in a pipe or tubular reaction under turbannular flow conditions. Reaction material having a glycol equivalents to carboxylic acid equivalents mole ratio of from 1.0:1 to 1.2:1, together with a superatmospheric dense gaseous medium are fed co-currently to the reactor. Dicarboxylic acid and /or diol raw materials may be injected into any of the reaction zones in the process during operation to achieve the overall desired mole ratio balance. The process operates at temperatures of from about 220 DEG C to about 320 DEG C, with turbannular flow achieved before the polymer product and gas exit the reactor process. The pressure in the reaction zones can be in the range from 15 psia to 2500 psia ( SIMILAR 103-17238 kPa). A polymer product having a DP of a greater than 40, more preferably at least about 70, is achieved by the transfer of water from the reacting ma terial polymer melt to the gaseous medium in the reactor.

Description

POLYMERIZATION BY POLYESTER PRESSURE BACKGROUND OF THE INVENTION 1) FIELD OF THE INVENTION The invention pertains to the field of polyester synthesis from mixtures of monomer and / or oligomer of hydroxyalkyl dicarboxylic acid ester and the polymerization thereof to form a polyester. In particular, the present invention is a polyesterification process that preferably utilizes substantially monohydroxyalkyl ester monomer as monohydroxyethyleneterephthalate, with preferably little or no bishydroxyalkyl ester monomer. The invention also relates to the field of synthesis of polyester from traditional starting materials of dicarboxylic acid and diol, said starting materials reacting to form mixtures of monomer and / or oligomer of hydroxyalkyl dicarboxylic acid ester (mainly monohydroxyalkyl ester monomer with little or no ester monomer bishydroxyalkyl) and polyesterifying these to form a polyester. This reaction process and the aforementioned reaction procedure occur in a tube or tube reactor under pressure its peratmosphere where a turbanular flow occurs. 2) PREVIOUS TECHNIQUE The manufacture of polyester is practiced globally, and a variety of methods are taught. The direct esterification of a dicarboxylic acid with a diol, for example terephthalic acid (TA) and ethylene glycol (EG) form a reactive monomeric material that releases water. The monomeric material is known to contain the monohydroxyethyl terephthalate, bishydroxyethyl terephthalate, and longer chain length oligomers of the same structural type having an average polymerization (DP) degree of 1 to 6. The DP is further increased by polymerization of melting of the monomeric material under vacuum conditions, a process referred to as polycondensation. A polymerization process is described in the U.S. Patent. No. 3,480,587 to Porter where at least part of the polycondensation takes place while the liquid reaction mixture flows in a long, narrow tube with an inert gas as nitrogen in a two-phase flow regime called turban flow. The reaction mixture entering the tube has an average DP of 27 to 40, and the product leaving the tube has an intrinsic viscosity ratio between 1.7 and 2.0 (measured as a 1% solution in orthochlorophenol under standard conditions) corresponding to DP of 65-100. Porter teaches that in a pressure procedure that uses an inert gas "the ratio of the cross-sectional area of the tube divided by the length of the wetted perimeter should be less than 2.5 cm." To maintain this requirement for a tubular reactor, the diameter of the reactor should not exceed 11.6 cm. As taught by Porter, a high weight ratio of gas to polymer is required for polymerization to proceed. High gas velocities are required to achieve a high weight ratio in the reactor tube. If a reaction material has a DP of 20 or less, particularly or less, extremely high and impractical gas velocities would be required to achieve a practical increase in DP under turban flow.

As taught by Porter, then, it would be economically unfeasible to manufacture polyester from lower DP reaction materials in a tubular reaction zone under turban flow. The Patent of E.U.A. 5,434,239 to Bhatia discloses an atmospheric pressure procedure for the continuous production of polyester through a bishydroxyethylterephthalate (BHET) smelter or its low molecular weight oligomer. The BHEF is intimately contacted with nitrogen gas flowing countercurrent to the smelter, to facilitate polymerization and removal of the volatile reaction byproducts. A degree of polymerization of 15 to 30 is achieved in the prepolymer stage. The DP is further increased to 50-100 DP in the finishing stage with a countercurrent flow of nitrogen. Polymerization occurs in both stages without resorting to a vacuum. Bhatia teaches that nitrogen velocity is critical to the success of the procedure, and that the nitrogen velocity should be between 0.61 and 1.52 m / sec. The countercurrent flow procedure that Bhatia describes does not include operation in the turban flow regime, which by definition is a co-current procedure. WO 96/22318 to Iwasyk et al. Describes a multi-step process for producing polyester oligomers without vacuum. In the first step a polyol is added to material supplied with esterified oligomer in a pipe reactor. The inert gas is injected into the oligomeric product at the end of the first stage, to bring the oligomeric material to the following stages. The inert gas also serves to provide a pressure drop along the tubular reactor, which aids in the removal of volatile reaction byproducts. The amount of inert gas used in the Iwasyk and other process is less than .908 kg per .454 kg of oligomer, and the flow rate in the tubular reactors is not turbanular. It is taught that a prepolymer with a DP of 2 to 40 and a balance of carboxyl group to final hydroxyl group between 1: 2 and 1: 8 occurs at the exit of the final stage of the pipe reactor. It would be desirable to efficiently and economically obtain a polymer with relatively high DP (40 or more) with a carboxyl to hydroxyl end group ratio that prevents premature interruption of the polymerization. WO 96/39456 to DeSimone describes a transesterification-polycondensation process for polyester using carbon dioxide as the polymerization medium. The polycondensation of BHET conducted in an autoclave with flowing supercritical carbon dioxide is exemplified. Ethylene glycol is released for each step of growth in DP. An excess of 10 - 50 mole percent glycol is recommended. The highest DP achieved in the examples is 33. As illustrated by DeSimone, the excess diol is removed by flowing carbon dioxide or by including a carbon dioxide surfactant capable of purifying the condensate inside the reactor without removing the carbon dioxide . Higher DP increases would be desirable without requiring the purification of the diol to remove the surfactant. In addition, as demonstrated by DeSimone, a DP of 33 was achieved in a continuous flow procedure starting with BHET. However, this required a relatively large amount of carbon dioxide, that is, a weight ratio of carbon dioxide to BHET of 43: 1. In general, it is known that dicarboxylic acids and diols react with the removal of water to form polyesters, which under favorable conditions will increase in polymer chain length. More specifically, with respect to conventional esterification of polyester monomers, the degree of polymerization (DP) obtained is a function of operating pressure for the repeating unit: where R is contributed by the diol, which for polyethylene terephthalate (PET) is -CH2CH2-, and n is the degree of polymerization. The DP of the oligomer is determined by dividing the number average molecular weight by the molecular weight of the repeating unit, which for PET is 192. By characterizing the mole percentage of components of a reaction material that may contain monomeric and oligomeric components, The average DP number of any oligomer must be determined to find the molecular weight of the oligomers.

