CN112409513A

CN112409513A - Method for preparing ethylene polymer by microreactor and microreactor

Info

Publication number: CN112409513A
Application number: CN201910770279.4A
Authority: CN
Inventors: 李化毅; 刘卫卫; 李倩; 胡友良
Original assignee: Institute of Chemistry CAS
Current assignee: Institute of Chemistry CAS
Priority date: 2019-08-20
Filing date: 2019-08-20
Publication date: 2021-02-26
Anticipated expiration: 2039-08-20
Also published as: CN112409513B

Abstract

The invention belongs to the technical field of polyolefin, and discloses a method for preparing an ethylene polymer by high-pressure polymerization of a microreactor. The method comprises adding a reaction mixture into a microreactor through a split-flow channel; the reaction mixture passes through each reaction zone of the microreactor, and an initiator is injected to carry out polymerization reaction; and (4) the polymerized liquid enters a separation device, and ethylene monomers are recovered after separation to obtain a polymer. Also provides a microreactor device of the preparation method. The device can quickly transfer heat, control the uniformity of local reaction environment, flexibly realize parallel reaction, and ensure the maximum safety of production; the prepared polymerization product has excellent performance, and high-throughput quantification of polymerization is realized.

Description

Method for preparing ethylene polymer by microreactor and microreactor

Technical Field

The invention belongs to the technical field of polyolefin, and particularly relates to a method for preparing ethylene polymer in a microreactor and the microreactor.

Background

The high-pressure polyethylene (LDPE) has low density and low melting temperature, and is widely applied to the fields of films, pipes, wire and cable coating layers and the like. An ultra-high pressure reaction research group of ICI corporation in England in 1933 prepares white waxy polyethylene under high temperature and high pressure, the white waxy polyethylene is applied to high-frequency insulating materials, a 100t/a pilot plant is built in 1939, and the industrial production of polyethylene by a high-pressure method is realized. The autoclave process for polyethylene was introduced in 1942, and the tubular reactor process was developed in 1943 by Du Pont and UCC. To date, the processes and equipment of both methods have been greatly improved, and the industry still uses them to produce high-pressure polyethylene with almost equivalent production capacity, even the combined application of the two methods is appeared. Commercial LDPE plants typically consist of ethylene compression, initiator injection, polymerization, separation systems, and extrusion pelletization. In the high pressure polymerization process of ethylene, polyethylene is prepared by free radical polymerization in supercritical ethylene. The average residence time of the ethylene in the tank reactor is generally from 3 to 5min and the residence time in the tubular reactor is from 0.5min to several minutes. Compared with a kettle type reactor, the materials in the tubular reactor flow in a plunger shape in the tube, and the back mixing phenomenon does not occur; the polymerization reactor has simple structure, high design flexibility, convenient manufacture and maintenance and lower device construction cost than the kettle type method. Domestic LDPE devices coexist in kettle type and tubular type methods, and the kettle type production is mainly used.

As is well known, the reaction pressure of a high-pressure polyethylene device is up to 220-350MPa, and the reaction temperature is up to 270-330 ℃. Under such high temperature and high pressure conditions, there are great potential safety hazards in industrial production. And the inner diameter range of the pipe of the currently used tubular reactor is 20-100mm, and the pipe diameter is large, thus being not beneficial to heat transfer. Furthermore, the prior art also has difficulty in achieving the particular requirements of ethylene polymerization for a specific residence time of the fluid in the microreactor channel (polymerization time).

Disclosure of Invention

In view of the deficiencies of the prior art, the present invention provides a method for preparing ethylene polymers in a microreactor.

A method for preparing an ethylene polymer in a microreactor, said method comprising the steps of:

(1) adding the reaction mixture into each main reaction module of the microreactor in a split-flow manner; each main reaction module is in parallel connection;

(2) the reaction mixture passes through each reaction zone of the main reaction module, and an initiator is respectively injected into each reaction zone to carry out polymerization reaction;

(3) merging the crude products obtained by the polymerization reaction in each reaction zone, and separating to obtain unreacted monomers and ethylene polymers;

the reaction mixture includes a reactive monomer that is ethylene or a mixture of ethylene and other comonomers.

Preferably, chain transfer agents and/or other processing aids are also optionally included in the reaction mixture.

According to the invention, the ethylene polymer is an ethylene homopolymer or an ethylene copolymer.

