JP5072637B2 - Propylene polymer production method - Google Patents

Description

  The present invention relates to the production of homopolymers based on propylene and copolymers having a high comonomer content and impact-modified propylene polymers. In particular, the present invention relates to a method for producing propylene polymer in a reactor system comprising a combination of at least one slurry reactor and at least one gas phase reactor.

  The softness, impact resistance, and heat sealability of the propylene-based polymer can be improved by copolymerizing propylene with other olefins such as ethylene and isobutylene. Both bulk and gas phase processes are used. However, the comonomer used in the polymerization causes swelling of the polymer in the bulk polymerization medium. As a result, when swollen soft polymer particles are flushed after polymerization, the particle morphology is destroyed and the bulk density of the powdered polymer is greatly reduced. At the same time, amorphous material is deposited on the surface of the powder. The sticky low density material easily agglomerates on the walls of the flash tank, causing problems during transportation. Increasing the proportion of comonomer increases these problems.

  For these reasons, in the prior art, polymerization is mainly performed using a gas phase method. These processes have been proposed for the production of sticky but flowable products (EP 0237 003) and rubber products such as EPR and EPDM (EP 0614917). In the process, the gas velocity of the fluidized bed reactor is sufficient to separate the particles and act as a fluid. However, the polymer in the fluidized bed reactor is essentially carried in plug flow mode.

  The gas phase method is also preferred for high comonomer content products. See EP 06749991, EP 058574, EP 0605001, and EP 0704464.

  However, problems with gas phase reactors are caused by long residence times, which means long transition times and potential production losses. This is especially true in multireactor processes. Catalyst productivity in the gas phase process is low, which means higher catalyst usage and higher production costs.

  In order to benefit from the different advantages of slurry bulk reactors and gas phase reactors, several combinations of bulk and gas phase reactors for producing propylene copolymers have been proposed in the art. However, at present, none of the prior art methods meet the requirements of flexibility and low manufacturing cost required for the production of many different polyolefin qualities using the same process arrangement. In particular, the circulation of a significant amount of unreacted monomer to the slurry reactor, which is a general feature of known processes, impairs loop reactor dynamics and delays the transition to new product quality.

  An improved two-stage process for the polymerization of propylene in a combination loop and gas phase reactor is disclosed in US Pat. No. 4,740,550. The main purpose of US Pat. No. 4,740,550 is to provide a method for producing high quality block copolymers by feeding the homopolymer to the block copolymerization stage with a narrow residence time distribution. The disclosed method comprises the following stages: a first stage consisting of homopolymerization in a bulk loop reactor; a second stage homopolymerization in a gas phase reactor; in a cyclone between the first and second stages Fine material removal; and finally, impact copolymerization in an addition gas phase reactor.

  Prior to feeding the polymerization product of the loop reactor to the gas phase, the fines fraction is removed and recycled to the loop reactor. Along with the fine material, a portion of the monomer from the gas phase reactor is circulated directly to the first stage loop reactor.

  There are several significant problems with this prior art. For example, if all the fine material is removed from the reactor outlet of the loop reactor and recycled to the loop reactor, the loop reactor will eventually become an inactive catalyst or slightly polymerized dead fine material (dead There is a significant risk that is met with fines). On the other hand, inhomogeneity problems in the final product arise when a portion of the fine material stream is combined with the product from the final reactor. Furthermore, as proposed in US Pat. No. 4,740,550, a complicated and economically unacceptable operation is required when a portion of the fine material stream is collected separately and blended with the separated homopolymer.

The object of the present invention is to solve the problems associated with prior art single and multi-reactor processes and to provide a new process for producing propylene homo and copolymers.
Another object of the present invention is to provide a very versatile process that can be used to produce a wide variety of propylene (co) polymer products.

  These and other objects, as well as advantages over known methods apparent from the following description, are achieved by the present invention described below.

  The method of the present invention is based on a combination of at least one slurry reactor and at least one gas phase reactor connected in series in that order to form a cascade. Propylene (co) polymer is produced in the presence of a catalyst at elevated temperature and pressure. In accordance with the present invention, the polymerization product of at least one slurry reactor containing unreacted monomer is passed to the first gas phase reactor with or without minimal circulation of monomer to the slurry reactor. Carry. In the context of the present invention, a high quality impact copolymer is produced by a two-stage homopolymerization followed by an impact copolymerization step without fine material removal and circulation after the first and second stages of copolymerization. It has been found that it can. In the present invention, the amount of circulation can be minimized by using a specific arrangement of reactors and by selecting the relative amounts produced in each reactor in view of its purpose.

  According to another aspect of the invention, at least one slurry reactor and at least one gas phase reactor connected in series are used as a reactor system, wherein the at least one slurry reactor is at an elevated temperature or supercritical temperature. The contents of the slurry reactor comprising a reaction medium containing a copolymer product and unreacted monomer are gasified using a conduit connecting the slurry reactor to the gas phase reactor. Delivered directly to the fluidized bed in the phase reactor.

  In particular, the method according to the invention mainly has the features described in the characterizing part of claim 1.

  The present invention has many important advantages. With the arrangement of the present invention, it has been found that the monomer fed to the first reactor is largely or completely consumed in the gas phase reactor after the slurry reactor. This is possible by gas phase operation where a small amount of gas is discharged along with the polymer product. Loop reactor dynamics in the cascade provide fast transitions and high productivity. Fast start-up is also possible because the gas phase layer material is available directly from the loop reactor. Loops and gas phase reactor cascades can provide a wide variety of broad molecular weight distributions or bimodal products. The at least one gas phase reactor provides high flexibility in the reaction rate ratio of the first and second parts of the product due to the adjustable layer level and reaction rate. In addition, gas phase reactors that do not have solubility limitations make it possible to produce high or very high comonomer content polymers.

  The loop-gas phase reactor combination significantly reduces residence time and production loss compared to the gas-gas multi-reactor method.

Definitions In the present invention, “slurry reactor” means a reactor, such as a continuous or simple stirred tank reactor or loop reactor, operated in bulk or slurry and the polymer is formed in particulate form. “Bulk” means polymerization in a reaction medium comprising at least 60% by weight of monomers. According to a preferred embodiment, the slurry reactor comprises a bulk loop reactor.

  “Gas phase reactor” means a mechanical mixing or fluidized bed reactor. The gas phase reactor preferably comprises a mechanically stirred fluidized bed reactor having a gas velocity of at least 0.2 m / sec.

  “High temperature polymerization” means a polymerization temperature above the 80 ° C. limit temperature known to be detrimental to related prior art high yield catalysts. At high temperatures, catalyst stereospecificity and polymer powder morphology can be lost. Such is not the case with the particularly preferred types of catalysts used in the present invention described below. High temperature polymerization occurs at temperatures above the limiting temperature and below the corresponding critical temperature of the reaction medium.

  “Supercritical polymerization” means polymerization that occurs above the corresponding critical temperature and pressure of the reaction medium.

  “Direct feed” means the process in which the contents of the slurry reactor comprising the polymerization product and the reaction medium are conveyed directly to the fluidized bed of the gas phase reactor.

  “Reaction zone” means several reactors of the same type, connected in series or in series, that produce polymers of the same type or nature.

  The expressions “essentially no monomer circulation” and “minimum or zero monomer circulation” mean that about 30% by weight or less, preferably 20% by weight or less, in particular zero monomer is recycled to the slurry process. means. In contrast, in a normal slurry process, 50 weight percent or more monomer is circulated.