BRIEF DESCRIPTION OF THE INVENTION A method using a gaseous compound has been discovered as the reaction medium wherein the degree of chain growth occurs beyond a DP of 40, in a super atmospheric pressure process using reaction materials that preferably have a molar ratio from glycol equivalents to carboxylic acid equivalents from 1: 1 to 1.2: 1, considering all possible points of reagent additions to the process, and surprisingly without added catalyst at certain pressures and temperatures. According to a fundamental aspect of the invention, the polymerization at super atmospheric pressure (> 1.0545 kg / cm2 absolute) to form polyester starting from a reaction material having a DP of 1 is provided in one or more polymerization steps. to 40, preferably 1 to 20, said process consists in contacting the reaction material with an atmospheric low temperature boiling (gaseous) compound (<100 ° C), with the polymerization reaction involving water. According to another aspect of the invention, a continuous process for polymerizing polyester precursors in the presence of a gaseous compound having a boiling point of low temperature atmospheric is provided in two or more stages using elongated tubular reaction zones ( < 100 ° C). The reaction material is supplied to the first zone, which operates from 220 ° C to 320 ° C, which has an average DP of 1 and consists of 50-100 mole percent monohydroxyalkyl ester of a dicarboxylic acid and 0- 50 mole percent bishydroxyalkyl ester of a dicarboxylic acid. In any subsequent zone the reaction product of the previous zone is further reacted. Preferably the ratio of the cross-sectional area of the zones divided by the moistened perimeter is greater than 2.5 cm. According to another aspect of the invention, in a preferred embodiment, a process for polymerizing polyester reaction material of an average DP of 1 to 10 is provided to a polyester product having an average DP of greater than 40, and preferably larger than 90, which has no metal catalyst. Such catalysts may include, for example, metal oxides and other suitable compounds formed of, but not limited to, metals such as antimony, titanium, tin or germanium.

According to another aspect of the invention, there is provided a process for reacting the traditional starting materials of dicarboxylic acid and diol- which by definition has a DP of zero- to form mainly monohydroxyalkyl ester with the total reaction products having a DP from 1 to 40 and preferably 1 to 20, and further polyesterifying the reaction products to form a polyester.

