WO2024002393A1 - 烯烃自由基聚合的方法与烯烃自由基聚合装置 - Google Patents

烯烃自由基聚合的方法与烯烃自由基聚合装置 Download PDF

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Publication number
WO2024002393A1
WO2024002393A1 PCT/CN2023/115693 CN2023115693W WO2024002393A1 WO 2024002393 A1 WO2024002393 A1 WO 2024002393A1 CN 2023115693 W CN2023115693 W CN 2023115693W WO 2024002393 A1 WO2024002393 A1 WO 2024002393A1
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Prior art keywords
stage high
unit
pressure polymerization
pressure
tubular reactor
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PCT/CN2023/115693
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English (en)
French (fr)
Inventor
王靖岱
林华杰
范小强
史绽春
杨遥
刘国强
任聪静
田保政
阳永荣
柳兆坤
Original Assignee
中国石油化工股份有限公司
浙江大学
中石化宁波新材料研究院有限公司
中国石油化工股份有限公司镇海炼化分公司
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Priority claimed from CN202210775378.3A external-priority patent/CN117362494A/zh
Priority claimed from CN202210774665.2A external-priority patent/CN117358152A/zh
Application filed by 中国石油化工股份有限公司, 浙江大学, 中石化宁波新材料研究院有限公司, 中国石油化工股份有限公司镇海炼化分公司 filed Critical 中国石油化工股份有限公司
Publication of WO2024002393A1 publication Critical patent/WO2024002393A1/zh

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    • C—CHEMISTRY; METALLURGY
    • C08—ORGANIC MACROMOLECULAR COMPOUNDS; THEIR PREPARATION OR CHEMICAL WORKING-UP; COMPOSITIONS BASED THEREON
    • C08F—MACROMOLECULAR COMPOUNDS OBTAINED BY REACTIONS ONLY INVOLVING CARBON-TO-CARBON UNSATURATED BONDS
    • C08F110/00—Homopolymers of unsaturated aliphatic hydrocarbons having only one carbon-to-carbon double bond
    • C08F110/02—Ethene
    • C—CHEMISTRY; METALLURGY
    • C08—ORGANIC MACROMOLECULAR COMPOUNDS; THEIR PREPARATION OR CHEMICAL WORKING-UP; COMPOSITIONS BASED THEREON
    • C08F—MACROMOLECULAR COMPOUNDS OBTAINED BY REACTIONS ONLY INVOLVING CARBON-TO-CARBON UNSATURATED BONDS
    • C08F2/00—Processes of polymerisation
    • C08F2/01—Processes of polymerisation characterised by special features of the polymerisation apparatus used

Definitions

  • the invention relates to the field of high-pressure polymerization of olefins, and in particular to an olefin free radical polymerization method and an olefin free radical polymerization device.
  • Low-density polyethylene (LDPE) is produced through high-pressure free radical polymerization. Since tubular reactors are easier to scale up during the polymerization process and are more economical, tubular technology gradually dominates.
  • the existing high-pressure tubular process compresses the reaction materials to above 200MPa, passes them into a preheater and heats them to 170°C, and then enters the reactor to initiate the reaction.
  • the outlet material of the tubular reactor is separated by a high-pressure separator and a low-pressure separator, in which ethylene, telogen, and some oligomers enter the high-circulation loop and the low-circulation loop, while LDPE with a small amount of ethylene dissolved enters the extruder for granulation.
  • LDPE polymers produced in high-pressure tubular reactors usually have narrow product molecular weight distribution (MWD) and low long chain branching (LCB), while different downstream products have different effects on the molecular weight distribution (MWD) and relatively low length of polyethylene.
  • MFD molecular weight distribution
  • LCB long chain branching
  • the requirements for high long chain branching (LCB) are different.
  • medical grade/food grade LDPE resin requires a narrow molecular weight distribution (MWD), while the production of heavy-duty packaging bags, floor heating pipes and other products with excellent chemical properties requires a wide molecular weight distribution (MWD). Therefore, The production of products with different molecular chain structures on one device will have better economic benefits.
  • the existing method in this field to adjust the molecular weight distribution (MWD) and long chain branching (LCB) of LDPE polymer products is to change the feeding position of the telogen, including the entrance of the secondary machine, the stage of the secondary machine, and the secondary machine. Secondary machine outlet, preheater, reactor, reactor side feed point upstream, etc.
  • injecting telogen into the compression system can lead to premature polymerization and fouling in the compression system, resulting in a decrease in production load. Injecting the telogen into the reactor or reactor side feed will cause the telogen to mix with the initiator, reducing the initiator efficiency, and the mixing behavior of the telogen additional stream and the mainstream may create cold spots and reduce heat transfer.
  • the purpose of the present invention is to overcome the inability to realize the production of products with different molecular chain structures on the same device in the existing high-pressure olefin polymerization process, and the injection of telogen into the reactor or the side feed of the reactor will cause the telogen and initiator to Mixing reduces the initiator initiation efficiency, and the mixing behavior of the telogen additional flow and the mainstream may produce cold spots and reduce heat transfer defects.
  • the molecular weight distribution of polyethylene produced by the existing high-pressure tubular method of high-pressure polymerization is narrow and long.
  • the branch content is low, and the same device cannot produce thin film polyethylene products with narrow molecular weight distribution and low long chain branch content and coating polyethylene products with higher branching degree and wider molecular weight.
  • the invention provides an olefin radical polymerization method and an olefin radical polymerization device.
  • a first aspect of the present invention provides a method for olefin free radical polymerization.
  • the method includes: introducing at least two reaction monomer streams containing olefin sources into at least two parallel tubular reactors, and performing a first-stage high-pressure reactor respectively. polymerization, and then flow the obtained first-stage high-pressure polymerization product into one or more tubular reactors connected in series to perform multi-stage high-pressure polymerization; wherein at least one free radical polymerization initiator is introduced to participate in the first-stage high-pressure polymerization and/or Or multi-stage high-pressure polymerization, the pressure of the reaction monomer flow is greater than 100MPa.
  • the method includes: introducing a reaction monomer stream containing an ethylene source into at least two parallel tubular reactors to perform the reaction in the presence of an initiator; introducing at least one of the at least two parallel tubular reactors into Part of the material from the outlet of the tubular reactor is recycled back to at least one of at least two parallel tubular reactors for reaction; the remaining materials from the outlet of at least two parallel tubular reactors are recycled Continue to introduce one or more serially connected tubular reactors to perform the reaction in the presence of an initiator.
  • a second aspect of the present invention provides a device for the method of olefin free radical polymerization of the present invention, the device comprising:
  • One-stage high-pressure polymerization unit and multi-stage high-pressure polymerization unit among which,
  • the one-stage high-pressure polymerization unit is connected in series upstream of the multi-stage high-pressure polymerization unit;
  • the first-stage high-pressure polymerization unit includes at least two parallel tubular reactors for conducting one-stage high-pressure polymerization on at least two reaction monomer streams containing olefin sources;
  • the multi-stage high-pressure polymerization unit includes one or more tubular reactors connected in series for multi-stage high-pressure polymerization of the product from the first-stage high-pressure polymerization unit;
  • At least one tubular reactor in the one-stage high-pressure polymerization unit and/or the multi-stage high-pressure polymerization unit is provided with an initiator feed port.
  • the device further includes: a fluid suction and delivery unit, wherein the fluid suction and delivery unit includes one or at least two fluid suction and delivery devices arranged in parallel for sucking and delivering at least one reaction unit containing an ethylene source. Bulk flow and part of the material from the outlet of at least one tubular reactor in the first-stage high-pressure polymerization unit;
  • the initiator supply unit is used to deliver initiators to the one-stage high-pressure polymerization unit and the multi-stage high-pressure polymerization unit.
  • the present invention at least has the following beneficial effects:
  • the present invention proposes that at least two reaction monomer streams containing olefin sources are introduced into at least two parallel tubular reactors, and each performs one-stage high-pressure polymerization.
  • the parallel arrangement can better control the tubes.
  • the feed temperature, pressure and other parameters of the reactor can be adjusted to achieve product control while ensuring the conversion rate;
  • the present invention proposes that at least two reaction monomer streams containing olefin sources are introduced into at least two parallel tubular reactors, and each performs one-stage high-pressure polymerization, which can be achieved without changing the reaction section of the tubular reactor.
  • the concentration distribution of the adjusting agent such as the initiator in the tubular reactor can be better controlled, thereby realizing the adjustment of the molecular chain structure such as the number average molecular weight and molecular weight distribution (MWD) of the product, and the method of the present invention can be used to produce Obtain downstream products that match different fields;
  • the method proposed by the present invention is not only suitable for olefin homopolymerization initiated by free radical polymerization initiators, but also suitable for copolymerization of olefins and other olefinic monomers, thereby producing a variety of olefin homopolymerization and copolymerization products, improving the use
  • the method of the invention has good device utilization rate and applicability, and has good economic benefits.
  • the method of the present invention can improve the utilization efficiency of olefin raw materials, increase the conversion rate and increase the output;
  • the molecular weight distribution width and long-chain branching degree of polyethylene can be significantly increased, and a narrower molecular weight distribution and lower degree of polyethylene can be produced using the method of the present invention.
  • the molecular weight distribution range refers to the width of the molecular weight distribution range of the ethylene product that can be produced by the method of the present invention
  • the method of the present invention can also prepare polyethylene products with a wide molecular weight distribution. This is specifically reflected in the low-density polyethylene prepared by the method of the present invention.
  • the ratio of the molecular weight distribution to the conversion rate of the ethylene source is equal to or greater than 0.018 and equal to or less than 0.048.
  • Figure 1 is an olefin radical polymerization device according to a preferred embodiment of the present invention.
  • Figure 2 is an olefin radical polymerization device according to another preferred embodiment of the present invention.
  • Figure 3 is a reaction flow diagram of an ethylene free radical polymerization method according to some embodiments of the present invention.
  • Figure 4 is a reaction flow diagram of an ethylene free radical polymerization method according to other embodiments of the present invention.
  • downstream refers to the flow direction of materials in the device.
  • the first aspect of the present invention provides a method for olefin free radical polymerization.
  • the method includes: introducing a reaction monomer stream containing an olefin source into at least two parallel tubular reactors, each performing one-stage high-pressure polymerization, and then The obtained first-stage high-pressure polymerization product flows into one or more serially connected tubular reactors for multi-stage high-pressure polymerization; wherein at least one free radical polymerization initiator is introduced to participate in the first-stage high-pressure polymerization and/or multi-stage high-pressure polymerization.
  • the pressure of the reaction monomer stream containing the olefin source is greater than or equal to 100 MPa.
  • polyolefin products with wider molecular distribution and higher polymer dispersion index (PDI) can be produced.
  • the inventor speculates that the first-stage high-pressure polymerization unit of at least two parallel tubular reactors The setting can better control the reaction time at high and low temperatures during polymerization, thereby increasing the polymer dispersion index (PDI); at the same time, it can better control the feed port of the tubular reactor when using the device of the present invention. Temperature, pressure and other parameters, and then by setting the feed of the free radical polymerization initiator, the product can be controlled without increasing the fouling of the device while ensuring the conversion rate of the reaction monomer stream containing the olefin source.
  • the number of strands of the material containing the olefin source there is no limit to the number of strands of the material containing the olefin source. Compression is generally performed using a compression unit.
  • the number of strands of the material containing the olefin source corresponds to the number of compression units.
  • the number of strands of the material containing the olefin source is less than Equal to the number of strands of the reactive monomer stream.
  • the number of strands of the material containing the olefin source is less than the number of strands of the reactive monomer stream, it can be compressed by the compression unit and then divided into the required number of strands of the reactive monomer stream containing the olefin source. . For example, after a stream of material containing an olefin source is compressed to 100 MPa or more through a compression unit, it is divided into two streams of reaction monomer streams containing an olefin source.
  • the high-pressure polymerization conditions in the first-stage high-pressure polymerization and the second-stage high-pressure polymerization are that the reactive monomer stream can be polymerized under high pressure.
  • the pressure of the reactive monomer stream containing the olefin source is 110-400 MPa. (For example, 110MPa, 130MPa, 150MPa, 170MPa, 200MPa, 250MPa, 300MPa, 330MPa, 350MPa, and any value within the range of any of the above values); further preferably, it is 170-330MPa. It should be understood that the pressure of each of the reaction monomer streams containing the olefin source may be the same or different.
  • the pressure of the reactive monomer stream containing the olefin source is the inlet pressure of the reactive monomer stream containing the olefin source entering the first-stage high-pressure polymerization unit.
  • One-stage high-pressure polymerization is carried out under pressure.
  • both the one-stage high-pressure polymerization and the multi-stage high-pressure polymerization are carried out in a tubular reactor.
  • a pressure drop in the length direction of the tubular reactor In the present invention, it is called one-stage high-pressure polymerization.
  • the pressure drop before and after and the pressure drop before and after multi-stage high-pressure polymerization Preferably, the sum of the pressure drop before and after one-stage high-pressure polymerization and the pressure drop before and after multi-stage high-pressure polymerization: the pressure drop before and after one-stage high-pressure polymerization is 3:1- 30:1, preferably 6:1-8:1.
  • the deviation of materials can be reduced.
  • biasing flow refers to the deviation between the ratio of material flow rates in different parallel tubular reactors and the ratio of material flow rates calculated according to Bernoulli's equation to avoid the defect of excessive local temperature and ensure the conversion rate. It can realize the adjustment of molecular chain structure such as product molecular weight distribution (MWD) and long chain branching (LCB).
  • MWD product molecular weight distribution
  • LCB long chain branching
  • each one-stage high-pressure polymerization can be performed at the same time or not at the same time, as long as the first-stage high-pressure polymerization product flows into one or more pipes connected in series.
  • each stage of high pressure polymerization is performed simultaneously.
  • the molecular chain structure such as the molecular weight distribution (MWD) and long chain branching (LCB) of the product can be adjusted while ensuring the conversion rate.
  • the temperature of each of the reaction monomer streams containing the olefin source is 100-200°C (for example, 100°C, 120°C, 150°C, 170°C, 200°C, and any of the above values) any value within the range), preferably 150-200°C, and the sum of the reaction monomer streams containing the olefin source at the inlet of each parallel tubular reactor each satisfies the correlation expression: 10000 ⁇ 1 / ⁇ 1 ⁇ 1500, preferably 6000 ⁇ ⁇ 1 / ⁇ 1 ⁇ 3000; the unit of density ⁇ 1 is: kg/m 3 , and the unit of viscosity ⁇ 1 is: centipoise (cP).
  • Viscosity is measured at 25°C.
  • the reaction monomer stream containing the olefin source be heated to a temperature that can initiate polymerization, but also the product molecular weight and product molecular weight distribution can be better realized by controlling conditions such as preheating ( Adjustment of molecular chain structure such as MWD) and long chain branching (LCB).
  • preheating Adjustment of molecular chain structure such as MWD
  • LCB long chain branching
  • the temperature of the one-stage high-pressure polymerization and each multi-stage high-pressure polymerization can be selected as needed.
  • the respective temperatures of each one-stage high-pressure polymerization and each multi-stage high-pressure polymerization are 100-350°C (such as 100°C, 120°C, 125°C, 135°C, 150°C, 164°C, 170°C, 176°C, 180°C, 190°C, 192°C, 203°C, 211°C, 224°C, 225 °C, 295°C, 300°C, 320°C, 350°C, and any value within the range of any of the above values).
  • it is possible to control the molecular structure of the product such as molecular weight distribution and branch chain distribution while ensuring the conversion rate.
  • free radical polymerization is the main method.
  • the reaction temperature changes during the one-stage high-pressure polymerization and the multi-stage high-pressure polymerization, but the temperature changes are all within the range of 100-350°C.
  • the addition of free radical polymerization initiator will affect the temperature of polymerization.
  • the temperature of the materials in the reactor where the free radical polymerization initiator is injected through the initiator feed port is recorded as the "inlet temperature" ;Also record the peak temperature in the tubular reactor where the free radical polymerization initiator is introduced.
  • the feed amount of each olefin source-containing reaction monomer stream is not limited and can be selected according to needs.
  • the maximum feed amount of each olefin source-containing reaction monomer stream is The ratio of the feed amount to the minimum feed amount is (20-1):1, such as 20:1, 15:1, 10:1, 5:1, 3:1, 1:1, and any of the above values Any value within the range; preferably (5-1):1. Ratios are by weight.
  • different tubular reactors of one-stage high-pressure polymerization can be used to produce polymers with different molecular structural characteristics, thereby regulating the molecular structure of the final product. At the same time, it can reduce the difficulty of equipment design of tubular reactors for one-stage high-pressure polymerization.
  • the feed amount of each olefin source-containing reaction monomer stream refers to the feed amount of each olefin source-containing reaction monomer flow flowing into the tubular reactor in the first-stage high-pressure polymerization unit. .
  • the at least two olefin source-containing reaction monomer flows entering at least two parallel tubular reactors have a certain flow rate.
  • the flow rate of each of the olefin source-containing reaction monomer flows is greater than or equal to 5 m.
  • /s and less than or equal to 30m/s, such as 5m/s, 6m/s, 7m/s, 7.24m/s, 8m/s, 10m/s, 11m/s, 12m/s, 13m/s, 14m /s, 15m/s, 16m/s, 17m/s, 18m/s, 19m/s, 20m/s, 21m/s, 22m/s, 23m/s, 24m/s, 25m/s, 26m/s , 27m/s, 28m/s, 29m/s, 30m/s, and any value within the range of any of the above values, preferably greater than or equal to 8m/s and less than or equal to 20m/s.
  • Adopting the aforementioned embodiments can reduce the problem of polymers in parallel tubular reactors adhering to the inner walls of the reaction tubes, ensuring the safety of the reaction tubes, thereby improving the heat transfer efficiency and production efficiency of the tubular reactors. It can control the molecular structure of the product such as molecular weight distribution and branch chain distribution while ensuring the conversion rate.
  • the number of the first-stage high-pressure polymerization is not limited. In some preferred embodiments, the number of the first-stage high-pressure polymerization is 2-4. indivual. Under a certain flow rate of the olefin source-containing reaction monomer flow, the greater the number of first-stage high-pressure polymerizations, the smaller the inner diameter of the reactor that needs to be performed for the first-stage high-pressure polymerization, which imposes stricter requirements on equipment.
  • the aforementioned preferred embodiments can not only achieve molecular weight distribution (MWD) and long chain branch (LCB) content of the product, etc. The adjustment of the chain structure and the requirements for the reaction equipment are not so stringent.
  • At least one free radical initiator is introduced to participate in one-stage high-pressure polymerization; at least one initiator is introduced to participate in multi-stage high-pressure polymerization.
  • the free radical polymerization initiator is introduced in an intermittent or continuous injection manner to participate in one-stage high-pressure polymerization and/or multi-stage high-pressure polymerization.
  • the feed amount of the radical polymerization initiator in each strand can be selected as needed, and there is no particular restriction in the present invention.
  • the molecular weight of the prepared product can be changed by adding a telogen.
  • the method further includes feeding at least one telogen to participate in the first-stage high-pressure polymerization and multi-stage polymerization. high-pressure polymerization.
  • MWD product molecular weight distribution
  • LCB long chain branching
  • the feed amount of the telogen in each strand can be selected as needed, and there is no particular restriction in the present invention.
  • the olefin copolymer can be prepared by adding comonomers.
  • the method further includes feeding at least one comonomer to participate in the one-stage high-pressure polymerization and the multi-stage high-pressure polymerization. polymerization.
  • the method of the present invention is not only suitable for olefin homopolymerization initiated by free radical polymerization initiators, but also suitable for the copolymerization of olefins and comonomers, thereby producing a variety of olefin homopolymerization and copolymerization products, improving device utilization and applicability, and having relatively high efficiency. good economic effect.
  • the materials obtained by the multi-stage high-pressure polymerization are cooled under reduced pressure, and then unreacted monomers and polymer products are separated.
  • the olefin can be a monoolefin or a diolefin with a carbon number of 1 to 6. Specific examples include ethylene, propylene, butylene, isobutylene, 1,3-butadiene, pentadiene, and isoprene. One or more types of alkenes.
  • the product prepared by the method of the present invention is linear low density polyethylene.
  • the type of the comonomer can be selected according to needs. It can be understood that the type of the comonomer is different from that of the olefin source, and the comonomers that can be free-radically copolymerized with the olefin source under high pressure are all the same. Applicable to the system of the present invention.
  • examples of the comonomers are ⁇ , ⁇ -unsaturated C 3 -C 8 carboxylic acids, particularly acrylic acid, methacrylic acid, maleic acid, and fumaric acid ; and/or derivatives of ⁇ , ⁇ -unsaturated C 3 -C 8 carboxylic acid, for example, ⁇ , ⁇ -unsaturated C 3 -C 5 carboxylic acid ester or ⁇ , ⁇ -unsaturated C 3 -C 5 carboxylic acid anhydride , especially methyl methacrylate, n-butyl methacrylate, tert-butyl methacrylate, methyl acrylate, ethyl acrylate, n-butyl acrylate, tert-butyl acrylate, methacrylic anhydride and maleic anhydride; and/or 1-olefins, for example, propylene, 1-butene, 1-pentene, 1-hexene
  • the ratio of olefin monomers and comonomers is not limited and can be specifically selected according to actual needs.
  • the type of the free radical polymerization initiator is not limited. Any substance that can generate free radicals in one-stage high-pressure polymerization and/or multi-stage high-pressure polymerization can be used as the free radical polymerization initiator in the present invention.
  • the free radical polymerization initiator includes one or more of oxygen, air, azo compounds, organic peroxides, and hydrocarbons of CC initiators.
  • organic peroxides include peroxyesters, peroxyketals, peroxyketones and peroxycarbonates, such as di(2-ethylhexyl)peroxydicarbonate, dicyclohexyl peroxydicarbonate, Diacetyl peroxydicarbonate, peroxyisopropyl tert-butyl carbonate, di-tert-butyl peroxide, di-tert-amyl peroxide, Dicumyl peroxide, 2,5-dimethyl-2,5-di-tert-butylperoxyhexane, tert-butylcumyl peroxide, 2,5-dimethyl-2,5-di (tert-butylperoxy)hex-3-yne, 1,3-diisopropyl monohydroperoxide or tert-butyl hydroperoxide, didecanoyl peroxide, 2,5-dimethyl-2 ,5-di(2-ethylhexan
  • the radical polymerization initiator can be introduced in any state, such as liquid, dissolved state, and supercritical state.
  • a gaseous radical polymerization initiator such as oxygen or air
  • the gaseous radical polymerization initiator is introduced in a supercritical state.
  • the free radical polymerization initiator is a dissolved free radical polymerization initiator ; More preferably, the concentration of the free radical polymerization initiator in the dissolved free radical polymerization initiator is 5-80 wt%.
  • dissolved free radical polymerization initiator refers to a mixture of a solvent capable of dissolving a free radical polymerization initiator and a corresponding free radical polymerization initiator.
  • the type of solvent is not limited and can dissolve the corresponding free radical.
  • All polymerization initiator solvents are suitable for use in the system of the present invention. Examples of suitable solvents include ketones, aliphatic hydrocarbons (such as octane, decane, isododecane, etc.) and other saturated C 8 -C 25 hydrocarbons. Adopting the aforementioned preferred embodiment not only avoids the phenomenon of pyrolysis caused by overheating of the free radical polymerization initiator, making the reaction safer, but also improves the efficiency of the initiator and reduces the cost of using the initiator.
  • the telomerization agent includes one or more of aliphatic hydrocarbons, olefins, ketones, aldehydes, aliphatic alcohols, or hydrogen.
  • Aliphatic hydrocarbons that can be listed include propane, butane, pentane, hexane, cyclohexane, etc.; alkenes that can be listed include propylene, 1-pentene or 1-hexene; ketones that can be listed include acetone, methylethyl Ketone (2-butanone), methyl isobutyl ketone, methyl isopentyl ketone, diethyl ketone, dipentyl ketone, etc.; aldehydes that can be listed include formaldehyde, acetaldehyde or propionaldehyde; lipids that can be listed Family alcohols include methanol, ethanol, propanol, isopropyl alcohol, butanol, etc.
  • the telomerization agent is one or more of aliphatic aldehydes (such as propionaldehyde), 1-olefins (such as propylene or 1-hexene) and aliphatic hydrocarbons (such as propane).
  • aliphatic aldehydes such as propionaldehyde
  • 1-olefins such as propylene or 1-hexene
  • aliphatic hydrocarbons such as propane
  • the pressures involved are all absolute pressures.
  • the second aspect of the present invention provides a device for the method of olefin free radical polymerization of the present invention.
  • the device includes:
  • One-stage high-pressure polymerization unit and multi-stage high-pressure polymerization unit among which,
  • the one-stage high-pressure polymerization unit is connected in series upstream of the multi-stage high-pressure polymerization unit;
  • the first-stage high-pressure polymerization unit includes at least two parallel tubular reactors for conducting one-stage high-pressure polymerization on at least two reaction monomer streams containing olefin sources;
  • the multi-stage high-pressure polymerization unit includes one or more tubular reactors connected in series for multi-stage high-pressure polymerization of the product from the first-stage high-pressure polymerization unit;
  • At least one tubular reactor in the one-stage high-pressure polymerization unit and/or the multi-stage high-pressure polymerization unit is provided with an initiator feed port.
  • the device of the present invention through the arrangement of a one-stage high-pressure polymerization unit including at least two parallel tubular reactors and a multi-stage high-pressure polymerization unit including one or more tubular reactors connected in series, it is in line with the current situation.
  • a one-stage high-pressure polymerization unit including at least two parallel tubular reactors and a multi-stage high-pressure polymerization unit including one or more tubular reactors connected in series
  • PDI polymer dispersion index
  • the tube By adjusting the temperature, pressure and other parameters of the feed port of the reactor, and then setting the initiator feed ports at different positions, the product can be realized without increasing the fouling of the device while ensuring the conversion rate of the reaction monomer stream containing the olefin source. control.
  • the position of the initiator feed port is not limited and can be selected according to needs.
  • the reaction monomers of at least one tubular reactor in the first-stage high-pressure polymerization unit An initiator feed port is provided at the flow feed port end; in some embodiments, at least one tubular reactor in the first-stage high-pressure polymerization unit is provided with at least one initiator feed port along its length direction (for example, 1, 2, 3, 4, etc.), preferably 1-3 (referring to 1-3 respectively provided along the length direction of any one or more tubular reactors in the first-stage high-pressure polymerization unit) Initiator feed port); in some embodiments, at least one tubular reactor in the multi-stage high-pressure polymerization unit is provided with at least one initiator feed port (for example, 1, 2, 3, 4 Among them, “on the tubular reactor” includes the material inlet end of the tubular reactor and any position along the length of the tubular reactor.
  • “Feeding port” refers to the inlet end that receives the material of the first-stage high-pressure polymerization unit and the inlet end that receives the material flowing out of the adjacent tubular reactor), preferably 1-5 (referring to any one of the multi-stage high-pressure polymerization units) Or 2-5 initiator feed ports respectively provided on multiple tubular reactors).
  • the device of the present invention can better adjust the molecular chain structure of the product such as molecular weight distribution (MWD) and long chain branching (LCB), and can use the device of the present invention to produce downstream products that match different fields. product.
  • MWD molecular weight distribution
  • LCB long chain branching
  • reaction monomer flow feed port end refers to the reaction monomer flow feed port at one end of the tubular reactor in the first-stage high-pressure polymerization unit.
  • tubular reaction The other end of the reactor is the outlet; the inlet end of the tubular reactor in the multi-stage high-pressure polymerization unit refers to the end of the product flowing into the first-stage high-pressure polymerization unit in the first tubular reactor in the multi-stage high-pressure polymerization unit.
  • the first-stage high-pressure polymerization unit includes 2-4 parallel tubular reactors.
  • the molecular chain structure of the product such as molecular weight distribution (MWD) and long chain branching (LCB), can be adjusted.
  • the device in order to ensure that each reaction monomer stream containing an olefin source has an inlet pressure entering the first-stage high-pressure polymerization unit, in some embodiments, the device further includes a device located in the first-stage high-pressure polymerization unit. and at least one compression unit upstream of the high-pressure polymerization unit.
  • the device in order to ensure that each reaction monomer stream containing an olefin source has an inlet temperature entering the first-stage high-pressure polymerization unit, in some embodiments, the device further includes a device located in the first-stage high-pressure polymerization unit. At least one preheater upstream of the first-stage high-pressure polymerization unit; preferably, the preheater is located between the compression unit and the first-stage high-pressure polymerization unit. Using the aforementioned embodiment, one-stage high-pressure polymerization can occur more smoothly.
  • each reaction monomer stream containing an olefin source has an inlet pressure to enter the first-stage high-pressure polymerization unit.
  • the compression unit It includes one (or more serially connected, for example, 2 serially connected, 3 serially connected, 4 serially connected, 5 serially connected) compressors.
  • the compression unit includes 2-4 Compressors connected in series.
  • the compression unit includes two compressors connected in series, namely a first-stage compressor 2 and a second-stage compressor 3; in other embodiments, the compression unit includes three compressors connected in series.
  • the compressors are, in order, one-stage compressor 2, two-stage compressor 3 and three-stage compressor. Using the aforementioned embodiment, each reaction monomer flow can reach the inlet pressure of the first-stage high-pressure polymerization unit.
  • At least one compression unit is arranged in series upstream of at least two parallel tubular reactors in the first-stage high-pressure polymerization unit; in some embodiments, At least one compression unit is arranged in series at the corresponding upstream of the tubular reactors in the first-stage high-pressure polymerization unit.
  • each reaction monomer stream containing the olefin source has an inlet temperature for entering the first-stage high-pressure polymerization unit.
  • at least one preheater is arranged in series upstream of at least two parallel tubular reactors in the first-stage high-pressure polymerization unit; in some embodiments, at least one preheater is arranged in series in a Each of the tubular reactors in the high-pressure polymerization unit has its own corresponding upstream.
  • the device further includes at least one telogen feed inlet (for example, 1, 2, 3, 4, 5, 6, etc.).
  • the telogen feed port can be used to introduce the telogen into the device to participate in one-stage high-pressure polymerization and/or multi-stage high-pressure polymerization. Polymerization can better adjust the molecular weight of the product.
  • the number and position of the telogen feed ports are not limited and can be selected according to needs.
