CN114989340A - Olefin polymerization method - Google Patents

Abstract

The present invention relates to an olefin polymerization method in which ethylene is copolymerized with at least one C3 to C12 alpha-olefin. The method adopts the series connection mode of two-stage reactors, and then enters the separation system to obtain polymer products and circulating streams. The specific steps are: (a) feeding the reaction material and catalyst into the tank reactor, and discharging the material after the reaction to obtain the first polymer solution; (b) feeding the first polymer solution into the tubular reactor Continue the reaction inside, and the obtained second polymer solution is fed to the solution preheater to obtain the third polymer solution with the same or higher temperature; (c) the third polymer solution is subjected to separation operation in the separation system to obtain polymerization product and a recycle stream containing ethylene, solvent, alpha-olefin comonomer. The method can effectively utilize the reaction heat, save the consumption of public works, and effectively increase the yield of the polymerization reaction, and has the advantages of good economic benefit and high production efficiency.

Description

A kind of olefin polymerization method

technical field

The present invention relates to the field of olefin polymerization, and more particularly, to a method for producing ethylene and a copolymerization of alpha-olefin, the method comprising forming a tank reactor in a first stage and a second alpha-olefin with at least one C3 to C12 The polymerization reaction occurs in a series combination of tubular reactors, and then enters the separation system to obtain the product and part of the recyclable components.

Background technique

Ethylene and α-olefin copolymers with high comonomer content are a class of high-performance thermoplastic elastomers that are extremely widely used. Compared with ordinary polyolefin plastics, their intramolecular comonomer content is higher and the density is lower. It is widely used in automotive, packaging, wire and cable, polymer modification, seals, medical and other fields. At present, there are gas phase method, slurry method and solution method for the production of ethylene and α-olefin copolymer in industry. Compared with the first two methods, the ethylene and α-olefin polymers produced by solution polymerization have the advantages of lower product density and better elasticity. "Solution polymerization" refers to a polymerization process in which a polymer is dissolved in a liquid polymerization system, such as an inert solvent or one or more monomers or blends thereof, solution polymerization is a homogeneous liquid polymerization system. Conventional solution polymerization processes are carried out at temperatures of 40-160° C. and pressures below 13 MPa, and operate in polymerization systems in the presence of more than 65 wt% inert solvent.

In solution polymerization, the ethylene monomer, comonomer, catalyst and co-catalyst are all dissolved in the solvent for the reaction, the resulting polymer product also remains dissolved in the solvent, and a polymer solution is formed after the reaction occurs. After the reaction is completed, the outlet stream of the reactor is sent to the separation system for separation and recovery to obtain a polymer product, and the available components such as ethylene monomer, solvent and comonomer are recycled to the reaction part to continue to participate in the reaction. One of the main differences between olefin polymerization production processes is reflected in the different reactor combinations. Different reactor combinations have different production efficiency and polymer product performance characteristics. The reaction temperature and component concentration in the ideal fully mixed-flow tank reactor are uniform and consistent with the outlet stream. Whereas a plug flow tubular reactor temperature and component concentrations exist along the distribution of the reactor, as for an adiabatic tubular reactor, the tubular reactor temperature gradually increases from the inlet to the outlet as the polymerization proceeds.