DETAILED DESCRIPTION OF THE INVENTION The present invention is directed to an esterification polymerization process for increasing the DP of supplied reaction material having an average DP from 1 to 40, and increasing the DP. However, it is preferred that the reaction material have a DP of 1 to 20, and that the product have a DP of at least 70, preferably 90, and more preferably 180. The main polymer chain growth mechanism is polyesterification and , on a molar basis, yields an excess of not more than 20 mole percent equivalents of diol or glycol, based on moles of dicarboxylic acid supplied. The reaction materials of the present invention are characterized by a molar ratio of glycol equivalents to carboxylic acid equivalents in the range of 1.0: 1 to 1.2: 1, considering all possible points of reagent additions to the process, preferably 1 : 1 to 1.1: 1, wherein the alkyl group can be ethyl as monohydroxyethyl terephthalate- (MHET): also propylene, butylene, etc. or a mixture thereof to form copolymers; or bishydroxyalkyl terephthalate wherein the alkyl may be the same as above, that is, for ethyl the monomer is bishydroxyethyl terephthalate (BHET): also propylene, butylene, etc., or mixtures thereof when the copolymers are being formed. When bishydroxyalkyl terephthalate is used, it can not be present in an amount greater than the mole percentage of the monohydroxyalkyl terphthalate. Because bishydroxyalkyl terephthalate has two (2) glycol equivalents, and one equivalent of dicarboxylic acid, it may be necessary to add dicarboxylic acid so that the reaction materials have a molar ratio of glycol equivalents to carboxylic acid equivalents on the scale from 1.0: 1 to 1.2: 1, considering all possible points of reagent additions to the procedure. For example, if the molar ratio is 1: 1 then the reaction material can be 100% monohydroxyalkyl terephthalate (which has 1 glycol equivalent and 1 carboxylic acid equivalent); or the reaction material may be 50 mole percent monohydroxyalkyl terephthalate, and 25 mole percent bishydroxyalkyl terephthalate and mole percent dicarboxylic acid. If the desired molar ratio of glycol equivalents to carboxylic acid equivalents is 1.2 for example, the reaction material can be substantially monohydroxyalkyl terephthalate with only enough bishydroxyalkyl terephthalate to yield the desired molar ratio of 1. 2. The reaction material may also consist of a majority of monohydroxyalkyl terephthalate (such as 60 mole percent) with less than 40 mole percent being bishydroxyalkyl terphthalate, and only enough dicarboxylic acid to obtain the desired molar ratio of 1.2. Similar variations of reaction materials can also be used, where beneficial, when it is desired to have a starting molar ratio of glycol equivalents to carboxylic acid equivalents of less than 1.0: 1, for example 0.92. It is well known to those skilled in the art that the polymerization reaction between a diol and dicarboxylic acid, selected from all materials named subsequently, will not proceed to high DP if an excess of the acid species is present through the process. Consequently in this invention the reactive materials are added at some later point in order that the total molar ratio of glycol equivalents to carboxylic acid equivalents, considering all possible points of addition, is at least 1.0: 1. In the given example where the reaction materials have a molar ratio of glycol equivalents to carboxylic acid equivalents of 0.92: 1, an adjustment must then be made where 0.08 molar equivalents of glycol are added to the process. Under the correct operating conditions the reaction materials undergo a polyesterification reaction in the presence of a gaseous compound to produce polyester, with the evolution of water, small amounts of glycol (especially when the molar ratio is greater than 1.0: 1), and some amount of unreacted starting reaction material. In turban flow, gas flows along the core of the tubular reactor, while polyesterification occurs in the polymer, flowing along the reactor wall. The gas flowing in the center of the tube allows the removal of polyesterification byproducts, and any byproduct resulting from the polycondensation, of the molten polymer. This occurs on the surface of the molten polymer by diffusion in the gas phase. Suitable dicarboxylic acids include but are not limited to: oxalic, malonic, succinic, glutaric, adipic, pimelic, suberic, azelaic, sebacic, maleic, fumaric, phthalic, isophthalic, terephthalic, those naphthalene derivatives, anthracene, anthraquinone, biphenyl, and hemimelítico, or a mixture of them. Additionally, dicarboxylic anhydrides can be used in this process. Such suitable compounds include, but are not limited to: succinic anhydride, maleic anhydride, phthalic anhydride, and those anhydrides arising from dicarboxylic acids derived from naphthalene, anthracene, anthraquinone and biphenyl, or a mixture thereof. Suitable materials that provide glycol equivalents (such as diols and glycols) include but are not limited to: ethylene glycol, 1,3-propanediol, 1,4-butanediol, cyclohexyldimethanol, bisphenol-A and hydroquinone, or a mixture thereof. The tube or tubular reactor (s) contemplated for use in the process of the present invention must be operated at a sufficient temperature with which the reaction materials are pumpable. For the reaction materials set forth herein the temperature is in the range of 220 ° C to 320 ° C. Operation on this scale is necessary to melt the reaction materials, without degradation caused by higher temperatures. Additionally, operation at super atmospheric pressures is generally limited only at the upper end by operating costs. For some conditions, the higher the pressure the more DP increases that can be obtained. However, a pressure reactor operating above absolute 175.75 kg / cm2 is more expensive to operate, and also has a higher capital equipment cost. The preferred pressure scale for this process is from 7.03 to 175.75 kg / cm2 absolute and more preferably from 7.03-70.3 kg / cm2 absolute. The polyester polymers produced according to the methods of the present invention include polyester homopolymers, or as in the embodiment wherein one or more comonomers are used in combination with polyester-forming monomers, the resulting polyester polymers can be copolymers. The comonomers used as reaction materials in the method of the present invention may be any of a wide variety of comonomers conventionally used for the production of useful copolyesters. Suitable reaction materials capable of copolymerization with the reaction materials of the present invention to produce copolymers according to the process of the present invention The invention include but are not limited to copolyesters such as those conventionally based on: terephthalic acid / isophthalic acid / ethylene glycol; anthracene dicarboxylic acid / terephthalic acid / ethylene glycol; terephthalic acid / isophthalic acid / cyclohexyldimethanol; hydroxybenzoic acid / terephthalic acid / bisphenol-A; terephthalic acid / hydroxybenzoic acid / hydroquinone; terephthalic acid / hydroxybenzoic acid / naphthalenedicarboxylic acid / hydroquinone; terephthalic acid / ethylene glycol / 1,3 propanediol; and terephthalic acid / ethylene glycol / 1,4-butanediol. The compressed gaseous medium is selected from the group gases having boiling point temperatures of atmospheric pressure (at sea level) that are lower than the boiling point of water. This is critical for separating the gas from the polymer at the outlet of each stage while preventing the dispersion or re-entry of the polymer, and for removing the water from the reaction of the gaseous medium. Another requirement of the gaseous medium is the inertness to eliminative reactions with the polymer. The gaseous medium must also be stable at polymerization temperatures and form little or no degradation product. Illustrative gaseous media that exhibit these characteristics include but are not limited to ethers, ketones, carbon dioxide, and nitrogen, or mixtures thereof. Although it is preferred to operate the process of the present invention without catalysts known in the art, those skilled in the art may wish to achieve a balance between increased throughput (using conventional catalysts) and polymer purity (without using catalysts). The scope of the present invention covers both procedures, however, the examples are without catalyst because that is unconventional. The following working modes are simply designed as examples of the invention and do not define limits on the scope of the invention.

