JP2752458B2 - Method for producing methacrylic polymer - Google Patents

Description

The present invention relates to an improved method for producing a methacrylic polymer. More specifically, the present invention employs special reaction conditions in a single stirred tank reactor (completely mixed reactor) in polymerizing a monomer mixture containing methyl methacrylate as a main component. The present invention relates to a method for obtaining a methacrylic polymer excellent in molding processability by industrially and advantageously performing bulk (bulk) polymerization.

(Prior art) Methacrylic polymer (hereinafter referred to as PMMA) has excellent transparency and surface gloss, excellent light resistance, mechanical strength,
Due to moldability and machinability, it has been used for signboards, nameplates, vehicle parts, and the like. In recent years, it has been used in a large amount in the optical field such as an optical disk substrate, and a high-purity polymer free from inclusion of minute foreign matter, which has not been a problem in the past, has been required.

The start of industrial production of PMMA began in the 1930s with the oldest polymer currently in use, along with polyvinyl chloride, polystyrene and polyethylene. After that, while the other three polymers grew rapidly and came to be called general-purpose resins, the biggest reason that only PMMA was left behind was that the other polymers quickly converted monomer raw materials into petrochemical products. In contrast to the conversion, only the methyl methacrylate (hereinafter referred to as MMA), a monomer of PMMA, was limited to a special raw material called hydrocyanic acid, which is extremely difficult to handle, until very recently.

However, a few years ago, in Japan, the conversion of raw materials to isobutylene, one of the representative petrochemical products, came to be carried out ahead of the world. Furthermore, with the development of an MMA synthesis method from ethylene and propylene in Europe and its industrialization planned, it was assured that the biggest factor that prevented PMMA from becoming a general-purpose resin would disappear. Thus, the rapid growth of PMMA, which has various excellent properties unrivaled by others, is expected.

The conventional industrial production method of PMMA has been a batch polymerization method using a suspension polymerization method. This method was suitable for high-mix low-volume production like the PMMA industry so far, but is not suitable for mass production following the generalization of PMMA to general-purpose resin. That is, disadvantages such as the complexity of batch operation, the increase in energy consumption due to the repetition of heating and cooling, and the necessity of a large amount of wastewater treatment become more pronounced as the size increases. These disadvantages have usually been solved by converting batch operation to continuous operation, but continuation of suspension polymerization is difficult due to the problem of sedimentation and deposition of slurry in the slurry transfer line, and MMA industrial continuous As a polymerization method, a so-called bulk (bulk) polymerization method is employed. This method does not require the addition of dispersants or stabilizers as in the suspension polymerization method, so it is superior to the suspension polymerization method, even for high-purity products, which have recently become increasingly demanding. . Therefore, it can be considered that a bulk polymerization method will be adopted as an industrial continuous polymerization method of MMA in the future.

Conventionally proposed continuous bulk polymerization methods are roughly classified into the following two types. One is JP-B 40-22,200, JP-B 55-7,
845, JP-B-54-42,035 and JP-A-58-132,002, by adding a solvent in an amount of 1 to 40% and using a complete mixing type reactor and a plug flow type reactor in a series to obtain a monomer conversion of 90%. As described above, the polymerization method (hereinafter, referred to as a high conversion method) is strictly a solution polymerization method. The other one is a method such as Japanese Patent Publication No. 52-32,665 and Japanese Patent Application Laid-Open No. 53-147,788, etc., in which the conversion is reduced to about 60% or less by a complete mixing type reactor without adding a solvent, and the unreacted monomer is used as a solvent for polymerization. (Hereinafter referred to as a low conversion rate method). Historically, the high conversion method is older and the low conversion method is relatively new. The reason is that the development of the bulk polymerization of MMA was modeled on the bulk process of styrene, which was already in large scale at that time. The industrialization of bulk polymerization is the earliest, and currently the largest is carried out on a large scale is the polymerization of styrene, but the polymerization method adopted there is almost a high conversion method.

Both have pros and cons. In the high conversion method, high viscosity liquids can be handled by specially devising the second-stage piston flow type reactor. The polymer content can be higher than in a perfectly mixed reactor. As a result, the one-pass productivity is increased, so that the main device is more compact. Also, the amount of monomers and solvents that need to be collected and recycled is reduced, so that the recovery device is compact and the amount of utility used there is reduced. However, there is a disadvantage that a special reactor is required in the second stage, which increases equipment costs and complicates the operation. The technical point of this method is how to achieve high monomer conversion with a small amount of solvent added, but it depends on how to develop a technology that can handle even highly viscous reaction liquids in the second stage reactor. I have. However, this method has an essential disadvantage that the polymerization degree distribution of the final product is slightly widened due to the difference in the concentration of the radical initiator (hereinafter simply referred to as initiator) in the first and second stages. Further, a solvent separation device is required.

