Improving catalyst performance in multi-stage polyolefin production

Abstract

The present disclosure relates to a method for polymerizing olefins in a multi-stage polymerization process configuration, the method comprising a) polymerizing a first olefin monomer in a first polymerization step, optionally in the presence of at least one other alpha olefin comonomer, in the presence of a metallocene polymerization catalyst to form a first polymer component (A); and b) transferring the first polymer component (A) to a separation unit to remove low molecular weight penetrants and obtain separated solid polyolefin particles (A*) of the first polymer component, and c) polymerizing a second olefin monomer in a second polymerization step in the gas phase, optionally in the presence of at least one other alpha olefin comonomer, in the presence of the separated solid polyolefin particles (A*) of step b) to form a second polymer component (B). The present disclosure also relates to a method for improving the performance of a metallocene polymerization catalyst in a multi-stage olefin polymerization, wherein the first polymer component (A) produced in the first polymerization step is transferred to a separation unit to remove low molecular weight penetrants and obtain separated solid polyolefin particles (A*) of the first polymer component, and then the obtained separated solid polyolefin particles (A*) of the first polymer component are transferred to another polymerization step.

Description

Technical Field

[0001] The present disclosure relates to the copolymerization of olefins and, more particularly, to a multi-stage polyolefin production process. The present disclosure also relates to a method for improving the performance of metallocene polymerization catalysts by selecting flash distillation operating conditions for a high-pressure flash distillation used to remove unreacted components and diluent from polymer particles before entering subsequent polymerization stages in a multi-stage olefin polymerization process. The present disclosure also relates to the effect of particle size distribution progression in a gas phase reactor (GPR), which in turn can eliminate solids transfer, sheeting, and agglomeration. Background Art

[0002] Multi-stage polyolefin production processes (e.g. Borstar PE, PP and Spheripol PP) consist of a multi-stage reactor configuration to provide multi-modal capabilities for achieving easy processing of resins with desired mechanical properties. In this type of process, a series of slurry loop reactors are used in combination, followed by a gas phase reactor, to produce various polyolefins.

[0003] One of the key features of a multi-stage olefin polymerization process is to ensure appropriate catalyst performance in all stages of the multi-stage polymerization process, more specifically, the appropriate selection of gas phase reactor operating conditions that will lead to smooth operability in the GPR. This can be challenging when utilizing single site catalysts that have superior comonomer incorporation capabilities compared to 1st generation catalysts. The presence of small sized particles (also known as Stocke's particles: particles with a buoyancy higher than gravity in a gas-solid fluidized environment) that tend to be entrained by the fluidizing gas can give rise to problems associated with reactor fouling (polymer coating on the reactor walls), sheeting and agglomeration, as well as fouling of the recycle gas compressor and heat exchanger units. In this context, optimizing catalyst performance in terms of eliminating the overall number of small sized particles in the GPR is of paramount importance and represents a key aspect for the successful implementation of catalysts in multi-stage ethylene copolymerization processes. Summary of the Invention

[0004] It is an object of the present disclosure to provide a process for polymerizing olefins in a multi-stage polymerization process configuration that overcomes the above-mentioned problems.

[0005] The objects of the present disclosure are achieved by a method which is characterized by what is stated in the independent claim. Preferred embodiments of the present disclosure are disclosed in the dependent claims.

[0006] The present disclosure is based on the idea of ​​transferring a first polymer component produced in a first polymerization step to a separation unit to remove low molecular weight permeants and obtain separated solid polyolefin particles (A*) of the first polymer component, and then transferring this first polymer component to a second polymerization step where further polymerization takes place. BRIEF DESCRIPTION OF THE DRAWINGS

[0007] The present disclosure will be described in more detail below by way of preferred embodiments with reference to the accompanying drawings, in which

[0008] Figure 1 An embodiment of the polymerization process according to the invention is shown in a fluidized bed reactor with a fluidization grid. DETAILED DESCRIPTION

[0009] The present disclosure relates to a method for polymerizing olefins in a multi-stage polymerization process configuration, the method comprising:

[0010] (a) polymerizing a first olefin monomer in a first polymerization step, optionally in the presence of at least one other alpha olefin comonomer, in the presence of a metallocene polymerization catalyst to form a first polymer component (A); and

[0011] (b) transferring the first polymer component (A) to a separation unit to remove the low molecular weight penetrant and obtain separated solid polyolefin particles (A*) of the first polymer component, and

[0012] (c) polymerizing in a second polymerization step in gas phase a second olefin monomer, optionally in the presence of at least one other alpha olefin comonomer, in the presence of the separated solid polyolefin particles (A*) of step b) to form a second polymer component (B).

[0013] The present method defines the optimal setting of operating conditions (e.g. pressure and residence time of the particulate polymer material in the high-pressure flash separator) during the flashing of polyethylene particulate material in a flash separator (e.g. a loop and a GPR) positioned between the first polymerization step and the second polymerization step in a multi-stage polymerization process, in order to eliminate the risk of fines generation and to achieve a high initial polymerization rate in the GPR due to adsorbed 1-butene, thereby improving the catalyst performance during the polymerization process.

[0014] Therefore, the present disclosure also relates to a method for improving the performance of a metallocene polymerization catalyst in a multi-stage olefin polymerization process, wherein the first polymer component (A) produced in the first polymerization step is transferred to a separation unit to remove a low-molecular penetrant and obtain separated solid polyolefin particles (A*) of the first polymer component, and then the obtained separated solid polyolefin particles (A*) of the first polymer component are transferred to another polymerization step.