EXPERIMENTS OF 100 ml REACTOR The reactor used in this set of experiments was a Parr stainless steel stirred reactor with an internal volume of 100 ml. A stirring rod and thermocouple were placed through holes drilled in the upper part of the reactor. The reactor was heated with an aluminum heater block, and the heater and reactor interior temperatures were monitored.

The reactor vessel and components were cleaned and dried with methanol between the experiments. The reaction materials (MHET, BHET, and TA) listed in the following tables were weighed, added to the container, and agitated. The TA listed in the table was present in the MHET reaction material as an impurity. The level of TA in the MHET reaction material was found to be 3% by weight, based on an analysis of several MHET samples used in the experiments. BHET was therefore added to the MHET / TA reaction material in order to balance the molar ratio of glycol equivalents to carboxylic acid equivalents, recognizing TA impurity. The reactor was sealed and evacuated for 10 minutes. An initial amount of approximately 13.4 gm of diethyl ether (DME) was added through a tap port. Different weight ratios of DME to the reaction material were used during this set of experiments. With the temperature of the heating block set at 310 ° C, the internal temperature of the reactor reached 260 ° C. At 260 ° C, additional DME was added to bring the pressure in the reactor to the desired pressure. (The total amount of DME added to the reactor for each experiment is shown in Table I). The experiment started when the desired set point pressure and temperature were achieved. The experiment was completed after the desired reaction time by turning off the agitator and heater, removing the container from the heat source, avoiding further reaction by rapidly cooling, and then allowing the reactor to cool below 60 ° C. The reactor was then depressurized, opened and the product removed. The degree of polymerization (DP) achieved during an experiment was determined by GPC (gel petion chromatography) analysis of the product. A 10 mg sample was dissolved in 1 ml of (CF3) 2CHOH, (HFIP-hexafluoroisopropanol), filtered, and injected into a conventional GPC column. A calibration chart was constructed using polystyrene standards to relate the GPC retention time to the molecular weight, from which the DP of the polymer was determined. Table I lists the experimental conditions and results obtained with the 100 ml reactor. The reaction temperatures were 260 ° C and the reaction times were 60 minutes in each example. No polymerization catalyst was added to the reactor.

TABLE 1 100 ml reactor } Reactor experiments of 25 ml. The apparatus, experimental procedure and analyzes used in the following examples were identical to those previously described, with the exception of the internal volume of the reactor, which for this set of experiments was 25 ml. Table II lists the experimental conditions and results obtained with the 25 ml reactor. The reaction temperatures were again 260 ° C and the reaction times were 90 minutes for each example in the table. No polymerization catalyst was added to the reactor.

TABLE 2 Reactor of 25 m I The results in Table II can be compared with those in Table I. The lowest DP value of the listed product is related to the smallest reactor volume used in the experiments and the largest amounts of materials added to the reactor, on a from mass to volume.

Experiments in CO Table III lists the experimental conditions and results obtained when CO2 was used as the gaseous medium instead of EMD.

For these experiments, the intermittent production reactor was used. 100 ml. The reaction temperatures were 270 ° C and the reaction times were 60 minutes for the examples in the table. No polymerization catalyst was added to the reactor.

TABLE III Reactor of 100 ml with CO2 Experiments in high polymer DP stages were achieved by using low molar ratio reaction materials (within the claimed molar ratio scale) by performing the reaction in a series of successive steps. At the end of each stage, the gaseous medium, containing water, a minimum amount of glycol and any low molecular weight byproduct impurities were expelled and the reactor filled with the clean gaseous medium. The invention therefore provides unexpectedly purer polymers of commercially usable molecular weight by means of reactions carried out successively using fresh gaseous medium in each step. The method to operate the procedure in successive stages is described below. The first stage in the reaction series was carried out in two parts. After the completion of the first part of the reaction stage 1 in the required time, the pressure of the reactor was reduced, the seal was removed to remove the gaseous medium used, and a fresh gas charge was added. At all stages in the experiments listed in Table IV approximately 23 g of DME were used as the gaseous medium each time the gas was exchanged. The reactor was then resealed and reheated, and the second part of stage 1 was started. The product of the second part of stage 1 was then pooled as previously described and part of the sample was used to measure the DP of the product (10.4). The second and each subsequent stage had only one part. A known product weight from the previous stage, known DP, was added back to the reactor and the material ground to a fine powder in the reactor. The experiment was then performed again with fresh gaseous medium and repeated cycles until the necessary program was completed. After the first cycle, the product was removed at the end of each cycle for analysis. In order to better quantify the DP of the product samples from these staged experiments, the GPC calibration was modified using known molecular weight PET standards, to improve the accuracy of molecular weight measurement in the high DP region. All subsequent data reported herein are referenced against this new GPC calibration curve. Table IV lists the experimental conditions and results of the stepwise reactions conducted with 23 grams of DME. In those experiments, the 100 mL reactor was used. All the experiments described in Table IV were performed at 260 ° C for 60 minutes without any added catalyst.

TABLE IV Experiments performed in 100 ml Reactor with DME Note * - The PET used is the product of the previous stage Table V describes the experiments in stages conducted with CO2 in the 100 ml reactor using procedures similar to those described for Table IV. The reaction temperatures were 260 ° C and the residence times were 60 minutes. No polymerization catalyst was added.