Overall Method The present invention comprises a bulk reaction zone having at least one slurry reactor, and at least one gas phase reactor in a cascade after the at least one slurry reactor, with minimal or zero monomer being the first reaction. Relates to a multi-stage process which is circulated in a vessel and fed directly or indirectly to the gas phase for the homo- or copolymerization of propylene.

  In the direct feed process, the contents of the slurry reactor, the polymerization product and the reaction medium are conveyed directly to the fluidized bed reactor. The product output of the slurry reactor is discontinuous or preferably continuous. The gas or particle stream is not separated based on various particle sizes, but the slurry is carried as it is. Do not return the particles to the previous reactor. Optionally, the line between the slurry reactor and the gas phase reactor can be heated to evaporate some or all of the reaction medium prior to entering the gas phase reactor layer.

  The reaction is continued in the gas phase reactor. All and substantially all (at least 90%) of the monomer entering the gas phase reactor from the slurry reactor is part of the reactor gas inventory until it is converted to polymer.

  In the operation of the two reactors, the polymer exiting from the gas phase reactor with the exit system enters the solid / gas separation unit. The polymer from the bottom is fed to a further processing stage, the gas is compressed and recycled to the gas phase reactor after the purification stage. In general, light inert materials such as methane and ethane, and heavy inert materials such as propane and oligomers are removed in these purification steps. Purification can be carried out by distillation or semipermeable membrane separation. In the case of distillation, the monomer is mainly circulated as a liquid in the gas phase reactor.

  In the operation of the three reactors, the polymer exiting from the first gas phase reactor with the exit system enters the solid / gas separation unit. The polymer from the bottom is further fed to the second gas phase reactor, the gas is compressed and circulated to the first gas phase reactor after the purification step. In general, light inert materials such as methane and ethane, and heavy inert materials such as propane and oligomers are removed in these purification steps. Purification can be carried out by distillation or semipermeable membrane separation. In the case of distillation, the monomer is mainly circulated as a liquid in the gas phase reactor.

  In the operation of the three reactors, optionally, the polymer exiting from the first gas phase reactor having the outflow system can enter the second gas phase reactor directly with the accompanying gas.

  In the operation of the three reactors, the polymer exiting from the second gas phase reactor with the exit system enters the solid / gas separation unit. The polymer from the bottom is fed to a further processing stage, the gas is compressed, partly directly and partly recycled to the second gas phase reactor after the purification stage. In general, light inert materials such as methane and ethane, and heavy inert materials such as propane and oligomers are removed in these purification steps. Purification can be carried out by distillation or semipermeable membrane separation. For distillation, a high ethylene content stream is circulated to the second gas phase reactor and the propylene-propane stream is fed to the propane and oligomer removal stage.

  A polymerization product is obtained by using a catalyst. The catalyst can be any catalyst that provides adequate activity at elevated temperatures. The preferred catalyst system used comprises a high yield Ziegler-Natta catalyst containing catalyst components, a cocatalyst component, an external donor (donor), and optionally an internal donor. Other preferred catalyst systems are, for example, metallocene based catalysts which have a bridged ligand structure conferring high stereoselectivity and are impregnated on a support or support in the form of an active complex.

  The polymerization temperature is at least 60 ° C, preferably at least 65 ° C. The slurry reactor is operated at a high pressure of at least 35 bar to 100 bar and the gas phase reactor is operated at a pressure of at least 10 bar to dew point. Alternatively, any reactor in a series reactor can be operated above the critical temperature and pressure.

In a plurality of polymerization reactors connected in series, propylene and optionally one or more C _{2 to} C ₁₆ olefins such as ethylene, 1-butene, 4-methyl-1-pentene, 3-methyl-1 -Butene, 1-hexene, 1-octene, 1-decene, diene or cyclic olefins such as vinylcyclohexane or cyclopentene are subjected to polymerization and copolymerization, respectively. Comonomer olefins can be used in any of the reactors. Various amounts of hydrogen can be used as molecular weight modifiers or regulators in any or all reactors.

  The desired propylene (co) polymer can be recovered from the product separation means in the gas phase reaction zone.

Catalyst A polymerization product is obtained by using a catalyst. As catalysts, stereospecific catalysts for propylene can be used that have high yields and useful polymer properties such as isotacticity and morphology at high temperatures and possible supercritical polymerization. The preferred catalyst system used comprises a high yield Ziegler-Natta catalyst containing a catalyst component, a cocatalyst component, optionally an external donor and an internal donor. Another preferred catalyst system is a metallocene catalyst having an active complex having a bridged ligand structure that imparts high stereoselectivity and impregnated on a support. Finally, the catalyst is preferably a catalyst that imparts adequate activity at elevated temperatures.

  Examples of suitable systems are disclosed, for example, in FI patents 86866, 96615, 88047, 88048, and 88049.

  One particularly preferred catalyst that can be used in the present invention is disclosed in FI Patent No. 88047. Other preferred catalysts are disclosed in FI Patent Application No. 963707.

  Other preferred catalysts are disclosed in PCT / FI97 / 00191 and PCT / FI97 / 00192.

Prepolymerization The catalyst can be subjected to prepolymerization before being fed to the first polymerization reactor in series. Prior to feeding to the reactor, the catalyst component is contacted with a monomer, such as an olefin monomer, during prepolymerization. An example of a suitable system is disclosed, for example, in FI patent application FI961152.

  Prepolymerization can also be performed in the presence of a viscous material such as an olefin wax to provide a prepolymerization catalyst that is stable during storage and handling. The catalyst prepolymerized in wax facilitates the loading of the catalyst into the polymerization reactor. An example of a suitable system is disclosed, for example, in FI Patent 95387. In general, about 1 part of catalyst is used for up to 4 parts of polymer.

  The monomers used for the prepolymerization consist of propylene, 1-butene, 4-methyl-1-pentene, 3-methyl-1-butene, vinylcyclohexane, cyclopentene, 1-hexene, 1-octene and 1-decene. You can choose from a group.

  The prepolymerization can be carried out batchwise in wax, in a continuous initial polymerization reactor, or in a continuous plug flow type prepolymerization reactor.

Polymerization The present invention is based on a combination of at least one slurry reactor and at least one gas phase reactor connected in series, called a cascade.

  The polymerization stage apparatus comprises a suitable type of polymerization reactor. A slurry reactor is a continuous or simple stirred tank reactor or loop reactor operated in bulk or slurry, where the polymer is formed in particulate form in the reactor. Bulk means polymerization in a reaction medium comprising at least 60% (w / w) monomer. The gas phase reactor is a mechanical mixing or fluidized bed reactor. According to the invention, the slurry reactor is preferably a bulk loop reactor and the gas phase reactor is a fluidized bed reactor having a mechanical stirrer.

  The reactor in the process may be a supercritical polymerization reactor.

  The production split between the slurry reactor and the first gas phase reactor is generally 67:33 to 50:50 when monomer circulation to the slurry reactor is performed. On the other hand, if circulation to the slurry reactor is not required, the production volume in the slurry reactor is less than or equivalent to the production volume in the first gas phase reactor. In all cases, the production in the slurry reactor is greater than 10%. Thus, in a preferred embodiment, 10-70 wt%, preferably 20-65 wt%, especially 40-60 wt% of polymer is produced in the slurry reaction zone and no monomer is circulated to the slurry reaction zone. If 50% to 67% polymer is produced in the slurry reaction zone, a small amount of monomer can be circulated from the gas phase reaction zone to the slurry reactor.