There are several types of reactors and designs of reactor combinations for processes related to olefin production. CN111630071A discloses a method for copolymerization of ethylene and α-olefin in a solution polymerization reactor. The polymerization reaction is first performed in a kettle type reactor, and then the effluent is discharged into a solution preheater. After reaching the set temperature , which feeds the heated stream into the separation system. CN 112500510A relates to a polymerization method for producing polyethylene by solution method. The method adopts two kettle type reactors in series to carry out polymerization reaction, and then enters into a separation system; adding reactors is a well-known method for improving production efficiency and regulating products. KR 101590998B discloses a method for continuous solution polymerization of olefins. The method is prepolymerized in a first-stage reactor (tank reactor), and then sent to a second-stage reactor (tubular reactor) in series with the tank reactor. The main polymerization reaction is carried out in the reactor), and the purpose of the invention is to add a static mixer composed of a Kenics mixer and a Sulzer mixer in the tubular reactor to strengthen the mixing effect, thereby obtaining a more homogeneous polymer. The polymerization temperature of the kettle reactor and the tubular reactor is the same. For example, in Example 1, the temperature of the first stage and the second stage reactor are both 120 ° C, and the reaction heat of the tubular reactor cannot be fully utilized to reduce the Utilities consumption. Patent CN 107614541A describes a method for continuous solution polymerization, which obtains a polymer solution through a polymerization reactor, passes through two heat exchangers successively before feeding it to the separator, and heats it to a certain temperature, wherein the second heat exchanger Use utilities. Patent CN105377902A proposes a method for improving the energy utilization of solution polymerization facilities. After the polymer solution obtained in the reactor is increased in temperature through a heat exchanger, the pressure is reduced and the solution is fed to a gas-liquid separator to reduce energy consumption by directly reusing part of the gas phase. Patent CN1283204A proposes a treatment design for increasing the polymer content in an olefin solution polymerization process, and a similar heat exchanger is also provided to increase the temperature of the polymer stream entering the separator. The olefin solution polymerization process uses a lot of solvents. Before the polymer solution discharged from the reactor is separated, the current design requires a heater at the reactor outlet to increase the temperature to enhance the separation effect, resulting in high energy consumption in the solution polymerization process.

Aiming at the problems that heat and utilities are consumed before solvent separation of the olefin polymer solution and the reaction yield is still low, a new process needs to be found, which can effectively increase the reaction yield and reduce the utility energy consumption of the process.

definition:

"At least one C3 to C12 alpha-olefin" refers to a comonomer, ie, in addition to ethylene monomers, may be selected from alpha-olefins having 3 to 12 carbon atoms.

"Alpha-olefin" refers to a single olefin whose double bond is at the end of the molecular chain, such as 1-butene, 1-hexene, 1-octene.

A "polymer" is a polymer of ethylene with one or more C3 to C12 alpha-olefin comonomers, including terpolymers.

"Continuous" refers to a system that operates without interruption. For example, the reactants are continuously introduced into one or more reactors, and the polymer product is continuously withdrawn.

"Solution preheater" refers to a heat exchanger arranged between the reactor outlet and the separation system inlet for changing the polymer solution.

A "heat exchanger" is a device used to transfer heat from a hot fluid to a cold fluid to meet specified process requirements. It occupies an important position in chemical, petroleum, power, food and many other industrial productions, especially in chemical production, heat exchangers can be used as heaters, coolers, condensers, evaporators and reboilers, etc., and are widely used.

"Separation system" is a system including multi-step separation and recovery operations, the purpose of which is to separate the polymer obtained by the polymerization reaction from the polymer solution and obtain recyclable monomers, comonomers, and solvents. components, and/or remove impurity components. The separation system of olefin polymer production process usually includes multi-stage flash evaporation, devolatilization, circulation, extrusion and other processes. For example, when only two-stage flashing is considered, the heated stream from the outlet of the solution preheater is fed to the first-stage flash tank for separation, and the reconstituted stream of the still liquid is sent to the second-stage flash tank to continue. to separate.

"Flash separation" means a separation step that results in phase separation by pressure reduction.

A "tank reactor" is a cylindrical reactor with a low aspect ratio, and a stirring (eg, mechanical stirring) device is usually provided in the reactor. When the height-diameter ratio is large, multi-layer stirring blades can be used. During the reaction process, the material needs to be heated or cooled, a jacket can be set at the wall of the reactor, or a heat exchange surface can be set in the reactor, or heat exchange can be carried out through external circulation.

A "tubular reactor" is a continuous operation reactor with a tubular shape and a large aspect ratio, which belongs to a plug flow reactor.

"Adiabatic reactor" refers to a reactor that does not exchange heat with the outside world. The material is adjusted to a specified temperature and then sent to the reactor. The reaction temperature is only related to the feed flow rate, feed temperature and reaction heat, and is not affected by the external environment. influences.