TABLE V Experiments in stages in Reactor of 100 ml with CO2 Note * - The PET used is the product of the previous stage The increment in DP terminated experimentally "within the stage" tended generally downwards with an increasing number of stages, after the first two stages, through the experiment. In order to achieve a more commercially useful polymer (ie a higher DP) from this procedure, than those described in Tables IV and V, the polymerization process was carried out more frequently, using a further amount. small of material. This effectively reduces the amount of sub products generated in the enclosed and confined space of the intermittent production reactor at each stage, and better stimulates a continuous process. For this example a total of 0.050 g of reaction material was used consisting of 0.043 g MHET, 0. 006 g BHET, and 0.001 g TA, representing a molar ratio of glycol equivalent to carboxylic acid equivalent in the reaction materials of 1. 08: 1 This material was placed in the intermittent production reactor of 100 ml without catalyst, pressurized with CO2 at 35.15 kg / cm2 a, and heated to 280 ° C. The polymerization was carried out 15 times successively, the first 5 stages for 30 minutes each and the last 10 stages for 20 minutes each for a total of 350 minutes of reaction time in stages. At the end of this experiment, the material had a DP of 91. All reactor experiments described previously were the rationale for proving that the polyesterification reaction of the invention is possible. However, these reactions were all reactions of a single intermittent or stepwise production, successive intermittent production reactions requiring a larger gas volume. From a commercial point of view intermittent production reactions are impractical. In the following examples a polyesterification process is described using a tube or tubular reactor having a continuous flow, with at least the reactor outlet area being in turban flow.

Tubular reactors of continuous flow In turbanular flow, the gas flows along the core of the tubular reactor, while polyesterification occurs in the polymer, flowing along the wall of the reactor. The gas flowing in the center of the tubular reactor facilitates the removal of the polymer melt from the polyesterification sub-products, and any by-products that result from the polycondensation. This occurs in the melting surface of the polymer by diffusion in the gas phase. In the process of the present invention, the turban flow is maintained at high gas mass flow rates and low gas velocities operating at high gas densities. In any part of the tubular reactor the partial pressure of water or glycol, if present, in the gas phase, must be low enough to promote the diffusion of the sub product of the molten polymer to the gas phase, thus acting as both to promote the growth of polymer chain. Simultaneously the gas and polymer flows in the tubular reactor should be such that the turban flow develops. If the velocity of the gas is very large, the flow pattern will become "scattered" instead of turbanular. If the gas velocity is very small, the flow pattern will become unstable and "chaotic pieces" of polymer and gas will form, resulting in unstable conditions in the tubular reactor. At very low gas velocities the flow pattern becomes "stratified". Turban flow is the preferred flow regime for efficient polymerization and this condition develops over the last part of the tubular reactor, but in general for the initial reaction stage in the reactor intake, the turban flow pattern is not established completely.

Under suitable operating conditions, the turban flow is always present near the exit stage of the reactor. The concept of different flows in two-phase systems is well known and characterized by a 'Baker' plane. Such a diagrammatic image of the flow rates is shown in Figure 1. The data points illustrated on the drawing refer to those procedures known and / or operated by the inventors herein. As a partial explanation to Baker's plan, the data described on the ordinate axis are calculated from the data that refer to the gas phase component of the procedure, and those on the abscissa of the data are related to the phase component of casting the procedure.

EXAMPLE 1 Table VI illustrates the advantage of using a high pressure (dense) gas over a conventional low pressure system, when used in a tubular reactor. The gas velocities are based on the total cross-sectional area and the current gas velocities (not including the cross-sectional area of the polymer) will be larger than the tabulated values. In the turban flow, the current gas velocities can be up to twice the tabulated surface values depending on the thickness of the polymer ring. At the maximum reactor diameter given for the Porter low pressure gas system (11.6 cm) the high gas velocity required limits the polymer yield to 136 kg / hr. The tubular reactor defined by the method of the present invention, however, achieves a performance of 454 kg / hr. or more, preferably 908 kg / hr. or more, with significantly lower gas velocity under turban flow conditions. The lower gas velocity provides a higher polymer yield and allows an efficient gas / polymer separation step to be achieved while avoiding any appreciable polymer entry into the effluent gas stream. This performance effect is seen in Table VI. The conditions given under nitrogen gas in Table VI are based on data from the Porter reference ('587). Example # 5. In Porter Example # 5, the reaction temperature is 292 ° C. The pressure at the reactor outlet was 3.72 atm and at the outlet it was 1 atm. The material supplied to the reactor had an IV of 0.29, which rose to 0.63 in the product. For the first row of table 6, the temperature and IVs of the input and output polymer of example no are used. 5 of Porter, and a constant pressure is assumed along the 1 atm tube. The reaction diameter is set at 11.6 cm, which is the maximum allowable reactor diameter as taught by Porter. The production of PET in kg / hr as well as the gas velocity are extrapolated from the values in the Porter patent. The chain length growth of the polymer is measured by intrinsic viscosity (IV). An increase in IV from 0.29 to 0.63 corresponds approximately to an increase in DP from 41 to 104, calculated from the ratio: IV = 1.7x W4 (PM number) 083 where PM number is the number average molecular weight of the polymer produced and the repeated unit is assumed to have a molecular weight for polyethylene terephthalate of 192.

TABLE VI Comparison of turban reactor Note * - Surface gas velocity under turban flow.

It has been discovered that by practicing the invention using a superatmospheric pressure gas under conditions where polymerization occurs mainly by water removal, an elongated tubular polymerization reaction zone is not limited to a diameter of 10.16 cm previously identified by Porter. Consequently, larger-scale tubular reactors with diameters of 10.16 cm and larger are operable for commercial processes.

