JPH0362725B2 - - Google Patents

Description

[Detailed description of the invention]

Linear polyesters, consisting essentially of bisphenols and dicarboxylic acid residues, are important engineering thermoplastics recognized for their excellent performance, especially at high temperatures. The transesterification of bisphenols and diaryl dicarboxylic esters, preferably in the presence of small effective amounts of catalyst, for producing thermoplastic linear polyesters is described, for example, in Imperial Chemical, published April 24, 1963. This is disclosed in Industries UK Patent No. 924697. Reactions that are generally carried out in bulk, i.e., without solvents or diluents and under melting conditions of the reactants, generally employ temperatures from about 100° C. or higher to about 400° C. or higher, e.g., bisphenol A-isophthalic ester- For terephthalic acid ester polyester, temperatures from about 160°C or higher to about 350°C are used. Since the reaction is reversible, when using a diphenyl ester as the diaryl ester reactant, there is a Turn on the device. Monohydroxy aromatics can be produced by gradually raising the temperature of the reaction mass from the minimum temperature required to initiate transesterification to about 300-400°C, while gradually lowering the reaction pressure from about atmospheric pressure to about 0.1 mm Hg or less. Removal of group compounds is generally easier. The viscosity of this polyester product is generally prohibitively high; as the reaction approaches completion, the viscosity of the reaction mass, measured at reaction conditions, i.e., at about 32° C. at a shear rate of about 24 sec ^-1 , is about 1×10 ⁶ From the above, it approaches a value of about 10 × 10 ⁶ centipoise or more. Since the above indicates a wide variation in the conditions used for the reaction, such as temperature and pressure, and since a wide variation in the type and degree of agitation caused by the above viscosity increase during the reaction is required, the prior art The reactions were generally carried out in two steps, a prepolymerization step at a moderate temperature to produce a polyester oligomer prepolymer, and a subsequent polymerization step to condense the oligomer with the reactants to yield the polyester product. . The prepolymerization step is carried out at about 100 to about 350° C. and under moderate agitation conditions to produce a relatively low molecular weight prepolymer or oligomer corresponding to a low molecular weight polyester, i.e., an intrinsic viscosity of about 0.1 or less to about 0.3 dl/g. to make. The polymerization stage uses extreme agitation conditions at approximately 400°C.
It is carried out at temperatures ranging up to Due to the difficulty of dealing with the above-mentioned high viscosities of the molten reaction mass encountered in transesterification reactions, batch or semi-continuous processes (prepolymerization is batchwise and polymerization is continuous) are preferred for the production of these polyesters by transesterification. Only the formula ) has been proposed. Batch transesterification is generally not commercially attractive for producing large amounts of polyester, and the products generally exhibit undesirably wide variation in product properties from batch to batch. In the batch transesterification process, changes in reaction conditions result in product polydispersity (discussed further below).
It is generally not possible to substantially alter the molecular weight distribution of the product when measured by . West German Patent Application No. 2232877 (published January 17, 1974) to K. Eise et al. describes the semi-continuous transesterification polymerization of polyesters to form oligomers (referred to as precondensates in the above patent). The prepolymerization step is carried out in batch mode, and the subsequent polymerization step is carried out in continuous mode using an exhaust extruder reactor, ie, an extruder with a device for removing the vapors of the phenolic by-product of the reaction. The use of an extruder according to Eise et al. has the disadvantage that the reaction mass in the extrusion reactor is in the form of a thick mass, for example about 20 mm thick or more, and adheres primarily to the screw element of the extruder. The residence time of the reaction mass in the extruder is relatively long compared to thin film reactors, resulting in an undesirably long thermal history for the polyester. Such a long thermal history results in a more colored product, and in extreme cases, the discoloration of the product makes it unacceptable for many applications. In normal operation, the new surface exposure rate is limited, thus limiting the rate at which phenol can be removed from the system. As a result, the removal of phenolic by-products requires relatively expensive energy expenditures to maintain satisfactory transesterification rates. As an alternative to overcome the above-mentioned difficulties in the preparation of polyesters by transesterification, in particular those arising from the above-mentioned high melt viscosity of bisphenol-dicarboxylic acid polyesters, conventional techniques of the art, e.g.
Hemisier Verke Bitten issued on August 27, 1968
U.S. Patent No. F. Blasiyuke et al., assigned to AG.
3399170) are alkylene and cycloalkylene glycols, such as ethylene glycol,
We propose to modify the structure of polyester by introducing small amounts of polybutylene glycol residues into the polyester. Modification of the structure of the polyester in this manner tends to lower the melt viscosity of the polyester, thus overcoming the process difficulties described above, but with the high temperature performance characteristic of the unmodified bisphenol-dicarboxylic acid polyesters of the present invention. This results in a non-modified polyester product. According to the technology developed by Teijin for solid state polymerization, this method results in an increase in molecular weight, but this method is limited to materials that crystallize. The improvements of the present invention overcome the above-mentioned difficulties of the prior art in the production of linear aromatic polyesters by transesterification polymerization of monomeric reactants consisting essentially of bisphenols and diaryl dicarboxylic esters in the polymerization reaction zone. , the method involves reacting the bisphenol and the diaryl ester in a first reaction step in the zone to form a polyester oligomer, which is further polymerized in a second reaction step in the zone to increase the molecular weight and thereby form a polyester. It consists of producing a product. According to this improvement, the transesterification polymerization is carried out batchwise or substantially continuously in the prepolymerization stage, the transesterification polymerization is carried out substantially continuously in the polymerization stage, and the polymerization reaction in the polymerization stage is carried out in a batchwise or substantially continuous manner. It is carried out in a reaction zone equipped with a mechanically stirred thin film reactor. According to the present invention, substantially continuous means that the reactants are introduced into a particular reaction stage in a substantially continuous manner and the reaction products of that reaction stage are removed in a substantially continuous manner. A countercurrent mechanically stirred thin film reactor is defined as a reactor equipped with a motor-driven rotating screw or rotor with suitable vanes, and is used in the art for the continuous processing or reaction of viscous molten materials such as polymers. It is known that More particularly, as described below, the viscous material (in the present invention the transesterified mass)
is moved through the container by the rotation of the screw in the form of a thin film, said thin film adhering primarily to the inner wall of the container as it moves through the container. In currently known types of thin film reactors, the direction of mass movement through the vessel is opposite to the direction of gas or vapor flow emanating from the moving mass. In the process of the invention, the vapors removed from the transesterification mass are
The above-mentioned monohydroxy aromatic, ie, phenol compound, is a by-product of the transesterification reaction. According to one aspect of the invention, a semi-continuous process of transesterification polymerization is contemplated, in which the prepolymerization step is carried out in batches and the subsequent polymerization step is carried out in the countercurrent mechanically stirred thin film reactor, i.e. wiped membrane reactor. (wiped film) is carried out substantially continuously in one or more reactors. Another aspect of the invention contemplates a fully continuous transesterification polymerization, in which the prepolymerization is carried out in one or more stirred tank reactors, and the polymerization step is preferably carried out in one or more of the above-mentioned wipe membranes. Carry out in a reactor. According to another embodiment of the invention, an improvement in the above process technique for the transesterification production of linear aromatic polyesters is contemplated, in which the polymerization is carried out in a first stage batchwise or substantially continuously. , the polymerization is carried out in a continuous manner in the second stage, but if the polymerization in the first reaction stage is carried out in a batch manner, the polymerization in the second stage is carried out in the wipe membrane reaction zone and then in the exhaust extrusion reaction zone, i.e. The polymerization consists of an exhaust double-screw extruder as the reactor. Preferably, according to this embodiment, the continuous production of polyester also uses this form of reactor. This embodiment eliminates unsightly brown and black high molecular weight gel particles and burnt decomposition particulate products from the polyester, respectively. If the wiped membrane reactor described above is used in continuous or semi-continuous mode to produce high product molecular weights corresponding to a product intrinsic viscosity (measured as described in this application) of about 0.5 dl/g or more, It has been found that the gel and burnt particles are generally present in the product in trace amounts (the amount of the particles in the product increases with increasing molecular weight of the product). It is believed that the unsightly particles of polyester gel and polyester scorch result from entrainment of the polyester onto the wipe membrane reactor rotor shaft and vanes. The entrained polyester is thus not forced through the membrane reactor with the rest of the polymer mass. At the high temperatures reached in the final stages of polymerization, statically entrained polyesters polymerize to excessively high molecular weights and even to scorch. Additionally, entrained particles that accumulate in the wipe membrane reactor during semi-continuous and fully continuous polymerization are irregularly discharged together with the polymerized mass and appear as gel particles. Also, high viscosity polyesters in the final stages of polymerization can undergo excessive shear heating in the polymer oil sliding bottom bearings of wipe membrane reactors. Local temperature increases in the polymer due to heat from shear mechanical forces can also result in sometimes excessive polymer molecular weight increases and char formation, sometimes undesirable changes in the molecular weight of products obtained in semi-continuous or fully continuous processes. fluctuations, and contamination of the product with brown and black particles. This variation in product molecular weight and brown and black particle contamination of the product is overcome by the practice of this embodiment of the invention. The multi-screw extrusion reactor contemplated by this embodiment of the invention is a known device that uses multiple (generally two) rotating screws to move the molten mass through zones or stages in which a reaction comparison volatile organic by-products (corresponding to the relatively volatile monohydroxy organic by-products of the present invention) can be removed. This device ultimately extrudes the molten reaction mass through a suitable orifice, such as a die or tip plate. It is preferred to carry out the transesterification reaction in the presence of a catalyst, including acidic, neutral and especially basic catalysts for transesterification. Particular preference is given to introducing the basic catalyst in liquid form. It is particularly preferred to introduce the basic catalyst into the polymerization reaction mass simultaneously with the introduction of the molten reactants into the prepolymerization, so as to avoid substantial heating of the reactants by contacting the catalyst prior to initiation of the transesterification reaction. The present invention uses a semi-continuous or fully continuous mode of reaction to achieve desirable light colors and high melt viscosities, such as from less than about 1 x 10 ⁶ centipoise to about 10 x 10 ⁶ centipoise or more (about 24 sec ^{-1 )} . The present invention provides linear aromatic polyesters consisting essentially of bisphenols and residues of dicarboxylic acids, having viscosities in the range (measured at a shear rate of about 320° C.). Diphenyl isophthalate as a typical reactant,
Using diphenyl terephthalate, bisphenol A, the semi-continuous process of the present invention was carried out in the first, second and fourth stages.