[0015] method

[0016] The present disclosure relates to a multi-stage polymerization process using a metallocene polymerization catalyst, the process comprising an optional but preferred prepolymerization step followed by a first polymerization step and a second polymerization step.

[0017] Preferably, the same catalyst is used in each step and ideally, the catalyst is transferred sequentially from the prepolymerization step to the subsequent polymerization step in a known manner. A preferred process configuration is based on Type cascades, especially 2G type cascade, preferably 3G type cascade.

[0018] Thus, the process of the present invention for polymerizing olefins in a multi-stage polymerization process configuration comprises:

[0019] (a) polymerizing a first olefin monomer in a first polymerization step, optionally in the presence of at least one other alpha olefin comonomer, to form a first polymer component (A); and

[0020] (b) transferring the first polymer component (A) to a separation unit to remove the low molecular weight penetrant and obtain separated solid polyolefin particles (A*) of the first polymer component, and

[0021] (c) polymerizing in a second polymerization step in gas phase a second olefin monomer, optionally in the presence of at least one other alpha olefin comonomer, in the presence of the separated solid polyolefin particles of step b) to form a second polymer component (B).

[0022] Typically, the first olefin monomer polymerized in the first polymerization step and the second olefin monomer polymerized in the second polymerization step are the same olefin monomer, in particular ethylene. At least one olefin comonomer may optionally be present in either step, and it may be the same or different comonomer in both steps or present in only one of the steps as described below.

[0023] Prepolymerization step

[0024] The polymerization step may be preceded by a prepolymerization step. The purpose of the prepolymerization is to polymerize a small amount of polymer onto the catalyst at low temperatures and / or low monomer concentrations. Prepolymerization may improve the performance of the catalyst in the slurry and / or change the characteristics of the final polymer. The prepolymerization step is preferably carried out in a slurry and the amount of the polymer produced in the optional prepolymerization step is calculated as the amount (weight %) of the ethylene polymer component (A).

[0025] When a prepolymerisation step is present, preferably the catalyst components are all introduced into the prepolymerisation step. Preferably, the reaction product of the prepolymerisation step is then introduced into the first polymerisation step.

[0026] However, in the case where the solid catalyst component and the cocatalyst can be fed separately, it is possible to introduce only a portion of the cocatalyst into the prepolymerization stage and the remainder into the subsequent polymerization stage. In such a case as well, it is necessary to introduce into the prepolymerization stage so much cocatalyst that a sufficient polymerization reaction is achieved.

[0027] It is understood that within the scope of the present invention the amount of polymer produced in the prepolymerisation is in the range of 1 to 7 wt% relative to the final multimodal copolymer. This can be considered as part of the first ethylene polymer component (A) produced in the first polymerisation step a).

[0028] First polymerization step a)

[0029] In the present process, the first polymerization step a) involves polymerizing an olefin monomer and optionally at least one olefin comonomer.

[0030] In one embodiment, the first polymerization step involves polymerizing ethylene to produce an ethylene homopolymer.

[0031] In another embodiment, the first polymerization step involves polymerizing ethylene and at least one olefin comonomer to produce an ethylene copolymer.

[0032] The first polymerization step may take place in any suitable reactor or series of reactors. The first polymerization step may be carried out in one or more slurry polymerization reactors or in a gas phase polymerization reactor or a combination thereof. Preferably, the first polymerization step is carried out in one or more slurry polymerization reactors, more preferably in at least three slurry phase reactors, including a slurry phase reactor for carrying out prepolymerization.

[0033] The polymerization in the first polymerization zone is preferably carried out in a slurry. The polymer particles formed in the polymerization are then suspended in a fluid hydrocarbon together with the catalyst broken up and dispersed within the particles. The slurry is stirred to transfer the reactants from the fluid to the particles.

[0034] Slurry polymerization often occurs in an inert diluent (typically a hydrocarbon diluent such as methane, ethane, propane, n-butane, isobutane, pentane, hexane, heptane, octane, or a mixture thereof). Preferably, the diluent is a low-boiling hydrocarbon having 1 to 4 carbon atoms or a mixture of such hydrocarbons. A particularly preferred diluent is propane, which may contain small amounts of methane, ethane, and / or butane.

[0035] The ethylene content of the fluid phase of the slurry may be from 2 mol% to about 50 mol%, preferably from about 3 mol% to about 20 mol%, and particularly from about 5 mol% to about 15 mol%. The benefit of having a high ethylene concentration is increased catalyst productivity, but the disadvantage is that more ethylene needs to be recovered than at lower concentrations.

[0036] The temperature of the slurry polymerization is generally 50 to 115° C., preferably 60 to 110° C., in particular 70 to 100° C. The pressure is from 1 to 150 bar, preferably 10 to 100 bar.

[0037] The pressure in the first polymerization step is generally from 35 to 80 bar, preferably from 40 to 75 bar, in particular from 45 to 70 bar.

[0038] The residence time of the first polymerization stage is generally from 0.15 to 3.0 hours, preferably from 0.20 to 2.0 hours, in particular from 0.30 to 1.5 hours.

[0039] Sometimes it is advantageous to carry out a slurry polymerization at a pressure above the critical temperature and fluid mixture. Such an operation is described in US-A-5391654. In such an operation, the temperature is generally 85°C to 110°C, preferably 90°C to 105°C, and the pressure is 40 bar to 150 bar, preferably 50 bar to 100 bar.

[0040] Slurry polymerization can be carried out in any known reactor for slurry polymerization. Such reactors include continuous stirred tank reactors and loop reactors. Particularly preferred is carrying out the polymerization in a loop reactor. In such a reactor, the slurry is circulated at high speed along a closed tube using a circulation pump. Loop reactors are generally known in the art, and examples thereof are given, for example, in US-A-4582816, US-A-3405109, US-A-3324093, EP-A-479186 and US-A-5391654.