SUMMARY OF THE INVENTION

In order to overcome the deficiencies of the prior art, the present invention provides a method for copolymerizing ethylene with at least one C3 to C12 α-olefin to obtain a copolymer of ethylene and C3 to C12 α-olefin, which adopts a two-stage reaction The reactor is connected in series, the first stage reactor is a tank reactor, the second stage reactor is a tubular reactor, and then the tubular reactor discharge material enters the separation system through the solution preheater or directly discharges the material to the separation system, Obtain polymer product and recycle stream; Concrete steps are:

(a) feeding catalyst, co-catalyst, ethylene, α-olefin comonomer and solvent to the tank reactor, after reacting in the tank reactor, discharging to obtain the first polymer solution;

(b) feeding the first polymer solution into the tubular reactor to continue the reaction, and discharging to obtain the second polymer solution;

(c) feeding the second polymer solution to a solution preheater to obtain a third polymer solution with a higher temperature, and then feeding it to the separation system; or directly feeding the second polymer solution to the separation system ;

(d) The polymer solution entering the separation system is subjected to a separation operation in the separation system to obtain a polymer product and a recycle stream comprising ethylene, a solvent, and α-olefin comonomer.

The inlet of the kettle-type reactor is provided with a feeding pipe for feeding into the kettle-type reactor, including a main catalyst feeding pipe, a co-catalyst feeding pipe, a solvent feeding pipe, an ethylene feeding pipe, and a comonomer feeding pipe. The feed pipe is also provided with a circulation line for receiving the circulating stream from the separation system. These feed pipes or circulation lines can be connected to the reactor individually, or a part or all of them can be combined into a feed pipe to connect the tank reactor ,;

The tubular reactor is also provided with feed lines for replenishing fresh ethylene monomer, solvent, C3 to C12 alpha-olefin comonomers, and also with recycle lines for receiving recycle streams from the separation system, which feed The feed pipe or circulation line can be connected to the tubular reactor alone, or can be partially or fully combined into a feed pipe to connect the tubular reactor;

It is characterized in that the discharge temperature of the tubular reactor is higher than the discharge temperature of the tank reactor.

According to the method of the present invention, the tubular reactor is an adiabatic reactor.

According to the method of the present invention, the discharge temperature of the second polymerization solution of the tubular reactor is at least 5°C higher than the discharge temperature of the first sentence and the solution of the tank reactor.

According to the method of the present invention, the temperature of the tubular reactor and its temperature distribution can be controlled by regulating the flow rates of ethylene, solvent and α-olefin comonomer added.

According to the method of the present invention, the temperature of the tubular reactor and its temperature distribution are controlled by regulating the temperature of the added ethylene, solvent and α-olefin comonomer.

According to the process of the present invention, the feed to the tubular reactor is from fresh or recycled ethylene monomer, solvent, C3 to C12 alpha-olefin comonomer or a mixture of the above components, wherein the recycled stream is from the tubular reaction The separation system downstream of the device.

According to the method of the present invention, the mass ratio of the total amount of ethylene, solvent, and α-olefin comonomer added to the tubular reactor to the output of the first-stage reactor is between 1/50 and 1/2, preferably 1/50 to 1/2. Between 10 and 1/5.

According to the method of the present invention, the mass ratio of the polymer product of the tank reactor to the newly added polymer product of the tubular reactor is between 1/5 and 40, preferably between 1/2 and 20, more preferably between 2 and 10.

According to the method of the present invention, the pressure of the tubular reactor is between 30 and 200 bar, preferably between 35 and 50 bar.

According to the method of the present invention, the outlet temperature of the tubular reactor is 125-240°C, preferably 140-180°C.

According to the process of the present invention, the outlet temperature of the tubular reactor is controlled by the flow of fresh ethylene entering the tubular reactor through a control loop.

According to the method of the present invention, the ratio of the residence time of the tubular reactor to the residence time of the tank reactor is between 1/50 and 1/2, preferably between 1/20 and 1/5.

According to the method of the present invention, the α-olefin is a C3-C12 olefin, preferably a C4-C8 α-olefin.

According to the method of the present invention, the solvent is selected from C4-C12 chain alkanes or cycloalkanes, C6-C9 aromatic hydrocarbons or mixed solvents thereof, preferably C5-C8 chain alkanes or cycloalkanes.

According to the method of the present invention, the residence time of the tank reactor is 2-60 min, preferably 5-40 min, more preferably 8-20 min.

According to the method of the present invention, the pressure in the tank reactor is 30-200 bar, preferably 35-50 bar.

According to the method of the present invention, the temperature in the tank reactor is 120-240°C, preferably 125-155°C.

According to the method of the present invention, the temperature rise of the kettle reactor is 80-220°C, that is, the reaction discharge temperature is 80-220°C higher than the reaction feed temperature, preferably 130-180°C.

