CN114989340A - Olefin polymerization method - Google Patents
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- C08F—MACROMOLECULAR COMPOUNDS OBTAINED BY REACTIONS ONLY INVOLVING CARBON-TO-CARBON UNSATURATED BONDS
- C08F210/00—Copolymers of unsaturated aliphatic hydrocarbons having only one carbon-to-carbon double bond
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- B01J8/00—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes
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Abstract
The present invention relates to a process for the polymerization of olefins in which ethylene is copolymerized with at least one alpha-olefin having from C3 to C12. The method adopts a two-stage reactor series connection mode, and then the polymer product and the circulating stream are obtained after entering a separation system. The method comprises the following specific steps: (a) feeding the reaction materials and the catalyst into a kettle type reactor, and discharging after reaction to obtain a first polymer solution; (b) feeding the first polymer solution into a tubular reactor to continue the reaction, and feeding the obtained second polymer solution into a solution preheater to obtain a third polymer solution at an equal or higher temperature; (c) the third polymer solution is subjected to a separation operation in a separation system to obtain a polymer product and a recycle stream comprising ethylene, solvent, alpha-olefin comonomer. The invention can effectively utilize reaction heat, save the public engineering consumption, effectively improve the polymerization reaction yield and has the advantages of good economic benefit and high production efficiency.
Description
Technical Field
The invention relates to the field of olefin polymerization, and more particularly relates to a copolymerization method for producing ethylene and alpha-olefin, which comprises the steps of carrying out polymerization reaction on the ethylene and at least one alpha-olefin with the C3-C12 in a series combination of a first stage of a tank reactor and a second stage of a tubular reactor, and then feeding the ethylene and the at least one alpha-olefin into a separation system to obtain a product and part of recyclable components.
Background
The ethylene and alpha-olefin copolymer with high comonomer content is a high-performance thermoplastic elastomer with extremely wide application, and compared with common polyolefin plastics, the ethylene and alpha-olefin copolymer with high comonomer content has higher comonomer content and lower density in a molecular chain, and is widely applied to the fields of automobiles, packaging, wires and cables, polymer modification, sealing pieces, medical treatment and the like. Currently, there are gas phase, slurry and solution processes for the industrial production of ethylene and α -olefin copolymers. Compared with the former two methods, the ethylene and alpha-olefin polymer produced by solution polymerization has the advantages of lower product density and better elasticity. By "solution polymerization" is meant 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, the solution polymerization being a homogeneous liquid polymerization system. The conventional solution polymerization process is carried out at a temperature of 40 to 160 ℃ and a pressure of less than 13MPa and operates in a polymerization system in which more than 65% by weight of an inert solvent is present.
In the solution polymerization, ethylene monomer, comonomer, catalyst and cocatalyst are all dissolved in a solvent for reaction, the generated polymer product is also kept dissolved in the solvent, and a polymer solution is formed after the reaction. And after the reaction is finished, sending the outlet stream of the reactor into a separation system for separation and recovery to obtain a polymer product, and circulating available components such as ethylene monomer, solvent, comonomer and the like to the reaction part for continuously participating in the reaction. One of the main differences between olefin polymerization production processes is represented by the differences in reactor combinations, and the production efficiency and the polymer product performance characteristics corresponding to different reactor combinations are different. The ideal full mixed flow kettle type reactor has uniform reaction temperature and component concentration and is consistent with the outlet stream. While a plug flow tubular reactor temperature and component concentration exists along the reactor, as in the case of an insulated tubular reactor, the tubular reactor temperature increases gradually from the inlet to the outlet as the polymerization proceeds.
