CN1140554C - Process for preparing propylene polymers - Google Patents

Abstract

Described herein is a process for preparing propylene copolymers. The process comprises the steps of polymerizing propylene with comonomers in at least one slurry reactor and at least one gas phase reactor, at least 10 % of the polymer product being produced in the gas phase reactor(s); recovering from the slurry reactor a copolymerization product containing unreacted monomers and conducting the copolymerization product to a first gas phase reactor essentially without recycling of the unreacted monomers to the slurry reactor before the gas phase reactor. The process will provide high randomness copolymers, which are very soft, and copolymers having improved impact resistance.

Description

Process for preparing propylene polymers

Background of the invention

field of invention

The present invention relates to the preparation of propylene based homopolymers and copolymers with high comonomer content and modified impact resistance propylene polymers. In particular, the present invention relates to a process for the preparation of propylene polymers in a combined reactor system comprising at least one slurry reactor and at least one gas phase reactor.

Description of related technologies

The softness, impact resistance and heat sealing properties of propylene-based polymers can be improved by copolymerizing propylene with other olefins such as ethylene, isobutylene, etc. Both bulk and gas phase methods have been used. However, the comonomers used in the polymerization cause the polymer to swell in the polymerization medium of the bulk process. As a result, when the swollen and soft polymer particles flash off after polymerization, the morphology of the particles is disrupted and the bulk density of the powdered polymer becomes very low. At the same time, amorphous substances accumulate on the powder surface. Sticky, low density materials tend to accumulate on the flash tank walls and create transfer problems. These problems increase with increasing proportion of comonomer.

For this reason, such polymerizations are mainly carried out in the prior art by using gas phase methods. Therefore, these methods have been proposed for the preparation of gummy, but flowable products (EP 0237003) and rubbery 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 behave like a fluid. However, the polymer in the fluidized bed reactor passes substantially in a plug flow manner.

Gas phase methods are also suitable for high comonomer content products, see EP 0674991, EP0584574, EP 0605001 and EP 0704464.

However, a problem associated with gas phase reactors arises from their long residence times, which imply long transition times and possible production losses. This is even more the case for multi-reactor processing. Catalyst productivity is low in gas phase processes, which means higher catalyst and production costs.

In order to extract different advantages from the slurry bulk reactor and the gas phase reactor, some people in the art have proposed some combinations of the bulk reactor and the gas phase reactor for the preparation of propylene copolymers. However, to date, none of the prior art processes has met the requirements of flexibility and low production costs for the production of various polyolefin grades using the same process flow. In particular, the recycle of relatively large amounts of unreacted monomer to the slurry reactor, which is a typical feature of known processes, impairs the dynamics of the loop reactor and slows down the transition to new product grades.

US Patent 4740550 discloses an improved two-stage process for the polymerization of propylene in a combination of loop reactor and gas phase reactor. The main objective of US4740550 is to provide a process for preparing high quality block copolymers by sending a homopolymer with a narrow residence time distribution to the block copolymerization stage. The disclosed process comprises the following stages: a first stage homopolymerization in a bulk loop reactor, a second stage homopolymerization in a gas phase reactor, a cyclone separator between the first stage and the second stage Fines removal and final impact copolymerization in another gas phase reactor.

Before the polymerization product of the loop reactor is fed to the gas phase, its fines fraction is removed and recycled back to the loop reactor. A portion of the monomer from the gas phase reactor is recycled directly to the first stage loop reactor together with the fines.

This prior art suffers from some serious problems. Therefore, if all the fines are removed from the reactor outlet of the loop reactor and recycled back to the loop reactor, there is a great risk that the loop reactor will eventually be filled with deactivated catalyst or slightly polymerized waste fines. On the other hand, if a portion of the fines stream is mixed with the product from the latter reactor, this can lead to problems with inhomogeneity of the final product. Also, if a portion of the fines stream is collected separately and blended with a separate homopolymer, as also proposed in US4740550, operational complexity and economic burdens result.

SUMMARY OF THE INVENTION

It is an object of the present invention to eliminate the problems of the prior art of single and multi-reactor processes and to provide a new process for the preparation of propylene homopolymers and copolymers.

Another object of the present invention is to provide a highly versatile process that can be used to prepare a wide range of different propylene (co)polymers.

These and other objects, together with their advantages over known methods, are achieved by the invention hereinafter described and claimed, as will be apparent from the following description of the patent specification.

The process according to the invention is based on the combination of at least one slurry reactor and at least one gas phase reactor connected in series (thus forming a cascade). Propylene (co)polymers are prepared at elevated temperature and pressure in the presence of a catalyst. According to the present invention, the unreacted monomer-containing polymerization product of at least one slurry reactor is sent to a first gas phase reactor with little or no monomer recycling back to the slurry reactor. In the context of the present invention, it has been found that high quality impact copolymers can be prepared by two stage homopolymerization followed by an impact copolymerization step without any fines removal and recycling steps after either the first or second stage copolymerization. In the present invention, the circulation volume is minimized according to the object of the invention by using a specific reactor sequence and by selecting the corresponding throughput of each reactor.

According to another aspect of the present invention, at least one slurry reactor and at least one gas phase reactor connected in series are used as a reactor system, said at least one slurry reactor being a bulk operated at high or supercritical temperature a loop reactor, and the contents of said slurry reactor, including copolymerization product and reaction medium containing unreacted monomer, are introduced directly into the gas phase reactor fluidized bed using piping interconnecting the slurry reactor and the gas phase reactor .

More precisely, the method according to the invention is essentially characterized by what is stated in the characterizing part of claim 1 .

The present invention has numerous important advantages. As far as its arrangement is concerned, it has been found that the monomer fed to the first reactor may be largely or completely consumed in the gas phase reactor following the slurry reactor. This may be due to the gas phase operation which vents a small amount of gas with the polymerization product. The dynamics of the loop reactors in the cascade provide fast transitions and high production rates. Fast start-up is also possible, as the gas phase bed material can come directly from the loop reactor. A variety of broad molecular weight distribution or bimodal products can be produced using loop reactors and gas phase reactor cascades. Because of adjustable bed levels and reaction rates, its at least one gas phase reactor offers a high degree of flexibility in the reaction rate ratio between the first and second fraction of products. Furthermore, the absence of solubility limitations in gas phase reactors makes it possible to produce high and very high comonomer content polymers.

The loop-gas-phase reactor combination greatly reduces residence time and production losses compared to gas-gas-phase multi-reactor processes.

Brief description of the drawings

Fig. 1 is the schematic process flow diagram of the first preferred embodiment of the present invention; With

Figure 2 is a schematic process flow diagram of the second preferred embodiment of the present invention.

Detailed Description of the Invention

definition

For the purposes of the present invention, "slurry reactor" means any reactor, such as a continuous or simple stirred tank reactor or a loop reactor, operating in bulk or slurry and forming polymer in granular form. "Bulk" means polymerization in a reaction medium comprising at least 60% by weight of monomer. According to a preferred embodiment, the slurry reactor comprises a bulk loop reactor.