It is found in the present invention that by using the series combination form of the tank reactor in the first stage and the tubular reactor in the second stage, the temperature of the polymer solution at the outlet of the reactor can be increased by the heat of polymerization while the reaction yield is increased, thereby reducing the separation rate. Utility consumption for the solution preheater upstream of the system. It is well known that there is an optimal reaction temperature range for polymerization catalysts. As shown in the graph of catalyst activity and temperature in Figure 1, the activity of the catalyst increases with the increase of the reaction temperature, and at first when the temperature is low, it increases continuously until a certain temperature is reached. After the lower catalyst reaches the optimum activity, the catalyst activity begins to decrease with the continuous increase of temperature. Therefore, in order to improve the production efficiency, it is necessary to optimize the selection of the reactor, the combination of the reactor and the reaction temperature. The temperature of the tank reactor was uniform and equal to the outlet temperature. The reaction temperature of the tank reactor can be set in the optimal activity zone, and its space-time yield (the amount of reaction products obtained per unit residence time and unit reaction volume) is higher than that of the tubular reactor. However, it is impossible to take into account the high reaction outlet temperature and the high reaction rate only by setting up the tank reactor. The first-stage reactor of the present invention adopts a tank-type reactor and the reaction temperature is set in the temperature range with higher activity. The second-stage reactor adopts a tubular reactor instead of a tank reactor, because there is a temperature distribution in the tubular reactor. As the reaction proceeds, the temperature gradually increases. By adjusting the reaction temperature, the catalyst activity can be achieved in the tubular reactor. The reactor discharge temperature was higher while still maintaining a higher level. The present invention further finds that the reaction temperature distribution of the tubular reactor and the temperature of the polymer solution at the reaction outlet can be controlled by adjusting the flow rate and/or temperature of the feeds such as ethylene, comonomer and solvent entering the tubular reactor.

Combining the above technical solutions, the solution of the present invention has the following advantages:

According to the characteristics of the copolymerization reaction process of ethylene and α-olefin, the reaction yield is improved and public works are saved. The tank reactor reaction stream is combined with supplemental ethylene monomer, alpha-olefin comonomer, solvent, or a combination of the above in the form of a series combination of tank reactors in the first stage and tubular reactors in the second stage. The mixture is mixed and entered into the tubular reactor to continue the reaction, so that the reaction yield can be as high as possible, and the reaction heat can be used to provide energy for the subsequent separation system.

Description of drawings

Fig. 1 is a graph showing the variation of catalyst activity with temperature;

Fig. 2 is the schematic flow chart of the method of the present invention;

Figure 3 is a graph of the reaction rate as a function of residence time.

Detailed ways

The present invention will be further elaborated and described below in conjunction with specific embodiments. The described embodiments are merely exemplary of the present disclosure and do not delineate the scope of limitation. The technical features of the various embodiments of the present invention can be combined correspondingly on the premise that there is no conflict with each other.

As shown in FIG. 2 , the present invention provides a solution polymerization method for copolymerizing ethylene and α-olefin. The polymerization device used in the method mainly includes a tank reactor (1), a tubular reactor (2), and a solution preheater. (3) and the separation system (4), the kettle type reactor is provided with a feed pipe, and the ethylene monomer, solvent and α-olefin comonomer are all merged into one feed pipe to connect the reactor, and the kettle type reactor is also A recycle line is provided for receiving the recycle stream from the separation system; the tank reactor outlet is connected to the tubular reactor inlet, and the first polymerization solution obtained from the tank reactor outlet is combined with fresh and/or recycled After the mixed streams of ethylene monomer, α-olefin comonomer and solvent are mixed, they enter the tubular reactor for reaction; the second polymer solution obtained from the outlet of the tubular reactor passes through the solution preheater to obtain a higher temperature The third polymer solution obtained is then transported to the separation system, and the separation system performs a series of operations on the mixed materials to separate, recover and recycle the materials, and finally obtain polymer products and circulating streams.

The chemical reaction rate refers to how fast a chemical reaction proceeds. It is usually expressed in terms of the change (decrease or increase) of the concentration of reactants or products per unit time (such as per minute, per hour), and the reaction rate is related to the nature and concentration of reactants, temperature, pressure, catalyst, etc. , if the reaction is carried out in solution, it is also related to the nature and amount of the solvent. The reaction rate can be controlled by controlling the reaction conditions to achieve certain goals, such as improving the reaction yield. The reaction yield is an important indicator for evaluating the chemical production process, that is, the target product quality (or molar amount) obtained per unit time.