  • the telogen feed ports are provided in multiple locations.
  • the polymerization agent inlet can be set at any position upstream of the outlet of the multi-stage high-pressure polymerization unit, that is, according to the needs of the product, the polymerization agent inlet can be set at any position upstream of the outlet of the last tubular reactor in the multi-stage high-pressure polymerization unit.
  • each of the telogen feed ports can be arranged in various ways.
  • each of the telogen feed inlets is respectively arranged at: the inlet of the compression unit; and/or the outlet of the compression unit; and/or the connecting pipe of any two adjacent compressors in the compression unit; and/or, the reaction monomer flow inlet side of at least one tubular reactor in the first-stage high-pressure polymerization unit (referring to the connection between at least one tubular reactor in the first-stage high-pressure polymerization unit and its corresponding upstream compression unit at the pipeline); and/or, on at least one tubular reactor in the first-stage high-pressure polymerization unit (including the reaction monomer flow feed port end of the tubular reactor and any position along the length of the tubular reactor ); and/or, at the connecting pipe between the one-stage high-pressure polymerization unit and the multi-stage high-pressure polymerization unit; and/or, on at least one tubular reactor in the multi-stage high-pressure polymerization unit (including the
  • the device further includes at least one comonomer feed port (for example, 1, 2, 3, 4, 5, 6, etc.).
  • the arrangement of the comonomer feed port can be used to introduce the comonomer into the device to participate in one-stage high-pressure polymerization and/or multi-stage high-pressure polymerization, so as to utilize the device to produce polyolefin copolymers.
  • the number and position of the comonomer feed ports are not limited and can be selected according to needs.
  • the comonomer feed ports are arranged in a The comonomer feed port can be set at any position upstream of the outlet of the first-stage high-pressure polymerization unit, that is, according to the needs of the product, the comonomer feed port can be set at any position upstream of the outlet of at least one tubular reactor at the outlet of the first-stage high-pressure polymerization unit.
  • each comonomer feed port can be arranged in various ways.
  • each of the comonomer feed ports is respectively provided at: the inlet of the compression unit; and/or at the connecting pipe of any two adjacent compressors in the compression unit; and/or with the first-stage high-pressure polymerization
  • the reaction monomer flow inlet side of at least one tubular reactor in the unit; and/or, the connecting pipe between the preheater and the compression unit.
  • the device further includes a separation circulation unit located downstream of the multi-stage high-pressure polymerization unit for separating the materials obtained by the multi-stage high-pressure polymerization to obtain polymerized products and unreacted monomers. body.
  • a separation circulation unit located downstream of the multi-stage high-pressure polymerization unit for separating the materials obtained by the multi-stage high-pressure polymerization to obtain polymerized products and unreacted monomers. body.
  • the structure of the separation and circulation unit is not limited.
  • the separation and circulation unit includes a separator and a circulation loop.
  • the separator is used to separate the products from the multi-stage high-pressure polymerization unit, and the circulation loop is used to circulate unreacted monomers to the upstream of the first-stage high-pressure polymerization unit; further preferably, the separator includes high-pressure units connected in series. Separator 9 and low pressure separator 11.
  • the circulation loop includes a high circulation loop 10 connected to the high pressure separator 9 and a low circulation loop 12 connected to the low pressure separator 11. More preferably, the high circulation loop 10 is connected at one end.
  • the other end of the high-pressure separator 9 is connected to the suction side of the two-stage compressor 3; one end of the low-pressure separator 11 is connected to the low-pressure separator 11, and the other end is connected to the suction side of the first-stage compressor 2.
  • the separation circulation unit also includes a high-pressure relief valve 7 and a cooler 8 between the second-stage reactor and the low-pressure separator 11, for converting the high-pressure water from the multi-stage reactor to the low-pressure separator 11.
  • the high circulation loop 10 also includes a cooler and a separator for removing some components that are not unreacted monomers from the gaseous fraction in the high-pressure separator 9 (e.g. oligomers).
  • the low circulation loop 12 also includes a cooler and a separator for removing some components that are not unreacted monomers from the gaseous fraction in the low-pressure separator 11 (such as oligomers); further preferably, the low circulation loop 12 also includes at least one circulating material compressor 1 (for example, 1 circulating material compressor 1, 2 circulating material compressors 1, 3 Compressor 1 for circulating materials, etc.), further preferably, the compressor 1 for circulating materials is provided downstream of the cooler and separator on the low circulation circuit 12.
  • This method can not only obtain polymer products, but also better realize the recycling of unreacted monomers, which has better economic effects.
  • the device of the present invention is used to pass the product obtained by the multi-stage high-pressure polymerization in the method of the present invention in sequence through the high-pressure relief valve 7 and the cooler in the separation circulation unit and then flow into the separation
  • the gaseous fraction A and the liquid fraction A are separated in the high-pressure separator in the circulation unit, and the liquid fraction A flows into the low-pressure separator 11 in the separation circulation unit to separate the gaseous fraction B and the polymer product; further preferably, the gaseous fraction A After A is cooled and separated by the cooler and separator on the high circulation loop 10, the unreacted monomer flows into the suction side of the secondary compressor 3; and/or the gaseous fraction B passes through the cooler and separation on the low circulation loop 12 The unreacted monomer obtained after cooling and separation is compressed by the compressor on the low circulation circuit 12 and then flows into the suction side of the first-stage compressor 2; and/or the polymer is sent to be granulated to obtain the corresponding product.
  • the cooling and separation conditions of the coolers and separators on the high-pressure separator, low-pressure separator 11, high circulation circuit 10 and low circulation circuit 12 are not limited. Technology in the art Personnel can choose according to their needs.
  • the specific structures of the compressor, tubular reactor, high-pressure relief valve 7, cooler, high-pressure separator 9, low-pressure separator 11, and separator are not particularly limited, and they can be respectively Various compressors, tubular reactors, high-pressure relief valves 7, coolers, high-pressure separators, and low-pressure separators 11 commonly used in this field are all well known to those skilled in the art and will not be described in detail here.
  • each material inlet or material outlet can be cross-connected, and each equipment is connected through pipelines.
  • the device is also equipped with valves and other components to realize the circulation of materials.
  • the tubular reactor selects a cooling jacket structure to realize heat exchange. The invention has no special requirements for this, so no further details will be given here.
  • the present invention will be described in detail below through examples.
  • the number average molecular weight M n the weight average molecular weight M w and the polymer dispersion index PDI were measured by high temperature gel permeation chromatography HT-GPC method.
  • Olefin radical polymerization is carried out using an apparatus as shown in Figure 1, wherein in the apparatus of Figure 1: two parallel tubular reactors 5a and The length of the tubular reactor 5b is 560m and the inner diameter is 0.045m; the lengths of the three serially connected tubular reactors 6a, 6b, and 6c in the multi-stage high-pressure polymerization unit are all 400m, the inner diameter is 0.045m; the reaction monomer flow feed port end of the tubular reactor 5a and the tubular reactor 5b is equipped with an initiator feed port; the tubular reactor 6a, the tubular reactor 6b, the tube Each feed port end of the reactor 6c is provided with an initiator feed port; the outlet of the compression unit is provided with a telogen feed port;
  • a stream of material containing olefin source is sequentially compressed by the first-stage compressor 2 and the second-stage compressor 3 in the compression unit.
  • a stream of telogen is fed to the outlet of the compression unit through the telogen feed port and is then compressed.
  • the olefin source material is fully mixed, it is divided into two reaction monomer streams containing the olefin source in equal amounts.
  • the two reaction monomer streams containing the olefin source pass through the preheater 4a and the tubular reactor located upstream of the tubular reactor 5a respectively.
  • the preheater 4b upstream of 5b is introduced into the two parallel tubular reactors 5a and 5b in the first-stage high-pressure polymerization unit included in the device, each of which performs one-stage high-pressure polymerization, and then the obtained first-stage high-pressure polymerization
  • the polymerization product flows into three tubular reactors 6a, tubular reactions 6b, and tubular reactions 6c in the multi-stage high-pressure polymerization unit included in the device for multi-stage high-pressure polymerization; among them, five strands of free radical polymerization initiators pass through the device including
  • the initiator feed ports are respectively introduced to participate in the corresponding one-stage high-pressure polymerization or multi-stage high-pressure polymerization; the multi-stage high-pressure polymerization products pass through the high-pressure pressure relief valve 7 and the cooler 8 in the separation cycle unit in sequence and then flow into the high-pressure water in the separation cycle unit.
  • the gaseous fraction A and the liquid fraction A are separated in the separator 9, and the liquid fraction A flows into the low-pressure separator 11 in the separation cycle unit to separate the gaseous fraction B and the polymer product; the gaseous fraction A passes through the cooler on the high circulation loop 10 After being cooled and separated by the separator, the unreacted monomer is obtained and flows into the suction side of the secondary compressor; the gaseous fraction B is cooled and separated by the cooler and separator on the low circulation loop 12 and the unreacted monomer is obtained after being cooled and separated by the low circulation circuit 12.
  • the circulating material compressor 1 on the loop is compressed and flows into the suction side of the first-stage compressor; the polymer product is sent to the low-density polyethylene product obtained by granulation; among which:
  • the material containing the olefin source is ethylene; the inlet pressures of the two reaction monomer streams entering the tubular reactor 5a and the tubular reactor 5b respectively are: 220MPa;
  • the inlet temperatures of the two reaction monomer streams entering the tubular reactor 5a and the tubular reactor 5b are respectively: 170°C and 180°C;
  • the telogen is propylene; the feeding amount of the telogen is: 250kg/h;
  • the inlet temperatures of preheater 4a and preheater 4b are both: 92°C;
  • the feed amount of the reaction monomer flow at the reaction monomer flow feed port end of the tubular reactor 5a and the tubular reactor 5b is both 21750kg/h; the flow rates of the reaction monomer flow are 7.24m/s and 7.31m/s respectively. , the densities are 527kg/m 3 and 522kg/m 3 respectively; the density and viscosity values of the reaction monomer flow of the tubular reactor 5a and the tubular reactor 5b are 5383 and 5398 respectively;
  • compositions of the first free radical polymerization initiator and the second free radical polymerization initiator are: di-tert-butyl peroxide, peroxide mixed in a mass ratio of 1:2:2:2 A mixture of tert-butyl benzoate, tert-butyl peroxy-2-ethylhexanoate and tert-butyl peroxypivalate;
  • the compositions of the third to fifth strands of free radical polymerization initiators are respectively: according to the mass ratio: A 10:2:1:1 mixture of di-tert-butyl peroxide, tert-butyl peroxybenzoate, tert-butyl peroxy-2-ethylhexanoate and tert-butyl peroxypivalate;
  • the feed amounts of the radical polymerization initiator at the initiator feed ports of the reaction monomer flow feed ports of the tubular reactor 5a and the tubular reactor 5b are respectively: 6.78kg/h and 6.19kg/h;
  • the feed amounts of the radical polymerization initiator at the initiator feed ports of the tubular reactor 6a, the tubular reactor 6b, and the tubular reactor 6c are 7.33kg/h, 7.44kg/h, and 7.80kg respectively. /h;
  • tubular reactor 6a tubular reactor 6a
  • tubular reactor 6b tubular reactor 6b
  • tubular reactor 6c tubular reactor 6c
  • tubular reactor 5a The peak temperatures in tubular reactor 5a, tubular reactor 5b, and tubular reactor 6a, tubular reactor 6b, and tubular reactor 6c are all 295°C;
  • the separation conditions of the high-pressure separator are: 25MPa, 235°C;
  • the separation conditions of the low-pressure separator are: 2bar, 220°C;
  • the pressure drop of the first-stage high-pressure polymerization unit is 4.8MPa, and the pressure drop of the multi-stage high-pressure polymerization unit is 30.6MPa.
  • Olefin radical polymerization is carried out using an apparatus as shown in Figure 2.
  • a preheater 4 is located in two parallel tubular reactors 5a and tubular reactor 5b in the first-stage high-pressure polymerization unit.
  • the olefin source-containing material is fully mixed and passes through the preheater 4, and is divided into two equal amounts of reaction monomer streams containing olefin source and introduced into two parallel tubular reactors 5a in the first-stage high-pressure polymerization unit included in the device. and tubular reactor 5b; where:
  • the inlet temperatures of the two reaction monomer flows entering the tubular reactor 5a and the tubular reactor 5b respectively are both: 170°C; the flow rates of the reaction monomer flows are both 7.24m/s, and the densities are both 527kg/m 3 ; the tube
  • the telogen is propylene; the feed rate of the telogen is: 180kg/h;
  • the inlet temperature of preheater 4 is: 92°C;
  • the feed amounts of the radical polymerization initiator at the initiator feed ports of the reaction monomer flow feed ports of the tubular reactor 5a and the tubular reactor 5b are respectively: 6.80kg/h and 6.80kg/h;
  • the feed amounts of the radical polymerization initiator at the initiator feed ports of the tubular reactor 6a, the tubular reactor 6b, and the tubular reactor 6c are 7.41kg/h, 7.26kg/h, and 7.36kg respectively. /h;
  • tubular reactor 6a tubular reactor 6a
  • tubular reactor 6b tubular reactor 6b
  • tubular reactor 6c tubular reactor 6c
  • the pressure drop of the first-stage high-pressure polymerization unit is 5.2MPa, and the pressure drop of the multi-stage high-pressure polymerization unit is 33.5MPa.
  • the telogen is propylene; the feeding amount of the telogen is: 195kg/h;
  • the feed amounts of the radical polymerization initiator at the initiator feed ports of the reaction monomer flow feed ports of the tubular reactor 5a and the tubular reactor 5b are respectively: 6.47kg/h and 6.22kg/h;
  • the peak temperatures in the tubular reactor 5a and tubular reactor 5b are both 295°C;
  • the separation conditions of the high-pressure separator are: 23MPa, 200°C
  • the separation conditions of the low-pressure separator are: 2bar, 190°C;
  • the pressure drop of the first-stage high-pressure polymerization unit is 5.0MPa, and the pressure drop of the multi-stage high-pressure polymerization unit is 28.1MPa.
  • the tubular reactor 5a is provided with a telogen feed port on the side of the reaction monomer flow feed port, and there is no telogen feed port set in other positions; a stream of telogen agent passes through The feed from the telogen feed port is mixed with the reaction monomer flow flowing into the tubular reactor 5a and then introduced into the tubular reactor 5a to perform the corresponding one-stage high-pressure polymerization;
  • the inlet temperatures of the two reaction monomer flows entering the tubular reactor 5a and the tubular reactor 5b are: 170°C and 190°C respectively; the flow rates of the reaction monomer flows are 7.24m/s and 7.37m/s respectively, and the densities are They are 527kg/m 3 and 518kg/m 3 respectively; the density and viscosity values of the reaction monomer flow of the tubular reactor 5a and the tubular reactor 5b are 5383 and 5424 respectively;
  • the telogen is propylene; the feed rate of the telogen is: 170kg/h;
  • the feed amounts of the radical polymerization initiator at the initiator feed ports of the reaction monomer flow feed ports of the tubular reactor 5a and the tubular reactor 5b are respectively: 6.71kg/h and 6.25kg/h;
  • the feed amounts of the radical polymerization initiator at the initiator feed ports of the tubular reactor 6a, the tubular reactor 6b, and the tubular reactor 6c are 7.32kg/h, 7.42kg/h, and 7.78 respectively. kg/h;
  • tubular reactor 6a tubular reactor 6a
  • tubular reactor 6b tubular reactor 6b
  • tubular reactor 6c tubular reactor 6c
  • the pressure drop of the first-stage high-pressure polymerization unit is 5.1MPa, and the pressure drop of the multi-stage high-pressure polymerization unit is 33.8MPa.
  • the tubular reactor 5a and the tubular reactor 5b are respectively provided with a telogen feed port on the reaction monomer flow feed port side, and there is no telogen feed port provided in other positions of the device. ; After the two strands of telogen are respectively fed into the reaction monomer flow inlet side of the tubular reactor 5a and the reaction monomer flow inlet side of the tubular reactor 5b through the telogen feed inlet, they are respectively connected with The two reaction monomer streams are fully mixed and then flow into the tubular reactor 5a and the tubular reactor 5b to perform one-stage high-pressure polymerization respectively; where:
  • the telogen is propylene; the feeding amount of both telogens is: 97.5kg/h;
  • the feed amounts of the radical polymerization initiator at the initiator feed port of the reaction monomer flow feed port of the tubular reactor 5a and the tubular reactor 5b are respectively: 6.78kg/h and 6.19kg/h;
  • tubular reactor 6a tubular reactor 6a
  • tubular reactor 6b tubular reactor 6b
  • tubular reactor 6c tubular reactor 6c
  • the pressure drop of the first-stage high-pressure polymerization unit is 5.0MPa, and the pressure drop of the multi-stage high-pressure polymerization unit is 32.9MPa.
  • a telogen feed port is provided at the connecting pipe between the one-stage high-pressure polymerization unit and the multi-stage high-pressure polymerization unit. There is no telogen feed port provided in other locations of the device.
  • the material is fed through the telogen feed port to the connecting pipe of the first-stage high-pressure polymerization unit and the multi-stage high-pressure polymerization unit, and is thoroughly mixed with the first-stage high-pressure polymerization product and then flows into the multi-stage high-pressure polymerization unit;
  • the inlet temperatures of the two reaction monomer streams entering the tubular reactor 5a and the tubular reactor 5b respectively are: 170°C and 150°C; the flow rates of the reaction monomer streams are 7.24m/s and 7.37m/s respectively, and the densities They are 527kg/m 3 and 537kg/m 3 respectively; the density and viscosity values of the reaction monomer flow of the tubular reactor 5a and the tubular reactor 5b are 5383 and 5317 respectively;
  • the telogen is propylene; the feeding amount of the telogen is: 195kg/h;
  • the feed amounts of the radical polymerization initiator at the initiator feed ports of the reaction monomer flow feed ports of the tubular reactor 5a and the tubular reactor 5b are respectively: 6.72kg/h and 2.40kg/h;
  • the feed amounts of the radical polymerization initiator at the initiator feed ports of the tubular reactor 6a, the tubular reactor 6b, and the tubular reactor 6c are 7.35kg/h, 7.42kg/h, and 7.81kg respectively. /h;
  • tubular reactor 6a tubular reactor 6a
  • tubular reactor 6b tubular reactor 6b
  • tubular reactor 6c tubular reactor 6c
  • the temperature peak of tubular reactor 5b is 250°C, and the temperature peak of other tubular reactors is 295°C;
  • the pressure drop of the first-stage high-pressure polymerization unit is 5.6MPa, and the pressure drop of the multi-stage high-pressure polymerization unit is 34.5MPa.
  • Olefin radical polymerization is carried out using an apparatus as shown in Figure 1, wherein in the apparatus of Figure 1: two parallel tubular reactors 5a and The length of the tubular reactor 5b is 560m and the inner diameter is 0.030m;
  • the feed amount of the reaction monomer flow at the reaction monomer flow feed port end of the tubular reactor 5a and the tubular reactor 5b is both 21750kg/h; the flow rates of the reaction monomer flow are 16.29m/s and 16.45m/s respectively. ;
  • the feed amounts of the radical polymerization initiator at the initiator feed ports of the reaction monomer flow feed ports of the tubular reactor 5a and the tubular reactor 5b are respectively: 6.78kg/h and 6.19kg/h;
  • the feed amounts of the radical polymerization initiator at the initiator feed ports of the tubular reactor 6a, the tubular reactor 6b, and the tubular reactor 6c are 8.00kg/h, 5.05kg/h, and 8.49kg respectively. /h;
  • tubular reactor 6a tubular reactor 6a
  • tubular reactor 6b tubular reactor 6b
  • tubular reactor 6c tubular reactor 6c
  • the pressure drop of the first-stage high-pressure polymerization unit is 21.8MPa, and the pressure drop of the multi-stage high-pressure polymerization unit is 30.3MPa.
  • Olefin radical polymerization is carried out using an apparatus as shown in Figure 1, wherein in the apparatus of Figure 1: two parallel tubular reactors 5a and The lengths of the tubular reactor 5b are 560m and 325m respectively, and the inner diameters are 0.045m and 0.024m respectively; the feed amounts of the reaction monomer flow at the reaction monomer flow feed port ends of the tubular reactor 5a and the tubular reactor 5b are respectively are 37285kg/h and 6215kg/h; the flow rates of the reaction monomer flow are 12.41m/s and 7.47m/s respectively;
  • the feed amounts of the radical polymerization initiator at the initiator feed ports of the reaction monomer flow feed ports of the tubular reactor 5a and the tubular reactor 5b are respectively: 10.62kg/h and 2.45kg/h;
  • the feed amounts of the radical polymerization initiator at the initiator feed ports of the tubular reactor 6a, the tubular reactor 6b, and the tubular reactor 6c are 7.98kg/h, 4.92kg/h, and 8.32kg respectively. /h;
  • tubular reactor 6a tubular reactor 6a
  • tubular reactor 6b tubular reactor 6b
  • tubular reactor 6c tubular reactor 6c
  • the pressure drop of the first-stage high-pressure polymerization unit is 10.2MPa, and the pressure drop of the multi-stage high-pressure polymerization unit is 30.5MPa.
  • the one-stage high-pressure polymerization unit connected in series upstream of the multi-stage high-pressure polymerization unit is a tubular reactor 5a (the length of the tubular reactor 5a is 560m and the inner diameter is 0.045m).
  • the tubular reactor 5b and preheater 4b There is no Tubular reactor 5b and preheater 4b; four initiator feed ports are provided in the device, which are respectively located at the reaction monomer flow feed port end of the tubular reactor 5a, the tubular reactor 6a, and the tubular reactor.
  • a stream of material containing olefin source is sequentially compressed by the first-stage compressor 2 and the second-stage compressor 3 in the compression unit.
  • a stream of telogen is fed to the outlet of the compression unit through the telogen feed port and is then compressed.
  • the materials of the olefin source are mixed and passed through the preheater 4a to obtain a reaction monomer stream containing the olefin source.
  • Four radical polymerization initiators are introduced through the initiator feed port included in the device to participate in the corresponding first-level high-pressure polymerization or Multi-stage high pressure polymerization; where:
  • the inlet pressure of the reaction monomer flow into the tubular reactor 5a is: 220MPa
  • the inlet temperature of the reaction monomer flow into the tubular reactor 5a is: 175°C;
  • the feed amount of the reaction monomer flow at the monomer flow feed port end of the tubular reactor 5a is 43500kg/h;
  • Free radical polymerization initiator feed to the reaction monomer flow feed port end of the tubular reactor 5a, the initiator feed port of the tubular reactor 6a, tubular reactor 6b, and tubular reactor 6c The quantities are: 13.37kg/h, 6.56kg/h, 7.71kg/h and 7.39kg/h;
  • tubular reactor 6a tubular reactor 6a
  • tubular reactor 6b tubular reactor 6b
  • tubular reactor 6c tubular reactor 6c
  • tubular reactor 5a The peak temperatures in tubular reactor 5a, tubular reactor 6a, tubular reactor 6b, and tubular reactor 6c are all 295°C;
  • the total pressure drop of the tubular reactor is 44.5MPa.
  • the one-stage high-pressure polymerization unit connected in series upstream of the multi-stage high-pressure polymerization unit is a tubular reactor 5a (the length of the tubular reactor 5a is 560m and the inner diameter is 0.045m).
  • the device is provided with four initiator feed ports, which are respectively provided at the reaction monomer flow feed port end of tubular reactor 5a, tubular reactor 6a, tubular reactor 6b, tubular reactor The feed port end of reactor 6c; no initiator feed port is provided in other locations of the device;
  • a stream of olefin source-containing material is compressed by the primary compressor 2 and the secondary compressor 3 in the compression unit in sequence, and a stream of telogen is passed through the conditioning unit.
  • the polymerization agent feed port is fed to the outlet of the compression unit and mixed with the compressed material containing olefin sources. After passing through the preheater 4a, a reaction monomer stream containing olefin sources is obtained.
  • Four streams of free radical polymerization initiators pass through the device respectively.
  • the included initiator feed ports are each introduced to participate in the corresponding one-stage high-pressure polymerization or multi-stage high-pressure polymerization; where:
  • the inlet pressure of the reaction monomer flow into the tubular reactor 5a is: 220MPa;
  • the inlet temperature of the reaction monomer flow into the tubular reactor 5a is: 175°C;
  • the feed amount of the reaction monomer flow at the monomer flow feed port end of the tubular reactor 5a is 43500kg/h;
  • Free radical polymerization initiator feed to the reaction monomer flow feed port end of the tubular reactor 5a, the initiator feed port of the tubular reactor 6a, tubular reactor 6b, and tubular reactor 6c The quantities are: 13.34kg/h, 6.54kg/h, 7.36kg/h and 6.82kg/h;
  • tubular reactor 6a tubular reactor 6a
  • tubular reactor 6b tubular reactor 6b
  • tubular reactor 6c tubular reactor 6c
  • tubular reactor 5a The peak temperatures in tubular reactor 5a, tubular reactor 6a, tubular reactor 6b, and tubular reactor 6c are all 295°C;
  • the total pressure drop of the tubular reactor is 46.1MPa.
  • the one-stage high-pressure polymerization unit connected in series upstream of the multi-stage high-pressure polymerization unit is a tubular reactor 5a (the length of the tubular reactor 5a is 560m and the inner diameter is 0.045m).
  • the device is provided with an initiator feed port, which is located at the reaction monomer flow feed port end of the tubular reactor 5a; no initiator feed port is provided in other locations of the device ;
  • a stream of material containing olefin source is sequentially compressed by the first-stage compressor 2 and the second-stage compressor 3 in the compression unit.
  • a stream of telogen is fed to the outlet of the compression unit through the telogen feed port and is then compressed.
  • the materials of the olefin source are mixed and passed through the preheater 4a to obtain a reaction monomer stream containing the olefin source.
  • a free radical polymerization initiator is introduced through the initiator feed port included in the device to participate in the corresponding first-level high-pressure polymerization.
  • the inlet pressure of the reaction monomer flow into the tubular reactor 5a is: 220MPa;
  • the inlet temperature of the reaction monomer flow into the tubular reactor 5a is: 175°C;
  • the feed amount of the reaction monomer flow at the reaction monomer flow feed port end of the tubular reactor 5a is 43500kg/h;
  • the feed amount of the free radical polymerization initiator at the initiator feed port of the reaction monomer flow feed port of the tubular reactor 5a is: 12.70kg/h;
  • the peak temperature of tubular reactor 5a is 295°C;
  • the total pressure drop of the tubular reactor is 41.2MPa.
  • a tube reactor with a total length of 1760m and an inner diameter of 0.045m is used to replace the one-stage high-pressure polymerization unit and the multi-stage high-pressure polymerization unit.
  • An initiator feed port is provided at 560m, 960m, and 1360m along its length direction; the reaction monomer flow containing olefin sources is introduced into the tubular reactor for high-pressure polymerization, and the free radical polymerization initiator passes through the initiator feed port Continuously feed into the device to participate in high-pressure polymerization (the product obtained by high-pressure polymerization is equivalent to the product obtained by multi-stage high-pressure polymerization in Example 1);
  • the inlet pressure of the reaction monomer flow entering the tubular reactor is: 220MPa;
  • the inlet temperature of the reaction monomer flow entering the tubular reactor is: 175°C;
  • the telogen is propylene; the feed rate of the telogen is: 170kg/h;
  • the feed amount of the reaction monomer flow at the feed port of the tubular reactor is: 43500kg/h;
  • the free radical polymerization initiator feed amounts at the reaction monomer flow feed port end of the tubular reactor and the four initiator feed ports at 560m, 960m, and 1260m along the length of the tubular reactor are 14.29kg/ h, 7.40kg/h, 7.27kg/h, 7.32kg/h;
  • the inlet temperatures at the reaction monomer flow inlet end of the tubular reactor and at 560m, 960m, and 1260m along the length of the tubular reactor are 193°C, 211°C, and 229°C respectively;
  • the peak temperatures of the reaction monomer flow inlet end of the tubular reactor and the four zones divided into 560m, 960m and 1260m along the length of the tubular reactor are all 295°C respectively;
  • the total pressure drop of the tubular reactor is 47.6MPa.
  • the one-stage high-pressure polymerization unit connected in series upstream of the multi-stage high-pressure polymerization unit is a tubular reactor 5a (the length of the tubular reactor 5a is 560m and the inner diameter is 0.045m).
  • the tubular reactor 5b and preheater 4b There is no Tubular reactor 5b and preheater 4b; four initiator feed ports are provided in the device, which are respectively located at the reaction monomer flow feed port end of the tubular reactor 5a, the tubular reactor 6a, and the tubular reactor.
  • telogen feed port is provided at the connecting pipe between tubular reactor 5a and the multi-stage high-pressure polymerization unit , no telogen feed port is provided in other locations of the device; a stream of telogen is fed into the tubular reactor 5a in the first-stage high-pressure polymerization unit and the connecting pipe of the multi-stage high-pressure polymerization unit through the telogen feed port After being fully mixed with the first-stage high-pressure polymerization product, it flows into the multi-stage high-pressure polymerization unit;
  • the inlet pressure of the reaction monomer flow into the tubular reactor 5a is: 220MPa;
  • the inlet temperature of the reaction monomer flow entering the tubular reactor 5a is: 175°C;
  • the feed amount of the reaction monomer flow at the monomer flow feed port end of the tubular reactor 5a is 43500kg/h;
  • Free radical polymerization initiator feed to the reaction monomer flow feed port end of the tubular reactor 5a, the initiator feed port of the tubular reactor 6a, tubular reactor 6b, and tubular reactor 6c The quantities are: 13.19kg/h, 6.58kg/h, 7.68kg/h and 7.38kg/h;
  • tubular reactor 6a tubular reactor 6a
  • tubular reactor 6b tubular reactor 6b
  • tubular reactor 6c tubular reactor 6c
  • tubular reactor 5a The peak temperatures in tubular reactor 5a, tubular reactor 6a, tubular reactor 6b, and tubular reactor 6c are all 295°C;
  • the total pressure drop of the tubular reactor is 48.3MPa.
  • the devices in Examples 1-8 of the present invention can control the inlet temperature and reactor temperature of at least two parallel tubular reactors in the first-stage high-pressure polymerization unit according to the performance requirements of downstream products. Peak, telogen feeding position and other parameters, so as to control the molecular weight distribution (MWD) of the product while ensuring a certain conversion rate; and, at the same peak temperature, through Examples 1-8 and Comparative Examples 1-5
  • the device of the present invention can produce products with a wider molecular weight distribution, overcoming the shortcomings of the tubular method that cannot produce products of the kettle method.