On the other hand, in the low conversion method, a one-pass monomer conversion rate is at most about 60%, so that a large amount of monomer must be recycled. However, there is no need to separate the solvent and monomer, and in bulk polymerization of MMA, the difference between the impurities in the recovered monomer and the boiling point of the monomer is large,
The utility costs required for monomer recovery and recycling are not too high. The advantage of this method is that a general stirred tank can be used as it is as a reactor,
The sensible heat of the feed monomer can be effectively used for cooling the polymerization heat, so that the heat removal device can be reduced, and the distribution of the degree of polymerization is slightly sharper than that of the high conversion method using a two-stage reaction.

The low conversion method has been widely adopted as a method for producing a prepolymer (syrup) in a cast polymerization method. In this case, the monomer conversion rate is at most 30 to 40%. Also in the above-mentioned JP-A-53-147788, the monomer conversion is 25 to 60% in the claims,
Looking at the examples, the conversions are all below 40%. The above-mentioned JP-B-52-32,665 was the first to carry out the low conversion method at a relatively high monomer conversion of about 60%.

Japanese Patent Publication No. 52-32,665 discloses a method in which the type and amount of an initiator are limited and the monomer conversion is 50-78% at a polymerization reaction (hereinafter sometimes simply referred to as reaction) temperature of 130-160 ° C. An initiator having a half life at the reaction temperature of 2 minutes to 5.6 hours is claimed. The average residence time (hereinafter simply referred to as residence time) is not limited, but according to the examples is from 3 to 6 hours, the half-life of the initiator used at the reaction temperature being from 18 to 33 minutes. . From this, the ratio of the half-life of the initiator to the residence time is in the range of 1/6 to 1/45. The ratio of the minimum half-life in the claim to the maximum residence time in the examples is 1 /
180. According to this patent, the use of an initiator having a shorter half-life will lower the thermal decomposition temperature of the product polymer and reduce the temperature range in which molding can be performed (hereinafter, this is simply referred to as poor molding processability. But there is no mention of the reason.

(Problems to be Solved by the Invention) Under the polymerization conditions of Japanese Patent Publication No. 52-32,665, a so-called automatic accelerating effect (hereinafter, referred to as a gel effect) is developed and is characterized by utilizing it. However, according to experiments conducted in accordance with the examples of this patent, it was found that there was a problem that it was difficult to react at a stable monomer conversion rate, and that only operations with a large fluctuation range of the conversion rate could be performed. In addition, the fluctuation range becomes too large, and the reaction may run away. If the monomer conversion rate fluctuates, the reaction heat generation amount, the polymer production amount, the recovered monomer generation amount, etc. will fluctuate, which will hinder stable operation of the entire process.

Furthermore, if a trouble occurs in the process after the reactor, continuous operation will be temporarily stopped and repair will be performed, but the monomer conversion rate will increase steadily during the temporary suspension of the continuous reaction. It turned out that there was a problem of runaway.

An object of the present invention is to avoid the above-mentioned difficult problems in continuously bulk-polymerizing a monomer mixture containing methyl methacrylate as a main component without using a solvent by using a single complete mixing reactor. It is an object of the present invention to provide a method for obtaining a colorless and transparent polymer excellent in molding processability while achieving stable operation at an arbitrary monomer conversion ratio of 45 to 70%.

(Means for Solving the Problems) The inventors of the present invention have conducted intensive studies to achieve the above object, and as a result, after reducing the dissolved oxygen in the monomer to 1 ppm or less, an initiator having a specific half-life The present inventors have found that the object can be achieved by stirring at a specific temperature and a residence time at a specific monomer conversion rate with stirring at a specific power consumption, thereby completing the present invention.

That is, the present invention provides a method for continuously (bulkly) polymerizing a monomer mixture containing methyl methacrylate as a main component without using a solvent by using a complete mixing type reactor. To reduce the dissolved oxygen in the monomer to 1 ppm or less, (2) using a radical initiator having a half-life at the polymerization temperature of 0.5 to 120 seconds,
(3) reaction with stirring at 1 m ³ agitator having a stirring power consumption per 0.5~20KW, (4) the ratio of the half-life of the radical initiator at the polymerization temperature average residence time (= tau _1/2 / θ) is 1
Set the average residence time to be / 200 to 1 / 10,000,
(5) A process for producing a colorless and transparent methacrylic polymer excellent in moldability, characterized in that the polymerization is carried out at 130 to 160 ° C so that the monomer conversion is 45 to 70%.