According to the technical scheme of the invention, the other comonomer comprises at least one of alpha-olefin, diene, alpha, beta-unsaturated carboxylic acid, alpha, beta-unsaturated carboxylic ester and anhydride thereof. For example, the α -olefin is selected from at least one of propylene, butene and 1-hexene; the diene is selected from at least one of 1, 5-hexadiene, 1, 7-octadiene, 1, 9-decadiene and 1, 13-tetradecadiene; the alpha, beta-unsaturated carboxylic acid is selected from at least one of maleic acid, itaconic acid, crotonic acid, methacrylic acid and acrylic acid; the alpha, beta-unsaturated carboxylic acid ester and the anhydride thereof are selected from at least one of methyl methacrylate, butyl methacrylate, vinyl acetate, ethyl acrylate, ethylene glycol di (methacrylate), 1, 3-butylene glycol di (methacrylate), 1, 4-butylene glycol di (methacrylate), hexanediol di (methacrylate), dodecanediol di (methacrylate), trimethylolpropane tri (methacrylate), methacrylic anhydride, itaconic anhydride and maleic anhydride.

Further, the other comonomer content is 1-50%, preferably 3-40%, 4-30% of the mass of the reaction mixture, and exemplarily 5%, 8%, 10%, 15%, 20% of the mass of the reaction mixture.

According to the technical scheme of the invention, the chain transfer agent comprises but is not limited to at least one of aliphatic hydrocarbon, olefin, ketone, aldehyde and unsaturated aliphatic alcohol. For example, the unsaturated aliphatic hydrocarbon may be a saturated hydrocarbon having 6 or more carbon atoms, such as at least one of hexane, cyclohexane, octane, and the like. For example, the olefin may be at least one of propylene, pentane, and hexene. For example, the ketone may be at least one of acetone, diethyl ketone, and diamyl ketone. For example, the aldehyde may be formaldehyde and/or acetaldehyde. For example, the saturated aliphatic alcohol may be at least one of methanol, ethanol, propanol, and butanol. When the chain transfer agent is an olefin, it may be the same as or different from the comonomer and is selected independently from the olefins mentioned above.

Further, the chain transfer agent is used in an amount of 0.1 to 10%, preferably 0.5 to 5%, illustratively 0.8%, 1.0%, 1.5%, 2.0% by mass of the reaction mixture.

According to the technical scheme of the invention, the reaction mixture is gas or a mixture of gas and liquid. The flow dividing mode can be realized by a flow dividing flow channel. For example, the gas is compressed and then enters the feed split flow channel; optionally, the liquid is injected by an injection pump, and the liquid and the compressed gas are mixed and then enter the feed split flow channel.

Further, the system pressure of the reaction mixture after compression and before entering the flow dividing channel is 50-400MPa, preferably 100-300MPa, and exemplarily 150 MPa; the flow rate of the compressed reaction mixture fluid may be in the range of 1 to 30m/s, preferably 2 to 15 m/s.

According to the technical scheme of the invention, the number of the main reaction modules is at least 2, such as 3, 4, 5, 6 or more, and exemplarily, the number of the main reaction modules is 2. Further, the number of reaction zones within the main reaction module is at least 1, such as 2, 3, 4, 5, 6 or more; when the number of the reaction zones is more than or equal to 2, the reaction zones are in a series connection relationship; illustratively, each main reaction module contains 3 reaction zones in series. Further, the inlet temperature of the reaction mixture into the first reaction zone (i.e., into the first reaction zone) of the microreactor is 100-.

According to the solution of the present invention, in step (1), after splitting, the mass of the reaction mixture of each stream can be the same or different, thereby providing flexibility to fine-tune the volume of the reaction mixture entering each reaction zone.

According to the technical scheme of the invention, in the step (2), the initiator comprises at least one of oxygen and organic peroxide; wherein the organic peroxide comprises at least one of a monofunctional organic peroxide and a difunctional organic peroxide.

Further, the monofunctional organic peroxide comprises at least one of peroxyester, ketone peroxide, peroxyacetal and peroxycarbonate; illustratively, the monofunctional organic peroxide is selected from at least one of t-butyl peroxymaleate, t-butyl peroxybenzoate, t-butyl peroxyisononanoate, t-butyl peroxypivalate, t-butyl peroxyneodecanoate, t-amyl peroxypivalate, t-butyl peroxy-2-ethylhexanoate, cumyl peroxyneodecanoate, di-t-butyl peroxide, t-butyl hydroperoxide, cumene hydroperoxide, diisopropylbenzene hydroperoxide, diisononanoyl peroxide, didecanoyl peroxide, 2-bis (t-butylperoxy) butane, methyl isobutyl ketone hydroperoxide, dicyclohexyl peroxydicarbonate, diacetyl peroxydicarbonate, and di-2-ethylhexyl peroxydicarbonate; tert-butyl peroxypivalate is preferred.