  • the device of the present invention has a higher ethylene conversion rate.
  • the method and device of the present invention have the effects of convenient control, easy operation, and wide product coverage.
  • Some preferred embodiments of the present invention also provide a method and device for free radical polymerization of ethylene.
  • the invention also provides a method for free radical polymerization of ethylene, which method includes: introducing a reaction monomer stream containing an ethylene source into at least two parallel tubular reactors to react in the presence of an initiator; Part of the material from the outlet of at least one of the tubular reactors is recycled back to at least one of at least two parallel tubular reactors for reaction; the at least two parallel tubular reactors are The remaining material from the outlet of the reactor continues to flow into one or more tubular reactors connected in series to react in the presence of an initiator; the pressure of the reaction monomer stream containing the ethylene source is greater than or equal to 100 MPa.
  • the kettle reactor can produce polyethylene products with a higher degree of branching and a wider molecular weight.
  • the polyethylene products produced by the tubular reactor have a narrower molecular weight distribution and lower long-chain branch content.
  • the inventor found that , compared with the existing technology, the method of the present invention can be used to produce polyethylene products with high degree of branching and wider molecular weight that cannot be produced by existing tubular reactors.
  • a reaction monomer stream containing an ethylene source is introduced into at least two parallel tubular reactors to react in the presence of an initiator, and at the same time, the output of at least one tubular reactor in the at least two parallel tubular reactors is Part of the material from the feed port is circulated back to at least one of at least two parallel tubular reactors for reaction, which enables the newly generated free radicals to be transferred to the molecular chain of the polymer, thereby producing a molecular weight distribution width and Polyethylene products with a wider range of long-chain branching degrees broaden the application fields of tubular process products, and can also improve the utilization efficiency of raw materials, resulting in an increase in conversion rate and output.
  • the circulation ratio of materials at the outlet of the at least two parallel tubular reactors is less than 1, preferably less than or equal to 0.3, and more preferably 0.04-0.2 , for example, 0.04, 0.05, 0.07, 0.09, 0.1, 0.12, 0.13, 0.15, 0.18, 0.2, and any value within the range consisting of any two of the above values.
  • polyethylene products with different molecular weight distributions can be obtained.
  • polyethylene products with a wide molecular distribution that cannot be produced by the tubular method in the prior art can be obtained, while ensuring that better ethylene can be maintained. Conversion rate and yield.
  • the term "circulation ratio" refers to the total mass of materials recycled from the outlet of the tubular reactor back to at least one tubular reactor of at least two parallel tubular reactors to all at least The mass ratio of the total mass of materials from all outlets of two parallel tubular reactors; at the same time, the outlet of at least one tubular reactor is recycled back to at least one tubular reactor of at least two parallel tubular reactors.
  • the materials in the reactor can be selected from at least two parallel tubular reactors. Part of the material at the outlet of one of the tubular reactors or the total material collected from the partial materials at the outlets of multiple tubular reactors.
  • part of the material from the outlet of at least one of the at least two parallel tubular reactors is recycled back to at least one tubular reaction of the at least two parallel tubular reactors.
  • the fluid suction and transportation equipment may have a pressure drop loss.
  • the pressure drop is less than 60MPa, preferably less than 30MPa, and further preferably less than 15MPa (for example, the pressure drop is 15MPa, 12MPa, 8MPa, 6MPa, and any value within the range of any of the above values).
  • the pressure of the reactive monomer stream containing the ethylene source flowing into at least two parallel tubular reactors can enable the reactive monomer stream containing the ethylene source to undergo high-pressure free radical polymerization under the conditions of the initiator. That is, in some preferred embodiments, the pressure of the reaction monomer flow containing the ethylene source flowing into at least two parallel tubular reactors is 140-300MPa (such as 140MPa, 160MPa, 220MPa, 300MPa, and the above range of arbitrary values).
  • the chain growth rate and the ease of chain transfer during the reaction can be controlled as needed, thereby better controlling the distance between ethylene and ethylene molecules, and obtaining products with different densities and branch chain distributions, so that
  • the method of the invention can produce heavy packaging film materials, agricultural film materials, injection plastics, coating materials, medical packaging materials and other products.
  • the inlet pressure of the reaction monomer flow containing the ethylene source into the polymerization unit can be adjusted through a pressure reducing valve on the inlet side of the tubular reactor.
  • the inlet pressure of the reaction monomer stream containing the ethylene source introduced into the corresponding tubular reactor through the fluid suction conveying device is equal to the reaction through the ethylene source with a certain pressure.
  • the pressure of the monomer flow (that is, the pressure before entering the fluid suction and transportation equipment) minus the pressure drop existing in the fluid suction and transportation equipment; in addition, it can be understood that the content of the monomer flow that has not been introduced into the corresponding tubular reactor through the fluid suction and transportation equipment
  • the inlet pressure of the reaction monomer of the ethylene source is equal to the pressure of the reaction monomer flow containing the ethylene source having a certain pressure, or the inlet pressure of the reaction monomer flow containing the ethylene source is reduced through the pressure reducing valve on the inlet side of the tubular reactor.
  • the reaction temperature of the reaction in at least two parallel tubular reactors and one or more serially connected tubular reactors there is no limit to the reaction temperature of the reaction in at least two parallel tubular reactors and one or more serially connected tubular reactors, as long as high-pressure free radical polymerization can occur.
  • the reaction temperatures in the one-stage high-pressure polymerization unit and the multi-stage high-pressure polymerization unit are each 100-350°C (for example, 100°C, 130°C, 150°C, 170°C, 200°C, 250°C, 295°C, 300°C °C, 350°C, and any value within the range of any of the above values).
  • the reaction temperature in at least two parallel tubular reactors and one or more serially connected tubular reactors is a temperature within a range, specifically including, The inlet temperature, outlet temperature and peak temperature of the reaction zone in at least two parallel tubular reactors and one or more serially connected tubular reactors are all within the reaction temperature range.
  • the temperature at the outlet of each reaction zone is less than or equal to the peak temperature of the corresponding reaction zone;
  • the temperature at the inlet of each reaction zone is less than the peak temperature of the corresponding reaction zone;
  • the at least two parallel tubular reactors are Tubular reactors, and one or more tubular reactors connected in series: the absolute value of the difference between the temperature at the outlet of each reaction zone and the peak temperature of the corresponding reaction zone is 0-150°C (for example, 20°C, 50°C °C, 70 °C, 100 °C, 120 °C, 150 °C, and any value within the range of any of the above values), more preferably 20-120 °C.
  • the temperature of the reaction monomer stream containing the ethylene source flowing into at least two parallel tubular reactors can be obtained after being preheated by a preheating unit.
  • the temperature of the reaction monomer flow flowing into at least two parallel tubular reactors is 140-190°C (such as 140°C, 150°C, 170°C, 180°C, 190°C, and the above Any value within the range of any numerical value); in some preferred embodiments, the at least two parallel tubular reactors, and the respective reaction zones in one or more serially connected tubular reactors
  • the peak temperature is 200-300°C (such as 200°C, 225°C, 260°C, 295°C, 300°C, and any value within the range of any of the above values); using the aforementioned embodiments, the reaction zone can be better controlled Temperature distribution, while using the aforementioned embodiments, polyethylene products with different molecular weight distributions can be obtained.
  • the weight ratio of the maximum feed amount to the minimum feed amount of the reaction monomer stream containing the ethylene source in the at least two parallel tubular reactors Is 1: (0.01-1), such as 1:0.1, 1:0.5, 1:1, and any value within the range of any of the above values.
  • At least two initiators each participate in the at least two parallel tubular reactors; at least one initiator participates in one or more serially connected tubular reactors. reaction in.
  • the newly generated free radicals can be better transferred to the molecular chain of the polymer, thereby producing a product with a wider range of molecular weight distribution width and long chain branching degree.
  • the wide range of polyethylene products broadens the application fields of tubular products, and can also improve the utilization efficiency of raw materials, resulting in an increase in conversion rate and output.
  • the method of the present invention further includes at least one chain transfer agent participating in the at least two parallel tubular reactors, and one or more tubular reactions in series. reaction in the device.
  • a chain transfer agent can transfer newly generated free radicals to the molecular chain of the polymer, thereby producing polyethylene products with a wider range of molecular weight distribution width and long-chain branching degree, and broadening the scope of the tubular method.
  • the application fields of the product can also improve the utilization efficiency of raw materials, resulting in an increase in conversion rate and increase in output.
  • the method of the present invention further includes at least one comonomer participating in at least two parallel tubular reactors, and one or more serially connected tubular reactors. Reaction.
  • the method of the present invention can also obtain different ethylene copolymers.
  • the outlet of the comonomer supply unit used to install the comonomer supply unit included in the method of the present invention can be introduced into different positions of the device according to the needs of the present invention.
  • the method of the present invention also includes separating the materials obtained in one or more tubular reactors connected in series to obtain polymerization products and unreacted monomers; in order to make the present invention
  • the invention has good economic effects. It is further preferred that the unreacted monomer is recycled back to the upstream of at least two parallel tubular reactors to continue the reaction.
  • the conditions for the cyclic separation can be selected as needed, and will not affect the purpose of the present invention, so they will not be described in detail here.
  • the specific selection of the initiator is not limited, and any initiator that can generate free radical substances under the reaction conditions of the present invention is suitable for use in the present invention.
  • the initiator is selected from one or more of azo compounds, organic peroxides, oxygen and air.
  • the initiator in the present invention can be used directly or dissolved in a solvent. used in.
  • azo compounds include azobisisobutyronitrile, azobisisovaleronitrile and azobisisoheptanitrile;
  • organic peroxides include 2,2-bis(tert-butylperoxy)propane, Tert-butyl peroxy-2-ethylhexanoate, bis-(2-ethylhexyl peroxydicarbonate), di-tert-butyl peroxide, dialkyl peroxide, tert-butyl peroxybenzoate , 1,1-bis(tert-butylperoxy)cyclohexane, tert-butyl peroxypivalate.
  • the initiator can be introduced in any state, such as liquid, dissolved state or supercritical state.
  • a gaseous radical polymerization initiator such as oxygen or air
  • the initiator is introduced in a supercritical state.
  • the mixture of di-tert-butyl peroxide, tert-butyl peroxybenzoate, tert-butyl peroxy-2-ethylhexanoate and tert-butyl peroxypivalate is used as an example to illustrate the present invention. advantages, but the invention is not limited thereto.
  • the type of the chain transfer agent is not limited.
  • the chain transfer agent is selected from one or more of aliphatic hydrocarbons, olefins, ketones, aldehydes, aliphatic alcohols and hydrogen.
  • aliphatic hydrocarbons include propane, butane and cyclohexane
  • examples of olefins include propylene and hexene.
  • any monomer that can copolymerize with ethylene under high-pressure radical polymerization conditions can be used as the comonomer of the present invention, such as propylene and/or vinyl acetate.
  • the ratio of the molecular weight distribution of the low-density polyethylene obtained by the method of the present invention to the conversion rate of the ethylene source is greater than or equal to 0.01 and less than or equal to 0.05, preferably greater than or equal to 0.018 and less than or equal to 0.018. 0.048.
  • the invention also provides a device for ethylene free radical polymerization, which device includes: a polymerization unit, a fluid suction and transportation unit, and an initiator supply unit; the polymerization unit includes a first-level high-pressure polymerization unit and a first-level high-pressure polymerization unit connected in series downstream of the first-level high-pressure polymerization unit.
  • the multi-stage high-pressure polymerization unit; the first-stage high-pressure polymerization unit includes at least two parallel tubular reactors; the multi-stage high-pressure polymerization unit includes one or more tubular reactors connected in series; the fluid suction and transportation The unit includes one or at least two fluid suction and conveying devices arranged in parallel for inhaling and conveying at least one reaction monomer stream containing an ethylene source and the outlet of at least one tubular reactor in the first-stage high-pressure polymerization unit. Part of the materials; the initiator supply unit is used to deliver initiator to the polymerization unit.
  • the initiator supply unit supplies the initiator into the device through the outlet of the initiator supply unit, and the outlet of the initiator is in the device of the invention.
  • the location and the number of outlets of the initiator supply unit can be determined based on The selection is made according to the location and number of the required reaction zones.
  • those skilled in the art can understand that when the initiator supply unit transports the initiator to the polymerization unit, one of the initiator supply units
  • the connection point between the discharge port and any part of the device is the "reaction zone entrance".
  • reaction zone a zone where the temperature rises
  • a section of the area is the corresponding "cooling zone”; when a unit has only one reaction zone in the direction of material flow, the outlet in the direction of material flow is the “reaction zone outlet” or “cooling zone outlet”; when a unit When the unit has multiple reaction zones connected in series, along the flow direction of the material, the connection point between the last outlet in the initiator supply unit and the downstream position of the adjacent previous "reaction zone entrance” is the next “"Reaction zone entrance” is also the “reaction zone exit” or “cooling zone exit” of the previous reaction zone, and so on. The last "reaction zone exit” or “cooling zone exit” in the initiator supply unit is the end of the reaction.
  • the last outlet; for example, the multi-stage high-pressure polymerization unit in the device of the present invention includes two reaction zones connected in series, that is, the first outlet in the initiator supply unit and the outlet in the multi-stage high-pressure polymerization unit
  • the first connection point is the "first reaction zone entrance”.
  • the connection point is the “second reaction zone entrance” and also the “first reaction zone exit” or the “first cooling zone exit”; from the “second reaction zone entrance” there is a zone of rising temperature along the flow direction of the material. It is called the “second reaction zone”.
  • the second reaction zone When the temperature rises to the peak temperature, the area where the temperature begins to decrease is the corresponding "second cooling zone”.
  • the outlet after the multi-stage unit reaction is completed is the “second reaction zone outlet.” ” or “Second Cooling Zone Exit”.
  • the tubular reactor in the one-stage high-pressure polymerization unit, the tubular reactor in the multi-stage high-pressure polymerization unit, the fluid suction conveying equipment and the initiator supply unit can introduce materials and discharge materials according to the materials. It is required to set up one or more feed ports and discharge ports, which will not be described in detail here; the tubular reactor in the one-stage high-pressure polymerization unit and the multi-stage high-pressure polymerization unit is not limited, and is preferably a tubular reactor with a sleeve structure.
  • the reactor can be a single casing structure or a multi-casing structure. The reactor using the casing structure can better control the pressure in the one-stage high-pressure polymerization unit and the multi-stage high-pressure polymerization unit when the device of the present invention is used. temperature reflex.
  • the kettle reactor can produce polyethylene products with a higher degree of branching and a wider molecular weight.
  • the polyethylene products produced by the tubular reactor have a narrower molecular weight distribution and lower long-chain branch content.
  • the device of the present invention can be used to produce polyethylene products with high degree of branching and wider molecular weight that cannot be produced by existing tubular reactors.
  • the inventor speculates that at least one of the devices of the present invention Two parallel tubular reactors, and a fluid suction for sucking and delivering at least one reaction monomer stream containing an ethylene source and a part of the material outlet of at least one tubular reactor in the first-stage high-pressure polymerization unit
  • the newly generated free radicals can be transferred to the molecular chain of the polymer, thereby producing polyethylene products with a wider range of molecular weight distribution width and long-chain branching degree, broadening the scope of tubular It can also improve the application field of raw materials, increase the conversion rate and increase the output.
  • the number of tubular reactors included in the first-stage high-pressure polymerization unit is not limited.
  • the first-stage high-pressure polymerization unit includes 2 -4 parallel tubular reactors.
  • the number of reaction zones included in the multi-stage high-pressure polymerization unit is not limited.
  • the multi-stage high-pressure polymerization unit includes 1 or 2-6 reaction zones connected in series; further preferably, 2-4 reaction zones connected in series. Using the aforementioned preferred embodiments, the molecular weight distribution of the polyethylene product can be effectively adjusted as needed.
  • the main function of the fluid suction and transportation equipment is to suck in and transport materials.
  • equipment including but not limited to jet pumps; the number of fluid suction and transportation equipment It can be set as needed.
  • the number of fluid suction delivery devices is less than or equal to the reaction monomer flow containing the ethylene source introduced into the tubular reactor in the first-stage high-pressure polymerization unit.
  • the number of strands, the number of tubular reactors in the first-stage high-pressure polymerization unit is greater than or equal to the number of strands of the reaction monomer stream containing the ethylene source.
  • At least one of the fluid suction and delivery devices is arranged in series upstream of at least two parallel tubular reactors in the first-stage high-pressure polymerization unit; in some embodiments, at least one of the fluid suction and delivery equipment is arranged in series upstream of at least two parallel tubular reactors in the first-stage high-pressure polymerization unit.
  • the device of the present invention in order to enable the reaction monomer flow containing the ethylene source to have an inlet pressure into the polymerization unit, in some embodiments, the device of the present invention further includes a fluid suction and delivery unit located upstream of the polymerization unit.
  • compression unit preferably, the compression unit includes at least a 2-stage compressor.
  • the specific number of stages of the compressor in the compression unit is not limited, as long as it can enable the reaction monomer stream containing the ethylene source to have an inlet pressure into the polymerization unit.
  • those skilled in the art can set a pressure reducing valve on the inlet side of the tubular reactor in the first-stage high-pressure polymerization unit to better regulate the reaction monomer flow containing the ethylene source into the polymerization unit. inlet pressure.
  • the device of the present invention in order to ensure that the reaction monomer stream containing the ethylene source has a feed temperature into the polymerization unit, in some embodiments, the device of the present invention further includes a preheating unit located upstream of the polymerization unit. .
  • At least two-stage compressors refer to at least two compressors connected in series.
  • Each stage compressor is provided with an air inlet and an air outlet, and the upstream compressor has an air inlet and an air outlet.
  • the air outlet is connected to the air outlet of its adjacent downstream compressor; at the same time, the term "compression stage" refers to the connecting pipe between two adjacent compressors.
  • At least one compression unit is disposed in series upstream of at least two parallel tubular reactors in the fluid suction and delivery unit; in some embodiments, at least one compression unit Fluid suction and delivery devices are arranged in series at respective upstreams of the fluid suction and delivery units.
  • the preheating unit includes one or more preheating devices arranged in parallel; the type of the preheating device is not limited as long as it can achieve the purpose of preheating.
  • At least one of the preheating devices is arranged in series upstream of at least two parallel tubular reactors in the first-stage high-pressure polymerization unit; in some embodiments, at least one The preheating equipment is arranged in series at the corresponding upstream of the tubular reactors in the first-stage high-pressure polymerization unit.
  • At least one of the preheating devices is arranged in series upstream of each corresponding tubular reactor in the first-stage high-pressure polymerization unit; in some preferred embodiments, at least one of the The preheating device is located between the corresponding tubular reactor in the compression unit and the first-stage high-pressure polymerization unit; in some preferred embodiments, at least one of the preheating devices is located between the corresponding suction fluid device in the fluid suction delivery unit and a between corresponding tubular reactors in the high-pressure polymerization unit.
  • the reaction monomer flow containing the ethylene source can be divided into at least two tubular reactors in the first-stage high-pressure polymerization unit without passing through the fluid suction and delivery unit. , without affecting production, and at the same time, it can better regulate the flow rate of each reaction monomer stream containing ethylene source entering the first-stage high-pressure polymerization unit.
  • the device of the present invention further includes a chain transfer agent supply unit for delivering the chain transfer agent into the device.
  • the chain transfer agent supply unit can transfer newly generated free radicals to the molecular chain of the polymer, thereby producing polyethylene products with a wider range of molecular weight distribution width and long chain branching degree, and broadening the pipeline. Application fields of formula products.
  • At least one outlet of the chain transfer agent supply unit is connected to the corresponding reaction monomer flow inlet side of the fluid suction and delivery device; in some preferred embodiments, In the method, at least one outlet of the chain transfer agent supply unit is connected to any position on the tubular reactor in the first-stage high-pressure polymerization unit (including the feed port end of the reaction monomer flow of the tubular reactor and The tubular reactor is connected (any position along the material flow direction); in some preferred embodiments, at least one outlet of the chain transfer agent supply unit is connected to the corresponding reaction monomer flow inlet side of the preheating device.
  • At least one outlet of the chain transfer agent supply unit is connected to the common upstream of at least two preheating devices arranged in parallel; in some preferred In the embodiment, at least one outlet of the chain transfer agent supply unit is connected to the common upstream of at least two fluid suction and delivery devices arranged in parallel; in some preferred embodiments, the chain transfer agent supply unit At least one outlet is connected to the inlet side of the multi-stage high-pressure polymerization unit (the side that flows into the inlet end of the material from the first-stage high-pressure polymerization unit); in some preferred embodiments, at least one of the chain transfer agent supply units The discharge port is connected to any position in the multi-stage high-pressure polymerization unit; in some preferred embodiments, at least one discharge port of the chain transfer agent supply unit is connected to the inlet of the compression unit; in some preferred embodiments, the At least one outlet of the chain transfer agent supply unit is connected to any position of the compression interstage connection pipeline of the compression unit.
  • the device of the present invention in order to be able to produce the ethylene copolymer, in some embodiments, the device of the present invention further includes a comonomer supply unit to provide comonomer to the device.
  • the position and number of the outlet of the comonomer supply unit in the device can be selected according to needs.
  • at least one outlet of the comonomer supply unit Connected to the respective corresponding reaction monomer flow inlet ports of the tubular reactors in the first-stage high-pressure polymerization unit; in some embodiments, at least one outlet of the comonomer supply unit is connected to a preheating device The respective corresponding reaction monomer flow inlet ports are connected side by side; in some embodiments, at least one outlet of the comonomer supply unit is connected to the common upstream of at least two preheating devices arranged in parallel; in some implementations In this way, at least one outlet of the comonomer supply unit is connected to the respective corresponding reaction monomer flow inlet side of the fluid suction and delivery device; in some embodiments, at least one outlet of the comonomer supply unit The outlet is connected to the common upstream of at least two fluid suction and delivery devices arranged in parallel;
  • the device of the present invention also includes a separation circulation unit located downstream of the polymerization unit to separate polymer products and unreacted monomers; in some preferred embodiments, the Unreacted monomers can be recycled as circulating materials to the upstream of the compression unit as needed, so that the device of the present invention has better economy.
  • the separation and circulation unit can be selected and configured according to specific needs, and will not be described in detail here.
  • a stream of chain transfer agent 209 is connected to the inlet of the compression unit 201 through an outlet of the chain transfer agent supply unit.
  • the chain transfer agent 209 is supplied and mixed with a stream of material C 208 (which is fresh ethylene) and then compressed by the compression unit 201. Then it is divided into two reactive monomer streams containing ethylene source, which are reactive monomer stream A 210 containing ethylene source and reactive monomer stream B 211 containing ethylene source;
  • the reactive monomer stream A 210 containing the ethylene source is passed through the fluid suction conveying device 202 to obtain the reactive monomer stream C 212 containing the ethylene source.
  • the reactive monomer stream C 212 containing the ethylene source and the reactive monomer stream B 211 containing the ethylene source are obtained.
  • Part of the reaction monomer stream D 213 containing the ethylene source is mixed and then preheated by the preheating device A 3 and then enters the tubular reactor A in the first-stage high-pressure polymerization unit for reaction;
  • the remaining part of the reactive monomer stream B 211 containing the ethylene source is preheated by the preheating device B 204 to obtain the reactive monomer stream E 214 containing the ethylene source.
  • the reactive monomer stream E 214 containing the ethylene source passes through the tubular reactor
  • the pressure reducing valve on the inlet side of B reduces the pressure of the reaction monomer stream E 214 containing the ethylene source to the inlet pressure required by the tubular reactor B and then enters the tubular reactor B in the first-stage high-pressure polymerization unit for reaction;
  • Material A 207 is circulated back to the fluid suction conveying device 202 and mixed with the reaction monomer stream A 210 containing the ethylene source that flows into the fluid suction conveying device 202. After flowing out, it continues to be mixed with a part of the reaction monomer stream B 211 containing the ethylene source.
  • the reaction monomer stream D 213 of the ethylene source is mixed and preheated by the preheating device A 203 and then enters the tubular reactor 205a in the first-stage high-pressure polymerization unit for reaction;
  • Material B 215 flows into the tubular reactor in the multi-stage high-pressure polymerization unit for reaction;
  • the materials flowing out after reaction in the multi-stage high-pressure polymerization unit are separated through the separation and circulation unit to obtain polymer and circulating materials, and the circulating materials are returned to Back to the inlet side of compression unit 201;
  • the first initiator I1 and the second initiator I2 respectively enter the corresponding first reaction zone 205a and the second reaction zone 205b from the first and second outlets of the initiator supply unit to participate in the first-stage high-pressure polymerization unit.
  • reaction in; the third initiator I3, the fourth initiator I4, and the fifth initiator I5 enter the corresponding third reaction zone 206a and the fourth from the third, fourth, and fifth outlet of the initiator supply unit respectively.
  • the reaction zone 206b and the fifth reaction zone 206c participate in the reaction in the multi-stage high-pressure polymerization unit.
  • a stream of material C 208 (which is fresh ethylene) is compressed by the compression unit 201 and divided into two reactive monomer streams containing an ethylene source, which are reactive monomer stream A 210 containing an ethylene source and reactive monomer stream B containing an ethylene source. 211;
  • the reactive monomer stream A 210 containing the ethylene source is passed through the fluid suction conveying device 202 to obtain the reactive monomer stream C 212 containing the ethylene source.
  • the reactive monomer stream C 212 containing the ethylene source and the reactive monomer stream B 211 containing the ethylene source are obtained.
  • Part of the reaction monomer stream D 213 containing the ethylene source is mixed and then preheated by the preheating device 203 and then enters the tubular reactor A in the first-stage high-pressure polymerization unit for reaction;
  • a stream of chain transfer agent 209 is supplied to the inlet side of the reaction monomer flow corresponding to the preheating device B 204 in the preheating unit through an outlet of the chain transfer agent supply unit, and is mixed with the reaction monomer flow B containing the ethylene source.
  • the remaining part in 211 is mixed and preheated by the preheating device B 204 to obtain the reaction monomer flow E 214 containing the ethylene source.
  • the reaction monomer flow E 214 of the ethylene source enters the tubular reactor B in the first-stage high-pressure polymerization unit. react in;
  • Material A 207 is circulated back to the fluid suction conveying device 202 and mixed with the reaction monomer stream A 210 containing the ethylene source that flows into the fluid suction conveying device 202. After flowing out, it continues to be mixed with a part of the reaction monomer stream B 211 containing the ethylene source.
  • the reaction monomer stream D 213 of the ethylene source is mixed and preheated by the preheating device A 203 and then enters the tubular reactor 205a in the first-stage high-pressure polymerization unit for reaction;
  • Material B 215 flows into the tubular reactor in the multi-stage high-pressure polymerization unit for reaction;
  • the materials flowing out after the reaction in the multi-stage high-pressure polymerization unit are separated through the separation and circulation unit to obtain polymer and circulating materials, and the circulating materials are returned to the inlet side of the compression unit 201;
  • the first initiator I1 and the second initiator I2 respectively enter the corresponding first reaction zone 205a and the second reaction zone 205b from the first and second outlets of the initiator supply unit to participate in the first-stage high-pressure polymerization unit.
  • reaction in; the third initiator I3, the fourth initiator I4, and the fifth initiator I5 enter the corresponding third reaction zone 2206a and the fourth from the third, fourth, and fifth outlet of the initiator supply unit respectively.
  • the reaction zone 206b and the fifth reaction zone 206c participate in the reaction in the multi-stage high-pressure polymerization unit.
  • the pressures involved are all absolute pressures.
  • M n number average molecular weight
  • M w weight average molecular weight
  • PDI molecular weight distribution index
  • Ethylene radical polymerization is carried out using an ethylene radical polymerization device as shown in Figure 3.
  • the device includes: a polymerization unit, a fluid suction and delivery unit, an initiator supply unit, a chain transfer agent supply unit, Compression unit 201, preheating unit and separation cycle unit;
  • the fluid suction and delivery unit includes a fluid suction and delivery device 202 (which is a jet pump) for sucking and delivering at least a reaction monomer stream containing an ethylene source and an outlet of at least one tubular reactor in the first-stage high-pressure polymerization unit.
  • a fluid suction and delivery device 202 which is a jet pump
  • the initiator supply unit is used to transport the initiator to the polymerization unit
  • the chain transfer agent supply unit is used to transport the chain transfer agent 209 to the device
  • the compression unit 201 is used to convert each unit of ethylene source-containing
  • the reaction monomer streams each have an inlet pressure entering the first-stage high-pressure polymerization unit
  • the compression unit 201 is located upstream of the fluid suction delivery unit and the polymerization unit;
  • the polymerization unit includes a primary high-pressure polymerization unit and a multi-stage high-pressure polymerization unit connected in series downstream of the primary high-pressure polymerization unit;
  • the primary high-pressure polymerization unit includes two parallel tubular reactors A and tubular reactor B (tubular reactor The length of tubular reactor A and tubular reactor B are both 560m and the inner diameter is 0.045m).
  • the inlet side of tubular reactor B is equipped with a pressure reducing valve, which is used to supply reverse initiator in the multi-stage high-pressure polymerization unit.
  • the equipment required for the unit is three tubular reactors C1, tubular reactor C2 and tubular reactor C3 connected in series (the lengths of tubular reactor C1, tubular reactor C2 and tubular reactor C3 are all 400m, inner diameter is 0.045m);
  • the first-stage high-pressure polymerization unit includes two reaction zones: the first and second outlets of the initiator supply unit are respectively connected with the ethylene source in the tubular reactor A and tubular reactor B in the first-stage high-pressure polymerization unit.