According to the invention, the polymerization process comprises at least the following steps:
Subjecting propylene and optionally other olefins to polymerization or copolymerization in the first slurry polymerization zone or reactor;
Recovering the first polymerization product from the first reaction zone with the reaction medium;
Feed the first polymerization product directly or indirectly to the gas phase polymerization zone or reactor;
Optionally supplying additional propylene and / or comonomers to the second reaction zone;
Excess propylene and / or comonomer from the first zone and additional propylene and / or comonomer are subjected to a second polymerization reaction in the presence of the first polymerization product to obtain a second polymerization product;
• recovering the polymerization product from the second reaction zone; and • separating and recovering the polypropylene from the second reaction product.
Further, the method can comprise one or more of the following additional steps:
Prepolymerizing the catalyst with one or more monomers;
Separating the gas from the product of the second reaction zone;
Feed the initial zone recovered polymerization product to the third or fourth reaction zone or reactor;
Optionally supplying additional propylene and / or comonomers to the third and fourth reaction zones;
Excess propylene and / or comonomer and additional propylene and / or comonomer are subjected to a third and fourth polymerization reaction in the presence of the polymerization product in the initial zone to obtain a third or fourth polymerization product;
• recovering the polymerization product from the third or fourth reaction zone; and • separating and recovering the polypropylene from the third or fourth reaction product.

  In the first stage of the process, propylene and optional comonomers are fed to the first polymerization reactor along with the active catalyst complex and optional cocatalyst and other auxiliary components. The catalyst can be prepolymerized or prepolymerized before being fed to the process. Along with the above components, hydrogen can be fed into the reactor as a molecular weight regulator in the amount necessary to obtain the desired molecular weight of the polymer. In embodiments where there is no circulation to the slurry reactor, only fresh monomer is fed to the first reactor.

  Alternatively, in embodiments where minimal recycle of monomer to the slurry reactor occurs, the feed to the reactor is recycled monomer from the next reactor passed through the recovery system, if present. New monomer, hydrogen, optional comonomer, and catalyst components.

  In all embodiments, the active catalyst complex polymerizes in the presence of propylene, optional comonomers, cocatalysts and other auxiliary components and is suspended in particulate form in the slurry reactor, i.e., in the fluid circulated to the reactor. The product is formed in the form of turbid polymer particles.

  The polymerization medium generally comprises monomers and optionally hydrocarbons and the fluid is a liquid or a gas. In the case of a slurry reactor, in particular a loop reactor, the fluid is a liquid and a suspension of polymer is continuously circulated to the slurry reactor, whereby the polymer in particulate form in a hydrocarbon medium or monomer. More suspension is formed. According to a preferred embodiment, the first polymerization or copolymerization reaction is carried out in a reaction medium mainly consisting of propylene. It is preferred that at least 60% by weight of the medium is propylene.

  The conditions of the slurry reactor are selected such that at least 10%, preferably at least 12% by weight of the total product is polymerized in the first slurry reactor. The temperature is from 40 ° C. to 110 ° C., preferably from 50 ° C. to 100 ° C., more preferably from 80 ° C. to 100 ° C. for homopolymers and highly random copolymers, and from 60 ° C. for copolymers with a high comonomer content. 75 ° C. The reaction pressure is 30 to 100 bar, preferably 35 to 80 bar, based on the vapor pressure of the reaction medium.

  In the slurry polymerization zone, two or more reactors can be used in series. In such a case, the inert hydrocarbon or polymer suspension in the monomer formed in the first slurry reactor is operated at a lower pressure than the slurry reactor without being separated from the inert components and monomers. The slurry reactor is periodically or continuously fed.

  The heat of polymerization is removed by cooling the reactor using a cooling jacket. In order to obtain a sufficient degree of polymerization, the residence time in the slurry reactor is at least 10 minutes, preferably 20 to 100 minutes. This is necessary to obtain a polymer yield of at least 40 kg PP / g catalyst. When the particles swell, it is also beneficial to operate the slurry reactor at a high solids concentration, for example 50% for homopolymers and 35% or 40% for copolymers. If the solids concentration in the loop reactor is too low, the amount of reaction medium delivered to the second reaction zone or gas phase reactor is increased.

  The contents of the slurry reactor, the polymerization products and the reaction medium are conveyed directly to the gas phase reactor fluidized bed.

  The second reactor is preferably a gas phase reactor, in which propylene and optionally comonomer are polymerized in a reaction medium consisting of gas or steam.

  The gas phase reactor may be a conventional fluidized bed reactor, but other types of gas phase reactors can also be used. In a fluidized bed reactor, the layer consists of polymer particles formed or about to be formed, as well as still active catalyst and polymer parts. By introducing a gas component, for example a monomer at a flow rate (at least 0.2 m / sec), which causes the particles to act as a fluid, the bed is maintained in a fluid state. The flowing gas can also contain an inert gas such as nitrogen and hydrogen as a modifier. In the present invention, it is not recommended to use unnecessary inert gas that may cause problems in the recovery section.

  The gas phase reactor used can be operated at a temperature of 50 ° C. to 115 ° C., preferably 60 ° C. to 110 ° C., and a reaction pressure of 10 bar to 40 bar, with a partial pressure of monomer of 2 bar to 30 Bar or higher is preferred and is always below the dew point pressure.

  According to one preferred embodiment, no fresh propylene is fed to the first gas phase reactor except as required for various flushings.

  The pressure of the second polymerization product containing the gaseous reaction medium is then reduced after the first gas phase reactor, eg, in a flash tank, with the product gas and possible volatile components (eg heavy Separate the portion of the comonomer and the compound used in the catalyst feed). The overhead gas stream is circulated by the recovery system to the first gas phase reactor, or part to the first gas phase reactor and part to the slurry reactor. Some monomer, generally heavier comonomer, can be recycled to the bulk reaction zone.

  If desired, the polymerization product can be fed to a second gas phase reactor and subjected to a third polymerization reaction to obtain a modified polymerization product, from which polypropylene is separated and recovered. To do. The third polymerization reaction can be conducted in a gas phase reactor in the presence of a comonomer that imparts properties such as improved impact resistance, ductility, or softness to the third polymerization product. In general, some of the gas flowing in from the first gas phase reactor is removed in a decompression stage before the second gas phase reactor. The removed gas is compressed and transported to the recovery section and processed as described above in the case of two reactors. Alternatively, the second product can be delivered directly to the third reactor.

  In general, when copolymers are produced by the process of the present invention, they contain at least 0.5% by weight of comonomer, in particular at least about 2% by weight, preferably at most 20% by weight of at least one comonomer. The typical comonomer content of the copolymer fed to the first gas phase reactor is about 2 to 16% by weight. The copolymer produced exhibits high randomness (very soft copolymer).

  In particular, to produce a rubbery copolymer by a third (co) polymerization reaction, the polymerization product is fed to a second gas phase reactor to obtain a modified polymerization product. The third polymerization reaction imparts properties such as improved impact strength to the polymerization product. The step of producing the elastomer can be performed by various methods. Therefore, it is preferable to produce an elastomer by copolymerizing at least propylene and ethylene to obtain an elastomer. The conditions for copolymerization are in the range of normal EPM production conditions as described, for example, in Encyclopedia of Polymer Science and Engineering, 2nd edition, volume 6, pages 545-558. A rubber-like product is formed when the ethylene repeating unit content in the polymer is within a predetermined range. Accordingly, it is preferred to copolymerize ethylene and propylene to form an elastomer in such a proportion that the copolymer contains 10-70 wt% ethylene units. In particular, the ethylene unit content is 30-50% by weight of the copolymer propylene / ethylene elastomer. In other words, ethylene and propylene are polymerized to obtain an elastomer having an ethylene / propylene molar ratio of 30/70 to 50/50.