During the polymerization reaction, the rate of chemical reaction is proportional to the activity of the catalyst. Catalyst activity refers to the quantity (or mass) of the raw material reactants converted per unit volume (or mass) of catalyst per unit time. The higher the catalyst activity, the faster the reaction rate. The catalyst activity is mainly determined by the temperature. As mentioned above, the relationship between the catalyst activity and the temperature can be known, as shown in FIG. 1 .

The polymerization catalyst can be any catalyst known in the art that can copolymerize ethylene and α-olefin comonomer, including ZN catalyst, TiCl ₃ -A1Et ₂ Cl catalytic system, wherein the main catalyst component is an amino titanium compound ((R ¹ R ² N) _4-n TiYn, the co-catalyst component is aluminoxane compound, and the main catalyst component is a transition metal compound MR _I (OR') _m X _n-(1+m) , and the co-catalyst component is aluminoxane Alkane compound and the third component of a dihydroxy-containing organic compound, etc.; also include metallocene catalysts, including the main catalyst is a constrained geometry catalyst (CGC), and the co-catalyst is methylaluminoxane (MAO) and the like.

As known in the art, a small amount of a substance with a large chain transfer constant is often added to the polymerization system to reduce the molecular weight of the polymer, ie a molecular weight regulator, which can optionally be used in a solution polymerization reactor to control the molecular weight of the copolymer. There are many suitable chain transfer agents, such as hydrogen.

As known in the art, after the polymerization reaction enters the separation stage, the reaction needs to be terminated first, that is, the catalyst is deactivated, in order to avoid the polymer molecular weight being too large or the uncontrollable exotherm of the reaction in the separation stage, which can be optionally used for solution polymerization. The substance in the reactor to deactivate the catalyst. There are many deactivators, such as water.

A large amount of solvent is required in solution polymerization to dissolve the ethylene monomer, alpha-olefin comonomer, catalyst, molecular weight regulator, and polymer product under polymerization conditions. And the solvent needs to be inert to the catalyst system and reactants, and be stable during the reaction, so straight-chain, cyclic or branched alkyl groups with 6 carbon atoms or two or more of them can be selected. of mixtures, such as n-hexane.

In the chemical reaction process, there are many types of reactors, and they have their own characteristics. The kettle type reactor has the characteristics of uniform temperature in the whole kettle, and the outlet temperature is the same as the reaction temperature. The flow rate and temperature of the inlet stream can be adjusted to control the reaction temperature in the kettle, so that the catalyst is at a high activity temperature. It has the characteristics of temperature distribution for a long time. Not only there are some areas to make the catalyst activity high, but also the heat of polymerization reaction is fully utilized, which can take into account the high outlet temperature and the reaction yield.

According to the characteristics of each reactor, this method selects two-stage reactors in series, the first stage is an adiabatic kettle type reactor, and the second stage is an adiabatic tubular reactor, which is then sent to the separation system after passing through the solution preheater.

A solution preheater is a heat exchanger, a device used to transfer heat from a hot fluid to a cold fluid to meet specified process requirements. It occupies an important position in chemical, petroleum, power, food and many other industrial productions, especially in chemical production, heat exchangers can be used as heaters, coolers, condensers, evaporators and reboilers, etc., and are widely used. In this method, a heat exchanger is used to control the second polymerization solution from the outlet of the tubular reactor, so that this stream can accurately reach the desired temperature of the separation system.

The separation system is a system including multi-step separation and recovery operations, usually including multi-stage flash evaporation, devolatilization, circulation, extrusion and other processes. When only two-stage flash evaporation is considered, the heating flow from the outlet of the solution preheater is divided The stock enters the first-stage flash tank for separation, and the reconstituted fraction of the still liquid is sent to the second-stage flash tank for further separation.