Further, the bifunctional organic peroxide is selected from the group consisting of 1, 4-di-t-butylperoxycarbocyclohexane, 1-di-t-butylperoxycyclohexane, 1-di-t-amylperoxy cyclohexane, 2-di-t-butylperoxybutane, 2, 5-dimethyl-2, 5-di-t-butylperoxyhexane, 2, 2-di-4, 4-di-tert-butylperoxy cyclohexyl propane, 2, 5-dimethyl-2, 5-di-2-ethylhexanoylperoxyhexane, 2, 5-dimethyl-2, 5-tert-peroxyhexyne, n-butyl-4, 4-di-tert-butylperoxyvalerate, ethyl-3, 3-di-tert-butylperoxybutyrate and ethyl-3, 3-di-tert-amylperoxy butyrate.

Further, when the initiator comprises oxygen, the oxygen and the reaction mixture enter the microreactor after passing through a compressor; the other initiators are injected in the form of solutions, for example metered in by means of injection pumps at different initiator injection points in the microreactors. For example, the initiator is dissolved in a solvent selected from at least one of ketones and aliphatic hydrocarbons, preferably saturated C₈-C₂₅Hydrocarbons such as octane, decane, and isododecane, illustratively isododecane. Further, the initiator content in the initiator solution is 2 to 65% by mass, preferably 5 to 40% by mass, particularly preferably 8 to 30% by mass, and exemplarily 20% by mass.

Further, the initiator is injected in the polymerization system in an amount of 1 to 1000ppm, for example 5 to 500ppm, 8 to 100ppm, illustratively 10ppm, based on the fluid injection of the reaction mixture.

According to the technical scheme of the invention, in the step (2), the temperature of the polymerization reaction can be 150-350 ℃, and is preferably 250-330 ℃. For example, the peak temperature of each reaction zone is 200-; preferably, in at least one reaction zone, the peak temperature is 280-; preferably, the reaction mixture is cooled to at least 20 ℃, more preferably at least 40 ℃ and most preferably at least 50 ℃ below the peak temperature of the reaction zone in each reaction zone upstream of the initiator injection point before the reaction mixture reaches the next initiator injection point.

According to the technical scheme of the invention, the step (3) specifically comprises the following steps: the separation may include high pressure separation and low pressure separation. Further, the pressure of the high-pressure separation is 100-300bar, and the pressure of the low-pressure separation is 1-20 bar. Preferably, the high pressure separation is performed first, followed by the low pressure separation. Preferably, before separation, a cooling process of the crude product is also included, for example, the crude product is cooled to 150-280 ℃, for example to 170-260 ℃, exemplary to 180 ℃ before entering the high-pressure separation; the temperature is further reduced to 90-110 deg.C, for example to 100 deg.C, before the low-pressure separation by the high-pressure separation.

According to the technical scheme of the invention, the ethylene polymer obtained in the step (3) is an ethylene homopolymer or an ethylene copolymer. For example, the polymer has a density of 0.91 to 0.95g/cm³Molecular weight distribution M thereof_w/M_nIs 2 to 35; preferably, the polymer has a density of 0.92 to 0.93g/cm³Molecular weight distribution M thereof_w/M_nIs 5 to 15; illustratively, the polymer has a density of 0.918, 0.919, 0.920, 0.921, 0.923, 0.931g/cm³Molecular weight distribution M thereof_w/M_n5.1, 6.3, 6.8, 7.9, 8.5 and 12.3. For example, the number average molecular weight of the polymer is (100-270). times.10³E.g., (105-250). times.10³Illustratively, the number average molecular weight is 110 × 10³、143×10³、177×10³、188×10³、213×10³、230×10³。

Further, the present invention provides an ethylene polymer prepared by the above process. The ethylene polymer has the meaning as described above.

Further, the present invention provides the use of the ethylene polymer prepared by the above method in packaging, construction, industry or agriculture. For example, the ethylene polymers are applied individually, in admixture or by coextrusion.