Describe the diagram. Referring first to FIG. 1, these reactants pass through conduit 10 and into melting tank 12.
The tank is maintained at approximately 180°C to maintain the reactants in a molten state. The mixture is passed through conduit 14,
Good mixing of the reactants can be obtained by circulating through 18 and back to melting tank 12.
The mixture in melt tank 12 is continuously purged with dry nitrogen to exclude air. No appreciable reaction occurs at this stage. According to the semi-continuous process, the batch of monomer mixture is transferred through conduit 16 to agitated reactor 20. The catalyst is added either directly to the stirred reactor 20 or through conduit 17 with the molten mixture of monomers so that the catalyst does not come into contact with the molten monomers for a significant period of time prior to initiation of the polymerization reaction in the stirred reactor 20. Make it. The reaction is initiated by raising the temperature and lowering the pressure in the stirring reactor 20. In a stirred reactor, the intrinsic viscosity of the resulting polymer is approximately
The prepolymerization reaction is carried out to a degree of 0.15 to about 0.25, preferably about 0.18, which equates to a degree of reactivity of just over 90%. This prepolymer product is transferred to surge tank 44 under melt conditions. The phenol vapor produced in the reaction and some unreacted monomer pass from the stirred reactor 20 via an overhead conduit 24 to a distillation column 26, where the overhead mixture passes through a conduit 31 to a distillation column 26.
contact with recycled phenol that enters the system. The distillation column removes monomer entrained in the phenol vapor and returns the monomer through conduit 29 to the agitator reactor 20. Phenol vapor exiting overhead from distillation column 26 passes through conduit 28 to condenser 30. A portion of the recycled phenol passes through conduit 31 to distillation column 2.
6 and the remainder is sent through conduit 32 to distillate tank 34. Distillate is transferred from tank 34 to conduit 3
8 to the distillate tank 34. Vapors from the distillate tank are discharged through conduit 36 to a refrigeration trap (not shown). Liquid product from distillate tank 34 is removed from the process through conduit 40. The prepolymer is maintained at molten conditions in a surge tank 44 under a nitrogen blanket and passed through conduit 46.
It is pumped continuously at the rate required for the next polymerization step, which is operated in a continuous manner as described with respect to FIG. As shown in FIG. 2, the molten prepolymer is introduced through conduit 46 into the top of wipe membrane reactor 48;
The reactor 48 is propelled downwardly through the reaction zone in response to rotary impellers (not shown) in the reactor 48.
It is operated at a temperature range of 310°C and a pressure of about 1 to about 10 mm Hg absolute. The polymer product from the bottom of reactor 48 passes through conduit 64 to the top of wipe membrane reactor 66 and then travels down reactor 66 in response to the rotating blades of wipe membrane reactor 66. Polymer product is removed from the process from the bottom of reactor 66 through conduit 84. Reactor 48 includes a spray condenser 52 which receives overhead vapor through conduit 50 and discharges condensed phenol through conduit 54 to distillate tank 56. The effluent from distillate tank 56 is split, a portion being discharged from the process through conduit 60 and the remainder being atomized through conduit 58 for contact with further vapors coming in through conduit 50 from reactor 48. Recirculate to the top of condenser 52. Vapor from the spray condenser passes overhead through conduit 62 to a refrigeration trap (not shown). In a similar manner, the vapor product of reactor 66 passes through conduit 68 to a spray condenser 70, from where the condensate passes through conduit 7.
2 to the distillate tank 74. The liquid effluent from distillate tank 74 is partially recycled to distillate tank 56 and the remainder is recycled through conduit 78 to spray condenser 70 for contact with vapor distillate from reactor 66. be done. Overhead vapor product from the spray condenser is discharged through conduit 80 to a refrigeration trap (not shown). Ethylene glycol is introduced through conduit 82 into distillate tank 74, from where it is recycled through conduit 78 to the top of spray condenser 70, where the ethylene glycol helps condense and dissolve the phenol vapor. Another portion of the ethylene glycol is recycled through conduit 76 to distillate tank 56 and then through recirculation conduit 58 to the top of spray condenser 52 to perform a similar function. Meanwhile, the molten prepolymer is wiped from conduit 46 to membrane reactor 86, as shown in FIG.
is introduced into the top of the reactor 86 and travels downward through the reaction zone in response to rotating impellers (not shown) in reactor 86, which reactor 86 has a temperature of about 1 to about 10 mm at a temperature range of about 280 to 310°C.
Operated with absolute Hg pressure. The polymer product from the bottom of reactor 86 passes through conduit 102 to the head of evacuated extrusion reactor 104 and travels horizontally through the reactor in response to a rotating screw (not shown) in reactor 104 where it is reacted. The vessel 104 is operated at a temperature range of about 300 to about 350°C and an absolute pressure of about 0.1 mm Hg or more to about 1 mm Hg or less. The polymer product is in reactor 104
is removed from the process through conduit 106. Reactor 86 includes a spray condenser 90 which receives overhead vapor through conduit 88 and discharges condensed phenol through conduit 92 to distillate tank 94. The liquid effluent from distillate tank 94 is split, a portion of which is discharged from the process through conduit 98 and the remainder is passed through conduit 96 for contact with additional vapor coming from reactor 86 through conduit 88. Recirculated to the top of the spray condenser 90. Steam from the spray condenser passes overhead through conduit 100 to a refrigeration trap (not shown). Overhead vapor from reactor 104 is exhausted from reactor 104 through conduit 108 to a refrigeration trap (not shown). Diphenyl isophthalate as a typical reactant,
Using diphenyl terephthalate, bisphenol A, the continuous process of the present invention is described with reference to FIGS. 2, 3, and 4. Referring first to FIG. 3, these reactants are introduced through conduit 10 into a melting tank 12 which is maintained at about 160 DEG to 210 DEG C. to maintain the reactants in a molten state. Thorough mixing of the reactants is obtained by circulating the mixture through conduits 14 and 18 and back to melt tank 12. The mixture in melt tank 12 is continuously purged with dry nitrogen to exclude oxygen. No appreciable reaction occurs at this stage. According to a continuous process, the batch of monomer mixture is transferred through conduit 16 to agitated reactor 20. The catalyst can be added directly to the stirred reactor or in conduit 17.
with the molten mixture of monomers to prevent the catalyst from contacting the molten monomers for any significant amount of time prior to initiation of the polymerization reaction in the stirred reactor 20.