"Gas phase reactor" means any mechanically mixed reactor or fluidized bed reactor. Preferably said gas phase reactor comprises a mechanically agitated fluidized bed reactor having a gas velocity of at least 0.2 m/s.

By "high temperature polymerisation" is meant a polymerisation temperature above the defined temperature of 80°C which is known to be detrimental to the corresponding prior art high yield catalysts. At high temperatures, the stereoregularity of the catalyst and the morphology of the polymer powder can be compromised. This is not the case with the particularly preferred types of catalysts used in the present invention described below. The high temperature polymerization is carried out at a temperature above a defined temperature and below the corresponding critical temperature of the reaction medium.

"Supercritical polymerisation" means a polymerisation carried out above the corresponding critical temperature and pressure of the reaction medium.

"Direct feed" refers to a process in which the contents of a slurry reactor including polymerization product and reaction medium are introduced directly into the fluidized bed of a gas phase reactor.

"Reaction zone" means one reactor or several reactors of similar type connected in series for the production of polymers of the same type or of the same characteristics.

The phrase "essentially no monomer recycle" has the same meaning as "little or no monomer recycle" and is used to mean less than about 30% by weight, preferably less than 20% by weight, especially no monomer recycle Back to slurry handling. In contrast, in normal slurry processes, 50% by weight or more of the monomers are recycled.

overall approach

The present invention relates to a multi-stage process comprising a bulk reaction zone comprising at least one slurry reactor and a gas phase reaction zone comprising at least one gas phase reactor cascaded after the at least one slurry reactor, with little or no single The bulk is recycled back to the first reactor and fed directly or indirectly into the gas phase for homopolymerization or copolymerization of propylene.

In the direct feed process, the contents of the slurry reactor, ie the polymer product and the reaction medium, are introduced directly into the fluidized bed reactor. The product outflow from the slurry reactor may be intermittent, but is preferably continuous. The slurry is withdrawn as such without any separation of gases or separation of particle streams based on different particle sizes. No particulate matter is sent back to the forereactor. The line between the slurry reactor and the gas phase reactor may optionally be heated in order to evaporate only part or all of the reaction medium before it enters the gas phase reactor polymer bed.

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

In dual reactor operation, the polymer leaving the gas phase reactor with discharge system enters a solids/gas separation unit. The polymer is sent from the bottom to the next processing step and the gas is compressed and recycled back to the gas phase reactor after the purification step. Generally light inerts such as methane and ethane, as well as heavier inerts such as propane and oligomers are removed in these purification steps. The purification can be performed using distillation or membrane separation. In the case of distillation, the monomers are recycled back to the gas phase reactor mainly in liquid form.

In a three-reactor operation, polymer leaving the first gas phase reactor with a discharge system enters a solids/gas separation unit. The polymer is sent further from the bottom to the second gas phase reactor and the gas is compressed and recycled back to the first gas phase reactor after the purification step. Generally light inerts such as methane and ethane, as well as heavier inerts such as propane and oligomers are removed in these purification steps. The purification can be performed using distillation or membrane separation. In the case of distillation, the monomers are recycled back to the gas phase reactor mainly in liquid form.

Optionally in a three-reactor operation, the polymer leaving the first gas phase reactor with a discharge system enters the second gas phase reactor directly, together with accompanying gas.

In a three-reactor operation, polymer leaving a second gas phase reactor with a discharge system enters a solids/gas separation unit. The polymer is sent from the bottom to further processing steps and the gas is compressed and recycled partly directly back to the second gas phase reactor and partly after the purification step. Generally light inerts such as methane and ethane, as well as heavier inerts such as propane and oligomers are removed in these purification steps. The purification can be performed using distillation or membrane separation. In the case of distillation, an ethylene-rich stream is recycled back to the second gas phase reactor, while a propylene-propane stream is sent to the propane and oligomer removal steps.

Catalysts are used to obtain polymeric products. The catalyst can be any catalyst having sufficient activity at elevated temperatures. Preferred catalyst systems for use include high yield Ziegler-Natta catalysts having a catalyst component, a cocatalyst component, an external donor and optionally an internal donor. Another preferred catalyst system is a metallocene-based catalyst, for example with a bridged ligand structure providing high stereoselectivity and impregnated on the support in the form of an active complex.

The polymerization temperature is at least 60°C and preferably at least 65°C. Slurry reactors are operated at elevated pressures of at least 35 bar up to 100 bar and gas phase reactors are operated at pressures of at least 10 bar up to dew point pressure. Or any reactor in the series of reactors can be operated above the critical temperature and pressure.

Propylene and optionally one or more other _C2 to _C16 olefins such as ethylene, 1-butene, 4-methyl-1-pentene, 3-methyl-1-butene, 1-hexene , 1-octene, 1-decene, diolefins or cycloolefins such as vinylcyclohexane or cyclopentene are polymerized and copolymerized respectively in multiple polymerization reactors connected in series. Olefins as comonomers can be used in any reactor. Different amounts of hydrogen can be used as molar mass regulator in any one of the reactors.

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

catalyst

Polymerization products are obtained by using catalysts. As catalysts, any stereospecific catalyst of propylene with high yields and useful polymer properties such as isotacticity and morphology at high temperature and possibly supercritical polymerization can be used. The preferred catalyst system used comprises a high yield Ziegler-Natta catalyst having a catalyst component, a cocatalyst component, optionally an external donor and an internal donor. Another preferred catalyst system is a metallocene catalyst having a bridged ligand structure providing high stereoselectivity and having the active complex impregnated on the support. Ultimately, the catalyst is preferably any other catalyst that provides sufficient activity at elevated temperatures.

Examples of suitable catalyst systems are described, for example, in Finnish patents 86866, 96615 and 88047, 88048 and 88049.

Finnish patent 88047 discloses a particularly preferred catalyst for use in the present invention. Another preferred catalyst is disclosed in Finnish patent application 963707.

Further preferred catalysts are disclosed in PCT/FI 97/00191 and PCT/FI 97/00192.

pre-polymerization

The catalyst may be prepolymerized before being fed to the first polymerization reactor in the series. In prepolymerization, the catalyst components are contacted with monomers, such as olefin monomers, prior to feeding into the reactor. Examples of suitable systems are described, for example, in Finnish patent application FI 961152.

The prepolymerization may also be carried out in the presence of a viscous substance such as olefinic wax to provide a prepolymerized catalyst which is stable in storage and handling. Catalyst prepolymerized in wax allows easy metering of the amount of catalyst into the polymerization reactor. Examples of suitable systems are described eg in Finnish patent 95387. Typically about 1 part catalyst is used for up to 4 parts polymer.

The monomers used for prepolymerization can be selected from propylene, 1-butene, 4-methyl-1-pentene, 3-methyl-1-butene, vinylcyclohexane, cyclopentene, 1-hexene ene, 1-octene and 1-decene.