  Elastomers can also be obtained by adding ready-made or natural elastomers to the polymer product of the first gas phase reactor.

  Impact modified polypropylene generally contains about 5-50% by weight of the elastomer, especially about 10-45% by weight, preferably about 15-40% by weight.

  Generally, to produce a high molecular weight product to obtain improved impact resistance, the hydrogen concentration of the second reaction product is reduced before feeding the product to the second gas phase.

  It is also possible to carry the product of the second gas phase reaction to a third (such as fourth) polymerization reaction zone where the copolymerization is carried out in the presence of a comonomer that imparts properties to the third polymerization modified product.

  The third and fourth gas phase reactors can be operated at a temperature of 60 ° C. to 80 ° C., and the reaction pressure can be maintained at 10-30 bar.

In summary, one particularly preferred embodiment of the present invention (FIG. 1) is:
-Polymerizing propylene in a loop reactor at a pressure of 40-80 bar, a temperature of 80-100C, and using hydrogen to adjust the molecular weight of the polymerization product;
-Recovering the polymerization product from the loop reactor and transporting the product to the gas phase reactor fluidized bed;
Optionally feeding additional propylene and optional comonomers to the gas phase reactor;
Optionally supplying additional hydrogen to the gas phase reactor to adjust the hydrogen / propylene ratio to obtain a polymerization product of the desired molecular weight;
-Recovering the polymerization product from the gas phase reactor, transporting the product to a flash tank, where the product pressure is reduced, overhead product essentially containing unreacted propylene and hydrogen, and polymerization Obtaining a bottom product mainly containing solids;
-Circulating the overhead product or at least a majority of it through the recovery section to the gas phase reactor; and-recovering the polypropylene polymer as the bottom product of the flash tank;
Comprising that.
According to a second particularly preferred embodiment (FIG. 1):
-Propylene and copolymers, such as ethylene or 1-butene or both, are polymerized in a loop reactor at a pressure of 40-80 bar, a temperature of 60 ° C-80 ° C, and using hydrogen and desired Obtaining a molecular weight polymerization product;
-Carrying the polymerization product from the loop reactor directly to the gas phase reactor fluidized bed;
Optionally feeding additional propylene and comonomers to the gas phase reactor;
Optionally supplying additional hydrogen to the gas phase reactor to adjust the hydrogen / propylene ratio to obtain a polymerization product of the desired molecular weight;
-Recovering the polymerization product from the gas phase reactor and transporting it to a flash tank where the pressure is reduced, mainly containing overhead products essentially containing unreacted monomer and hydrogen, and polymerization solids Obtaining a bottom product;
Circulate the overhead product, or at least most of it, through the recovery section to the gas phase reactor; and-recover the polypropylene polymer as the bottom product of the flash tank.
According to a third particularly preferred embodiment (FIG. 2),
-Propylene and optionally the copolymer are polymerized in a loop reactor at a pressure of 40-80 bar, a temperature of 60C-100C, and hydrogen is used to adjust the mass of the polymerization product;
-Recovering the polymerization product from the loop reactor and transporting it to the gas phase reactor fluidized bed;
Optionally supplying additional propylene and additional comonomer to the gas phase reactor;
Optionally supplying additional hydrogen to the gas phase reactor to adjust the hydrogen / propylene ratio to obtain a polymerization product of the desired molecular weight;
-The polymerization product from the first gas phase reactor is recovered and transported to an intermediate flash tank, where the product pressure is reduced, overhead product essentially containing unreacted monomer and hydrogen, and polymerization Obtaining a bottom product mainly containing solids;
Circulating the overhead product, or at least a majority of it, through the recovery section to the first gas phase reactor; and-polypropylene polymer from the bottom of the intermediate flash tank to the third polymerization reactor via the polymer feed system. Supply;
The third polymerization reaction is carried out in the presence of a comonomer in a gas phase reactor;
-Recovering the polymerization product from the second gas phase reactor and transporting it to a flash tank, where the product pressure is reduced, overhead product essentially containing unreacted monomer and hydrogen, and polymerization Obtaining a bottom product mainly containing solids;
-Optionally, the polymerization product from the third polymerization can be transported directly or via a flash tank to a third (such as fourth) gas phase polymerization reactor in the presence of a comonomer. Polymerization takes place.

The two preferred embodiments are also shown in the drawing, which shows the specific arrangement of the process equipment used. The numbers indicate the following parts of the device.
1; 101 Prepolymerization reactor 30; 130 Catalyst reservoir 31; 131 Feed device 32; 132 Diluent (optional)
33; 133 catalyst / diluent mixture 34; 134 monomer 35; 135 catalyst and possible donor 40; 140 loop reactor 42; 142 diluent supply (optional)
43; 143 monomer supply 44; 144 hydrogen supply 45; 145 comonomer supply (optional)
46; 146 Loop reactor through line 46; 146
40; return to 140 47; 147 one or several discharge valves 150b flash separator 152b take-off line 60; 160; 160b gas phase reactor 61; 161; 161b gas transfer line 62; 162; 162b compressor 63; 163; 163b monomer feed 64; 164; 164b comonomer feed 65; 165; 165b hydrogen feed 66; 166; 166b transfer line 67; 167 product transfer line 68; 168 polymer product recovery system, eg flash
Tank 69; 169 Recovery line 70; 170 Monomer recovery system

  According to FIG. 1, the catalyst from the reservoir 30 is fed to the feeding device 31 along with any diluent from the line 32. Feed device 31 feeds the catalyst / diluent mixture to polymerization chamber 1 via line 33. Monomer can be fed by 34 and the catalyst and possible donor can be fed to reactor 1 by conduit 35 or, preferably, the catalyst and donor are mixed and fed by line 35.

  The prepolymerized catalyst is preferably discharged directly from the polymerization chamber 1 via line 36 and carried to the loop reactor 40. In loop reactor 40, the polymerization is continued by adding via line 46 any diluent from line 42, monomer from line 43, hydrogen from line 44, and any comonomer from line 45. Any catalyst can also be introduced into the loop reactor 40.

  The polymer-hydrocarbon mixture is fed from the loop reactor 40 through one or several discharge valves 47 as described in FI patent application Nos. 971368 or 971367. There is a direct product transfer 67 from the loop reactor 40 to the gas phase reactor 60.

  From the upper part of the reactor 60, through the line 61, the compressor 62, and a heat exchanger (not shown), the gas removed in the usual way is circulated to the lower part of the reactor 60 in the usual way. A fluidized bed consisting of polymer particles that are maintained in a fluidized state is present at the bottom of the gas phase reactor 60. Although it is not essential, it is advantageous to attach a mixer to reactor 60 (described in FI Patent Application No. 933073, not shown). Monomer from line 63, optionally comonomer from line 64, and hydrogen from line 65 can be introduced into the lower portion of reactor 60 in a known manner. Product is withdrawn continuously or periodically from reactor 60 through transfer line 66 to flash tank 68. The overhead product of the recovery system is circulated to the gas phase reactor via the monomer recovery system.