The reaction is initiated in the stirred reactor 20 by lowering the pressure and increasing the temperature. The stirred reactor 20 is operated at about 200-300°C, 40-300 mm Hg, and a residence time of 1 hour. The total amount of phenolic by-products is approx.
60% is removed from the agitated reactor 20. The viscosity of the reaction mixture is generally about 1000 cp at 240°C. The phenol vapors produced in the reaction and some unreacted monomer pass overhead from the stirred reactor 20 through conduit 24 to distillation column 26, where the overhead mixture passes through conduit 31 to distillation column 26.
contact with recycled phenol entering the phenol. The distillation column removes the monomer entrained in the phenol vapor and returns the monomer through conduit 29 to the agitated reactor 20.
Phenol vapor exiting overhead from distillation column 26 passes through conduit 28 to condenser 30. A portion of the recycled phenol is recycled through conduit 31 to distillation column 26 and the remainder is sent through conduit 32 to distillate tank 34. Distillate is transferred from tank 34 to conduit 3
8 to the distillate tank 34. Vapors from the distillate tank are discharged through conduit 36 to a refrigeration trap (not shown). Liquid product from distillate tank 34 is removed from the process through conduit 40. The prepolymerization reaction is continued in a stirred reactor 45,
The reactor has a temperature in the range of about 250 to about 300°C and about
Operated at 10-50 mmHg absolute pressure and 1 hour residence time. 30% of the total removable phenols are removed during this step, resulting in a reaction mixture with a melt viscosity of about 10,000 cp at 280°C. The intrinsic viscosity of the reaction mixture at this stage is approximately 0.18 dl/g. The reaction mixture or prepolymer is removed from reactor 45 through conduit 46. The phenol vapors produced in reactor 45 and some unreacted monomer are passed overhead in conduit 21 to distillation column 25 where the overhead mixture is contacted with recycled phenol entering distillation column 25 through conduit 35. The distillation column removes monomer entrained in the phenol vapor and returns the monomer through conduit 23 to reactor 45. Phenol vapor exiting overhead from distillation column 25 passes through conduit 27 to condenser 33.
go to A portion of the recycled phenol is returned to the distillation column 25 through conduit 35, and the remainder is returned to the distillation column 25 through conduits 37 and 32.
through the distillate tank 34 and processed as described above with respect to FIG. Referring again to FIG. 2, all polymerization products are collected in conduit 46.
is continuously sent to the wipe membrane reactor 48 through the membrane reactor 48.
In reactor 48, the molten prepolymer is introduced at the top of the reactor and travels downwardly through the reaction zone in response to slowly rotating impellers in wipe membrane reactor 48.
The reactor operates at temperatures ranging from about 280 to 310°C.
Operated at approximately 10mmHg absolute pressure. The polymer product from the bottom of reactor 48 passes through conduit 64 to the top of wipe membrane reactor 66 where the temperature is about 290-330°C and the pressure is 0.1-0.5 mm Hg and then wiped. It travels down the reactor 66 in response to the slowly rotating blades of the membrane reactor 66. Polymer product is removed from the process from the bottom of reactor 66 through conduit 84. Reactor 48 includes a spray condenser 52 which receives overhead vapor through conduit 50 and directs condensed phenol through conduit 54 to distillate tank 56.
discharge to. The liquid effluent from distillate tank 56 is treated as described for the semi-continuous process.
In a similar manner, vapor product from reactor 66 passes through conduit 68 to spray condenser 70, from where the condensate passes through conduit 72 to distillate tank 74.
The liquid effluent from distillate tank 74 is treated as described for the semi-continuous process. Ethylene glycol is introduced through conduit 82 into distillate tank 74, from where it is recycled through conduit 78 to the top of spray condenser 70, where the ethylene glycol assists in condensing and dissolving the phenol vapor. Another portion of the ethylene glycol is recycled through conduit 76 to distillate tank 56 and then through recirculation conduit 58 to spray condenser 5.
It is circulated to the top of 2, where it performs a similar function. On the other hand, as shown in FIG. 4, which illustrates an embodiment of the present invention employing an extruder in the polymerization step of the present invention, the molten prepolymer is transferred from conduit 46 to wipe membrane reactor 86.
is continuously introduced into the top of the reactor 86 and travels downward through the reaction zone in response to rotating impellers (not shown) in reactor 86 at a temperature range of about 280 to 310° C. and about 1 to about 10 mm Hg absolute. Operated by pressure. Reactor 86
The polymer product from the bottom of the reactor passes through conduit 102 to the head of evacuated extrusion reactor 104 and
The reactor 104 is moved horizontally through the reactor in response to a rotating screw (not shown) at a
From about 0.1mmHg or more to about 1 in the temperature range of ~350℃
Operated at absolute pressures below mmHg. Polymer product is removed from the process from the end of reactor 104 through conduit 106. Reactor 86 is equipped with a spray condenser 90 which receives overhead vapor through conduit 88 and discharges condensed phenol through conduit 92 to distillate tank 94. The liquid effluent from distillate tank 94 is divided, a portion of which is discharged from the process through conduit 98 and the remainder is passed through conduit 96 for contact with additional vapor coming from reactor 86 through conduit 88. Recirculate to the top of the spray condenser 90. Vapor from the spray condenser passes overhead to a refrigeration trap (not shown) through conduit 100. Overhead vapor from reactor 104 is exhausted from reactor 104 through conduit 108 to a refrigeration trap (not shown). A Reactant The reaction of the method of the present invention is a dicarboxylic acid with 6 to 20
Diesters with monohydroxy aromatic compounds of the benzene or naphthalene series of 5 carbon atoms are used. Examples of monohydroxy aromatic compounds are given in JP-A-Sho.
This is disclosed in Publications No. 56-8430 and No. 56-8429. Monohydroxy compounds of the benzene series, especially phenols, are preferred. Dicarboxylic acids useful in the method of the invention are also well known and can be represented by the following structural formula. However, X is oxygen or sulfur, and Z is -
Ar- or -Ar-Y-Ar-, where Ar has the same definition as given for biphenol and Y is alkylene of 1 to 10 carbon atoms, high alkylene, -O-, -S-, −SO−, _−SO2
−, −SO ₃ −, −CO−,

or GN〓, where n is 0 or 1, and G is alkyl, haloalkyl, aryl, haloaryl, alkylaryl, haloalkylaryl, arylalkyl, haloarylalkyl, cycloalkyl, halocycloalkyl. Examples of aromatic and aliphatic dicarboxylic acids are disclosed in US Pat. No. 4,126,602, the description of which is hereby incorporated by reference. Most preferably, the dicarboxylic acid component comprises a mixture of about 60 to about 100 mole percent isophthalic acid and about 40 to 0 mole percent terephthalic acid. Generally, aromatic dicarboxylic acid diesters which are normally solid, ie, solid at room temperature, are preferred. When the process is carried out in a semi-continuous or fully continuous manner in which the liquid reaction mass is passed from one reactor to another, the liquid reaction mass consists of about 15 to about 85 mole percent isophthalic acid and about 85 to about 15 mole percent terephthalic acid. Much preference is given to using dicarboxylic acid components. According to the invention, the use of the above proportions suppresses or prevents the precipitation, ie crystallization, of low molecular weight polyesters, ie polyester oligomers, from the molten transesterification mass. The precipitated oligomers can clog the conduits between the reactors through which the molten mass flows. Bisphenols useful in the method of the present invention are disclosed in US Pat. No. 4,126,602, the description of which is incorporated herein by reference. In general, bisphenols that can be used in the method of the invention are known in the art and correspond to the following general structural formula. However, Ar is preferably aromatic containing 6 to 18 carbon atoms (including phenyl, biphenyl, naphthyl), and G is alkyl, haloalkyl, aryl, haloaryl, alkylaryl, haloalkylaryl, arylalkyl, haloaryl alkyl, cycloalkyl,
halocycloalkyl, E is divalent (or disubstituted) alkylene, haloalkylene, cycloalkylene, halocycloalkylene, arylene, or haloarylene, -O-, -S-, -
SO−, −SO ₂ −, −SO ₃ −, −CO−,

[Formula] or GN〓, where T and T' are chlorine or a halogen such as bromine,
are independently selected from the group consisting of G, OG, and m is O.
b is an integer from O to the number of substitutable hydrogen atoms on Ar, and x is 0 or 1.