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

polymerization

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

Equipment for the polymerization step may include any suitable type of polymerization reactor. The slurry reactor can be any continuous or simple stirred tank reactor or loop reactor operating in bulk or slurry and in which the polymer is formed in particulate form. Bulk means polymerization in a reaction medium comprising at least 60% by weight of monomer. The gas phase reactor can be any mechanically mixed reactor or fluidized bed reactor. According to the present invention, the slurry reactor is preferably a bulk loop reactor, and the gas phase reactor is preferably a fluidized bed reactor with a mechanical stirrer.

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

The production split between the slurry reactor and the first gas phase reactor is typically 67:33 to 50:50 when monomer recycling is allowed back to the slurry reactor. Conversely, when recycling back to the slurry reactor is not required, the throughput in the slurry reactor is less than or equal to the throughput in the first gas phase reactor. In all cases the throughput in the slurry reactor was greater than 10%. Thus, according to a preferred embodiment, 10-70% by weight, preferably 20-65% by weight, especially 40-60% by weight, of the polymer is produced in the slurry reaction zone and no monomer is Recycle back to the slurry reactor zone. When 50-67% of the polymer is produced in the slurry reaction zone, a small amount of monomer can be recycled from the gas phase reaction zone to the slurry reactor.

According to the present invention, said polymerization process comprises at least the following steps:

- polymerizing or copolymerizing propylene and optionally other olefins in a first slurry polymerization zone or reactor,

- recovering the first polymerization product and the reaction medium from the first reaction zone,

- direct or indirect delivery of the first polymerisation product to a gas phase polymerisation zone or reactor,

- optionally feeding further propylene and/or comonomer to the second reaction zone,

- subjecting excess propylene and/or comonomer from the first zone and additional propylene and/or comonomer to a second polymerization reaction in the presence of the first polymerization product to produce a second polymerization product ,

- recovering the polymer product from the second reaction zone, and

- Separation and recovery of polypropylene from the second reaction product.

In addition, the method may also include one or more of the following additional steps

- prepolymerization of the catalyst with one or more monomers,

- separation of gases from the second reaction zone product,

- sending recovered polymerized product from the preceding zone to a third or fourth reaction zone or reactor,

- optionally feeding additional propylene and/or comonomer to the third and fourth reaction zone,

- third and fourth polymerization reactions of excess propylene and/or comonomer and additional propylene and/or comonomer in the presence of the polymerization product of the previous zone to produce third or fourth polymerization products, and

- recovering the polymer product from the third or fourth reaction zone, and

- separation and recovery of polypropylene from the third or fourth reaction product.

In the first step of the process, propylene and optional comonomers are fed into a first polymerization reactor together with the activated catalyst complex and optional cocatalyst and other auxiliary components. The catalyst can be pre-polymerized or pre-polymerized before being sent to processing. As molar mass regulator, hydrogen can also be fed into the reactor together with the aforementioned components in the amount required to obtain the desired molar mass of the polymer. In embodiments where no recycle is returned to the slurry reactor, only fresh monomer is sent to the first reactor.

Alternatively, in embodiments where very little monomer is recycled back to the slurry reactor, the feed to the reactor may consist of monomer, if any, recycled from the reactor below via a recovery system, as well as fresh monomer added. body, hydrogen, optional comonomers and catalyst components.

In all embodiments, in the presence of propylene, optional comonomers, co-catalyst and other co-agent components, the activated catalyst complex will polymerize in a slurry reactor and form to be suspended in the reactor The product in the form of particles in the circulating fluid is the polymer particles.

The polymerization medium generally comprises monomers and optionally a hydrocarbon and the fluid is either liquid or gaseous. In the case of slurry reactors, especially in the case of loop reactors, the fluid is liquid and the suspension of the polymer is continuously circulated in the slurry reactor whereby more polymeric Suspensions of substances in hydrocarbon media or monomers. According to a preferred embodiment, the first polymerization or copolymerization is carried out in a reaction medium consisting essentially of propylene. Preferably at least 60% by weight of the medium is propylene.

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

In the slurry polymerization zone, more than one reactor may be used in series. In this case, the polymer suspension in inert hydrocarbons or monomers produced in the first slurry reactor is sent to the next without periodic or continuous separation of inert components and monomers In a slurry reactor operating at a lower pressure than the previous slurry reactor.

The heat of polymerization was removed by cooling the reactor with a cooling jacket. The residence time in the slurry reactor must be at least 10 minutes and preferably 20 to 100 minutes to obtain a sufficient degree of polymerization. This is necessary to obtain a polymer yield of at least 40 kg polypropylene per gram catalyst. It is advantageous to operate slurry reactors at high solids concentrations such as 50% for homopolymers and 35 or 40% for some copolymers when the particles swell. If the solids concentration in the loop reactor is too low, the amount of reaction medium sent to the second reaction zone or gas phase reactor is increased.

The contents of the slurry reactor, ie the polymer product and reaction medium, are directed into the fluidized bed of the gas phase reactor.

Said second reactor is preferably a gas phase reactor in which propylene and optional comonomers are polymerized in a reaction medium consisting of gas or vapour.

The gas phase reactor may be a conventional fluidized bed reactor, although other types of gas phase reactors may be used. In a fluidized bed reactor, the bed consists of formed and increasing polymer particles together with the polymer fraction still active catalyst. The bed is maintained in a fluidized state by introducing a gas component such as a monomer at a flow rate (at least 0.2 m/s) such that the particles are fluidized. The fluidization gas may also comprise inert gases such as nitrogen and hydrogen also as regulator. In the present invention, it is not recommended to use unnecessary inert gas which may cause problems in the recovery section.

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

According to a preferred embodiment, no fresh propylene is sent to the first gas phase reactor other than that required for various flushes.

In order to separate part of the gaseous and possibly volatile components (e.g. re-comonomers and compounds used as catalyst feedstock), for example in a flash tank, the pressure of the second polymerization product, including the gaseous reaction medium, is at the first After the gas phase reactor is lowered. The overhead gas stream is recycled back to the first gas phase reactor or partly back to the first gas phase reactor and partly back to the slurry reactor through the recovery system. Some monomer, generally the heavier comonomers, may be recycled to the bulk reaction zone.

If desired, the polymer product can be fed to a second gas phase reactor and subjected to a third polymerisation reaction to produce a modified polymer product from which polypropylene can be separated and recovered. The third polymerization reaction is carried out in a gas phase reactor in the presence of comonomers which impart properties to the third polymerization product such as improved impact strength, ductility or softness. Typically, part of the gas from the first gas phase reactor is removed in a depressurization step before the second gas phase reactor. The removed gas is compressed to the recovery section and treated as described for the dual reactor case. Alternatively, the second product can be transferred directly to the third reactor.

Generally, if the copolymers are prepared according to the process of the present invention, they contain at least 0.5% by weight of comonomer, especially at least about 2% by weight of comonomer and preferably up to 20% by weight. % by weight of at least one comonomer. The typical comonomer content of the copolymer sent to the first gas phase reactor is about 2-16% by weight. The resulting copolymers can exhibit highly random properties (very soft copolymers).