Referring to Fig. 2, the process flow of the present invention is introduced into the tank reactor as the first-stage reactor, because the whole tank temperature of the tank reactor is the same as the outlet temperature, and the inlet ethylene monomer, comonomer, The temperature of the mixed stream such as solvent and catalyst is used to control the temperature in the tank reactor, so that the optimal activity temperature of the catalyst can be reached in the tank reactor and the production efficiency can be improved; The type reactor has the characteristics of temperature distribution. It is not necessary to reduce the temperature of the inlet stream of the tubular reactor, that is, the outlet temperature of the tank reactor to an excessively low level. It only needs to mix the stream with fresh ethylene monomer, comonomer and solvent. After mixing, it is sent to the tubular reactor for reaction, so that some areas in the tubular reactor are in the catalyst high activity temperature zone, and the outlet temperature is as close as possible to the temperature of the separation system, so as to make full use of the reaction heat and take into account the reaction yield .

Please continue to refer to FIG. 2 , in which only a schematic diagram of the separation system is made, and the specific process of the separation system is not drawn, and it is considered that a reusable circulating stream has been obtained. According to this method, the subsequently separated ethylene monomer, comonomer, solvent mixed stream should be recycled to the tank reactor and the tubular reactor to continue the reaction, or only recycled to the tank reactor or the tubular reactor.

As mentioned earlier, Figure 1 illustrates that the catalyst activity first increases to the optimum activity temperature as the temperature increases, and then decreases as the temperature continues to increase. Referring to Fig. 3, the graph shows the relationship between reaction rate and time in the tubular reactor and the tank reactor under the same feed temperature, flow rate and composition.

As shown by the curve represented by the solid line in Fig. 3, the tubular reactor controls the inlet stream temperature, so that the reaction rate is at a high level at the initial reaction time t ₀ , and then continues the reaction along the length of the tube to reach the time t ₁ , During the period from t ₀ to t ₁ , the reaction rate was at a high level, and then as the reaction temperature continued to increase, the reaction rate began to decrease, but the outlet stream temperature met the set requirements.

The tank reactor, as shown by the straight line in the form of dot-line in Figure 3, maintains the same residence time as the tubular reactor. The temperature is constant and at high temperature, the reaction rate remains uniform, the catalyst activity is low, and the reaction rate is always at a low level, and the production effect is obviously not as good as that of the tubular reactor.

In order to make the purpose and advantages of the present invention clearer, the present invention will be further described below with reference to the embodiments; it should be understood that the specific embodiments described herein are only used to explain the present invention, but not to limit the present invention.

Example 1, carried out according to the described method, the first stage adopts an adiabatic kettle type reactor with a stirring part inside, the reaction pressure is 40bar, and the residence time is 10min, and the second stage adopts an adiabatic tubular reactor, the reaction pressure is 40bar, and the residence time is 10 minutes. The separation system adopts a two-stage flash evaporation method, the first stage is medium pressure flash evaporation, the flash tank is set at 16 bar, and the separation temperature is 200 ℃; the second stage is low pressure flash evaporation, the flash tank is set at 3 bar, and the separation temperature is 195°C. The feed composition selects ethylene monomer, the α-olefin comonomer selects octene, the solvent selects n-hexane, the catalyst selects CGC as the main catalyst, MAO as the cocatalyst, the catalyst optimum activity temperature is 140 ℃, the molecular weight regulator selects hydrogen, The deactivating agent is water. The feed composition of the first-stage reactor is as follows: 800kg/h of ethylene monomer feed, 450kg/h of octene feed, 4400kg/h of n-hexane feed, 0.05kg/h catalyst and 0.5kg/h cocatalyst are added, 0.05kg/h of hydrogen was introduced, and all the feed components were mixed and then heat-exchanged to -25°C and then fed to the kettle reactor. The outlet temperature of the kettle reactor was 140°C, and the temperature of the reactor was increased by 165°C. The composition of the feed mixed with the output of the first stage reactor and supplemented to the second stage reactor is as follows: ethylene monomer feed 220kg/h, octene feed 180kg/h, n-hexane feed 500kg/h. Notably, the ethylene monomer feed was passed through a controlled reflux, controlled by the outlet temperature of the second stage reactor. When the outlet temperature is stable above the set point, reducing the ethylene flow can reduce the outlet temperature of the second stage reactor, and vice versa. The feed component adopts one feed after all mixing, the feed temperature is 50 ° C, the temperature of the stream entering the tubular reactor after mixing with the first polymerization solution is 128 ° C, and the second polymerization solution at the outlet of the tubular reactor is 128 ° C. The temperature is 180 °C, and then heated by the solution preheater to reach the specified temperature of the first-stage flash tank of the separation system, that is, 200 °C. At this time, the polymer output of the first-stage reactor is 809kg, the polymer output of the second-stage reactor is 332kg, the output ratio of the first-stage reactor and the second-stage reactor is 2.4, the total polymer output is 1141kg, and the energy consumption of the heat exchanger is is 110kW.