There are several reactor types and reactor combinations designed for the current processes associated with olefin production. CN111630071A discloses a method for copolymerization of ethylene and α -olefins in a solution polymerization reactor, wherein the polymerization reaction is first carried out in a tank reactor, then the effluent is discharged into a solution preheater, and after reaching a set temperature, the heated stream is fed into a separation system. CN 112500510A relates to a polymerization method for preparing polyethylene by a solution method, which adopts two kettle type reactors connected in series to carry out polymerization reaction and then enters a separation system; adding reactors is a well known method to increase production efficiency and regulate product. KR 101590998B discloses a continuous solution polymerization method of olefin by pre-polymerizing in a first stage reactor (tank reactor) and feeding into a second stage reactor (tubular reactor) connected in series with the tank reactor to perform main polymerization reaction, aiming at adding a static mixer composed of Kenics mixer and Sulzer mixer in the tubular reactor to enhance mixing effect and obtain polymer with more uniform property. The polymerization temperature of the kettle type reactor and the polymerization temperature of the tubular reactor are the same, and as in the first embodiment, the temperature of the first stage reactor and the second stage reactor are both 120 ℃, so that the reaction heat of the tubular reactor cannot be fully utilized to reduce the consumption of public works. Patent CN 107614541a describes a continuous solution polymerization process in which a polymer solution is obtained from a polymerization reactor and passed successively through two heat exchangers before being fed to a separator, and heated to a certain temperature, wherein the second heat exchanger uses utilities. Patent CN105377902A proposes a method for improving energy utilization of solution polymerization facilities, in which the polymer solution obtained from the reactor is passed through a heat exchanger to raise the temperature, then fed to a gas-liquid separator under reduced pressure, and part of the gas phase is directly recycled to reduce energy consumption. Patent CN1283204A proposes a process design for increasing the polymer content in an olefin solution polymerization process, also provided with a similar heat exchanger for increasing the temperature of the polymer stream entering the separator. More solvents are used in the olefin solution polymerization process, and a heater needs to be arranged at the outlet of the reactor in the current design before the separation operation is carried out on the polymer solution discharged from the reactor, so that the temperature is increased to strengthen the separation effect, and the energy consumption of the solution polymerization process is high.
Aiming at the problems that the heat public engineering is required to be consumed and the reaction yield is still low before the olefin polymer solution is subjected to solvent separation, a new process is required to be found, and the energy consumption of the public engineering of the process is reduced while the reaction yield is effectively improved.
Defining:
by "at least one C3 to C12 alpha olefin" is meant a comonomer, i.e. in addition to ethylene monomer, which may be selected from alpha olefins having 3 to 12 carbon atoms.
"alpha-olefin" means a monoolefin having a double bond at the end of the molecular chain, such as 1-butene, 1-hexene, 1-octene.
"polymers" are polymers, including terpolymers, of ethylene and one or more alpha-olefin comonomers from C3 to C12.
"continuous" refers to a system that operates without interruption. For example, reactants may be introduced into one or more reactors continuously and polymer product withdrawn continuously.
"solution preheater" means a heat exchanger disposed 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. The heat exchanger plays an important role in chemical industry, petroleum industry, power industry, food industry and other industrial production, especially in chemical industry, the heat exchanger can be used as a heater, a cooler, a condenser, an evaporator, a reboiler and the like, and the application is wide.
The "separation system" is a system comprising a plurality of separation and recovery operations, and is intended to separate a polymer obtained by polymerization from a polymer solution and obtain a monomer, a comonomer, a solvent component which can be recycled, and/or remove an impurity component. The separation system of the olefin polymer production process typically comprises multiple stages of flashing, devolatilization, recycle, extrusion, and the like. For example, when only two-stage flash is contemplated, the heated stream from the outlet of the solution preheater is fed to a first stage flash tank for separation and the bottoms heavies stream is fed to a second stage flash tank for continued separation.
"flash separation" means a separation step that results in phase separation by a reduction in pressure.
A "tank reactor" is a cylindrical reactor with a low aspect ratio, in which stirring (e.g., mechanical stirring) means are usually provided. When the height-diameter ratio is larger, a plurality of layers of stirring blades can be used. In the reaction process, materials need to be heated or cooled, a jacket can be arranged on the wall of the reactor, or a heat exchange surface is arranged in the reactor, or heat exchange can be carried out through external circulation.
The tubular reactor is a tubular continuous operation reactor with large length-diameter ratio, belonging to a plug flow reactor.
The "adiabatic reactor" refers to a reactor which does not exchange heat with the outside, and the material is adjusted to a specified temperature and then fed into the reactor, and the reaction temperature is only related to the feeding flow, the feeding temperature and the reaction heat and is not affected by the outside environment.