  • the inlet ends of the reaction monomer streams are connected to obtain the first reaction zone 205a and the second reaction zone 205b;
  • the multi-stage high-pressure polymerization unit includes three reaction zones connected in series: the third, fourth, and fifth outlet of the initiator supply unit are respectively connected with the inlet end of the tubular reactor C1 in the multi-stage high-pressure polymerization unit (referring to the inlet into The inlet of the product from the first-stage high-pressure polymerization unit), the inlet end of the tubular reactor C2 (referring to the inlet of the product flowing into the tubular reactor C1), the inlet end of the tubular reactor C3 (referring to the inlet of the product flowing from the tubular reactor C1 (the inlet of the product of vessel C2) is connected to obtain the third reaction zone 206a, the fourth reaction zone 206b, and the fifth reaction zone 206c;
  • the fluid suction and conveying equipment 202 is arranged in series upstream of the tubular reactor A in the first-stage high-pressure polymerization unit;
  • the preheating unit includes two preheating equipment A 203 and preheating equipment A 204 arranged in parallel;
  • the preheating equipment A 203 is located Between the fluid suction and delivery equipment 202 and the corresponding tubular reactor A in the first-stage high-pressure polymerization unit;
  • the preheating equipment A 204 is located between the compression unit 201 and the corresponding tubular reactor B in the first-stage high-pressure polymerization unit;
  • An outlet of the chain transfer agent supply unit is connected to the outlet of the compression unit 201 for transporting the chain transfer agent into the device;
  • the separation circulation unit is located downstream of the polymerization unit to separate polymer products and recycling materials (unreacted monomers);
  • a stream of chain transfer agent 209 (which is propylene, with a flow rate of 0.205t/h) is introduced into the inlet of the compression unit 201 through an outlet of the chain transfer agent supply unit and is combined with a stream of material C 208 (which is fresh ethylene, with a flow rate of 0.205t/h). 43.5t/h) is mixed and compressed and divided into two reactive monomer streams containing ethylene source, namely reactive monomer stream A 210 containing ethylene source (flow rate 16.1t/h) and reactive monomer stream B containing ethylene source. 211;
  • the reactive monomer stream A 210 containing the ethylene source is passed through the fluid suction conveying device 202 (pressure drop is 28 MPa) to obtain the reactive monomer stream C 212 containing the ethylene source.
  • the reactive monomer stream C 212 containing the ethylene source is mixed with the reactive monomer stream C 212 containing the ethylene source.
  • a part of the reaction monomer flow D 213 containing the ethylene source in the reaction monomer flow B 211 is mixed (the flow rate of the mixed reaction monomer flow is 26.72t/h) and is preheated by the preheating device A 203 before entering the first-level high pressure
  • the reaction is carried out in the tubular reactor A in the polymerization unit;
  • the remaining part of the reactive monomer stream B 211 containing the ethylene source is preheated by the preheating device B 204 to obtain the reactive monomer stream E 214 containing the ethylene source.
  • the reactive monomer stream E 214 containing the ethylene source passes through the tubular reactor
  • the pressure reducing valve on the inlet side of B reduces the pressure of the reaction monomer stream E 214 containing the ethylene source to the inlet pressure required by the tubular reactor B and then enters the tubular reactor B in the first-stage high-pressure polymerization unit for reaction;
  • tubular reactor A and tubular reactor B After the materials flowing out of tubular reactor A and tubular reactor B are collected, they are divided into two streams of materials, namely material A 207 and material B 215;
  • Material A 207 is circulated back (circulation ratio is 0.10) to the fluid suction and conveying equipment 202, mixed with the reaction monomer flow A 210 containing the ethylene source that flows into the fluid suction and transportation equipment 202, and then flows out and continues to be mixed with the reaction monomer flow B containing the ethylene source.
  • Part of the reaction monomer stream D 213 containing the ethylene source in 211 is mixed and preheated by the preheating device A 203 and then enters the tubular reactor A in the first-stage high-pressure polymerization unit for reaction;
  • Material B 215 flows into the tubular reactor in the multi-stage high-pressure polymerization unit for reaction;
  • the first initiator and the second initiator respectively enter the corresponding first reaction zone 205a and the second reaction zone 205b from the first and second outlets of the initiator supply unit to participate in the first-stage high-pressure polymerization unit.
  • reaction; the third, fourth, and fifth initiators enter the corresponding third reaction zone 206a, fourth reaction zone 206b, and fifth reaction zone 206c from the third, fourth, and fifth outlet of the initiator supply unit respectively to participate.
  • the inlet pressures of tubular reactor A and tubular reactor B are respectively: the pressure is 220MPa;
  • the flow rates of the first initiator, the second initiator, the third initiator, the fourth and the fifth initiator are 6.72kg/h, 6.80kg/h, 8.34kg/h and 7.44kg/h respectively. ,7.48kg/h;
  • compositions of the first initiator and the second initiator are: di-tert-butyl peroxide, tert-butyl peroxybenzoate, and peroxy-2- mixed in a mass ratio of 1:2:2:2 A mixture of tert-butyl ethylhexanoate and tert-butyl peroxypivalate;
  • compositions of the third initiator, the fourth initiator and the fifth initiator are: di-tert-butyl peroxide, tert-butyl peroxybenzoate, and peroxybenzoate mixed in a mass ratio of 10:2:1:1.
  • the inlet temperature of the first reaction zone 205a is 170°C and the outlet temperature is 194°C; the inlet temperature of the second reaction zone 205b is 170°C and the outlet temperature is 192°C; the third reaction zone 206a, the fourth reaction zone 206b, The temperatures at the inlet of the fifth reaction zone 206c are 193°C, 211°C, and 229°C respectively, and the temperature at the outlet of the fifth reaction zone 206c is 233°C; the first reaction zone 5a, the second reaction zone 5b, the third reaction zone 206a, and the The peak temperatures of the fourth reaction zone 206b and the fifth reaction zone 206c are both 295°C;
  • the materials flowing out after reaction in the multi-stage high-pressure polymerization unit are separated through the separation and circulation unit to obtain polymer and circulating materials, and the circulating materials are returned to the inlet side of the compression unit 201;
  • the materials flowing out after the reaction in the multi-stage high-pressure polymerization unit are separated through the separation and circulation unit to obtain polymer products (low-density polyethylene) and unreacted monomers (i.e., recycled materials).
  • Example 201 the difference is that: the method is carried out in an ethylene free radical polymerization device as shown in Figure 4,
  • An outlet of the chain transfer agent supply unit is connected to the reaction monomer flow inlet side of the tubular reactor B for transporting the chain transfer agent 209 into the device;
  • a stream of material C 208 (fresh ethylene, flow rate 43.5t/h) is compressed by the compression unit 201 and divided into two reaction monomer streams containing an ethylene source, namely the reaction monomer stream A 210 containing an ethylene source and the reaction monomer stream A 210 containing ethylene.
  • the reactive monomer stream A 210 containing the ethylene source is passed through the fluid suction conveying device 202 (pressure drop is 28 MPa) to obtain the reactive monomer stream C 212 containing the ethylene source.
  • the reactive monomer stream C 212 containing the ethylene source is mixed with the reactive monomer stream C 212 containing the ethylene source.
  • a part of the reaction monomer flow D 213 containing the ethylene source in the reaction monomer flow B 211 is mixed (the flow rate of the mixed reaction monomer flow is 26.72t/h) and is preheated by the preheating device 203 before entering the first-stage high-pressure polymerization.
  • the reaction is carried out in the tubular reactor A in the unit;
  • a stream of chain transfer agent 209 (propylene, with a flow rate of 0.205t/h) is introduced into the inlet side of the reaction monomer flow corresponding to the preheating device 204 in the preheating unit through an outlet of the chain transfer agent supply unit, and After being mixed with the remaining part of the reactive monomer stream B 211 containing the ethylene source and preheated by the preheating device 204, the reactive monomer stream E 214 containing the ethylene source is obtained.
  • the reactive monomer stream E 214 containing the ethylene source is passed through the pipe
  • the pressure reducing valve on the inlet side of reactor B reduces the pressure of the reaction monomer stream E 214 containing the ethylene source to the inlet pressure required by the tubular reactor B and then enters the tubular reactor B in the first-stage high-pressure polymerization unit. reaction;
  • the first initiator and the second initiator respectively enter the corresponding first reaction zone 205a and the second reaction zone 205b from the first and second outlets of the initiator supply unit to participate in the first-stage high-pressure polymerization unit.
  • reaction; the third, fourth, and fifth initiators enter the corresponding third reaction zone 206a, fourth reaction zone 206b, and fifth reaction zone 206c from the third, fourth, and fifth outlet of the initiator supply unit respectively to participate.
  • the flow rates of the first initiator, the second initiator, the third initiator, the fourth and the fifth initiators are 6.54kg/h, 6.29kg/h, 8.26kg/h, and 7.67kg respectively. /h, 8.01kg/h;
  • the inlet temperature of the first reaction zone 205a is 170°C and the outlet temperature is 193°C; the inlet temperature of the second reaction zone 205b is 190°C and the outlet temperature is 192°C; the third reaction zone 206a, the fourth reaction zone 206b, The temperatures at the inlet of the fifth reaction zone 206c are 193°C, 211°C, and 225°C respectively, and the temperature at the outlet of the fifth reaction zone 206c is 234°C.
  • One outlet of the chain transfer agent supply unit is not connected to the outlet of the compression unit 201, but is connected to the inlet side of the multi-stage high-pressure polymerization unit (i.e., the inlet side of the tubular reactor C1) for conveying chain transfer to the device. agent 210;
  • a stream of material C 208 (fresh ethylene, with a flow rate of 43.5t/h) is compressed by the compression unit 201 and divided into two reaction monomer streams containing an ethylene source, which are respectively a reaction monomer stream A 210 containing an ethylene source (the flow rate is 21t/h) and reaction monomer stream B 211 containing ethylene source;
  • the reaction monomer flow A 210 containing the ethylene source is passed through the fluid suction conveying device 202 (pressure drop is 30 MPa) to obtain the reaction monomer flow C 212 containing the ethylene source.
  • the reaction monomer flow C 212 containing the ethylene source is mixed with the reaction monomer flow C 212 containing the ethylene source.
  • a part of the reaction monomer flow D 213 containing the ethylene source in the reaction monomer flow B 211 is mixed (the flow rate of the mixed reaction monomer flow is 26.72t/h) and then enters the first-level high pressure after being preheated by the preheating device A 3
  • the reaction is carried out in the tubular reactor A in the polymerization unit;
  • the remaining part of the reactive monomer stream B 211 containing the ethylene source is preheated by the preheating device B 204 to obtain the reactive monomer stream E 214 containing the ethylene source.
  • the reactive monomer stream E 214 containing the ethylene source passes through the tubular reactor
  • the pressure reducing valve on the inlet side of B reduces the pressure of the reaction monomer stream E 214 containing the ethylene source to the inlet pressure required by the tubular reactor B and then enters the tubular reactor B in the first-stage high-pressure polymerization unit for reaction;
  • Material A 207 is circulated back (circulation ratio is 0.15) to the fluid suction and conveying equipment 202, mixed with the reaction monomer flow A 210 containing the ethylene source that flows into the fluid suction and transportation equipment 202, and then flows out and continues to be mixed with the reaction monomer flow B containing the ethylene source.
  • Part of the reaction monomer stream D 213 containing the ethylene source in 211 is mixed and preheated by the preheating device A 203 and then enters the tubular reactor A in the first-stage high-pressure polymerization unit for reaction;
  • tubular reactor A and tubular reactor B After the materials flowing out of tubular reactor A and tubular reactor B are collected, they are divided into two streams of materials, namely material A 207 and material B 215;
  • a stream of chain transfer agent 209 (propylene, flow rate 0.205t/h) is introduced into the inlet side of the multi-stage high-pressure polymerization unit through an outlet of the chain transfer agent supply unit and mixed with material B and then flows into the multi-stage high-pressure polymerization unit.
  • the reaction is carried out in a tubular reactor;
  • the flow rates of the first initiator, the second initiator, the third initiator, the fourth and the fifth initiators are 6.48kg/h, 6.11kg/h, 8.84kg/h, and 7.79kg respectively. /h, 8.11kg/h;
  • the inlet temperature of the first reaction zone 205a is 170°C and the outlet temperature is 196°C; the inlet temperature of the second reaction zone 205b is 185°C and the outlet temperature is 192°C; the third reaction zone 206a, the fourth reaction zone 206b, The temperatures at the inlet of the fifth reaction zone 206c are 194°C, 210°C, and 225°C respectively, and the temperature at the outlet of the fifth reaction zone 206c is 234°C.
  • the inlet pressures of tubular reactor A and tubular reactor B are both 270MPa; the flow rates of the first, second, third, fourth and fifth initiators are 3.76 respectively. kg/h, 3.79kg/h, 4.66kg/h, 4.22kg/h, 4.27kg/h;
  • the inlet temperature of the first reaction zone 205a is 170°C and the outlet temperature is 194°C; the inlet temperature of the second reaction zone 205b is 170°C and the outlet temperature is 192°C; the third reaction zone 206a, the fourth reaction zone 206b, The temperatures at the inlet of the fifth reaction zone 206c are 193°C, 209°C, and 229°C respectively, and the temperature at the outlet of the fifth reaction zone 206c is 233°C.
  • the flow rates of the first initiator, the second initiator, the third initiator, the fourth and the fifth initiator are 7.89kg/h, 7.95kg/h, 9.77kg/h and 8.97kg/h respectively. , 9.41kg/h; the inlet temperature of the first reaction zone 205a is 170°C, and the outlet temperature is 194°C; the inlet temperature of the second reaction zone 205b is 170°C, and the outlet temperature is 192°C; the third reaction zone 206a, The temperatures at the entrances of the fourth reaction zone 206b and the fifth reaction zone 206c are 193°C, 210°C, and 229°C respectively, and the temperatures at the outlet of the fifth reaction zone 206c are 235°C; the first reaction zone 205a, the second reaction zone 205b, and the The peak temperatures of the third reaction zone 206a, the fourth reaction zone 206b, and the fifth reaction zone 206c are all 300°C.
  • the equipment used for the reverse initiator supply unit in the multi-stage high-pressure polymerization unit is a tubular reactor C (the length of the tubular reactor C is 1200m and the inner diameter is 0.045m);
  • the one-stage high-pressure polymerization unit includes two Reaction zone:
  • the first and second outlets of the initiator supply unit are respectively connected to the inlet ends of the reaction monomer flow containing the ethylene source in the tubular reactor A and tubular reactor B in the first-stage high-pressure polymerization unit. , to obtain the first reaction zone 205a and the second reaction zone 205b;
  • the multi-stage high-pressure polymerization unit includes two reaction zones connected in series: the third outlet of the initiator supply unit and the inlet end of the multi-stage high-pressure polymerization unit ((referring to the port that flows into the product from the first-stage high-pressure polymerization unit) Connect to obtain the third reaction zone 206a; the fourth outlet of the initiator supply unit is connected to a position 400m away from the inlet end along the material flow direction in the multi-stage high-pressure polymerization unit to obtain the fourth reaction zone 206b;
  • One stream of initiator and the second stream of initiator respectively enter the corresponding first reaction zone 205a and the second reaction zone 205b from the first and second outlets of the initiator supply unit to participate in the reaction in the first-stage high-pressure polymerization unit;
  • Three and four initiators respectively enter the corresponding third reaction zone 206a and fourth reaction zone 206b from the third and fourth outlet of the initiator supply unit to participate in the reaction in the multi-stage high-pressure polymerization unit;
  • the flow rates of the first initiator, the second initiator, the third initiator and the fourth initiator are 6.75kg/h, 6.81kg/h, 8.33kg/h and 7.46kg/h respectively;
  • the inlet temperature of the first reaction zone 205a is 170°C and the outlet temperature is 194°C; the inlet temperature of the second reaction zone 205b is 170°C and the outlet temperature is 192°C; the entrances of the third reaction zone 206a and the fourth reaction zone 206b The temperatures are 193°C and 211°C respectively, and the temperature at the outlet of the fourth reaction zone 206b is 230°C;
  • Material A 207 is circulated back (circulation ratio is 0.03) to the fluid suction and conveying equipment 202, mixed with the reaction monomer flow A 210 containing the ethylene source that flows into the fluid suction and transportation equipment 202, and then flows out and continues to be mixed with the reaction monomer flow B containing the ethylene source.
  • Part of the reaction monomer stream D 213 containing the ethylene source in 211 is mixed and preheated by the preheating device A 203 and then enters the tubular reactor A in the first-stage high-pressure polymerization unit for reaction;
  • tubular reactor A and tubular reactor B After the materials flowing out of tubular reactor A and tubular reactor B are collected, they are divided into two streams of materials, namely material A 207 and material B 215;
  • the flow rates of the first initiator, the second initiator, the third initiator, the fourth and the fifth initiators are 6.83kg/h, 6.80kg/h, 7.68kg/h and 7.35kg respectively. /h, 7.43kg/h;
  • the inlet temperature of the first reaction zone 205a is 170°C and the outlet temperature is 193°C; the inlet temperature of the second reaction zone 205b is 170°C and the outlet temperature is 192°C; the third reaction zone 206a, the fourth reaction zone 6b, The temperatures at the inlet of the fifth reaction zone 206c are 192°C, 211°C, and 229°C respectively, and the temperature at the outlet of the fifth reaction zone 206c is 233°C.
  • the first-stage high-pressure polymerization unit includes a tubular reactor A (the length of the tubular reactor A is 560m and the inner diameter is 0.045m); the preheating unit includes a preheating unit.
  • Equipment A 203; preheating equipment A 203 is located between the compression unit 201 and the corresponding tubular reactor A in the first-stage high-pressure polymerization unit;
  • the first-stage high-pressure polymerization unit includes a reaction zone: the first outlet of the initiator supply unit is connected to the inlet end of the reaction monomer flow containing the ethylene source in the tubular reactor A in the first-stage high-pressure polymerization unit. , to obtain the first reaction zone 205a;
  • the multi-stage high-pressure polymerization unit includes three reaction zones connected in series: the second, third, and fourth outlet of the initiator supply unit are respectively connected with the inlet end of the tubular reactor C1 in the multi-stage high-pressure polymerization unit (referring to the inlet into the inlet of the product from the first-stage high-pressure polymerization unit), the inlet end of the tubular reactor C2 (referring to the inlet of the product flowing into the tubular reactor C1), the inlet end of the tubular reactor C3 (referring to the inlet of the product from the tubular reactor C1 (the inlet of the product from the tubular reactor C2) is connected to obtain the second reaction zone 206a, the third reaction zone 206b, and the fourth reaction zone 206c;
  • a stream of chain transfer agent 209 (which is propylene, with a flow rate of 0.205t/h) is introduced into the inlet of the compression unit 201 through an outlet of the chain transfer agent supply unit and is combined with a stream of material C 208 (which is fresh ethylene, with a flow rate of 0.205t/h). 43.5t/h) after mixing and compression, it is preheated by preheating equipment A 23 and then sent to tubular reactor A for reaction.
  • the material flowing out of tubular reaction A flows into the tubular reactor in the multi-stage high-pressure polymerization unit for reaction. ;
  • the first initiator enters the corresponding first reaction zone 205a from the first outlet of the initiator supply unit to participate in the reaction in the first-stage high-pressure polymerization unit;
  • the second, third, and fourth initiators are supplied from the initiator respectively.
  • the second, third, and fourth discharge ports of the unit enter the corresponding second reaction zone 206a, third reaction zone 206b, and fourth reaction zone 206c to participate in the reaction in the multi-stage high-pressure polymerization unit;
  • the inlet pressure of tubular reactor A is: pressure 220MPa;
  • the flow rates of the first initiator, the second initiator, the third initiator and the fourth initiator are 14.66kg/h, 7.42kg/h, 7.36kg/h and 7.36kg/h respectively;
  • the temperature at the entrance of the first reaction zone 205a is 170°C and the temperature at the outlet is 194°C; the temperatures at the entrances of the second reaction zone 206a, the third reaction zone 206b and the fourth reaction zone 206c are 194°C, 210°C and 229°C respectively.
  • the temperature at the outlet of the fourth reaction zone 206c is 233°C.
  • the circulation ratio is 0; the flow rates of the first initiator, the second initiator, the third initiator, the fourth and the fifth initiator are 6.80kg/h, 6.80kg/h, and 7.42kg/h respectively. , 7.26kg/h and 7.36kg/h;
  • the entrance temperatures of the first reaction zone 205a and the second reaction zone 205b are both 170°C, and the outlet temperatures are both 192°C; the entrance temperatures of the third reaction zone 206a, the fourth reaction zone 206b, and the fifth reaction zone 206c are respectively 192°C. °C, 211 °C and 229 °C, the temperature at the outlet of the fifth reaction zone 206c is 233 °C.
  • Embodiment 204 Others are the same as Embodiment 204, except that:
  • the first-stage high-pressure polymerization unit includes a tubular reactor A (the length of the tubular reactor A is 560m and the inner diameter is 0.045m); the preheating unit includes a preheating unit.
  • Equipment A 203; preheating equipment A 203 is located between the compression unit 201 and the corresponding tubular reactor A in the first-stage high-pressure polymerization unit;
  • the first-stage high-pressure polymerization unit includes a reaction zone: the first outlet of the initiator supply unit is respectively connected to the inlet end of the reaction monomer flow containing the ethylene source in the tubular reactor A in the first-stage high-pressure polymerization unit. , to obtain the first reaction zone 205a;
  • the multi-stage high-pressure polymerization unit includes three reaction zones connected in series: the second, third, and fourth outlet of the initiator supply unit are respectively connected with the inlet end of the tubular reactor C1 in the multi-stage high-pressure polymerization unit (referring to the inlet into The inlet of the product from the first-stage high-pressure polymerization unit), the inlet end of the tubular reactor C2 (referring to the inlet of the product flowing into the tubular reactor C1), the inlet end of the tubular reactor C3 (referring to the inlet of the product flowing from the tubular reactor C1 (the inlet of the product of vessel C2) is connected to obtain the second reaction zone 206a, the third reaction zone 206b, and the fourth reaction zone 206c;
  • a stream of chain transfer agent 209 (which is propylene, with a flow rate of 0.205t/h) is introduced into the inlet of the compression unit 201 through an outlet of the chain transfer agent supply unit and is combined with a stream of material C 208 (which is fresh ethylene, with a flow rate of 0.205t/h). 43.5t/h) is mixed and compressed and sent to the tubular reactor A for reaction after being preheated by the preheating equipment A 203.
  • the material flowing out of the tubular reaction A flows into the tubular reactor in the multi-stage high-pressure polymerization unit for reaction. ;
  • the first initiator enters the corresponding first reaction zone 205a from the first outlet of the initiator supply unit to participate in the reaction in the first-stage high-pressure polymerization unit;
  • the second, third, and fourth initiators are supplied from the initiator respectively.
  • the second, third, and fourth discharge ports of the unit enter the corresponding second reaction zone 206a, third reaction zone 206b, and fourth reaction zone 206c to participate in the reaction in the multi-stage high-pressure polymerization unit;
  • the inlet pressure of tubular reactor A is: the pressure is 270MPa; the flow rates of the first initiator, second initiator, third initiator and fourth initiator are 8.12kg/h, 4.14 kg/h, 4.12kg/h, 4.18kg/h; the temperature of the inlet of the first reaction zone 5a is 170°C and the temperature of the outlet is 193°C; the second reaction zone 206a, the third reaction zone 206b and the fourth reaction zone 206c The temperatures at the inlets are 193°C, 210°C and 229°C respectively, and the temperature at the outlet of the fourth reaction zone 6c is 233°C.
  • Embodiment 205 Others are the same as Embodiment 205, except that:
  • the first-stage high-pressure polymerization unit includes a tubular reactor A (the length of the tubular reactor A is 560m and the inner diameter is 0.045m); the preheating unit includes a preheating unit.
  • Equipment A 203; preheating equipment A 203 is located between the compression unit 201 and the corresponding tubular reactor A in the first-stage high-pressure polymerization unit;
  • the first-stage high-pressure polymerization unit includes a reaction zone: the first outlet of the initiator supply unit is connected to the inlet end of the reaction monomer flow containing the ethylene source in the tubular reactor A in the first-stage high-pressure polymerization unit. , to obtain the first reaction zone 205a;
  • the multi-stage high-pressure polymerization unit includes three reaction zones connected in series: the second, third, and fourth outlet of the initiator supply unit are respectively connected with the inlet end of the tubular reactor C1 in the multi-stage high-pressure polymerization unit (referring to the inlet into The inlet of the product from the first-stage high-pressure polymerization unit), the inlet end of the tubular reactor C2 (referring to the inlet of the product flowing into the tubular reactor C1), the inlet end of the tubular reactor C3 (referring to the inlet of the product flowing from the tubular reactor C1 (the inlet of the product of vessel C2) is connected to obtain the second reaction zone 206a, the third reaction zone 206b, and the fourth reaction zone 206c;
  • a stream of chain transfer agent 209 (which is propylene, with a flow rate of 0.205t/h) is introduced into the inlet of the compression unit 201 through an outlet of the chain transfer agent supply unit and is combined with a stream of material C 208 (which is fresh ethylene, with a flow rate of 0.205t/h). 43.5t/h) is mixed and compressed and sent to the tubular reactor A for reaction after being preheated by the preheating equipment A 203.
  • the material flowing out of the tubular reaction A flows into the tubular reactor in the multi-stage high-pressure polymerization unit for reaction. ;
  • the first initiator enters the corresponding first reaction zone 205a from the first outlet of the initiator supply unit to participate in the reaction in the first-stage high-pressure polymerization unit; the second, third, and fourth initiators are supplied from the initiator respectively.
  • the second, third, and fourth discharge ports of the unit enter the corresponding second reaction zone 206a, third reaction zone 206b, and fourth reaction zone 206c to participate in the reaction in the multi-stage high-pressure polymerization unit; where: tubular reactor A
  • the inlet pressures are: pressure 220MPa;
  • the flow rates of the first initiator, the second initiator, the third initiator and the fourth initiator are 16.85kg/h, 8.65kg/h, 8.76kg/h and 9.15kg/h respectively;
  • the temperature at the entrance of the first reaction zone 205a is 170°C and the temperature at the outlet is 193°C; the temperatures at the entrances of the second reaction zone 206a, the third reaction zone 206b and the fourth reaction zone 206c are 193°C, 211°C and 230°C respectively.
  • the temperature at the outlet of the fourth reaction zone 206c is 235°C.
  • Embodiment 206 Others are the same as Embodiment 206, except that:
  • the circulation ratio is 0; the flow rates of the first initiator, the second initiator, the third initiator and the fourth initiator are 6.79kg/h, 6.80kg/h, 7.43kg/h and 7.26kg/ respectively. h;
  • the inlet temperature of the first reaction zone 205a is 170°C and the outlet temperature is 192°C; the inlet temperature of the second reaction zone 205b is 170°C and the outlet temperature is 192°C; the entrances of the third reaction zone 206a and the fourth reaction zone 206b The temperatures are 192°C and 211°C respectively, and the temperature at the outlet of the fourth reaction zone 206b is 229°C.

