JPS6363563B2 - - Google Patents

Description

DETAILED DESCRIPTION OF THE INVENTION The present invention provides a high-pressure ethylene polymer or copolymer production process using a tubular reactor, in which the amount of polymerization initiator injected into the reactor is determined to stabilize the polymerization reaction. The present invention relates to a method for controlling and polymerizing.

Ethylene or ethylene and a comonomer mixture copolymerizable with ethylene (hereinafter referred to as comonomer mixture) are heated in a tube reactor or a tank reactor to
A widely known method is to obtain an ethylene polymer or copolymer (hereinafter referred to as ethylene (co)polymer) by polymerizing in the presence of a polymerization initiator at a pressure of 4000 kg/cm ² and a temperature of 120 to 350°C. It is being

In the polymerization method using a tank reactor, the reactor is equipped with a stirrer and kept in a state close to complete mixing, so that the injected polymerization initiator is almost uniformly dispersed and does not disperse in localized high temperature areas ( This prevents the formation of hot spots) and avoids runaway decomposition reactions.

On the other hand, in the polymerization method using a tubular reactor, ethylene or the comonomer mixture and the ethylene (co)polymer flow in the reaction tube in a state close to a plug flow (pulsating flow), so the ethylene is attached to the wall of the reaction tube. The (co)polymer may adhere and form hot spots, leading to runaway decomposition reactions. In order to prevent such a situation, an outlet valve (hereinafter referred to as an L/D valve) is provided at the outlet of the reactor to apply pressure pulsations.

However, pressure pulsations change the mass flow rate of ethylene, comonomer mixture, or ethylene (co)polymer mixture in the reaction tube, quantitatively delivering the polymerization initiator to points along the length of the reactor. Even if the polymerization initiator is injected at Product quality of the polymer or copolymer may be impaired. In more extreme cases, it may even progress to a decomposition reaction.

An object of the present invention is to provide a method for polymerizing ethylene or a comonomer mixture using a polymerization initiator, in which the polymerization initiator is The object of the present invention is to provide a method for stably performing a polymerization reaction.

As a result of intensive efforts, the present inventors have discovered that the polymerization reaction can be stabilized by the following method.

That is, in the present invention, ethylene or a comonomer mixture is introduced into the inlet of a tubular reactor, and a polymerization initiator is injected at at least one point along the length of the reactor to increase the reaction pressure to 1000 to 4000 kg/kg. cm ² , reaction temperature
In a method for obtaining an ethylene (co)polymer by polymerizing at 120 to 350°C, a polymerization initiator control valve is provided in the polymerization initiator introduction pipe, and a L/L/L valve is provided at the outlet of the reactor.
Detect the time (delay time) from the time when valve D is opened until the drop in reaction pressure spreads to the polymerization initiator injection point, and after the delay time has elapsed, open the polymerization initiator control valve to an arbitrary opening degree, Polymerization initiator is injected in proportion to the increased amount of ethylene or comonomer mixture generated by opening the L/D valve, and the reaction pressure at the polymerization initiator injection point begins to drop and reaches the lowest point. After a certain period of time, close the polymerization initiator control valve to the desired opening.
The amount of polymerization initiator injection is reduced according to the amount of decrease in ethylene or comonomer mixture caused by closing the L/D valve, and the time required for the reaction pressure at the polymerization initiator injection point to recover to the normal value. By returning the opening degree of the polymerization initiator control valve to its original position after the lapse of time,
This is a method for producing an ethylene (co)polymer, characterized by making the concentration of a polymerization initiator per ethylene or comonomer mixture in the reactor uniform.

In the present invention, a polymerization initiator compensation controller such as a computer can be used, for example, so that the opening degree and time of the polymerization initiator control valve can be set to arbitrary optimal values. The polymerization initiator control value can also be automatically controlled by setting the opening degree and time of the polymerization initiator controller to zero. It is desirable to provide the polymerization initiator control valve in the middle of the polymerization initiator injection pipe close to the polymerization initiator injection point so that pressure changes at the polymerization initiator injection point of the reactor can be quickly dealt with.

Polymerization initiators used in the present invention include oxygen; diisopropyl peroxydicarbonate;
t-Butyl neodecanate, t-butyl perpivalate, dilauryl peroxide, dibenzoyl peroxide, di-t-butyl peroxide, t-butyl peroxy 2-ethylhexenoate, t-butyl perisobutyl Organic peroxides such as ester, t-butyl peracetate, t-butyl perbenzoate, t-butyl hydroperoxide; azo compounds such as azobisisobutyronitrile, azobis-2,2-diphenylacetonitrile, etc. Species or mixtures of two or more species can be used.