[0041] The slurry can be drawn from the reactor continuously or intermittently. A preferred mode for intermittent drawing is to use settling legs that allow the slurry to concentrate and then draw a batch of concentrated slurry from the reactor. The use of settling legs is disclosed in US-A-3374211, US-A-3242150, and EP-A-1310295, etc. Continuous drawing is disclosed in EP-A-891990, EP-A-1415999, EP-A-1591460, and WO-A-2007 / 025640, etc. Continuous drawing is advantageously combined with a suitable concentration method, as disclosed in EP-A-1310295, EP-A-1591460, and EP3178853B1.

[0042] Hydrogen can be fed into the reactor to control the molecular weight of the polymer, as is known in the art. Additionally, one or more alpha-olefin comonomers can be added to the reactor to control the density of the polymer product. The actual amount of such hydrogen and comonomer feed depends on the catalyst used and the desired melt index (or molecular weight) and density (or comonomer content) of the resulting polymer.

[0043] Hydrogen can be fed into the reactor to control the molecular weight of the polymer, as is known in the art. Additionally, one or more alpha-olefin comonomers can be added to the reactor to control the density of the polymer product. The actual amount of such hydrogen and comonomer feed depends on the catalyst used and the desired melt index (or molecular weight) and density (or comonomer content) of the resulting polymer.

[0044] Separation step b)

[0045] The term "low molecular weight penetrant" refers to C2-6 hydrocarbons, particularly α-C2-6 olefins (eg, ethylene, 1-butene, 1-hexene) and C2-6 alkanes (eg, propane) and chain transfer agents (eg, hydrogen).

[0046] The isolated solid polyolefin particles obtained exhibit a particle size distribution with an average particle size (D50) between 100 μm and 600 μm, preferably between 120 μm and 550 μm, most preferably between 150 μm and 500 μm, and a span (span: (d90-d10) / d50) between 0.8 and 3.5, more preferably between 1.0 and 2.5, and most preferably between 1.2 and 2.0. The concentration of unreacted components, diluents and chain transfer agents in the polymer particles is 1*10 -6 g / gampol (grams of components per gram of amorphous polymer) to 0.9 g / gpol, more preferably 1*10 -5 g / gampol to 0.8g / gampol, most preferably 5*10 -5 g / gampol to 0.7g / gampol.

[0047] The separation step can be carried out in any known separator for removing volatile hydrocarbons from solids. Such units include flash separators.

[0048] The operating pressure of the separation unit, in particular the flash separator, is in the range of 1 to 25 bar, preferably from 2 to 22 bar, more preferably from 3 to 20 bar.

[0049] The operating temperature of the separation unit, in particular the flash separator, is in the range of 40 to 90°C, preferably from 50 to 85°C, more preferably from 60 to 80°C.

[0050] The average residence time of the particulate matter in the separation unit, in particular the flash separator, is in the range of 2 minutes to 30 minutes, preferably from 3 minutes to 25 minutes, more preferably from 5 minutes to 20 minutes.

[0051] It has been surprisingly discovered that the residual 1-butene concentration of the polymer particles exiting the loop reactor and entering the gas phase reactor is a key factor in determining the initial particle growth rate in the GPR, particularly when a "comonomer-sensitive" single-site polyethylene catalyst is used in the process. Furthermore, operating the flash separator at high pressure helps maintain a high comonomer concentration in the polymer particles while reducing stresses induced in the polymer particles due to the pressure difference between the operating pressures in the loop and flash separator. This leads to the conclusion that the flash conditions in the high-pressure flash play an important role in determining the particle size distribution in the gas phase, and these flash conditions are associated with operational challenges such as sheeting, agglomeration, and excessive solids transfer.

[0052] Second polymerization step c)

[0053] The first polymer component (A) is transferred as separated solid polyolefin particles (A*) from the first polymerization step a) to the second polymerization step c) via a separation step b).

[0054] In the present process, the second polymerization step c) involves polymerizing an olefin monomer and optionally at least one olefin comonomer. The olefin monomer of the second polymerization step b) (i.e., the second olefin monomer) is typically the same as the olefin monomer of the first polymerization step a) (i.e., the first olefin monomer). Preferably, the olefin monomer is ethylene.

[0055] In one embodiment, the second polymerization step involves polymerizing ethylene and optionally at least one olefin comonomer to produce an ethylene homopolymer or an ethylene copolymer, respectively.

[0056] The second polymerization step takes place in one or more gas phase polymerization reactors.

[0057] Gas phase polymerization can be carried out in any known reactor for gas phase polymerization. Such reactors include fluidized bed reactors, fast fluidized bed reactors, or settled bed reactors, or any combination of these reactors. When a combination of reactors is used, polymer is transferred from one polymerization reactor to another. In addition, some or all of the polymer from a polymerization stage can be returned to the previous polymerization stage.

[0058] Gas phase polymerization is carried out in gas-solid fluidized beds, which are also called gas phase reactors (GPRs). Gas-solid olefin polymerization reactors are commonly used for the polymerization of α-olefins such as ethylene and propylene because they allow relatively high flexibility in polymer design and the use of various catalyst systems. A common gas-solid olefin polymerization reactor variant is the fluidized bed reactor.

[0059] In a gas-solid olefin polymerization reactor, polymerization is carried out using gaseous olefin monomers in which polymer particles grow.

[0060] The present method is applicable to any type of gas-solid olefin polymerization reactor suitable for the polymerization of α-olefin homopolymers or copolymers. Suitable reactors are, for example, continuous stirred tank reactors or fluidized bed reactors. Both types of gas-solid olefin polymerization reactors are well known in the art.