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Abstract

本发明涉及烯烃高压聚合领域，具体涉及一种烯烃自由基聚合方法与烯烃自由基聚合装置。该方法包括：将至少两股含烯烃源的反应单体流分别引入至少两个并联的管式反应器中，各自进行一级高压聚合，再将得到的一级高压聚合产物流入一个或多个依次串联的管式反应器中进行多级高压聚合；其中，将至少一股自由基聚合引发剂各自引入参与一级高压聚合和/或多级高压聚合，所述反应单体流的压力大于等于100MPa。本发明中将含烯烃源的反应单体流引入至少两个并联的管式反应器中，各自进行一级高压聚合，以更好地控制管式反应器进料温度、压力等参数，从而在保证转化率的前提下实现产品的调控。能够显著提高聚乙烯的分子量分布宽度和长链支化度。

Description

烯烃自由基聚合的方法与烯烃自由基聚合装置
相关申请的交叉引用
本申请要求2022年7月1日提交的中国专利申请202210774665.2和202210775378.3的权益，该申请的内容通过引用被合并于本文。
技术领域
本发明涉及烯烃高压聚合领域，具体涉及一种烯烃自由基聚合的方法与烯烃自由基聚合装置。
背景技术
低密度聚乙烯(LDPE)通过高压自由基聚合方法生产，由于管式反应器在聚合过程中较容易按比例放大且经济性合理，管式技术逐步占据主导地位。
现有高压管式工艺流程将反应物料经压缩至200MPa以上，通入预热器中加热到170℃，而后进入反应器引发反应。管式反应器出口物料经过高压分离器和低压分离器分离，其中乙烯、调聚剂、部分低聚物进入高循回路和低循回路，而溶解少量乙烯的LDPE进入挤压机造粒。但是高压管式反应器生产出来的LDPE聚合物通常具有较窄的产品分子量分布(MWD)和较低长链分支度(LCB)，而不同的下游产品对聚乙烯的分子量分布(MWD)和较高长链分支度(LCB)要求不同，例如医用级/食品级LDPE树脂需要较窄分子量分布(MWD)，而生产力学性能优良的重型包装袋、地暖管等产品需要宽分子量分布(MWD)，因此，在一个装置上实现具有不同分子链结构产品的生产将具有更好的经济效益。
目前，本领域现有调整LDPE聚合物产品分子量分布(MWD)和长链分支度(LCB)的方法为对调聚剂的进料位置进行改变，包括二次机入口、二次机级间、二次机出口、预热器、反应器、反应器侧线进料点上游等。但是将调聚剂注入压缩系统会导致过早聚合，在压缩系统中产生结垢，导致生产负荷下降。将调聚剂注入反应器或反应器侧线进料会导致调聚剂与引发剂混合，降低引发剂引发效率，且调聚剂附加流与主流的混合行为可能产生冷点，降低热传递。
因此，研究和开发一种制备低密度聚乙烯的方法具有重要的意义。
发明内容
本发明的目的在于克服现有高压烯烃聚合过程中，不能够在同一装置上实现具有不同分子链结构产品的生产、调聚剂注入反应器或反应器侧线进料会导致调聚剂与引发剂混合，降低引发剂引发效率，且调聚剂附加流与主流的混合行为可能产生冷点，降低热传递的缺陷，如现有高压聚合的高压管式法生产的聚乙烯分子量分布较窄和长支链含量较低、且同一个装置中不能生产出较窄的分子量分布和较低的长支链含量薄膜类聚乙烯产品和支化度更高、分子量更宽的涂覆料类聚乙烯产品的缺陷，提供了一种烯烃自由基聚合的方法与烯烃自由基聚合装置。
本发明第一方面提供了一种烯烃自由基聚合的方法，该方法包括：将至少两股含烯烃源的反应单体流分别引入至少两个并联的管式反应器中，各自进行一级高压聚合，再将得到的一级高压聚合产物流入一个或多个依次串联的管式反应器中进行多级高压聚合；其中，将至少一股自由基聚合引发剂各自引入参与一级高压聚合和/或多级高压聚合，所述反应单体流的压力大于100MPa。
优选地，所述方法包括：将含乙烯源的反应单体流引入至少两个并联的管式反应器中在引发剂存在下进行反应；将至少两个并联的管式反应器中的至少一个管式反应器的出料口的部分物料循环回至少两个并联的管式反应器的至少一个管式反应器中进行反应；将至少两个并联的管式反应器的出料口的剩余物料继续引入一个或多个依次串联的管式反应器中在引发剂的存在下进行反应。
本发明第二方面提供了用于本发明烯烃自由基聚合的方法的装置，该装置包括：
一级高压聚合单元和多级高压聚合单元；其中，
所述一级高压聚合单元串联于多级高压聚合单元的上游；
所述一级高压聚合单元包括至少两个并联的管式反应器，用于将至少两股含烯烃源的反应单体流各自进行一级高压聚合；
所述多级高压聚合单元包括一个或多个依次串联的管式反应器，用于将来自于一级高压聚合单元的产物进行多级高压聚合；
所述一级高压聚合单元和/或多级高压聚合单元中的至少一个管式反应器上设置有引发剂进料口。
优选地，所述装置还包括：流体吸入输送单元，其中，所述流体吸入输送单元包括一个或至少两个并联设置的流体吸入输送设备，用于吸入并输送至少一股含乙烯源的反应单体流和一级高压聚合单元中的至少一个管式反应器的出料口的部分物料；
引发剂供给单元用于向所述一级高压聚合单元和多级高压聚合单元中输送引发剂。
与现有技术相比，本发明至少具有下述有益效果：
(1)本发明所提出的将至少两股含烯烃源的反应单体流分别引入至少两个并联的管式反应器中，各自进行一级高压聚合，其中并联的设置可以更好地控制管式反应器进料温度、压力等参数，从而在保证转化率的前提下实现产品的调控；
(2)本发明所提出的将至少两股含烯烃源的反应单体流分别引入至少两个并联的管式反应器中，各自进行一级高压聚合，能够在不改变管式反应器反应段温度的情况下，更好的控制调整剂如引发剂在管式反应器的浓度分布，从而实现产品数均分子量、分子量分布(MWD)等分子链结构的调整，并利用本发明的方法能够生产得到匹配不同领域的下游产品；
(3)本发明所提出的方法不仅适用于自由基聚合引发剂引发烯烃均聚，且适用于烯烃与其他烯属单体地共聚，从而生产得到多种烯烃均聚、共聚产品，提高用于本发明方法的装置利用率与适用性，具有较好的经济效益。
(4)采用本发明的方法能够提升烯烃原料的利用效率，使得转化率上升，产量提升；
(5)采用本发明的方法用于生产聚乙烯时，能够显著提高聚乙烯的分子量分布宽度和长链支化度，并且能利用本发明的方法中生产出较窄的分子量分布和较低的长支链含量薄膜类聚乙烯产品和支化度更高、分子量更宽的涂覆料类聚乙烯产品，拓宽管式法产品的应用领域，例如能够生产得到分子量分布指数范围为5-16的低密度聚乙烯，分子量分布范围是指本发明中的方法能够生产出的乙烯产品的分子分子量分布范围的宽度；
(6)采用本发明的方法在保证乙烯转化率和低密度聚乙烯产量的同时，还能够制备得到分子量分布宽的聚乙烯产品，具体表现在本发明制备的方法制备得到的低密度聚乙烯的分子量分布与乙烯源的转化率之比大于等于0.018，且小于等于0.048。
附图说明
图1是根据本发明的一种优选的实施方式中烯烃自由基聚合装置；
图2是根据本发明的另一种优选的实施方式中烯烃自由基聚合装置；
图3是根据本发明的一些实施方式中乙烯自由基聚合方法的反应流程图；
图4是根据本发明另一些实施方式中乙烯自由基聚合方法的反应流程图。
附图标记说明
1、循环用料压缩机                         2、一级压缩机
3、二级压缩机                             4/4a/4b、预热器
5a/5b/5c/6a/6b/6c、管式反应器             7、高压释压阀
8、冷却器                                 9、高压分离器
10、高循回路                              11、低压分离器
12、低循回路
201、压缩单元                             202、流体吸入输送设备
203、预热设备A                            204、预热设备B
205a/205b/206a/206b/206c、反应区          207、物料A
208、物料C                                209、链转移剂
210、含乙烯源的反应单体流A                 211、含乙烯源的反应单体流B
212、含乙烯源的反应单体流C                 213、含乙烯源的反应单体流D
214、含乙烯源的反应单体流E                 215、物料B
I1/I2/I3/I4/I5、引发剂
具体实施方式
在本文中所披露的范围的端点和任何值都不限于该精确的范围或值，这些范围或值应当理解为包含接近这些范围或值的值。对于数值范围来说，各个范围的端点值之间、各个范围的端点值和单独的点值之间，以及单独的点值之间可以彼此组合而得到一个或多个新的数值范围，这些数值范围应被视为在本文中具体公开。
在未作相反说明的情况下，使用的方位词如“上游、下游”是对物料在装置中的流动方向而言的。
本发明第一方面提供了一种烯烃自由基聚合的方法，该方法包括：将含烯烃源的反应单体流引入至少两个并联的管式反应器中，各自进行一级高压聚合，再将得到的一级高压聚合产物流入一个或多个依次串联的管式反应器中进行多级高压聚合；其中，将至少一股自由基聚合引发剂各自引入参与一级高压聚合和/或多级高压聚合，所述含烯烃源的反应单体流的压力大于等于100MPa。
使用本发明的方法生产聚烯烃时能够生产得到分子分布更宽、聚合物分散性指数(PDI)更高的聚烯烃产品，发明人推测至少两个并联的管式反应器的一级高压聚合单元的设置能够较好的控制聚合时在高低温下的反应的时间，从而增加聚合物分散性指数(PDI)；同时能够更好的控制在使用本发明的装置时管式反应器进料口的温度、压力等参数，再通过设置自由基聚合引发剂的进料，在保证所述含烯烃源的反应单体流的转化率和不增加装置结垢的前提下实现产品的调控。
在本发明中，可选择将至少一股含烯烃源的物料经过压缩后，和/或至少两股含烯烃源的物料分别经过压缩后使得含烯烃源的物料压缩为压力大于100MPa的反应单体流；其中，含烯烃源的物料的温度没有限制，可以根据需要进行选择。
在本发明中，含烯烃源的物料的股数没有限制，压缩一般是使用压缩单元进行，其中，含烯烃源的物料的股数与压缩单元的数量对应，含烯烃源的物料的股数小于等于反应单体流的股数，当含烯烃源的物料的股数小于反应单体流的股数时，其可以通过压缩单元压缩后再分成所需股数的含烯烃源的反应单体流。例如，将一股含烯烃源的物料经过一个压缩单元压缩至大于等于100MPa后，分成两股含烯烃源的反应单体流。
在本发明中，一级高压聚合和二级高压聚合中的高压聚合条件为反应单体流能够在高压下进行聚合，优选地，所述含烯烃源的反应单体流的压力为110-400MPa(例如110MPa、130MPa、150MPa、170MPa、200MPa、250MPa、300MPa、330MPa、350MPa，以及上述任意数值组成的范围内的任意值)；进一步优选为170-330MPa。应当理解的是，每股所述含烯烃源的反应单体流的压力可以相同，也可以不相同。
在本发明中，本领域技术人员可以理解的是，所述含烯烃源的反应单体流的压力即为所述含烯烃源的反应单体流进入一级高压聚合单元的入口压力，在该压力下进行一级高压聚合。
在本发明中，其一级高压聚合和多级高压聚合均在管式反应器中进行，在管式反应器的长度方向上会存在压降，本发明中将其称之为一级高压聚合前后的压降以及多级高压聚合前后的压降，优选地，一级高压聚合前后的压降与多级高压聚合前后的压降的总和：一级高压聚合前后的压降为3:1-30:1，优选为6:1-8:1。在前述实施方式下，可以降低物料发生偏流。其中，“偏流”指的是不同的并联管式反应器中物料流量之比和根据伯努利方程计算的物料流量之比发生偏离，避免导致局部温度过高的缺陷，在保证转化率的前提下能实现产品分子量分布(MWD)和长链分支量(LCB)等分子链结构的调整。
在本发明中，只要能实现本发明的目的，每个一级高压聚合的先后顺序没有限制，可以同时进行，也可以不同时进行，只要一级高压聚合产物流入一个或多个依次串联的管式反应器中进行多级高压聚合 即可。在一些优选的实施方式中，每个一级高压聚合同时进行。采用前述优选的实施方式，在保证转化率的前提下能实现产品分子量分布(MWD)和长链分支量(LCB)等分子链结构的调整。
根据本发明，优选地，每股所述含烯烃源的反应单体流的温度各自为100-200℃(例如100℃、120℃、150℃、170℃、200℃，以及上述任意数值组成的范围内的任意值)，优选为150-200℃，且每股每个并联的管式反应器的入口的所述含烯烃源的反应单体流的和各自满足关联式：10000≥ρ1/μ1≥1500，优选6000≥ρ1/μ1≥3000；密度ρ1的单位为：kg/m3，粘度μ1的单位为：厘泊(cP)。粘度为25℃下测定。采用前述优选的实施方式，不但能将所述含烯烃源的反应单体流加热到能够引发聚合的温度，同时通过对预热等条件的控制还能够更好的实现产品分子量现产品分子量分布(MWD)和长链支化度(LCB)等分子链结构的调整。
在本发明中，所述一级高压聚合和每个多级高压聚合的温度可以根据需要进行选择，在一些优选的实施方式中，每个一级高压聚合和每个多级高压聚合各自的温度为100-350℃(例如100℃、120℃、125℃、135℃、150℃、164℃、170℃、176℃、180℃、190℃、192℃、203℃、211℃、224℃、225℃、295℃、300℃、320℃、350℃，以及上述任意数值组成的范围内的任意值)。采用前述优选的实施方式，能够在保证转化率的前提下实现产品的分子量分布和支链分布等分子结构调控。
在本发明中，以自由基聚合为主，在反应过程中，一级高压聚合和多级高压聚合过程中，反应温度是在变化的，但是温度的变化均在100-350℃范围内。自由基聚合引发剂的加入会影响聚合的温度，在本发明的一些实施方式中，记录自由基聚合引发剂通过引发剂进料口注入位置的反应器内物料的温度，记为“入口温度”；同时记录引入自由基聚合引发剂的管式反应器中的峰值温度，另外可以理解的是，当管式反应器中没有自由基聚合引发剂引入的管式反应器中时，没有自由基聚合反应的发生，引入该管式反应器中的物流的温度变化不大，在实验过程中不需记录其对应的“入口温度”和“峰值温度”。例如图1、2所示，管式反应器6a、管式反应器6b、管式反应器6c的进料口端没有设置引发剂进料口，即没有自由基聚合引发剂引入管式反应器6a、管式反应器6b、管式反应器6c中，即管式反应器6a、管式反应器6b、管式反应器6c中对应的“入口温度”和“峰值温度”均不需记录。
在本发明中，各股含烯烃源的反应单体流的进料量没有限制，可以根据需要进行选择，在一些优选的实施方式中，各股所述含烯烃源的反应单体流的最大进料量与最小进料量的比值为(20-1)：1，例如20:1、15:1、10:1、5:1、3:1、1:1，以及上述任意数值组成的范围内的任意值；优选为(5-1)：1。比值为重量比。采用前述优选的技术方案，可以实现一级高压聚合的不同管式反应器生产具有不同分子结构特征的聚合物，从而调控最终产品的分子结构。与此同时，能够降低一级高压聚合的管式反应器的设备设计难度。
在本发明中，各股所述含烯烃源的反应单体流的进料量指各股所述含烯烃源的反应单体流流入一级高压聚合单元中的管式反应器的进料量。
在进入至少两个并联的管式反应器中的至少两股含烯烃源的反应单体流具有一定的流速，优选地，每股所述含烯烃源的反应单体流的流速各自大于等于5m/s，且小于等于30m/s，例如为5m/s、6m/s、7m/s、7.24m/s、8m/s、10m/s、11m/s、12m/s、13m/s、14m/s、15m/s、16m/s、17m/s、18m/s、19m/s、20m/s、21m/s、22m/s、23m/s、24m/s、25m/s、26m/s、27m/s、28m/s、29m/s、30m/s，以及上述任意数值组成的范围内的任意值，优选大于等于8m/s，且小于等于20m/s。采用前述实施方式，能够减少并联的管式反应器中的聚合物粘附在反应管内壁，保证反应管的良好的问题，从而提高管式反应器的传热效率和生产效率。能够在保证转化率的前提下实现产品的分子量分布和支链分布等分子结构调控。
在本发明所述的方法中，只要能实现本发明的目的，所述一级高压聚合的个数没有限制，在一些优选的实施方式中，所述一级高压聚合的个数为2-4个。在一定流速的所述含烯烃源的反应单体流下，一级高压聚合的个数越多，需要进行一级高压聚合的反应器的内径就会比较小，这样对设备的要求比较苛刻，使用前述优选的实施方式，不但能够实现产品分子量分布(MWD)和长链分支(LCB)量等分子 链结构的调整，同时对反应设备的要求并不是那么苛刻，但是，并不表示大于4个一级高压聚合不能适用于本发明，根据本发的发明构思，只要大于等于两个的一级高压聚合均能够实现本发明的发明目的。
在本发明中，优选地，至少一股自由基引发剂引入参与一级高压聚合；至少一股引发剂引入参与多级高压聚合。采用前述实施方式，能够在保证含烯烃源的反应单体流的转化率和不增加装置结垢的前提下实现产品的调控。
在本发明中，所述自由基聚合引发剂采用间歇或连续的注入方式引入参与一级高压聚合和/或多级高压聚合。
在本发明中，只要能实现本发明的目的，各股所述自由基聚合引发剂的进料量可根据需要进行选择，本发明中无特别限制。
在本发明中，可通过添加调聚剂来改变制备得到的产品的分子量，在一些实施方式中，所述方法还包括将至少一股调聚剂各自进料参与所述一级高压聚合和多级高压聚合。采用前述实施方式能够达到在不增加压缩机系统结垢与不改变管式反应器反应段温度的情况下，更好地控制调整剂沿管式反应器的浓度分布，从而实现产品分子量分布(MWD)和长链分支量(LCB)等分子链结构的调整，并可以利用该方法在同一个装置中得到匹配不同领域的下游产品。
在本发明中，只要能实现本发明的目的，各股所述调聚剂的进量料可根据需要进行选择，本发明中无特别限制。
在本发明中，可通过添加共聚单体来制备得到烯烃共聚物，在一些实施方式中，所述方法还包括将至少一股共聚单体各自进料参与所述一级高压聚合和多级高压聚合。本发明的方法不但适用于自由基聚合引发剂引发烯烃均聚，且适用于烯烃与共聚单体的共聚，从而生产得到多种烯烃均聚、共聚产品，提高装置利用率与适用性，具有较好的经济效应。
在本发明中，为了得到聚合物产品，在一些实施方式中，将所述多级高压聚合得到的物料经过减压冷却后，分离得到未反应的单体和聚合物产品。
在本发明中，所述烯烃源中的烯烃包括R2C＝CR2型单烯类化合物、共轭双烯烃、非共轭双烯烃中的一种或多种，其中R各自为H、烃基或卤素。例如，所述烯烃可以是碳原子数为1-6的单烯烃或双烯烃，具体的例如可以为乙烯、丙烯、丁烯、异丁烯、1,3-丁二烯、戊二烯、异戊二烯中的一种或多种。
在本发明中，优选地，当烯烃源为乙烯、没有共聚单体时，本发明所述的方法制备得到的产品为线性低密度聚乙烯。
在本发明中，所述共聚单体的种类可以根据需要进行选择，可以理解的是，所述共聚单体与烯烃源的种类不同，能够与烯烃源在高压下自由基共聚的共聚单体均适用本发明体系。在一些实施方式中，当烯烃源为乙烯时，所述共聚单体的实例有α,β-不饱和C3-C8羧酸，特别是丙烯酸、甲基丙烯酸、马来酸和富马酸；和/或α,β不饱和C3-C8羧酸的衍生物，例如，α,β-不饱和C3-C5羧酸酯或α,β-不饱和C3-C5羧酸酐，特别是甲基丙烯酸甲酯、甲基丙烯酸正丁酯、甲基丙烯酸叔丁酯、丙烯酸甲酯、丙烯酸乙酯、丙烯酸正丁酯、丙烯酸叔丁酯、甲基丙烯酸酐和马来酸酐；和/或1-烯烃，例如，丙烯、1-丁烯、1-戊烯、1-己烯、1-辛烯和1-癸烯；优选为乙酸乙烯酯、丙烯、1-己烯、丙烯酸、丙烯酸正丁酯、丙烯酸叔丁酯、丙烯酸-2-乙基己酯、乙酸乙烯酯或丙烯酸乙烯酯中的一种或多种为共聚单体。
在本发明中，在制备烯烃共聚物的情况下，烯烃单体与共聚单体的比例没有限制，可以根据实际需要进行具体选择。
在本发明中，所述自由基聚合引发剂的种类没有限制，在一级高压聚合和/或多级高压聚合中能够产生自由基的任何物质均可作为本发明中的自由基聚合引发剂。在一些实施方式中，所述自由基聚合引发剂包括氧、空气、偶氮化合物、有机过氧化物及C-C引发剂的烃中的一种或多种。可以列举的有机过氧化物有过氧酯、过氧缩酮、过氧酮和过氧碳酸酯，例如过氧二碳酸二(2-乙基己基)酯、过氧二碳酸二环己酯、二乙酰基过氧二碳酸酯、过氧异丙基碳酸叔丁基酯、二叔丁基过氧化物、二叔戊基过氧化物、 二枯基过氧化物、2,5-二甲基-2,5-二叔丁基过氧己烷、叔丁基枯基过氧化物、2,5-二甲基-2,5-二(叔丁基过氧基)己-3-炔、1,3-二异丙基单氢过氧化物或叔丁基氢过氧化物、二癸酰基过氧化物、2,5-二甲基-2,5-二(2-乙基己酰过氧基)己烷、过氧-2-乙基己酸叔戊酯、二苯甲酰过氧化物、过氧化-2-乙基己酸叔丁酯、过氧二乙基乙酸叔丁酯、过氧二乙基异丁酸叔丁酯、过氧-3,5,5-三甲基己酸叔丁酯、1,1-二(叔丁基过氧基)-3,3,5-三甲基环己烷、1,1-二(叔丁基过氧基)环己烷、过氧乙酸叔丁酯、过氧新癸酸枯基酯、过氧新癸酸叔戊酯、过氧新戊酸叔戊酯、过氧新癸酸叔丁酯、过马来酸叔丁酯、过氧新戊酸叔丁酯、过氧异壬酸叔丁酯、过氧化氢二异丙苯、氢过氧化枯烯、过氧苯甲酸叔丁酯、甲基异丁基酮氢过氧化物、3,6,9-三乙基-3,6,9-三甲基三过氧环壬烷和2,2-二(叔丁基过氧基)丁烷等；可以列举的偶氮化合物有：偶氮烷烃(二氮烯)﹑偶氮二羧酸酯、偶氮二羧酸二腈、偶氮双异丁腈等；可以列举的C-C引发剂的烃有1,2-二苯基-1,2-二甲基乙烷衍生物和1,1,2,2-四甲基乙烷衍生物等。本发明的自由基聚合引发剂可单独使用，也可多种不同种类的自由基聚合引发剂混合使用。
在本发明中，自由基聚合引发剂可以以任何状态引入，例如液体、溶解态、超临界状态。优选地，当使用气态自由基聚合引发剂时(例如氧或空气)，将气态自由基聚合引发剂以超临界状态状态引入。
在本发明中，优选地，当引发剂为偶氮化合物、有机过氧化物及C-C引发剂的烃中的一种或多种时，所述自由基聚合引发剂为溶解态自由基聚合引发剂；更优选的，溶解态的自由基聚合引发剂中自由基聚合引发剂的浓度为5-80wt％。
在本发明中，术语“溶解态自由基聚合引发剂”指的是能够溶解自由基聚合引发剂的溶剂与相应自由基聚合引发剂的混合物，其中溶剂的种类没有限制，能够溶解相应的自由基聚合引发剂的溶剂均适用于本发明体系，合适的溶剂的实例有酮、脂族烃(例如辛烷、癸烷、异十二烷等)及其他饱和C8-C25烃等。采用前述优选的实施方式，不但避免了自由基聚合引发剂过热出现热解的现象，使得反应的安全性更高，同时还提高引发剂的效率，降低引发剂的使用成本。
本发明中，只要能够实现本发明的目的，对调聚剂的种类没有限制，能够改变制备产品分子量的调聚剂均可用于本发明体系。在一些实施方式中，所述调聚剂包括脂肪烃、烯烃、酮、醛、脂族醇或氢的其中一种或多种。可以列举的脂肪烃有丙烷、丁烷、戊烷、己烷、环己烷等；可以列举的烯烃有丙烯、1-戊烯或1-己烯；可以列举的酮有丙酮、甲基乙基酮(2-丁酮)、甲基异丁基酮、甲基异戊基酮、二乙基酮、二戊基酮等；可以列举的醛有甲醛、乙醛或丙醛；可以列举的脂族醇有甲醇、乙醇、丙醇、异丙醇、丁醇等。优选的，所述调聚剂为脂族醛(例如丙醛)、1-烯烃(例如丙烯或1-己烯)及脂肪烃(例如，丙烷)中的一种或多种。
本发明中，涉及的压力均为绝压。
如图1-2所示，本发明第二方面提供了用于本发明烯烃自由基聚合的方法的装置，该装置包括：
一级高压聚合单元和多级高压聚合单元；其中，
所述一级高压聚合单元串联于多级高压聚合单元的上游；
所述一级高压聚合单元包括至少两个并联的管式反应器，用于将至少两股含烯烃源的反应单体流各自进行一级高压聚合；
所述多级高压聚合单元包括一个或多个依次串联的管式反应器，用于将来自于一级高压聚合单元的产物进行多级高压聚合；
所述一级高压聚合单元和/或多级高压聚合单元中的至少一个管式反应器上设置有引发剂进料口。
在本发明所述的装置中，通过包括至少两个并联的管式反应器的一级高压聚合单元和包括一个或多个依次串联的管式反应器的多级高压聚合单元的设置，与现有的管式反应器生产聚烯烃的技术相比，使用本发明的装置生产聚烯烃时能够生产得到分子分布更宽、聚合物分散性指数(PDI)更高的聚烯烃产品，发明人推测至少两个并联的管式反应器的一级高压聚合单元的设置能够较好的控制聚合时在高低温下的反应的时间，从而增加聚合物分散性指数(PDI)；同时能够更好的控制在使用本发明的装置时管 式反应器进料口的温度、压力等参数，再通过设置不同位置的引发剂进料口，使在保证含烯烃源的反应单体流的转化率和不增加装置结垢的前提下实现产品的调控。
在本发明所述的装置中，引发剂进料口的位置没有限制，可以根据需要进行选择，在一些实施方式中，所述一级高压聚合单元中的至少一个管式反应器的反应单体流进料口端设置有引发剂进料口；在一些实施方式中，所述一级高压聚合单元中的至少一个管式反应器上沿其长度方向设置有至少一个引发剂进料口(例如1个、2个、3个、4个等)，优选为1-3个(指一级高压聚合单元中的任意一个或多个管式反应器上沿其长度方向各自设置的1-3个引发剂进料口)；在一些实施方式中，所述多级高压聚合单元中的至少一个管式反应器上设置有至少一个引发剂进料口(例如1个、2个、3个、4个等，其中，“管式反应器上”包括管式反应器的物料进料口端和管式反应器上沿其长度方向的任意位置，多级高压聚合单元中管式反应器的“物料进料口端”指接收一级高压聚合单元物料的入口端和接收相邻管式反应器中流出的物料的入口端)，优选为1-5个(指多级高压聚合单元中的任意一个或多个管式反应器上各自设置的2-5个引发剂进料口)。采用前述实施方式，使本发明的装置能够较好的实现产品分子量分布(MWD)和长链分支量(LCB)等分子链结构的调整，并可以利用本发明的装置生产得到匹配不同领域的下游产品。
其中，“反应单体流进料口端”指的是一级高压聚合单元中的管式反应器一端的反应单体流进料口，本领域的技术人员应当理解的是，其管式反应器的另一端为出料口；多级高压聚合单元中的管式反应器的入口端指多级高压聚合单元中的第一个管式反应器中流入一级高压聚合单元的产物的一端。
在本发明所述的装置中，只要能实现本发明的目的，所述一级高压聚合单元中并联的管式反应器个数没有限制，在一些优选的实施方式中，所述一级高压聚合单元包括2-4个并联的管式反应器，使用前述优选的实施方式，能够实现产品分子量分布(MWD)和长链分支量(LCB)等分子链结构的调整。
在本发明所述的装置中，为了使每股含烯烃源的反应单体流各自具备进入所述一级高压聚合单元的入口压力，在一些实施方式中，所述装置还包括位于所述一级高压聚合单元的上游的至少一个压缩单元。
在本发明所述的装置中，为了使每股含烯烃源的反应单体流各自具备进入所述一级高压聚合单元的入口温度，在一些实施方式中，所述装置还包括位于所述一级高压聚合单元的上游的至少一个预热器；优选地，所述预热器位于所述压缩单元与一级高压聚合单元之间。采用前述实施方式，能够使得一级高压聚合更顺利的发生。
在本发明所述的装置中，压缩单元的个数没有限制，只要能够使每股含烯烃源的反应单体流各自具备进入所述一级高压聚合单元的入口压力即可，所述压缩单元包括一个(或多个依次串联的，例如2个依次串联的、3个依次串联的、4个依次串联的、5个依次串联的)压缩机，优选地，所述压缩单元包括2-4个依次串联的压缩机。在一些实施方式中，所述压缩单元包括2个依次串联的压缩机，依次为一级压缩机2和二级压缩机3；在另一些实施方式中，所述压缩单元包括3个依次串联的压缩机，依次为一级压缩机2、二级压缩机3和三级压缩机。采用前述实施方式，能够使每股反应单体流达到一级高压聚合单元的入口压力。
在本发明所述的装置中，在一些优选的实施方式中，至少一个压缩单元串联设置于一级高压聚合单元中的至少两个并联的管式反应器共同的上游；在一些实施方式中，至少一个压缩单元串联设置于一级高压聚合单元中的管式反应器各自对应的上游。
在本发明所述的装置中，所述预热器的个数没有限制，只要能够使每股含烯烃源的反应单体流各自具备进入所述一级高压聚合单元的入口温度即可，在一些优选的实施方式中，至少一个预热器串联设置于一级高压聚合单元中的至少两个并联的管式反应器共同的上游；在一些实施方式中，至少一个预热器串联设置于一级高压聚合单元中的管式反应器各自对应的上游。
在本发明所述的装置中，优选地，所述装置还包括至少一个调聚剂进料口(例如1个、2个、3个、4个、5个、6个等)。调聚剂进料口可用于将调聚剂引入所述装置中参与一级高压聚合和/或多级高压 聚合，能够更好的调节产品的分子量。
在本发明所述的装置中，所述调聚剂进料口的个数和位置没有限制，可以根据需要进行选择，在一些优选的实施方式中，所述调聚剂进料口设置于多级高压聚合单元出口上游的任意位置，即可以根据产品的需要，选择在多级高压聚合单元中的最后一个管式反应器出料口的上游的任意位置设置调聚剂进料口。
在本发明所述的装置中，每个所述调聚剂进料口的设置可以有多种方式。优选地，每个所述调聚剂进料口各自设置于：压缩单元的入口；和/或，压缩单元的出口；和/或，压缩单元中任意两个相邻压缩机的连接管道处；和/或，一级高压聚合单元中的至少一个管式反应器的反应单体流进料口侧(指一级高压聚合单元中的至少一个管式反应器和其对应上游的压缩单元的连接管道处)；和/或，一级高压聚合单元中的至少一个管式反应器上(包括管式反应器的反应单体流进料口端和管式反应器上沿其长度方向的任意位置)；和/或，一级高压聚合单元和多级高压聚合单元的连接管道处；和/或，多级高压聚合单元中的至少一个管式反应器上(包括管式反应器的进料口端和管式反应器上沿其长度方向的任意位置)。采用前述各种实施方式，能够实现根据需要得到生产不同分子量宽度的聚烯烃。
