WO2019090883A1 - 聚丙烯或丙烯乙烯共聚物的制备方法 - Google Patents

聚丙烯或丙烯乙烯共聚物的制备方法 Download PDF

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WO2019090883A1
WO2019090883A1 PCT/CN2017/115399 CN2017115399W WO2019090883A1 WO 2019090883 A1 WO2019090883 A1 WO 2019090883A1 CN 2017115399 W CN2017115399 W CN 2017115399W WO 2019090883 A1 WO2019090883 A1 WO 2019090883A1
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propylene
liquid phase
gas
polymerization
homopolymerization
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English (en)
French (fr)
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李泽民
刘立新
崔忠
陈红
刘利妍
吴霞
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北京华福工程有限公司
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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
    • C08F10/00—Homopolymers and copolymers of unsaturated aliphatic hydrocarbons having only one carbon-to-carbon double bond
    • C08F10/04—Monomers containing three or four carbon atoms
    • C08F10/06—Propene
    • 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
    • C08F110/04—Monomers containing three or four carbon atoms
    • C08F110/06—Propene
    • 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
    • C08F2/02—Polymerisation in bulk
    • 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
    • C08F2/34—Polymerisation in gaseous state
    • 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
    • C08F210/00—Copolymers of unsaturated aliphatic hydrocarbons having only one carbon-to-carbon double bond
    • C08F210/02—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
    • C08F210/00—Copolymers of unsaturated aliphatic hydrocarbons having only one carbon-to-carbon double bond
    • C08F210/04—Monomers containing three or four carbon atoms
    • C08F210/06—Propene
    • 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
    • C08F255/00—Macromolecular compounds obtained by polymerising monomers on to polymers of hydrocarbons as defined in group C08F10/00
    • C08F255/02—Macromolecular compounds obtained by polymerising monomers on to polymers of hydrocarbons as defined in group C08F10/00 on to polymers of olefins having two or three carbon atoms
    • 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
    • C08F285/00—Macromolecular compounds obtained by polymerising monomers on to preformed graft polymers
    • 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
    • C08F4/00—Polymerisation catalysts
    • C08F4/42—Metals; Metal hydrides; Metallo-organic compounds; Use thereof as catalyst precursors
    • C08F4/44—Metals; Metal hydrides; Metallo-organic compounds; Use thereof as catalyst precursors selected from light metals, zinc, cadmium, mercury, copper, silver, gold, boron, gallium, indium, thallium, rare earths or actinides
    • C08F4/60—Metals; Metal hydrides; Metallo-organic compounds; Use thereof as catalyst precursors selected from light metals, zinc, cadmium, mercury, copper, silver, gold, boron, gallium, indium, thallium, rare earths or actinides together with refractory metals, iron group metals, platinum group metals, manganese, rhenium technetium or compounds thereof
    • C08F4/62—Refractory metals or compounds thereof
    • C08F4/64—Titanium, zirconium, hafnium or compounds thereof
    • C08F4/646—Catalysts comprising at least two different metals, in metallic form or as compounds thereof, in addition to the component covered by group C08F4/64
    • 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
    • C08F6/00—Post-polymerisation treatments
    • C08F6/26—Treatment of polymers prepared in bulk also solid polymers or polymer melts
    • C08F6/28—Purification

Definitions

  • the present invention relates to the field of industrialization of propylene polymerization, and in particular to a process for preparing a polypropylene or propylene ethylene copolymer in a three-stage process.
  • Polypropylene is a polymer made of propylene as a monomer. It is a very important variety in general-purpose plastics. It is rich in raw materials, low in price, easy to process and form, non-toxic, etc. It can also be copolymerized, blended, Modification and other measures for modification have a very broad application field. Among polyolefin resins, it is second only to polyethylene and the second largest plastic. Further, propylene can be copolymerized with ethylene to prepare a propylene-ethylene random copolymer.
  • the process technologies for producing polypropylene at home and abroad mainly include a liquid phase bulk method, a gas phase bulk method, and a polymerization method in which a liquid phase body and a gas phase body are connected in series.
  • a liquid phase body and a gas phase body are connected in series.
  • it can be divided into Spheripol, Gasoline Fluidized Bed (Unipol), Vertical Stirred Tank Process (Novolen, Hypol), Horizontal Stirred Tank Process (Inoes, JPP) and different The method of combining the reactor phases.
  • the polymerization reaction takes place in the liquid phase.
  • prepolymerization is provided, and in the gas phase polymerization process, the catalyst and the raw material propylene are directly added to the reaction. There is no pre-polymerization.
  • the propylene monomer is polymerized in a polymerization reactor under the action of a catalyst, an activator and an electron donor to obtain a polypropylene slurry or a solid material, and the resulting polypropylene is further subjected to flash degassing. , catalyst deactivation, drying and nitrogen replacement, to obtain polypropylene powder products, the propylene gas entrained in the reactor discharge process, after low pressure washing, gas compression, light gas separation, return to the reaction system for recycling, thereby reducing raw materials Consumption of propylene and hydrogen.
  • INEOS's Innovene process using a horizontal stirred reactor and a high performance INcat CDi catalyst, the reactor is close to the plug flow reactor, using liquid phase propylene vaporization to remove heat.
  • the utility model has the advantages of high heat removal efficiency of the reactor, stable production operation, uniform material residence time during the reaction period, wide range of melting index and low energy consumption of the produced product; disadvantages thereof are that hot spots and plasticized blocks and reactors are easily generated in the reaction kettle. The processing is difficult, the key equipment needs to be imported, and the investment is high.
  • the object of the present invention is to solve the problem of how to reduce the unit energy consumption of polypropylene production, improve the quality and stability of the product, and the safety of the operation of the device, and reduce the loss of propylene, and propose a preparation method of polypropylene or propylene ethylene copolymer.
  • the method is coupled with propylene prepolymerization, propylene liquid phase bulk polymerization and propylene vapor phase bulk polymerization, which can realize the mild prepolymerization of propylene, reduce the amount of catalyst, meet the requirements of propylene polymer products, and reduce the unit energy consumption and propylene in the industrial implementation of the whole process. Loss, improve the economics of industrial production of propylene.
  • the present invention provides a method for preparing a polypropylene or propylene ethylene copolymer, the method comprising:
  • the prepolymerization temperature is 40 to 45 ° C, and the prepolymerization pressure is 3.2 to 3.9 MPaG;
  • the polymerization catalyst comprises an olefin polymerization catalyst, an activator and an electron donor, the olefin polymerization catalyst being used in an amount of 0.04 to 0.06% by weight of the liquid phase propylene;
  • the liquid phase polymerization raw material is sent to a liquid phase polymerization tank and ethylene is randomly copolymerized in a liquid phase to obtain a copolymer slurry; wherein the ethylene feed amount is 5% by weight or less of the liquid phase propylene;
  • the product containing the propylene homopolymer or the product containing the propylene ethylene random copolymer is subjected to gas-solid separation, and the separated solid is dried to obtain a polypropylene or propylene ethylene copolymer, and the separated gas recovers propylene and hydrogen.
  • the activator and the electron donor are used in an amount of 0.2 to 0.4% by weight and 0.04 to 0.06% by weight, respectively, of the liquid phase propylene.
  • the liquid phase propylene pressure is 4 to 4.5 MPaG, and the liquid phase propylene temperature is 40 to 45 °C.
  • the residence time of the propylene prepolymerization is 4 to 5 min.
  • the propylene prepolymer is a polypropylene having a polymerization ratio of 50 to 100 times.
  • the concentration of the polypropylene in the liquid phase polymerization raw material is 150-300 g / L; the amount of the recovered liquid phase propylene is 25-30% by weight of the liquid phase propylene; The amount of hydrogen added is 0.04 to 0.3 kg with respect to 1000 kg of polypropylene in the liquid phase polymerization raw material.
  • the liquid phase homopolymerization temperature of the propylene is 65 to 70 ° C
  • the liquid phase homopolymerization pressure of the propylene is 3 to 3.8 MPaG
  • the residence time of the liquid phase homopolymerization of the propylene is 35 to 45 minutes.
  • the liquid phase random copolymerization temperature is 65 to 70 ° C
  • the liquid phase random copolymerization pressure is 3 to 3.8 MPaG
  • the residence time of the liquid phase random copolymerization is 35 to 45 minutes.
  • the liquid phase homopolymerization or liquid phase random copolymerization of propylene is carried out in a liquid phase polymerization vessel with an external cooler, and vaporization of a part of the liquid phase propylene is carried out.
  • the heat of partial reaction of the propylene liquid phase homopolymerization or liquid phase random copolymerization is carried out.
  • the vaporized propylene gas is condensed or compressed and returned to the propylene liquid phase homopolymerization or the liquid phase random copolymerization as a first condensate or a first recycle propylene gas.
  • the gas phase homopolymerization temperature of the propylene is 80 to 95 ° C
  • the gas phase homopolymerization pressure of the propylene is 2.5 to 2.8 MPaG
  • the residence time of the gas phase homopolymerization of the propylene is 45 to 60 minutes.
  • the gas phase random copolymerization temperature is 80 to 95 ° C
  • the gas phase random copolymerization pressure is 2.5 to 2.8 MPaG
  • the gas phase random copolymerization residence time is 45 to 60 min.
  • the gas phase polymerization reactor is a horizontal reactor with an external cooler, and the polypropylene slurry or copolymer slurry is in the gas phase polymerization reactor.
  • the loading is 35 to 60% by volume of the gas phase polymerization reactor.
  • the unreacted propylene gas discharged from the gas phase polymerization reactor is condensed or compressed and returned to the propylene vapor phase homopolymerization or the gas phase random copolymerization as a second condensate or a second recycle propylene gas.
  • the recovered propylene gas is returned to the step (2) as the recovered liquid phase propylene, and the recovered hydrogen is mixed with the fresh hydrogen as the recycled hydrogen gas to return to the step (2).
  • the present invention provides a three-stage polymerization production process coupled with liquid phase prepolymerization, liquid phase bulk polymerization and gas phase bulk polymerization of propylene, which is used for industrial production of propylene homopolymerization or random copolymerization.
  • the conditions of propylene prepolymerization and feed control can be improved.
  • propylene prepolymerization can be carried out at 40 to 45 ° C and 3.2 to 3.9 MPa G, which is more than the conditions of 5 to 10 ° C of the prior art. It is mild; thus, in the feed control, the propylene feed does not need to condense to below zero.
  • the addition of the polymerization catalyst is carried out all at once from the prepolymerization, and the prior art batch addition of the catalyst is not required, which simplifies the operation steps and reduces the fluctuation of the polymerization process.
  • the process provided by the invention can realize that the primary feed of the liquid phase propylene is all involved in the prepolymerization of propylene, and the raw material slurry containing the better dispersed polypropylene prepolymer is obtained, which is favorable for the uniform reaction of the liquid phase and gas phase polymerization of the subsequent coupling. To improve the quality of polymer products.
  • the entire process recycles unreacted propylene gas without the need to additionally replenish fresh propylene during the process.
  • the overall process provided by the present invention can reduce energy consumption per unit product and loss of propylene.
  • the unit energy consumption of the product is less than 50kg standard oil/ton PP powder, and the minimum can reach 40kg standard oil/ton PP powder.
  • the processing cost per unit product is about 1/2 of the same type of imported technology products of the same scale.
  • the method provided by the invention can avoid the occurrence of hot spots and explosions in the reaction, reduce the probability of plasticized blocks in the product, and the product particles are not easily broken; the polymerization reaction adopts liquid phase propylene vaporization to remove heat, and the propylene cycle
  • the recycling amount is small, the system does not need large-scale circulation equipment, which can effectively reduce the investment of the device, the construction period is short, and the economy is better.
  • FIG. 1 is a schematic view showing the process flow of a method for preparing polypropylene according to the present invention
  • FIG. 2 is a schematic view showing the process flow of a method for preparing a propylene ethylene copolymer of the present invention.