[0061] Preferably, the gas-solid olefin polymerization reactor is a fluidized bed reactor.

[0062] In a fluidized bed reactor, polymerization is carried out in a fluidized bed formed by polymer particles growing in an upward moving gas stream. In the fluidized bed, the polymer particles containing the active catalyst are contacted with a reaction gas (such as monomer, comonomer and hydrogen), which causes the polymer to be produced on the particles.

[0063] Therefore, in a preferred embodiment, the fluidized bed reactor may include a fluidizing grid located below the fluidized bed, thereby separating the bottom zone and the middle zone of the reactor. The upper limit of the fluidized bed is usually defined by the separation zone, where, due to its enlarged diameter compared to the middle zone, the fluidizing gas expands and the gas is separated from the polyolefin powder. Fluidized bed reactors with a separation zone and a fluidizing grid are well known in the art. Such a fluidized bed reactor suitable for the process of the present invention is shown in FIG. Figure 1 middle.

[0064] In another preferred embodiment, the fluidized bed reactor does not include a fluidizing grid. Polymerization is carried out in a reactor comprising a bottom zone, a middle zone, and a top zone. The bottom zone, having a generally conical shape, forms the lower portion of the reactor and forms the base of the fluidized bed in this bottom zone. The base of the bed is formed in the bottom zone in the absence of a fluidizing grid or a gas distribution plate. Above the bottom zone and in direct contact with it is the middle zone, which has a generally cylindrical shape. The upper portion of the bottom zone and the middle zone contain the fluidized bed. Because there is no fluidizing grid, a free exchange of gases and particles occurs between the different zones within the bottom zone and between the bottom zone and the middle zone. Finally, above the middle zone and in direct contact with it is the top zone, which has a generally conical shape that tapers upward.

[0065] The bottom zone of the reactor has a generally conical shape that tapers downward. Due to the shape of this zone, the gas velocity gradually decreases along the height in the bottom zone. The gas velocity in the lowest part is greater than the conveying velocity, and the particles ultimately contained in the gas are transported upward with the gas. At a certain height in the bottom zone, the gas velocity becomes less than the conveying velocity, and the fluidized bed begins to form. When the gas velocity becomes smaller, the bed becomes denser, and the polymer particles distribute the gas over the entire cross-section of the bed. EP-A-2 495 037, EP-A-2 495 038, EP3103818B1, EP2913346B1, EP3418308B1 and EP3642246B1 describe this fluidized bed reactor without a fluidizing grid.

[0066] In gas-solid olefin polymerization reactor, by from the top area of ​​reactor, usually at the highest position, draw the fluidizing gas stream as the second gas stream to set up the gas stream that moves upward.Then usually the second gas stream that will be drawn from reactor is cooled and reintroduced into the bottom area of ​​reactor as the first fluidizing gas stream.In a preferred embodiment, the fluidizing gas of the second gas stream is also compressed in a compressor.More preferably, compressor is positioned at the upstream of cooler.Preferably, before gas is sent to compressor, gas is filtered.Other olefin monomer, final comonomer, hydrogen and inert gas are suitably introduced in the recycle gas pipeline.Preferably for example by using the composition of online gas chromatography analysis recycle gas, and regulating the adding of gaseous component so that their content remains on desired level.

[0067] The temperature of the gas phase polymerization is usually 50 to 100°C, preferably 65 to 90°C.

[0068] The pressure of the gas phase polymerization is generally from 5 bar to 40 bar, preferably from 10 bar to 35 bar, preferably from 15 bar to 30 bar.

[0069] The residence time for the gas phase polymerization is from 1.0 to 4.5 hours, preferably from 1.5 to 4.0 hours, in particular from 2.0 to 3.5 hours.

[0070] The polymer production rate in the gas phase reactor can be from 10 tn / h to 65 tn / h, preferably from 12 tn / h to 58 tn / h, especially from 13 tn / h to 52.0 tn / h, and thus the total polymer uptake rate from the gas phase reactor can be from 15 tn / h to 100 tn / h, preferably from 18 tn / h to 90 tn / h, especially from 20 tn / h to 80.0 tn / h.

[0071] The production split (% of second polymer component (B) / % of first polymer component (A)) may be from 0.65 to 2.5, preferably from 0.8 to 2.3, most preferably from 1.0 to 1.65.

[0072] Figure 1 An embodiment of the polymerization process according to the invention is shown in a fluidized bed reactor with a fluidization grid.

[0073] Reference Mark

[0074] 1 Fluidized bed reactor

[0075] 2 Bottom area

[0076] 3. Middle Zone

[0077] 4 Top area (separation area)

[0078] 5 Fluidized bed (dense zone)

[0079] 6. First fluidizing gas flow

[0080] 7 Second fluidizing gas flow

[0081] 8 Compressor

[0082] 9 Compressed second fluidizing gas flow

[0083] 10 Cooler

[0084] 11 Cooled Second Fluidizing Gas Stream

[0085] 12 Feed line for induced swelling agent

[0086] 13 Feed lines for fresh monomer, comonomer, chain transfer agent and diluent