在本发明所述的装置中，优选地，所述装置还包括至少一个共聚单体进料口(例如1个、2个、3个、4个、5个、6个等)。共聚单体进料口的设置可以用于将共聚单体引入装置中参与一级高压聚合和/或多级高压聚合，以利用该装置生产聚烯烃共聚物。
在本发明所述的装置中，所述共聚单体进料口的个数和位置没有限制，可以根据需要进行选择，在一些优选的实施方式中，所述共聚单体进料口设置于一级高压聚合单元出口上游的任意位置，即可以根据产品的需要，选择在一级高压聚合单元出口的至少一个管式反应器出料口的上游的任意位置设置共聚单体进料口。
在本发明所述的装置中，每个所述共聚单体进料口的设置可以有多种方式。优选地，每个所述共聚单体进料口各自设置于：压缩单元的入口；和/或，压缩单元中任意两个相邻压缩机的连接管道处；和/或，与一级高压聚合单元中的至少一个管式反应器的反应单体流进料口侧；和/或，预热器与压缩单元的连接管道处。
在本发明所述的装置中，优选地，所述装置还包括位于多级高压聚合单元下游的分离循环单元，用于分离多级高压聚合得到的物料，以得到聚合产品和未反应完的单体。采用前述实施方式不但能够得到聚合产品，还能够将未反应的单体循环到一级高压聚合和多级高压聚合中。
在本发明所述的装置中，只要能实现本发明的目的，所述分离循环单元的结构没有限制，在一些优选的实施方式中，所述分离循环单元包括分离器和循环回路，所述分离器用于将来自于多级高压聚合单元的产物进行分离，所述循环回路用于将未反应的单体到循环到一级高压聚合单元上游；进一步优选地，所述分离器包括依次串联的高压分离器9和低压分离器11，所述循环回路包括与高压分离器9连接的高循回路10和与低压分离器11连接的低循回路12，更优选的，所述高循回路10一端连接高压分离器9，另一端连接二级压缩机3的吸入侧；所述低压分离器11一端连接低压分离器11，另一端连接一级压缩机2的吸入侧。
在本发明所述的装置中，优选地，所述分离循环单元还包括第二段反应器与低压分离器11之间的高压释压阀7与冷却器8，用于将来自于多级高压聚合单元的物料减压、冷却后再进入分离循环单元中进行分离循环。在本发明所述的装置中，优选地，所述高循回路10上还包括冷却器和分离器，用于将来自高压分离器9中的气态馏分除去一些不是未反应的单体的组分(例如低聚物)。在本发明所述的装置中，优选地，所述低循回路12上还包括冷却器和分离器，用于将来自低压分离器11中的气态馏分除去一些不是未反应的单体的组分(例如低聚物)；进一步优选地，所述低循回路12上还包括至少一个循环料用压缩机1(例如1个循环料用压缩机1、2个循环料用压缩机1、3个循环料用压缩机1等)，更进一步优选地，所述循环料用压缩机1设于低循回路12上的冷却器和分离器的下游。采用前述实施 方式，不但能够得到聚合物产品，还能更好的实现未反应的单体的循环利用，具有更好的经济效应。
在本发明中，具体地，使用本发明的装置将本发明的方法中所述多级高压聚合得到的产物依次经过所述分离循环单元中的高压释压阀7与冷却器后流入所述分离循环单元中的高压分离器中分离得到气态馏分A和液态馏分A，液态馏分A流入所述分离循环单元中的低压分离器11中分离得到气态馏分B和聚合物产品；进一步优选地，气态馏分A经过高循回路10上的冷却器和分离器冷却、分离后得到未反应的单体流入二级压缩机3的吸入侧；和/或气态馏分B经过低循回路12上的冷却器和分离器冷却、分离得到未反应的单体经过低循回路12上的压缩机压缩后流入一级压缩机2的吸入侧；和/或聚合物送去造粒得到相应的产品。
在本发明中，只要能实现本发明的目的，所述高压分离器、低压分离器11、高循回路10和低循回路12上冷却器和分离器的冷却分离条件没有限定，本领域的技术人员可以根据需要进行选择。
在本发明所述的装置中，对压缩机、管式反应器、高压释压阀7、冷却器、高压分离器9、低压分离器11、分离器的具体结构没有特别的限定，可以分别为本领域常用的各种压缩机、管式反应器、高压释压阀7、冷却器、高压分离器、低压分离器11，此均为本领域技术人员所熟知，在此不再赘述，同时本发明中各个物料入口或物料出口可以交叉连接，各个设备之间通过管道实现连接、装置中还设置有阀门等部件来实现物料的流通、管式反应器选择冷却夹套结构来实现换热，本发明对此无特殊要求，在此不多加赘述。
在本发明中，本领域的技术人员可以根据本发明中不同实施方式所需的装置的变化来调整本发明的方法，在此不多加赘述。
以下将通过实施例对本发明进行详细描述。以下实施例中：数均分子量Mn、重均分子量Mw和聚合物分散性指数PDI通过高温凝胶渗透色谱HT-GPC方法测得。
实施例1
采用如图1所示的装置进行烯烃自由基聚合，其中，在图1的装置中：串联于多级高压聚合单元的上游的一级高压聚合单元中的两个并联的管式反应器5a和管式反应器5b的长度均为560m、内径均为0.045m；多级高压聚合单元中的三个依次串联的管式反应器6a、管式反应器6b、管式反应器6c的长度均为400m，内径均为0.045m；管式反应器5a和管式反应器5b的反应单体流进料口端均设置有引发剂进料口；管式反应器6a、管式反应器6b、管式反应器6c的进料口端各设置有一个引发剂进料口；压缩单元的出口设置有一个调聚剂进料口；
一股含烯烃源的物料依次经过压缩单元中的一级压缩机2和二级压缩机压缩3，一股调聚剂通过调聚剂进料口进料到压缩单元的出口处与压缩后含烯烃源的物料充分混合后等量分成两股含烯烃源的反应单体流，两股含烯烃源的反应单体流分别经过位于管式反应器5a上游的预热器4a和管式反应器5b上游的预热器4b后引入装置包括的一级高压聚合单元中的两个并联的管式反应器5a和管式反应器5b中，各自进行一级高压聚合，再将得到的一级高压聚合产物流入装置包括的多级高压聚合单元中的三个管式反应器6a、管式反应6b、管式反应6c中进行多级高压聚合；其中，五股自由基聚合引发剂分别通过装置包括的引发剂进料口各自引入参与相应的一级高压聚合或多级高压聚合；多级高压聚合产物依次经过分离循环单元中的高压释压阀7与冷却器8后流入分离循环单元中的高压分离器9中分离得到气态馏分A和液态馏分A，液态馏分A流入分离循环单元中的低压分离器11中分离得到气态馏分B和聚合物产品；气态馏分A经过高循回路10上的冷却器和分离器冷却、分离后得到未反应的单体流入二级压缩机的吸入侧；气态馏分B经过低循回路12上的冷却器和分离器冷却、分离后得到未反应的单体经过低循回路上的循环用料压缩机1压缩后流入一级压缩机的吸入侧；聚合物产品送去造粒得到的低密度聚乙烯产品；其中：
含烯烃源的物料为乙烯；两股反应单体流分别进入管式反应器5a和管式反应器5b的入口压力均为：220MPa；
两股反应单体流分别进入管式反应器5a和管式反应器5b的入口温度分别为：170℃和180℃；
调聚剂为丙烯；调聚剂的进料量为：250kg/h；
预热器4a和预热器4b的入口温度均为：92℃；
管式反应器5a和管式反应器5b反应单体流进料口端的反应单体流的进料量均为21750kg/h；反应单体流的流速分别为7.24m/s和7.31m/s，密度分别为527kg/m3和522kg/m3；管式反应器5a和管式反应器5b的反应单体流的密度：粘度的值分别为5383和5398；
沿物料流动方向，第一股自由基聚合引发剂和第二股自由基聚合引发剂的组成分别为：按照质量比为1:2:2:2混合的二叔丁基过氧化物、过氧苯甲酸叔丁酯、过氧化-2-乙基己酸叔丁酯和过氧新戊酸叔丁酯的混合物；第三至第五股自由基聚合引发剂的组成分别为：按照质量比为10:2:1:1混合的二叔丁基过氧化物、过氧苯甲酸叔丁酯、过氧化-2-乙基己酸叔丁酯和过氧新戊酸叔丁酯的混合物；
管式反应器5a与管式反应器5b的反应单体流进料口端的引发剂进料口的自由基聚合引发剂进料量分别为：6.78kg/h、6.19kg/h；
管式反应器6a、管式反应器6b、管式反应器6c的进料口端的引发剂进料口的自由基聚合引发剂进料量分别为7.33kg/h、7.44kg/h、7.80kg/h；
管式反应器6a、管式反应器6b、管式反应器6c的入口温度分别为192℃、211℃、225℃；
管式反应器5a中、管式反应器5b中及管式反应器6a、管式反应器6b和管式反应器6c中的峰值温度均为295℃；
高压分离器的分离条件为：25MPa，235℃；
低压分离器的分离条件为：2bar，220℃；
一级高压聚合单元的压降为4.8MPa，多级高压聚合单元的压降为30.6MPa，一级高压聚合单元的压降与多级高压聚合单元的压降的总和：一级高压聚合单元的压降为7.38:1。
如表1所示，为该实施例中低密度聚乙烯的数均分子量Mn、重均分子量Mw、分子量分布指数PDI、产量和乙烯转化率。
实施例2
按照实施例1的方法，不同之处在于：
采用如图2所示的装置进行烯烃自由基聚合，在图2的装置中，一个预热器4位于一级高压聚合单元中的两个并联的管式反应器5a与管式反应器5b共同的上游，一股含烯烃源的物料依次经过压缩单元中的一级压缩机2和二级压缩机压缩3，一股调聚剂通过调聚剂进料口进料到压缩单元的出口处与压缩后含烯烃源的物料充分混合后经过预热器4后等量分成两股含烯烃源的反应单体流分别引入装置包括的一级高压聚合单元中的两个并联的管式反应器5a和管式反应器5b中；其中：
两股反应单体流分别进入管式反应器5a和管式反应器5b的入口温度均为：170℃；反应单体流的流速均为7.24m/s，密度均为527kg/m3；管式反应器5a和管式反应器5b的反应单体流的密度：粘度的值均为5383；
调聚剂为丙烯；调聚剂的进料量为：180kg/h；
预热器4的入口温度为：92℃；
管式反应器5a与管式反应器5b的反应单体流进料口端的引发剂进料口的自由基聚合引发剂进料量分别为：6.80kg/h、6.80kg/h；
管式反应器6a、管式反应器6b、管式反应器6c的进料口端的引发剂进料口的自由基聚合引发剂进料量分别为7.41kg/h、7.26kg/h、7.36kg/h；
管式反应器6a、管式反应器6b、管式反应器6c的入口温度分别为192℃，211℃，229℃；
一级高压聚合单元的压降为5.2MPa，多级高压聚合单元的压降为33.5MPa，一级高压聚合单元的压降与多级高压聚合单元的压降的总和：一级高压聚合单元的压降为7.44:1。
如表1所示，为该实施例中低密度聚乙烯的数均分子量Mn、重均分子量Mw、分子量分布指数PDI、产量和乙烯转化率。
实施例3
按照实施例1的方法，不同之处在于：
在图1的装置中，所述管式反应器6a、管式反应器6b、管式反应器6c的进料口端没有设置引发剂进料口，两股自由基聚合引发剂分别通过所述装置包括的引发剂进料口各自引入参与相应的一级高压聚合；其中：
调聚剂为丙烯；调聚剂的进料量为：195kg/h；
管式反应器5a与管式反应器5b的反应单体流进料口端的引发剂进料口的自由基聚合引发剂进料量分别为：6.47kg/h、6.22kg/h；
所述管式反应器5a、管式反应器5b中的峰值温度均为295℃；
高压分离器的分离条件为：23MPa，200℃
低压分离器的分离条件为：2bar，190℃；
一级高压聚合单元的压降为5.0MPa，多级高压聚合单元的压降为28.1MPa，一级高压聚合单元的压降与多级高压聚合单元的压降的总和：一级高压聚合单元的压降为6.62:1。
如表1所示，为该实施例中低密度聚乙烯的数均分子量Mn、重均分子量Mw、分子量分布指数PDI、产量和乙烯转化率。
实施例4
按照实施例1的方法，不同之处在于：
在图1的装置中：所述管式反应器5a的反应单体流进料口侧设置有一个调聚剂进料口，其他位置没有设置调聚剂进料口；一股调聚剂通过调聚剂进料口进料与流入管式反应器5a的反应单体流混合后引入管式反应器5a中进行相应的一级高压聚合；其中：
两股反应单体流分别进入管式反应器5a和管式反应器5b的入口温度分别为：170℃和190℃；反应单体流的流速分别为7.24m/s和7.37m/s，密度分别为527kg/m3和518kg/m3；管式反应器5a和管式反应器5b的反应单体流的密度：粘度的值分别为5383和5424；
调聚剂为丙烯；调聚剂的进料量为：170kg/h；
管式反应器5a与管式反应器5b的反应单体流进料口端的引发剂进料口的自由基聚合引发剂进料量分别为：6.71kg/h、6.25kg/h；
管式反应器6a、管式反应器6b、管式反应器6c的进料口端的的引发剂进料口的自由基聚合引发剂进料量分别为7.32kg/h、7.42kg/h、7.78kg/h；
管式反应器6a、管式反应器6b、管式反应器6c的入口温度分别为192℃、211℃、224℃；
一级高压聚合单元的压降为5.1MPa，多级高压聚合单元的压降为33.8MPa，一级高压聚合单元的压降与多级高压聚合单元的压降的总和：一级高压聚合单元的压降为7.63:1。
如表1所示，为该实施例中低密度聚乙烯的数均分子量Mn、重均分子量Mw、分子量分布指数PDI、产量和乙烯转化率。
实施例5
按照实施例4的方法，不同之处在于：
在图1的装置中：管式反应器5a和管式反应器5b的反应单体流进料口侧分别设置有一个调聚剂进料口，装置的其他位置未设置调聚剂进料口；两股调聚剂分别通过调聚剂进料口进料到管式反应器5a的反应单体流进料口侧和管式反应器5b的反应单体流进料口侧后，分别与两股反应单体流充分混合后流入管式反应器5a和管式反应器5b中各自进行一级高压聚合；其中：
调聚剂为丙烯；两股调聚剂的进料量均为：97.5kg/h；
管式反应器5a与管式反应器5b的反应单体流进料口端引发剂进料口的自由基聚合引发剂进料量分别为：6.78kg/h、6.19kg/h；
管式反应器6a、管式反应器6b、管式反应器6c的入口温度分别为192℃、211℃、225℃；
一级高压聚合单元的压降为5.0MPa，多级高压聚合单元的压降为32.9MPa，一级高压聚合单元的压降与多级高压聚合单元的压降的总和：一级高压聚合单元的压降为7.58:1。
如表1所示，为该实施例中低密度聚乙烯的数均分子量Mn、重均分子量Mw、分子量分布指数PDI、产量和乙烯转化率。
实施例6
按照实施例4的方法，不同之处在于：
在图1的装置中：一级高压聚合单元和多级高压聚合单元的连接管道处设置有一个调聚剂进料口，装置的其他位置未设置调聚剂进料口，一股调聚剂通过调聚剂进料口进料到一级高压聚合单元和多级高压聚合单元的连接管道处与一级高压聚合产物充分混合后流入多级高压聚合单元；其中：
两股反应单体流分别进入管式反应器5a和管式反应器5b的入口温度均为：170℃和150℃；反应单体流的流速分别为7.24m/s和7.37m/s，密度分别为527kg/m3和537kg/m3；管式反应器5a和管式反应器5b的反应单体流的密度：粘度的值分别为5383和5317；
调聚剂为丙烯；调聚剂的进料量为：195kg/h；
管式反应器5a与管式反应器5b的反应单体流进料口端的引发剂进料口的自由基聚合引发剂进料量分别为：6.72kg/h、2.40kg/h；
管式反应器6a、管式反应器6b、管式反应器6c的进料口端的引发剂进料口的自由基聚合引发剂进料量分别为7.35kg/h、7.42kg/h、7.81kg/h；
管式反应器6a、管式反应器6b、管式反应器6c的入口温度分别为192℃、211℃、224℃；
管式反应器5b的温峰为250℃，其他管式反应器的温峰为295℃；
一级高压聚合单元的压降为5.6MPa，多级高压聚合单元的压降为34.5MPa，一级高压聚合单元的压降与多级高压聚合单元的压降的总和：一级高压聚合单元的压降为7.16:1。
如表1所示，为该实施例中低密度聚乙烯的数均分子量Mn、重均分子量Mw、分子量分布指数PDI、产量和乙烯转化率。
实施例7
按照实施例1的方法，不同之处在于：
采用如图1所示的装置进行烯烃自由基聚合，其中，在图1的装置中：串联于多级高压聚合单元的上游的一级高压聚合单元中的两个并联的管式反应器5a和管式反应器5b的长度均为560m、内径均为0.030m；
管式反应器5a和管式反应器5b反应单体流进料口端的反应单体流的进料量均为21750kg/h；反应单体流的流速分别为16.29m/s和16.45m/s；
管式反应器5a与管式反应器5b的反应单体流进料口端的引发剂进料口的自由基聚合引发剂进料量分别为：6.78kg/h、6.19kg/h；
管式反应器6a、管式反应器6b、管式反应器6c的进料口端的引发剂进料口的自由基聚合引发剂进料量分别为8.00kg/h、5.05kg/h、8.49kg/h；
管式反应器6a、管式反应器6b、管式反应器6c的入口温度分别为201℃、235℃和215℃；
一级高压聚合单元的压降为21.8MPa，多级高压聚合单元的压降为30.3MPa，一级高压聚合单元的压降与多级高压聚合单元的压降的总和：一级高压聚合单元的压降为2.39:1。
如表1所示，为该实施例中低密度聚乙烯的数均分子量Mn、重均分子量Mw、分子量分布指数PDI、产量和乙烯转化率。
实施例8
按照实施例1的方法，不同的是：
采用如图1所示的装置进行烯烃自由基聚合，其中，在图1的装置中：串联于多级高压聚合单元的上游的一级高压聚合单元中的两个并联的管式反应器5a和管式反应器5b的长度分别为560m和325m、内径分别为0.045m和0.024m；管式反应器5a和管式反应器5b反应单体流进料口端的反应单体流的进料量分别为37285kg/h和6215kg/h；反应单体流的流速分别为12.41m/s和7.47m/s；
管式反应器5a与管式反应器5b的反应单体流进料口端的引发剂进料口的自由基聚合引发剂进料量分别为：10.62kg/h、2.45kg/h；
管式反应器6a、管式反应器6b、管式反应器6c的进料口端的引发剂进料口的自由基聚合引发剂进料量分别为7.98kg/h、4.92kg/h、8.32kg/h；
管式反应器6a、管式反应器6b、管式反应器6c的入口温度分别为199℃、236℃、215℃；
一级高压聚合单元的压降为10.2MPa，多级高压聚合单元的压降为30.5MPa，一级高压聚合单元的压降与多级高压聚合单元的压降的总和：一级高压聚合单元的压降为3.99:1。
如表1所示，为该实施例中低密度聚乙烯的数均分子量Mn、重均分子量Mw、分子量分布指数PDI、产量和乙烯转化率。
对比例1
按照实施例1的方法，不同之处在于：
在图1的装置中，串联于多级高压聚合单元的上游的一级高压聚合单元为一个管式反应器5a(管式反应器5a的长度为560m、内径均为0.045m)，装置中无管式反应器5b和预热器4b；装置中设置有四个引发剂进料口，分别设置于管式反应器5a的反应单体流进料口端、管式反应器6a、管式反应器6b、管式反应器6c的进料口端；装置的其他位置未设置引发剂进料口；
一股含烯烃源的物料依次经过压缩单元中的一级压缩机2和二级压缩机压缩3，一股调聚剂通过调聚剂进料口进料到压缩单元的出口处与压缩后含烯烃源的物料混合后经预热器4a后得到一股含烯烃源的反应单体流，四股自由基聚合引发剂分别通过装置包括的引发剂进料口各自引入参与相应的一级高压聚合或多级高压聚合；其中：
反应单体流进入管式反应器5a的入口压力为：220MPa
反应单体流进入管式反应器5a的入口温度为：175℃；
管式反应器5a单体流进料口端的反应单体流的进料量为43500kg/h；
管式反应器5a的反应单体流进料口端、管式反应器6a、管式反应器6b、管式反应器6c的进料口端的引发剂进料口的自由基聚合引发剂进料量分别为：13.37kg/h、6.56kg/h、7.71kg/h和7.39kg/h；
管式反应器6a、管式反应器6b、管式反应器6c的入口温度分别为：206℃、207℃、228℃；
管式反应器5a、管式反应器6a、管式反应器6b、管式反应器6c中的峰值温度均为295℃；
管式反应器的总压降为44.5MPa。
如表1所示，为该对比例中低密度聚乙烯的数均分子量Mn、重均分子量Mw、分子量分布指数PDI、产量和乙烯转化率。
对比例2
按照实施例2的方法，不同之处在于，
在图2的装置中，串联于多级高压聚合单元的上游的一级高压聚合单元为一个管式反应器5a(管式反应器5a的长度为560m、内径均为0.045m)，装置中无管式反应器5b；装置中设置有四个引发剂进料口，分别设置于管式反应器5a的反应单体流进料口端、管式反应器6a、管式反应器6b、管式反应器6c的进料口端；装置的其他位置未设置引发剂进料口；
一股含烯烃源的物料依次经过压缩单元中的一级压缩机2和二级压缩机压缩3，一股调聚剂通过调 聚剂进料口进料到压缩单元的出口处与压缩后含烯烃源的物料混合后经预热器4a后得到一股含烯烃源的反应单体流，四股自由基聚合引发剂分别通过装置包括的引发剂进料口各自引入参与相应的一级高压聚合或多级高压聚合；其中：
反应单体流进入管式反应器5a的入口压力为：220MPa；
反应单体流进入管式反应器5a的入口温度为：175℃；
管式反应器5a单体流进料口端的反应单体流的进料量为43500kg/h；
管式反应器5a的反应单体流进料口端、管式反应器6a、管式反应器6b、管式反应器6c的进料口端的引发剂进料口的自由基聚合引发剂进料量分别为：13.34kg/h、6.54kg/h、7.36kg/h和6.82kg/h；
管式反应器6a、管式反应器6b、管式反应器6c的入口温度分别为206℃、207℃、232℃；
管式反应器5a、管式反应器6a、管式反应器6b、管式反应器6c中的峰值温度均为295℃；
管式反应器的总压降为46.1MPa。
如表1所示，为该对比例中低密度聚乙烯的数均分子量Mn、重均分子量Mw、分子量分布指数PDI、产量和乙烯转化率。
对比例3
按照实施例3的方法，不同之处在于：
在图1的装置中，串联于多级高压聚合单元的上游的一级高压聚合单元为一个管式反应器5a(管式反应器5a的长度为560m、内径均为0.045m)，装置中无管式反应器5b和预热器4b；装置中设置有一个引发剂进料口，设置于管式反应器5a的反应单体流进料口端；装置的其他位置未设置引发剂进料口；
一股含烯烃源的物料依次经过压缩单元中的一级压缩机2和二级压缩机压缩3，一股调聚剂通过调聚剂进料口进料到压缩单元的出口处与压缩后含烯烃源的物料混合后经预热器4a后得到一股含烯烃源的反应单体流，一股自由基聚合引发剂分别通过装置包括的引发剂进料口各自引入参与相应的一级高压聚合；其中：
反应单体流进入管式反应器5a的入口压力为：220MPa；
反应单体流进入管式反应器5a的入口温度为：175℃；
管式反应器5a的反应单体流进料口端的反应单体流的进料量为43500kg/h；
管式反应器5a的反应单体流进料口端的引发剂进料口的自由基聚合引发剂进料量为：12.70kg/h；
管式反应器5a的峰值温度为295℃；
管式反应器的总压降为41.2MPa。
如表1所示，为该对比例中低密度聚乙烯的数均分子量Mn、重均分子量Mw、分子量分布指数PDI、产量和乙烯转化率。
对比例4
按照实施例2的方法，不同之处在于：
在图2的装置中，使用全长1760m，内径0.045m的管反应器替代一级高压聚合单元和多级高压聚合单元，管式反应器的反应单体流进料口端、管式反应器上沿其长度方向的560m、960m、1360m分别设置有一个引发剂进料口；含烯烃源的反应单体流引入管式反应器中进行高压聚合，自由基聚合引发剂通过引发剂进料口连续进料至装置中参与高压聚合(高压聚合得到的产物相当于实施例1中多级高压聚合得到的产物)；
反应单体流进入管式反应器的入口压力均为：220MPa；
反应单体流进入管式反应器的入口温度均为：175℃；
调聚剂为丙烯；调聚剂的进料量为：170kg/h；
管式反应器的反应单体流进料口端的反应单体流的进料量为：43500kg/h；
管式反应器的反应单体流进料口端、管式反应器上沿其长度方向560m、960m、1260m处四个引发剂进料口的自由基聚合引发剂进料量分别为14.29kg/h、7.40kg/h、7.27kg/h、7.32kg/h；
管式反应器的反应单体流进料口端、管式反应器上沿其长度方向560m、960m、1260m处的入口温度分别为193℃、211℃、229℃；
管式反应器的反应单体流进料口端、管式反应器上沿其长度方向560m、960m、1260m处分成的四个区的峰值温度分别均为295℃；
管式反应器的总压降为47.6MPa。
如表1所示，为该对比例中低密度聚乙烯的数均分子量Mn、重均分子量Mw、分子量分布指数PDI、产量和乙烯转化率。
对比例5
按照实施例6的方法，不同之处在于：
在图1的装置中，串联于多级高压聚合单元的上游的一级高压聚合单元为一个管式反应器5a(管式反应器5a的长度为560m、内径均为0.045m)，装置中无管式反应器5b和预热器4b；装置中设置有四个引发剂进料口，分别设置于管式反应器5a的反应单体流进料口端、管式反应器6a、管式反应器6b、管式反应器6c的进料口端；装置的其他位置未设置引发剂进料口；管式反应器5a和多级高压聚合单元的连接管道处设置有一个调聚剂进料口，装置的其他位置未设置调聚剂进料口；一股调聚剂通过调聚剂进料口进料到一级高压聚合单元中的管式反应器5a和多级高压聚合单元的连接管道处与一级高压聚合产物充分混合后流入多级高压聚合单元；其中：
反应单体流进入管式反应器5a的入口压力为：220MPa；
反应单体流进入管式反应器5a的入口温度均为：175℃；
管式反应器5a单体流进料口端的反应单体流的进料量为43500kg/h；
管式反应器5a的反应单体流进料口端、管式反应器6a、管式反应器6b、管式反应器6c的进料口端的引发剂进料口的自由基聚合引发剂进料量分别为：13.19kg/h、6.58kg/h、7.68kg/h和7.38kg/h；
管式反应器6a、管式反应器6b、管式反应器6c的入口温度分别为205℃、208℃、228℃；
管式反应器5a、管式反应器6a、管式反应器6b、管式反应器6c中的峰值温度均为295℃；
管式反应器的总压降为48.3MPa。
如表1所示，为该对比例中低密度聚乙烯的数均分子量Mn、重均分子量Mw、分子量分布指数PDI、产量和乙烯转化率。
表1

通过表1的结果可以看出，本发明实施例1-8中的装置可以根据下游产品性能的需求，控制一级高压聚合单元中至少两个并联的管式反应器的入口温度与反应器温峰、调聚剂进料位置等参数，从而在保证转化率一定的情况下，控制产品的分子量分布(MWD)；并且，在相同峰值温度下，通过实施例1-8与对比例1-5的对比，采用本发明的装置能够生产得到分子量分布更宽的产品，克服了管式法不能生产釜式法产品的缺点，另一方面本发明的装置乙烯转化率更高。此外，本发明所述的方法和装置具有控制方便，操作容易，产品覆盖范围广的效果。
本发明的一些优选实施方式还提供一种乙烯自由基聚合的方法与装置。
本发明还提供一种乙烯自由基聚合的方法，该方法包括：将含乙烯源的反应单体流引入至少两个并联的管式反应器中在引发剂存在下进行反应；将至少两个并联的管式反应器中的至少一个管式反应器的出料口的部分物料循环回至少两个并联的管式反应器的至少一个管式反应器中进行反应；将至少两个并联的管式反应器的出料口的剩余物料继续流入一个或多个依次串联的管式反应器中在引发剂的存在下进行反应；所述含乙烯源的反应单体流的压力大于等于100MPa。
如前述，釜式反应器能够生产支化度更高，分子量更宽的聚乙烯产品，管式反应器生产的聚乙烯产品具有较窄的分子量分布和较低的长支链含量，发明人发现，与现有技术相比，利用本发明所述的方法能够生产得到现有管式反应器不能生产的支化度高、分子量更宽的聚乙烯产品，发明人推测本发明的方法中至少两股含乙烯源的反应单体流分别引入至少两个并联的管式反应器中在引发剂存在下进行反应，同时将至少两个并联的管式反应器中的至少一个管式反应器的出料口的部分物料循环回至少两个并联的管式反应器的至少一个管式反应器中进行反应，能够使得新生成的自由基向聚合物的分子链上转移，从而生产得到分子量分布宽度和长链支化度范围更宽的聚乙烯产品，拓宽了管式法产品的应用领域，同时还能够提升原料的利用效率，使得转化率上升，产量提升。
根据本发明的方法，在一些优选的实施方式中，所述至少两个并联的管式反应器中的出料口的物料的循环比小于1，优选为小于等于0.3，更优选为0.04-0.2，例如为0.04、0.05、0.07、0.09、0.1、0.12、0.13、0.15、0.18、0.2，以及以上任意两个数值组成的范围内的任意值。采用前述优选的实施方式能够得到不同分子量分布的聚乙烯产品，尤其是能够得到现有技术中管式法不能生产得到的宽分子分布量的聚乙烯产品，同时在保证还能保持较好的乙烯转化率和产量。
在本发明所述的方法中，术语“循环比”指管式反应器的出料口的循环回至少两个并联的管式反应器的至少一个管式反应器中的物料总质量与所有至少两个并联的管式反应器所有的出料口的物料总质量的质量比；同时，至少一个管式反应器的出料口的循环回至少两个并联的管式反应器的至少一个管式反应器中的物料可以根据需要选择至少两个并联的管式反应器中的一个管式反应器的出料口的部分物料或多个管式反应器出料口的部分物料汇集后的总物料，循环回至少两个并联的管式反应器的至少一个管式反应器中；也可以选择多个出料口的部分物料分别循环回至少两个并联的管式反应器的多个管式反应器中，还可以根据需要以其他方式循环回至少两个并联的管式反应器的至少一个管式反应器中。
在本发明所述的方法中，至少两个并联的管式反应器中的至少一个管式反应器的出料口的部分物料循环回至少两个并联的管式反应器的至少一个管式反应器中进行反应时，可以理解的是，需要借助一些设备实现循环，例如吸入流体吸入输送设备，流体吸入输送设备可能会存在压降损失，一般为压降小于60MPa，优选小于30MPa，进一步优选小于15MPa(例如压降为15MPa、12MPa、8MPa、6MPa，以及上述任意数值组成的范围内的任意值)。