As the chain transfer agent, all those commonly used in the production of polyethylene can be used, such as ethane, propane, butane, pentane, hexane, heptane, propylene, butene-1, isobutylene, hexene, etc. alone or in combination of two or more. Can be used in combination.

Comonomers that can be copolymerized with ethylene in the method of the present invention include vinyl compounds, alkyl esters of acrylic acid and methacrylic acid, and the like.

It is not necessary to introduce the entire amount of the ethylene to be fed from one end of the reactor, and a portion thereof can be introduced unheated or cooled to a point in the middle of the reactor.

The present invention will be explained below with reference to the drawings.

FIG. 1 is a flow diagram showing an example of an ethylene (co)polymer manufacturing process using a tubular reactor according to the present invention.

The mixture of feed ethylene 1, chain transfer agent 3, and optionally comonomer 4 enters compressor 5 and is compressed from 1000 to 4000
It is compressed to kg/cm ² and passed through a pipe 6 to a tubular reactor 8.
to go into.

A portion of the compressed ethylene mixture is passed through pipe 7 without heating or, if necessary, by cooling, to terminate the polymerization reaction along the length of the pipe and, although not shown in this figure, pass through the reactor. It is cooled by the refrigerant flowing through the enclosing outer tube, and can be introduced into at least one location where the temperature inside the reactor decreases.

After the supplied ethylene is preheated to the polymerization starting temperature, the polymerization initiator is injected from the polymerization initiator supply tank 9 through the polymerization initiator injection pump 10 through the polymerization initiator control valve 11 to start polymerization. After the reaction temperature reaches the first peak, it is further cooled by the refrigerant flowing through the outer tube and by the unheated ethylene supplied from the pipe 7, and the reaction temperature drops. Further, the polymerization initiator that has passed from the polymerization initiator supply tank 12 through the polymerization initiator injection pump 13 is supplied to the polymerization initiator control valve 1.
4, and after reaching the second peak as well, the reaction temperature gradually decreases.

When oxygen is used as a polymerization initiator, it is mixed with ethylene and supplied to the compressor 5 via piping 2.
The ethylene (co)polymer and unreacted ethylene produced in the reactor pass through the L/D valve 18 and the pipe 19, and are separated by a high pressure separator and a low pressure separator (not shown in this figure). The ethylene (co)polymer is discharged from the system, and unreacted ethylene is recycled and used. In order to prevent the ethylene (co)polymer from adhering to the inner pipe of the reactor 8, the L/D valve 18 is operated by a periodic signal from the L/D valve controller 15, as shown in FIG. This creates pressure pulsations by opening and closing the valve.

A signal to open the L/D valve from the L/D valve controller is transmitted to the polymerization initiator compensation controllers 16 and 17.

The broken line (b) in FIG. 2 shows how the pressure at the reactor outlet located upstream of the L/D valve 18 changes. The solid line in the figure shows how the reaction pressure changes at the polymerization initiator injection point. Further, C in FIG. 2 shows the inlet of the polymerization initiator control valve. There is a delay in the propagation of reaction pressure fluctuations upstream of the reactor, and this delay is generally between 0.1 and 0.2 seconds per 100 m of reactor.

This delay time TM1 is set in the polymerization initiator compensation controllers 16 and 17, and after TM1, the polymerization initiator control valves 11 and 14 are opened to the depth of the arbitrary opening degree D1, so that the reaction pressure at the polymerization initiator injection point is The polymerization initiator control valve is closed to the desired opening degree (D1 plus D2) after the time when it starts to drop and reaches the lowest value (TM2). Furthermore, after the time TM3 during which the reaction pressure at the polymerization initiator injection point recovers to the normal value, the polymerization initiator control valve is opened to the opening degree D2 and returned to its original position.

There is generally a slight propagation delay in transmitting the output from the polymerization initiator compensation controllers 16 and 17 to the polymerization initiator control valve, and therefore the timing to open and close the polymerization initiator control valve is made earlier by the propagation delay time. It is necessary to output to the polymerization initiator control valve. If the propagation delay time is larger than TM1, a signal to open the previous L/D valve is output to the polymerization initiator controller, and the signal is calculated from the sum of the L/D valve opening/closing cycle and the delay time TM1. It is also possible to output to the polymerization initiator control valve after a time (corresponding to TM1') minus the propagation delay time.