 Hereinafter, the present invention will be described in more detail.

The methacrylic polymer of the present invention is a homopolymer of MMA or 1 to 15% by weight of methyl acrylate (hereinafter referred to as MA).
Or acrylate (hereinafter referred to as EA).

The present invention comprises the steps of monomer preparation, polymerization, heating, degassing extrusion and monomer recovery in the order of the steps. Hereinafter, the steps will be described in order.

In the monomer compounding step, MMA and a predetermined amount of MA or EA, a chain transfer agent and an initiator are mixed and dissolved in a tank (mixing tank) with a stirrer, and are continuously supplied to a polymerization reactor. MM
Since the polymerization activities of A and MA or EA are different, the monomer composition used is slightly different from the copolymer composition in the polymer, but the relationship between the two can be easily determined by testing in advance. The amount of the chain transfer agent to be added to 100 g of the monomer is selected from the range of 0.01 to 0.1 g so as to have a target average molecular weight. 0.001 ~
From the range of 0.1 g, it is selected so that the target monomer conversion is obtained at the polymerization temperature at that time.

As the chain transfer agent to be used, mercaptans, especially n
-Butyl, n-octyl, n-dodecylmercaptan are preferred.

The initiator used has a half life at any reaction temperature between 130 and 160 ° C. of 0.5 to 120 seconds, preferably 1 to 60 seconds.
Seconds radical initiator, which is the most important point of the present invention. When an initiator having a half-life of more than 120 seconds is used, the fluctuation range of the monomer conversion becomes large and it becomes difficult to secure a stable operation. Also, if the half-life is less than 0.5 seconds,
The use of too much initiator causes problems in the colorless transparency of the product polymer. As the initiator used in the present invention, 2,2-azobisisobutyronitrile, 2,2-azobis (2,4-dimethylvaleronitrile), 2,2-azohis (2
Azo compounds such as -methylbutyronitrile), and
Organic peroxides such as 1,1-bis (t-butylperoxy) 3,3,5-trimethylcyclohexane, t-betylperoxyisobutyrate, benzoyl peroxide and lauroyl peroxide are preferred. In the present invention, the half-life value of the initiator is calculated according to the technical bulletin issued by Wako Pure Chemical Industries, Ltd. for azo compounds and the catalog (12th edition) issued by NOF Corporation for organic peroxides.
Was adopted.

There are two types of monomer blending systems, batch and continuous, any of which may be employed. In the batch method, two mixing tanks are installed and batch mixing is alternately performed and used. In the continuous blending, the flow rates of the respective components are automatically controlled to predetermined amounts in separate lines, and then mixed and supplied to the reactor. Of these, only the initiator is continuously supplied as a 0.1 to 5% monomer solution.

The method of reducing the dissolved oxygen in the monomer to 1 ppm or less may be carried out by bubbling an inert gas into the mixing tank for a certain period of time in the case of batch mixing, but it is installed in the line when charging the monomer into the mixing tank. An effective method is to mix an inert gas with a prepared line mixer and to separate gas and liquid in the mixing tank. It is preferable to use a motionless mixer such as a static mixer as the line mixer. In the case of continuous preparation, after mixing each component, a line mixer and a small gas-liquid separator are provided, and the inert gas is continuously supplied to the line mixer and discharged from the gas-liquid separator, thereby dissolving dissolved oxygen in the monomer. Should be removed. The amount of the inert gas is preferably selected from the range of 2 to 10 parts by volume of gas per 1 part by volume of monomer. 1ppm dissolved oxygen in monomer
If it exceeds, the monomer conversion rate of the reaction becomes unstable.

The mixed solution prepared as described above is preferably passed through a filter of 0.5 micron or less, and then fed to the reactor.

In the polymerization step, the polymerization reaction is carried out at a temperature in the range of 130 to 160 ° C., preferably 140 to 155 ° C., so that the monomer conversion is substantially constant in the range of 45 to 70%, preferably 55 to 65%. . If the reaction temperature exceeds 160 ° C., undesirably, high boiling substances such as oligomers such as dimers and trimers and compounds of monomers and mercaptans increase. Also, 130
If the temperature is lower than 0 ° C., the viscosity of the reaction solution increases, and the contribution of the gel effect increases, making it difficult to maintain the monomer conversion rate stably, and disadvantageous in removing the reaction heat. If the monomer conversion rate exceeds 70%, the viscosity rises and it becomes difficult to maintain a stable operation. If the monomer conversion rate is less than 45%, the load on the degassing step and the monomer recovery step increases, which is disadvantageous. Since the self-controllability of the reaction described disappears, it is disadvantageous for stable operation.