In particular, the polymerization product may be sent to a second gas phase reactor to provide a rubbery copolymer by a third (co)polymerization reaction forming a modified polymerization product. This third polymerization reaction will provide polymer product properties such as improved impact strength. This step of providing an elastomer can be performed in various ways. Accordingly, it is preferred that the elastomer is prepared by copolymerizing at least propylene and ethylene to form an elastomer. The conditions for the copolymerization are within the limits of the conventional EPM production conditions disclosed, for example, on pages 545 to 558 of the Encyclopedia of Polymer Science and Engineering, Volume 6, Second Edition. If the content of ethylene repeating units in the polymer is within a certain range, a rubbery product is formed. Therefore, it is preferable to copolymerize ethylene and propylene into an elastomer in such a ratio that the copolymer finally contains 10 to 70% by weight of ethylene units. Specifically, the ethylene unit content is 30 to 50% by weight of the copolymer propylene/ethylene elastomer. In other words, ethylene and propylene are copolymerized at an ethylene/propylene molar ratio of 30/70 to 50/50 to form an elastomer.

The elastomers may also be provided by adding prefabricated or natural elastomers to the polymerization product of the first gas phase reactor.

The impact-modified polypropylene generally comprises about 5-50% by weight, especially about 10-45% by weight and preferably about 15-40% by weight of the abovementioned elastomers.

Generally, in order to achieve higher molar mass for improved impact properties, the hydrogen concentration of the second reaction product is reduced before the product is sent to the second gas phase.

It is also possible to transfer the product of the second gas phase reaction to a third (fourth etc.) polymerization reaction zone, wherein the copolymerization is carried out in the presence of a comonomer which imparts improved properties to the third polymerization product.

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

Summarizing the above, a particularly preferred embodiment of the present invention comprises (Fig. 1)

- polymerisation of propylene in a loop reactor at a pressure of 40 to 80 bar and a temperature of 80 to 100 °C, hydrogen is used to control the molar mass of the polymerisation product,

- recovery of polymerized product from the loop reactor and transfer to the fluidized bed of the gas phase reactor,

- optionally feeding additional propylene and optionally comonomers to the gas phase reactor,

- optionally feeding additional hydrogen to the gas phase reactor to control the hydrogen/propylene ratio to provide the desired molecular mass of the polymerized product,

- recovering the polymerized product from the gas phase reactor and passing it to a flash tank, where the pressure of the product is reduced to produce a tank top product comprising essentially unreacted propylene and hydrogen and a bottoms product comprising mainly polymerized solids,

- recycling the tank top product, or at least a major part thereof, to the gas phase reactor by means of the recovery part, and

- Recovery of polypropylene polymer as flash tank bottom product.

According to a second particularly preferred embodiment (FIG. 1):

- In a loop reactor at a pressure of 40 to 80 bar and a temperature of 60 to 80°C, propylene and comonomers such as ethylene or 1-butene or both are polymerized and hydrogen is used to provide polymerized product,

- transfer of the polymerization product from the loop reactor directly to the fluidized bed of the gas phase reactor,

- feeding optionally additional propylene and comonomers to the gas phase reactor,

- feeding optionally additional hydrogen to the gas phase reactor to control the hydrogen/propylene ratio to provide the desired molar mass of polymerized product,

- recovering the polymerized product from the gas phase reactor and passing it to a flash tank, where the pressure is reduced to produce a tank top product substantially comprising unreacted monomer and hydrogen and a bottom product comprising mainly polymerized solids,

- recycling the tank top product, or at least a major part thereof, to the gas phase reactor by means of the recovery part, and

- Recovery of polypropylene polymer as flash tank bottom product.

According to a third particularly preferred embodiment (FIG. 2):

- polymerization of propylene and optional comonomers in a loop reactor at a pressure of 40 to 80 bar and a temperature of 60 to 100° C. and hydrogen is used to control the molar mass of the polymerization product,

- recovery of polymerized product from the loop reactor and transfer to the fluidized bed of the gas phase reactor,

- feeding optionally additional propylene and optional comonomers to the gas phase reactor,

- optionally feeding additional hydrogen to the gas phase reactor to control the hydrogen/propylene ratio to provide a polymer product of desired molecular mass,

- Polymerization product is recovered from the first gas phase reactor and sent to an intermediate flash tank where the pressure of the product is reduced to produce a tank top product consisting essentially of unreacted monomer and hydrogen and comprising mainly polymerized solids bottom product of

- recycle the tank top product, or at least the main part thereof, to the first gas phase reactor by means of the recovery part,

- the polypropylene polymer from the bottom of the intermediate flash tank is sent to the third polymerization via the polymer feed system,

- a third polymerization in a gas phase reactor in the presence of comonomers,

- Polymerization product from the second gas phase reactor is recovered and sent to a flash tank where the pressure of the product is reduced to produce a top product containing essentially unreacted monomer and hydrogen and a bottom containing mostly polymerized solids product,

- Optionally the polymer product from the third polymerisation can be transferred directly or via a flash tank to a third (fourth etc.) gas phase polymerisation reactor where the polymerisation is carried out in the presence of comonomers.

The two preferred embodiments described above are also shown in the accompanying drawings illustrating the specific arrangement of the processing equipment used. Its number corresponds to the following specific equipment:

1;101 Pre-polymerization reactor

30; 130 Catalyst memory

31;131 Feeding device

32; 132 diluent (optional)

33;133 Catalyst/diluent mixtures

34; 134 Monomer

35;135 Cocatalysts and possible donors

40; 140 Loop Reactor

42;142 Diluent feed (optional)

43;143 monomer feed

44;144 Hydrogen feed

45;145 comonomer feed (optional)

46; 146 Return loop reactor 40; 140 via line 46; 146

47;147 One or more exhaust valves

150b flash separator

152b Remove pipeline

60; 160; 160b Gas phase reactor

61;161;161b Gas delivery lines

62; 162; 162b Compressor

63;163;163b monomer feed

64;164;164b comonomer feed

65; 165; 165b Hydrogen feed

66; 166; 166b Transmission pipeline

67; 167 Product delivery pipeline

68;168 Polymerization product recovery systems such as flash tanks

69; 169 Take-up line

70; 170 Monomer recovery system

Referring to FIG. 1 , it may be noted that catalyst from storage 30 is sent to feed 31 with optional diluent from line 32 . Feed device 31 feeds the catalyst/diluent mixture into prepolymerization chamber 1 via line 33 . The monomer is fed through 34, the cocatalyst and possibly the donor can be fed into the reactor 1 through line 35 or preferably the cocatalyst and the donor body are intermixed and fed in line 35.

The prepolymerized catalyst is preferably removed from the prepolymerization chamber 1 directly via line 36 and conveyed to the loop reactor 40 . In loop reactor 40, the polymerization is carried out by adding optional diluent from line 42, monomer from line 43, hydrogen from line 44 and optional comonomer from line 45 via line 46. Ongoing. Optional co-catalysts may also be introduced into loop reactor 40 .

From the loop reactor 40 the polymer-hydrocarbon mixture is sent through one or several vent valves 47 such as those described in Finnish patent application 971368 or 971367. Between loop reactor 40 and gas phase reactor 60 there is a direct product transfer line 67 .