Because of this propagation delay, TM1, TM2,
It is considered that TM3 or TM1' can be set at any time. In addition, it is necessary to adjust the valve opening degree of the polymerization initiator control valve depending on the variation range of the reaction temperature, and the valve opening degrees D1 and D2 can also be set arbitrarily from the polymerization initiator compensation controllers 16 and 17. It has been devised.

According to the method of the present invention, the concentration of the polymerization initiator in the ethylene or comonomer mixture became uniform, the range of change in reaction temperature downstream of the polymerization initiator injection point became significantly narrower, and the reaction state became extremely stable.

By making the polymerization initiator concentration uniform, it is possible to prevent the formation of hot spots in the reactor.
This has the advantage of avoiding coloration of the ethylene (co)polymer and runaway decomposition reactions.

The concentration of the polymerization initiator in the reactor can be homogenized to some extent by narrowing the opening of the polymerization initiator control valve and making the polymerization initiator injection pressure considerably higher than the reactor pressure. In this case, the discharge pressure of the polymerization initiator injection pump becomes high, and the polymerization initiator often leaks from the pump, resulting in a disadvantage that it becomes difficult to continue the normal reaction. Moreover, the risk of leakage from the joints of the polymerization initiator injection pipe increases, and increasing the discharge pressure of the pump increases the power consumption of the pump, which is not preferable.

Furthermore, according to the method of the present invention, by narrowing the range of change in reaction temperature, products with high commercial value and excellent haze and gloss of the resulting ethylene (co)polymer can be produced. .

Furthermore, according to the method of the present invention, the concentration of the polymerization initiator in ethylene or the comonomer mixture becomes uniform, so that the value of the polymerization initiator efficiency is the same as that of the present invention under the same reaction pressure and temperature conditions as the present invention. It can be made smaller than without it. In other words, there is an economical advantage that the manufacturing cost of the ethylene (co)polymer can be reduced.

Next, the present invention will be specifically explained with reference to Examples. However, the present invention is not limited thereto.

Example 1 Per million parts by weight of ethylene from the inlet of the compressor,
Inject 30 parts by weight of oxygen, use propylene as a chain transfer agent, maintain the concentration in the reaction system at 1.5 mol%, compress it to 2600 kg/cm ² with a compressor, supply it to a reactor with a pipe length of 500 m, and preheat it. 10 parts by weight of t-butyl perpivalate was compressed to 2800 kg/cm ² using a polymerization initiator injection pump and fed from the polymerization initiator injection point located 100 m from the reactor entrance. Continuous polymerization was carried out at a maximum reaction temperature of 300°C.

Since the propagation delay time from the polymerization initiator compensation controller to the polymerization initiator control valve is 2.2 seconds, and the pressure pulsation delay from the L/D valve to the polymerization initiator injection point is 0.7 seconds,
The setting time of the polymerization initiator compensation controller is TM1′=13.5
seconds, TM2 = 1.5 seconds, TM3 = 3.0 seconds. In addition, the operating output to the polymerization initiator control valve is D1 = 12%, D2 =
By setting the temperature to 3%, the reaction temperature fluctuation reached its maximum at a position of 150 to 200 m from the reactor inlet, but the fluctuation range was about 8°C. The L/D valve controller is set so that the pressure pulsation at the inlet of the reactor is 260Kg/ ^cm2 .
Output was made to the L/D valve every 15 seconds. At this time, 200,000 parts by weight of polyethylene was obtained, and the product density was 0.
921Kg/cm ³ , melting index is 4g/10min, and the optical properties of the product film are haze value 3.5% and gloss value
It was 13.2%. Polymerization initiator efficiency excluding oxygen is
It was 0.050g/1Kg polyethylene.

Comparative Example 1 Example 1 without using a polymerization initiator compensation controller
Polymerization of ethylene was carried out under the same conditions. At this time, t-butyl perpivalate as a polymerization initiator was required in an amount of 12 parts by weight per 1 million parts by weight of ethylene. As in Example 1, the variation in reaction temperature reached its maximum at a position of 50 to 100 m from the polymerization initiator injection point, and the range of variation reached approximately 25°C, resulting in an unstable reaction state. At this time, 200,000 parts by weight of polyethylene was obtained per 1,000,000 parts by weight of ethylene, and had a density of 0.921 g/cm ³ and a melting index of 4 g/10 min, which were the same as in Example 1, but a haze value of 3.9% and a gloss The optical properties of the product film deteriorated with a value of 13.0%. The polymerization initiator efficiency excluding oxygen at this time was 0.060 g/1 kg polyethylene.