The present invention is based on the finding that when polymerization is carried out under a specific condition using a specific initiator, a higher reaction temperature results in a lower monomer conversion, which is contrary to common knowledge of chemical reaction kinetics. With two autoclaves with double helical ribbon blade stirrer, 2,2-
Using azobisisobutyronitrile (hereinafter referred to as AIBN) at a reaction temperature of 150 ° C, a residence time of 3.0 hours, and a stirring rotation speed of 200 rpm
FIG. 1 shows the results of an experiment conducted at a power consumption of stirring of 9 KW / m ³ . At a monomer conversion of 45% or less, a common sense result is that the higher the temperature, the higher the conversion at the same initiator concentration, but the opposite occurs at a concentration of 45% or more. This figure can explain the self-controllability of the polymerization reaction which forms the basis of the present invention. That is, suppose that the reaction temperature rises for some reason. According to FIG. 1, as a result, the monomer conversion rate decreases, so that the amount of heat of reaction decreases and the reaction temperature automatically lowers to restore the original temperature. Conversely, the same applies when the reaction temperature decreases. According to FIG. 1, the higher the monomer conversion rate, the greater this reversal phenomenon. The reason for such a reversal phenomenon is that the gel effect occurs at a monomer conversion rate of 45% or more and a low temperature and high conversion rate occurs. It is considered that the degree increases as the rate increases. The stability of the monomer conversion in the continuous polymerization achieved by the present invention is due to such self-control.

The residence time of the reaction may be selected from the range of 1 to 6 hours. In this case, the ratio of the half life at the reaction temperature of the initiator used to the residence time (τ _1/2 / θ) is 1/200 to 1 /. It is essential to set it in the range of / 10,000. Such a small (τ
Examples ever carrying out the polymerization reaction using the value of _1/2 / theta), is the fact that shown in Figure 1 those for the first time obtained by adopting the value of such (τ _1/2 / _θ) It is.
Stable operation at a monomer conversion of 60 to 70%,
The short residence time was achieved by employing such a small (τ _1/2 / θ) value, and the present invention provides a process with unprecedented productivity. It is.

The polymerization pressure is not particularly limited, but when the reaction solution is boiled, the polymer adheres to the vessel wall of the gas phase in the reactor due to entrainment and the like, which hinders long-term operation. It is preferable to pressurize with an inert gas so that the reaction pressure becomes higher, and to carry out the reaction while suppressing the boiling of the reaction solution. The inert gas is introduced from immediately below the seal portion of the stirring shaft such as a mechanical seal or the like, and is used to prevent monomer vapor from rising to the seal portion and to prevent trouble of polymer adhesion in the seal mechanism.
This gas is discharged after condensing and removing monomer vapor via a reflux condenser, and after controlling the pressure to be constant by a pressure control valve. The pressure may be selected from 3 to 9 kgf / cm ² G, but it is not necessary to increase it too much because the purpose is to suppress boiling of the reaction solution. Excessive pressure reduces the monomer composition in the gas phase of the reactor, and consequently reduces the amount of monomer vapor condensed in the reflux condenser, which is disadvantageous in terms of heat removal.

Regarding the stirring of the reaction solution, the power consumed for stirring was 1 m ^{3 of the} reaction solution.
It is essential to set it to 0.5 to 20 kW, preferably 1 to 10 kW per unit. If it is 0.5 KW or less, the amount of the initiator required to obtain a predetermined monomer conversion rate increases, which adversely affects the moldability and colorless transparency of the polymer. On the other hand, when the power exceeds 20 KW, the power consumption becomes too large, which is not economical. As the stirring blade, it is preferable to use a general paddle blade, a double helical ribbon blade, or a Max Blend blade of Sumitomo Heavy Industries, Ltd. In the method of the present invention, since the monomer must be evaporated while suppressing the boiling of the reaction solution, a double helical ribbon blade or a max blend blade having a large liquid surface renewal effect is recommended. When these stirring blades are used, there is an advantage that the evaporation of the monomer can be drastically promoted by stirring with the liquid level being lower than the upper end of the blade.

Removal of heat of reaction is performed by three methods conventionally used,
That is, it is desirable to use a method that utilizes the sensible heat of the monomer to be supplied, a method that removes heat from the reaction solution using a heating medium, and a method that uses the latent heat of evaporation of the monomer in the reaction solution. According to the method of the present invention, the amount of heat removed by sensible heat can be increased by cooling the monomer to be supplied, so that the amount of heat removed which must be dependent on the other two methods is reduced. As a result, there is no need to insert a heat exchanger into the reactor or install an external heat exchanger to externally circulate the reaction solution by means of a circulation pump. The size of the reflux condenser becomes compact, and the equipment cost is reduced.