In the lower part of the gas phase reactor 60, there is a fluidized bed of polymer particles which are passed through the gas removed from the top of the reactor 60 in the normal way through the line 61, the compressor 62 and the heat exchanger ( not shown) to the lower part of the reactor 60 for circulation to maintain a fluidized state. Reactor 60 may advantageously be equipped with a mixer (described in Finnish patent application 933073, not shown in the figures), but this is not required. Into the lower part of reactor 60 are introduced monomer from line 63, optional comonomer from line 64 and hydrogen from line 65 in a known manner. Product may be continuously or periodically discharged from reactor 60 to flash tank 68 via transfer line 66 . The top product of the recovery system is recycled to the gas phase reactor through the monomer recovery system.

The embodiment shown in Figure 2 differs from Figure 1 only in the sense that the product from gas phase reactor 160 is sent to a further gas phase reactor 160b. Polymer particles are removed from flash tank 168 and polymer feed tank 150b via removal line 152b and sent to gas phase reactor 160b. The gas phase reactor is advantageously equipped with a mixer (not shown).

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

In above-mentioned two kinds of embodiments, numbering 70 and 170 represent that the circulating monomer of gas phase reactor (60, 160 and 160b) or separator (68, 168 and 168b) can be combined with hydrogen and/or general boiling point lower than single Separation devices such as membrane units or strippers for the separation of light inert hydrocarbons from solids.

polymer

Products produced according to the present invention include polypropylene copolymers including polypropylene terpolymers. In particular, very soft highly random copolymers can be produced by means of the present invention. The copolymer comprises at least 0.5% by weight of comonomer, especially at least about 2% by weight and preferably up to 20% by weight of comonomer. Typical comonomer levels are from about 2 to 12% by weight. An essential feature of the present invention is the high polymerization temperature used, preferably above 75°C, which provides a more uniform comonomer distribution in the copolymerization. The degree of randomness, as measured by Fourier transform infrared spectroscopy, was 69% at a polymerization temperature of 60°C and 71% at 65°C, at a polymerization temperature of 75°C in the first reactor and at a polymerization temperature of 80°C in the second reactor. 74% at the polymerization temperature of °C.

Additional products produced by the present invention include impact modified propylene polymers, preferably comprising rubbery copolymers, especially ethylene-propylene copolymers, which improve the impact resistance of the product. The elastomer has a proportion of polypropylene of about 5-40% by weight.

The following non-limiting examples serve to illustrate the mechanism of the invention.

Example 1

Simulation of a production-scale continuous production facility for polypropylene homopolymer. The production facility includes catalyst, alkanes, donors, propylene feed system, prepolymerization reactor, loop reactor and fluidized bed gas phase reactor (GPR).

Catalyst, alkanes, donors and propylene are fed into the prepolymerization reactor. The polymer slurry from the prepolymerization reactor is fed to the loop reactor, and hydrogen and more propylene are simultaneously fed to the loop reactor. The polymer slurry from the loop reactor was fed to the GPR with additional hydrogen and propylene. The production rates in the individual reactors were: 300 kg/h in the prepolymerization reactor, 15 t/h in the loop reactor and 10 t/h 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 MFR (2.16 kg, 230° C.) of the polypropylene homopolymer produced in the loop reactor was adjusted to 1 by controlling the hydrogen feed.

The GPR operates at a pressure of 35 bar and a temperature of 85°C. The MFR (2.16 kg, 230° C.) of the polypropylene homopolymer produced by GPR was adjusted to 13 by controlling the partial pressure of hydrogen. 5 t/h of propylene was recycled from the GPR outlet to the loop reactor. The per pass conversion of propylene was 83%.

Example 2

Simulation of a production-scale facility for the continuous production of polypropylene homopolymer with good impact properties. The facility included catalyst, alkanes, donors, propylene feed system, prepolymerization reactor, loop reactor and two fluidized bed gas phase reactors (see Figure 2).

Catalyst, alkanes, donors and propylene are fed into the prepolymerization reactor. The polymer slurry from the prepolymerization reactor is fed to the loop reactor, and hydrogen and more propylene are simultaneously fed to the loop reactor. The polymer slurry from the loop reactor was fed to the first GPR with additional hydrogen and propylene.

The polymer from the first GPR is depressurized before entering the second GPR. Ethylene and additional propylene are fed to the second GPR.

The production rates in each reactor are: 300 kg/h in the prepolymerization reactor, 15 t/h in the loop reactor, 10 t/h in the first GPR and 6 t/h 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 MFR (2.16 kg, 230° C.) of the polypropylene homopolymer produced in the loop reactor was adjusted to 20 by controlling the hydrogen feed.

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 polypropylene homopolymer produced by the first GPR was adjusted to 20 by controlling the partial pressure of hydrogen. 4.3 t/h of propylene was recycled from the GPR outlet to the loop reactor.

The second GPR operates at a pressure of 20 bar and a temperature of 70°C. The MFR (2.16 kg, 230° C.) of the polypropylene homopolymer produced by the second GPR was adjusted to 13 by controlling the partial pressure of hydrogen. 2.7 t/h of propylene was recycled from the second GPR outlet to the loop reactor and 1.6 t/h of ethylene was recycled to the second GPR.

Example 3

Simulation of a production-scale facility for the continuous production of random polypropylene polymers. The facility includes catalyst, alkanes, donors, propylene and ethylene feed systems, prepolymerization reactors, loop reactors and a fluidized bed gas phase reactor (GPR).

Catalyst, alkanes, donors and propylene are fed into the prepolymerization reactor. The polymer slurry from the prepolymerization reactor is fed to the loop reactor. Ethylene, hydrogen and more propylene are fed to the loop reactor. The polymer slurry from the loop reactor is fed to the GPR with additional hydrogen, ethylene and propylene. The production rates in the individual reactors were: 300 kg/h in the prepolymerization reactor, 15 t/h in the loop reactor and 10 t/h in the 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 atactic polypropylene produced in the loop reactor was adjusted to 7 by controlling the hydrogen feed and the ethylene content to 3.5% by weight by ethylene feed.

The GPR operates at a pressure of 35 bar and a temperature of 80°C. The MFR (2.16 kg, 230° C.) of the GPR-produced atactic polypropylene was adjusted to 7 by controlling the partial pressure of hydrogen, and the ethylene content was set at 3.5% by weight by adjusting the partial pressure of ethylene. 5 t/h of propylene and 33 kg/h of ethylene are recycled from the GPR outlet to the loop reactor. The per pass conversions of propylene and ethylene were 83% and 96%, respectively.

Example 4

Simulation of a production-scale facility for the continuous production of polypropylene copolymers with good impact and creep properties. The facility includes catalyst, alkanes, donors, ethylene and propylene feed systems, prepolymerization reactor, loop reactor, flash tank and two fluidized bed gas phase reactors.

Catalyst, alkanes, donors and propylene are fed into the prepolymerization reactor. The polymer slurry from the prepolymerization reactor is fed to the loop reactor, and hydrogen and more propylene are simultaneously fed to the loop reactor. The polymer slurry from the loop reactor is sent to a flash tank where propylene and polymer are separated.