Further, the invention also provides a microreactor for preparing ethylene polymer, which can comprise a feeding module, a main reaction module, an initiator injection module and a product separation module;

the feeding module is connected with the main reaction module, the main reaction module is connected with the product separation module, and the initiator injection module is connected with the main reaction module;

the feed module comprises a first splitter and the initiator injection module comprises a second splitter;

the number of the main reaction modules is at least two, and the main reaction modules are in parallel connection;

the main reaction module comprises at least one reaction area, and the reaction area comprises an independent reaction flow channel, an initiator injection point and a polymerization liquid discharge flow channel; the first flow divider is connected with each reaction zone through a flow dividing flow channel, and the initiator injection point is arranged at a material inlet of the reaction flow channel; and the second flow divider is connected with the initiator injection point through a flow dividing flow passage.

According to the embodiment of the invention, each reaction zone corresponds to the same or different temperature control systems; the temperature control system comprises a constant temperature insulation box and a cold and hot circulator component. Further, the main reaction module is placed in a constant temperature incubator.

According to an embodiment of the present invention, when the number of reaction zones in each main reaction module is 2 or more, the reaction zones are connected in series.

According to an embodiment of the present invention, the feeding module further comprises a compressor and a first injection pump, and the compressor and the first injection pump are respectively connected with the first flow divider. The first flow divider and the flow dividing channels are arranged to realize multi-point feeding and parallel connection of the main reaction modules. Specifically, the compressor is connected to the first flow divider, each flow dividing flow passage is independently connected to the first flow divider, and each flow dividing flow passage is provided with a metering pump.

According to an embodiment of the invention, the number of compressors of the feed module is selected from one or more than two; the compressor is preferably a multi-stage compressor, more preferably a multi-stage metal diaphragm compressor.

According to the embodiment of the invention, the number of the flow dividing channels for connecting the feeding module and the main reaction module is at least two, and each flow dividing channel is in a parallel relation. Preferably, the number of the flow dividing flow channels, the number of the reaction zones and the number of the initiator injection points are kept consistent.

Furthermore, a heating component is arranged on a connecting pipeline of the first shunting flow channel and the first reaction zone so as to achieve a proper inlet temperature.

According to an embodiment of the invention, the first reaction zone in the reaction module is defined as the following part of the microreactor: the reaction mixture delivered from the first divergent channel is brought into mixing contact with the initiator delivered from the first initiator injection point to the point of the second initiator injection point. Similarly, the second reaction zone, the third reaction zone, the Nth reaction zone are defined with reference to the first reaction zone.

Further, the initiator injection points are arranged on the reaction flow channel, and the number of the initiator injection points is at least 1, preferably 2 to 10, and more preferably 2 to 6.

Further, the side stream fed in from the first splitter is usually fed into the microreactor upstream of the initiator injection point (preferably cooled to, for example, 10 to 20 ℃) in order to lower the temperature of the reaction mixture.

According to an embodiment of the present invention, the number of the feeding module, the main reaction module, the initiator injection module, the separation module may be the same or different and is selected from one or more than two independently from each other.

According to the embodiment of the invention, a feed inlet and a discharge outlet are arranged on the first flow divider; the number of feed openings may be 1, preferably 2: a gas feed port and a liquid feed port; the number of the discharge ports may be 1 or more, for example, 2 or more.

According to the embodiment of the invention, each discharge hole of the first flow divider is connected with a corresponding flow dividing flow passage; preferably, a fastener is arranged between the discharge port and the flow dividing channel and used for connecting the discharge port of the flow divider with each flow dividing channel.

Wherein, a check valve and/or a stop valve can be arranged on the flow dividing channel.

Wherein the main reaction module comprises a multi-way valve, and the multi-way valve is arranged at the downstream in the reaction flow channel and in front of the next initiator injection point. The multi-way valve can realize multi-stage discharging.

Further, the number of the multi-way valves may be one, two or more, preferably two or more;

furthermore, the multi-way valve can be a three-way valve, two ends of the three-way valve are connected with the reaction flow channel, and the other end of the three-way valve is connected with other three-way valves through pipelines.

According to an embodiment of the invention, the initiator injection module further comprises a second syringe pump connected with the second flow splitter. Preferably, the number of the diversion flow channels connecting the initiator injection module and the main reaction module is at least two, such as 2-10, preferably 2-6; preferably, each split flow channel corresponds to an outlet of the initiator splitter.

Further, the syringe pump is a pulseless metering syringe pump.

Further, cooling within the reaction zone is achieved by a temperature control system or a combination of a temperature control system and the introduction of an ethylene sidestream.

According to an embodiment of the invention, the product separation module comprises a pressure controller, a product cooler, and a separation device; the pressure controller is used for controlling the pressure in the reaction flow channel, and preferably selects a high-pressure discharge valve; one end of the pressure controller is connected with the multi-way valve of the main reaction module, and the other end of the pressure controller is connected with the product separation device through the product cooler.