Disclosure of Invention
In order to overcome the defects of the prior art, the invention provides a method for copolymerizing ethylene and at least one alpha-olefin of C3-C12 to obtain the copolymer of the ethylene and the alpha-olefin of C3-C12, which adopts a series connection mode of two stages of reactors, wherein the first stage of reactor is a kettle type reactor, the second stage of reactor is a tubular reactor, and then the discharge of the tubular reactor enters a separation system through a solution preheater or directly discharges 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 reactor, reacting in the kettle 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 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, solvent, alpha-olefin comonomer.
The inlet of the tank reactor is provided with feed pipes for feeding materials into the tank reactor, the feed pipes comprise a main catalyst feed pipe, a cocatalyst feed pipe, a solvent feed pipe, an ethylene feed pipe and a comonomer feed pipe, and a circulating pipeline for receiving circulating streams from a separation system is also arranged, and the feed pipes or the circulating pipelines can be independently connected with the reactor or can be partially or completely combined into one feed pipe to be connected with the tank reactor;
the tubular reactor is also provided with feed pipes for make-up of fresh ethylene monomer, solvent, C3 to C12 alpha-olefin comonomer and with recycle lines for receiving recycle streams from the separation system, which feed pipes or recycle lines can be connected to the tubular reactor individually or can be combined partially or totally into one feed pipe connected to the tubular reactor;
it is characterized in that the discharge temperature of the tubular reactor is higher than that of the kettle reactor.
According to the process of the invention, the tubular reactor is an adiabatic reactor.
According to the process of the invention, the second polymerization solution discharge temperature of the tubular reactor is at least 5 ℃ higher than the discharge temperature of the first polymerization solution and the solution of the tank reactor.
According to the process of the present invention, the tubular reactor temperature and its temperature profile can be controlled by regulating the flow of the ethylene, solvent, alpha-olefin comonomer feed.
According to the process of the present invention, the temperature of the tubular reactor and its temperature profile is controlled by regulating the temperature of the ethylene, solvent, alpha-olefin comonomer feed.
According to the process of the present invention, the feed to the tubular reactor is derived from fresh or recycled ethylene monomer, solvent, alpha-olefin comonomer from C3 to C12, or a mixture of the above components, wherein the recycle stream is derived from a separation system downstream of the tubular reactor.
According to the method, the mass ratio of the total amount of the ethylene, the solvent and the alpha-olefin comonomer fed into the tubular reactor to the discharge amount of the first-stage reactor is 1/50-1/2, preferably 1/10-1/5.
According to the method, the mass ratio of the polymer product in the kettle reactor to the polymer product newly added in the tubular reactor is 1/5-40, preferably 1/2-20, and more preferably 2-10.
According to the process of the invention, the pressure in the tubular reactor is from 30 to 200bar, preferably from 35 to 50 bar.
According to the method, the outlet temperature of the tubular reactor is 125-240 ℃, and preferably 140-180 ℃.
According to the process of the invention, the tubular reactor outlet temperature is controlled by the fresh ethylene flow entering the tubular reactor via a control loop.
According to the inventive method, the ratio of the tubular reactor residence time to the tank reactor residence time is between 1/50 and 1/2, preferably between 1/20 and 1/5.
According to the process of the present invention, the alpha-olefin is a C3-C12 olefin, preferably a C4-C8 alpha-olefin.
According to the method of the invention, the solvent is selected from chain alkane or cycloalkane of C4-C12, aromatic hydrocarbon of C6-C9 or a mixed solvent thereof, and preferably C5-C8 chain alkane or cycloalkane.
According to the method, the residence time of the kettle type reactor is 2-60 min, preferably 5-40 min, and more preferably 8-20 min.
According to the method of the invention, the pressure in the tank reactor is 30-200 bar, preferably 35-50 bar.
According to the method, the temperature in the kettle reactor is 120-240 ℃, preferably 125-155 ℃.
According to the method, the temperature rise of the kettle type reactor is 80-220 ℃, namely the temperature of the reaction discharge material is 80-220 ℃ higher than that of the reaction feed material, and preferably 130-180 ℃.