根据本发明的方法，流入至少两个并联的管式反应器的所述含乙烯源的反应单体流的压力能使含乙烯源的反应单体流在引发剂的条件能发生高压自由基聚合即可，在一些优选的实施方式中，流入至少两个并联的管式反应器的所述含乙烯源的反应单体流的压力为140-300MPa(例如140MPa、160MPa、220MPa、300MPa，以及上述任意数值组成的范围)。采用前述实施方式，能够根据需要控制反应时链增长速率的速率和链转移的难易程度，从而更好的控制乙烯与乙烯分子之间的间距，得到不同密度和支链分布的产品，以使得本发明的方法能够生产重包装膜料、农膜用料、注塑料、涂覆料以及医用包装料等产品。
根据本发明的方法，可通过管式反应器入口侧的减压阀，调节所述含乙烯源的反应单体流进入聚合单元中的入口压力。
在本发明所述的方法中，可以理解的是，通过流体吸入输送设备引入相应的管式反应器中的含乙烯源的反应单体流的入口压力等于通过具有一定压力的含乙烯源的反应单体流的压力(即进入流体吸入输送设备前的压力)减去流体吸入输送设备存在的压降；另外，可以理解的是，未经过流体吸入输送设备引入相应的管式反应器中的含乙烯源的反应单体的入口压力等于具有一定压力的含乙烯源的反应单体流的压力或是通过管式反应器入口侧的减压阀降低含乙烯源的反应单体流的入口压力。
根据本发明的方法，至少两个并联的管式反应器，以及一个或多个依次串联的管式反应器中进行反应的反应温度没有限制，只要能发生高压自由基聚合即可，在一些实施方式中，所述一级高压聚合单元和多级高压聚合单元中的反应温度各自为100-350℃(例如100℃、130℃、150℃、170℃、200℃、250℃、295℃、300℃、350℃，以及上述任意数值组成的范围内的任意值)。
在本发明所述的方法中，可以理解的是，至少两个并联的管式反应器，以及一个或多个依次串联的管式反应器中的反应温度为一个范围内的温度，具体包括，至少两个并联的管式反应器，以及一个或多个依次串联的管式反应器中的反应区入口的温度、出口的温度、峰值温度均在该范围反应温度内。
根据本发明的方法，所述至少两个并联的管式反应器，和一个或多个依次串联的管式反应器中：各自的反应区出口的温度小于等于相应反应区的峰值温度；所述至少两个并联的管式反应器，和一个或多个依次串联的管式反应器中：各自的反应区入口的温度小于相应反应区的峰值温度；进一步优选地，所述至少两个并联的管式反应器，和一个或多个依次串联的管式反应器中：各自的反应区出口的温度和相应的反应区的峰值温度差值的绝对值为0-150℃(例如20℃、50℃、70℃、100℃、120℃、150℃，以及上述任意数值组成的范围内的任意值)，更进一步优选为20-120℃。采用前述实施方式，能够较好的控制反应区中温度的分布，同时采用前述实施方式能够得到不同分子量分布的聚乙烯产品。
在本发明所述的方法，流入至少两个并联的管式反应器中的所述含乙烯源的反应单体流的温度可以由预热单元预热后得到。在一些优选的实施方式中，流入至少两个并联的管式反应器中的反应单体流的温度为140-190℃(例如140℃、150℃、170℃、180℃、190℃，以及上述任意数值组成的范围内的任意值)；在一些优选的实施方式中，所述至少两个并联的管式反应器中，以及一个或多个依次串联的管式反应器中各自的反应区的峰值温度为200-300℃(例如200℃、225℃、260℃、295℃、300℃，以及上述任意数值组成的范围内的任意值)；采用前述实施方式，能够较好的控制反应区中温度的分布，同时采用前述实施方式能够得到不同分子量分布的聚乙烯产品。
根据本发明的方法，在一些优选的实施方式中，所述至少两个并联的管式反应器中的所述含乙烯源的反应单体流的最大进料量与最小进料量的重量比为1：(0.01-1)，例如1:0.1、1：0.5、1：1，以及上述任意数值组成的范围内的任意值。采用前述优选的实施方式，能够根据需要得到较宽范围的分子分布量的不同产品。
根据本发明的方法，在一些优选的实施方式中，至少两股引发剂各自参与所述至少两个并联的管式反应器；至少一股引发剂参与一个或多个依次串联的管式反应器中的反应。采用前述实施方式，能够更好地使得新生成的自由基向聚合物的分子链上转移，从而生产得到分子量分布宽度和长链支化度范围更 宽的聚乙烯产品，拓宽了管式法产品的应用领域，同时还能够提升原料的利用效率，使得转化率上升，产量提升。
根据本发明的方法，在一些实施方式中，本发明所述的方法还包括至少一股链转移剂参与所述至少两个并联的管式反应器，和一个或多个依次串联的管式反应器中的反应。采用前述实施方式，配合链转移剂够使新生成的自由基向聚合物的分子链上转移，从而生产得到分子量分布宽度和长链支化度范围更宽的聚乙烯产品，拓宽了管式法产品的应用领域，同时还能够提升原料的利用效率，使得转化率上升，产量提升。
根据本发明的方法，在一些实施方式中，本发明所述的方法还包括至少一股共聚单体参与至少两个并联的管式反应器，和一个或多个依次串联的管式反应器中的反应。采用前述实施方式中，本发明的方法还能够得到不同的乙烯共聚物。
在本发明所述的方法中，可以根据本发明的需要从用于装置本发明所述的方法包括的共聚单体供给单元的出料口引入装置的不同位置。
根据本发明的方法，在一些优选的实施方式中，本发明所述的方法还包括一个或多个依次串联的管式反应器中得到的物料分离得到聚合产品和未反应单体；为了使本发明具有较好的经济效应，进一步优选的，所述未反应单体循环回至少两个并联的管式反应器上游继续进行反应。
在本发明所述的方法中，其循环分离的条件可以根据需要进行选择，对本发明的目的并不会产生影响，在此不多加赘述。
根据本发明的方法，所述引发剂的具体选择没有限制，在本发明的反应条件下能产生自由基物质的任何引发剂均适用于本发明。在一些实施方式中，所述引发剂选自偶氮化合物、有机过氧化物、氧气和空气中的一种或多种，同时，本发明中的引发剂可以直接使用也可以将其溶解在溶剂中使用。可以列举的偶氮化合物偶氮二异丁腈、偶氮二异戊腈和偶氮二异庚腈；可以列举的有机过氧化物有2,2-双(叔丁基过氧基)丙烷、过氧化-2-乙基己酸叔丁酯、过氧化二碳酸双-(2-乙基己酯)、二叔丁基过氧化物、二烷基过氧化物、过氧苯甲酸叔丁酯、1,1-双(叔丁基过氧基)环己烷、过氧新戊酸叔丁酯。
在本发明中，引发剂可以以任何状态引入，例如液体、溶解态或超临界状态。例如，当使用氧或空气等气态自由基聚合引发剂时，引发剂以超临界状态状态引入。
本发明中以二叔丁基过氧化物、过氧苯甲酸叔丁酯、过氧化-2-乙基己酸叔丁酯和过氧新戊酸叔丁酯的混合物作为示例性说明本发明的优势，但本发明不局限于此。
根据本发明的方法，所述链转移剂的种类没有限制，在一些实施方式中，所述链转移剂选自肪脂烃、烯烃、酮、醛、脂族醇和氢中的一种或多种。可以列举的肪脂烃有丙烷、丁烷和环己烷；可以列举的烯烃有丙烯和己烯。
在本发明所述的方法中，能够在高压自由基聚合的条件下与乙烯发生共聚的单体均能够作为本发明的共聚单体，例如：丙烯和/或乙酸乙烯酯。
在本发明所述的方法中，利用本发明所述的方法得到的低密度聚乙烯的分子量分布与乙烯源的转化率之比大于等于0.01，且小于等于0.05，优选大于等于0.018，且小于等于0.048。
本发明还提供了一种乙烯自由基聚合的装置，该装置包括：聚合单元、流体吸入输送单元、引发剂供给单元；所述聚合单元包括一级高压聚合单元和串联于一级高压聚合单元下游的多级高压聚合单元；所述一级高压聚合单元包括至少两个并联的管式反应器；所述多级高压聚合单元包括一个或多个依次串联的管式反应器；所述流体吸入输送单元包括一个或至少两个并联设置的流体吸入输送设备，用于吸入并输送至少一股含乙烯源的反应单体流和一级高压聚合单元中的至少一个管式反应器的出料口的部分物料；所述引发剂供给单元用于向所述聚合单元中输送引发剂。
在本发明所述的装置中，可以理解的是，所述引发剂供给单元通过引发剂供给单元的出料口向装置中供给引发剂，其引发剂的出料口在本发明所述装置中的位置和引发剂供给单元的出料口的个数可以根 据所需反应区的位置和个数进行选择，同时，本领域的技术人员可以理解的是，当所述引发剂供给单元向所述聚合单元中输送引发剂时，引发剂供给单元中的一个出料口与装置中的任意一个连接处为“反应区入口”，沿物料的流向方向，有一段温度升高的区域，称为“反应区”；当温度升高至峰值温度，开始降低的一段区域为相应的“冷却区”；当某个单元在物料的流向方向上只有一个反应区时，物料的流向方向的出料口为“反应区出口”或“冷却区出口”；当某个单元有多个依次串联的反应区时，沿物料的流向方向，引发剂供给单元中的后一个出料口、且位于相邻的前一个“反应区入口”下游位置的连接处为后一个“反应区入口”，同时为前一反应区的“反应区出口”或“冷却区出口”，依次类推，引发剂供给单元中的最后一个“反应区出口”或“冷却区出口”为该反应结束后的出料口；例如，本发明装置中的多级高压聚合单元中包括2个依次串联的反应区，即表示，引发剂供给单元中的第一个出料口与多级高压聚合单元中的第一个连接处为“第一反应区入口”，沿物料的流向方向有一段温度升高的区域，称为“第一反应区”，当温度升高至峰值温度，开始降低的一段区域为相应的“第一冷却区”；引发剂供给单元中的第二个出料口与沿物料的流向方向、且位于“第一反应区入口”下游的多级高压聚合单元中的第二个连接处为“第二反应区入口”也为“第一反应区出口”或“第一冷却区出口”；从“第二反应区入口”处沿物料的流向方向有一段温度升高的区域，称为“第二反应区”，当温度升高至峰值温度，温度开始降低的一段区域为相应的“第二冷却区”，多级单元反应结束后的出料口为“第二反应区出口”或“第二冷却区出口”。
在本发明的装置中，一级高压聚合单元中的管式反应器、多级高压聚合单元中的管式反应器、流体吸入输送设备和引发剂供给单元等可根据物料引入物料和排出物料的需求设置一个或多个进料口和出料口，在此不多加赘述；所述一级高压聚合单元和多级高压聚合单元中的管式反应器没有限制，优选为套管结构的管式反应器，可以为单套管结构，也可以为多套管结构，采用套管结构的反应器可以较好的控制本发明的装置在使用时一级高压聚合单元和多级高压聚合单元中的反应温度。
如前述，釜式反应器能够生产支化度更高，分子量更宽的聚乙烯产品，管式反应器生产的聚乙烯产品具有较窄的分子量分布和较低的长支链含量，发明人发现，与现有技术相比，利用本发明所述的装置能够生产得到现有管式反应器不能生产的支化度高、分子量更宽的聚乙烯产品，发明人推测，本发明装置中的至少两个并联的管式反应器、与用于吸入并输送至少一股含乙烯源的反应单体流和一级高压聚合单元中的至少一个管式反应器的出料口的部分物料的流体吸入输送单元的装置中进行生产产品时，能够使新生成的自由基向聚合物的分子链上转移，从而生产得到分子量分布宽度和长链支化度范围更宽的聚乙烯产品，拓宽了管式法产品的应用领域，同时还能够提升原料的利用效率，使得转化率上升，产量提升。
根据本发明的装置，只要能实现本发明的目的，所述一级高压聚合单元包括的管式反应器的个数没有限定，在一些优选的实施方式中，所述一级高压聚合单元包括2-4个并联的管式反应器。采用前述优选的实施方式，能够更灵活的调控聚乙烯产品的分子链结构，有利于开发更多的聚乙烯新产品，实现柔性生产。
根据本发明的装置，只要能实现本发明的目的，所述多级高压聚合单元中包括的反应区的个数没有限制，在一些优选的实施方式中，所述多级高压聚合单元中包括1个或2-6个依次串联的反应区；进一步优选为2-4个依次串联的反应区。采用前述优选的实施方式，能够根据需要有效的调整聚乙烯产品分子量的分布。
在本发明的装置中，如无特殊说明，所述流体吸入输送设备主要的功能是用于吸入并输送物料，设备的具体选择没有限制，包括但不限于射流泵；流体吸入输送设备的个数可以根据需要进行设置，同时，本领域的技术人员应当理解的是，流体吸入输送设备的个数小于等于引入一级高压聚合单元中的管式反应器中的含乙烯源的反应单体流的股数，一级高压聚合单元中的管式反应器的个数大于等于含乙烯源的反应单体流的股数。
根据本发明的装置，在一些实施方式中，至少一个所述流体吸入输送设备串联设置于所述一级高压聚合单元中的至少两个并联的管式反应器共同的上游；在一些实施方式中，至少一个所述流体吸入输送设备串联设置于所述一级高压聚合单元中的至少两个并联的管式反应器共同的上游。
根据本发明的装置，为了使含乙烯源的反应单体流具备进入聚合单元中的入口压力，在一些实施方式中，本发明所述的装置还包括位于所述流体吸入输送单元与聚合单元上游的压缩单元；优选地，所述压缩单元包括至少2级压缩机。
根据本发明的装置，压缩单元中的压缩机的具体级数没有限定，只要其能使含乙烯源的反应单体流具备进入聚合单元中的入口压力即可。
根据本发明所述的装置，本领域的技术人员可在一级高压聚合单元中的管式反应器的入口侧设置减压阀，以更好的调节含乙烯源的反应单体流进入聚合单元中的入口压力。
根据本发明的装置，为了使含乙烯源的反应单体流具备进入聚合单元中的进料温度，在一些实施方式中，本发明所述的装置还包括位于所述聚合单元上游的预热单元。
在本发明的装置中，本领域技术人员可以理解的是，至少2级压缩机指至少两个依次串联的压缩机，每一级压缩机均设有进气口和出气口，上游压缩机的出气口与其相邻的下游压缩机的出气口连通；同时术语“压缩级间”指两个相邻压缩机的连接管道间。
根据本发明的装置，在一些实施方式中，至少一个压缩单元串联设置于所述流体吸入输送单元中的至少两个并联的管式反应器共同的上游；在一些实施方式中，至少一个压缩单元串联设置于所述流体吸入输送单元中的流体吸入输送设备各自对应的上游。
根据本发明的装置，在一些实施方式中，所述预热单元包括一个或多个并联设置的预热设备；所述预热设备的种类没有限制，只要能实现预热的目的即可。
根据本发明的装置，在一些实施方式中，至少一个所述预热设备串联设置于一级高压聚合单元中的至少两个并联的管式反应器共同的上游；在一些实施方式中，至少一个所述预热设备串联设置于一级高压聚合单元中的管式反应器各自对应的上游。
根据本发明的装置，在一些实施方式中，至少一个所述预热设备串联设置于一级高压聚合单元中的管式反应器各自对应的上游；在一些优选的实施方式中，至少一个所述预热设备位于压缩单元和一级高压聚合单元中对应的管式反应器之间；在一些优选的实施方式中，至少一个所述预热设备位于流体吸入输送单元中对应的吸入流体设备与一级高压聚合单元中对应的管式反应器之间。采用前述实施方式，当装置中的流体吸入输送单元出现故障时，含乙烯源的反应单体流可装置可不经过流体吸入输送单元分成进入一级高压聚合单元中的至少两个管式反应器中，不影响生产，同时还能够更好的调控进入一级高压聚合单元中的各股含乙烯源的反应单体流的流量。
在本发明所述的装置中，可以理解的是，如果含乙烯源的反应单体流自身已具备进入聚合单元中的进料温度，可以不需要设置预热单元。
根据本发明的装置在一些优选的实施方式中，本发明所述的装置还包括链转移剂供给单元，用于向装置中输送链转移剂。采用前述实施方式，配合链转移剂供给单元能够使新生成的自由基向聚合物的分子链上转移，从而生产得到分子量分布宽度和长链支化度范围更宽的聚乙烯产品，拓宽了管式法产品的应用领域。
根据本发明的装置，在一些优选的实施方式中，所述链转移剂供给单元的至少一个出料口与流体吸入输送设备各自对应的反应单体流进料口侧连接；在一些优选的实施方式中，所述链转移剂供给单元的至少一个出料口与所述一级高压聚合单元中的管式反应器上任意位置(包括管式反应器的反应单体流的进料口端和管式反应器沿物料流动方向的任意位置)连接；在一些优选的实施方式中，所述链转移剂供给单元的至少一个出料口与预热设备各自对应的反应单体流进料口侧连接；在一些优选的实施方式中，所述链转移剂供给单元的至少一个出料口与至少两个并联设置的预热设备的共同上游连接；在一些优选 的实施方式中，所述链转移剂供给单元的至少一个出料口与至少两个并联设置的流体吸入输送设备的共同上游连接；在一些优选的实施方式中，所述链转移剂供给单元的至少一个出料口与多级高压聚合单元的入口侧(流入来自一级高压聚合单元中的物料入口端的一侧)连接；在一些优选的实施方式中，所述链转移剂供给单元的至少一个出料口与多级高压聚合单元中任意位置连接；在一些优选的实施方式中，所述链转移剂供给单元的至少一个出料口与压缩单元入口连接；在一些优选的实施方式中，所述链转移剂供给单元的至少一个出料口与压缩单元的压缩级间连接管路的任意位置连接。采用前述优选的实施方式，能够调控一级高压聚合单元出料口的初始量的聚合物的分子量分布，从而更好的调节最终产品的分子量分布。
根据本发明的装置，为了能够生产乙烯共聚物，在一些实施方式中，本发明所述的装置还包括共聚单体供给单元，以向装置中提供共聚单体。
根据本发明的装置，所述共聚单体供给单元的出料口在装置中的位置和个数可以根据需要进行选择，在一些实施方式中，所述共聚单体供给单元的至少一个出料口与所述一级高压聚合单元中的管式反应器各自对应的反应单体流进料口侧连接；在一些实施方式中，所述共聚单体供给单元的至少一个出料口与预热设备各自对应的反应单体流进料口侧连接；在一些实施方式中，所述共聚单体供给单元的至少一个出料口与至少两个并联设置的预热设备的共同上游连接；在一些实施方式中，所述共聚单体供给单元的至少一个出料口与流体吸入输送设备各自对应的反应单体流进料口侧连接；在一些实施方式中，所述共聚单体供给单元的至少一个出料口与至少两个并联设置的流体吸入输送设备的共同上游连接；在一些实施方式中，所述共聚单体供给单元的至少一个出料口与压缩单元入口连接；在一些实施方式中，所述共聚单体供给单元的至少一个出料口与压缩单元的压缩级间连接管路的任意位置连接。在本发明所述的装置中，采用前述实施方式更进一步扩宽了本发明装置的应用范围。
根据本发明的装置，在一些实施方式中，本发明所述装置还包括位于聚合单元下游的分离循环单元，以分离得到聚合物产品和未反应单体；在一些优选的实施方式中，所述未反应完的单体可根据需要作为循环物料循环至压缩单元上游，使本发明所述装置具有更好的经济性。
在本发明所述装置中，只要能实现本发明的目的，所述分离循环单元可以根据具体需要进行选择设置，在此不多加赘述。
在本发明中，在一些实施方式中，结合图3，说明本发明的方法在本发明的装置中的运行操作过程：
一股链转移剂209通过链转移剂供给单元的一个出料口与压缩单元201入口处连接，供给链转移剂209，并与一股物料C 208(为新鲜乙烯)混合后通过压缩单元201压缩后分成两股含乙烯源的反应单体流，分别为含乙烯源的反应单体流A 210和含乙烯源的反应单体流B 211；
含乙烯源的反应单体流A 210通过流体吸入输送设备202后得到含乙烯源的反应单体流C 212，含乙烯源的反应单体流C 212与含乙烯源的反应单体流B 211中的其中一部分含乙烯源的反应单体流D 213混合后经过预热设备A 3预热后进入一级高压聚合单元中的管式反应器A中进行反应；
含乙烯源的反应单体流B 211中的剩余部分经过预热设备B 204预热后得到含乙烯源的反应单体流E 214，含乙烯源的反应单体流E 214通过管式反应器B入口侧的减压阀将含乙烯源的反应单体流E 214的压力降低至管式反应器B所需的入口压力后进入一级高压聚合单元中的管式反应器B中进行反应；
管式反应器205a和管式反应器205b流出的物料汇集后，分成两股物料，分别为物料A 207和物料B 215；
物料A 207循环回流体吸入输送设备202中与流入流体吸入输送设备202中的含乙烯源的反应单体流A 210混合流出后继续与含乙烯源的反应单体流B 211中的其中一部分含乙烯源的反应单体流D 213混合并经过预热设备A 203预热后进入一级高压聚合单元中的管式反应器205a中进行反应；
物料B 215流入多级高压聚合单元中管式反应器中进行反应；
多级高压聚合单元中反应后流出的物料经过分离循环单元分离得到聚合物和循环物料，循环物料返 回压缩单元201入口侧；
其中，第一股引发剂I1和第二股引发剂I2分别从引发剂供给单元的第一、二个出料口进入相应的第一反应区205a和第二反应区205b参与一级高压聚合单元中的反应；第三引发剂I3、第四引发剂I4、第五股引发剂I5分别从引发剂供给单元的第三、四、五个出料口进入相应的第三反应区206a、第四反应区206b、第五反应区206c中参与多级高压聚合单元中的反应。
根据本发明的方法，在另一些实施方式中，结合图4，说明本发明的方法的运行操作过程：
一股物料C 208(为新鲜乙烯)通过压缩单元201压缩后分成两股含乙烯源的反应单体流，分别为含乙烯源的反应单体流A 210和含乙烯源的反应单体流B 211；
含乙烯源的反应单体流A 210通过流体吸入输送设备202后得到含乙烯源的反应单体流C 212，含乙烯源的反应单体流C 212与含乙烯源的反应单体流B 211中的其中一部分含乙烯源的反应单体流D 213混合后经过预热设备203预热后进入一级高压聚合单元中的管式反应器A中进行反应；
一股链转移剂209通过链转移剂供给单元的一个出料口供给预热单元中的预热设备B 204对应的反应单体流进料口侧，并与含乙烯源的反应单体流B 211中的剩余部分混合、经过预热设备B 204预热后得到含乙烯源的反应单体流E 214，乙烯源的反应单体流E 214进入一级高压聚合单元中的管式反应器B中进行反应；
管式反应器205a和管式反应器205b流出的物料汇集后，分成两股物料，分别为物料A 207和物料B 215；
物料A 207循环回流体吸入输送设备202中与流入流体吸入输送设备202中的含乙烯源的反应单体流A 210混合流出后继续与含乙烯源的反应单体流B 211中的其中一部分含乙烯源的反应单体流D 213混合并经过预热设备A 203预热后进入一级高压聚合单元中的管式反应器205a中进行反应；
物料B 215流入多级高压聚合单元中管式反应器中进行反应；
多级高压聚合单元中反应后流出的物料经过分离循环单元分离得到聚合物和循环物料，循环物料返回压缩单元201入口侧进行；
其中，第一股引发剂I1和第二股引发剂I2分别从引发剂供给单元的第一、二个出料口进入相应的第一反应区205a和第二反应区205b参与一级高压聚合单元中的反应；第三引发剂I3、第四引发剂I4、第五股引发剂I5分别从引发剂供给单元的第三、四、五个出料口进入相应的第三反应区2206a、第四反应区206b、第五反应区206c中参与多级高压聚合单元中的反应。
本发明中，涉及的压力均为绝压。
以下将通过实施例对本发明进行详细描述。以下实施例中数均分子量(Mn)、重均分子量(Mw)、分子量分布指数(PDI)均采用GB/T 36214.4-2018标准通过高温凝胶渗透色谱HT-GPC测试得到。
实施例201
采用如图3所示的乙烯自由基聚合装置进行乙烯自由基聚合，其中，在图3的装置中，该装置包括：聚合单元、流体吸入输送单元、引发剂供给单元、链转移剂供给单元、压缩单元201、预热单元和分离循环单元；
流体吸入输送单元包括一个流体吸入输送设备202(为射流泵)，用于吸入并输送至少至一股含乙烯源的反应单体流和一级高压聚合单元中的至少一个管式反应器的出料口的部分物料；引发剂供给单元用于向所述聚合单元中输送引发剂，链转移剂供给单元用于向装置中输送链转移剂209；压缩单元201用于将每股含乙烯源的反应单体流各自具备进入所述一级高压聚合单元的入口压力；压缩单元201位于流体吸入输送单元和聚合单元的上游；
聚合单元包括一级高压聚合单元和串联于一级高压聚合单元下游的多级高压聚合单元；一级高压聚合单元包括两个并联的管式反应器A和管式反应器B(管式反应器A和管式反应器B的长度均为560m、内径均为0.045m)，管式反应器B的入口侧设置有减压阀，多级高压聚合单元中用于进行反引发剂供 给单元应的设备为三个依次串联的管式反应器C1、管式反应器C2和管式反应器C3(管式反应器C1、管式反应器C2和管式反应器C3的长度均为400m、内径为0.045m)；
一级高压聚合单元中包括两个反应区：引发剂供给单元的第一、二个出料口分别与一级高压聚合单元中的管式反应器A和管式反应器B中的含乙烯源的反应单体流的入口端连接，以得到第一反应区205a和第二反应区205b；
多级高压聚合单元中包括三个依次串联的反应区：引发剂供给单元的第三、四、五个出料口分别与多级高压聚合单元中的管式反应器C1的入口端(指流入来自一级高压聚合单元的产物的入口)、管式反应器C2的入口端(指流入来自管式反应器C1的产物的入口)、管式反应器C3的入口端(指流入来自管式反应器C2的产物的入口)连接，以得到第三反应区206a、第四反应区206b、第五反应区206c；
流体吸入输送设备202串联设置于一级高压聚合单元中管式反应器A对应的上游；预热单元包括两个并联设置的预热设备A 203与预热设备A 204；预热设备A 203位于流体吸入输送设备202和一级高压聚合单元中对应的管式反应器A之间；预热设备A 204位于压缩单元201和一级高压聚合单元中对应的管式反应器B之间；
链转移剂供给单元的一个出料口与压缩单元201的出口连接，用于向装置中输送链转移剂；
分离循环单元位于聚合单元下游的分离循环单元，以分离得到聚合物产品和循环物料(未反应单体)；
其中，一股链转移剂209(为丙烯、流量为0.205t/h)通过链转移剂供给单元的一个出料口引入压缩单元201的入口并与一股物料C 208(为新鲜乙烯，流量为43.5t/h)混合压缩后分成两股含乙烯源的反应单体流，分别为含乙烯源的反应单体流A 210(流量为16.1t/h)和含乙烯源的反应单体流B 211；
含乙烯源的反应单体流A 210通过流体吸入输送设备202(压降为28MPa)后得到含乙烯源的反应单体流C 212，含乙烯源的反应单体流C 212与含乙烯源的反应单体流B 211中的其中一部分含乙烯源的反应单体流D 213混合后(混合的反应单体流的流量为26.72t/h)经过预热设备A 203预热后进入一级高压聚合单元中的管式反应器A中进行反应；
含乙烯源的反应单体流B 211中的剩余部分经过预热设备B 204预热后得到含乙烯源的反应单体流E 214，含乙烯源的反应单体流E 214通过管式反应器B入口侧的减压阀将含乙烯源的反应单体流E 214的压力降低至管式反应器B所需的入口压力后进入一级高压聚合单元中的管式反应器B中进行反应；
管式反应器A和管式反应器B流出的物料汇集后，分成两股物料，分别为物料A 207和物料B 215；
物料A 207循环回(循环比为0.10)流体吸入输送设备202中与流入流体吸入输送设备202中的含乙烯源的反应单体流A 210混合流出后继续与含乙烯源的反应单体流B 211中的其中一部分含乙烯源的反应单体流D 213混合并经过预热设备A 203预热后进入一级高压聚合单元中的管式反应器A中进行反应；
物料B 215流入多级高压聚合单元中管式反应器中进行反应；
其中，第一股引发剂和第二股引发剂分别从引发剂供给单元的第一、二个出料口进入相应的第一反应区205a和第二反应区205b参与一级高压聚合单元中的反应；第三、四、五股引发剂分别从引发剂供给单元的第三、四、五个出料口进入相应的第三反应区206a、第四反应区206b、第五反应区206c中参与多级高压聚合单元中的反应；
其中：管式反应器A、管式反应器B的入口压力分别为：压力均为220MPa；
第一股引发剂、第二股引发剂、第三股引发剂、第四股和第五股引发剂的流量分别为6.72kg/h、6.80kg/h、8.34kg/h、7.44kg/h、7.48kg/h；
第一股引发剂、第二股引发剂的组成均为：按质量比为1:2:2:2混合的二叔丁基过氧化物、过氧苯甲酸叔丁酯、过氧化-2-乙基己酸叔丁酯和过氧新戊酸叔丁酯的混合物；
第三股引发剂、第四股和第五股引发剂的组成均为：按质量比为10:2:1:1混合的二叔丁基过氧化物、过氧苯甲酸叔丁酯、过氧化-2-乙基己酸叔丁酯和过氧新戊酸叔丁酯的混合物。
第一反应区205a入口的温度为170℃、出口的温度为194℃；第二反应区205b的入口温度为170℃、出口的温度为192℃；第三反应区206a、第四反应区206b、第五反应区206c入口的温度分别为193℃、211℃、229℃，第五反应区206c出口的温度为233℃；第一反应区5a、第二反应区5b、第三反应区206a、第四反应区206b、第五反应区206c的峰值温度均为295℃；
多级高压聚合单元中反应后流出的物料经过分离循环单元分离得到聚合物和循环物料，循环物料返回压缩单元201入口侧；
多级高压聚合单元中反应后流出的物料经过分离循环单元后分离得到聚合物产品(低密度聚乙烯)和未反应单体(即循环物料)。
如表2所示，为该实施例中低密度聚乙烯的数均分子量(Mn)、重均分子量(Mw)、分子量分布指数(PDI)、产量和乙烯转化率。
实施例202
按照实施例201的方法，不同之处在于：该方法在如图4所示的乙烯自由基聚合的装置中进行，
链转移剂供给单元的一个出料口与管式反应器B反应单体流进料口侧连接，用于向装置中输送链转移剂209；
一股物料C 208(为新鲜乙烯，流量为43.5t/h)通过压缩单元201压缩后分成两股含乙烯源的反应单体流，分别为含乙烯源的反应单体流A 210和含乙烯源的反应单体流B 211；
含乙烯源的反应单体流A 210通过流体吸入输送设备202(压降为28MPa)后得到含乙烯源的反应单体流C 212，含乙烯源的反应单体流C 212与含乙烯源的反应单体流B 211中的其中一部分含乙烯源的反应单体流D 213混合后(混合的反应单体流的流量为26.72t/h)经过预热设备203预热后进入一级高压聚合单元中的管式反应器A中进行反应；