Embodiment 2, carry out polymerization according to the method of embodiment 1, the difference is to increase the second-stage tubular reactor to increase the residence time of the tubular reactor, so that the temperature of the second polymerization solution at the outlet of the tubular reactor is 200 ° C, At this time, the polymer output of the first-stage reactor was 809 kg, the polymer output of the second-stage reactor was 465 kg, the total polymer output was 1274 kg, and the energy consumption of the solution preheater was 0 kW.

Example 3, the polymerization is carried out according to the method of Example 1, the difference is that 50% of the gas-phase stream flow rate obtained from the first-stage flash evaporation of the separation system is recycled back to the reaction system, so the fresh ethylene, 1 - Decrease in both octene and solvent. After the circulating stream is cooled, it is mixed with the fresh material in the liquid phase and sent to the tank reactor, and the second stage tubular reactor is enlarged as in Example 2 to increase the residence time of the tubular reactor, so that the The temperature of the second polymerization solution at the outlet of the reactor is 200°C, the polymer output of the first-stage reactor is 809kg, the polymer output of the second-stage reactor is 465kg, the total polymer output is 1274kg, and the solution preheater can Consumption is 0kW.

Comparative example 1, this comparative example and Example 1 use a similar method to carry out olefin polymerization, the difference is that there is no second stage reactor, and the polymerization is carried out in a tank reactor according to Example 1, and the pipe in Example 1 is added. The fresh materials of the reactor are added to the tank reactor, the obtained first polymerization solution stream is 163 ° C, and the temperature is directly heated to 200 ° C through the solution preheater. At this time, the polymer output of the first stage reactor is 1031kg, the total polymer output is 1031kg, and the solution preheater energy consumption is 220kW.

Comparative Example 2, which uses a similar method to carry out olefin polymerization with Example 1, the difference is that the second stage reactor adopts a tank reactor instead of a tubular reactor, and the second polymerization solution at the outlet of the tank reactor is at this time. When the temperature is 140°C, the polymer output of the second-stage reactor is 809 kg, the polymer output of the second-stage reactor is 229 kg, the total polymer output is 1108 kg, and the energy consumption of the solution preheater is 176 kW.

Comparative example 3, carry out polymerization according to the method of comparative example 2, the difference lies in adjusting the residence time of the second-stage tank reactor, so that the temperature of the second polymerization solution at the outlet of the tank-type reactor is 180 ° C, and the second stage reactor is at this moment. The output polymer output was 809kg, the second stage reactor polymer output was 323kg, the total polymer output was 1132kg, and the solution preheater energy consumption was 106kW.

Compared with Example 1 and Comparative Example 1, by adding a tubular reactor, the energy consumption of the solution preheater required before the polymer solution enters the separation system is greatly reduced, and the polymer output is increased at the same time.

Compared with Example 1 and Comparative Example 2, the effect that the second-stage reactor is a tubular reactor is better than that the second-stage reactor is a tank-type reactor, which is embodied as: larger polymer output and smaller solution Preheater energy consumption.

The comparison between Example 1 and Comparative Example 3 is at the expense of increasing the equipment cost of the tank reactor, so that the outlet temperature is the same as that of the tubular reactor, and the energy consumption and polymer output of the solution preheater are similar to those of Example 1.

Example 2 Compared with Example 1, the polymer output of the second-stage reactor is larger, and the polymer output is higher. The biggest feature is that by adding a larger tubular reactor as the second-stage reactor, the outlet is The temperature meets the requirements of the downstream separation system, and no additional solution preheater is required.

Compared with Example 2, the output of the polymer and the energy consumption of the solution preheater are the same in Example 3, and its biggest feature is that the gas phase stream part obtained from the first-stage flash evaporation of the separation system is recycled back to the reaction part, which reduces the consumption of raw materials. consume.

The above-mentioned embodiments only represent several embodiments of the present invention, and the descriptions thereof are specific and detailed, but should not be construed as a limitation on the scope of the patent of the present invention. For those of ordinary skill in the art, without departing from the concept of the present invention, several modifications and improvements can be made, which all belong to the protection scope of the present invention.