When there is more than one G substituent on the bisphenol, the substituents can be the same or different. The T and T' substituents can be in ortho, meta, or para positions to the hydroxyl group. The hydrocarbon group preferably has carbon atoms as follows. Alkyl of 1 to 14 carbon atoms, haloalkyl, alkylene, haloalkylene, aryl of 6 to 14 carbon atoms, haloaryl, arylene, haloarylene, 7 to 14
alkylaryl, haloalkylaryl, arylalkyl, haloarylalkyl of 4 to 14 carbon atoms, cycloalkyl, halocycloalkyl, cycloalkylene of 4 to 14 carbon atoms,
Halocycloalkylene. Furthermore, mixtures of the above bisphenols can be used to obtain polymers with particularly desired properties. Bisphenols generally contain 12 to about 30 carbon atoms, preferably 12 to about 25 carbon atoms. Bisphenols and biphenols are generally solids. The preferred bisphenols are commonly referred to as bisphenol A. The usual reaction procedure for carrying out transesterification production of the present ester is described in JP-A-56-8430 and JP-A-56-8429. Generally, a molar amount of dicarboxylic acid reactant, ie, diaryl ester, is used that is stoichiometrically equivalent to the molar amount of hydroxyl reactant, ie, bisphenol, used. When a polyester product with carboxylic ester end groups is desired, a molar excess of the diaryl ester over the hydroxyl reactant may be used, for example up to about 5 mole percent or more excess over the stoichiometric equivalent proportion of the diaryl ester. Use diaryl ester. Similarly, when a polyester with hydroxy end groups is desired, a similar molar excess of bisphenol and glycol modifier is used over the above stoichiometric equivalent proportions of the hydroxyl reactants. Unexamined Japanese Patent Publication 1987-
When a particularly end-capped polyarylate is desired, as in US Pat. Since one or more of the above reactants are usually solids, it is generally necessary to melt the reactants before introducing them into a polymerization reaction carried out in the molten state. Usually the solid reactants can be individually melted and mixed before addition to the polymerization, but since mixtures of solid reactants generally melt below the melting points of the individual solid reactants, it is preferable to mix the reactants and then melt them. preferable. Use of the latter preferred embodiment generally saves the energy normally used to melt the solid reactants. 1980
As disclosed in the pending application of J.C. Rosenfeld et al., U.S. Pat.
The invention is advantageously practiced using an excess of diaryl ester up to 1.8 mole percent. B Catalyst Also according to the practice of ordinary transesterification reactions, the preferred catalysts present in the transesterification reaction of the present invention are acidic, neutral, or basic catalysts; Based on the reaction of a conventional acid-base indicator with a catalyst when dissolved in an ionized solvent. It is preferable to use a basic catalyst as described in JP-A-56-136819. As described in JP-A-56-136819, the basic catalyst is brought into liquid form before introduction into the reaction mass, for example by melting or by dissolving in a liquid or normally solid but low melting point solvent. It is preferable to change. The use of catalysts, especially basic catalysts, which are normally solid at room temperature, in liquid form in the molten or solution state as described above accelerates the initial part of the transesterification reaction and also facilitates metering of the catalyst into the polymerization reaction. Therefore, excessive contact between the catalyst and atmospheric moisture can be avoided. In addition, introducing the basic catalyst in the liquid state and the reactants in the molten state into the polymerization reaction substantially simultaneously avoids contact of the catalyst with the reactants before the start of the transesterification reaction, resulting in superior quality. A polyester product (for example, less colored or produced with black spots, and excellent in terms of polydispersity as described below) can be obtained. Suitable basic catalysts include lithium, sodium,
Alkali metals such as potassium, rubidium, cesium, francium, and their carbonates, hydroxides, hydrides, borohydrides, phenolates, bisphenolates (i.e. bisphenols or salts of biphenols), carboxylic acids Salts such as acetates or benzoates, including oxides. In place of the above set of alkali metals and their compounds, groups and group elements such as calcium, magnesium,
Aluminum metal and its compounds can also be used. Other bases include trialkyl or triaryltin hydroxides, acetates, phenolates, and the like. Examples of catalysts are lithium, sodium, potassium, rubidium, cesium, francium metal, potassium or rubidium carbonates, potassium hydride, lithium hydride, sodium borohydride, potassium borohydride, calcium acetate, magnesium acetate, aluminum trichloride. Isopropoxide, triphenyltin hydroxide. Phenol is the preferred solvent for the normally solid catalyst. Substituted phenols that can be used include those having the following structural formula. However, R is alkyl of 1 to 10 carbon atoms, alkoxy of 1 to 10 carbon atoms, 6 to 10
aryl, chloro, bromo, or a mixture thereof, of 5 carbon atoms, and n is 1 or 2. Typical solvents are o-benzylphenol, o-
Bromophenol, m-bromophenol, m
-chlorophenol, p-chlorophenol,
2,4-dibromophenol, 2,6-dichlorophenol, 3,5-dimethoxyphenol, o-ethoxyphenol, m-ethylphenol, p-ethylphenol, o-isopropylphenol, m-methoxyphenol, m
-Including propylphenol, p-propylphenol, etc. Other particularly useful solvents are those of the ether type, such as tetrahydrofuran and various glymes, such as ethylene glycol dimethyl ether, liquid phosphites such as triphenyl phosphite and tris(nonylphenyl) phosphite, and phosphoric acid. Phosphate esters such as trimethyl or triethyl are also useful as diluents. Combinations of catalysts and phenols or phenols and solvents can also be used. Particularly preferred liquid basic catalysts are charged in a dissolved, usually solid, low melting point organic solvent such as phenol. Particularly preferred catalysts that support excellent results are the basic catalysts rubidium phenoxide, potassium phenoxide, potassium borophenoxide, each dissolved in molten phenol, as described in JP-A-56-136819. . Preferably also when using a basic catalyst,
As described in JP-A-56-136819, the catalyst is introduced into the polymerization simultaneously with the molten reactants to avoid heating the reactants in the presence of the catalyst before the transesterification begins. Catalysts that are less preferred than the above include metal oxides,
Contains metal acetate, titanium, and tin compounds. Suitable metal oxides include antimony trioxide, germanium oxide, arsenic trioxide, lead oxide, magnesium oxide, zinc oxide. Suitable lead metal acetates include cobalt acetate, zinc acetate, cadmium acetate, manganese acetate. Suitable titanium compounds include titanate esters such as tetrabutyl titanate, tetraisopropyl titanate. Suitable tin compounds include dibutyltin oxide, dibutyltin methoxide, dibutyltin dilaurate. In accordance with conventional transesterification practices, a catalytically effective amount, e.g.