In the heating step, the reaction solution obtained by the above polymerization reaction is heated to 220 to 270 ° C. in the shortest possible time, and the latent heat of vaporization necessary for gasifying unreacted monomers in the next degassing step is reacted. This is a step of applying sensible heat to the liquid. The heating step and the degassing step are separated by a needle valve, and the pressure in the heating step is set slightly higher than the saturated vapor pressure of the monomer at the heating temperature. Otherwise, gas may be generated in the heater, which not only reduces the heat exchange capacity, but also generates carbide in the heater.
If the heating temperature is lower than 220 ° C., the temperature of the polymer after being flushed by the needle valve becomes too low, and the load of deaeration of the deaeration extruder is undesirably increased. On the other hand, if the temperature is higher than 270 ° C., the amount of high-boiling by-products such as dimers and trimers increases, which is not preferable.

In the heating step, the temperature of the reaction solution is raised as high as 60 to 120 ° C., so that it is inevitable that the polymerization reaction and the side reaction to generate dimers, trimers and the like proceed to some extent.
When the polymerization reaction proceeds in the heater, the degree of polymerization of the polymer is slightly different from that of the polymer produced in the reactor.
Also, if dimers or trimers are formed in the heater, the purity of the polymer is undesirably reduced. However,
In the present invention, the initiator concentration in the reaction solution in the heater,
Since the concentration of the initiator in the monomer supplied to the reactor is several thousandths, the progress of the polymerization reaction in the heater becomes negligibly small. This is one of the advantages of the present invention. In order to suppress the formation of dimers and trimers in the heater, it is necessary to reduce the volume of the heater as much as possible and raise the temperature to a target temperature with a short residence time. For this purpose, it is preferable to use a heating device having a large heat transfer coefficient, and it is advantageous to use a double-tube or multi-tube heat exchanger of a motionless mixer such as a Noritake static mixer. This allows the residence time to be between 1 and 3 minutes, minimizing by-product formation in the heater.

In the degassing step, the reaction liquid heated in this manner is flushed with the needle valve to separate unreacted monomers from the polymer as gas. Degassing is preferably performed in two stages. The needle valve installed at the outlet of the heater is connected to the first-stage deaerator. It is preferable to use an extruder having the following structure as this device. A needle valve is connected to the barrel of the screw deep groove (feed zone) to flush the reaction from the heater, and the polymer is extruded from the die through the compression and metering zones that follow the feed zone. The gasified monomer is withdrawn at normal pressure from a vent (back vent) hole provided in the barrel of the screw deep groove opposite to the polymer flow direction with respect to the needle valve. This gas may be fed as it is to the distillation tower. At this stage, the monomer in the polymer is usually 1-5%.

The second stage degassing reduces the monomer content of 1-5% to 0.1%
The purpose is to reduce it below. As this apparatus, a vacuum tank having a polymer discharge screw at the bottom or an extruder having one to three vacuum vent holes (front vents) is used. In any case, the degree of vacuum is selected from the range of 1 to 30 Torr. When a vacuum evaporation tank is used, a wiper type falling film type evaporator is preferable. In this case, it is preferable that the clearance between the blades of the wiper and the tank wall be small, but they should not be in contact with each other. Practically, a clearance in the range of 2 to 5 mm is used. When an extruder is used, the first stage and the second stage deaerator may be integrated. In any case, the deaerator is provided with a jacket and heated using a heat medium. The wall temperature of the deaerator is preferably selected from the range of 200 to 260 ° C. for both the first and second stages. If the temperature is below 200 ° C, degassing will be insufficient, and if it is above 260 ° C, the polymer will be thermally decomposed in the degassing unit. In any case, the content of the monomer in the product polymer will be 0.1% or less. Will not be. The polymer thus obtained is finally extruded from the die as a strand or a plate. At this time, it is preferable to attach a filter such as a 400 mesh wire mesh to the die breaker plate.

In the monomer recovery step, mercaptan and high-boiling compounds such as dimers and trimers are removed from the vapor extracted from the first-stage degasifier, and recovered as a monomer that can be reused in the polymerization reaction. A general distillation column is used as the high-boiling cut device, but it is preferable to introduce a small amount of air from a reboiler and discharge it from a condenser at the top of the column in order to prevent polymerization in the column. The storage tank for the recovered monomer is preferably cooled to 0 to 20 ° C. to form an air atmosphere.

(Examples) Next, the present invention will be described in more detail by way of examples, but these examples do not limit the present invention at all. FIG. 2 shows a flow sheet of the apparatus used in the examples. The specifications of the main equipment are as follows.