Claims (10)

1. A method for copolymerizing ethylene and at least one alpha-olefin of C3-C12 to obtain the copolymer of ethylene and alpha-olefin of C3-C12 is characterized in that the method adopts a series connection mode of two stages of reactors, wherein the first stage reactor is a kettle type reactor, the second stage reactor is a tubular reactor, and then the discharge of the tubular reactor enters a separation system through a solution preheater or is directly discharged to the separation system to obtain a polymer product and a circulating stream; the method comprises the following specific steps:

a) feeding a catalyst, a cocatalyst, ethylene, an alpha-olefin comonomer and a solvent into a kettle type reactor, reacting in the kettle type reactor, and discharging to obtain a first polymer solution;

b) feeding the first polymer solution into a tubular reactor for continuous reaction, and discharging to obtain a second polymer solution;

c) feeding the second polymer solution to a solution preheater to obtain a third polymer solution having a higher temperature, which is fed to a separation system; or directly feeding the second polymer solution to a separation system;

separating the polymer solution entering the separation system in the separation system to obtain a polymer product and a circulating stream containing ethylene, a solvent and an alpha-olefin comonomer;

the inlet of the tank reactor is provided with feeding pipes for feeding materials into the tank reactor, wherein the feeding pipes comprise a main catalyst feeding pipe, a cocatalyst feeding pipe, a solvent feeding pipe, an ethylene feeding pipe and a comonomer feeding pipe, and a circulating pipeline for receiving a circulating stream from a separation system is also arranged, and the feeding pipes or the circulating pipelines can be independently connected with the tank reactor or can be partially or completely combined into one feeding pipe to be connected with the tank reactor;

said tubular reactor is also provided with feed pipes for make-up of fresh ethylene monomer, solvent, alpha-olefin comonomer from C3 to C12, and with a recycle line for receiving a recycle stream from the separation system, these feed pipes or recycle lines being either individually connected to the tubular reactor or combined partly or totally into one feed pipe connected to the tubular reactor;

the discharging temperature of the tubular reactor is higher than that of the kettle type reactor.

2. The process of claim 1, wherein the tubular reactor is an adiabatic reactor having a discharge temperature at least 5 ℃ greater than the discharge temperature of the tank reactor.

3. The process of claim 1, wherein the tank reactor temperature is controlled by regulating the flow of the added ethylene, solvent, alpha-olefin comonomer; and/or, controlling the tubular reactor temperature by regulating the temperature of the added ethylene, solvent, alpha-olefin comonomer.

4. The process according to claim 1, wherein the mass ratio of the total amount of ethylene, solvent, alpha-olefin comonomer fed to the tubular reactor to the output of the first stage reactor is between 1/50 and 1/2, preferably between 1/10 and 1/5.

5. The process according to claim 1, wherein the pressure in the tubular reactor is between 30 and 200bar, preferably between 35 and 50 bar; the outlet temperature of the tubular reactor is 125-240 ℃, and preferably 140-180 ℃; the tubular reactor outlet temperature is controlled by the fresh ethylene flow entering the tubular reactor through a control loop.

6. The process according to claim 1, characterized in that the ratio of the tube reactor residence time to the tank reactor residence time is between 1/50 and 1/2, preferably between 1/20 and 1/5.

7. The process of claim 1, wherein the mass ratio of the kettle reactor polymer product to the tubular reactor polymer product is between 1/5 and 40, preferably between 1/2 and 20, more preferably between 2 and 10.

8. The method according to claim 1, wherein the solvent is selected from the group consisting of C4-C12 chain alkanes or cycloalkanes, C6-C9 aromatic hydrocarbons or their mixture solvents, preferably C5-C8 chain alkanes or cycloalkanes.

9. The method according to claim 1, wherein the tank reactor residence time is 2 to 60min, preferably 5 to 40min, more preferably 8 to 20 min; the pressure in the tank reactor is 30 to 200bar, preferably 35 to 50 bar.

10. The process according to claim 1, wherein the temperature in the tank reactor is between 120 and 240 ℃, preferably between 125 and 155 ℃; the temperature rise of the kettle type reactor is 80-220 ℃, namely the temperature of the reaction discharge is 80-220 ℃ higher than that of the reaction feed, and preferably 130-180 ℃.

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