  The embodiment shown in FIG. 2 differs from FIG. 1 only in that the product from the gas phase reactor 160 is fed to an additional gas phase reactor 160b. Polymer particles are transported from flash tank 168 and polymer supply tank 150b to gas phase reactor 160b via extraction line 152b. A mixer (not shown) is preferably attached to the gas phase reactor.

  The overhead of the flash separator 168b is circulated partly to the gas phase reactor 160b and partly to the monomer recovery system.

  In both of the above embodiments, the numbers 70 and 170 have a circulating monomer in the gas phase reactor (60, 160, 160b) or separator (68, 168, 168b) generally having a lower boiling point than hydrogen and / or the monomer. It means a separation means such as a semipermeable membrane unit or a removal column that can be liberated from light inert hydrocarbons.

Polymer The product produced according to the present invention comprises a polypropylene copolymer including a polypropylene terpolymer. In particular, very soft, highly random copolymers can be produced by the present invention. The copolymer contains at least 0.5% by weight of comonomer, in particular at least about 2% by weight, preferably up to 20% by weight of comonomer. A typical comonomer content is about 2-12% by weight. An essential feature of the present invention is the use of high polymerization temperatures, preferably above 75 ° C., which give a more homogeneous comonomer distribution during copolymerization. The randomness measured by FTIR at a polymerization temperature of 60 ° C. is 69%, 71% at 65 ° C., and 74 at a polymerization temperature of 75 ° C. in the first reactor and 80 ° C. in the second reactor. %.
Other products produced by the present invention include impact modified propylene polymers, preferably containing rubbery copolymers, particularly ethylene-propylene copolymers, that improve the impact resistance of the product. The proportion of elastomer is about 5-40% by weight of propylene.

  The following examples illustrate the principles of the present invention.

A production scale facility for continuous production of PP homopolymer was simulated. The equipment includes a catalyst, alkyl, donor, propylene feed system, prepolymerization reactor, loop reactor and fluidized bed gas phase reactor (GPR).
Catalyst, alkyl, donor and propylene were fed to the prepolymerization reactor. The polymer slurry obtained from the prepolymerization reactor was fed to the loop reactor, and hydrogen and further propylene were also fed to the loop reactor. The polymer slurry obtained from the loop reactor and additional hydrogen and propylene were fed to the GPR. The production rate in these reactors was 300 kg / hour in the prepolymerization, 15 t / hour in the loop and 10 t / hour in the GPR.
The prepolymerization loop reactor was operated at a pressure of 56 bar and a temperature of 20 ° C. The loop reactor was operated at a pressure of 55 bar and a temperature of 85 ° C. The PP homopolymer MFR (2.16 kg, 230 ° C.) produced in the loop reactor was adjusted to 1 by controlling the hydrogen feed rate.
The GPR was operated at a pressure of 35 bar and a temperature of 85 ° C. The PP homopolymer MFR (2.16 kg, 230 ° C.) removed from the GPR was adjusted to 13 by controlling the hydrogen partial pressure. 5 t / hr of propene was recycled from the GPR outlet back to the loop reactor. The conversion of propylene after one treatment was 83%.

A production scale facility for the continuous production of PP copolymers with good impact properties was simulated. The equipment includes a catalyst, alkyl, donor, propylene feed system, prepolymerization reactor, loop reactor and two fluidized bed gas phase reactors (GPR) (see FIG. 2).
Catalyst, alkyl, donor and propylene were fed to the prepolymerization reactor. The polymer slurry obtained from the prepolymerization reactor was fed to the loop reactor, and hydrogen and further propylene were also fed to the loop reactor. The polymer slurry obtained from the loop reactor and additional hydrogen and propylene were fed to the first GPR.
Prior to entering the second GPR, the pressure of the polymer obtained from the first GPR was released. Ethylene and additional propylene were fed to the second GPR.
The production rates in these reactors were 300 kg / hour in the prepolymerization, 15 t / hour in the loop, 10 t / hour in the first GPR, and 6 t / hour in the second GPR.
The prepolymerization loop reactor was operated at a pressure of 56 bar and a temperature of 20 ° C. The loop reactor was operated at a pressure of 55 bar and a temperature of 85 ° C. The PP homopolymer MFR (2.16 kg, 230 ° C.) produced in the loop reactor was adjusted to 20 by controlling the hydrogen feed rate.
The first GPR was operated at a pressure of 35 bar and a temperature of 85 ° C. The MFR (2.16 kg, 230 ° C.) of the PP homopolymer removed from the first GPR was set to 20 by controlling the hydrogen partial pressure. 4.3 t / hr of propene was recycled from the GPR outlet back to the loop reactor.
The second GPR was operated at a pressure of 20 bar and a temperature of 70 ° C. The MFR (2.16 kg, 230 ° C.) of the PP copolymer removed from the second GPR was adjusted to 13 by using hydrogen partial pressure as the control means. 2.7 t / hr of propene was recycled from the outlet of the second GPR back to the loop reactor and 1.6 t / hr of ethylene was recycled to the second GPR.

A production scale facility for continuous production of random PP polymers was simulated. Equipment includes catalysts, alkyls, donors, propylene and ethylene feed systems, prepolymerization reactors, loop reactors and fluidized bed gas phase reactors (GPR).
Catalyst, alkyl, donor and propylene were fed to the prepolymerization reactor. The polymer slurry obtained from the prepolymerization reactor was fed to the loop reactor. Ethylene, hydrogen and further propylene were also fed to the loop reactor. The polymer slurry obtained from the loop reactor and additional hydrogen ethylene and propylene were fed to the GPR. The production rate in these reactors was 300 kg / hour during prepolymerization, 15 t / hour in the loop reactor and 10 t / hour in GPR.
The prepolymerization reactor was operated at a pressure of 56 bar and a temperature of 20 ° C. The loop reactor was operated at a pressure of 55 bar and a temperature of 75 ° C. The random PP MFR (2.16 kg, 230 ° C.) produced in the loop reactor was adjusted to 7 by controlling the hydrogen feed, and the ethylene content was adjusted to 3.5% w / w by the ethylene feed.
The GPR was operated at a pressure of 35 bar and a temperature of 80 ° C. The MFR (2.16 kg, 230 ° C.) of random PP extracted from GPR was adjusted to 7 by controlling the hydrogen partial pressure, and the ethylene content was adjusted to 3.5% w / w by adjusting the ethylene partial pressure. 5 t / hr propene and 33 kg / hr ethylene were recycled from the GPR outlet back to the loop reactor. The conversion of propylene and ethylene in one treatment was 83% and 96%, respectively.