一股链转移剂209(为丙烯、流量为0.205t/h)通过链转移剂供给单元的一个出料口引入预热单元中的预热设备204对应的反应单体流进料口侧，并与含乙烯源的反应单体流B 211中的剩余部分混合、经过预热设备204预热后得到含乙烯源的反应单体流E 214，含乙烯源的反应单体流E 214通过管式反应器B入口侧的减压阀将含乙烯源的反应单体流E 214的压力降低至管式反应器B所需的入口压力后进入一级高压聚合单元中的管式反应器B中进行反应；
其中，第一股引发剂和第二股引发剂分别从引发剂供给单元的第一、二个出料口进入相应的第一反应区205a和第二反应区205b参与一级高压聚合单元中的反应；第三、四、五股引发剂分别从引发剂供给单元的第三、四、五个出料口进入相应的第三反应区206a、第四反应区206b、第五反应区206c中参与多级高压聚合单元中的反应；
其中：第一股引发剂、第二股引发剂、第三股引发剂、第四股和第五股引发剂的流量分别为6.54kg/h、6.29kg/h、8.26kg/h、7.67kg/h、8.01kg/h；
第一反应区205a入口的温度为170℃、出口的温度为193℃；第二反应区205b的入口温度为190℃、出口的温度为192℃；第三反应区206a、第四反应区206b、第五反应区206c入口的温度分别为193℃、211℃、225℃，第五反应区206c出口的温度为234℃。
如表2所示，为该实施例中低密度聚乙烯的数均分子量(Mn)、重均分子量(Mw)、分子量分布指数(PDI)、产量和乙烯转化率。
实施例203
按照实施例201的方法，不同之处在于：
链转移剂供给单元的一个出料口不与压缩单元201的出口连接，而与多级高压聚合单元的入口侧(即管式反应器C1的入口侧连接)，用于向装置中输送链转移剂210；
一股物料C 208(为新鲜乙烯，流量为43.5t/h)通过压缩单元201压缩后分成两股含乙烯源的反应单体流，分别为含乙烯源的反应单体流A 210(流量为21t/h)和含乙烯源的反应单体流B 211；
含乙烯源的反应单体流A 210通过流体吸入输送设备202(压降为30MPa)后得到含乙烯源的反应单体流C 212，含乙烯源的反应单体流C 212与含乙烯源的反应单体流B 211中的其中一部分含乙烯源的反应单体流D 213混合后(混合的反应单体流的流量为26.72t/h)经过预热设备A 3预热后进入一级高压聚合单元中的管式反应器A中进行反应；
含乙烯源的反应单体流B 211中的剩余部分经过预热设备B 204预热后得到含乙烯源的反应单体流E 214，含乙烯源的反应单体流E 214通过管式反应器B入口侧的减压阀将含乙烯源的反应单体流E 214的压力降低至管式反应器B所需的入口压力后进入一级高压聚合单元中的管式反应器B中进行反应；
物料A 207循环回(循环比为0.15)流体吸入输送设备202中与流入流体吸入输送设备202中的含乙烯源的反应单体流A 210混合流出后继续与含乙烯源的反应单体流B 211中的其中一部分含乙烯源的反应单体流D 213混合并经过预热设备A 203预热后进入一级高压聚合单元中的管式反应器A中进行反应；
管式反应器A和管式反应器B流出的物料汇集后，分成两股物料，分别为物料A 207和物料B 215；
一股链转移剂209(为丙烯、流量为0.205t/h)通过链转移剂供给单元的一个出料口引入多级高压聚合单元的入口侧并与物料B混合后流入多级高压聚合单元中管式反应器中进行反应；
其中：第一股引发剂、第二股引发剂、第三股引发剂、第四股和第五股引发剂的流量分别为6.48kg/h、6.11kg/h、8.84kg/h、7.79kg/h、8.11kg/h；
第一反应区205a入口的温度为170℃、出口的温度为196℃；第二反应区205b的入口温度为185℃、出口的温度为192℃；第三反应区206a、第四反应区206b、第五反应区206c入口的温度分别为194℃、210℃、225℃，第五反应区206c出口的温度为234℃。
如表2所示，为该实施例中低密度聚乙烯的数均分子量(Mn)、重均分子量(Mw)、分子量分布指数(PDI)、产量和乙烯转化率。
实施例204
按照实施例201的方法，不同之处在于：
管式反应器A、管式反应器B的入口压力均为270MPa；第一股引发剂、第二股引发剂、第三股引发剂、第四股和第五股引发剂的流量分别为3.76kg/h、3.79kg/h、4.66kg/h、4.22kg/h、4.27kg/h；
第一反应区205a入口的温度为170℃、出口的温度为194℃；第二反应区205b的入口温度为170℃、出口的温度为192℃；第三反应区206a、第四反应区206b、第五反应区206c入口的温度分别为193℃、209℃、229℃，第五反应区206c出口的温度为233℃。
如表2所示，为该实施例中低密度聚乙烯的数均分子量(Mn)、重均分子量(Mw)、分子量分布指数(PDI)、产量和乙烯转化率。
实施例205
按照实施例201的方法，不同之处在于：
第一股引发剂、第二股引发剂、第三股引发剂、第四股和第五股引发剂的流量分别为7.89kg/h、7.95kg/h、9.77kg/h、8.97kg/h、9.41kg/h；第一反应区205a入口的温度为170℃、出口的温度为194℃；第二反应区205b的入口温度为170℃、出口的温度为192℃；第三反应区206a、第四反应区206b、第五反应区206c入口的温度分别为193℃、210℃、229℃，第五反应区206c出口的温度为235℃；第一反应区205a、第二反应区205b、第三反应区206a、第四反应区206b、第五反应区206c的峰值温度均为300℃。
如表2所示，为该实施例中低密度聚乙烯的数均分子量(Mn)、重均分子量(Mw)、分子量分布指数(PDI)、产量和乙烯转化率。
实施例206
按照实施例201的方法，不同之处在于：
多级高压聚合单元中用于进行反引发剂供给单元应的设备为一个管式反应器C(管式反应器C的长度为1200m、内径为0.045m)；一级高压聚合单元中包括两个反应区：引发剂供给单元的第一、二个出料口分别与一级高压聚合单元中的管式反应器A和管式反应器B中的含乙烯源的反应单体流的入口端连接，以得到第一反应区205a和第二反应区205b；
多级高压聚合单元中包括两个依次串联的反应区：引发剂供给单元的第三个出料口与多级高压聚合单元的入口端((指流入来自一级高压聚合单元的产物的口)连接，以得到第三反应区206a；引发剂供给单元的第四个出料口与多级高压聚合单元中沿物料流动方向远离入口端的400m处的位置连接，以得到第四反应区206b；
一股引发剂和第二股引发剂分别从引发剂供给单元的第一、二个出料口进入相应的第一反应区205a和第二反应区205b参与一级高压聚合单元中的反应；第三、四股引发剂分别从引发剂供给单元的第三、四个出料口进入相应的第三反应区206a、第四反应区206b中参与多级高压聚合单元中的反应；
第一股引发剂、第二股引发剂、第三股引发剂和第四股引发剂的流量分别为6.75kg/h、6.81kg/h、8.33kg/h、7.46kg/h；
第一反应区205a入口的温度为170℃、出口的温度为194℃；第二反应区205b的入口温度为170℃、出口的温度为192℃；第三反应区206a、第四反应区206b入口的温度分别为193℃、211℃，第四反应区206b出口的温度为230℃；
如表2所示，为该实施例中低密度聚乙烯的数均分子量(Mn)、重均分子量(Mw)、分子量分布指数(PDI)、产量和乙烯转化率。
实施例207
按照实施例201的方法，不同之处在于：
物料A 207循环回(循环比为0.03)流体吸入输送设备202中与流入流体吸入输送设备202中的含乙烯源的反应单体流A 210混合流出后继续与含乙烯源的反应单体流B 211中的其中一部分含乙烯源的反应单体流D 213混合并经过预热设备A 203预热后进入一级高压聚合单元中的管式反应器A中进行反应；
管式反应器A和管式反应器B流出的物料汇集后，分成两股物料，分别为物料A 207和物料B 215；
其中：第一股引发剂、第二股引发剂、第三股引发剂、第四股和第五股引发剂的流量分别为6.83kg/h、6.80kg/h、7.68kg/h、7.35kg/h、7.43kg/h；
第一反应区205a入口的温度为170℃、出口的温度为193℃；第二反应区205b的入口温度为170℃、出口的温度为192℃；第三反应区206a、第四反应区6b、第五反应区206c入口的温度分别为192℃、211℃、229℃，第五反应区206c出口的温度为233℃。
如表2所示，为该实施例中低密度聚乙烯的数均分子量(Mn)、重均分子量(Mw)、分子量分布指数(PDI)、产量和乙烯转化率。
对比例201
按照实施例201的方法，不同之处在于：
乙烯自由基聚合的装置中无流体吸入输送单元，一级高压聚合单元包括一个管式反应器A(管式反应器A的长度为560m、内径均为0.045m)；预热单元包括一个预热设备A 203；预热设备A 203位于压缩单元201和一级高压聚合单元中对应的管式反应器A之间；
一级高压聚合单元中包括一个反应区：引发剂供给单元的第一个出料口分别与一级高压聚合单元中的管式反应器A中的含乙烯源的反应单体流的入口端连接，以得到第一反应区205a；
多级高压聚合单元中包括三个依次串联的反应区：引发剂供给单元的第二、三、四个出料口分别与多级高压聚合单元中的管式反应器C1的入口端(指流入来自一级高压聚合单元的产物的入口)、管式反应器C2的入口端(指流入来自管式反应器C1的产物的入口)、管式反应器C3的入口端(指流入来 自管式反应器C2的产物的入口)连接，以得到第二反应区206a、第三反应区206b、第四反应区206c；
其中，一股链转移剂209(为丙烯、流量为0.205t/h)通过链转移剂供给单元的一个出料口引入压缩单元201的入口并与一股物料C 208(为新鲜乙烯，流量为43.5t/h)混合压缩后经过预热设备A 23预热后送入管式反应器A中进行反应，管式反应A中流出的物料流入多级高压聚合单元中管式反应器中进行反应；
其中，第一股引发剂从引发剂供给单元的第一个出料口进入相应的第一反应区205a参与一级高压聚合单元中的反应；第二、三、四股引发剂分别从引发剂供给单元的第二、三、四个出料口进入相应的第二反应区206a、第三反应区206b、第四反应区206c中参与多级高压聚合单元中的反应；
其中：管式反应器A的入口压力分别为：压力为220MPa；
第一股引发剂、第二股引发剂、第三股引发剂、第四股引发剂的流量分别为14.66kg/h、7.42kg/h、7.36kg/h和7.36kg/h；
第一反应区205a入口的温度为170℃、出口的温度为194℃；第二反应区206a、第三反应区206b和第四反应区206c入口的温度分别为194℃、210℃和229℃，第四反应区206c出口的温度为233℃。
如表2所示，为该实施例中低密度聚乙烯的数均分子量(Mn)、重均分子量(Mw)、分子量分布指数(PDI)、产量和乙烯转化率。
对比例202
按照实施例201的方法，不同之处在于：
循环比为0；第一股引发剂、第二股引发剂、第三股引发剂、第四股和第五股引发剂的流量分别为6.80kg/h、6.80kg/h、7.42kg/h、7.26kg/h和7.36kg/h；
第一反应区205a和第二反应区205b入口的温度均为170℃、出口的温度均为192℃；第三反应区206a、第四反应区206b和第五反应区206c入口的温度分别为192℃、211℃和229℃，第五反应区206c出口的温度为233℃。
如表2所示，为该实施例中低密度聚乙烯的数均分子量(Mn)、重均分子量(Mw)、分子量分布指数(PDI)、产量和乙烯转化率。
对比例203
其他同实施例204，不同之处在于：
乙烯自由基聚合的装置中无流体吸入输送单元，一级高压聚合单元包括一个管式反应器A(管式反应器A的长度为560m、内径均为0.045m)；预热单元包括一个预热设备A 203；预热设备A 203位于压缩单元201和一级高压聚合单元中对应的管式反应器A之间；
一级高压聚合单元中包括一个反应区：引发剂供给单元的第一个出料口分别与一级高压聚合单元中的管式反应器A中的含乙烯源的反应单体流的入口端连接，以得到第一反应区205a；
多级高压聚合单元中包括三个依次串联的反应区：引发剂供给单元的第二、三、四个出料口分别与多级高压聚合单元中的管式反应器C1的入口端(指流入来自一级高压聚合单元的产物的入口)、管式反应器C2的入口端(指流入来自管式反应器C1的产物的入口)、管式反应器C3的入口端(指流入来自管式反应器C2的产物的入口)连接，以得到第二反应区206a、第三反应区206b、第四反应区206c；
其中，一股链转移剂209(为丙烯、流量为0.205t/h)通过链转移剂供给单元的一个出料口引入压缩单元201的入口并与一股物料C 208(为新鲜乙烯，流量为43.5t/h)混合压缩后经过预热设备A 203预热后送入管式反应器A中进行反应，管式反应A中流出的物料流入多级高压聚合单元中管式反应器中进行反应；
其中，第一股引发剂从引发剂供给单元的第一个出料口进入相应的第一反应区205a参与一级高压聚合单元中的反应；第二、三、四股引发剂分别从引发剂供给单元的第二、三、四个出料口进入相应的第二反应区206a、第三反应区206b、第四反应区206c中参与多级高压聚合单元中的反应；
其中：管式反应器A的入口压力分别为：压力为270MPa；第一股引发剂、第二股引发剂、第三股引发剂和第四股引发剂的流量分别为8.12kg/h、4.14kg/h、4.12kg/h、4.18kg/h；第一反应区5a入口的温度为170℃、出口的温度为193℃；第二反应区206a、第三反应区206b和第四反应区206c入口的温度分别为193℃、210℃和229℃，第四反应区6c出口的温度为233℃。
如表2所示，为该实施例中低密度聚乙烯的数均分子量(Mn)、重均分子量(Mw)、分子量分布指数(PDI)、产量和乙烯转化率。
对比例204
其他同实施例205，不同之处在于：
乙烯自由基聚合的装置中无流体吸入输送单元，一级高压聚合单元包括一个管式反应器A(管式反应器A的长度为560m、内径均为0.045m)；预热单元包括一个预热设备A 203；预热设备A 203位于压缩单元201和一级高压聚合单元中对应的管式反应器A之间；
一级高压聚合单元中包括一个反应区：引发剂供给单元的第一个出料口分别与一级高压聚合单元中的管式反应器A中的含乙烯源的反应单体流的入口端连接，以得到第一反应区205a；
多级高压聚合单元中包括三个依次串联的反应区：引发剂供给单元的第二、三、四个出料口分别与多级高压聚合单元中的管式反应器C1的入口端(指流入来自一级高压聚合单元的产物的入口)、管式反应器C2的入口端(指流入来自管式反应器C1的产物的入口)、管式反应器C3的入口端(指流入来自管式反应器C2的产物的入口)连接，以得到第二反应区206a、第三反应区206b、第四反应区206c；
其中，一股链转移剂209(为丙烯、流量为0.205t/h)通过链转移剂供给单元的一个出料口引入压缩单元201的入口并与一股物料C 208(为新鲜乙烯，流量为43.5t/h)混合压缩后经过预热设备A 203预热后送入管式反应器A中进行反应，管式反应A中流出的物料流入多级高压聚合单元中管式反应器中进行反应；
其中，第一股引发剂从引发剂供给单元的第一个出料口进入相应的第一反应区205a参与一级高压聚合单元中的反应；第二、三、四股引发剂分别从引发剂供给单元的第二、三、四个出料口进入相应的第二反应区206a、第三反应区206b、第四反应区206c中参与多级高压聚合单元中的反应；其中：管式反应器A的入口压力分别为：压力为220MPa；
第一股引发剂、第二股引发剂、第三股引发剂和第四股引发剂的流量分别为16.85kg/h、8.65kg/h、8.76kg/h、9.15kg/h；
第一反应区205a入口的温度为170℃、出口的温度为193℃；第二反应区206a、第三反应区206b和第四反应区206c入口的温度分别为193℃、211℃和230℃，第四反应区206c出口的温度为235℃。
如表2所示，为该实施例中低密度聚乙烯的数均分子量(Mn)、重均分子量(Mw)、分子量分布指数(PDI)、产量和乙烯转化率。
对比例205
其他同实施例206，不同之处在于：
循环比为0；第一股引发剂、第二股引发剂、第三股引发剂和第四股引发剂的流量分别为6.79kg/h、6.80kg/h、7.43kg/h、7.26kg/h；
第一反应区205a入口的温度为170℃、出口的温度为192℃；第二反应区205b的入口温度为170℃、出口的温度为192℃；第三反应区206a、第四反应区206b入口的温度分别为192℃、211℃，第四反应区206b出口的温度为229℃。
如表2所示，为该实施例中低密度聚乙烯的数均分子量(Mn)、重均分子量(Mw)、分子量分布指数(PDI)、产量和乙烯转化率。
表2

通过表2的结果可以看出，通过本发明所述的多级高压聚合中包括3个依次串联的反应区时的实施例201-205、207与对比例201-204的对比、多级高压聚合中包括2个依次串联的反应区时的实施例206与对比例205的对比可知，在通过链转移剂的进料位置、反应温度、引发剂的进料量等参数来调节产品的分子量分布指数时，本发明实施例201-207中的装置能够更好的实现到分子量分布指数范围更宽的调控效果，同时还能够得到分子量分布指数更宽和转化率更高的产品。
以上详细描述了本发明的优选实施方式，但是，本发明并不限于此。在本发明的技术构思范围内，可以对本发明的技术方案进行多种简单变型，包括各个技术特征以任何其它的合适方式进行组合，这些简单变型和组合同样应当视为本发明所公开的内容，均属于本发明的保护范围。

Claims (25)

  1. 一种烯烃自由基聚合的方法，其特征在于，该方法包括：将含烯烃源的反应单体流引入至少两个并联的管式反应器中，各自进行一级高压聚合，再将得到的一级高压聚合产物流引入一个或多个依次串联的管式反应器中进行多级高压聚合；
    其中，将至少一股自由基聚合引发剂各自引入参与一级高压聚合和/或多级高压聚合，所述含烯烃源的反应单体流的压力大于等于100MPa。
  2. 根据权利要求1所述的方法，其中，所述含烯烃源的反应单体流的压力为110-400MPa；
    和/或，一级高压聚合前后的压降与多级高压聚合前后的压降的总和：一级高压聚合前后的压降为3∶1-30∶1，优选为6∶1-8∶1。
  3. 根据权利要求1或2所述的方法，其中，每股所述含烯烃源的反应单体流的温度各自为100-200℃，且每股每个并联的管式反应器的入口的所述含烯烃源的反应单体流的和各自满足关联式：10000≥ρ1/μ1≥1500，优选6000≥ρ1/μ1≥3000；密度ρ1的单位为kg/m3，粘度μ1的单位为cP；
    和/或，每个一级高压聚合和每个多级高压聚合各自的温度为100-350℃。
  4. 根据权利要求1或2所述的方法，其中，各股所述含烯烃源的反应单体流的最大进料量与最小进料量的比值为(20-1)∶1，优选为(5-1)∶1；
    和/或，每股所述含烯烃源的反应单体流的流速各自大于等于5m/s，且小于等于30m/s，优选大于等于8m/s，且小于等于20m/s。
  5. 根据权利要求1或2所述的方法，其中，所述一级高压聚合的管式反应器的个数为2-4个；
    和/或，至少一股自由基引发剂引入参与一级高压聚合；至少一股引发剂引入参与多级高压聚合。
  6. 根据权利要求1或2所述的方法，其中，该方法还包括将至少一股调聚剂各自进料参与所述一级高压聚合和/或多级高压聚合；
    和/或，该方法还包括将至少一股共聚单体各自进料参与所述一级高压聚合和/或多级高压聚合；
    和/或，该方法还包括，将所述多级高压聚合得到的物料经过减压冷却后，分离得到未反应的单体和聚合物产品。
  7. 根据权利要求6所述的方法，其中，所述调聚剂包括脂肪烃、烯烃、酮、醛、脂族醇和氢中的一种或多种。
  8. 根据权利要求1或2所述的方法，其中，所述烯烃源中的烯烃包括R2C＝CR2型单烯类化合物、共轭双烯烃、非共轭双烯烃中的一种或多种，其中，R选自H、烃基或卤素；
    和/或，所述自由基聚合引发剂包括氧、空气、偶氮化合物、有机过氧化物及C-C引发剂的烃中的一种或多种。
  9. 根据权利要求1所述的方法，其中，所述方法包括：将含乙烯源的反应单体流引入至少两个并联的管式反应器中在引发剂存在下进行反应；将至少两个并联的管式反应器中的至少一个管式反应器的出料口的部分物料循环回至少两个并联的管式反应器的至少一个管式反应器中进行反应；将至少两个并联的管式反应器的出料口的剩余物料继续引入一个或多个依次串联的管式反应器中在引发剂的存在下进行反应。
  10. 根据权利要求9所述的方法，其中，所述至少两个并联的管式反应器中的出料口的物料的循环比小于1，优选为小于等于0.3，更优选为0.04-0.2；
    和/或，流入至少两个并联的管式反应器的所述含乙烯源的反应单体流的压力为140-300MPa；
    和/或，所述至少两个并联的管式反应器，和一个或多个依次串联的管式反应器中的反应温度各自为100-350℃；
    和/或，所述至少两个并联的管式反应器中的反应单体流的最大进料量与最小进料量的重量比为1∶(0.01-1)；
    和/或，将引发剂引入至少两个并联的管式反应器中和一个或多个依次串联的管式反应器中。
  11. 根据权利要求9或10所述的方法，其中，该方法还包括将链转移剂引入所述至少两个并联的管式反应器，和一个或多个依次串联的管式反应器中；
    和/或，该方法还包括将至少一股共聚单体引入至少两个并联的管式反应器，和一个或多个依次串联的管式反应器中。
  12. 根据权利要求11所述的方法，其中，所述链转移剂选自脂肪烃、烯烃、酮、醛、脂族醇和氢中的一种或多种；所述引发剂选自偶氮化合物、有机过氧化物、氧气和空气中的一种或多种。
  13. 一种烯烃自由基聚合装置，其特征在于，该装置包括：一级高压聚合单元和多级高压聚合单元；其中，所述一级高压聚合单元串联于多级高压聚合单元的上游；所述一级高压聚合单元包括至少两个并联的管式反应器，用于将至少两股含烯烃源的反应单体流各自进行一级高压聚合；
    所述多级高压聚合单元包括一个或多个依次串联的管式反应器，用于将来自于一级高压聚合单元的产物进行多级高压聚合；
    所述一级高压聚合单元和/或多级高压聚合单元中的至少一个管式反应器上设置有引发剂进料口。
  14. 根据权利要求13所述的装置，其中，所述一级高压聚合单元包括2-4个并联的管式反应器；
    和/或，所述一级高压聚合单元中的至少一个管式反应器的反应单体流进料口端设置有引发剂进料口；
    和/或，所述一级高压聚合单元中的至少一个管式反应器上沿其长度方向设置有至少一个引发剂进料口，优选设置有1-3个引发剂进料口；
    和/或，所述多级高压聚合单元中的至少一个管式反应器上设置有至少一个引发剂进料口，优选设置有2-5个引发剂进料口。
  15. 根据权利要求13或14所述的装置，其中，该装置还包括位于所述一级高压聚合单元的上游的至少一个压缩单元，用于将每股含烯烃源的反应单体流各自具备进入所述一级高压聚合单元的入口压力；其中，所述压缩单元包括一个或多个依次串联的压缩机；
    该装置还包括位于所述一级高压聚合单元的上游的至少一个预热器，用于将每股含烯烃源的反应单体流各自具备进入所述一级高压聚合单元的入口温度。
  16. 根据权利要求15所述的装置，其中，至少一个压缩单元串联设置于一级高压聚合单元中的至少两个并联的管式反应器共同的上游；
    和/或，至少一个压缩单元串联设置于一级高压聚合单元中的管式反应器各自对应的上游；
    和/或，所述预热器位于所述压缩单元与一级高压聚合单元之间；
    和/或，至少一个预热器串联设置于一级高压聚合单元中的至少两个并联的管式反应器共同的上游；
    和/或，至少一个预热器串联设置于一级高压聚合单元中的管式反应器各自对应的上游。
  17. 根据权利要求15所述的装置，其中，该装置还包括至少一个调聚剂进料口和/或至少一个共聚单体进料口；
    其中，所述调聚剂进料口设置于多级高压聚合单元出口上游的任意位置；
    其中，所述共聚单体进料口设置于一级高压聚合单元出口上游的任意位置。
  18. 根据权利要求17所述的装置，其中，每个所述调聚剂进料口各自设置于：压缩单元的入口；
    和/或，压缩单元的出口；
    和/或，压缩单元中任意两个相邻压缩机的连接管道处；
    和/或，一级高压聚合单元中的至少一个管式反应器的反应单体流进料口侧；
    和/或，一级高压聚合单元中的至少一个管式反应器上；
    和/或，一级高压聚合单元和多级高压聚合单元的连接管道处；
    和/或，多级高压聚合单元中的至少一个管式反应器上；
    和/或，每个所述共聚单体进料口各自设置于：压缩单元的入口；
    和/或，压缩单元中任意两个相邻压缩机的连接管道处；
    和/或，一级高压聚合单元中的至少一个管式反应器的反应单体流进料口侧；
    和/或，预热器与压缩单元的连接管道处。
  19. 根据权利要求13所述的装置，其中，所述装置还包括：流体吸入输送单元，其中，所述流体吸入输送单元包括一个或至少两个并联设置的流体吸入输送设备，用于吸入并输送至少一股含乙烯源的反应单体流和一级高压聚合单元中的至少一个管式反应器的出料口的部分物料；
    引发剂供给单元用于向所述一级高压聚合单元和多级高压聚合单元中输送引发剂。
  20. 根据权利要求19所述的装置，其中，所述一级高压聚合单元包括2-4个并联的管式反应器；
    和/或，至少一个所述流体吸入输送设备串联设置于所述一级高压聚合单元中的至少两个并联的管式反应器共同的上游；
    和/或，至少一个所述流体吸入输送设备串联设置于所述一级高压聚合单元中的管式反应器各自对应的上游。
  21. 根据权利要求19或20所述的装置，其中，该装置还包括位于所述流体吸入输送单元与聚合单元上游的压缩单元，用于将每股含乙烯源的反应单体流各自具备进入所述一级高压聚合单元的入口压力；其中，所述压缩单元包括至少2级压缩机；
    和/或，该装置还包括位于所述聚合单元上游的预热单元，用于将含乙烯源的反应单体流具备进入聚合单元中的进料温度，其中，所述预热单元包括一个或多个并联设置的预热设备。
  22. 根据权利要求21所述的装置，其中，至少一个压缩单元串联设置于所述流体吸入输送单元中的至少两个并联的管式反应器共同的上游；
    和/或，至少一个压缩单元串联设置于所述流体吸入输送单元中的流体吸入输送设备各自对应的上游；
    和/或，至少一个所述预热设备串联设置于一级高压聚合单元中的至少两个并联的管式反应器共同的上游；
    和/或，至少一个所述预热设备串联设置于一级高压聚合单元中的管式反应器各自对应的上游；
    和/或，至少一个所述预热设备位于压缩单元和一级高压聚合单元中对应的管式反应器之间；
    和/或，至少一个所述预热设备位于流体吸入输送单元中对应的吸入流体设备与一级高压聚合单元中对应的管式反应器之间。
  23. 根据权利要求21所述的装置，其中，该装置还包括：链转移剂供给单元，用于向装置中输送链转移剂；
    其中，所述链转移剂供给单元的至少一个出料口与所述一级高压聚合单元中的管式反应器上任意位置连接；
    和/或，所述链转移剂供给单元的至少一个出料口与预热设备各自对应的反应单体流进料口侧连接；
    和/或，所述链转移剂供给单元的至少一个出料口与至少两个并联设置的预热设备的共同上游连接；
    和/或，所述链转移剂供给单元的至少一个出料口与流体吸入输送设备各自对应的反应单体流进料口侧连接；
    和/或，所述链转移剂供给单元的至少一个出料口与至少两个并联设置的流体吸入输送设备的共同上游连接；
    和/或，所述链转移剂供给单元的至少一个出料口与多级高压聚合单元的入口侧连接；
    和/或，所述链转移剂供给单元的至少一个出料口与多级高压聚合单元中任意位置连接；
    和/或，所述链转移剂供给单元的至少一个出料口与压缩单元入口连接；
    和/或，所述链转移剂供给单元的至少一个出料口与压缩单元的压缩级间连接管路的任意位置连接。
  24. 根据权利要求21所述的装置，其中，该装置还包括共聚单体供给单元，以向装置中提供共聚单体；其中，所述共聚单体供给单元的至少一个出料口与所述一级高压聚合单元中的管式反应器各自对应的反应单体流进料口侧连接；
    和/或，所述共聚单体供给单元的至少一个出料口与预热设备各自对应的反应单体流进料口侧连接；
    和/或，所述共聚单体供给单元的至少一个出料口与至少两个并联设置的预热设备的共同上游连接；
    和/或，所述共聚单体供给单元的至少一个出料口与流体吸入输送设备各自对应的反应单体流进料口侧连接；
    和/或，所述共聚单体供给单元的至少一个出料口与至少两个并联设置的流体吸入输送设备的共同上游连接；
    和/或，所述共聚单体供给单元的至少一个出料口与压缩单元入口连接；
    和/或，所述共聚单体供给单元的至少一个出料口与压缩单元的压缩级间连接管路的任意位置连接。
  25. 根据权利要求13-24中任意一项所述的装置，其中，该装置还包括位于多级高压聚合单元下游的分离循环单元，用于分离多级高压聚合得到的物料，以得到聚合产品和未反应的单体。
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