~2 mol% or more, preferably about
Use 0.01-1 mol% of catalyst. C. Prepolymerization Stage for Semi-Continuous Transesterification According to the embodiment of the above semi-continuous process of the invention, the prepolymerization stage is carried out in batch mode and the polymerization stage is carried out in continuous mode. The reaction in the prepolymerization step for producing polyester oligomers having an intrinsic viscosity of about 0.1 or less to about 0.3 dl/g and a melt viscosity of up to about 100,000 cp is carried out using a conventional agitator connected in series. The design can be carried out with one, two, three or more reaction tanks. For carrying out the prepolymerization reaction step according to the semi-continuous process embodiment of the invention, it is preferred to use a single stirred reaction tank as the reactor. The reaction tank is equipped with a distillation device to remove the overhead of the phenolic compound by-product of the reaction which vaporizes at the temperature of the reaction mixture. Since some reactants can be entrained with the phenol byproduct vapor stream, the phenol byproduct removal system separates the phenol overhead vapor stream so that the entrained monomers are separated, condensed, and returned to the reaction mixture of the prepolymerization step. It is desirable to include a device for distillation and condensation. In semi-continuous operation, the phenol overhead vapors of the pre-polymerization stage and the phenol overhead vapors from the polymerization stage are condensed in the presence of ethylene glycol or an equivalent high-boiling phenolic solvent, as described below in connection with the operation of the polymerization stage. let The phenol overhead streams can be combined before the condensation step. The prepolymerization stage of the semi-continuous method of carrying out the process of this invention generally includes a holding or surge tank in series with the batch tank reactor described above. This tank can be equipped with an agitation device if desired, and serves to collect and hold batches of molten oligomer containing unreacted molten bisphenol and diaryl ester monomers for continuous feeding to the polymerization stage. To maintain the reactants in a molten state, the holding tank is maintained at substantially the same temperature as the immediately preceding tank. In a semi-continuous method of practicing the invention, the prepolymerization is carried out in batch reactors at temperatures from about 100°C or higher to about 350°C, preferably from about 160 to 330°C, especially from about 200 to about
The reaction is carried out at 300°C under reduced pressure from about normal pressure to about 10 mmHg absolute, especially about 1.0 mmHg or less, and the phenol byproduct is removed by increasing the temperature and decreasing the pressure as the reaction progresses. make it easier. Conveniently, in accordance with the semi-continuous embodiment of the present invention, the catalyst is added to the batch reactor in the prepolymerization stage while ramping the reaction temperature and reducing the reaction pressure stepwise. D. Prepolymerization Stage for Continuous Transesterification In fully continuous operation of the prepolymerization stage, a stirred reaction tank, such as that used in batch operation, is used to carry out the reaction in substantially continuous mode. Preferably two stirred tank reactors in series are used for continuous operation of the prepolymerization stage. If desired, one or all of the conventional stirred tank reactors can be replaced with a mechanically stirred thin film reactor of the type specifically described below which operates co-currently. In fully continuous operations, the prepolymerization stage preferably uses a separate distillation unit from that used in the polymerization stage for removal of the phenolic by-product of transesterification and separation of entrained monomers. In the phenol byproduct distillation removal system used for continuous prepolymerization, the phenolic compound vapor stream (containing traces of entrained monomer) from each reactor of the stage passes overhead to a fractionation column that carries the liquid entrainment. The monomers are separated and returned to the reactor, and the phenol vapor flows from the fractionation column, passes through a condenser and is collected as a liquid condensate, then flows to a freezing trap where it is solidified. The surge or holding tank used at the end of the batch prepolymerization step is vented so that the oligomers and unreacted bisphenols and diaryl esters are sent directly to the polymerization step below. Reaction conditions substantially similar to those described for the prepolymerization batch operation above are used, with the exception that gradients in reaction temperature and pressure are preferably avoided. E. Polymerization Step The polymerization step of the process, which is carried out in a substantially continuous manner, preferably employs one, two, three, or more series of countercurrent mechanically stirred thin film reactors, including the following: Other reactors may also be used in combination. As mentioned above, thin film reactors known for carrying out reactions of viscous substances in a continuous manner use a rotating screw for mechanical stirring and for moving the substances through the reactor as a thin film that adheres to the walls of the reactor. use. Thin film reactors for use in the present invention generally provide for the transfer of thin agitated membranes from about 0.1 mm or less to about 10 mm thick, preferably about 6 mm or less. Mechanically stirred thin film reactors are available of suitable construction and motor horsepower for reacting materials of various viscosities. From relatively low viscosity to about 0.1×10 ⁶
A mechanically stirred thin film reactor of relatively low horsepower and lightweight construction suitable for reactions of viscous materials up to centipoise is available as the Viscon Polymer Machines MK manufactured by Ruwa Corporation. High horsepower mechanically stirred thin film reactors of suitable weight construction are available for the reaction of more viscous materials, i.e. materials with a viscosity of 0.1 x ¹⁰⁶ centipoise or higher, and HM (for moderately viscous materials) and HS.
available as Filmtruder Polymer Machines, manufactured and sold by Ruwa Corporation under the name (particularly for viscous materials). Mechanically agitated thin film reactors of the type described above are available in both designs for countercurrent and cocurrent removal of vapor from the material being processed in the thin film reactor. In the polymerization step of the process of the invention, a countercurrent mechanically agitated thin film reactor is used such that the phenol byproduct is removed in countercurrent to the flow of the reaction mass through the thin film reactor. Thin-film reactors, especially mechanically stirred thin-film reactors (both co-flow and counter-flow types), are manufactured by Biurethan's Luwa Polymer Technology - Luwa Polymer Technology.
"Polymer Machines - HS-19", Ruwa Corporation and F. Widmer, "Processing of viscous substances in thin film equipment", Ruwa Corporation, 1970
Detailed and exemplified in 2007. Optionally, instead of the last wiped film reactor of the process of the invention according to the embodiments of the invention described above,
Using a dual-screw extrusion reactor, the product is virtually free of polyester gel and burnt material;
Also, products of consistent and uniform molecular weight can be obtained. Such extruder equipment typically includes an elongated, generally horizontal chamber or barrel, and a head or inlet end connected to a feed device. An extrusion section, such as an extrusion die (or tip plate) terminates the other end of the chamber.
In between the two ends is a device that applies vacuum and controls temperature. A vent is provided for venting steam. Within the chamber is a mechanical device such as one or more screws suitably positioned to force the molten polyester reaction mass forward through the extruder from the inlet end toward the tip plate for forward movement of the composition. is provided.
This screw propels the molten polyester reaction mass forward even at high viscosities. During vaporization of by-products,
This screw acts to knead the remaining polyester reaction mass within the chamber, facilitating its handling and forward movement through the extruder die, thereby avoiding substantial entrainment of the polyester within the extruder chamber. Generally, according to the present invention, a two-screw extruder is used in which the screws rotate in the same direction. Although one or more multi-stage extruders can be used in series, it is preferred to use a single extruder. Although one or more wiped film reactors can also be placed before the extruder, a combination of one wiped film reactor followed by a main extrusion reactor is preferred for the polymerization stage of the process. Extruders of the type used in the practice of this embodiment of the invention are generally described in US Pat.
No. 3335111, and this description is cited here. The preferred vented two-screw extruder contemplated by the present invention is similar to that described in K. Eise et al., cited above, DE 22328774. The description of this application is hereby cited. The preferred extruder of the present invention contains five vacuum zones, but corresponding extruders containing from three to six of the above zones can be used. On the other hand, although less preferred, in carrying out the fully continuous transesterification polymerization according to the invention, other reactors equipped with intensive agitation known in the art as suitable for the treatment of highly viscous masses can be used,
For example, a "dispack" polymer processing reactor can be used, such as that described by Z. Tadmore et al., Plastics Engineering, November 1979, pages 20-25. The polymerization step of the present semi-continuous transesterification process preferably uses two countercurrent mechanically stirred thin film reactors connected in series. Preferably both thin film reactors are of the high shear type described above. If convenient, the first thin film reactor at this stage is of the type described above suitable for handling moderately viscous materials, and the second
The thin film reactor is of the high shear type mentioned above.
In accordance with the present invention, the polymerization step of the process is generally at a temperature somewhat higher than that used in the prepolymerization, from about 150 to
in the range of about 400°C or higher, preferably
It is carried out at 225 to about 350°C, particularly about 150 to about 330°C, and the reaction pressure ranges from about 100 to about 0.7 mm Hg absolute, increasing the reaction temperature and decreasing the reaction pressure during this stage of the reaction. According to a particularly preferred operation of the process using two film reactors in the polymerization stage, the first film reactor is approximately
295 to about 305℃, especially about 300℃, about 5 to about 15
mmHg absolute pressure, especially about 10 mmHg absolute pressure, with a flow rate of melted reaction mixture through the reactor of about 4 to 16 Kg per hour, especially about 8 to 12 Kg per square meter of thin film reactor surface in the polymerization stage, from about 10 to about 200rpm
The rotor tip speed is preferably 0.5 to 1 m/sec. According to the above particularly preferred operation of the invention:
The temperature of the second thin film reactor during the polymerization stage is about 280 to about 325℃.
, more preferably about 300 to 315°C, especially about
at 305°C, with a reaction pressure in the range from about 0.5 or less to about 5 mmHg (absolute), preferably from about 0.5 or less to about 1.0 mmHg (absolute), as described for the thin film reactor prior to that step. Operating at substantially the same reaction mixture flow rate and screw tip speed as described above. The final thin film reactor of the polymerization stage is preferably equipped with a heated die to remove the molten polyester product. To facilitate product removal from the reactor, the die was heated approximately 10°C above the thin film reactor.