A production scale facility for continuous production of PP copolymers with good impact and creep properties was simulated. The equipment includes a catalyst, alkyl, donor, ethylene and propylene feed system, prepolymerization reactor, loop reactor, flash tank and two fluidized bed gas phase reactors.
Catalyst, alkyl, donor and propylene were fed to the prepolymerization reactor. The polymer slurry obtained from the prepolymerization reactor was fed to the loop reactor, and hydrogen and further propylene were also fed to the loop reactor. The polymer slurry obtained from the loop reactor was fed to a flash tank where propylene and polymer were separated.
The polymer obtained from the flash tank was fed to the first GPR. The propylene obtained from the flash tank was fed to the first GPR after removing the hydrogen. Ethylene and additional propylene were fed to the first GPR. The polymer obtained from the first GPR was fed to the second GPR. Ethylene, some hydrogen, and additional propylene were fed to the second GPR.
The production rates in these reactors were 300 kg / hour in the prepolymerization, 10 t / hour in the loop reactor, 10 t / hour in the first GPR, and 6 t / hour in the second GPR. It was.
The prepolymerization reactor was operated at a pressure of 56 bar and a temperature of 20 ° C. The loop reactor was operated at a pressure of 55 bar and a temperature of 85 ° C. The PP homopolymer MFR (2.16 kg, 230 ° C.) produced in the loop reactor was set to 100 by adjusting the hydrogen feed rate.
The GPR was operated at a pressure of 35 bar and a temperature of 80 ° C. The GPR PP MFR (2.16 kg, 230 ° C.) was set to 0.4 by controlling the production split between reactors and the hydrogen removal efficiency from the evaporation propene. The ethylene content was set to 2% w / w by adjusting the ethylene partial pressure and controlling the production split between the reactors.
The second GPR was operated at a pressure of 20 bar and a temperature of 70 ° C. The MFR (2.16 kg, 230 ° C.) of the PP copolymer removed from the second GPR was adjusted to 0.3 by controlling the hydrogen partial pressure and controlling the production split between the reactors. A small amount of propylene was circulated back to the loop reactor from the outlet of the second GPR.

A production scale facility for the continuous production of PP copolymers with good creep properties was simulated. Equipment includes catalysts, alkyls, donors, ethylene and propylene feed systems, prepolymerization reactors, loop reactors, flash tanks and fluidized bed gas phase reactors.
Catalyst, alkyl, donor and propylene were fed to the prepolymerization reactor. The polymer slurry obtained from the prepolymerization reactor was fed to the loop reactor, and ethylene and further propylene were also fed to the loop reactor. The polymer slurry obtained from the loop reactor was fed to a flash tank where the monomer and polymer were separated.
The polymer obtained from the flash tank was fed to the GPR. The propylene obtained from the flash tank was fed to the GPR after removing the ethylene. Hydrogen and additional propylene were fed to the GPR.
The production rate in these reactors was 300 kg / hour during the prepolymerization, 15 t / hour in the loop reactor and 10 t / hour in the first GPR.

The PP homopolymer was produced using a continuously operated pilot plant. The equipment includes a catalyst, alkyl, donor, propylene feed system, prepolymerization reactor, loop reactor and fluidized bed gas phase reactor (GPR).
Catalyst, alkyl, donor and propylene were fed to the prepolymerization reactor. The polymer slurry obtained from the prepolymerization reactor was fed to the loop reactor, which was also fed with hydrogen and further propylene. The polymer slurry obtained from the loop reactor and additional hydrogen and propylene were fed to the GPR.
The product polymer and unreacted propylene were separated after removing the polymerization product from the GPR.
The catalyst used was a highly active and highly stereospecific ZN (Ziegler-Natta) catalyst made according to US Pat. No. 5,234,879. The catalyst was contacted with triethylaluminum (TEA) and dicyclopentyldimethoxysilane (DCPDMS) before feeding to the prepolymerization reactor (Al / Ti ratio is 250, Al / Do is 40 (mol)). there were).
The catalyst was fed according to U.S. Pat. No. 5,385,992 and placed in a prepolymerization reactor with propylene. The prepolymerization reactor was operated at a pressure of 51 bar, a temperature of 20 ° C. and an average residence time of the catalyst of 7 minutes.
The prepolymerized catalyst propylene and other components were transferred to the loop reactor. The loop reactor was operated at a pressure of 50 bar, a temperature of 80 ° C., and an average catalyst residence time of 1 hour. The PP homopolymer MFR (2.16 kg, 230 ° C.) produced in the loop reactor was adjusted to 7 by using the hydrogen feed rate as the control means.
The polymer slurry obtained from the loop reactor was transferred to GPR. The GPR reactor was operated at a total pressure of 29 bar and a propylene partial pressure of 21 bar. The temperature was 90 ° C. and the average residence time of the catalyst was 1 hour. The MFR (2.16 kg, 230 ° C.) of the PP homopolymer removed from the GPR apparatus was 7, and was controlled by adjusting the hydrogen partial pressure. The production split between reactors was 1% in the prepolymerization, 49% in the loop unit, and 50% in GPR. The catalyst productivity was 32 kg PP / g catalyst.

The PP homopolymer was produced using a continuously operated pilot plant. The equipment includes a catalyst, alkyl, donor, propylene feed system, prepolymerization reactor, loop reactor and fluidized bed gas phase reactor (GPR).
Catalyst, alkyl, donor and propylene were fed to the prepolymerization reactor. The polymer slurry obtained from the prepolymerization reactor was fed to the loop reactor, which was also fed with hydrogen and further propylene. The polymer slurry obtained from the loop reactor and additional hydrogen and propylene were fed to the GPR.
The product polymer and unreacted propylene were separated after removal from the GPR.
The catalyst used was a highly active and highly stereospecific ZN catalyst made according to Finnish Patent Application No. 963707. The catalyst was contacted with triethylaluminum (TEA) and dicyclopentyldimethoxysilane (DCPDMS) before feeding to the prepolymerization reactor (Al / Ti ratio is 250, Al / Do is 40 (mol)). there were).
The catalyst was fed according to U.S. Pat. No. 5,385,992 and placed in a prepolymerization reactor with propylene. The prepolymerization reactor was operated at a pressure of 53 bar, a temperature of 20 ° C., and an average residence time of 7 minutes.
The prepolymerized catalyst propylene and other components were transferred to the loop reactor. The loop reactor was operated at a pressure of 52 bar, a temperature of 85 ° C., and an average residence time of catalyst of 1 hour. The PP homopolymer MFR (2.16 kg, 230 ° C.) produced in the loop reactor was adjusted to 7 by controlling the hydrogen feed rate.
The polymer slurry obtained from the loop reactor was transferred to GPR. The GPR reactor was operated at a total pressure of 29 bar and a propylene partial pressure of 21 bar. The temperature of GPR was 80 ° C., and the average residence time of the catalyst was 1 hour. The MFR (2.16 kg, 230 ° C.) of the PP homopolymer removed from the GPR apparatus was 7, and was controlled by adjusting the hydrogen partial pressure. The production split between reactors was 1% in the prepolymerization, 53% in the loop unit, and 48% in GPR. The catalyst productivity was 50 kg PP / g catalyst.