Nitrogen mixer: Noritake Static Mixer N10 type inner diameter 41.
2mmφ, 0.77m (12 elements) Monomer blending tank: 400, SUS-316, with paddle wings, with jacket, 2 units, Pall Enflon filter MCY-1001F
RE reactor: 200, SUS-316, with Max Blend wings, with jacket Heater: Noritake Static Mixer N10 type 21mm inside diameter
φ, 1m (30 elements) × 3 units in series, SUS-316 Deaerator: (First stage) Single screw extruder with back vent, (Second stage) Wiper type falling film evaporator inner volume
30 (SUS316), heat transfer area 0.2m ² , screw device for extraction at the bottom High boiling cut tower: 71mm inner diameter, 1.2m length, 1/2 inch SUS Raschig ring filling The polymer evaluation method is as follows .

(1) Measurement of thermal decomposition resistance Product pellets are continuously injection-molded under the following conditions and 200m
Ten plates of m × 70 mm × 3 mm were obtained, and the monomer concentration therein was measured by gas chromatography, and 0.5% or less was regarded as having good thermal decomposition resistance.

Injection molding machine; IS-80EPN-2A manufactured by Toshiba Machine Injection pressure 400kgf / cm ² Injection and dwell time 8 seconds Barrel temperature 270 ° C Mold temperature 55 ° C Cycle time 180 seconds (2) MFR measurement Takara Kogyo Co., Ltd. Using an indexer, ASTM-D1
230 ° C, load 3.80kg, extrusion distance (l) based on 238
2.54cm, cylinder diameter (D) 5.08cm, die nozzle diameter 0.20
The extrusion time (t seconds) was measured under the condition of 95 cm, and was obtained by the following equation.

MFR = (x / 4) D ² l × (specific gravity) × 600 / t (3) Measurement of molecular weight by GPC Three columns of Shimadzu LC-5A, Shimadzu HSG-15, -30 and -50, were used. From the elution curve obtained using a sample in which 0.12 g of the polymer was dissolved in 20 ml of tetrahydrofuran, the elution curve was determined by a calibration curve made using standard polystyrene.

(4) Measurement of Average Molecular Weight by Limiting Viscosity Method A 0.6 g sample was dissolved in 25 ml of chloroform, and the viscosity at 20 ° C. was measured using an Ubbelohde viscometer. The viscosity was determined at several points by diluting the liquid 2 to 6 times, the intrinsic viscosity at infinite dilution [η] was determined from these values, and the average molecular weight M was determined by the following equation.

[Η] = 0.485 × 10 ⁻⁴ × M ^0.8 (5) Measurement of hue Using a SM color computer manufactured by Suga Test Instruments Co., Ltd., the hue of a 3 mm thick plate is used as the b value of the chromaticity coordinate of the hunter. I asked.

(6) Measurement of minute foreign matter Using a HIAC-ROYCO Model 346 automatic particle counter, a solution of 6 g of a sample dissolved in 1000 g of acetone was used.
The number of foreign substances having a size of 0.5 to 10 μ was measured.

Example 1 91.5 parts by weight of methyl methacrylate (MMA), 8.5 parts by weight of ethyl acrylate (EA), 0.36 parts by weight of n-octyl mercaptan (OM), 2,2-azoisobutyronitrile (A)
270 kg of a preparation liquid consisting of 0.011 part by weight of IBN) was charged into a preparation tank which had been previously replaced with nitrogen and cooled to 0 ° C., and stirred and mixed. Of these, MMA and EA are each pumped from raw material tanks.
At a flow rate of 1,500 / h, it was received via a heat exchanger cooled with a -5 ° C. refrigerant and a nitrogen mixer, the nitrogen used in the nitrogen mixer being 5,000 Nl / h. MMA and
While stirring the mixed solution of EA, OM and AIBN were separately charged from the nozzle of the upper lid of the mixing tank and mixed and dissolved. The dissolved oxygen of this mixture was measured using DO meter UC-12-SOL manufactured by Central Science Co., Ltd.
Was below the detection limit (0.6 ppm). 33.3kg / h to the reactor by plunger pump
Thus, continuous feeding was performed while alternately switching the two mixing tanks every 8 hours. The piping from the preparation tank to the reactor was covered with a jacket and cooled with a refrigerant at -5 ° C, and the above-mentioned supercharger was installed to remove dust of 0.2 µ or more. The mixture was allowed to flow directly from the upper mirror nozzle of the reactor onto the surface of the reaction solution.