[0087] 14 Feed port for metallocene polymerization catalyst

[0088] 15 Polymer Absorption

[0089] 16 Fluidized grid

[0090] Figure 1 Description

[0091] Figure 1An embodiment of a gas-solid olefin polymerization reactor system according to the present invention is shown. A fluidized bed reactor (1) comprises a bottom zone (2), an intermediate zone (3) and a separation zone (4) as a top zone. The intermediate zone (3) and the bottom zone (2) are separated by a fluidizing grid (16). A first fluidizing gas stream (6) enters the fluidized bed reactor (1) through the bottom zone (2) and flows upward, thereby passing through the fluidizing grid (16) and entering the intermediate zone (3). Due to the roughly cylindrical shape of the intermediate zone (3), the gas velocity is constant, so that above the fluidizing grid (16), a fluidized bed (5) is established in the intermediate zone (3). Due to the conical shape of the separation zone (4), the gas entering the separation zone (4) expands, so that the gas is separated from the polyolefin product of the polymerization reaction, so that the fluidized bed (5) is confined to the lower part of the intermediate zone (3) and the separation zone (4). The metallocene polymerization catalyst is introduced into the fluidized bed reactor (1) together with the optional polyolefin powder polymerized in the previous polymerization stage directly into the fluidized bed (5) through the feed opening (14). The polyolefin product of the polymerization process is withdrawn from the fluidized bed reactor through the outlet (15).

[0092] Fluidizing gas is withdrawn from the separation zone (4) as a second fluidizing gas stream (7) and introduced into a compressor (8). The compressed second gas stream (9) is withdrawn from the compressor (8) and introduced into a cooler (10). An inducing swelling agent is introduced into the second cooling gas stream (11) via a feed line (12), and fresh monomer, comonomer, diluent, and chain transfer agent are introduced into the second cooling gas stream (11) via a feed line (13).

[0093] polymerization catalysts

[0094] The polymerization catalyst used in the present method is a metallocene catalyst. Metallocene polymerization catalysts generally comprise (i) a transition metal complex, (ii) a cocatalyst and optionally (iii) a support.

[0095] Preferably, the first polymerization step and the second polymerization step are performed using the same metallocene catalyst, ie in the presence of the same metallocene catalyst.

[0096] The present method preferably utilizes single-site catalysis. Unlike Ziegler-Natta catalysis, polyethylene copolymers prepared using single-site catalysis have properties that allow them to be distinguished from Ziegler-Natta materials. In particular, the comonomer distribution is more uniform. This can be shown using TREF or Crystaf techniques. Catalyst residues can also indicate the catalyst used. Ziegler-Natta catalysts do not contain, for example, Zr or Hf Group (IV) metals.

[0097] Due to the low C6 concentration used in the gas phase reactor of the second polymerization step b), the present invention is particularly important for single site catalysts.

[0098] Transition metal complexes (i)

[0099] The transition metal complex comprises a transition metal (M) from Groups 3 to 10 of the Periodic Table (IUPAC 2007) or an actinide or lanthanide transition metal.

[0100] The term "transition metal complex" according to the present invention includes any metallocene or non-metallocene compound of a transition metal which carries at least one organic (coordinating) ligand and which exhibits catalytic activity alone or together with a cocatalyst. Transition metal compounds are well known in the art and the present invention encompasses compounds of metals from Groups 3 to 10, for example Groups 3 to 7 or Groups 3 to 6, such as Groups 4 to 6, and the lanthanides or actinides of the Periodic Table (IUPAC 2007).

[0101] In one embodiment, the transition metal complex (i) has the following formula (ii):

[0102] (L) m R n MX q (iI)

[0103] in

[0104] "M" is a transition metal (M) of Groups 3 to 10 of the Periodic Table (IUPAC 2007),

[0105] Each "X" is independently a monoanionic ligand, such as a sigma-ligand,

[0106] Each "L" is independently an organic ligand coordinated to the transition metal "M",

[0107] "R" is a bridging group connecting the organic ligand (L),

[0108] "m" is 1, 2 or 3, preferably 2,

[0109] "n" is 0, 1 or 2, preferably 0 or 1,

[0110] "q" is 1, 2 or 3, preferably 2, and

[0111] m+q equals the valence of the transition metal (M).

[0112] "M" is preferably selected from zirconium (Zr), hafnium (Hf) or titanium (Ti), more preferably selected from zirconium (Zr) and hafnium (Hf). "X" is preferably a halogen, most preferably Cl.

[0113] More preferably, the transition metal complex (i) is a metallocene complex comprising a transition metal compound as defined above comprising a cyclopentadienyl, indenyl or fluorenyl ligand as a substituent "L". In addition, the ligand "L" may have one or more substituents such as an alkyl group, an aryl group, an arylalkyl group, an alkylaryl group, a silyl group, a siloxy group, an alkoxy group or other heteroatom groups. Suitable metallocene catalysts are known in the art and are disclosed in WO-A-95 / 12622, WO-A-96 / 32423, WO-A-97 / 28170, WO-A-98 / 32776, WO-A-99 / 61489, WO-A-03 / 010208, WO-A-03 / 051934, WO-A-03 / 051514, WO-A-2004 / 085499, EP-A-1752462 and EP-A-1739103, among others.

[0114] In one embodiment of the present invention, the metallocene complex is bis(1-methyl-3-n-butylcyclopentadienyl)zirconium(IV) dichloride.

[0115] In another embodiment, the transition metal complex (i) has the following formula (i-II):

[0116]

[0117] wherein each X is independently a halogen atom, a C1-6 alkyl group, a C1-6 alkoxy group, a phenyl group or a benzyl group;

[0118] Each Het is independently a monocyclic heteroaromatic group containing at least one heteroatom selected from O or S;

[0119] L is -R'2Si-, wherein each R' is independently a C1-20 hydrocarbon group or a C1-10 alkyl group substituted with an alkoxy group having 1 to 10 carbon atoms;

[0120] M is Ti, Zr or Hf;

[0121] Each R1 is the same or different and is a C1-6 alkyl group or a C1-6 alkoxy group;

[0122] Each n is 1 to 2;

[0123] Each R2 is the same or different and is a C1-6 alkyl group, a C1-6 alkoxy group or a -Si(R)3 group;

[0124] each R is a C1-10 alkyl or phenyl group optionally substituted with 1 to 3 C1-6 alkyl groups; and

[0125] Each p is 0 to 1.