EXAMPLE 2 To illustrate the improvements, table VII compares the increased returns determined through modeling calculations. A computer-based model was developed from the procedure data in the given examples and the data identified in the plan Typical BAKER for the defined systems, and these are used to describe what can be achieved with the present invention in larger diameter reactors. The modeling conditions were similar to those in the example no. 5 on Porter's' 587 as previously described. The entry oligomer of low molecular weight PET had an IV of 0.29 (40 DP) and using low pressure nitrogen in one stage, the DP is increased to 0. 63 IV (104 DP). This DP product is the same with DME and CO2. It was discovered that by practicing this invention, the polymer yield is more than 10 times higher than that resulting from the operation under limits imposed by low pressure turban flow.

Therefore the ratio of tubular cross-sectional area to wetted perimeter in the present invention is larger than 2.5 cm, and the reactor yield (kg / hr) of polyester is significantly higher.

TABLE 7 Large-diameter turban reactors Note * - Surface gas velocity under turban flow.

POLYMERIZATION SYSTEMS IN STAGES EXAMPLE 3 A single-stage turban polymerization reactor system of the invention having a diameter above 10.16 cm may be supplied with low molecular weight material either by a conventional primary or secondary esterifier or other tubular system, operated at in such a way that the partial pressures of glycol and water are low, as can be obtained through vacuum conditions or by dilution with an inert gas of different low pressure. A multi-stage process is a series of single-stage tubular polymerization systems, which can constitute the total production chain from raw materials to finished product. In a turban flow reaction starting with EG and TA, two or more steps are needed to provide sufficient polymer chain growth by polyesterification to efficiently produce commercially usable PET resin. According to the invention, the superatmospheric pressure gas entering the tubular reactor stages must contain at most only a very small amount of moisture and / or smaller amounts of glycol byproducts, corresponding to a maximum moisture level for the gas , which depends on the reaction temperature and pressure. The moisture level of the gas must be low enough so that the partial pressure of the water in the gas is lower than the partial pressure of water balance at the reaction conditions. If the partial pressure is lower than the equilibrium partial pressure, the water in the reaction system will be transported from the polymer to the gas and the polymerization will proceed. If by-product water is not transported from the polymer phase, the polymerization will be impeded and there will be very little increase in the DP of the polymer. A knowledge of the partial pressure of equilibrium water vapor allows us to define the required amount of gas in the system so that the reaction is carried out below this pressure in all parts of the process. The following table VIII specifies the quantities of DME required to reach the specified DP scale from materials EG and conventional TA. The experimental data were used to establish the response of DP to partial pressure of water. The experiments revealed that the ratio of equilibrium gas to PET needed to achieve the DP of desired polymer product. From these data the corresponding reactor diameter and the gas flow rates required for turban flow at, for example, polymer yield of 4,500 kg / hr were determined.

TABLE VIII Stage polymerization in DME starting with EG and TA Note * - Surface gas velocity under turban flow.

The successive stages preferably use a tubular area of larger diameter and higher temperatures in order to maintain the gas velocity within the turbanular flow scale while driving the increasingly viscous polymer melt along the the reaction zone.

EXAMPLE 4 In the same way, Table IX specifies the quantities of CO2 required to reach the DP scale specified from conventional EG and TA materials under conditions derived from the current experiments. These experiments revealed the mass ratio of equilibrium gas to PET necessary to achieve the DP of the desired polymer product. In the same way, the gas flow velocities required to maintain the turban flow using CO2 using the specified reactor diameters at, for example, polymer flow rate of 5.450 kg / hr were determined.

TABLE IX Stage polymerization in CO2 starting with EG and TA Note * - Surface gas velocity under turban flow.

As shown in Table VIII and Table IX, where the process operates from very low DP reaction materials, the optimum diameter of the tubular reactor increases as the reaction proceeds and the polymer DP increases. Therefore if we use the data in Table VIII, then stage 1 is the initial stage where the reaction material is introduced into the tube reactor; Stage 2 is a middle portion of the largest diameter reactor and Stage 3 is the largest diameter portion of the reactor at the outlet end where the high DP polymer is produced. Steps 1-3 can be joined in multiple reaction sections, with different diameters beginning each of the sections. In this embodiment, if desired, the conditions of pressure, temperature, and mass flow rate of gaseous medium to polymer melt mass can be changed in each different reactor section.

EXAMPLES 5 AND 6 Tables X and XI show examples of polymerization systems in stages in which a material of low molar ratio (DP of 1) was used as the initial reaction material. Such a reaction material, for example, consists of 40 to 90 mole percent of MHET, and 30 to 5 mole percent each of BHET and TA in such a way that the molar ratio is 0.90: 1.0 to 1.2: 1.0. As previously taught, for reaction materials with initial molar ratios of less than 1.0: 1, adjustment is made in the latter procedure to increase the total molar ratio of the reaction material to at least 1.0: 1. The gas velocity is based on the total cross-sectional area for turban flow in the reaction zone. For illustration, in tables X and XI, a polymer yield of 4.540 kg / hr was used to specify the diameter of the reactor. As required, the performance can be changed resulting in a corresponding change in reactor diameter.

TABLE X Polymerization by stages in DME starting with DP Monomer 1 Note * - Surface gas velocity under turban flow.

TABLE XI Phase polymerization in CO2 starting with DP Monomer 1 Note * - Surface gas velocity under turban flow As the above explanation illustrates, the reaction material for the pressure polymerization process having the specified molar ratio can be selected from the group consisting of (A) a diol and a carboxylic acid, (B) an ester or monomeric esters of DP equal or 1.0, (C) an oligomeric ester or esters with an average DP greater than 1, or some combination of (A), (B), and (C) with the proviso that the total molar ratio of the reacted glycol equivalents and unreacted to reacted and unreacted carboxylic acid equivalents is within the range of 1. 0: 1 to no greater than 1.2: 1, considering all addition points.