Preferably kept at high temperatures. According to embodiments of the invention that utilize a dual screw exhaust extruder as the final reactor of the polymerization, the extruder is generally operated at reduced pressure. The reduced pressure can follow a pressure profile or be substantially uniform.
When using a vacuum profile to operate an extruder,
The vacuum is highest at the feed end of the extruder and decreases along the extruder to the lowest vacuum at the exit end of the extruder. Advantageously, the vacuum profile ranges from less than about 1.0 mm Hg absolute at the feed end of the extruder to about 0.1 mm Hg absolute at the exit end of the extruder, preferably from about 0.8 mm Hg to about 0.2 mm Hg. It is. Preferably, the vacuum is substantially uniform in the extruder, advantageously in the range from about 1.0 or less to about 0.1 mm Hg absolute;
Preferably, the pressure is maintained at about 0.8 to about 0.2 mmHg absolute. The extruder may be operated at temperature conditions that provide a substantially uniform temperature for the molten reaction mass or melt therein, or may be operated at temperature conditions that provide a temperature profile for the melt within the extruder. A substantially uniform melting temperature in all zones of the extruder, ranging from at least about 250°C to about 390°C, preferably from about 300 to about 350°C, especially from about 320 to about 340°C, is advantageous. Preferably, according to the invention, a melting temperature profile is used in the operation of the extruder, the lowest melting temperature being at the feed end of the extruder, the melting temperature increasing gradually or in steps along the extruder, and the lowest melting temperature being at the extruder outlet end. to achieve the highest melting degree. According to this profile, the melting temperature advantageously ranges from about 250 to about 390°C at the feed end of the extruder, and preferably from about 300 to about 350°C, especially about 320°C at the outlet end of the extruder.
~about 340°C. The screw speed of this extruder is generally about 50~
200 rpm, preferably about 100-150 rpm. If desired, additional monomeric reactants can be introduced into the polymerization stage, especially by introducing monomeric reactants into the first reactor of that stage, but preferably any unreacted reactants present in the polymerization stage can be introduced. The monomers are those which have not reacted during the prepolymerization and are generally present in amounts up to about 10% by weight of the total monomers used. At the beginning of the continuous polymerization in the polymerization stage of the present invention, in order to avoid overhead loss of oligomer, the pressure is generally gradually lowered in the thin film reactor to reach the reaction pressure used in the reaction. Generally, at the beginning of the reaction of the polymerization step, the pressure is reduced from normal pressure to about 40 mmHg absolute at a rate of about 144 mmHg (absolute)/min, and then to about 10 mmHg (absolute) at a rate of about 15 mmHg (absolute)/hour. The reactor is then lowered at a speed while the filled reactor is heated to the speed used for the polymerization stage. When carrying out the method of the present invention in a completely continuous manner,
The fractionator used to remove phenol byproduct vapor from the thin film reactor at this stage is preferably separate from the corresponding fractionator at the prepolymerization stage. The phenolic compound vapor stream that is removed overhead from each of the thin film reactors in this stage is
Preferably, the phenolic solvent and vapor pressure depressant, such as ethylene glycol, or an equivalent solvent and vapor pressure depressant for the phenol vapor, is fed to a spray condenser operating in a circulating loop. Other solvents include glycols such as diethylene glycol and higher glycols. Preferably, each spray condenser glycol circulation loop operates independently. F Products of the Invention The methods of the invention provide products with improved reduced color. The product color of the semi-continuous process of the present invention is unexpected from the above-mentioned detrimental product color results obtained in semi-continuous polymerization practices using an extruder in the polymerization stage. According to the above-described preferred embodiment of the invention, in which it is intended to introduce the basic catalyst in liquid form into the polymerization simultaneously with the liquid reactants, polyesters of low polydispersity D are unexpectedly obtained. D is defined by the following formula: D=Mw/Mn, where Mw is the weight average molecular weight of the polyester product and Mn is the number average molecular weight of the polyester product. Polydispersity is also known as polydispersity index, heterogeneity ratio, heterogeneity coefficient, or dispersion ratio, and is F.
W. Billmeyer Jr., Textbook of Polymer Science, 2nd edition, Wiley Interscience, 1971.
78, and discussed in E. M. Matsukafei, Laboratory Preparation Methods for Polymer Chemists, McGraw-Hill Books, 1970, p. 332. The polydispersity of linear condensation polymers, such as polyesters of bisphenols and dicarboxylic acids, is a known measure of the breadth of the molecular weight distribution of the product molecules. Low values of D in the range from about 2.0 or higher to about 2.3 indicate a narrow molecular weight distribution and are not consistent with batchwise distribution as shown by KG Denby, Trans.Faraday Soc., vol. 43, p. 648 (1947). It is known that linear products of formula polycondensation are characteristic. High values of D, for example in the range from about 2.5 to about 6 or more, are known to be characteristic of linear products of continuous or semi-continuous polycondensation. Unexpectedly, the polyester products of the above-described preferred embodiments of the invention, although produced using semicontinuous polyesterification reaction techniques, exhibit low polydispersity, i.e., molecular weight distributions characteristic of products of batch polymerization. It has polydispersity. The low polydispersity polyester products obtained by the above-described preferred embodiments of the invention, which use a semi-continuous process of reaction in which liquid basic catalyst is added simultaneously with the addition of molten reactants, generally have increased impact strength, low gel concentration,
It is further characterized by a low free monomer concentration,
These properties could hitherto only be obtained with corresponding polyesters of low polydispersity produced by the batch process of transesterification polymerization. However, as is well known, batch production is commercially less desirable than semi-continuous or continuous polymerization processes for large scale production of polymers such as polyesters of bisphenols and dicarboxylic acids. Continuous transesterification polymerization processes also produce products with excellent properties. Furthermore, the continuous mode of the process of the invention, unlike its batch counterpart, allows the polydispersity of the product to be varied, ie increased or broadened, by varying the reaction conditions. By varying the reaction conditions, which affect the residence time distribution of the reaction mass, the polydispersity of the product can be varied. High molecular weight polyesters produced by the above embodiment of the invention using an evacuated double screw extruder reactor instead of a wiped film reactor as the final reactor of the polymerization, i.e. corresponding to a polyester intrinsic viscosity of about 0.5 dl/g or more. Polyesters of a molecular weight of about 100% are characterized by the advantageous properties mentioned above, and are further characterized by the substantial absence of brown gels and small unsightly particles of black char.
By the semi-continuous and continuous processes of the present invention using this embodiment, high molecular weight polyesters are generally produced with substantially consistent molecular weights. The polymer viscosity used here is from "Encyclopeia of Polymer Science and Technology", edited by Marc and Gaillord, vol. 14,
As defined in Interscience Publishers (Wiley & Sons), 1964, pp. 717-719. The following examples further illustrate the invention, but are not intended to limit the invention. Various modifications can be made to the invention without departing from the scope and spirit of the invention. Unless otherwise indicated herein, temperatures are in °C and all parts, percentages, and proportions are by weight. Examples 1 to 10 are reference examples conducted in batch mode, illustrating preliminary experiments regarding the semi-continuous or fully continuous production method of the present invention. Example 1 (Reference Example) A. Catalyst Preparation Approximately 800 c.c. of molten phenol was placed in a three-necked flask, mechanically stirred, placed under a blanket of dry nitrogen, and heated to reflux. The phenol was distilled through a 6-inch Viguru column at normal pressure, and approximately 500 c.c. of phenol center fraction was collected. 88.1 g of distilled phenol was transferred to a 250 c.c. three-necked flask with a nitrogen inlet, mechanical stirrer, Y-tube with thermometer, and drying tube (filled with dryerite). The system was placed under a blanket of dry nitrogen and heated to 70°C. KBH ₄ 1.00
g was added. After approximately 20 minutes, most of the solid had dissolved and the temperature was 55° C., and an additional 3.23 g of KBH ₄ was added. After another 2 hours, the reaction temperature was 71° C. and the solution was a clear off-white color and was allowed to cool. B. Production of prepolymer 4.61 mol (1052.5 g) of bisphenol A, 3.46 mol (1100.6 g) of diphenyl isophthalate, and 1.15 mol (366.9 g) of diphenyl terephthalate were dried in a vacuum oven at approximately 75°C and several mmHg overnight. Five resin kettles were equipped with a Vigre-column-distillation head combination with a nitrogen inlet, a mechanical stirrer, a thermometer, and three ice-water-cooled flasks as receivers. The reaction mixture was heated to 186° C. with a heating mantle so that all the reactants were liquid. Next, 2.5 cc of the above melted KBH ₄ -phenol solution
(0.002305 moles of catalyst) was added via a warm syringe under a blanket of dry nitrogen. At 190-255°C, phenol was continuously removed over a period of about 3 hours by gradually applying vacuum to about 1 mmHg. The clear pale yellow prepolymer was then poured into an aluminum foil lined glass tray. Intrinsic viscosity is 0.19dl/g
(sym-tetrachloroethane, 30°C).