The polymer from the flash tank is fed to the first GPR. Propylene from the flash tank is fed to the first GPR after hydrogen removal. Ethylene and additional propylene are fed to the first GPR. The polymer from the first GPR is fed to the second GPR. Ethylene, some hydrogen and additional propylene are fed to the second GPR.

The production rates in each reactor are: 300 kg/h in the pre-polymerization reactor, 10 t/h in the loop reactor, 10 t/h in the first GPR and 6 t/h in the second 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 85°C. The MFR (2.16 kg, 230° C.) of the polypropylene homopolymer produced in the loop reactor was adjusted to 100 by controlling the hydrogen feed.

The GPR operates at a pressure of 35 bar and a temperature of 80°C. The MFR (2.16 kg, 230° C.) of the polypropylene of the GPR was adjusted to 0.4 by controlling the throughput ratio between the reactors and the hydrogen removal efficiency of the flashed propylene. The ethylene content was set at 2% by weight by adjusting the ethylene partial pressure and controlling the throughput ratio between the reactors.

The second GPR operates at a pressure of 20 bar and a temperature of 70°C. The MFR (2.16 kg, 230°C) of the polypropylene copolymer produced by the second GPR was adjusted to 0.3 by controlling the partial pressure of hydrogen and by controlling the throughput ratio between the reactors. A small amount of propylene is recycled from the second GPR back to the loop reactor.

Example 5

Simulation of a production-scale facility for the continuous production of polypropylene copolymers with good creep properties. The facility includes catalyst, alkanes, donors, ethylene and propylene feed systems, prepolymerization reactors, loop reactors, flash tanks and fluidized bed gas phase reactors.

Catalyst, alkanes, donors and propylene are fed into the prepolymerization reactor. The polymer slurry from the prepolymerization reactor is fed to the loop reactor, and ethylene and more propylene are simultaneously fed to the loop reactor. The polymer slurry from the loop reactor is sent to a flash tank where monomer and polymer are separated.

The polymer from the flash tank is fed to the GPR. Propylene from the flash tank is fed to the GPR after removal of ethylene. Hydrogen and additional propylene are fed to the GPR.

The production rates in each reactor are: 300 kg/h in the prepolymerization reactor, 10 t/h in the loop reactor and 10 t/h in the first GPR.

Example 6

Polypropylene homopolymer is produced using a continuously operating pilot plant. The equipment includes catalyst, alkanes, donors, propylene feed system, prepolymerization reactor, loop reactor and fluidized bed gas phase reactor (GPR).

Catalyst, alkanes, donors and propylene are fed into the prepolymerization reactor. The polymer slurry from the prepolymerization reactor is fed to the loop reactor, and hydrogen and more propylene are simultaneously fed to the loop reactor. The polymer slurry from the loop reactor was fed to the GPR with additional hydrogen and propylene.

After removal of the polymer product from the GPR, the formed polymer and unreacted propylene were separated.

The catalyst used is a highly active and stereospecific ZN-catalyst prepared according to US Patent 5234879. The catalyst was mixed with triethylaluminum (TEA) and dicyclopentyldimethoxysilane (DCPDMS) (Al/Ti ratio 250, Al/Do 40 (mol) before being sent to the prepolymerization reactor. ))touch.

The catalyst was fed and flushed with propylene into the prepolymerization reactor according to US Pat. No. 5,385,992. The prepolymerization reactor was operated at a pressure of 51 bar and a temperature of 20°C with an average catalyst residence time of 7 minutes.

The prepolymerized catalyst, propylene and other components are transferred to the loop reactor. The loop reactor was operated at a pressure of 50 bar and a temperature of 80°C with an average catalyst residence time of 1 hour. The MFR (2.16 kg, 230° C.) of the polypropylene homopolymer produced in the loop reactor was adjusted to 7 by controlling the hydrogen feed.

The polymer slurry from the loop reactor was transferred to the 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 polypropylene homopolymer produced by the GPR was 7 and was controlled by adjusting the partial pressure of hydrogen. The throughput ratio between the reactors was: 1% for the prepolymerization reactor, 49% for the loop reactor and 50% for the GPR. The catalyst productivity was 32 kg of polypropylene per gram of catalyst.

Example 7

Polypropylene homopolymer is produced using a continuously operating pilot plant. The equipment includes catalyst, alkanes, donors, propylene feed system, prepolymerization reactor, loop reactor and fluidized bed gas phase reactor (GPR).

Catalyst, alkanes, donors and propylene are fed into the prepolymerization reactor. The polymer slurry from the prepolymerization reactor is fed to the loop reactor, and hydrogen and more propylene are simultaneously fed to the loop reactor. The polymer slurry from the loop reactor was fed to the GPR with additional hydrogen and propylene.

The formed polymer and unreacted propylene were separated after removal from the GPR.

The catalyst used was a highly active and stereospecific ZN-catalyst prepared according to Finnish patent application 963707. The catalyst was mixed with triethylaluminum (TEA) and dicyclopentyldimethoxysilane (DCPDMS) (Al/Ti ratio 250, Al/Do 40 (mol) before being sent to the prepolymerization reactor. ))touch.

The catalyst was fed according to US Patent 5385992 and flushed with propylene into the prepolymerization reactor. The prepolymerization reactor was operated at a pressure of 53 bar and a temperature of 20°C with an average catalyst residence time of 7 minutes.

The prepolymerized catalyst, propylene and other components are transferred to the loop reactor. The loop reactor was operated at a pressure of 52 bar and a temperature of 85°C with an average catalyst residence time of 1 hour. The MFR (2.16 kg, 230° C.) of the polypropylene homopolymer produced in the loop reactor was adjusted to 7 by controlling the hydrogen feed.

The polymer slurry from the loop reactor was transferred to the GPR. The GPR was operated at a total pressure of 29 bar and a propylene partial pressure of 21 bar. The temperature of the GPR was 80°C and the average residence time of the catalyst was 1 hour. The MFR (2.16 kg, 230°C) of the GPR-produced polypropylene homopolymer was 7 and was adjusted by controlling the partial pressure of hydrogen. The throughput ratio between the reactors was: 1% for the prepolymerization reactor, 53% for the loop reactor and 48% for the GPR. The catalyst productivity was 50 kg polypropylene per gram catalyst.

Example 8

Polypropylene homopolymer is produced using a continuously operating pilot plant. The equipment includes catalyst, alkanes, donors, propylene feed system, prepolymerization reactor, loop reactor and fluidized bed gas phase reactor (GPR).

Catalyst, alkanes, donors and propylene are fed into the prepolymerization reactor. The polymer slurry from the prepolymerization reactor is fed to the loop reactor, and hydrogen and more propylene are simultaneously fed to the loop reactor. The polymer slurry from the loop reactor was fed to the GPR with additional hydrogen and propylene.

The formed polymer and unreacted propylene were separated after removal from the GPR.