Example 2 The same reactor as shown in Example 1 was used, propylene was used as a chain transfer agent, the propylene concentration in the system was maintained at 2.1 mol%, and the compressor was used to generate 2700 Kg/cm ²
650,000 parts by weight per 1 million parts by weight of ethylene is supplied to the inlet of the reactor.
15 parts by weight of t-butyl perpivalate, 15 parts by weight of t-butyl perbenzoate, and 10 parts by weight of di-t-butyl peroxide were supplied from the injection point of the first polymerization initiator located at a point 100 m away from the injection point of the first polymerization initiator. The peak temperature is 300℃, and after the peak temperature has passed, the 180℃ refrigerant flows outside the reaction tube and is 290m from the reactor inlet.
350,000 parts by weight, 50℃ to the cold ethylene supply section located at
of ethylene and cooled to 205°C, then 28 parts by weight of t-butyl perbenzoate and 35 parts by weight of di-t-butyl peroxide were added from the second polymerization initiator injection port located 300 m from the reactor entrance. was injected and polymerized at a second peak temperature of 300°C.

The setting time and operation output of the first polymerization initiator compensation controller are TM1' = 13.3 seconds, TM2 = 1.5 seconds, TM3 = 3
seconds, D1 = 12%, D2 = 3%, and the setting time and operation output of the second polymerization initiator compensation controller are TM1' =
13.0 seconds, TM2=1.3 seconds, TM3=2.8 seconds, D1=15%,
By setting D2 = 5%, the reaction temperature fluctuation was maximum at 50 to 100 m from the first polymerization initiator injection point, and the fluctuation range was about 12°C. Moreover, the maximum temperature was reached at a position 50 m downstream from the second polymerization initiator injection point, and the fluctuation range was 15°C. At this time, the density is 0.922g/cm ³ ,
Melt index 3g/10min, Haze 3.5%, Gloss 13.2
% polyethylene was obtained in an amount of 240,000 parts by weight. Polymerization initiator efficiency was 0.43g/1Kg polyethylene.

Comparative Example 2 On the other hand, when carried out under the same conditions as Example 2 without using a polymerization initiator compensation controller, the maximum temperature fluctuation width in the first reaction zone was 28°C, and the maximum temperature fluctuation width in the second reaction zone was 28°C. The temperature reached 40℃, and the reaction continued to be unstable. The polyethylene obtained at this time has a density of
0.922g/cc, melting index 3g/10min, haze value 4.0
%, a gross value of 12.8% was obtained, and the polymerization initiator efficiency was
It was 0.57g/1Kg polyethylene.

The physical properties of the produced polyethylene were measured according to the following procedure.

Density: ASTM D 1505-68 Melting index: ASTM D 1238−70 Cloudiness: ASTM D 1003-61 Glossiness: ASTM D 2457-651

[Brief explanation of drawings]

FIG. 1 is a schematic diagram of the polymerization apparatus of the present invention, and the main symbols in the diagram are as follows. FIG. 2 shows the relationship between the L/D valve input and the polymerization initiator control valve input. 51 Ultra-high pressure compressor, 8: Tubular reactor, 9,1
2: Polymerization initiator supply tank, 10, 13: Polymerization initiator injection pump, 11, 14: Polymerization initiator control valve, 15: L/D valve controller, 16, 17: Polymerization initiator compensation controller.

Claims (1)

[Claims] 1. Introducing ethylene or a mixture containing ethylene and a comonomer copolymerizable with ethylene into the inlet of a tubular reactor, and injecting a polymerization initiator at at least one point along the length of the reactor. Reaction pressure 1000~
Polymerize at 4000Kg/cm ² and reaction temperature 120~350℃,
In the method for obtaining an ethylene polymer or copolymer, a polymerization initiator control valve is provided in the middle of the polymerization initiator introduction pipe, and the reaction pressure is controlled to decrease from the time when the outlet valve provided at the reactor outlet is opened. After the time when the polymerization initiator injection point is affected, the polymerization initiator control valve is forcibly opened to an arbitrary opening degree, and after the time when the pressure inside the reactor at the polymerization initiator injection point reaches the lowest point, the polymerization initiator control valve is opened. By closing the valve to an arbitrary opening and then opening the polymerization initiator control valve to the original opening after the time has elapsed for the pressure to recover to its normal value, the ethylene polymer or copolymer is released onto the inner wall of the reactor. The polymerization initiator is added to the reaction in response to changes in the mass flow rate of ethylene or a comonomer mixture copolymerizable with ethylene in the reactor caused by pressure pulsations caused by opening and closing of the reactor outlet valve to prevent adhesion. A method for producing an ethylene polymer or copolymer, characterized by injecting it into a container.