  • the invention provides a preparation method of a polypropylene or a propylene ethylene copolymer, as shown in FIG. 1 and FIG. 2, the method comprises:
  • the prepolymerization temperature is 40 to 45 ° C, and the prepolymerization pressure is 3.2 to 3.9 MPaG;
  • the polymerization catalyst comprises an olefin polymerization catalyst, an activator and an electron donor, the olefin polymerization catalyst being used in an amount of 0.04 to 0.06% by weight of the liquid phase propylene;
  • the liquid phase polymerization raw material is sent to a liquid phase polymerization tank and ethylene is randomly copolymerized in a liquid phase to obtain a copolymer slurry; wherein the ethylene feed amount is 5% by weight or less of the liquid phase propylene;
  • the product containing the propylene homopolymer or the product containing the propylene ethylene random copolymer is subjected to gas-solid separation, and the separated solid is dried to obtain a polypropylene or propylene ethylene copolymer, and the separated gas recovers propylene and hydrogen.
  • the propylene prepolymerization is carried out under the above conditions, and the combination with the liquid phase bulk polymerization and the gas phase bulk polymerization can be realized, the propylene raw material condensation, the batch feed and the catalyst amount can be reduced, and the liquid phase bulk polymerization can be assisted.
  • gas phase bulk polymerization utilizes propylene liquefaction to remove heat and reuse unreacted propylene, and control the polymerization conditions accordingly.
  • the whole process is used in the industrialization process to achieve sufficient progress of each stage of polymerization, reducing the unit energy consumption of polypropylene production and reducing the production process.
  • the loss of propylene raw materials reduces the production cost of propylene polymerization.
  • the step (1) first completes the prepolymerization of the liquid phase propylene.
  • the prepolymerized feed can realize that the propylene raw materials are all fed in the liquid phase without the prior art, in general, part of the propylene is condensed and then enters.
  • the prepolymerization and another portion enters the liquid phase bulk polymerization, which simplifies the feed control and allows the propylene feedstock to all undergo prepolymerization to form a more uniform prepolymer.
  • propylene is pressurized and liquefied into liquid phase propylene.
  • the liquid phase propylene pressure is 4 to 4.5 MPaG
  • the liquid phase propylene temperature is 40 to 45 °C.
  • the propylene prepolymerization may be carried out as long as it provides a polypropylene which satisfies the polymerization ratio.
  • the residence time of the propylene prepolymerization is 4 to 5 minutes.
  • the propylene prepolymer is better dispersed, and the propylene prepolymer is a polypropylene having a polymerization ratio of 50 to 100 times, which is advantageous for the subsequent step (2).
  • the reaction process of liquid phase homopolymerization of propylene and gas phase homopolymerization of propylene is more stable, which is beneficial to reduce production energy consumption and propylene loss.
  • the polymerization catalyst may include an olefin polymerization catalyst, an activator, and an electron donor.
  • the olefin polymerization catalyst is selected from a titanium-based propylene polymerization catalyst, such as a catalyst of the domestic grade CS-1; the activator is selected from triethylaluminum; and the electron donor is selected from cyclohexylmethyldimethoxysilane. (hereinafter referred to as "silane").
  • the liquid phase propylene can simultaneously serve as a carrier for the olefin polymerization catalyst, the activator and the electron donor, and the olefin polymerization catalyst, the activator and the electron donor are mixed into the liquid phase propylene to enter the prepolymerization kettle to carry out the propylene prepolymerization.
  • prepolymerization is carried out at 40 to 45 ° C and 3.2 to 3.9 MPaG, and the olefin polymerization catalyst can be used in a reduced amount.
  • the activator and the electron donor are used in an amount of 0.2 to 0.4% by weight and 0.04 to 0.06% by weight, respectively, of the liquid phase propylene.
  • the pre-polymerization kettle can be selected from a vertical tank reactor with stirring and jacketing, or a small loop tube with a cooling jacket, and the heat exchange area of the jacket area or the inner cooling tube must satisfy the pre-polymerization.
  • the heat removal requirement of the reaction can be.
  • the propylene prepolymerization completed under the condition of the step (1) the obtained raw material slurry can be further combined with liquid phase bulk polymerization and gas phase bulk polymerization to continue homopolymerization or random copolymerization of propylene to obtain a desired product.
  • the raw material slurry contains a prepolymerized propylene prepolymer, unreacted liquid phase propylene, and a non-reactive olefin polymerization catalyst, an activator, and an electron donor.
  • Step (2) may be used to prepare a material for subsequent liquid phase and gas phase polymerization, and the liquid phase propylene and hydrogen may be recovered into the raw material slurry to obtain the liquid phase polymerization raw material, and the composition thereof satisfies the continuous liquid phase polymerization and gas phase of propylene. Polymerization, the entire process reduces unit energy consumption and propylene loss in polypropylene production.
  • the concentration of the polypropylene in the liquid phase polymerization raw material is 150-300 g / L; the amount of the recovered liquid phase propylene is 25-30% by weight of the liquid phase propylene;
  • the amount of hydrogen added is 0.04 to 0.3 kg with respect to 1000 kg of polypropylene in the liquid phase polymerization raw material.
  • the amount of the recovered liquid phase propylene and hydrogen can be adjusted to satisfy the above conditions.
  • One embodiment of the present invention can perform liquid phase homopolymerization of propylene to produce a propylene homopolymerized product.
  • the liquid phase homopolymerization temperature of the propylene is 65 to 70 ° C
  • the liquid phase homopolymerization pressure of the propylene is 3 to 3.8 MPaG
  • the residence time of the liquid phase homopolymerization of the propylene is 35 to 45 minutes.
  • Another embodiment of the present invention can carry out liquid phase random copolymerization for producing a propylene-ethylene random copolymer product.
  • the liquid phase random copolymerization temperature is 65 to 70 ° C
  • the liquid phase random copolymerization pressure is 3 to 3.8 MPaG
  • the residence time of the liquid phase random copolymerization is 35 to 45 minutes.
  • the ethylene feed amount is from 1 to 3% by weight of the liquid phase propylene. That is, the liquid phase propylene fed to the prepolymerization tank in the step (1) is used as a basis for measuring the ethylene feed amount.
  • the liquid phase homopolymerization or liquid phase random copolymerization of propylene is carried out in a liquid phase polymerization vessel with an external cooler, and vaporization of a part of the liquid phase propylene is carried out.
  • the heat of partial reaction of the propylene liquid phase homopolymerization or liquid phase random copolymerization is carried out.
  • the vaporized propylene gas can be cooled by an external cooler or compressed by a compressor to be recycled.
  • the vaporized propylene gas is condensed or compressed and returned to the propylene liquid phase homopolymerization or liquid phase random copolymerization as a first condensate or a first recycle propylene gas.
  • the liquid phase polymerizer may be a vertical stirred reactor, and the stirring of the blades makes the reaction more uniform, and at the same time, the heat transfer between the materials is enhanced to prevent the local reaction from overheating and agglomeration.
  • the gas-liquid two-phase coexistence in the polymerization kettle eliminates the need for high-power hybrid equipment, and the equipment has high production intensity and easy control of reaction conditions.
  • the polymerization kettle can also be equipped with multiple parallels, which can flexibly adjust the capacity of the device, and can also produce multi-peak polypropylene products according to the owner's requirements.
  • steps (4a) and (4b) can continue the gas phase homopolymerization and gas phase random copolymerization of propylene, respectively.
  • the reaction pressure of the propylene gas phase homopolymerization carried out in the step (4a) is lower than the reaction pressure of the propylene liquid phase homopolymerization carried out in the step (3a).
  • the polypropylene slurry obtained in the step (3a) can be continuously introduced into the step (4a) by means of the pressure difference of the two-step reaction, and the unreacted gas phase propylene in the polypropylene slurry is continuously polymerized.
  • the pressure difference may be 0.4 to 1.2 MPaG.
  • the reaction pressure of the gas phase random copolymerization carried out in the step (4b) is lower than the reaction pressure of the liquid phase random copolymerization carried out in the step (3b).
  • the copolymer slurry obtained in the step (3b) can be continuously introduced into the step (4b) by means of the pressure difference of the two-step reaction, and the unreacted gas phase propylene and ethylene in the copolymer slurry are continuously subjected to copolymerization.
  • the pressure difference may be 0.4 to 1.2 MPaG.
  • One embodiment of the present invention can perform gas phase homopolymerization of propylene to produce a propylene homopolymerized product.
  • the gas phase homopolymerization temperature of the propylene is 80 to 95 ° C
  • the gas phase homopolymerization pressure of the propylene is 2.5 to 2.8 MPaG
  • the residence time of the gas phase homopolymerization of the propylene is 45 to 60 minutes.
  • Another embodiment of the present invention can carry out gas phase random copolymerization for producing a propylene-ethylene random copolymer product.
  • the gas phase random copolymerization temperature is 80 to 95 ° C
  • the gas phase random copolymerization pressure is 2.5 to 2.8 MPaG
  • the gas phase random copolymerization residence time is 45 to 60 min.
  • the gas phase polymerization reactor is a horizontal reactor with an external cooler, and the polypropylene slurry or copolymer slurry is in the gas phase polymerization reactor.
  • the loading is 35 to 60% by volume of the gas phase polymerization reactor.
  • the heat of polymerization generated by the gas phase homopolymerization of propylene in the step (4a) and the gas phase random copolymerization in the step (4b) can be removed by vaporization of the propylene chill liquid, that is, unreacted during the above reaction.
  • the gas such as propylene and hydrogen
  • a part of the gas is distributed to the propylene recovery system (such as the recovery tower) by adjustment, and the other part is externally cooled.
  • the device After cooling, the device enters the condensate separation tank, and the propylene chilling liquid is separated and returned to the horizontal reactor through the propylene condensate pump, and then the heat of the polymerization reaction is evaporated to remove the reaction heat.
  • the unreacted propylene gas during the steps (4a) and (4b) can be recycled back to the polymerization process of steps (4a) and (4b).
  • the propylene gas separated by the condensate separation tank is recycled to the reaction process, and the polypropylene powder inside the horizontal reactor can be fluidized for heat removal and reducing the stirring power of the reactor.
  • the horizontal reactor is relatively long, and can be divided into 6 to 8 temperature control zones by segmental temperature control.
  • the equipment for circulating the fluidized powder can be determined according to the specific conditions.
  • the unreacted propylene gas discharged from the gas phase polymerization reactor is condensed or compressed and returned to the propylene vapor phase homopolymerization or the gas phase random copolymerization as a second condensate or a second recycle propylene gas.
  • the recovered propylene gas is returned to the step (2) as the recovered liquid phase propylene, and the recovered hydrogen is mixed with the fresh hydrogen as the recycled hydrogen gas to return to the step (2).
  • the fresh propylene which has reached the polymerization requirement is introduced into the propylene buffer tank 1, and then the propylene is transferred from the sole propylene outlet of the propylene buffer tank 1 to the propylene compression pump 2 and pressurized to 4 to 4.5 MPaG to obtain a liquid phase propylene having a temperature of 40 to 45 °C. .
  • the liquid phase propylene is fed into the prepolymerization tank 3 through the line 100, and at the same time, the olefin polymerization catalyst, the activator and the electron donor are mixed into the liquid phase propylene through the respective inlet ports on the line 100, and enter the prepolymerization kettle under the carrying of the liquid phase propylene. 3.
  • the prepolymerization conditions used can make the propylene raw material free from freezing and cooling, and can be fed into the propylene prepolymerization finally, and finally propylene.
  • the entire preparation process of homopolymerization or random copolymerization reduces energy consumption and propylene loss.
  • Propylene prepolymerization In the prepolymerization tank 3, liquid phase propylene is subjected to propylene prepolymerization under the action of an olefin polymerization catalyst, an activator and an electron donor.
  • the prepolymerization tank 3 is equipped with a stirrer, the prepolymerization temperature is 40 to 45 ° C, the prepolymerization pressure is 3.2 to 3.9 MPaG, and the prepolymerization residence time is about 4 to 5 minutes, and the polymerization ratio of the propylene prepolymer is about 50 to 100. Times.