The invention finds that the utility consumption of a solution preheater at the upstream of a separation system can be reduced by increasing the temperature of a polymer solution at the outlet of a reactor by the heat of polymerization while increasing the reaction yield by using the series combination of a tank reactor as the first stage and a tubular reactor as the second stage. It is known that there exists an optimum reaction temperature range for a polymerization catalyst, as shown in the catalyst activity-temperature diagram of fig. 1, the activity of the catalyst increases with the increase of the reaction temperature, and increases at a lower temperature until the activity of the catalyst starts to decrease with the increase of the temperature after the catalyst reaches the optimum activity at a certain temperature. Therefore, in order to improve the production efficiency, it is necessary to optimize the type of the reactor, the combination of the reactors, 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 at an optimum activity zone, which is higher in the space-time yield (the amount of reaction product obtained per residence time and per reaction volume) than that of the tubular reactor. However, only the kettle reactor is arranged, and the high reaction outlet temperature and the high reaction rate cannot be considered at the same time. The first-stage reactor adopts a kettle type reactor and sets the reaction temperature in a temperature range with higher activity. The second-stage reactor adopts a tubular reactor instead of a kettle reactor, because the tubular reactor has temperature distribution, the temperature is gradually increased along with the reaction, and the higher level of the activity of the catalyst in the tubular reactor can be kept and the discharging temperature of the reactor is higher by regulating and controlling the reaction temperature. It has further been found that the reaction temperature profile of the tubular reactor and the temperature of the polymer solution at the reaction outlet can be controlled by adjusting the flow and/or temperature of the feeds of ethylene, comonomer and solvent into the tubular reactor.
By combining the technical scheme, the scheme of the invention has the following advantages:
aiming at the characteristics of the copolymerization process of ethylene and alpha-olefin, the reaction yield is improved, and the public works are saved. Through the form of series combination of the first stage of the kettle type reactor and the second stage of the tubular reactor, the reaction stream of the kettle type reactor is mixed with the supplemented ethylene monomer, the alpha-olefin comonomer, the solvent or the mixture of the components, and the mixture enters the tubular reactor for continuous reaction, so that the reaction yield is as high as possible, and the reaction heat can be used for providing energy for a subsequent separation system.
Drawings
FIG. 1 is a graph of catalyst activity as a function of temperature;
FIG. 2 is a schematic flow diagram of the process of the present invention;
FIG. 3 is a graph of reaction rate as a function of residence time.
Detailed Description
The invention will be further illustrated and described with reference to specific embodiments. The described embodiments are merely exemplary of the disclosure and are not intended to limit the scope thereof. The technical features of the embodiments of the present invention can be combined correspondingly without mutual conflict.
As shown in FIG. 2, the invention provides a solution polymerization method for copolymerizing ethylene and alpha-olefin, which uses a polymerization device mainly comprising a tank reactor (1), a tubular reactor (2), a solution preheater (3) and a separation system (4), wherein the tank reactor is provided with a feeding pipe, ethylene monomer, solvent and alpha-olefin comonomer are all combined into one feeding pipe to be connected with the reactor, and the tank reactor is also provided with a circulating pipeline for receiving a circulating stream from the separation system; the outlet of the kettle type reactor is connected with the inlet of the tubular reactor, and the first polymerization solution obtained from the outlet of the kettle type reactor is mixed with fresh and/or recycled ethylene monomer, alpha-olefin comonomer and a mixed stream of solvent and then enters the tubular reactor for reaction; and the second polymer solution obtained from the outlet of the tubular reactor passes through a solution preheater to obtain a third polymer solution with higher temperature, and then the third polymer solution is conveyed to a separation system, and the separation system performs series of operations on the mixed material to separate, recover and circulate the material, so as to finally obtain a polymer product and a circulating stream.
The chemical reaction rate means how fast the chemical reaction proceeds. Generally expressed as a change (decrease or increase) in the concentration of a reactant or a product per unit time (e.g., per minute, per hour), the reaction rate is related to the nature and concentration of the reactant, the temperature, the pressure, the catalyst, etc., and, if the reaction is carried out in solution, the nature and amount of the solvent. The reaction rate can be controlled by controlling the reaction conditions to achieve certain objectives, such as improving the reaction yield, which is an important index for evaluating the chemical production process, i.e., the quality (or molar quantity) of the target product obtained per unit time.