[0126] Preferably, the compound of formula (i-II) has structure (i-III):

[0127]

[0128] wherein each X is independently a halogen atom, a C1-6 alkyl group, a C1-6 alkoxy group, a phenyl group or a benzyl group;

[0129] L is Me2Si-;

[0130] Each R1 is the same or different and is a C1-6 alkyl group, such as methyl or t-Bu;

[0131] Each n is 1 to 2;

[0132] R2 is a -Si(R)3 alkyl group; each p is 1;

[0133] Each R is a C1-6 alkyl group or a phenyl group.

[0134] Highly preferred transition metal complexes of formula (i-II) are

[0135]

[0136] Co-catalyst (ii)

[0137] To form the polymerization catalyst, a cocatalyst, also known as an activator, is used, as is well known in the art. Cocatalysts comprising Al or B are well known and can be used herein. Preferably, aluminoxanes (e.g., MAO) or boron-based cocatalysts (such as borates) are used.

[0138] Suitable cocatalysts are metal alkyl compounds, especially alkylaluminum compounds known in the art. Particularly suitable activators for use with metallocene catalysts are alkylaluminumoxy compounds, such as methylaluminoxane (MAO), tetraisobutylaluminoxane (TIBAO) or hexaisobutylaluminoxane (HIBAO).

[0139] Preferably, the cocatalyst is methylaluminoxane (MAO).

[0140] Carrier (iii)

[0141] According to the scheme in WO03 / 051934, it is possible to use the present polymerization catalyst in solid but unsupported form. Preferably, the present polymerization catalyst is used in solid supported form. The particulate support material used may be an inorganic porous support such as silica, alumina or a mixed oxide such as silica-alumina, in particular silica.

[0142] Preference is given to using a silica support.

[0143] Particularly preferably, the support is a porous material so that the complex can be loaded into the pores of the particulate support, for example using methods similar to those described in WO94 / 14856, WO95 / 12622, WO2006 / 097497 and EP1828266.

[0144] The average particle size of a support such as a silica support may typically be from 10 μm to 100 μm. The average particle size (ie, median particle size, D 50 ) can be measured using a laser diffraction particle size analyzer Malvern Mastersizer 3000, sample dispersion: dry powder.

[0145] The average pore size of the support, such as a silica support, may be in the range of 10 nm to 100 nm, and the pore volume in the range of 1 mL / g to 3 mL / g.

[0146] Examples of suitable support materials are for example ES757 produced and sold by PQ Corporation, Sylopol 948 produced and sold by Grace or SUNSPERA DM-L-303 produced by AGC Si-Tech Co. The support may optionally be calcined before use in catalyst preparation in order to achieve an optimum silanol group content.

[0147] The catalyst may contain 5 to 500 μmol, such as 10 to 100 μmol, of transition metal per gram of support, such as silica, and 3 to 15 mmol of Al per gram of support, such as silica.

[0148] Multimodal polyethylene polymers

[0149] The present invention relates to the preparation of multimodal polyethylene homopolymer or copolymer. The density of the multimodal ethylene homopolymer or copolymer can be between 900 kg / m 3 Up to 980kg / m 3 between, preferably between 905kg / m 3 Up to 940kg / m 3 between, especially 910kg / m 3 Up to 935kg / m 3 between.

[0150] Preferably, the multimodal polyethylene polymer is a copolymer. More preferably, the multimodal polyethylene copolymer is LLDPE. Its density may be 905 kg / m 3 Up to 940kg / m 3 , preferably 910 kg / m 3 Up to 935kg / m 3 , more preferably 915kg / m3 Up to 930kg / m 3 , especially 916kg / m 3 Up to 928kg / m 3 In one embodiment, 910 kg / m 3 Up to 928kg / m 3 The term "LLDPE" as used herein means linear low density polyethylene. The LLDPE is preferably multimodal.

[0151] The term "multimodal" includes polymers that are multimodal with respect to the MFR and thus also includes bimodal polymers.The term "multimodal" may also mean multimodality with respect to the "comonomer distribution".

[0152] In general, polymers are said to be "multimodal" if they comprise at least two polyethylene fractions, which have been produced under different polymerisation conditions, resulting in the fractions having different (weight average) molecular weights and molecular weight distributions. The prefix "multi" relates to the number of different polymer fractions present in the polymer. Thus, for example, the term "multimodal polymer" includes so-called "bimodal" polymers consisting of two fractions. The form of the molecular weight distribution curve, i.e. the appearance of a plot of the polymer weight fraction as a function of its molecular weight, of a multimodal polymer such as LLDPE may show two or more maxima, or at least be significantly broadened compared to the curves for the individual fractions. Typically the final MWD curve will be broad, needle-shaped or show a shoulder.

[0153] Ideally, the molecular weight distribution curve of the multimodal polymer used in the present invention will show two distinct maxima. Alternatively, the polymer fractions have similar MFR and are bimodal with respect to comonomer content. A polymer comprising at least two polyethylene fractions produced under different polymerization conditions resulting in the fractions having different comonomer contents is also referred to as "multimodal."

[0154] For example, if a polymer is produced in a sequential multi-stage process using reactors connected in series and using different conditions in each reactor, the polymer fractions produced in the different reactors will each have their own molecular weight distribution and weight average molecular weight. When a molecular weight distribution curve for such a polymer is recorded, the individual curves from these fractions are superimposed on the molecular weight distribution curve for the total resulting polymer product, typically producing a curve with two or more distinct maxima.