The catalyst used is a highly active and stereospecific ZN-catalyst prepared according to US Patent 5234879. The catalyst was mixed with triethylaluminum (TEA) and dicyclopentyldimethoxysilane (DCPDMS) (Al/Ti ratio 250, Al/Do 40 (mol) before being sent to the prepolymerization reactor. ))touch.

The catalyst was fed according to US Patent 5385992 and flushed with propylene into the prepolymerization reactor. The prepolymerization reactor was operated at a pressure of 58 bar and a temperature of 20°C with an average catalyst residence time of 7 minutes.

The prepolymerized catalyst, propylene and other components are transferred to the loop reactor.

The loop reactor was operated at a pressure of 57 bar and a temperature of 80°C with an average catalyst residence time of 2 hours. The MFR (2.16 kg, 230° C.) of the polypropylene homopolymer produced in the loop reactor was adjusted to 375 by controlling the hydrogen feed.

The polymer slurry from the loop reactor was transferred to the GPR. The GPR was operated at a total pressure of 29 bar and a propylene partial pressure of 16 bar. The temperature of the reactor was 80°C and the average residence time of the catalyst was 2 hours. The MFR (2.16 kg, 230° C.) of the polypropylene homopolymer produced by the GPR was 450 and was adjusted by controlling the hydrogen partial pressure and by controlling the throughput ratio between reactors. The throughput ratios were adjusted as follows: 1% for the prepolymerization reactor, 50% for the loop reactor and 49% for the GPR.

Example 9

Production of polypropylene random polymers using a continuously operating pilot plant. The equipment includes catalyst, alkanes, donors, propylene and ethylene feed systems, loop reactors and a fluidized bed gas phase reactor (GPR).

Catalyst, alkanes, donors and propylene are fed into the prepolymerization reactor. The polymer slurry from the loop reactor is fed to the GPR with additional hydrogen, propylene and ethylene. The formed polymer and unreacted propylene were separated after removal from the GPR.

The catalyst used is a highly active and stereospecific ZN-catalyst prepared according to US Patent 5234879. The catalyst was prepolymerized batchwise with propylene according to Finnish patent 95387 (mass ratio polypropylene/catalyst 10). The prepolymerized catalyst was mixed with triethylaluminum (TEA) and dicyclopentyldimethoxysilane (DCPDMS) (Al/Ti ratio 140, Al/Do 10) before feeding into the loop reactor. (mole)) contact.

The catalyst was fed according to US Patent 5385992 and flushed with propylene into the loop reactor. The loop reactor was operated at a pressure of 50 bar and a temperature of 75°C with an average catalyst residence time of 1 hour. The MFR (2.16 kg, 230° C.) of the polypropylene random polymer produced in the loop reactor was adjusted to 4 by controlling the hydrogen feed. The ethylene content was controlled to 3.5% by weight by controlling the ethylene feed.

The polymer slurry from the loop reactor was transferred to the 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 the GPR was 80°C and the average residence time of the catalyst was 1.5 hours. The MFR (2.16 kg, 230° C.) of the polypropylene random polymer produced by the GPR was adjusted to 4 by controlling the partial pressure of hydrogen. The ethylene content was adjusted to 3.5% by weight by controlling the ethylene partial pressure. The throughput ratio between the reactors is: 55% for the loop reactor and 45% for the GPR.

Example 10

Production of polypropylene random polymers using a continuously operating pilot plant. The equipment includes catalyst, alkanes, donors, propylene and ethylene feed systems, loop reactors and a fluidized bed gas phase reactor (GPR).

Catalyst, alkanes, donors and propylene are fed into the prepolymerization reactor. The polymer slurry from the loop reactor was fed to the GPR with additional hydrogen and propylene. The formed polymer and unreacted propylene were separated after removal from the GPR.

The catalyst used is a highly active and stereospecific ZN-catalyst prepared according to US Patent 5234879. The catalyst was prepolymerized batchwise with propylene according to Finnish patent 95387 (mass ratio polypropylene/catalyst 10). The prepolymerized catalyst was mixed with triethylaluminum (TEA) and dicyclopentyldimethoxysilane (DCPDMS) (Al/Ti ratio 135, Al/Do 10) before feeding into the loop reactor. (mole)) contact.

The catalyst was fed according to US Patent 5385992 and flushed with propylene into the loop reactor. The loop reactor was operated at a pressure of 50 bar and a temperature of 75°C with an average catalyst residence time of 1 hour. The MFR (2.16 kg, 230° C.) of the polypropylene random polymer produced in the loop reactor was adjusted to 0.2 by controlling the hydrogen feed. The ethylene content was adjusted to 3.5% by weight by controlling the ethylene feed.

The polymer slurry from the loop reactor was transferred to the 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 the GPR was 80°C and the average residence time of the catalyst was 1.5 hours. The MFR (2.16 kg, 230° C.) of the polypropylene random polymer produced by the GPR was adjusted to 3 by controlling the partial pressure of hydrogen. The ethylene content was set to 1.8% by weight by adjusting the throughput ratio between the reactors. The desired ethylene content was obtained at a throughput ratio of 40% in the loop reactor and 60% in the 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 atactic polypropylene produced in the loop reactor was lower than 0.1 and the ethylene content was adjusted to 3.5% by weight by controlling the ethylene feed.

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 GPR-produced polypropylene copolymer was adjusted to 0.3 by controlling the hydrogen partial pressure. The ethylene content was set at 1.8% by weight by adjusting the throughput ratio between the reactors.

Ethylene at the loop reactor outlet is recovered from the flash gas and recycled back to the loop reactor. The propylene at the outlet of the GPR is recovered and sent to the loop reactor after hydrogen removal. The per-pass conversions of propylene and ethylene were 83% and 84%, respectively.

Example 11

A continuously operated pilot plant is used to produce polypropylene copolymers with good impact and creep properties. The plant included catalyst, alkanes, donors, propylene and ethylene feed systems, a prepolymerization reactor, a loop reactor and two fluidized bed gas phase reactors (GPR).

Catalyst, alkanes, donors and propylene are fed into the prepolymerization reactor. The polymer slurry from the prepolymerization reactor was fed to the loop reactor, and hydrogen, ethylene and additional propylene were simultaneously fed to the loop reactor.

The polymer slurry from the loop reactor was fed to the first GPR with additional hydrogen and propylene. The polymer from the first GPR is fed to the second GPR. Ethylene, some hydrogen and additional propylene are fed to the second GPR. The formed polymer and unreacted propylene are separated after discharge from the second GPR.

The catalyst used is a highly active and stereospecific ZN catalyst prepared according to US Patent 5,234,879. The catalyst was mixed with triethylaluminum (TEA) and dicyclopentyldimethoxysilane (DCPDMS) (Al/Ti ratio of 150, Al/Do of 10 (mol) before being sent to the prepolymerization reactor. ))touch.

The catalyst was fed according to US Patent 5385992 and flushed with propylene into the loop reactor. The prepolymerization reactor was operated at a pressure of 51 bar and a temperature of 20°C with an average catalyst residence time of 7 minutes.