During the polymerization reaction, the chemical reaction rate is proportional to the activity of the catalyst. The catalyst activity means the amount (or mass) of the reactant for converting the raw material per unit volume (or mass) of the catalyst in a unit time, and the higher the catalyst activity, the faster the reaction rate. The catalyst activity is mainly determined by the temperature, and the relationship between the catalyst activity and the temperature can be known as described above, and is shown in fig. 1.
The polymerization catalyst can be any catalyst known in the art capable of copolymerizing ethylene with an alpha-olefin comonomer, including Z-N catalysts in which TiCl is present 3 -A1Et 2 Cl catalyst system, in which the main catalyst component is amino titanium compound ((R) 1 R 2 N) 4-n TiYn, the cocatalyst component is aluminoxane compound, and the main catalyst component is transition metal compound MR I (OR’) m X n-(1+m) The cocatalyst component is an aluminoxane compound and a third component of an organic compound containing dihydroxy, etc.; also includes metallocene catalyst, the main catalyst is limited geometrical catalyst (CGC), the auxiliary catalyst is Methyl Aluminoxane (MAO) and so on.
As is known in the art, small amounts of chain transfer constant-large materials are often added to the polymerization system to reduce the polymer molecular weight, i.e., molecular weight regulators, may optionally be used in the solution polymerization reactor to control the molecular weight of the copolymer. Suitable chain transfer agents are many, such as hydrogen.
As is known in the art, the introduction of a separate stage after the end of the polymerization reaction requires the termination of the reaction, i.e., catalyst deactivation, to avoid excessive polymer molecular weight or difficult control of the exotherm of the reaction continuing in the separate stage, and may optionally be a catalyst deactivation material used in the solution polymerization reactor. Deactivating agents are numerous, such as water.
Solution polymerization requires large amounts of solvent for dissolving 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 stable during the reaction, a linear, cyclic or branched alkyl group having 6 carbon atoms or a mixture of two or more thereof, such as n-hexane, may be selected.
In the chemical reaction process, the types of the reactors are various and have respective characteristics. The kettle type reactor has the characteristics of uniform temperature of the whole kettle and same outlet temperature and reaction temperature, and can control the reaction temperature in the kettle by regulating and controlling the flow and temperature of inlet material flow so as to enable the catalyst to be at a high activity temperature; the tubular reactor has the characteristic of temperature distribution along the length of the tube, so that not only partial areas exist to ensure higher catalyst activity, but also polymerization reaction heat is fully utilized, and high outlet temperature and reaction yield can be considered.
According to the characteristics of each reactor, two stages of reactors are selected and connected in series, wherein the first stage is an adiabatic kettle type reactor, and the second stage is an adiabatic tube type reactor, and then the two stages are sent into a separation system after passing through a solution preheater.
A solution preheater is a heat exchanger, which is a device used to transfer heat from a hot fluid to a cold fluid to meet specified process requirements. The heat exchanger plays an important role in chemical industry, petroleum industry, power industry, food industry and other industrial production, particularly can be used as a heater, a cooler, a condenser, an evaporator, a reboiler and the like in chemical industry production, and is widely applied. In the process, a heat exchanger is used to control the second polymerization solution exiting from the tubular reactor so that the stream can reach exactly the temperature specified in the separation system.
The separation system is a system comprising a plurality of steps of separation and recovery operations, generally comprises processes of multi-stage flash evaporation, devolatilization, circulation, extrusion and the like, when only two-stage flash evaporation is considered, a heating stream from an outlet of a solution preheater enters a first-stage flash drum for separation, and a kettle heavy component stream is sent to a second-stage flash drum for continuous separation.
Referring to fig. 2, the process flow of the present invention is introduced into a tank reactor as a first-stage reactor, and since the full tank temperature of the tank reactor is the same as the outlet temperature, the temperature in the tank reactor can be controlled by controlling the temperature of the mixed streams of the inlet ethylene monomer, the comonomer, the solvent, the catalyst, etc., so that the optimal activity temperature of the catalyst in the tank reactor is reached, thereby improving the production efficiency; the second-stage reactor adopts a tubular reactor, and utilizes the characteristic that the tubular reactor has temperature distribution, the temperature of the stream at the inlet of the tubular reactor is not required to be reduced to a low level, namely the temperature at the outlet of the kettle-type reactor is only required to be mixed with the mixed stream of fresh ethylene monomer, comonomer and solvent and then sent into the tubular reactor for reaction, so that partial area in the tubular reactor is in a high-activity temperature area of the catalyst, and the outlet temperature is as close to the temperature of a separation system as possible, thereby fully utilizing the reaction heat and simultaneously considering the reaction yield.