The top of the reactor established the multi-tube heat exchanger of the heat transfer area of 5 m ^2, also the reactor upper mirror, provided with a jacket for all torso and the lower mirror, the top mirror with water, the body and the lower mirror Was cooled with a heating medium to control the reaction temperature to 150 ° C. A double mechanical seal was used for the shaft seal of the stirrer. Nitrogen was introduced at ³ Nm ³ / h from the connection between the seal box and the upper mirror, and the internal pressure was controlled to 4 kgf / cm ² G by a pressure control valve via the multitubular heat exchanger, and then released.

The amount of the reaction solution in the reactor was 100 kg. At this time, since the half life of the initiator is 3.6 seconds and the residence time is 3.0 hours, the ratio of the half life of the initiator to the residence time (τ _1/2 / θ) is 1/3000. The rotation speed of the stirrer was 200 rpm. The power consumed by stirring obtained by subtracting the aerodynamic power from the value of the power meter of the stirrer motor was 1.0 KW / m ³ . The reaction liquid was drawn out by a gear pump at the bottom of the reactor, and the liquid level in the reactor was controlled to be constant.

The reaction medium from the gear pump was heated to 235 ° C. by flowing a heat medium of 245 ° C. through the heater jacket with an alternating current. The pressure inside the heater was adjusted to 24 kgf / cm ² G by a needle valve at the outlet.

The needle valve was directly connected to the barrel of the third pitch (the number of pitches from the driving side) of the deep groove screw (lead 130 mm) of the single screw extruder, and the monomer vapor was extracted from the back vent hole provided at the second pitch. The polymer from which most of the monomer had been removed was advanced by a screw flight, and was extruded through a compression zone and a metering zone to a feed nozzle of a thin film evaporator. The heating medium temperature of this evaporator jacket is 230
℃, pressure is 10 Torr, clearance between wiper and wall is
At 3 mm, the rotation speed of the wiper was 150 rpm. The degassed polymer was extruded as a strand from the die with a screw at the bottom, cooled in a water bath, and then cut with a cutter to obtain a pellet-shaped molding material. A wire mesh of 400 mesh was attached to the breaker plate of this die.

The yield of the pellets was measured hourly and averaged over a 24-hour period.
kg / h (61.9 ± 0.6% in monomer conversion).

The monomer vapor extracted from the back vent is fed as it is to the lower end of the packed section of the high boiling cut tower, and a high boiling substance (having a high boiling matter concentration of about 40% and a monomer of about 60%) is refluxed at a reflux ratio of 0.5. Cut to the bottom and n
-Monomers containing no high boiling point, such as octyl mercaptan, dimer and trimer, were recovered. In addition, air was introduced at a rate of 1.5 / h from the bottom of the tower and discharged from the condenser at the top of the tower. The recovered monomer was placed in a jacketed tank, purged with air, stored, and reused.

The melt flow rate (MFR) of the molding material obtained as described above is 20.5, the EA in the pellet by gas chromatography analysis is 7.2% by weight, the residual monomer,
Dimers and trimers were 0.08, 0.09 and 0.08% by weight, respectively. The number average molecular weight measured using GPC is 36,
000, the weight average molecular weight was 67,000, and the value of molecular weight distribution Mw / Mn was 1.9. The average molecular weight measured by the limiting viscosity method was 84,000. 0.33 monomer in injection molded plate
% And the hue b value of the molded plate were 0.19, which were all good. The number of micro foreign substances of 0.5 to 10 microns per g of polymer was 10,800 / g.

Example 2 Example 1 was repeated except that 8.5 parts by weight of ethyl acrylate (EA) was replaced by 9.0 parts by weight of methyl acrylate (MA) and the amount of AIBN added was 0.0095 part by weight.
As a result, the monomer conversion was constant at 61.8 ± 0.5% in a continuous operation for 15 days.
In addition, MFR20.1 of the obtained pellet, MA8.2 in the pellet
Wt%, monomer 0.09 wt%, dimer 0.087 wt%,
0.084% by weight of trimer, Mw / Mn1.9, average molecular weight 84,000,
0.38% of monomer in injection molded board, hue of the molded board
The value is 0.19, the number of micro foreign substances of 0.5 to 10 microns is 12,500 / g
It became.

Example 3 Except that the type and amount of the initiator used in Example 1 were changed to 0.0825 parts by weight of lauroyl peroxide, the same operation as in Example 1 was performed. Became constant at 62.2 ± 0.4%. At this time, since the half-life of the initiator is 1.8 seconds, (τ
_1/2 / θ) is 1/6000. In addition, M of the obtained pellet
FR20.3, 7.2% by weight of EA in pellets, 0.094% by weight of monomer, 0.093% by weight of dimer, 0.080% by weight of trimer, Mw / M
n1.9, average molecular weight 84,000, monomer in injection molded plate is 0.
38%, the hue b value of the molded plate was 0.21, and the number of minute foreign matters of 0.5 to 10 microns was 14,000 / g.