The PP homopolymer was produced using a continuously operated pilot plant. The equipment includes a catalyst, alkyl, donor, propylene feed system, prepolymerization reactor, loop reactor and fluidized bed gas phase reactor (GPR).
Catalyst, alkyl, donor and propylene were fed to the prepolymerization reactor. The polymer slurry obtained from the prepolymerization reactor was fed to the loop reactor, which was also fed with hydrogen and further propylene. The polymer slurry obtained from the loop reactor and additional hydrogen and propylene were fed to the GPR.
The product polymer and unreacted propylene were separated after removing the product from the GPR.
The catalyst used was a highly active and highly stereospecific ZN catalyst made according to US Pat. No. 5,234,879. The catalyst was contacted with triethylaluminum (TEA) and dicyclopentyldimethoxysilane (DCPDMS) before feeding to the prepolymerization reactor (Al / Ti ratio is 250, Al / Do is 40 (mol)). there were).
The catalyst was fed according to U.S. Pat. No. 5,385,992 and placed in a prepolymerization reactor with propylene. The prepolymerization reactor was operated at a pressure of 58 bar, a temperature of 20 ° C., and an average residence time of 7 minutes.
The prepolymerized catalyst propylene and other components were transferred to the loop reactor.
The loop reactor was operated at a pressure of 57 bar, a temperature of 80 ° C. and a catalyst average residence time of 2 hours. The PP homopolymer MFR (2.16 kg, 230 ° C.) produced in the loop reactor was set to 375 depending on the hydrogen feed rate.
The polymer slurry obtained from the loop reactor was transferred to GPR. The GPR was operated at a total pressure of 29 bar and a propylene partial pressure of 16 bar. The reactor temperature was 80 ° C. and the average residence time of the catalyst was 2 hours. The MFR (2.16 kg, 230 ° C.) of the PP homopolymer removed from the GPR is 450, which was adjusted by controlling the hydrogen partial pressure and controlling the production split between the reactors. The production split between the reactors was adjusted to be 1% in the prepolymerization, 50% in the loop reactor and 49% in GPR.

PP random polymers were produced using a continuously operated pilot plant. The equipment includes catalysts, alkyls, donors, propylene and ethylene feed systems, loop reactors and fluidized bed gas phase reactors (GPR).
Catalyst, alkyl, donor and propylene were fed to the prepolymerization reactor. The polymer slurry obtained from the loop reactor and additional hydrogen, propylene and ethylene were fed to the GPR. The product polymer and unreacted propylene were separated after removal from the GPR.
The catalyst used was a highly active and highly stereospecific ZN catalyst made according to US Pat. No. 5,234,879. This catalyst was prepolymerized with propylene in batch mode according to Finnish Patent No. 95387 (PP / catalyst weight ratio was 10). The prepolymerized catalyst was contacted with triethylaluminum (TEA) and dicyclopentyldimethoxysilane (DCPDMS) before feeding to the loop reactor (Al / Ti ratio is 140, Al / Do is 10 (mol). Met).
The catalyst was fed according to US Pat. No. 5,385,992 and placed in a loop reactor with propylene. The loop reactor was operated at a pressure of 50 bar, a temperature of 75 ° C., and an average residence time of catalyst of 1 hour. The PP random polymer MFR (2.16 kg, 230 ° C.) produced in the loop reactor was adjusted to 4 by the hydrogen feed rate. The ethylene content was controlled to be 3.5% w / w by the ethylene feed rate.
The polymer slurry obtained from the loop reactor was transferred to GPR. The GPR reactor was operated at a total pressure of 29 bar and a propylene partial pressure of 21 bar. The operating temperature of GPR was 80 ° C., and the average residence time of the catalyst was 1.5 hours. The MFR (2.16 kg, 230 ° C.) of PP random polymer removed from GPR was adjusted to 4 by hydrogen partial pressure. The ethylene content was controlled to be 3.5% w / w by the ethylene feed rate. The production split between reactors was 55% in the loop reactor and 45% in GPR.

PP random polymers were produced using a continuously operated pilot plant. The equipment includes catalysts, alkyls, donors, propylene and ethylene feed systems, loop reactors and fluidized bed gas phase reactors (GPR).
Catalyst, alkyl, donor and propylene were fed to the prepolymerization reactor. The polymer slurry obtained from the loop reactor and additional hydrogen and propylene were fed to the GPR. The product polymer and unreacted propylene were separated after removal from the GPR.
The catalyst used was a highly active and highly stereospecific ZN catalyst made according to US Pat. No. 5,234,879. The catalyst was prepolymerized with propylene (PP / catalyst weight ratio was 10), batchwise, according to Finnish patent 95387. The prepolymerized catalyst was contacted with triethylaluminum (TEA) and dicyclopentyldimethoxysilane (DCPDMS) before feeding to the loop reactor (Al / Ti ratio is 135, Al / Do is 10 (mol). Met).
The catalyst was fed according to US Pat. No. 5,385,992 and placed in a loop reactor with propylene. The loop reactor was operated at a pressure of 50 bar, a temperature of 75 ° C., and an average residence time of catalyst of 1 hour. The PP random polymer MFR produced in the loop reactor (2.16 kg, 230 ° C.) was set to 0.2 by adjusting the hydrogen feed rate. The ethylene content was 3.5% w / w, which was adjusted by controlling the ethylene feed.
The polymer slurry obtained from the loop reactor was transferred to GPR. The GPR reactor was operated at a total pressure of 29 bar and a propylene partial pressure of 21 bar. The operating temperature was 80 ° C. and the average residence time of the catalyst was 1.5 hours. The PP random polymer MFR (2.16 kg, 230 ° C.) removed from the GPR was adjusted to 3 by controlling the hydrogen partial pressure. The ethylene content was set to 1.8% w / w by adjusting the production split between the reactors. The desired ethylene content was obtained when the production split between reactors was 40% in the loop unit and 60% in GPR.
The prepolymerization reactor was operated at a pressure of 56 bar and a temperature of 20 ° C. The loop reactor was operated at a pressure of 55 bar and a temperature of 75 ° C. The MFR (2.16 kg, 230 ° C.) of the random PP produced in the loop reactor was less than 0.1, and the ethylene content was adjusted to 3.5% w / w by controlling the ethylene feed rate. .
The GPR reactor was operated at a pressure of 35 bar and a temperature of 80 ° C. The MFR (2.16 kg, 230 ° C.) of the PP copolymer extracted from GPR was 0.3, which was adjusted by the hydrogen partial pressure. The ethylene content was set to 1.8% w / w by adjusting the production split between the reactors.
The ethylene in the loop discharge was recovered from the flash gas and returned to the loop reactor for circulation. Propylene in the GRP effluent was recovered and fed to the loop reactor after removing the hydrogen. The conversion of propylene and ethylene in one treatment was 83% and 84%, respectively.

A continuously operated pilot plant was used to produce PP copolymers with good impact and creep properties. The equipment includes a catalyst, alkyl, donor, propylene and ethylene feed system, a prepolymerization reactor, a loop reactor and two fluidized bed gas phase reactors (GPR).
Catalyst, alkyl, donor and propylene are fed to the prepolymerization reactor. The polymer slurry obtained from the prepolymerization reactor is fed to the loop reactor, which is also fed with hydrogen, ethylene and further propylene.
The polymer slurry obtained from the loop reactor and additional hydrogen and propylene are fed to the first GPR. A polymer obtained from the first GPR is fed to the second GPR. Ethylene, some hydrogen, and additional propylene were fed to the second GPR. The product polymer and unreacted propylene are separated after removal from the second GPR.
The catalyst used is a highly active and highly stereospecific ZN catalyst made according to US Pat. No. 5,234,879. The catalyst is contacted with triethylaluminum (TEA) and dicyclopentyldimethoxysilane (DCPDMS) before feeding to the prepolymerization reactor (Al / Ti ratio is 150, Al / Do is 10 (mole)). ).
The catalyst is fed according to US Pat. No. 5,385,992 and placed in a loop reactor with propylene. The prepolymerization reactor is operated at a pressure of 51 bar, a temperature of 20 ° C., and an average residence time of the catalyst of 7 minutes.
The loop apparatus is operated at a pressure of 50 bar, a temperature of 75 ° C., and an average catalyst residence time of 1 hour. The MFR (2.16 kg, 230 ° C.) of the PP random polymer produced in the loop reactor is set to 7 by controlling the hydrogen feed rate. The ethylene content is adjusted to 3.5% w / w by using ethylene feed as a control means.
The polymer slurry obtained from the loop reactor is transferred to the first GPR. The first GPR reactor is operated at a total pressure of 29 bar and a propylene partial pressure of 21 bar. The operating temperature is 80 ° C. and the average residence time of the catalyst is 1.5 hours. The MFR (2.16 kg, 230 ° C.) of the PP random polymer removed from the GPR is adjusted to 10 by using hydrogen partial pressure. The ethylene content is set to 2% w / w by adjusting the production split between the reactors.
The polymer obtained from the first GPR is transferred to the second GPR. The second GPR is operated at a total pressure of 10 bar and a monomer partial pressure of 7 bar. The operating temperature is 80 ° C. and the average residence time of the catalyst is 1.5 hours. The MFR (2.16 kg, 230 ° C.) of the PP copolymer removed from the GPR is adjusted to 7 by hydrogen partial pressure. The ethylene content is set to 10% w / w by adjusting the ethylene partial pressure and controlling the production split between reactors.
The desired properties are that the production split between reactors is 1% in the prepolymerization, 40% in the loop reactor, 40% in the first GPR, and 19% in the second GPR. Sometimes obtained.