With continued reference to fig. 2, fig. 2 is a schematic illustration of the separation system, and the specific process of the separation system is not shown and is considered to have achieved a recyclable recycle stream. According to the method, the ethylene monomer, comonomer and solvent mixed stream separated subsequently is recycled to the kettle reactor and the tubular reactor for continuous reaction, or only recycled to the kettle reactor or the tubular reactor.
As previously mentioned, fig. 1 illustrates that catalyst activity first increases with increasing temperature to an optimum activity temperature, and then decreases as temperature continues to increase. Referring to FIG. 3, a graph of reaction rate versus time for a tubular reactor versus a tank reactor is shown for the same feed temperature, flow rate, and composition.
As shown by the curve represented by the solid line in FIG. 3, the tubular reactor controlled the inlet stream temperature to make the reaction rate at the initial t of the reaction 0 At a higher level, and then continuously reacting along the length of the tube until t 1 At time t 0 -t 1 During the time, the reaction rates were all at a higher level, after which the reaction rates began to drop as the reaction temperature continued to rise, but the outlet stream temperature met the set requirements.
And the kettle type reactor is shown as a straight line represented by a dot-line form in figure 3, the residence time of the kettle type reactor is consistent with that of the tubular reactor, in order to enable the temperature of an outlet to be as close to the temperature required by the separation system as possible, the temperature in the reaction kettle is constant and high, the reaction rate is uniform, the activity of the catalyst is low, the reaction rate is always at a lower level, and the production effect is obviously inferior to that of the tubular reactor.
In order that the objects and advantages of the invention will be more clearly understood, the invention is further described below with reference to examples; it should be understood that the specific embodiments described herein are merely illustrative of the invention and are not intended to limit the invention.
Example 1, the procedure was as described, the first stage using an adiabatic tank reactor with a stirring element inside, a reaction pressure of 40bar and a residence time of 10min, the second stage using an adiabatic tube reactor with a reaction pressure of 40bar and a residence time of 2 min; the separation system adopts a two-stage flash mode, wherein the first stage is medium-pressure flash, the flash tank is provided with 16bar, and the separation temperature is 200 ℃; the second stage is low pressure flash evaporation, the flash tank is set at 3bar, and the separation temperature is 195 ℃. The ethylene monomer is selected as the feeding component, the octene is selected as the alpha-olefin comonomer, the normal hexane is selected as the solvent, the CGC is selected as the main catalyst, the MAO is selected as the auxiliary catalyst, the optimal activity temperature of the catalyst is 140 ℃, the hydrogen is selected as the molecular weight regulator, and the water is selected as the deactivator. The feed composition to the first stage reactor was as follows: feeding ethylene monomer 800kg/h, octene 450kg/h, n-hexane 4400kg/h, adding 0.05kg/h catalyst and 0.5kg/h cocatalyst, introducing hydrogen 0.05kg/h, mixing all the feeding components, exchanging heat until the temperature reaches-25 ℃, feeding one strand of the mixture into a kettle reactor, controlling the outlet temperature of the kettle reactor to be 140 ℃, and raising the temperature of the reactor to be 165 ℃. The feed composition mixed with the first stage reactor discharge and supplemented to the second stage reactor was as follows: ethylene monomer feed 220kg/h, octene feed 180kg/h, n-hexane feed 500 kg/h. It is worth noting that the ethylene monomer feed is controlled by the outlet temperature of the second stage reactor by a controlled reflux. When the outlet temperature is higher than the set value, the outlet temperature of the second-stage reactor can be reduced by reducing the flow of the ethylene, and vice versa. The feeding components are completely mixed and then fed, the feeding temperature is 50 ℃, the temperature of a stream entering the tubular reactor after being mixed with the first polymerization solution is 128 ℃, the temperature of a second polymerization solution at the outlet of the tubular reactor is 180 ℃, and the second polymerization solution is heated by a solution preheater to reach the specified temperature of a first-stage flash tank of the separation system, namely 200 ℃. At this time, the polymer yield of the first stage reactor was 809kg, the polymer yield of the second stage reactor was 332kg, the ratio of the first stage reactor yield to the second stage reactor yield was 2.4, the total polymer yield was 1141kg, and the energy consumption of the heat exchanger was 110 kW.