  • Prepolymerization tank feed the olefin polymerization catalyst (Ti catalyst) is about 0.04 to 0.06% by weight of the liquid phase propylene, and the activator (triethyl aluminum) is about 0.2 to 0.4% by weight of the liquid phase propylene, and the electron donor (ring)
  • the hexylmethyldimethoxysilane is about 0.04 to 0.06% by weight of the liquid phase propylene.
  • Liquid phase homopolymerization a raw material slurry containing a propylene prepolymer is obtained from the prepolymerization tank 3, and is mixed with a liquid phase propylene and hydrogen gas into a liquid phase polymerization raw material through a slurry line 101, and the concentration of the polypropylene in the liquid phase polymerization raw material is 150 to 300 g/L; the amount of the recovered liquid phase propylene is 25 to 30% by weight of the liquid phase propylene; and the amount of hydrogen added is 0.04 to 0.3 kg with respect to 1000 kg of the polypropylene in the liquid phase polymerization raw material.
  • the liquid phase polymerization vessel 4 is a vertical reactor with agitation.
  • the liquid phase polymerization raw material enters the liquid phase polymerization tank 4 to carry out liquid phase homopolymerization of propylene.
  • the liquid phase homopolymerization temperature of propylene is 65-70 ° C
  • the liquid-phase homopolymerization pressure of propylene is 3 to 3.8 MPaG
  • the residence time is about 40 min.
  • the level of the liquid phase polymerization raw material in the liquid phase polymerizer 4 is controlled to 45 to 57% by volume in the liquid phase polymerization reactor 4.
  • the heat of reaction of the liquid phase homopolymerization can be carried away by vaporization of liquid phase propylene and jacketed circulating water, wherein part of the vaporized propylene gas is cooled by the propylene condenser 6, and partially returned directly to the liquid phase polymerization tank 4, partially and uncooled.
  • the propylene gas (from the liquid phase polymerizer 4 and the propylene condenser 6) is mixed and then enters the condensate separation tank 5, and the separated liquid phase propylene is returned to the liquid phase polymerization tank 4; the separated gas phase is connected through the circulation line 102.
  • the circulating fan 7 is pressurized and returned to the liquid phase of the liquid phase polymerization tank 4 to bubbling, on the one hand, the gas and the reaction liquid phase are uniformly mixed, and on the other hand, the gas is lowered in the liquid phase temperature in the middle of the polymerization vessel, and the liquid phase polymerization kettle is
  • the pressure of 4 is also controlled by the amount of external circulation cooling system;
  • Random copolymerization ethylene is introduced into the circulation line 102, and the ethylene feed amount is 5% by weight or less, preferably 1 to 3% by weight, based on the liquid phase propylene, and propylene and ethylene can be obtained in the liquid phase polymerization tank 4.
  • the liquid phase random copolymerization is carried out to produce a random copolymer product.
  • Gas phase random copolymerization a copolymer slurry (containing propylene and propylene-ethylene random copolymer) discharged from the liquid phase polymerizer 4 is passed through a valve-equipped polypropylene slurry line 103 to a gas phase polymerization reactor by a pressure difference 8.
  • the unreacted propylene entrained in the gas phase is randomly copolymerized with ethylene;
  • the gas phase random copolymerization temperature is 80 to 95 ° C, the gas phase random copolymerization pressure is 2.5 to 2.8 MPaG; and the gas phase random copolymerization residence time is about 45 ⁇ 60 min; a product containing a propylene ethylene random copolymer was obtained.
  • the level of the material in the gas phase polymerization reactor 8 can be controlled by a radioactive level gauge or current, and the level is generally controlled within 35 to 60% by volume.
  • the gas phase polymerization reactor 8 can be selected as a horizontal reactor with a stirrer 9, the material in the reactor has a uniform residence time, the equipment has high production intensity, and has strong adaptability to a slightly sticky material such as a high melting index and a copolymer;
  • the "open" type structure can be used to mix the powder evenly.
  • the heat of polymerization in the gas phase polymerization reactor 8 can be carried away by the vaporization of the propylene chilling liquid and the jacketed circulating water; the unreacted gas (mainly propylene gas, the amount of ethylene gas is completely involved in the copolymerization) passes through the upper portion of the gas phase polymerization reactor 8 After the sedimentation section settles part of the powder, a part of the gas is distributed to the propylene recovery system by adjustment, and another part of the gas is cooled by the propylene condenser 10 of the gas phase reactor and then enters the propylene condensate tank 11, and the propylene chill liquid is separated by the propylene condensate.
  • the unreacted gas mainly propylene gas, the amount of ethylene gas is completely involved in the copolymerization
  • the pump 12 is returned to the gas phase polymerization reactor 8, and the heat of absorption of the polymerization reaction is evaporated to remove the heat of reaction.
  • the gas separated from the propylene condensate tank 11 is pressurized by the propylene circulation fan 13 and sent to the bottom of the gas phase polymerization reactor 8, that is, the propylene is recycled and the polypropylene powder inside the gas phase polymerization reactor 8 is fluidized for use. Help the system to remove heat and reduce the reactor agitation power. Whether or not the circulating air system is installed can be determined according to the scale of the gas phase polymerization reactor 8.
  • the gas phase polymerization reactor 8 can adopt a sectional temperature automatic control system, and can be divided into 6 to 8 temperature control zones according to the reactor scale.
  • the agitator 9 has the function of stirring and pushing the powder product forward, and the specific stirring blade angle varies according to the size and residence time of the reactor.
  • Gas-solid separation a product containing propylene homopolymer (containing propylene gas, hydrogen, polypropylene) or a random copolymer containing propylene ethylene from a gas phase polymerization vessel discharge pipe 14 (containing propylene gas, hydrogen gas) , propylene-ethylene copolymer), by means of pressure through the outlet powder control valve into the degassing chamber 15, the product is discharged in a pulsed manner in the gas phase polymerizer discharge pipe 14.
  • the cyclone 15 is internally provided with a cyclone 40 for separating and recovering dust in the propylene gas and a bag filter; the polymer powder separated by the degassing bin 15 is dropped by gravity to the deactuator 17, in the deactivator 17 An appropriate amount of steam is introduced to deactivate the catalyst entrained in the product, and the deactivated powder enters the dryer 16 for drying and degassing to further recover propylene.
  • the dryer 16 is a horizontal indirect heating paddle stirring dryer.
  • the hollow hot shaft and the outer jacket are all connected with low pressure steam, and the wet powder is heated and dried through the wall, and the stirring shaft can also move the wet material to the wet material.
  • the dryer 16 is operated at a temperature of 100 to 105 ° C and a pressure of a slight positive pressure.
  • (C) recovery gas the propylene gas released by the wet powder in the dryer 16 is heated and washed through the filter 18 into the water washing tower 22, and the water washing tower 22 uses demineralized water as a washing medium, and the gas contains a trace amount of hydrogen chloride decomposed by the catalyst, so Hydrochloric acid in an appropriate amount of alkali liquid and water is added to the desalted water, and the washed propylene gas is cooled by a water washing tower cooler 39, and then recovered by pressurization by a propylene recovery compressor 23, and can be used for external delivery.
  • the recovered propylene gas leaving the degassing tank 15 is washed into the oil washing tower 25, and then compressed by the propylene gas buffer tank 27 and the propylene gas compressor 28 to enter the high pressure propylene scrubbing tower 29; the top of the high pressure propylene scrubbing tower 29 is separated from the top of the tower.
  • the non-condensable gas enters the dehydrogenation column 30 to remove the hydrogen rich gas, and the hydrogen rich gas is condensed and separated by the recovered propylene condenser 33 to separate the liquid phase propylene and returned to the dehydrogenation column 30; the hydrogen rich gas is mixed with the metered fresh hydrogen gas, and then sequentially
  • the circulating hydrogen buffer tank 34, the circulating hydrogen compressor 35 is pressurized and the hydrogen buffer tank 36 is sent to the liquid phase polymerization tank 4 for use; the dehydrogenation tower 30 bottom condensate is buffered into the propylene condensate tank 31, and then the propylene condensate is recovered.
  • the oil washing tower 25 is a plate tower with a washing tower condenser 38 at the top, and the propylene gas is washed with white oil containing an antistatic agent in the interior thereof to remove the aluminum alkyl entrained in the propylene gas and Impurities such as oligomers.
  • the high-pressure propylene scrubber 29 is a sieve plate rectification column with a reboiler 37 at the bottom, and a propylene condensate of the propylene condensate tank 31 is used as a reflux liquid for separating propane in the propylene to prevent the accumulation of propane in the system;
  • the dehydrogenation column 30 is a sieve tray column connected in series after the high pressure propylene scrubber column 29, and the condensed liquid of the propylene condenser 33 is recovered as a reflux liquid of the dehydrogenation column 30 to cool the propylene in the dehydrogenation column 30 for separating the propylene gas.
  • the hydrogen-rich gas is contained, and the bottom of the dehydrogenation column 30 is in direct communication with the propylene condensate tank 31, and the reflux liquid is condensed and directly enters the propylene condensate tank 31.
  • the propylene is pressurized to a liquid phase propylene having a pressure of about 4.2 MPaG and a temperature of about 42 ° C; and a liquid catalyst of propylene is used as a carrier, and a Ti catalyst (CS-1) having a content of 0.04% by weight in the liquid phase propylene is separately added.
  • CS-1 Ti catalyst having a content of 0.04% by weight in the liquid phase propylene is separately added.
  • the raw material slurry obtained in (1) is added to the recovered liquid phase propylene and hydrogen mixed into a liquid phase polymerization raw material (wherein the concentration of the polypropylene is 200 g/L, and the amount of the recovered liquid phase propylene is about 25% by weight of the liquid phase propylene, and the amount of hydrogen added is For 0.08kg/1000kg polypropylene), the liquid phase homopolymerization of propylene is carried out at 68 ° C and 3 MPaG, and the level in the reaction kettle is 45 vol%, and the residence time is 40 min;
  • part of the liquid phase propylene vaporization carries away part of the polymerization heat.
  • the vaporized propylene gas is recovered and returned to the propylene liquid phase homopolymerization in a gas phase or a liquid phase.
  • the polypropylene slurry obtained in (2) was charged into a gas phase bulk polymerization reactor, and gas phase homopolymerization of propylene was carried out at 90 ° C and 2.6 MPaG for a residence time of 45 minutes, and the material level in the reactor was 55 vol%.
  • step (2) After completion of gas phase homopolymerization of propylene, the obtained product containing propylene homopolymer is subjected to drying and gas recovery to obtain a polypropylene product; and the separated propylene recovery is continued for the liquid phase homopolymerization of propylene in step (2).
  • the hydrogen mixed with fresh hydrogen is returned as recycle hydrogen for use in step (2).
  • the unit energy consumption for producing propylene homopolymer is 42 kg of standard oil per ton of PP powder. Production of 1000 kg of polypropylene, propylene loss of 4 kg.
  • the propylene is pressurized to a liquid phase propylene having a pressure of about 4 MPaG and a temperature of about 45 ° C; and a liquid catalyst of propylene is used as a carrier, and a Ti catalyst (CS-1) having a content of 0.06% by weight in the liquid phase propylene is separately added, and 0.2 weight is added.
  • a liquid phase propylene having a pressure of about 4 MPaG and a temperature of about 45 ° C
  • a liquid catalyst of propylene is used as a carrier, and a Ti catalyst (CS-1) having a content of 0.06% by weight in the liquid phase propylene is separately added, and 0.2 weight is added.
  • the raw material slurry obtained in (1) is added to recover the mixed raw material of propylene and hydrogen into a liquid phase polymerization raw material (wherein the concentration of polypropylene is 150 g/L, the amount of recycled propylene is about 27% by weight of liquid phase propylene, and the amount of hydrogen added is 0.12 kg/ 1000kg polypropylene), liquid phase homopolymerization of propylene at 70 ° C, 3.5MPaG, the reactor level is 60% by volume, staying for 35min;
  • part of the liquid phase propylene vaporization carries away part of the polymerization heat.