Further, the product cooler comprises a conduit having a cooling jacket; the product cooler has an internal diameter of at least 20mm and a length of at least 50 m.

The separation device is at least provided with a section of separation device; preferably a two stage separation device. Wherein the first stage separation device is a high pressure separator that effects a first separation of polymer from unreacted ethylene; the separated gas is fed to a compressor through a high-pressure circulating system, and the molten polymer at the bottom of the high-pressure separator is decompressed to a second-stage separation device. Wherein the second stage separation device is a low pressure separator and substantially all of the residual ethylene is separated from the polymer.

According to the embodiment of the invention, the cross sections of the reaction flow channel and each connecting pipeline are circular, and the pipe walls are uniform; preferably, the tube length of the single main reaction module may be 30 to 3000m, and the inner diameter may be 0.5 to 15 mm; for example, the tube length of a single main reaction module may be 35 to 500m, 40 to 100m, and the inner diameter may be 0.5 to 15mm, 1 to 10 mm; illustratively, the tube length of a single main reaction module may be 40m, 60m, and the inner diameter may be 2mm, 5 mm.

The invention has the beneficial effects that:

the micro-reactor is a device which controls the flow, transfer and reaction processes on the scale of tens of microns to thousands of microns and has extremely high mixing, heat transfer and mass transfer efficiency. The heat transfer/mass transfer coefficient in the microreactor is 1-3 orders of magnitude larger than that of the traditional chemical equipment, and the microreactor is particularly suitable for quick reaction, high exothermic reaction and the like, such as polymerization reaction. The method provided by the invention utilizes the high-efficiency heat transfer efficiency and high-precision reaction control characteristic of the microreactor to carry out ethylene high-pressure polymerization, overcomes the defects of uneven heat transfer, unsafe production process and the like of the traditional tubular high-pressure ethylene polymerization, and has the following advantages:

1. the flow divider device in the feeding module of the microreactor has ingenious structure, realizes the parallel connection of the main reaction modules, solves the problem of the residence time of the polymerization reaction in the prior art, and has more thorough polymerization reaction.

2. In the invention, a plurality of polymers with different molecular weights and distributions thereof are prepared simultaneously by adopting the parallel main reaction modules with independent temperature control, thereby realizing high-throughput quantification of polymerization.

3. The multi-stage discharging module of the microreactor realizes flexible control of polymerization reaction time by simply opening and closing corresponding valves, and is superior to a mode of manually connecting pipelines to adjust the length of a pipe in the prior art; meanwhile, the multi-stage discharging module is favorable for checking the blockage problem in the pipeline.

4. The polymerization device of the microreactor has small volume and is more flexible and safer to operate; the method is suitable for continuous production, has small amplification effect, and can form the industrial-grade reactor by copying the reaction channels at the laboratory level through the components and enabling the reaction channels to be parallel to each other.

Drawings

FIG. 1 is a schematic structural view of a microreactor means provided in example 1 of the present invention;

reference numerals: 1 is a feeding module; 11 is a compressor; 12 is a first syringe pump; 13 is a one-way valve; 14 is a first splitter; 15 is a stop valve; 2 is a main reaction module; 21a is a first reaction flow channel; 21b is a second reaction flow channel; 21c is a third reaction flow channel; 22 is a three-way valve; 23 is a multi-way valve; 3 is an initiator injection module; 31 is a second syringe pump; 32 is a second splitter, 32a is a first initiator injection port; 32b is a second initiator injection port; 32c is a third initiator injection port; 4 is a product separation module; 41 is a pressure controller; 42 is a product cooler; 43 is a high pressure separator; and 44 is a low pressure separator.

Detailed Description

The present invention will be described in further detail with reference to specific examples. It is to be understood that the following examples are only illustrative and explanatory of the present invention and should not be construed as limiting the scope of the present invention. All the technologies realized based on the above-mentioned contents of the present invention are covered in the protection scope of the present invention.

The polymer density was measured as follows: the density test specimens were first prepared according to ASTM D1928 and then tested within 1 hour of pressing the specimens according to ASTM D792.

The polymer melt index (MFI) was tested according to ASTM D1238. The amount of polymer extruded in units of g at 190 ℃ under a load of 2.16Kg over 10 min.

The molecular weight and molecular weight distribution of the polymerization product were measured using a high temperature Gel Permeation Chromatograph (GPC) PL-GPC-220. 8-12mg of polymer are added to 10ml of 1,2, 4-trichlorobenzene and dissolved completely in a shaker at 160 ℃. At 160 ℃ and a steady flow rate of 1ml/min, 200. mu.L of the sample solution was injected and measured using polystyrene as a standard.