[0155] In any multimodal polymer there may be a lower molecular weight component (LMW) and a higher molecular weight component (HMW). The LMW component has a lower molecular weight than the higher molecular weight component. The difference is preferably at least 5000 g / mol.

[0156] The multimodal polyethylene polymer produced by the present process preferably comprises at least one C4-10 comonomer. The comonomer may be present in the HMW component (or the second component (B) produced in the second polymerization step) or the LMW component (or the first component (A) produced in the first polymerization step), or in both. Hereinafter, the term "LMW / HMW component" will be used, but the embodiments described apply to both the first and second components, respectively.

[0157] It is preferred that the HMW component comprises at least one C4-10 comonomer. The LMW component may then be an ethylene homopolymer or may also comprise at least one C4-10 comonomer. In one embodiment, the multimodal polyethylene polymer comprises a single comonomer. In a preferred embodiment, the multimodal polyethylene polymer comprises at least two, for example exactly two, C4-10 comonomers.

[0158] The total comonomer content in the multimodal polyethylene polymer may be, for example, from 0.2 mol% to 14.0 mol%, preferably from 0.3 mol% to 12 mol%, more preferably from 0.5 mol% to 10.0 mol%, and most preferably from 0.6 mol% to 8.5 mol%.

[0159] 1-Butene may be present in an amount of 0.05 mol% to 6.0 mol%, such as 0.1 mol% to 5 mol%, more preferably 0.15 mol% to 4.5 mol%, and most preferably 0.2 mol% to 4 mol%.

[0160] The C6 to C10 alpha olefins may be present in an amount of 0.2 mol% to 6 mol%, preferably 0.3 mol% to 5.5 mol%, more preferably 0.4 mol% to 4.5 mol%.

[0161] Preferably, the LMW component has a smaller amount (mol %) of comonomer than the HMW component, for example the amount of comonomer (preferably 1-butene) in the LMW component is from 0.05 mol % to 0.9 mol %, more preferably from 0.1 mol % to 0.8 mol %, while the amount of comonomer (preferably 1-hexene) in the HMW component (B) is from 1.0 mol % to 8.0 mol %, more preferably from 1.2 mol % to 7.5 mol %.

[0162] The LMW component of the multimodal polyethylene polymer may have an MFR2 of from 0.5 g / 10 min to 3000 g / 10 min, more preferably from 1.0 g / 10 min to 1000 g / 10 min. In some embodiments, the MFR2 of the LMW component may be from 50 g / 10 min to 3000 g / 10 min, more preferably from 100 g / 10 min to 1000 g / 10 min, for example where cast film is the target.

[0163] The molecular weight (Mw) of the LMW component should preferably be in the range of 20,000 to 180,000, for example 40,000 to 160,000. Its density may be at least 925 kg / m3, for example at least 940 kg / m3. 3 930kg / m 3 Up to 950kg / m 3 , preferably 935kg / m 3 Up to 945kg / m 3 A range of densities is possible.

[0164] The HMW component of the multimodal polyethylene polymer may for example have an MFR2 of less than 1 g / 10 min, such as from 0.2 g / 10 min to 0.9 g / 10 min, preferably from 0.3 g / 10 min to 0.8 g / 10 min, and more preferably from 0.4 g / 10 min to 0.7 g / 10 min. It may have a density of less than 915 kg / m 3 , for example less than 910kg / m 3 , preferably less than 905kg / m 3 The Mw of the higher molecular weight component may be in the range of 70,000 to 1,000,000, preferably 100,000 to 500,000.

[0165] The LMW component may form 30 to 70 wt%, such as 35 to 65 wt%, especially 38 to 62 wt% of the multimodal polyethylene polymer.

[0166] The HMW component may form 30 to 70 wt%, such as 35 to 65 wt%, especially 38 to 62 wt% of the multimodal polyethylene polymer.

[0167] In one embodiment, there is 40 to 45 wt% of the LMW component and 60 to 55 wt% of the HMW component.

[0168] In one embodiment, the polyethylene polymer consists of HMW and LMW as the only polymer components.

[0169] The multimodal polyethylene polymer of the present invention may have an MFR2 of 0.01 g / 10 min to 50 g / 10 min, preferably 0.05 g / 10 min to 25 g / 10 min, especially 0.1 g / 10 min to 10 g / 10 min.

[0170] Example

[0171] catalyst

[0172] SiO2 loading:

[0173] 10 kg of silica (PQ Corporation ES757, calcined at 600° C.) was added from a drum feeder and inerted in the reactor until an O 2 level below 2 ppm was achieved.

[0174] Preparation of MAO / tol / MC:

[0175] At 25 ℃ (oil circulation temperature), a toluene solution (14.1 kg) of 30 wt % MAO was added to another reactor under equilibrium and stirred at 95 rpm, followed by addition of toluene (4.0 kg). After adding toluene, stirring speed was increased from 95 rpm to 200 rpm for 30 minutes. From a metal cylinder, 477 g of metallocene Rac-dimethylsilanediylbis{2-(5-(trimethylsilyl)furan-2-yl)-4,5-dimethylcyclopentadiene-1-yl}zirconium dichloride was added, followed by rinsing with 4 kg of toluene (total toluene amount was 8.0 kg). For MC feed, the reactor stirring speed was changed to 95 rpm and returned to 200 rpm within 3 hours of reaction time. After the reaction time, the MAO / tol / MC solution was transferred to a feed container.