A very soft PP copolymer was produced using a pilot plant operated continuously. Equipment includes catalysts, alkyls, donors, propylene and ethylene feed systems, prepolymerization reactors, loop reactors and fluidized bed gas phase reactors (GPR).
Catalyst, alkyl, donor and propylene were fed to the prepolymerization reactor. The polymer slurry obtained from the prepolymerization reactor was fed to the loop reactor, and hydrogen, ethylene and further propylene were also fed to the loop reactor.
The polymer slurry obtained from the loop reactor and additional ethylene, hydrogen and propylene were fed to the GPR. The product polymer and unreacted propylene are separated after removal from the GPR.
The catalyst used was a highly active and highly stereospecific ZN catalyst made according to US Pat. No. 5,234,879. The catalyst was contacted with triethylaluminum (TEA) and dicyclopentyldimethoxysilane (DCPDMS) before feeding to the prepolymerization reactor (Al / Ti ratio was 150, Al / Do was 10 (mole)). there were).
The catalyst was fed according to US Pat. No. 5,385,992 and placed in a loop reactor with propylene. The prepolymerization reactor was operated at a pressure of 51 bar, a temperature of 20 ° C. and an average residence time of the catalyst of 7 minutes.
The loop reactor was operated at a pressure of 50 bar, a temperature of 75 ° C., and an average residence time of catalyst of 1 hour. The PP random polymer MFR (2.16 kg, 230 ° C.) produced in the loop reactor was controlled at 4 by the hydrogen feed rate. The ethylene content was adjusted to 3.8% w / w by controlling the ethylene feed.
The polymer slurry obtained from the loop reactor was transferred to the first GPR. The first GPR reactor was operated at a total pressure of 29 bar and a propylene partial pressure of 21 bar. The operating temperature was 80 ° C. and the average residence time of the catalyst was 1.2 hours. The PP random polymer MFR (2.16 kg, 230 ° C.) removed from the GPR was set to 2.5 by adjusting the hydrogen partial pressure. The ethylene content was set to 8% w / w by adjusting the production split between reactors and the ethylene partial pressure.
The desired properties are obtained when the production split between the reactors is 1% in the prepolymerization, 45% in the loop reactor and 55% in the GPR.
An even softer PP copolymer can be produced by transferring the polymer obtained from GPR to another GPR to further increase the ethylene partial pressure in the second GPR.

A continuously operated pilot plant was used to produce PP copolymers with good creep properties. Equipment includes catalysts, alkyls, donors, propylene and ethylene feed systems, prepolymerization reactors, loop reactors and fluidized bed gas phase reactors (GPR).
Catalyst, alkyl, donor and propylene were fed to the prepolymerization reactor. The polymer slurry obtained from the prepolymerization reactor was fed to the loop reactor, which was also fed with hydrogen and further propylene.
The polymer slurry obtained from the loop reactor was fed to a flash tank where the monomer and polymer were separated. The polymer obtained from the flash tank was fed to the GPR. Propylene obtained from the flash tank was fed to GPR after removing hydrogen. Ethylene, additional hydrogen, and additional propylene were fed to the GPR.
The catalyst used was a highly active and highly stereospecific ZN catalyst made according to US Pat. No. 5,234,879. The catalyst was contacted with triethylaluminum (TEA) and dicyclopentyldimethoxysilane (DCPDMS) before feeding to the prepolymerization reactor (Al / Ti ratio was 150, Al / Do was 10 (mole)). there were).
The catalyst was fed according to US Pat. No. 5,385,992 and placed in a loop reactor with propylene. The prepolymerization reactor was operated at a pressure of 51 bar, a temperature of 20 ° C. and an average residence time of the catalyst of 7 minutes.
The loop reactor was operated at a pressure of 50 bar, a temperature of 75 ° C., and an average residence time of catalyst of 1 hour. The PP random polymer MFR (2.16 kg, 230 ° C.) produced in the loop reactor was set to 10 by adjusting the hydrogen feed rate.
The GPR reactor was operated at a total pressure of 29 bar and a propylene partial pressure of 16 bar. The operating temperature was 80 ° C. and the average residence time of the catalyst was 1.1 hours. The MFR (2.16 kg, 230 ° C.) of the PP copolymer removed from the GPR was adjusted to 5 by hydrogen partial pressure and production split between reactors. The ethylene content was adjusted to 3.5% w / w by controlling the production split between reactors and the ethylene partial pressure.
The desired properties are obtained when the production split between the reactors is 1% in the prepolymerization, 40% in the loop reactor and 59% in the GPR.
By transferring the polymer obtained from the GPR to another GPR and further increasing the ethylene partial pressure in the second GPR, a PP copolymer with even better impact resistance can be produced.

FIG. 1 is a schematic diagram of the process arrangement of the first preferred embodiment of the present invention. FIG. 2 is a schematic diagram of the process arrangement of the second preferred embodiment of the present invention.

Claims (3)

A method for producing a propylene copolymer comprising:
In at least one slurry reactor and at least two gas phase reactors selected from the group consisting of a loop reactor and a stirred tank reactor, a temperature of 40-110 ° C. and a pressure of 30-100 bar in the slurry reactor, In a gas phase reactor , propylene is polymerized with a comonomer in the presence of a catalyst at a temperature of 50 to 115 ° C. and a pressure of 10 to 40 bar , provided that 20 to 90% by weight of the polymer product is a slurry reactor as described above. Generated in;
Producing a first copolymerization product containing unreacted monomer in a slurry reactor; and without circulating the unreacted monomer to the slurry reactor in front of the gas phase reactor; The copolymerization product is fed directly to the first gas phase reactor to produce a second copolymerization product; and in the presence of further components for further copolymerization, the first gas phase reactor. Directing the second copolymerization product from to a second gas phase reactor to produce a first modified polymer having improved softness;
Wherein the second gas phase reactor is operated under a lower pressure than the first gas phase reactor. The process of claim 1, wherein the slurry reactor is operated at a temperature of 60C to 75C to produce random and ter-polymers. The process of claim 1, wherein the slurry reactor is operated at a temperature of 75C to 85C for improved activity and comonomer randomness.

Images