Example 1

The detailed description of fig. 1 is concerned. Two groups of main reaction modules which are connected in parallel are arranged in the microreactor device shown in the figure 1, the two groups of main reaction modules share the same initiator injection module 3, each group of main reaction modules consists of three reaction zones, and each group of main reaction modules is respectively connected with an independent product separation module. Taking the main reaction module 2 as an example, the connected product separation module is 4.

Fresh ethylene monomer and recycled ethylene streams (from the high pressure recycle system, not labeled in figure 1) enter compressor 11 in feed module 1. The compressed stream discharged from the compressor 11 flows into the first splitter 14 via a conduit having a one-way valve 13. Optionally, also flowing into the first splitter 14 is a stream of optional liquid comonomer, wherein liquid comonomer and/or chain transfer agent and/or other processing aid is supplied by a separate first injection pump 12. The material flow from the first splitter 14 passes through a conduit having a check valve 13 and a stop valve 15, passes through a preheating part, and then enters the reaction flow channel of the main reaction module 2, and the reaction flow channel includes

reaction flow channels

21a, 21b, and 21c, which are respectively disposed in the first reaction zone, the second reaction zone, and the third reaction zone. The shut-off valves 15 are selectively opened to allow for series and multi-point feeding of the reaction zones in the main reaction module 2.

Accordingly, 3 streams are split from first splitter 14 into main reaction module 2, stream 14a into

reaction channels

21a, 14b and 14c and ethylene side streams into

reaction channels

21b and 21c, respectively. Three initiator injection ports are also shown in each reaction channel: a first initiator injection port 32a, a second initiator injection port 32b, and a third initiator injection port 32 c. The initiator injection ports are arranged at intervals along each reaction flow channel, and are fed from the initiator injection module 3. The initiator injection module 3 includes a second injection pump 31 and a second flow divider 32 to inject the initiator solution. The first initiator injection port 32a is located just downstream of the front end of the reaction flow channel 21a and defines the beginning of the first reaction zone. The initiator entering through the first initiator injection port 32a is fluidly coupled to the hot reaction mixture from stream 14a and begins to polymerize as it proceeds down the reaction flow path 21a, causing the polymerization system to increase in temperature due to the exothermic heat of polymerization. As the temperature rises, the initiator decomposition and polymerization rate increases, accelerating heat generation and causing the temperature to rise further. A cooling device is provided at the reaction flow channel 21a to cool the reaction system. When the initiator is consumed, the initiation and the polymerization slow down, and the heat generated therein is equal to the heat conducted away from the reaction mixture, the temperature peaks (between 200 and 350 ℃) and then begins to drop. Downstream of this reaction zone, a three-way valve 22 is provided, from which the reaction can optionally be terminated by means of an outflow pressure controller 41, or the polymerization can optionally be continued along the reaction flow path.

As the reaction continues along the reaction flow path, the feed of side stream 14b may be adjusted to further cool the reaction system. The second initiator injection port 32b is just downstream of the entry point of the side stream 14b and determines the start of the second reaction zone. Again, the reaction system temperature rises, peaks and falls as it flows along the main reaction flow path 21b and the entry of the side stream 14c provides further rapid cooling, after which initiator is injected at the third initiator inlet 32c, which determines the start of the third reaction zone.

Downstream of the third reaction zone, the reaction flow path terminates in a pressure controller 41. A pressure controller 41 controls the pressure in the microreactor, downstream of which a product cooler 42 is connected. Once in the product cooler 42, the reaction mixture undergoes phase separation. And then into high pressure separator 43. The overhead gas from the high pressure separator 43 flows into the high pressure recycle system where the unreacted ethylene is cooled and returned to the compressor 11. The polymer product flows from the bottom of the high pressure separator 43 into the low pressure separator 44, separating almost all of the remaining ethylene from the polymer. The residual ethylene is recycled to the compressor 11 via a purification unit (not shown). The molten polymer exits the bottom of the low pressure separator 44 into an extruder (not shown) for extrusion, cooling, and pelletizing.

The microreactor is equipped with one or more sets of heating/cooling means (not shown) to adjust the temperature of the reaction system.

Fig. 1 is an exemplary diagram showing the positional relationship of the components, and those skilled in the art can change the number of the components as needed, and those skilled in the art can open or close the valve as needed to realize the parallel connection and multi-stage discharge of the components.