  • the vaporized propylene gas is recovered and returned to the propylene liquid phase homopolymerization in a gas phase or a liquid phase.
  • the polypropylene slurry obtained in (2) was charged into a gas phase bulk polymerization reactor, and gas phase homopolymerization of propylene was carried out at 80 ° C and 2.7 MPaG for a residence time of 60 min, and the level in the reactor was 45% by volume.
  • step (2) After completion of gas phase homopolymerization of propylene, the obtained product containing propylene homopolymer is subjected to drying and gas recovery to obtain a polypropylene product; and the separated propylene recovery is continued for the liquid phase homopolymerization of propylene in step (2).
  • the hydrogen mixed with fresh hydrogen is returned as recycle hydrogen for use in step (2).
  • the unit energy consumption for producing propylene homopolymer is 45 kg of standard oil per ton of PP powder. Production of 1000 kg of polypropylene, propylene loss of 5 kg.
  • the propylene is pressurized to a liquid phase propylene having a pressure of about 4.5 MPaG and a temperature of about 40 ° C; and a liquid catalyst of propylene is used as a carrier, and a Ti catalyst (CS-1) having a content of 0.05% by weight in the liquid phase propylene is separately added.
  • CS-1 Ti catalyst having a content of 0.05% by weight in the liquid phase propylene is separately added.
  • % by weight of triethylaluminum and 0.04% by weight of cyclohexylmethyldimethoxysilane then all directly into the prepolymerization kettle to form a polymerization slurry, followed by 40 ° C, 3.8 MPaG, residence time 5 min Propylene prepolymerization, the polymerization ratio of polypropylene in the obtained raw material slurry is 100 times;
  • the raw material slurry obtained in (1) is added to recover the mixed raw material of propylene and hydrogen to form a liquid phase polymerization raw material (wherein the concentration of polypropylene is 300 g/L, the amount of recycled propylene is about 30% by weight of liquid phase propylene, and the amount of hydrogen added is 0.2 kg/ 1000kg polypropylene), liquid phase homopolymerization of propylene at 69 ° C, 3.7 MPaG, the reactor level is 40% by volume, stay 45 minutes;
  • part of the liquid phase propylene vaporization carries away part of the polymerization heat.
  • the vaporized propylene gas is recovered and returned to the propylene liquid phase homopolymerization in a gas phase or a liquid phase.
  • the polypropylene slurry obtained in (2) was charged into a gas phase bulk polymerization reactor, and gas phase homopolymerization of propylene was carried out at 95 ° C and 2.8 MPaG for a residence time of 48 minutes, and the level in the reactor was 50% by volume.
  • step (2) After completion of gas phase homopolymerization of propylene, the obtained product containing propylene homopolymer is subjected to drying and gas recovery to obtain a polypropylene product; and the separated propylene recovery is continued for the liquid phase homopolymerization of propylene in step (2).
  • the hydrogen mixed with fresh hydrogen is returned as recycle hydrogen for use in step (2).
  • the unit energy consumption for producing propylene homopolymer is 50 kg of standard oil per ton of PP powder. Production of 1000 kg of polypropylene, propylene loss of 5 kg.
  • the propylene is pressurized to a liquid phase propylene having a pressure of about 4.2 MPaG and a temperature of about 42 ° C; and a liquid catalyst of propylene is used as a carrier, and a Ti catalyst (CS-1) having a content of 0.04% by weight in the liquid phase propylene is separately added.
  • CS-1 Ti catalyst having a content of 0.04% by weight in the liquid phase propylene is separately added.
  • % by weight of triethylaluminum and 0.05% by weight of cyclohexylmethyldimethoxysilane then all directly into the prepolymerization kettle to form a polymerization slurry, followed by 42 ° C, 3.2 MPaG, residence time 4 min Propylene prepolymerization, the polymerization ratio of polypropylene in the obtained raw material slurry is 75 times;
  • the raw material slurry obtained in (1) is added to recover the mixed raw material of propylene and hydrogen to form a liquid phase polymerization raw material (wherein the concentration of polypropylene is 200 g/L, the amount of recycled propylene is about 25% by weight of liquid phase propylene, and the amount of hydrogen added is 0.08 kg/ 1000kg polypropylene).
  • part of the liquid phase propylene vaporization carries away part of the polymerization heat.
  • the vaporized propylene gas is recovered and returned to the liquid phase for random copolymerization in a gas phase or a liquid phase.
  • the copolymer slurry obtained in (2) is added to the gas phase bulk polymerization reactor, and the gas phase random copolymerization of propylene and ethylene is carried out at 90 ° C and 2.6 MPaG, the residence time is 45 min, and the material level in the reactor is 55 vol%;
  • the obtained product of the propylene-containing random copolymer is subjected to drying and gas recovery to obtain a propylene ethylene random copolymer product; the separated propylene recovery is continued for the random copolymerization of propylene in the step (2), and the separated hydrogen is mixed and fresh. Hydrogen is returned as recycle hydrogen for use in step (2).
  • the unit energy consumption for producing the propylene ethylene random copolymer is 42 kg of standard oil per ton of PP-PE powder, and 1000 kg of propylene-ethylene random copolymer is produced, and the propylene loss is 4 kg.
  • the raw material slurry was mixed with the remaining 50% by weight of liquid phase propylene (concentration of polypropylene 50% by weight, hydrogen added to 0.08 kg / 1000 kg of polypropylene) into a liquid phase polymerization raw material, and then entered into the first loop reactor.
  • a part of the propylene in the liquid phase polymerization raw material is polymerized, and the remaining liquid is used as a diluent of the polymer to make the material in the reactor slurry, and the circulation is performed by the axial flow pump, and the slurry is kept flowing at a high speed and uniformly mixed in the reactor. ;
  • the slurry in the first loop reactor was continuously fed to the second loop reactor through a discharge-only line to continue liquid phase polymerization and to replenish fresh propylene (added in an amount of 25% by weight of liquid phase propylene).
  • the first and second loop reactors have a reaction temperature of about 70 to 73 ° C, a reaction pressure of about 3.8 MPaG, and a residence time of about 1 h.
  • the polypropylene slurry discharged from the second loop reactor is flashed, degassed, dried, and deactivated to obtain a polypropylene powder.
  • the gas obtained by degassing is recovered by propylene and sent to the reaction system for reuse.
  • the unit energy consumption for producing propylene homopolymer is about 70 kg of standard oil per ton of PP powder, and 1000 kg of polypropylene is produced, and the loss of propylene is 5 kg.
  • the above-mentioned loop reactor and process are currently used in many propylene polymerization reactors and processes at home and abroad. Since the heat removal of the loop reactor is completely realized by the jacket circulating water, the flow of the slurry in the loop tube is driven by the axial flow pump. To achieve, the polypropylene slurry is liquid phase flash discharge, and the steam heating system must be added to further increase the energy consumption of the process.
  • the liquid phase After pressurizing propylene to 3.5 MPa and condensing to -5 ° C, the liquid phase enters the prepolymerization vessel, and the polymerization catalyst (including Ti catalyst (CS-1), triethyl aluminum, cyclohexylmethyl dimethoxy silane
  • the polymerization catalyst including Ti catalyst (CS-1), triethyl aluminum, cyclohexylmethyl dimethoxy silane
  • the content of each component of the catalyst in the liquid phase propylene was 0.08% by weight of Ti catalyst (CS-1), 0.5% by weight of triethylaluminum and 0.08% by weight of cyclohexylmethyldimethoxy
  • the silane is mixed with a stirrer, and the active center of the catalyst is formed, and then the prepolymerization of propylene is started.
  • the residence time of the prepolymerization is 5 min, and the polymerization ratio of the polypropylene in the obtained raw material slurry is 75 times.
  • the prepolymerized slurry containing the active catalyst and the propylene mixture was introduced into a liquid phase polymerization vessel, and the reaction was continued at 69 ° C and 3.4 MPa for 1 to 1.6 hours.
  • the concentration of polypropylene in the slurry was 130 g/L, the total amount of propylene was 10 t/h, and the amount of hydrogen added was 150 L/min.
  • the level in the liquid phase reactor was 45% by volume.
  • a polymerization catalyst was further added to the liquid phase polymerization: Ti catalyst (CS-1) 0.4 g/h, triethyl aluminum 3 L/h, and cyclohexylmethyldimethoxysilane 0.4 L/h.
  • the slurry discharged from the liquid phase polymerizer enters the gas phase reaction vessel, and is subjected to gas phase bulk polymerization at 90 ° C and 2.8 MPa, and the residence time is 1.5 hours, and the level in the gas phase reactor is 40% by volume.
  • the obtained propylene homopolymer-containing product is subjected to drying and gas recovery to obtain a polypropylene product; the separated propylene recovery is continued for the liquid phase homopolymerization of propylene in the step (2), and the separated hydrogen gas is obtained.
  • the mixed fresh hydrogen is returned as the recycled recycle hydrogen for the step (2).
  • the unit energy consumption for producing propylene homopolymer is 60 kg of standard oil per ton of PP powder. Production of 1000 kg of polypropylene, propylene loss of 6 kg.
  • the method provided by the present invention can realize the combination of propylene prepolymerization, propylene liquid phase bulk polymerization and propylene vapor phase bulk polymerization, and the propylene prepolymerization process is simplified without propylene condensation, and can be all at once.
  • the liquid phase propylene and the polymerization catalyst are fed, and the olefin polymerization catalyst can reduce the addition amount, and all the propylene can participate in the prepolymerization and carry out the propylene prepolymerization at a mild temperature of 40 to 45 ° C to obtain a propylene prepolymer dispersed better raw material.
  • Slurry, improve product quality can reduce the energy consumption per unit of propylene polymerization and propylene loss.
  • propylene needs to be condensed to below zero, and only part of the liquid phase propylene can be pre-polymerized at a low temperature; fresh propylene and a polymerization catalyst need to be added during the polymerization, and the product needs to be flash treated, and the energy consumption of the propylene polymerization unit product in the whole process And propylene loss is high.