Example 4 The same operation as in Example 1 was carried out except that the type and amount of the initiator used in Example 1 were changed to 0.0085 parts by weight of t-butylperoxyisobutyrate, the result was that the initiator was continuously used for 15 days. In the operation, the monomer conversion became constant at 62.2 ± 0.5%. At this time, since the half-life of the initiator is 45 seconds, (τ _1/2 / θ) is 1/240. The MFR of the obtained pellets was 20.0, EA in the pellets was 7.2% by weight, and
085% by weight, 0.086% by weight of dimer, 0.077% by weight of trimer, Mw / Mn 1.9, average molecular weight 84,000, monomer in injection molded plate 0.40%, hue b value of molded plate is 0.20, 0.5 to 10
The number of microscopic foreign matter was 11,300 / g.

Comparative Example 1 Except that the type and amount of the initiator used in Example 1 were changed to 0.0026 parts by weight of di-t-butyl peroxide, the same operation as in Example 1 was performed. Since the half-life of the initiator at this time is 30 minutes, (τ _1/2 / θ) is 1/6. In 15 days of continuous operation, monomer conversion varied between 58% and 65% over a 20-30 hour cycle. The MFR of the obtained pellet was 19 to 23, the EA in the pellet was 7 to 7.5% by weight, the monomer was 0.08 to 0.1% by weight, the dimer was 0.08 to 0.1% by weight, the trimer was 0.08 to 0.12% by weight, and the Mw / Mn 2.0 ~ 2.2, average molecular weight 82,000 ~ 86,000, monomer in injection molded plate is 0.4
0.60.6%, and the hue b value of the molded plate varied between 0.20 and 0.25. In addition, the number of micro foreign substances of 0.5 to 10 microns is about 15,000
It became 0 pieces / g.

[Brief description of the drawings]

FIG. 1 is a graph showing a reversal phenomenon in which, when the special polymerization conditions of the present invention are employed, the conversion becomes lower as the temperature becomes higher at a monomer conversion of 45% or more. FIG. 2 shows an apparatus and a flow sheet used in the embodiment of the present invention, wherein reference numeral 1 denotes a static mixer for mixing nitrogen, 2 denotes a monomer mixing tank, 3 denotes a metering pump, and 4 denotes a metering pump. Filter, 5 is a reactor, 6 is a reflux condenser, 7
Is an automatic pressure control valve, 8 is a mechanical seal, 9 is a gear pump, 10 is a double tube heater with a built-in static mixer, 11 is a needle valve, 12 is a single screw extruder with a back vent, and 13 is a wiper type thin film evaporator. A vessel, 14 is an extraction screw, 15 is a die, 16 is a vacuum pump, 17 is a water tank, 18 is a cutter, 19 is a shifter, 20 is a distillation tower, 21 is a reboiler, 22 is a condenser, and 23 is a recovered monomer tank.

Continuation of the front page (72) Inventor Masaya Nagai 4-7 Kyowa-cho, Nakajo-machi, Kitakanbara-gun, Niigata Examiner in the Nakajo Plant of Kyowa Gas Chemical Industry Co., Ltd. Tsutomu Onodera (56) References JP-A-53-147788 JP-A-49-37993 (JP, A) JP-A-58-53901 (JP, A) JP-A-58-83010 (JP, A) JP-A-61-181813 (JP, A) 53-113882 (JP, A) JP-A-59-45310 (JP, A) JP-A-61-296010 (JP, A) JP-B-52-32665 (JP, B2) (58) Fields investigated (Int. Cl. ⁶ , DB name) C08F 20/14 C08F 2/02

Claims (1)

(57) [Claims] (1) When a monomer mixture containing methyl methacrylate as a main component is continuously bulk-polymerized without using a solvent by one complete mixing type reactor, (1) an inert gas is introduced into the monomer. After reducing the dissolved oxygen in the monomer to 1 ppm or less, (2) using a radical initiator having a half-life at a polymerization temperature of 0.5 to 120 seconds, and (3) 0.5 to 0.5 to 1 m ^{3 of a} reaction solution.
While stirring with a stirrer having stirring power of 20 KW,
(4) The average residence time is set so that the ratio of the half-life of the radical initiator to the average residence time at the polymerization temperature is 1/200 to 1 / 10,000. (5) The monomer conversion rate is 130 to 160 ° C.
A process for producing a colorless and transparent methacrylic polymer excellent in moldability, characterized in that the polymerization is carried out to 45 to 70%.