The prepolymer was very light in color, slightly cloudy in appearance, and brittle. 798.1 g of phenol was collected in the receiver. C Polymerization The above prepolymer was crushed and 1500 g thereof was charged into a 2 gallon oil heated stirred reactor. The reactor was placed under a nitrogen blanket and left overnight. Vacuum pressure was applied to the reactor. The reaction mixture was heated to 290 °C for 2.1 h (starting at 139 °C with 5 rpm stirring for 1.0 h), at which point the vacuum was 0.6
It was mmHg. This condition (temperature approximately 295℃, 0.6
Polymerization was continued for 4.3 hours at (mmHg, 5 rpm). The reactor was then opened and the viscous pale yellow transparent polymer was removed. When cooled, the polymer was clear, extremely tenacious, and had an intrinsic viscosity of 0.61 dl/g (sym-tetrachloroethane, 30°C). Examples 2-5 (Reference Examples) The method of Example 1 was repeated using different catalysts and conditions as shown in Table 1. Table 1 also shows the intrinsic viscosity of the polyester products produced.

Table: Examples 6-10 compare the results of adding the catalyst before melting the reactants (Examples 6 and 7) with the present method in which the catalyst and liquid reactant were added simultaneously. Example 6 (reference example) Neopentyl glycol 0.0500 mol (5.21 g),
0.0513 mol (11.60 g) of bisphenol A and 0.1000 mol (31.80 g) of diphenyl terephthalate were placed in a small glass tubular reactor equipped with a gas inlet, a mechanically driven glass stirrer with a flat annular tip, and a curved tube leading to a trap. (with oil heating jacket)
I prepared it in. The system was placed under a nitrogen blanket. 0.3 cc (0.00015 mol) of a solution of lithium metal in phenolic tetrahydrofuran solvent was added to the reaction mixture and the mixture was stirred and heated to 230°C.
Vacuum was gradually applied to 230° C. and 1.15 mm Hg over about 1 hour, followed by heating at about 1.13 mm Hg for an additional 30 minutes.
The product was then removed from the reactor. The product was slightly cloudy and light in color with some dark particles. Example 7 (Reference Example) A reaction system was prepared as in Example 6, heated to 230°C, and the appearance was examined. Black particles floating in the melted monomer were observed and the experiment was stopped. Example 8 (Reference Example) Monomers were charged as described above, but no catalyst was added. The reaction mixture was heated to 210°C, then catalyst was added to the molten reactant to initiate polymerization.
The phenol was removed at 230-240° C. with increasing vacuum for 1.25 hours, and the product was clear (slightly cloudy) and pale yellow in color. No black spots were observed.
The intrinsic viscosity of the polymer was 0.33 dl/g (sym-tetrachloroethane, 30°C). Add 0.7g of the above prepolymer (0.33dl/g) to 250c.c.
Distribute to the bottom of a round-bottomed flask and incubate at 310 °C in an oil bath for 10
Thin film polymerization was carried out by heating at approximately 0.1 mmHg for 1 min. The resulting polymer was scraped from the flask while it was still soft. The resulting polymer was clear, pale yellow, and tenacious, with an intrinsic viscosity of 0.59 dl/g (sym-tetrachloroethane, 30°C). Example 9 (Reference Example) 0.10 mol (22.81 g) of bisphenol A, 0.050 mol (15.92 g) of diphenyl isophthalate, and 0.050 mol (15.92 g) of diphenyl terephthalate were charged into a reactor as described above, and the mixture was heated at 210°C. Once dissolved, 0.3 cc of catalyst solution was added and polymerization was carried out as above. No black spots were observed. The product had the following properties: Prepolymer: transparent, pale yellow, 0.27dl/g. Polymer: After 10 minutes at 310℃, transparent, pale yellow 0.93 dl/
g, after 30 minutes at 310℃, transparent, pale yellow, 1.23 dl/
g. Example 10 (Reference Example) The operation was repeated as in Example 9 using 0.1 cc of the catalyst solution. The following results were obtained. Prepolymer: light color, cloudy, 0.29 dl/g. Polymer: After 10 minutes at 310℃, light color, slightly cloudy,
0.58dl/g. After 30 minutes at 310°, light color, slightly cloudy, 0.74
dl/g. Example 11 The following describes the present method using the semi-continuous method described in FIGS. 1 and 2. Diphenyl isophthalate 60.65 g-mol, diphenyl terephthalate 20.25 g-mol, bisphenol in an agitated 10-gallon melting tank equipped with a turbine agitator and inert gas inlet.
81.25 g-mol of A was melted. This vessel and all subsequent vessels/reactors were heated by circulating hot oil (Therminol 66) through the jacket in a controlled manner. The internal temperature of this melting tank was maintained at 180±5°C. A slow stream of dry nitrogen was purged into the melting tank during the experiment. The molten monomer was then transferred to a 10 gallon stainless steel stirred tank reactor.
The transfer was performed by pressurizing the melt tank to 2-3 atmospheres gauge and opening the bottom valve. KBH ₄
The liquid catalyst of this method was prepared by reacting purified phenol with purified phenol to form a solution. The resulting catalyst solution is expressed on a molar basis of bisphenol.
_KBH4 was used at a level of 0.05 mol%. The catalyst was added in liquid form to the stirred tank reactor. The internal temperature of the reactor was raised to 280℃ in 2 hours and 15 minutes, and at the same time the internal pressure was increased to 70℃.
Lowered to mmHg. During this time, by-product phenol was continuously distilled off. At the end of this time, a sample was collected and the intrinsic viscosity was determined to be 0.175±0.005 dl/g. By pressurizing the reactor and opening the bottom valve,
The above material was transferred to a stirred surge tank. The holding tank was maintained at normal pressure, 260±5° C., and under constant nitrogen purge. The prepolymer was then passed through a metering gear pump, a film truder with a heated surface of 5.4 square feet from Ruwa Corporation;
It was continuously supplied to the top supply port of type HM. The internal temperature, pressure, and rotor speed maintained in this first film truder were 295° C., 4 mmHg, and 140 rpm, respectively. The material was pumped out continuously from the bottom and fed into a second film truder of type HS of the same size.
The internal temperature, pressure, and rotor speed of the second film truder were 305° C., 1 mmHg, and 100 rpm, respectively. Under these conditions, weight average molecular weight 27.800, polydispersity 2.2
A polyester with a 20 lb/hr capacity was produced. The properties of this polyester product were as follows. Tensile strength, psi 10158 Tensile modulus, psi 3.44× ¹⁰⁵ Elongation at yield, % 7.3 Heating strain temperature, °C (unannealed) 155.6 Izod impact strength, ft/lb/notch inch 4.8 Intrinsic viscosity, dl/gspm - Tetrachloroethane During,
30° C. 0.51 Yellowness Index approximately 27 Example 12 The following describes the method using the continuous method as described in FIGS. 2 and 3. The essentially stoichiometric mixture of reactants was continuously charged to a melting tank and suitably purged with nitrogen to exclude air. A catalyst solution was prepared as in Example 11 and introduced into the first stirred reactor along with the liquid monomer. The reaction was carried out in four reactors of Table 2 operated in series. Table 2 summarizes the conditions used in each reactor at steady state.