Example 2

Polyethylene was prepared using the microreactor provided in example 1:

(1) compressing ethylene monomer to 150MPa by a compressor 11, introducing the ethylene monomer into a flow dividing flow channel through a flow divider, preheating the ethylene monomer and introducing the preheated ethylene monomer into a micro-reaction flow channel;

(2) the length of the tube of the single main reaction module of the used micro-reaction flow channel is 40m, and the inner diameter is 2 mm; dissolving initiator tert-butyl peroxypivalate in isododecane (the mass concentration of the initiator is 20%) and feeding the isododecane into the micro-reaction flow channel through a metering pump through a second flow divider, wherein the initiator is injected at a speed of 10ppm relative to the fluid; the inlet temperature of the first reaction zone is 130 ℃; the total residence time of the polymerization reaction is 15 s;

(3) the polymerization liquid enters a high-pressure separator through a product cooler with the inner diameter of 30mm and the length of 60m, the temperature is reduced to about 180 ℃, the pressure is about 250bar, and unreacted gas is fed into a compressor again through a high-pressure circulating system; the other materials enter a low-pressure separator, the temperature is reduced to about 100 ℃, the pressure is about 10bar, most of unreacted ethylene is separated, and the ethylene enters a compressor again; the polymer is discharged from the bottom of the low-pressure separator.

In this embodiment, the shunt channels (14b and 14c) connected to the main reaction channel are closed, and the corresponding initiator injection channel is also closed, and at this time, the feeding in the main reaction channel is not performed, that is, the main reaction channel has only one reaction region.

Example 3

The microreactor means used in this example were the same as in example 2. Except that the amount of ethylene monomer entering the first reaction zone via 14a was 30 wt%; 14b and 14c side streams are opened, and ethylene monomer enters the second reaction zone and the third reaction zone respectively by 35 wt% through the two flow channels; the corresponding initiator injection flow channel was opened and the initiator tert-butyl peroxypivalate dissolved in isododecane (20% by mass of initiator) was injected at a rate of 10ppm with respect to the flow of the fluids in the

flow channels

32a, 32b and 32c controlled by a second injection pump.

Example 4

The same as in example 1, except that methyl methacrylate was used as a comonomer. Methyl methacrylate (5 wt% of total monomer) was injected into the splitter by the first syringe pump 13 of the feed module and mixed with the compressed ethylene and fed into the micro-reaction flow channels.

Example 5

The same procedure as in example 1 was repeated, except that the inlet temperature of the first reaction zone was 125 ℃.

Example 6

The same procedure as in example 1 was repeated, except that the inner diameter of the micro reaction channel was set to 5 mm.

Example 7

In this example, except that the parallel mode of 4 main reaction modules was adopted, the microreactor means and the polymerization conditions were the same as those in example 2. At this time, the polymerization residence time was 30 s.

The properties of the ethylene polymers obtained in examples 2 to 7 are shown in Table 1.

Table 1.

Examples	2	3	4	5	6	7
							MFI(g/10min)	6.8	6.1	9.3	8.7	4.3	4.7
Density (g/cm)³)	0.920	0.923	0.919	0.921	0.918	0.931
							M_n(×10³)	177	213	110	143	188	230
M_w/M_n	5.1	8.5	6.8	7.9	12.3	6.3

The embodiments of the present invention have been described above. However, the present invention is not limited to the above embodiment. Any modification, equivalent replacement, or improvement made within the spirit and principle of the present invention should be included in the protection scope of the present invention.

Claims

1. A method for preparing an ethylene polymer in a microreactor, said method comprising the steps of:

2. The process of claim 1, wherein a chain transfer agent and/or other processing aids are optionally also included in the reaction mixture;

preferably, the ethylene polymer is an ethylene homopolymer or an ethylene copolymer;

preferably, the other comonomer comprises at least one of an alpha-olefin, a diene, an alpha, beta-unsaturated carboxylic acid ester, and anhydrides thereof; the content of the other comonomers is 1-50% of the mass of the reaction mixture;

preferably, the chain transfer agent includes, but is not limited to, at least one of aliphatic hydrocarbons, olefins, ketones, aldehydes, and unsaturated aliphatic alcohols; the dosage of the chain transfer agent is 0.1-10% of the mass of the reaction mixture.