DESCRIPTION OF THE EQUIPMENT USED TO OPERATE THE PROCEDURE OF THE INVENTION With reference to Figure 2 wherein similar numbers describe similar structures, a continuous polymerization process is provided which is not a continuous tube reactor, but a staged tubular reactor. The construction materials for the containers are beyond the scope of the description, but stainless steel 304 and 316, nickel alloys including types C-276, 265, 825 and 620, and ceramic coated metals are suitable. The reaction material is supplied through line 1 to vessel 20. The compressed gas compound is also supplied to vessel 20 through line 2. A dense gas / reaction material paste is supplied through line 3 to the reactor tubular 21 where the reaction occurs and the water of the reaction is transferred to the gas phase, increasing the degree of polymerization of the polymeric material. From the reactor 21, the two-phase material enters the separation vessel 22 where the dense gas is separated from the polymer by conventional means such as gravimetric separation. The polymer leaves the container from the bottom through the conduit 5, and the gas flows out of the upper part of the container 22 through a pressure regulating device, for example a control valve. The depressurized gas is supplied through line 8 to the cleaning unit , which may be composed of one or more containers, wherein the water and other compounds are removed from the gas by conventional means.

Any reaction material removed is discharged from the base of the container 25 through the conduit 10 for recovery, and any sub product of gas through other conduits, which are not shown. The clean gas is then supplied through conduit 9 to compressor 26. The clean pressurized gas enters the second stage tubular reactor, 23, with polymer from the first stage tubular reactor through conduits 11 and 5. The DP of the product The final polymer is increased to within the range of 40-180 or higher in the tubular reactor 23 and the polymer is separated from the gas in the container 24 with the gas passing through the conduit 2 to the container 20. If necessary, the Gas evolving from the reaction vessel 24, which exits in the conduit 2, can be cleaned of any harmful byproduct, in a manner similar to that used in the gas stream leaving the container 22, before the container 20 is returned. The polymer product is discharged from the system through the conduit 7. The pressures in the tubular reactors 21 and 23 are maintained above atmospheric pressure. The preferred pressures are in the range of 7.03 kg / cm2 to 175.75 kg / cm2 a, and may be different in the two reactors 21 and 23, or in the given example.

In one embodiment the gaseous medium consists of dimethyl ether.

As an example, reactor 23 is operated at 14.06 kg / cm2 at and at a temperature of 270 ° C. The temperature is maintained by a sleeve on the tubular reactor which contains a heating or cooling medium such as a heat transfer oil. The heated tubular reaction zone is of a predetermined length based on performance considerations, residence time, and preferred reaction conditions. As a specific example, a reaction zone of 300 meters long heated to 270 ° C and operating at almost 1,360 kg of DME per .450 kg of oligomeric reaction material is expected to have a residence time of 1 hr at a speed of gas of 8.23 meters / second. The reaction of the process results in an increase in DP of the polymer from an average of 9 for the supplied current to 50 or more for the polymer discharged from the tubular reactor 23. Higher DP increases are achieved when the system is operated with mass ratios from gas to higher reaction material. The highest limit in mass ratios of gas to reaction materials is reached when the turban flow in the tubular polymerization reactor becomes unstable. In the previous example, the reaction materials for the tubular reactor 23 are obtained by reacting oligomers in the reactor 21 at 14.06 kg / cm2 and 260 ° C, prepared from the reaction material composed essentially of EG and TA entering through the conduit 1 , at an average DP of 9. The oligomers leave the reactor 21 through the conduit 4, which is an extension of the tubular reactor 21, pass to a separation vessel 22, through the conduit 5, then towards the tubular rector 23. The gaseous medium and the gaseous by-products of the reaction in the tubular reactor 23 enter the separation vessel 24, which operates at a lower pressure than the reactor 23. If required, the gaseous stream leaving the vessel 24 may be cleaned using a container (not shown) that is similar to the gas cleaning unit 25, before recycling the gas stream through line 2 to the container 20. The polymer product discharged from the gas tubular actor 23 is controlled by adjusting the flow rate of the compressed gas supply to the reactor 23. Control and monitoring of process temperatures, pressures and flows is achieved using conventional means as would be used with conventional pumps, flow meters, and sensors as is well known in the art. In the procedure described as part of the previous example, the flow of the gaseous medium is concurrent with the melting of the polymer or oligomer phase at all times, both inside and between stages. Especially the gas enters with the reaction materials and passes through the successive stages before leaving the final reactor and being recycled. In another embodiment of the invention, the gas travels concurrently within a stage but against stages between stages. That is, it moves sequentially from the last stage of the procedure to the stage prior to the last stage, until it reaches the first stage before being recycled.

Within a stage the gas flow must be concurrent with the flow of the polymer to achieve a turbo-circular flow regime. The gas flow against interstage can be used to improve the efficiency of the total process and the economy of the operation by minimizing the total amount of compressed gas required for the polymerization. In this way, it is evident that, according to the invention, a method has been provided that fully satisfies the objects, objects and advantages set forth above. Although the invention has been described in conjunction with specific embodiments thereof, it is evident that many alternatives, modifications, and variations will be apparent to those skilled in the art in light of the aforementioned description. Accordingly, it is designed to cover all these alternatives, modifications and variations as they fall within the spirit and scope of the appended claims.