The loop reactor was operated at a pressure of 50 bar and a temperature of 75°C with an average catalyst residence time of 1 hour. The MFR (2.16 kg, 230° C.) of the polypropylene random polymer produced in the loop reactor was adjusted to 7 by controlling the hydrogen feed. The ethylene content was adjusted to 3.5% by weight by controlling the ethylene feed.

The polymer slurry 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. Its 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 polypropylene random polymer produced by the GPR was adjusted to 10 by controlling the partial pressure of hydrogen. The ethylene content was set to 2% by weight by adjusting the throughput ratio between the reactors.

Transfer polymer from the first GPR to the second GPR. The second GPR was operated at a total pressure of 10 bar and a monomer partial pressure of 7 bar. Its 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 polypropylene copolymer produced by the GPR was adjusted to 7 by controlling the partial pressure of hydrogen. The ethylene content was set to 10% by weight by adjusting the partial pressure of ethylene and by controlling the throughput ratio between the reactors.

The desired properties were obtained at a throughput ratio of 1% in the prepolymerization reactor, 40% in the loop reactor, 40% in the first GPR and 19% in the second GPR.

Example 12

Production of very soft polypropylene copolymers using a continuously operating pilot plant. The equipment includes catalyst, alkanes, donors, propylene and ethylene feed systems, prepolymerization reactors, loop reactors and a fluidized bed gas phase reactor (GPR).

Catalyst, alkanes, donors and propylene are fed into the prepolymerization reactor. The polymer slurry from the prepolymerization reactor was sent to the loop reactor and simultaneously hydrogen, ethylene and additional propylene were fed to the loop reactor.

The polymer slurry from the loop reactor is sent to the GPR with additional ethylene, hydrogen and propylene. The formed polymer and unreacted monomer were separated after removal from the GPR.

The catalyst used is a highly active and stereospecific ZN catalyst prepared according to US Patent 5,234,879. The catalyst was mixed with triethylaluminum (TEA) and dicyclopentyldimethoxysilane (DCPDMS) (Al/Ti ratio of 150, Al/Do of 10 (mol) before being sent to the prepolymerization reactor. ))touch.

The catalyst was fed according to US Patent 5385992 and flushed with propylene into the loop reactor. The prepolymerization reactor was operated at a pressure of 51 bar and a temperature of 20° C. and an average catalyst residence time of 7 minutes.

The loop reactor was operated at a pressure of 50 bar and a temperature of 75°C with an average catalyst residence time of 1 hour. The MFR (2.16 kg, 230° C.) of the polypropylene random polymer produced in the loop reactor was adjusted to 4 by controlling the hydrogen feed. The ethylene content was adjusted to 3.8% by weight by controlling the ethylene feed.

The polymer slurry 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. Its operating temperature is 80° C. and the average residence time of the catalyst is 1.2 hours. The MFR (2.16 kg, 230° C.) of the polypropylene random polymer produced by the GPR was adjusted to 2.5 by controlling the partial pressure of hydrogen. The ethylene content was set to 8% by weight by adjusting the throughput ratio between the reactors and the partial pressure of ethylene.

The desired properties were obtained at a throughput ratio of 1% in the prepolymerization reactor, 45% in the loop reactor and 55% in the GPR.

The polymer from the GPR can be transferred to another GPR to produce a softer polypropylene copolymer by having a higher ethylene partial pressure in the second GPR.

Example 13

A continuously operating pilot plant is used to produce polypropylene copolymers with good creep properties. The equipment includes catalyst, alkanes, donors, propylene and ethylene feed systems, prepolymerization reactors, loop reactors and a fluidized bed gas phase reactor (GPR).

Catalyst, alkanes, donors and propylene are fed into the prepolymerization reactor. The polymer slurry from the prepolymerization reactor was sent to the loop reactor along with hydrogen and additional propylene to the loop reactor.

The polymer slurry from the loop reactor is sent to a flash tank where monomer and polymer are separated. The polymer from the flash tank is sent to GPR. The propylene from the flash tank is sent to the GPR after hydrogen removal. Ethylene, additional hydrogen and additional propylene are sent to the GPR.

The catalyst used is a highly active and stereospecific ZN-catalyst prepared according to US Patent 5234879. The catalyst was mixed with triethylaluminum (TEA) and dicyclopentyldimethoxysilane (DCPDMS) (Al/Ti ratio of 140, Al/Do of 10 (mol) before being sent to the prepolymerization reactor. ))touch.

The catalyst was fed according to US Patent 5385992 and flushed with propylene into the loop reactor. The prepolymerization reactor was operated at a pressure of 51 bar and a temperature of 20° C. and an average catalyst residence time of 7 minutes.

The loop reactor was operated at a pressure of 50 bar and a temperature of 75°C with an average catalyst residence time of 1 hour. The MFR (2.16 kg, 230° C.) of the polypropylene random polymer produced in the loop reactor was adjusted to 10 by controlling the hydrogen feed.

The GPR reactor was operated at a total pressure of 29 bar and a propylene partial pressure of 16 bar. Its operating temperature is 80° C. and the average residence time of the catalyst is 1.1 hours. The MFR (2.16 kg, 230° C.) of the polypropylene copolymer produced by the GPR was adjusted to 5 by controlling the partial pressure of hydrogen and by controlling the throughput ratio between the reactors. The ethylene content was adjusted to 3.5% by weight by controlling the throughput ratio between the reactors and the partial pressure of ethylene.

The desired properties were obtained at a throughput ratio of 1% in the prepolymerization reactor, 40% in the loop reactor and 59% in the GPR.

Polymer from a GPR can be transferred to another GPR to produce a polypropylene copolymer with better impact properties through higher ethylene partial pressure in the second GPR.

Claims (7)

1. A method for preparing propylene copolymer, the steps it comprises are as follows: - polymerizing propylene with comonomers in at least one slurry reactor and at least one gas phase reactor at an elevated temperature of 60-85° C. and an elevated pressure of 4000-8000 kPa in the presence of a Zn catalyst, at least 10% by weight of polymerized product is produced in said gas phase reactor, - recovery of copolymerization product comprising unreacted monomer from the slurry reactor; and - The copolymerization product is introduced into the first gas phase reactor, essentially no unreacted monomer is recycled back to the slurry reactor preceding the gas phase reactor. 2. The process according to claim 1, wherein the slurry reactor is operated at a temperature of 60-75°C to produce random copolymers or terpolymers. 3. The process according to claim 1, wherein the slurry reactor is operated at a temperature of 75-85°C. 4. A process according to claim 1, wherein the polymerized product from the first gas phase reactor is copolymerized by subjecting it to a further copolymerization step in the presence of additional comonomer. 5. The method of claim 1, wherein a temperature of at least 80°C; At least 40% by weight of the polymerization product is produced in the gas phase reactor. 6. The method according to claim 5, wherein said polymerized product is copolymerized in the presence of a comonomer, thereby providing a polymerized product with improved impact properties. 7. Process according to claim 5, wherein said copolymerization is carried out in a second gas phase reactor arranged in series with the first gas phase reactor.

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