Abstract

一种聚丙烯或丙烯乙烯共聚物的制备方法，包括：（1）将液相丙烯携带聚合催化剂进行丙烯预聚合为含有丙烯预聚物的原料浆液；预聚合温度为40～45℃，预聚合压力为3.2～3.9MPaG，烯烃聚合催化剂为液相丙烯的0.04～0.06重量％；（2）将原料浆液、回收液相丙烯和回收循环氢气混为液相聚合原料；（3a）将液相聚合原料进行丙烯液相均聚、丙烯气相均聚为含丙烯均聚物的产物；或者（3b）将液相聚合原料与乙烯进行液相无规共聚、气相无规共聚为含丙烯乙烯无规共聚物的产物，乙烯进料量为液相丙烯的5重量％以下；将产物进行气固分离，分离出聚合物和回收丙烯、氢气。该方法可以降低丙烯聚合中单位产品能耗和丙烯损耗。

Description

聚丙烯或丙烯乙烯共聚物的制备方法 技术领域
本发明涉及丙烯聚合工业化领域，具体地，涉及一种以三段式工艺进行聚丙烯或丙烯乙烯共聚物的制备方法。
背景技术
聚丙烯是以丙烯为单体聚合而成的聚合物，是通用塑料中非常重要的一个品种，其具有原料丰富、价格低廉、容易加工成型、无毒等特性，亦可通过共聚、共混、增强等措施进行改性，具有十分广阔的应用领域，在聚烯烃树脂中，是仅次于聚乙烯第二大塑料。而且丙烯还可以与乙烯进行共聚，制备丙烯-乙烯无规共聚物。
目前国内外生产聚丙烯的工艺技术主要有：液相本体法、气相本体法及液相本体与气相本体相串联的聚合方法。根据反应器形式的不同，又可分为环管法(Spheripol)、气相流化床法(Unipol)、立式搅拌釜法(Novolen、Hypol)、卧式搅拌釜法(Inoes、JPP)及不同反应釜相组合的方法。在有液相聚合的工艺中，聚合反应是在液相中发生的，为了控制反应温度和聚合度，都设有预聚，而在气相聚合的工艺中，催化剂和原料丙烯均直接加入到反应中，不设预聚。
在聚丙烯生产过程中，丙烯单体在催化剂、活化剂、给电子体的作用下，在聚合反应器中进行聚合反应得到聚丙烯浆料或固体料，生成的聚丙烯再经过闪蒸脱气、催化剂失活、干燥及氮气置换，得到聚丙烯粉料产品，在反应器出料过程中夹带的丙烯气体，经过低压洗涤、气体压缩、轻气体分离后回到反应系统回收利用，从而降低原料丙烯及氢气的消耗。
目前国内外聚丙烯工艺，各自有不同的优缺点：
间歇法本体聚合工艺：由国内自主开发的丙烯聚合工艺，其优点有流程短、投资低、见效快，其缺点为工艺落后、自动化程度低、产品质量不稳定、运行成本高、产品质量差。
Lyondell Basell的Spheripol工艺：环管反应釜，采用轴流泵强制循环，液相本体聚合，靠环管外的夹套撤热。其优点为反应条件较易控制、反应器传热系数大、单位体积产率高、产品性能稳定，牌号范围覆盖广；其缺点投资大、反应单程转化率低、有大型转动设备、能耗高、建设周期长、后处理系统复杂。
Grace的Unipol工艺：气相流化床反应器、聚合热靠气体的显热带走，气体循环量大，需要大功率的增压风机才能实现，反应釜内为全混型，产品质量不太均匀。其优点有适合乙烯和丙烯共聚、流程简单、产品物理性能好、牌号覆盖广；其缺点有能耗高、设备尺寸大、加工难度高、投资高、设备效率低。
INEOS公司的Innovene工艺：使用卧式搅拌反应器和高性能的INcat CDi催化剂，该反应器接近活塞流式反应器，采用液相丙烯汽化撤热。其优点有反应器撤热效率高、生产操作平稳、反应期内物料停留时间均匀，生产的产品熔指范围宽、能耗较低；其缺点有反应釜内容易产生热点和塑化块、反应器加工难度高、关键设备需要进口、投资高。
日本三井的Hypol工艺：液相本体和气相本体组合工艺生产聚丙烯，前两个反应釜为立式搅拌釜，后两个为立式气体流化床聚合釜；其优点可生产无规共聚和嵌段共聚产品；其缺点流程长、设备效率低、气相反应为全混型、产品质量不均匀、催化剂适应性较差、装置规模小；近几年国内没有新上的Hypol工艺装置。
但是在上述各种工艺的工业化实施过程中，均存在丙烯聚合物生产单位能耗高的缺陷，并进而影响聚合物产品的质量及稳定性，以及装置运行的安全性，增加项目工业化的投资，影响项目的经济效益。
发明内容
本发明的目的是为了解决如何降低聚丙烯生产的单位能耗，提高产品的质量和稳定性以及装置运行的安全性，减少丙烯的损耗的问题，提出聚丙烯或丙烯乙烯共聚物的制备方法，该方法耦合丙烯预聚合、丙烯液相本体聚合和丙烯气相本体聚合，可以实现丙烯的缓和预聚合，减少催化剂用量，满足丙烯聚合物产物的要求，降低整个工艺工业化实施中的单位能耗以及丙烯损耗，提高丙烯工业化生产的经济性。
为了实现上述目的，本发明提供一种聚丙烯或丙烯乙烯共聚物的制备方法，该方法包括：
(1)将含有聚合催化剂的液相丙烯全部进料到预聚合釜中进行丙烯预聚合，得到含有丙烯预聚物的原料浆液；
其中，预聚合温度为40～45℃，预聚合压力为3.2～3.9MPaG；
所述聚合催化剂包括烯烃聚合催化剂、活化剂和给电子体，所述烯烃聚合催化剂的用量为所述液相丙烯的0.04～0.06重量％；
(2)将所述原料浆液、回收液相丙烯和回收循环氢气混合为液相聚合原料；
(3a)将所述液相聚合原料送入液相聚合釜中进行丙烯液相均聚，得到聚丙烯浆 液；
(4a)将所述聚丙烯浆液连续输入气相聚合反应器中，使所述聚丙烯浆液中的丙烯进行丙烯气相均聚，得到含丙烯均聚物的产物；
或者
(3b)将所述液相聚合原料送入液相聚合釜中与乙烯进行液相无规共聚，得到共聚物浆液；其中乙烯进料量为所述液相丙烯的5重量％以下；
(4b)将所述共聚物浆液连续输入气相聚合反应器中，使所述共聚物浆液中的乙烯和丙烯进行气相无规共聚，得到含丙烯乙烯无规共聚物的产物；
(5)将所述含丙烯均聚物的产物或含丙烯乙烯无规共聚物的产物进行气固分离，分离出的固体进行干燥得到聚丙烯或者丙烯乙烯共聚物，分离出的气体回收丙烯和氢气。
优选地，在步骤(1)中，所述活化剂和给电子体的用量分别为所述液相丙烯的0.2～0.4重量％和0.04～0.06重量％。
优选地，在步骤(1)中，液相丙烯压力为4～4.5MPaG，液相丙烯温度为40～45℃。
优选地，进行所述丙烯预聚合的停留时间为4～5min。
优选地，所述丙烯预聚物为聚合倍数为50～100倍的聚丙烯。
优选地，在步骤(2)中，所述液相聚合原料中聚丙烯的浓度为150～300g/L；所述回收液相丙烯的加入量为所述液相丙烯的25～30重量％；相对于所述液相聚合原料中的1000kg聚丙烯，氢气的加入量为0.04～0.3kg。
优选地，在步骤(3a)中，丙烯液相均聚温度为65～70℃，丙烯液相均聚压力为3～3.8MPaG，进行所述丙烯液相均聚的停留时间为35～45min。
优选地，在步骤(3b)中，液相无规共聚温度为65～70℃，液相无规共聚压力为3～3.8MPaG，进行所述液相无规共聚的停留时间为35～45min。
优选地，步骤(3a)和(3b)中，所述丙烯液相均聚或液相无规共聚在带有外置冷却器的液相聚合釜中进行，并通过部分所述液相丙烯汽化的方式将所述丙烯液相均聚或液相无规共聚的部分反应热进行撤热。
优选地，汽化的丙烯气被冷凝或压缩后作为第一凝液或第一循环丙烯气返回所述丙烯液相均聚或所述液相无规共聚。
优选地，在步骤(4a)中，丙烯气相均聚温度为80～95℃，丙烯气相均聚压力为2.5～2.8MPaG；进行所述丙烯气相均聚的停留时间为45～60min。
优选地，在步骤(4b)中，气相无规共聚温度为80～95℃，气相无规共聚压力为 2.5～2.8MPaG；进行所述气相无规共聚的停留时间为45～60min。
优选地，在步骤(4a)和(4b)中，所述气相聚合反应器为带有外冷却器的卧式反应器，所述聚丙烯浆液或共聚物浆液在所述气相聚合反应器中的装量为所述气相聚合反应器的35～60体积％。
优选地，所述气相聚合反应器排出的未反应的丙烯气被冷凝或压缩后作为第二凝液或第二循环丙烯气返回所述丙烯气相均聚或所述气相无规共聚。
优选地，步骤(5)中，回收的丙烯气作为回收液相丙烯返回步骤(2)，回收的氢气混合新鲜氢气作为回收循环氢气返回步骤(2)。
通过上述技术方案，本发明提供了一种耦合了丙烯液相预聚合、液相本体聚合及气相本体聚合的三段聚合生产工艺，用于丙烯均聚或无规共聚的工业化生产。通过该工艺，可以改进丙烯预聚合的条件和进料控制，具体地，丙烯预聚合可以在40～45℃、3.2～3.9MPaG下进行，相比于现有技术的5～10℃的条件更为温和；由此在进料控制上，丙烯进料不需要冷凝至零下低温。
此外，还可以控制烯烃聚合催化剂的添加量减少，降低催化剂的使用量和成本。而且聚合催化剂的加入采取一次全部从预聚合加入，无需现有技术分批加注催化剂，既简化了操作步骤，又减少了聚合过程的波动。
再有，本发明提供的工艺可以实现液相丙烯一次进料全部参与丙烯预聚合，得到含有更好分散的聚丙烯预聚合物的原料浆液，有利于后续耦合的液相、气相聚合的均匀反应，改善聚合物产品的质量。
整个工艺过程将未反应的丙烯气回收再利用，而无需在工艺过程中另外再补充新鲜的丙烯。
本发明提供的整个工艺可以降低单位产物能耗和丙烯损耗。
经测定，产品单位能耗低于50kg标油/吨PP粉料，最低可达40kg标油/吨PP粉料，单位产品加工成本约为同规模同类型引进技术产品的1/2。
另外本发明提供的方法在实施中能避免在反应局部出现热点和爆聚的情况，减少产品出现塑化块的几率，产品颗粒不易破碎；聚合反应采用液相丙烯汽化的方式撤热，丙烯循环回收量小，系统无需大型循环设备，能有效降低装置投资，建设周期短，经济性更好。
附图说明
图1为本发明的聚丙烯的制备方法的工艺流程示意图；
图2为本发明的丙烯乙烯共聚物的制备方法的工艺流程示意图。
附图标记说明
1-丙烯缓冲罐            2-丙烯压缩泵                  3-预聚釜
4-液相聚合釜            5-凝液分离罐                  6-丙烯冷凝器
7-循环风机              8-气相聚合反应器              9-搅拌器
10-气相釜丙烯冷凝器     11-丙烯凝液罐                 12-丙烯凝液泵
13-丙烯循环风机         14-气相聚合釜出料管           100-管线
101-浆料管线            102-循环管线                  103-聚丙烯浆液管线
15-脱气仓               16-干燥器                     17-失活器
18-过滤器               19-置换釜                     20-氮气加热器
21-水封罐               22-水洗塔                     23-丙烯回收压缩机