Example 2 polymerization was carried out as in example 1, except that the second stage tubular reactor was enlarged 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 was 200 ℃, at which time the polymer production in the first stage reactor was 809kg, the polymer production in the second stage reactor was 465kg, the total polymer production was 1274kg, and the solution preheater energy consumption was 0 kW.
Example 3 polymerization was carried out according to the method of example 1 except that 50% of the flow rate of the vapor phase stream obtained from the first stage flash of the separation system was recycled back to the reaction system, so that the fresh ethylene, 1-octene and solvent required for the feed were all reduced. After the circulating flow strand is cooled, the circulating flow strand is mixed with fresh materials in a liquid phase mode and then is sent into a kettle type reactor, a second stage tubular reactor is enlarged as in example 2 to increase the retention time of the tubular reactor, so that the temperature of a second polymerization solution at the outlet of the tubular reactor is 200 ℃, the polymer yield of a first stage reactor is 809kg, the polymer yield of a second stage reactor is 465kg, the total polymer yield is 1274kg, and the energy consumption of a solution preheater is 0 kW.
Comparative example 1, which was a similar process to that of example 1 except that no second stage reactor was provided, polymerization was carried out in a tank reactor as in example 1, and fresh materials added to the tubular reactor in example 1 were added together to the tank reactor, and the obtained first polymerization solution stream was 163 ℃, and was directly heated to 200 ℃ through a solution preheater, at which time the first stage reactor had a polymer production of 1031kg, a total polymer production of 1031kg, and a solution preheater had an energy consumption of 220 kW.
Comparative example 2, which was conducted to polymerize olefins using a similar method to that of example 1 except that the second stage reactor used a tank reactor instead of a tubular reactor, where the temperature of the second polymerization solution at the outlet of the tank reactor was 140 ℃, where the polymer production of the second stage reactor was 809kg, the polymer production of the second stage reactor was 229kg, the total polymer production was 1108kg, and the solution preheater energy consumption was 176 kW.
Comparative example 3 polymerization was carried out according to the method of comparative example 2, except that the residence time of the second stage tank reactor was adjusted so that the temperature of the second polymerization solution at the outlet of the tank reactor was 180 ℃, at which time the polymer production at the outlet of the second stage reactor was 809kg, the polymer production in the second stage reactor was 323kg, the total polymer production was 1132kg, and the solution preheater power consumption was 106 kW.
Example 1 in comparison to comparative example 1, the addition of a tubular reactor greatly reduced the energy consumption of the solution preheater required before the polymer solution entered the separation system, while increasing the polymer throughput.
Compared with the comparative example 2, the effect of the second-stage reactor being a tubular reactor is better than that of the second-stage reactor being a kettle reactor, which is specifically shown as follows: greater polymer production and less solution preheater energy consumption.
Example 1 is compared with comparative example 3, the solution preheater energy consumption and polymer production are closer than in example 1, at the expense of increased equipment cost for the tank reactor, making the outlet temperature the same as for the tubular reactor scheme.
Example 2 compared with example 1, the second stage reactor has higher polymer output and higher polymer output, and the biggest characteristic is that the outlet temperature meets the requirements of a downstream separation system by adding a larger tubular reactor as the second stage reactor, and no additional solution preheater is needed.
Example 3 compared with example 2, the polymer production and the solution preheater energy consumption are the same, and the biggest characteristic is that the gas phase stream obtained from the first-stage flash evaporation of the separation system is partially recycled to the reaction part, so that the consumption of raw materials is reduced.
The above-mentioned embodiments only express several embodiments of the present invention, and the description thereof is more specific and detailed, but not construed as limiting the scope of the present invention. It will be apparent to those skilled in the art that various changes and modifications can be made without departing from the spirit and scope of the 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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