3. The method according to claim 1 or 2, wherein the reaction mixture is a gas or a mixture of a gas and a liquid;

preferably, the flow division is realized by a flow dividing flow channel, the system pressure of the reaction mixture before entering the flow dividing flow channel is 50-400MPa, and the flow velocity of the fluid is 1-30 m/s;

preferably, the number of the main reaction modules is at least 2, and the number of the reaction zones in the main reaction modules is at least 1; when the number of the reaction zones is more than or equal to 2, the reaction zones are in a series connection relationship;

preferably, the inlet temperature of the reaction mixture into the first reaction zone of the microreactor is 100-135 ℃;

preferably, in step (1), after splitting, the mass of the reaction mixture of each stream is the same or different.

4. The method of any one of claims 1-3, wherein in step (2), the initiator comprises at least one of oxygen and an organic peroxide; wherein the organic peroxide comprises at least one of a monofunctional organic peroxide and a difunctional organic peroxide;

preferably, the monofunctional organic peroxide comprises at least one of peroxyester, peroxyketone, peroxyacetal, and peroxycarbonate;

preferably, the bifunctional organic peroxide is selected from the group consisting of 1, 4-di-tert-butylperoxycarbocyclohexane, 1-di-tert-butylperoxycyclohexane, 1-di-tert-amylperoxy cyclohexane, 2-di-tert-butylperoxybutane, 2, 5-dimethyl-2, 5-di-tert-butylperoxyhexane, at least one of 2, 2-bis-4, 4-di-tert-butylperoxycyclohexylpropane, 2, 5-dimethyl-2, 5-di-2-ethylhexanoylperoxyhexane, 2, 5-dimethyl-2, 5-tert-peroxyhexyne, n-butyl-4, 4-di-tert-butylperoxyvalerate, ethyl-3, 3-di-tert-butylperoxybutyrate, and ethyl-3, 3-di-tert-amylperoxybutyrate;

preferably, the injection amount of the initiator in the polymerization reaction system is 1-1000ppm of the injection amount of the reaction mixture fluid;

preferably, in the step (2), the temperature of the polymerization reaction is 150-350 ℃; preferably, the reaction mixture is cooled to at least 20 ℃ below the peak temperature of the reaction zone in each reaction zone upstream of the initiator injection point before the reaction mixture reaches the next initiator injection point.

5. The method according to any one of claims 1 to 4, wherein step (3) comprises in particular: the separation comprises high-pressure separation and low-pressure separation, the pressure of the high-pressure separation is 100-300bar, and the pressure of the low-pressure separation is 1-20 bar;

preferably, high pressure separation is performed first, followed by low pressure separation;

preferably, before the separation, a cooling process of the crude product is also included;

preferably, the polymer obtained in step (3) is an ethylene homopolymer or a polyethylene copolymer;

preferably, the polymer has a density of 0.91 to 0.95g/cm³Molecular weight distribution M thereof_w/M_nIs 2-35.

6. An ethylene polymer produced by the process of any one of claims 1 to 5.

7. Use of the ethylene polymer prepared by the process of any one of claims 1 to 5 in packaging, construction, industrial or agricultural applications.

8. Microreactor for use in a method according to any of claims 1 to 5, wherein said microreactor comprises a feed module, a main reaction module, an initiator injection module and a product separation module;

the main reaction module comprises at least one reaction zone; each reaction zone comprises an independent reaction flow channel, an initiator injection point and a polymerization liquid discharge flow channel; the first flow divider is connected with each reaction zone through a flow dividing flow channel, and the initiator injection point is arranged at a material inlet of the reaction flow channel; and the second flow divider is connected with the initiator injection point through a flow dividing flow passage.

9. The microreactor of claim 8, wherein each reaction zone corresponds to the same or different temperature control system; the temperature control system comprises a constant temperature insulation box and a cold and hot circulator component;

preferably, the main reaction module is placed in a constant temperature incubator;

preferably, when the number of reaction zones in each main reaction module is greater than or equal to 2, the reaction zones are connected in series.

10. The microreactor according to claim 8 or 9, wherein the feeding module further comprises a compressor, a first injection pump, the compressor being connected to the first flow divider, and each of the split flow channels being connected to the first flow divider independently of each other;

preferably, the initiator injection module further comprises a second injection pump connected with the second flow divider;

preferably, the product separation module comprises a pressure controller, a product cooler, and a separation device; the pressure controller is used for controlling the pressure in the reaction flow channel, one end of the pressure controller is connected with the multi-way valve of the main reaction module, and the other end of the pressure controller is connected with the product separation device through the product cooler;

preferably, the tube length of the single main reaction module may be 30 to 3000m, and the inner diameter may be 0.5 to 15 mm;

preferably, the product cooler has an internal diameter of at least 20mm and a length of at least 50 m.