24-水洗泵               25-油洗塔                     26-油洗泵
27-丙烯气缓冲罐         28-丙烯气压缩机               29-高压丙烯洗涤塔
30-脱氢塔               31-丙烯凝液罐                 32-回收丙烯凝液泵
33-回收丙烯冷凝器       34-循环氢气缓冲罐             35-循环氢气压缩机
36-氢气缓冲罐           37-再沸器                     38-油洗塔冷凝器
39-水洗塔冷凝器         40-旋风分离器
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。
本发明提供一种聚丙烯或丙烯乙烯共聚物的制备方法，如图1、图2所示，该方法包括：
(1)将含有聚合催化剂的液相丙烯全部进料到预聚合釜中进行丙烯预聚合，得到含有丙烯预聚物的原料浆液；
其中，预聚合温度为40～45℃，预聚合压力为3.2～3.9MPaG；
所述聚合催化剂包括烯烃聚合催化剂、活化剂和给电子体，所述烯烃聚合催化剂的用量为所述液相丙烯的0.04～0.06重量％；
(2)将所述原料浆液、回收液相丙烯和回收循环氢气混合为液相聚合原料；
(3a)将所述液相聚合原料送入液相聚合釜中进行丙烯液相均聚，得到聚丙烯浆液；
(4a)将所述聚丙烯浆液连续输入气相聚合反应器中，使所述聚丙烯浆液中的丙烯进行丙烯气相均聚，得到含丙烯均聚物的产物；
或者
(3b)将所述液相聚合原料送入液相聚合釜中与乙烯进行液相无规共聚，得到共聚物浆液；其中乙烯进料量为所述液相丙烯的5重量％以下；
(4b)将所述共聚物浆液连续输入气相聚合反应器中，使所述共聚物浆液中的乙烯和丙烯进行气相无规共聚，得到含丙烯乙烯无规共聚物的产物；
(5)将所述含丙烯均聚物的产物或含丙烯乙烯无规共聚物的产物进行气固分离，分离出的固体进行干燥得到聚丙烯或者丙烯乙烯共聚物，分离出的气体回收丙烯和氢气。
本发明中，限定丙烯预聚合在上述条件下进行，可以实现与液相本体聚合和气相本体聚合的结合，可以减少丙烯原料的冷凝、分批进料和催化剂用量，再辅助以液相本体聚合和气相本体聚合利用丙烯液化撤热和回用未反应的丙烯，以及相应控制聚合条件，整个工艺用于工业化过程实现各段聚合反应的充分进行，降低聚丙烯生产的单位能耗和减少生产过程中丙烯原料的损耗，减少丙烯聚合的生产成本。
本发明中，步骤(1)先完成液相丙烯的预聚合。在该步骤中，为实现预聚合与液相本体聚合和气相本体聚合的结合，预聚合的进料上可以实现丙烯原料以液相形式全部进料，而无需现有技术一般部分丙烯冷凝后进入预聚合而另一部分进入液相本体聚合，简化进料控制，且允许丙烯原料全部经历预聚合，形成分散更为均匀的预聚物。首先将丙烯加压液化为液相丙烯，优选地，在步骤(1)中，液相丙烯压力为4～4.5MPaG，液相丙烯温度为40～45℃。
本发明中，所述丙烯预聚合只要提供满足聚合倍数的聚丙烯即可。优选地，在步骤(1)中，进行所述丙烯预聚合的停留时间为4～5min。
本发明中，不仅全部液相丙烯均参与预聚合反应，使丙烯预聚物更好地分散，而且所述丙烯预聚物为聚合倍数为50～100倍的聚丙烯，有利于后续步骤(2)中的丙烯液相均聚和丙烯气相均聚的反应过程更平稳，有利于降低生产能耗和丙烯损耗。
本发明中，聚合催化剂可以包括烯烃聚合催化剂、活化剂和给电子体。所述烯烃聚合催化剂选自钛系丙烯聚合催化剂，例如国产牌号为CS-1的催化剂；所述活化剂选自三乙基铝；所述给电子体选自环己基甲基二甲氧基硅烷(以下简称“硅烷”)。液相丙烯可以同时作为烯烃聚合催化剂、活化剂和给电子体的载体，将烯烃聚合催化剂、活化剂及给电子体混入液相丙烯进入预聚釜中进行所述丙烯预聚合。本发明提供的方法中，采取在40～45℃、3.2～3.9MPaG下进行预聚合，所述烯烃聚合催化剂可以减少用量。优选地，在步骤(1)中，所述活化剂和给电子体的用量分别为所述液相丙烯的0.2～0.4重量％和0.04～0.06重量％。
本发明中，预聚合釜可以选用带搅拌及夹套的立式釜式反应器，亦可选用带有冷却加套的小环管，夹套面积或内冷管的换热面积需满足预聚合反应的撤热要求即可。
本发明中，在步骤(1)条件下完成的所述丙烯预聚合，得到的原料浆液可以进一步结合液相本体聚合和气相本体聚合继续进行丙烯的均聚或无规共聚，制取需要的产品。所述原料浆液中包含经预聚合而得的丙烯预聚物、未反应完的液相丙烯，以及未失去活性的烯烃聚合催化剂、活化剂和给电子体。步骤(2)可以配制进行后续液相、气相聚合的物料，可以向所述原料浆液中加入回收液相丙烯和氢气，得到所述液相聚合原料，其组成满足丙烯连续进行液相聚合和气相聚合，实现整个工艺降低聚丙烯生产的单位能耗和丙烯损耗。优选地，在步骤(2)中，所述液相聚合原料中聚丙烯的浓度为150～300g/L；所述回收液相丙烯的加入量为所述液相丙烯的25～30重量％；相对于所述液相聚合原料中的1000kg聚丙烯，氢气的加入量为0.04～0.3kg。可以调节所述回收液相丙烯和氢气的量以满足上述条件。
本发明的一种实施方式可以进行生产丙烯均聚产品的丙烯液相均聚。优选地，在步骤(3a)中，丙烯液相均聚温度为65～70℃，丙烯液相均聚压力为3～3.8MPaG，进行所述丙烯液相均聚的停留时间为35～45min。
本发明的另一种实施方式可以进行生产丙烯-乙烯无规共聚物产品的液相无规共聚。优选地，在步骤(3b)中，液相无规共聚温度为65～70℃，液相无规共聚压力为3～3.8MPaG，进行所述液相无规共聚的停留时间为35～45min。优选地，乙烯进料量为所述液相丙烯的1～3重量％。即以步骤(1)中进料到预聚合釜中的液相丙烯为乙烯进料量的计量基准。
本发明中，步骤(3a)所进行的丙烯液相均聚或者步骤(3b)所进行的液相无规共聚，在聚合过程中释放的反应热得到及时撤除有利于生产合格的丙烯均聚物或无规共 聚物。优选地，步骤(3a)和(3b)中，所述丙烯液相均聚或液相无规共聚在带有外置冷却器的液相聚合釜中进行，并通过部分所述液相丙烯汽化的方式将所述丙烯液相均聚或液相无规共聚的部分反应热进行撤热。汽化后的丙烯气可以经过外置冷却器冷却或经压缩机压缩再回收利用。优选地，汽化的丙烯气被冷凝或压缩后作为第一凝液或第一循环丙烯气返回所述丙烯液相均聚或液相无规共聚。
本发明中，所述液相聚合釜可以选用立式带搅拌反应釜，桨叶的搅拌使得反应更均匀，同时加强物料间的传热，防止局部反应过热而结块。聚合釜中为气液两相共存，无需大功率的混合动力设备，设备生产强度大，反应工况容易控制。根据生产产品或规模大小不同，聚合釜亦可设置多台并联，即能灵活调整装置产能，亦能根据业主要求生产多峰的聚丙烯产品。
本发明中，步骤(4a)和(4b)可以分别继续进行丙烯气相均聚和气相无规共聚。其中，步骤(4a)进行的所述丙烯气相均聚的反应压力，低于步骤(3a)进行的丙烯液相均聚的反应压力。可以借助两步反应的压力差实现步骤(3a)得到的所述聚丙烯浆液连续进入步骤(4a)，使所述聚丙烯浆液中未反应的气相丙烯继续聚合。所述压力差可以为0.4～1.2MPaG。同样，步骤(4b)进行的所述气相无规共聚的反应压力，低于步骤(3b)进行的液相无规共聚的反应压力。可以借助两步反应的压力差实现步骤(3b)得到的所述共聚物浆液连续进入步骤(4b)，使所述共聚物浆液中未反应的气相丙烯和乙烯继续进行共聚反应。所述压力差可以为0.4～1.2MPaG。
本发明的一种实施方式可以进行生产丙烯均聚产品的丙烯气相均聚。优选地，在步骤(4a)中，丙烯气相均聚温度为80～95℃，丙烯气相均聚压力为2.5～2.8MPaG；进行所述丙烯气相均聚的停留时间为45～60min。
本发明的另一种实施方式可以进行生产丙烯-乙烯无规共聚物产品的气相无规共聚。优选地，在步骤(4b)中，气相无规共聚温度为80～95℃，气相无规共聚压力为2.5～2.8MPaG；进行所述气相无规共聚的停留时间为45～60min。
本发明中，步骤(4a)和(4b)进行的反应中仍有未反应的丙烯，可以通过丙烯回收自循环利用。优选地，在步骤(4a)和(4b)中，所述气相聚合反应器为带有外冷却器的卧式反应器，所述聚丙烯浆液或共聚物浆液在所述气相聚合反应器中的装量为所述气相聚合反应器的35～60体积％。
本发明中，步骤(4a)的丙烯气相均聚和步骤(4b)的气相无规共聚的过程产生的聚合反应热可以采用丙烯激冷液汽化的方式撤热，即上述反应过程中未反应的气体 (如丙烯和氢气)通过反应器上部的沉降段沉降去除部分夹带出的聚合物粉料后，一部分气体通过调节被分送至丙烯回收系统(如回收塔)，另一部分气体经过外置冷却器冷却后进入凝液分离罐，分离出丙烯激冷液经过丙烯凝液泵打回到卧式反应器内，再吸收聚合反应热蒸发撤走反应热。同样，步骤(4a)和(4b)过程中未反应的丙烯气可以循环返回步骤(4a)和(4b)的聚合过程。优选地，凝液分离罐分出的丙烯气被压缩循环返回反应过程，可以对卧式反应器内部的聚丙烯粉料进行流化，用以撤热和减少反应釜搅拌功率。卧式反应器比较长，可以采用分段温控如分为6～8个控温区。根据装置规模及卧式反应器搅拌功率的不同，循环气流化粉料的相关设备可视具体情况确定是否设置。优选地，所述气相聚合反应器排出的未反应的丙烯气被冷凝或压缩后作为第二凝液或第二循环丙烯气返回所述丙烯气相均聚或所述气相无规共聚。
本发明中，优选地，步骤(5)中，回收的丙烯气作为回收液相丙烯返回步骤(2)，回收的氢气混合新鲜氢气作为回收循环氢气返回步骤(2)。
结合图1、图2对本发明作进一步详细说明。
(A)备料：
将达到聚合要求的新鲜丙烯进入丙烯缓冲罐1，然后从丙烯缓冲罐1的唯一丙烯出口将丙烯输送进丙烯压缩泵2加压至4～4.5MPaG而得到温度为40～45℃的液相丙烯。液相丙烯通过管线100输送进预聚釜3，同时，烯烃聚合催化剂、活化剂和给电子体通过管线100上的各自的加入口混入液相丙烯，在液相丙烯的携带下进入预聚釜3。
此过程中因为本发明丙烯预聚合与丙烯液相本体聚合、丙烯气相本体聚合相结合，使用的预聚合反应条件可以使丙烯原料无需冷冻降温，并且可以全部进料进过丙烯预聚合，最终丙烯均聚或无规共聚的整个制备过程降低能耗和丙烯损耗。
(B)聚合系统：
(1)丙烯预聚合：在预聚釜3中，液相丙烯在烯烃聚合催化剂、活化剂和给电子体的作用下进行丙烯预聚合反应。预聚釜3带有搅拌器，预聚合温度为40～45℃，预聚合压力为3.2～3.9MPaG，预聚合停留时间约为4～5min，得到丙烯预聚物的聚合倍数约为50～100倍。
预聚釜进料：烯烃聚合催化剂(Ti催化剂)约为液相丙烯的0.04～0.06重量％、活化剂(三乙基铝)约为液相丙烯的0.2～0.4重量％、给电子体(环己基甲基二甲氧基硅烷)约为液相丙烯的0.04～0.06重量％。