Claims (22)

NOVELTY OF THE INVENTION CLAIMS

1. - A process for preparing a polyester from monomers and / or oligomers of hydroxyalkyldicarboxylic acid ester consisting of: a) introducing reaction materials having DP of 1-40 of monomers and / or oligomers of hydroxyalkyldicarboxylic acid ester into a tube or tubular reactor, wherein said monomer and / or oligomers have a molar ratio of glycol equivalents to carboxylic acid equivalents of 1.0: 1 to 1.2: 1 considering all addition points; b) introducing a dense gaseous medium into said reactor; c) operating said reactor at superatmospheric pressure and at a temperature sufficient to: 1) achieve a turbo-air flow regime comprising said reaction materials and said gaseous medium before leaving the reactor; 2) causing the polyesterification to occur whereby the reaction material is polymerized to produce a polyester, with the DP of the polymer of reaction material increasing from the start of the reactor to the end; and 3) producing water with little or no reaction by-product or little or no unreacted monomer or oligomer which are separated from the polymer product through the gaseous medium phase.

2. The process according to claim 1, wherein said hydroxyalkyldicarboxylic acid ester monomer is an ester monomer of monohydroxyalkyldicarboxylic acid, or a mixture of monohydroxyalkyldicarboxylic acid ester monomer and a minor amount of acid ester monomer bishydroxyalkyldicarboxylic.

3. The process according to claim 2, wherein said monohydroxyalkyldicarboxylic acid ester can be monohydroxyethylterephthalate, monohydroxy (n-propyl) terephthalate, or monohydroxy (n-butyl) terephthalate.

4. The process according to claim 2, wherein said ester of bishydroxyalkyldicarboxylic acid can be bishydroxyethylterephthalate, bishydroxy (n-propyl) terephthalate or bishydroxy (n-butyl) terephthalate.

5. The process according to claim 1, wherein said dense gaseous medium is carbon dioxide, or nitrogen, or a mixture thereof.

6. The process according to claim 5, wherein said dense gaseous medium is carbon dioxide.

7. The process according to claim 5, wherein said gaseous dense medium is nitrogen.

8. The method according to claim 1, wherein said superatmospheric pressure is in the range of 1.05 kg / cm2a to 175.75 kg / cm2a.

9. The process according to claim 1, wherein said temperature is in the range of 220 ° C to 320 ° C.

10. The process according to claim 1, further comprising the step of introducing sufficient dicarboxylic acid with said reaction materials, as needed to obtain a total molar ratio of 1.0: 1 to 1.2: 1.

11. The process according to claim 1, wherein said reaction material and the dense gas medium are in concurrent flow within said reactor.

12. The process according to claim 3, wherein said reaction material is only monohydroxyethylterephthalate.

13. The process according to claim 3, wherein said reaction material is a mixture of monohydroxyethylterephthalate and bishydroxyethylterephthalate.

14. The process according to claim 2, wherein said reaction materials consist of 5-100 mole percent of monohydroxyalkyldicarboxylic acid esters and 0-50 mole percent of bishydroxyalkyldicarboxylic acid esters.

15. The process according to claim 1, wherein said reaction material has a degree of polymerization (DP) on the scale of 1-40.

16. The process for preparing a polyester from dicarboxylic acid and diol consisting of: a) introducing dicarboxylic acid and diol into a tube or tubular reactor wherein the molar ratio of glycol equivalents to carboxylic acid equivalents is 1.0 : 1 to 1.2: 1 considering all the points of addition; b) introducing a dense gaseous medium into said reactor; c) operating said reactor at superatmospheric pressure and at a temperature sufficient to react said dicarboxylic acid and said diol: 1) produce monomers and / or oligomers of hydroxyalkydicarboxylic acid ester having a DP of 1-40; 2) further reacting said hydroxyalkydicarboxylic acid ester monomers and / or oligomers by polyesterification to achieve a required turbo-annular flow in at least the outlet portion of the reactor; 3) producing water with little or no reaction by-product, or little or no unreacted monomer or oligomer, which are separated from the polymer product through the gaseous medium during the reaction.

17. The process according to claim 16, wherein the dicarboxylic acid is selected from a group consisting of oxalic, malonic, succinic, glutaric, adipic, pimelic, suberic, azelaic, sebasic, maleic, fumaric, phthalic, isophthalic acids. , terephthalic, those derivatives of naphthalene, anthracene, anthraquinone, biphenyl, and hemimelitic acids, or a mixture of them.

18. The process according to claim 16, wherein the dicarboxylic acid component is selected from a group of dicarboxylic anhydrides consisting of succinic anhydride, maleic anhydride, phthalic anhydride, and those derived from dicarboxylic acid or naphthalene, anthracene, anthraquinone, and biphenyl, or a mixture of them.

19. The process according to claim 16, wherein the diol component is selected from a group consisting of cyclohexyldimethanol, bisphenol-A, hydroquinone, ethylene glycol, 1-3-propanediol, 1-4-butanediol, or a mixture thereof.

20. The process according to claim 16, wherein all reactions occur without conventional catalyst as antimony, titanium, tin or germanium compounds, and the DP of the polymer is greater than 70. 21.- The procedure in accordance with Claim 16, wherein the reactor pressure is between 1.05-175.75 kg / cm2a and the reaction temperature is between 220-320 ° C. 22. The process according to claim 16, wherein said reactions occur with one or more conventional catalysts such as antimony, titanium, tin or germanium compounds.