Table 2 also includes the sequential degree of phenol removal performed in each reactor. This method is about 0.60~0.65
Approximately 20 pounds/hour of polyester was produced with a final product intrinsic viscosity of dl/g (sym-tetrachloroethane, 30°C). Examples 13, 14 and 15 A mixture of diphenyl isophthalate/diphenyl terephthalate 75/15 and bisphenol A was oven dried in a vacuum oven at 2 mm Hg for approximately 3 hours in a large flat tray and then placed in a polyethylene lined 5 gallon pail. It was weighed and sealed until use. The amounts used in this example are shown in Table 3. Diphenyl isophthalate, diphenyl terephthalate, and bisphenol A were placed in a stainless steel melting tank and heated to approximately 180°C under a blanket of dry nitrogen.
It melted. The molten monomer mixture was transferred through a heating line to a continuously stirred oil heated reactor. Catalyst (potassium phenoxide, 0.041 molar solution in phenol, 0.0081 molar, based on bisphenol A)
0.05 mol%) was added. A vacuum was applied and the temperature (initially 220°C) was gradually raised. After about 2.5 hours, the pressure was about 70 mmHg and the temperature was 280°C. The intrinsic viscosity of the polyester prepolymer at this point is approximately 0.17
It was dl/g. This prepolymer was transferred through a heating line to a second stirred reactor, and another prepolymer was similarly produced in the first stirred reactor. The prepolymer from the second stirred reactor is
It was continuously fed at a rate of about 20 to 25 pounds per hour to a vertical wiped membrane reactor maintained at an internal temperature of 300°C and a pressure of 1.50 mmHg (absolute). The reaction mass fed to the top of the wiped film reactor flowed downward through the reactor, propelled by gravity and by the downward thrust of the pitched rotor blades. This vane also agitated and renewed the polymer on the heated reaction surface of the film reactor. The material leaving the thin film reactor had an intrinsic viscosity of approximately 0.4 dl/g. The polymer was pumped out the bottom of the wiped film reactor and fed into a 5-stage, 2-screw horizontal extruder with 5 vacuum vents (one for each stage). The extruder is approximately 0.8 mmHg (absolute pressure) and the screw speed is
Operated at 125 rpm. The pressure was kept uniform throughout the barrel, ie, the extruder. The temperature of the melt at different stages of the extruder is kept between about 320-340℃,
The former temperature is the melt temperature at the feed end of the extruder, and the latter temperature is the melt temperature at the exit end of the extruder. The following temperature zones outside the extruder body:
(1) (close to the feed end of the extruder) to 305°C, zone (2)
By keeping zone (near the center of the extruder) at 310°C and zone (3) (near the exit end of the extruder) at 315°C,
The melt temperature profile within the extruder was controlled. Under these operating conditions, a pale yellow, clear, tenacious polymer was produced at a rate of 20 pounds/hour. The product was essentially free of black char and brown polyester gel particles in accordance with the teachings of the present invention. The semi-continuous polyester production process described above was initially carried out using a 0.5 mole percent excess of diaryl ester reactant over the stoichiometric proportion corresponding to the molar amount of bisphenol used. After reaching steady state reaction,
Five samples of the polyester product were removed at two hour intervals and the polyester was tested for intrinsic viscosity and yellow index color. The statistical deviations of the results of intrinsic viscosity and burn index color of the samples were also determined. This data is reported in Table 3 as Example 13. The proportion of diaryl ester reactor charged to the process was adjusted to provide a 0.75 mole percent excess diaryl ester. After the reaction reached steady state, the product sampling and testing procedure described above was repeated substantially as described above. The intrinsic viscosity and yellowness index color data obtained are reported in Table 3 as Example 14. The proportion of diaryl ester reactant charged to the process was adjusted to provide a 1.0 mole percent excess of diaryl ester monomer. After the semi-continuous reaction again reached steady state operating conditions, the product sampling and testing procedures described above were repeated substantially as described above. The data obtained are reported in Table 3 as Example 15. Example 13,
14,15 The low standard deviation obtained from the average intrinsic viscosity results is indicative of the product molecular weight achieved when carrying out the semi-continuous production of polyester using an extrusion reactor as the last reactor of the polymerization according to the present invention. It shows a match. The invention has been described above and illustrated with respect to specific embodiments in the Examples. However, these embodiments are not intended to limit the invention. Changes may be made in the specific details described above without departing from the scope and spirit of the invention.

【table】 [Brief explanation of drawings]

FIG. 1 shows the first stage or prepolymerization stage of the semi-continuous transesterification process of the present invention. FIG. 2 depicts the second or polymerization stage of the semi-continuous and fully continuous transesterification process of the present invention. FIG. 3 shows the first stage or prepolymerization stage of the fully continuous transesterification process of the present invention. FIG. 4 shows the second or polymerization stage of the semi-continuous and fully continuous transesterification process of the present invention, where the second stage polymerization consists of polymerization in the wipe membrane reaction zone and then polymerization in the exhaust extrusion reaction zone. It's summery.

Claims (1)

[Scope of Claims] 1. Linear aromatic polymerization by transesterification of a monomeric reactant consisting essentially of a bisphenol compound and an aromatic dicarboxylic acid diaryl ester in a polymerization reaction zone consisting of a series of at least two reaction stages. An improved process for producing polyester, characterized in that the polymerization reaction in the final stage of the series is carried out substantially continuously in a reaction zone equipped with at least one mechanically stirred thin film reactor. 2 the polymerization reaction zone comprises two stages in which the monomer reactants are reacted in a first reaction stage to produce a polyester oligomer and then continued to react in a second reaction stage to produce a polyester product;
2. The method of claim 1, wherein said second reaction step is carried out substantially continuously in a reaction zone comprising at least one mechanically stirred thin film reactor. 3. The method of claim 2, wherein the first reaction step is carried out in batches. 4. The method of claim 2, wherein the first reaction step is carried out substantially continuously. 5. characterized in that the monomeric reactants are substantially continuously introduced into the polymerization reaction zone and the polyester product is substantially continuously removed from the polymerization zone,
The method according to claim 4. 6. A method according to any one of claims 1 to 5, wherein the polymerization is carried out in the presence of a polymerization catalyst, which is normally solid at room temperature and is introduced into the polymerization process in liquid form. 7. The method of claim 6, wherein the catalyst is a basic catalyst. 8. The method according to any one of claims 1 to 7, wherein the monomer reactant is introduced into the polymerization reaction in a liquid state. 9. The method of claim 8, wherein the polymerization catalyst in liquid form is introduced into the polymerization reaction substantially simultaneously with the introduction of the reactants in liquid form into the polymerization reaction. 10 A polymerization catalyst in a liquid form is introduced into a polymerization reaction substantially simultaneously with the introduction of the reactant in a liquid state into the polymerization reaction, and both the catalyst and the reactant are introduced into the polymerization reaction substantially continuously to form a polyester. 9. The method of claim 8, wherein material is removed substantially continuously from the polymerization reaction. 11. The method of any of claims 1-10, wherein the mechanically stirred thin film reaction zone provides countercurrent flow of the reaction mixture and the vapor by-products of the reaction. 12. The method according to any one of claims 1 to 11, wherein the bisphenol is bisphenol A. 13. The method according to any one of claims 1 to 12, wherein the diaryl ester is diphenyl isophthalate, diphenyl terephthalate, or a mixture thereof. 14. Claims 6 to 13, wherein the liquid catalyst is a solution obtained by melting a substance that is normally solid at room temperature, or by dissolving a substance that is normally solid at room temperature in a solvent for the catalyst. The method described in any of the paragraphs. 15. The method of claim 14, wherein the liquid catalyst is a solution of metal phenoxide in phenol. 16. The manufacturing method according to claim 15, wherein the metal phenoxide is potassium phenoxide. 17. The manufacturing method according to claim 15, wherein the metal phenoxide is rubidium phenoxide. 18. The manufacturing method according to claim 15, wherein the metal phenoxide is lithium phenoxide. 19. The manufacturing method according to claim 15, wherein the metal phenoxide is sodium phenoxide. 20. The production method according to claim 14, wherein the liquid catalyst is a reaction product of potassium borohydride and phenol. 21 The diaryl ester is approximately 15% of isophthalic acid.
21. A method according to any one of claims 1 to 20, wherein the diaryl ester is a mixture of from about 85 mole percent to about 85 mole percent to about 15 mole percent terephthalic acid. 22. The method according to any one of claims 1 to 21, wherein the diaryl ester is a diphenyl ester.