(2)丙烯液相聚合：
(a)液相均聚：由预聚釜3得到含有丙烯预聚物的原料浆液，经浆液管线101混入回收液相丙烯、氢气成为液相聚合原料，液相聚合原料中聚丙烯的浓度为150～300g/L；回收液相丙烯的加入量为液相丙烯的25～30重量％；相对于液相聚合原料中的1000kg聚丙烯，氢气的加入量为0.04～0.3kg。液相聚合釜4是带搅拌的立式反应器。
液相聚合原料进入液相聚合釜4进行丙烯液相均聚。丙烯液相均聚温度为65～70℃，丙烯液相均聚压力为3～3.8MPaG，停留时间约40min。液相聚合釜4中液相聚合原料的料位控制在液相聚合釜4的45～57体积％。液相均聚的反应热可以通过液相丙烯的汽化及夹套循环水带走，其中汽化后的部分丙烯气体经过丙烯冷凝器6冷却后，部分直接返回液相聚合釜4，部分与未冷却的丙烯气(来自液相聚合釜4和丙烯冷凝器6)混合后进入凝液分离罐5，经分离出的液相丙烯返回液相聚合釜4中；分离出的气相经过经循环管线102连通到循环风机7，增压后返回液相聚合釜4的液相中鼓泡，一方面能使气体和反应液相混合均匀，一方面使气体降低聚合釜中间的液相温度，液相聚合釜4的压力也通过外循环冷却系统的量来控制；
(b)无规共聚：在循环管线102上通入乙烯，乙烯进料量为液相丙烯的5重量％以下，优选为1～3重量％，可以实现丙烯与乙烯在液相聚合釜4中进行液相无规共聚，以生产无规共聚产品。
(3)丙烯气相聚合：
(i)气相均聚：从液相聚合釜4排出的聚丙烯浆液(含有丙烯和聚丙烯)依靠压差通过带有阀门的聚丙烯浆液管线103进入到气相聚合反应器8，其中夹带的未反应的丙烯继续聚合反应为聚丙烯；气相聚合温度为80～95℃，反应压力为2.5～2.8MPaG，聚丙烯浆液在气相聚合反应器8中的停留时间约为45～60min，得到含丙烯均聚物的产物。或者，
(ii)气相无规共聚：从液相聚合釜4排出的共聚物浆液(含有丙烯和丙烯-乙烯无规共聚物)依靠压差通过带有阀门的聚丙烯浆液管线103进入到气相聚合反应器8，其中夹带的未反应的丙烯与乙烯进行气相无规共聚；气相无规共聚温度为80～95℃，气相无规共聚压力为2.5～2.8MPaG；气相无规共聚的停留时间约为45～60min；得到含丙烯乙烯无规共聚物的产物。
气相聚合反应器8内物料的料位可以通过放射性料位计或电流来控制，料位一般控制在35～60体积％内。气相聚合反应器8可选用带搅拌器9的卧式反应器，反应器内物料停留时间均匀、设备生产强度大、对高熔指及共聚物等稍发粘的物料适应性强； 搅拌桨9可以采用“开”型结构，使粉料混合均匀。气相聚合反应器8内的聚合反应热可以通过丙烯激冷液的汽化和夹套循环水带走；未反应的气体(主要是丙烯气，乙烯气量少完全参与共聚)通过气相聚合反应器8上部的沉降段沉降部分粉料后，一部分气体通过调节被分送至丙烯回收系统，另一部分气体经过气相釜丙烯冷凝器10冷却后进入丙烯凝液罐11，分离出丙烯激冷液经过丙烯凝液泵12打回至气相聚合反应器8，再吸收聚合反应热蒸发撤走反应热。丙烯凝液罐11分离出的气体经过丙烯循环风机13增压后送至气相聚合反应器8的底部，即回用丙烯又对气相聚合反应器8内部的聚丙烯粉料进行流化，用以帮助系统撤热和减少反应釜搅拌功率。此循环风系统是否设置，可根据气相聚合反应器8的规模确定。气相聚合反应器8可以采用分段温度自动控制系统，根据反应器规模可分为6～8个控温区。搅拌器9同时具有搅拌及推动粉料产品向前移动的功能，具体的搅拌叶角度根据反应釜规模及停留时间有所不同。
(C)含丙烯均聚物的产物或含丙烯乙烯无规共聚物的产物的后处理：
(甲)气固分离：将来自气相聚合釜出料管14的含丙烯均聚物的产物(含有丙烯气、氢气、聚丙烯)或含丙烯乙烯无规共聚物的产物(含有丙烯气、氢气、丙烯-乙烯共聚物)，依靠压力经出口粉料控制阀进入脱气仓15，所述产物在气相聚合釜出料管14中为脉冲方式出料。
脱气仓15内部设置用以分离回收丙烯气中的粉尘的旋风分离器40及布袋除尘器；脱气仓15分离下来的聚合物粉料依靠重力下落至失活器17，在失活器17中通入适量的蒸汽对所述产物内夹带的催化剂进行失活，失活过的粉料进入干燥器16进行干燥脱气，进一步回收丙烯。
干燥器16是一台卧式间接加热桨叶搅拌干燥器，空心热轴及外部夹套均通入低压蒸气，通过器壁对湿粉料进行加热干燥，同时搅拌轴也能将湿物料移至物料出口，干燥器16操作温度为100～105℃，压力为微正压。
(乙)收获聚合物产品：失活、干燥后的粉料靠重力从干燥器16下落到置换釜19，在置换釜19中用氮气加热器20提供的热氮气进一步脱出粉料中夹带的极微量丙烯气体；排除的气体经水封罐21、阻火器后高空达标排放；经过脱气后的粉料通过氮气送风系统送至后续工段，得到最终的聚丙烯或丙烯-乙烯共聚物产品。
(丙)回收气体：干燥器16中湿粉料加热释放出的丙烯气体，经过滤器18进入水洗塔22洗涤，水洗塔22采用脱盐水作为洗涤介质，气体中夹带催化剂分解的极微量氯化氢，因此在脱盐水加入适量碱液中和水中的盐酸，洗涤后的丙烯气经水洗塔冷却器 39冷却后，再经丙烯回收压缩机23加压后回收，可以用于外送。
离开脱气仓15的回收丙烯气体进入油洗塔25洗涤，再经丙烯气缓冲罐27、丙烯气压缩机28压缩后进入高压丙烯洗涤塔29；高压丙烯洗涤塔29塔顶分离出的塔顶不凝气进入脱氢塔30脱除富氢气，所述富氢气经回收丙烯冷凝器33冷凝分离出液相丙烯返回脱氢塔30；所述富氢气与计量后的新鲜氢气混合，再依次经循环氢气缓冲罐34、循环氢气压缩机35加压和氢气缓冲罐36后送至液相聚合釜4利用；脱氢塔30塔底凝液进丙烯凝液罐31缓冲，然后用回收丙烯凝液泵32加压后一部分返回液相聚合釜4利用，一部分作为高压丙烯洗涤塔29的塔顶回流液；离开高压丙烯洗涤塔29塔底含有大量丙烷的丙烯(丙烷含量约19重量％)经过滤后与丙烯回收压缩机23加压的含水丙烯混合后送出界区处理。
其中，油洗塔25为一台顶部带洗油塔冷凝器38的板式塔，在其内部用含抗静电剂的白油对丙烯气体进行洗涤，用以脱出丙烯气体中夹带的烷基铝和低聚物等杂质。高压丙烯洗涤塔29是一台底部带有再沸器37的筛板精馏塔，用丙烯凝液罐31的丙烯凝液做回流液，用以分离丙烯中的丙烷，防止系统中丙烷累积；脱氢塔30是在高压丙烯洗涤塔29后串联的筛板塔，以回收丙烯冷凝器33的凝液体作为脱氢塔30的回流液冷却脱氢塔30内的丙烯，用以分离丙烯气中含有的富氢气，脱氢塔30底部与丙烯凝液罐31直连通，回流液冷凝后直接进入到丙烯凝液罐31。
以下将通过实施例对本发明进行详细描述。
实施例1
(1)预聚合
将丙烯加压得压力约4.2MPaG、温度约为42℃的液相丙烯；以液相丙烯作为载体，分别加入在液相丙烯中的含量为0.04重量％的Ti催化剂(CS-1)、0.3重量％的三乙基铝和0.05重量％的环己基甲基二甲氧基硅烷；然全部直接输入到预聚釜中形成聚合浆料，接着在42℃、3.2MPaG、停留时间4min条件下进行丙烯预聚合，得到的原料浆液中聚丙烯的聚合倍数为75倍；
(2)液相本体聚合
将(1)得到的原料浆液加入回收液相丙烯和氢气混成液相聚合原料(其中聚丙烯的浓度为200g/L，回收液相丙烯加入量约为液相丙烯的25重量％，氢气加入量为0.08kg/1000kg聚丙烯)，在68℃、3MPaG下进行丙烯液相均聚，反应釜内料位为45体积％，停留40min；
丙烯液相均聚过程中，部分液相丙烯汽化带走部分聚合反应热。汽化的丙烯气经回收，以气相或液相返回丙烯液相均聚。
(3)气相本体聚合
将(2)得到的聚丙烯浆液加入气相本体聚合反应器中，在90℃、2.6MPaG下进行丙烯气相均聚，停留时间为45min，反应器中料位为55体积％。
丙烯气相均聚过程中，排出的部分未反应的丙烯气被回收，以气相或液相返回丙烯气相均聚。
(4)完成丙烯气相均聚后，得到的含丙烯均聚物的产物进行干燥和气体回收，得到聚丙烯产品；分离出的丙烯回收继续用于步骤(2)的丙烯液相均聚，分离出的氢气混合新鲜氢气返回作为回收循环氢气用于步骤(2)。
计算上述整个工艺过程中，生产丙烯均聚物的单位能耗为42kg标油/吨PP粉料。生产1000kg聚丙烯，丙烯损耗为4kg。
实施例2
(1)预聚合
将丙烯加压得压力约4MPaG、温度约为45℃的液相丙烯；以液相丙烯作为载体，分别加入在液相丙烯中的含量为0.06重量％的Ti催化剂(CS-1)、0.2重量％的三乙基铝和0.06重量％的环己基甲基二甲氧基硅烷；然后全部直接输入到预聚釜中形成聚合浆料，接着在45℃、3.6MPaG、停留时间4min条件下进行丙烯预聚合，得到的原料浆液中聚丙烯的聚合倍数为50倍；
(2)液相本体聚合
将(1)得到的原料浆液加入回收丙烯和氢气混成液相聚合原料(其中聚丙烯的浓度为150g/L，回收丙烯加入量约为液相丙烯的27重量％，氢气加入量为0.12kg/1000kg聚丙烯)，在70℃、3.5MPaG下进行丙烯液相均聚，反应釜内料位为60体积％，停留35min；
丙烯液相均聚过程中，部分液相丙烯汽化带走部分聚合反应热。汽化的丙烯气经回收，以气相或液相返回丙烯液相均聚。
(3)气相本体聚合
将(2)得到的聚丙烯浆液加入气相本体聚合反应器中，在80℃、2.7MPaG下进行丙烯气相均聚，停留时间为60min，反应器中料位为45体积％。
丙烯气相均聚过程中，排出的部分未反应的丙烯气被回收，以气相或液相返回丙 烯气相均聚。
(4)完成丙烯气相均聚后，得到的含丙烯均聚物的产物进行干燥和气体回收，得到聚丙烯产品；分离出的丙烯回收继续用于步骤(2)的丙烯液相均聚，分离出的氢气混合新鲜氢气返回作为回收循环氢气用于步骤(2)。
计算上述整个工艺过程中，生产丙烯均聚物的单位能耗为45kg标油/吨PP粉料。生产1000kg聚丙烯，丙烯损耗为5kg。
实施例3
(1)预聚合
将丙烯加压得压力约4.5MPaG、温度约为40℃的液相丙烯；以液相丙烯作为载体，分别加入在液相丙烯中的含量为0.05重量％的Ti催化剂(CS-1)、0.4重量％的三乙基铝和0.04重量％的环己基甲基二甲氧基硅烷；然后全部直接输入到预聚釜中形成聚合浆料，接着在40℃、3.8MPaG、停留时间5min条件下进行丙烯预聚合，得到的原料浆液中聚丙烯的聚合倍数为100倍；
(2)液相本体聚合
将(1)得到的原料浆液加入回收丙烯和氢气混成液相聚合原料(其中聚丙烯的浓度为300g/L，回收丙烯加入量约为液相丙烯的30重量％，氢气加入量为0.2kg/1000kg聚丙烯)，在69℃、3.7MPaG下进行丙烯液相均聚，反应釜内料位为40体积％，停留45min；
丙烯液相均聚过程中，部分液相丙烯汽化带走部分聚合反应热。汽化的丙烯气经回收，以气相或液相返回丙烯液相均聚。
(3)气相本体聚合
将(2)得到的聚丙烯浆液加入气相本体聚合反应器中，在95℃、2.8MPaG下进行丙烯气相均聚，停留时间为48min，反应器中料位为50体积％。
丙烯气相均聚过程中，排出的部分未反应的丙烯气被回收，以气相或液相返回丙烯气相均聚。
(4)完成丙烯气相均聚后，得到的含丙烯均聚物的产物进行干燥和气体回收，得到聚丙烯产品；分离出的丙烯回收继续用于步骤(2)的丙烯液相均聚，分离出的氢气混合新鲜氢气返回作为回收循环氢气用于步骤(2)。
计算上述整个工艺过程中，生产丙烯均聚物的单位能耗为50kg标油/吨PP粉料。 生产1000kg聚丙烯，丙烯损耗为5kg。
实施例4
(1)预聚合
将丙烯加压得压力约4.2MPaG、温度约为42℃的液相丙烯；以液相丙烯作为载体，分别加入在液相丙烯中的含量为0.04重量％的Ti催化剂(CS-1)、0.3重量％的三乙基铝和0.05重量％的环己基甲基二甲氧基硅烷；然后全部直接输入到预聚釜中形成聚合浆料，接着在42℃、3.2MPaG、停留时间4min条件下进行丙烯预聚合，得到的原料浆液中聚丙烯的聚合倍数为75倍；
(2)液相本体聚合
将(1)得到的原料浆液加入回收丙烯和氢气混成液相聚合原料(其中聚丙烯的浓度为200g/L，回收丙烯加入量约为液相丙烯的25重量％，氢气加入量为0.08kg/1000kg聚丙烯)。在回收丙烯气的循环管线102上加入液相丙烯的3重量％的乙烯，在68℃、3MPaG下进行丙烯、乙烯液相无规共聚，反应釜内料位为45体积％，停留40min；
液相无规共聚过程中，部分液相丙烯汽化带走部分聚合反应热。汽化的丙烯气经回收，以气相或液相返回液相无规共聚。
(3)气相本体聚合
将(2)得到的共聚物浆液加入气相本体聚合反应器中，在90℃、2.6MPaG下进行丙烯、乙烯气相无规共聚，停留时间为45min，反应器中料位为55体积％；
丙烯、乙烯气相无规共聚过程中，排出的部分未反应的丙烯气被回收，以气相或液相返回气相无规共聚。
得到的含丙烯乙烯无规共聚物的产物进行干燥和气体回收，得到丙烯乙烯无规共聚物产品；分离出的丙烯回收继续用于步骤(2)的丙烯无规共聚，分离出的氢气混合新鲜氢气返回作为回收循环氢气用于步骤(2)。
计算上述整个工艺过程中，生产丙烯乙烯无规共聚物的单位能耗为42kg标油/吨PP-PE粉料，生产1000kg丙烯-乙烯无规共聚物，丙烯损耗为4kg。
对比例1
(1)预聚合
取-5℃、3.81MPaG的液相丙烯的50重量％与聚合催化剂混合，聚合催化剂中含有以全部液相丙烯为基准的0.08重量％的Ti催化剂(CS-1)、0.5重量％的三乙基铝和0.08 重量％的环己基甲基二甲氧基硅烷，加入预聚环管中进行低温预聚合，预聚合温度约为10℃，预聚合压力约为3.8MPaG，停留时间为12min，得到的原料浆液中聚丙烯的聚合倍数约为60倍；
(2)液相本体聚合
将原料浆料与剩下的50重量％液相丙烯(聚丙烯的浓度为50重量％，加入0.08kg/1000kg聚丙烯的氢气)混合为液相聚合原料，然后进入第一环管反应器，液相聚合原料中的一部分丙烯发生聚合，剩余的液体作为聚合物的稀释剂使反应器中的物料呈淤浆状，通过轴流泵打循环，在反应器中保持淤浆高速流动和混合均匀；
再通过出料专用线将第一环管反应器中的淤浆连续送至第二环管反应器中继续进行液相聚合并补充新鲜丙烯(加入量为液相丙烯的25重量％)。第一、二环管反应器反应温度约为70～73℃，反应压力约为3.8MPaG，停留时间约1h。
从第二环管反应器排出的聚合淤浆经过闪蒸、脱气、干燥、失活后得到聚丙烯粉料。脱气得到的气体经过丙烯回收后送反应系统回用。
通过计算上述环管丙烯聚合过程，生产丙烯均聚物的单位能耗约为70kg标油/吨PP粉料，生产1000kg聚丙烯，丙烯损耗为5kg。
上述环管反应器及工艺是目前国内外采用较多的丙烯聚合反应器及工艺，由于环管反应器撤热全部由夹套循环水来实现，浆料在环管内流动是由轴流泵推动来实现，聚丙烯浆料为液相闪蒸出料，须增加蒸汽加热系统，进一步增加了该工艺反应过程的能耗。
对比例2
将丙烯加压达到3.5MPa并冷凝至-5℃后液相进入预聚合聚釜内，与聚合催化剂(包括Ti催化剂(CS-1)、三乙基铝、环己基甲基二甲氧基硅烷)在0℃接触，催化剂各组分在液相丙烯中的含量为0.08重量％的Ti催化剂(CS-1)、0.5重量％的三乙基铝和0.08重量％的环己基甲基二甲氧基硅烷，并用搅拌器捏物料混匀，在此生成催化剂的活性中心，然后开始丙烯的预聚，预聚合的停留时间为5min，得到的原料浆液中聚丙烯的聚合倍数为75倍。
预聚合得到的含有活性催化剂和丙烯混合物淤浆进入液相聚合釜内，在69℃、3.4MPa下停留1～1.6h继续反应。淤浆中聚丙烯浓度为130g/L、丙烯总量为10t/h、氢气加入量为150L/min。液相反应釜内料位为45体积％。
液相聚合中还补加聚合催化剂：Ti催化剂(CS-1)0.4g/h、三乙基铝3L/h、环己基 甲基二甲氧基硅烷0.4L/h。
液相聚合釜排出的浆料进入气相反应釜，在90℃、2.8MPa下进行气相本体聚合，停留时间为1.5h，气相反应釜中料位为40体积％。
完成丙烯气相均聚后，得到的含丙烯均聚物的产物进行干燥和气体回收，得到聚丙烯产品；分离出的丙烯回收继续用于步骤(2)的丙烯液相均聚，分离出的氢气混合新鲜氢气返回作为回收循环氢气用于步骤(2)。
计算上述整个工艺过程中，生产丙烯均聚物的单位能耗为60kg标油/吨PP粉料。生产1000kg聚丙烯，丙烯损耗为6kg。
通过上述实施例和对比例的结果可以看出，本发明提供的方法可以实现丙烯预聚合、丙烯液相本体聚合和丙烯气相本体聚合的结合，丙烯预聚合过程简化而无需丙烯冷凝，可以一次全部液相丙烯和聚合催化剂进料，且烯烃聚合催化剂可以降低加入量，可以全部丙烯都参与预聚合并在40～45℃的温和温度下进行丙烯预聚合，得到丙烯预聚物分散更好的原料浆液，提高产品质量，可以降低丙烯聚合的单位产品能耗和丙烯损耗。
对比例1中需要丙烯冷凝至零下，且只能部分液相丙烯进行低温预聚合；聚合过程中需要补加新鲜丙烯和聚合催化剂，产物需闪蒸处理，整个工艺过程的丙烯聚合单位产品能耗和丙烯损耗高。
对比例2中，现有技术需要丙烯冷凝至零下，并且预聚合温度低，催化剂加入量高，也要补加聚合催化剂，完成整个工艺过程的丙烯聚合单位产品能耗和丙烯损耗高。
以上结合附图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 (11)

  1. 一种聚丙烯或丙烯乙烯共聚物的制备方法，该方法包括：
    (1)将含有聚合催化剂的液相丙烯全部进料到预聚合釜中进行丙烯预聚合，得到含有丙烯预聚物的原料浆液；
    其中，预聚合温度为40～45℃，预聚合压力为3.2～3.9MPaG；
    所述聚合催化剂包括烯烃聚合催化剂、活化剂和给电子体，所述烯烃聚合催化剂的用量为所述液相丙烯的0.04～0.06重量％；
    (2)将所述原料浆液、回收液相丙烯和回收循环氢气混合为液相聚合原料；
    (3a)将所述液相聚合原料送入液相聚合釜中进行丙烯液相均聚，得到聚丙烯浆液；
    (4a)将所述聚丙烯浆液连续输入气相聚合反应器中，使所述聚丙烯浆液中的丙烯进行丙烯气相均聚，得到含丙烯均聚物的产物；
    或者
    (3b)将所述液相聚合原料送入液相聚合釜中与乙烯进行液相无规共聚，得到共聚物浆液；其中乙烯进料量为所述液相丙烯的5重量％以下；
    (4b)将所述共聚物浆液连续输入气相聚合反应器中，使所述共聚物浆液中的乙烯和丙烯进行气相无规共聚，得到含丙烯乙烯无规共聚物的产物；
    (5)将所述含丙烯均聚物的产物或含丙烯乙烯无规共聚物的产物进行气固分离，分离出的固体进行干燥得到聚丙烯或者丙烯乙烯共聚物，分离出的气体回收丙烯和氢气。
  2. 根据权利要求1所述的方法，其中，在步骤(1)中，所述活化剂和给电子体的用量分别为所述液相丙烯的0.2～0.4重量％和0.04～0.06重量％。
  3. 根据权利要求1所述的方法，其中，在步骤(1)中，液相丙烯压力为4～4.5MPaG，液相丙烯温度为40～45℃；
    优选地，进行所述丙烯预聚合的停留时间为4～5min；
    优选地，所述丙烯预聚物为聚合倍数为50～100倍的聚丙烯。
  4. 根据权利要求1所述的方法，其中，在步骤(2)中，所述液相聚合原料中聚 丙烯的浓度为150～300g/L；所述回收液相丙烯的加入量为所述液相丙烯的25～30重量％；相对于所述液相聚合原料中的1000kg聚丙烯，氢气的加入量为0.04～0.3kg。
  5. 根据权利要求1-4中任意一项所述的方法，其中，在步骤(3a)中，丙烯液相均聚温度为65～70℃，丙烯液相均聚压力为3～3.8MPaG，进行所述丙烯液相均聚的停留时间为35～45min。
  6. 根据权利要求1-4中任意一项所述的方法，其中，在步骤(3b)中，液相无规共聚温度为65～70℃，液相无规共聚压力为3～3.8MPaG，进行所述液相无规共聚的停留时间为35～45min。
  7. 根据权利要求1-4中任意一项所述的方法，其中，步骤(3a)和(3b)中，所述丙烯液相均聚或液相无规共聚在带有外置冷却器的液相聚合釜中进行，并通过部分所述液相丙烯汽化的方式将所述丙烯液相均聚或液相无规共聚的部分反应热进行撤热；
    优选地，汽化的丙烯气被冷凝或压缩后作为第一凝液或第一循环丙烯气返回所述丙烯液相均聚或所述液相无规共聚。
  8. 根据权利要求1-4中任意一项所述的方法，其中，在步骤(4a)中，丙烯气相均聚温度为80～95℃，丙烯气相均聚压力为2.5～2.8MPaG；进行所述丙烯气相均聚的停留时间为45～60min。
  9. 根据权利要求1-4中任意一项所述的方法，其中，在步骤(4b)中，气相无规共聚温度为80～95℃，气相无规共聚压力为2.5～2.8MPaG；进行所述气相无规共聚的停留时间为45～60min。
  10. 根据权利要求1-4中任意一项所述的方法，其中，在步骤(4a)和(4b)中，所述气相聚合反应器为带有外冷却器的卧式反应器，所述聚丙烯浆液或共聚物浆液在所述气相聚合反应器中的装量为所述气相聚合反应器的35～60体积％；
    优选地，所述气相聚合反应器排出的未反应的丙烯气被冷凝或压缩后作为第二凝液或第二循环丙烯气返回所述丙烯气相均聚或所述气相无规共聚。
  11. 根据权利要求1-4中任意一项所述的方法，其中，步骤(5)中，回收的丙烯气作为回收液相丙烯返回步骤(2)，回收的氢气混合新鲜氢气作为回收循环氢气返回步骤(2)。
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