[0176] Preparation of catalyst:

[0177] Reactor temperature is set to 10 ℃ (oil circulation temperature) and stirs with 40rpm after MAO / tol / MC adds.In 205 minutes, add MAO / tol / MC solution (target value 22.5kg, actual value 22.2kg), 60 minutes stirring time (oil circulation temperature is set to 25 ℃) subsequently.After stirring, make " dry mix " stable 12 hours under 25 ℃ (oil circulation temperature), stir with 0rpm.Make reactor return 20 ℃ (repeatedly) and start to stir several rounds with 5rpm, one hour once.

[0178] After stabilization, the catalyst was dried at 60°C (oil circulation temperature) under a 2 kg / h nitrogen flow for 2 hours, followed by drying under vacuum for 13 hours (same nitrogen flow with stirring at 5 rpm). The dried catalyst was sampled and the HC content was measured using thermogravimetric methods with a Sartorius moisture analyzer (model MA45) in a glove box. The target HC level was <2% (actual value was 1.3%).

[0179] Example 1 (Comparative Example)

[0180] The initial size was 25 μm, the span was 1.6 and the apparent density was 1.8 kg / m 3The LLDPE film was prepared using a single-site catalyst. The catalyst was first prepolymerized in a prepolymerization reactor at T = 50 ° C and P = 65 barg. More specifically, 900 kg / h of ethylene, 95 kg of 1-butene / tn ethylene, 0.27 kg of hydrogen / tn propane and 6.50 tn propane / h (diluent) were fed into the prepolymer reactor with an average residence time of 30 minutes. The product was transferred to a volume equal to 80m 3 A split loop reactor is prepared. Ethylene (C2), propane (diluent), 1-butene (C4) and hydrogen (H2) are fed to the reactor, the polymerization conditions being T = 85°C, P = 64 barg and an average residence time equal to 1.0 h. The molar ratios of H2 / C2 and C4 / C2 are 2 mol / kmol and 100 mol / kmol, respectively, and the total production rate in the loop reactor is 25 tn / h (total yield 2.5 kg / gcat). The material is then flashed off in a high-pressure separator, the operating pressure being chosen to be equal to 2 barg and the residence time being estimated to be equal to 5 minutes. Subsequently, the polymer particles are transferred to a reactor with a total volume equal to 350 m 3 A gas phase reactor (including the separation zone) was installed, the reactor being operated at a total pressure of 20 barg, a temperature of 75° C., and having a gas phase composition of 52.5 mol% propane, 10 mol% nitrogen, 32.5 mol% ethylene, 5 mol% C₆, and H₂ / C₂=0.5 mol / kmol. The total residence time in the GPR was 2.8 hours. A superficial gas velocity in the gas phase reactor was selected to be 0.45 m / s.

[0181] A cyclone separator is placed at the outlet of the separation zone (recycle gas duct) to collect entrained particles (estimate particle transport) and to prevent small sized particles from passing through the gas compressor and heat exchanger.

[0182] The total catalyst productivity in the GPR was 3.5 kg / gcat (averaged over three days). The production cutoff was 55%. Solids transfer was measured at 220 kg / h. Operation of the GPR was interrupted and ultimately stopped after three days of operation due to sheeting and caking issues.

[0183] Example 2 (Comparative Example)

[0184] The procedure of Example 1 was repeated, except that the operating pressure in the high-pressure separator was chosen equal to 7 barg and the residence time was estimated to be equal to 5 minutes.

[0185] The total catalyst productivity in the GPR was 3.5 kg / gcat (average over 12 days). The production cutoff was 55%. Solids transfer was measured at 160 kg / h. Operation of the GPR was interrupted and ultimately stopped after 12 days of operation due to sheeting and caking issues.

[0186] Example 3 (embodiment of the present invention)

[0187] The procedure of Example 1 was repeated, except that the operating pressure in the high-pressure separator was chosen equal to 10 barg and the residence time was estimated to be equal to 5 minutes.

[0188] The total catalyst productivity in the GPR was 3.5 kg / gcat. The production cutoff was 55%. The solids transfer was measured at 22 kg / h. The GPR operation was stable during the 20-day operation period.

[0189] Example 4 (embodiment of the present invention)

[0190] The procedure of Example 1 was repeated, except that the operating pressure in the high-pressure separator was chosen equal to 15 barg and the residence time was estimated to be equal to 5 minutes.

[0191] The total catalyst productivity in the GPR was 3.5 kg / gcat. The production cutoff was 55%. The solids transfer was measured at 9 kg / h. The GPR operation was stable during the 20-day operation period.

[0192] Table 1 summarizes the results of the examples.

[0193] Table 1: Summary of results.

[0194]

[0195] Example 5 (Comparative Example)

[0196] The initial size was 25 μm, the span was 1.6 and the apparent density was 1.8 kg / m 3 The LLDPE film was prepared using a single-site catalyst. The catalyst was first prepolymerized in a prepolymerization reactor at T = 50 ° C and P = 65 barg. More specifically, 900 kg / h of ethylene, 95 kg of 1-butene / tn ethylene, 0.27 kg of hydrogen / tn propane and 6.50 tn propane / h (diluent) were fed into the prepolymer reactor with an average residence time of 30 minutes. The product was transferred to a volume equal to 80m 3 A split loop reactor was prepared. Ethylene (C2), propane (diluent), 1-butene (C4) and hydrogen (H2) were fed to the reactor. The polymerization conditions were T = 85°C, P = 64 barg, and an average residence time of 1.0 hour. The molar ratios of H2 / C2 and C4 / C2 were 5 mol / kmol and 40 mol / kmol, respectively, and the total production rate in the loop reactor was 25 tn / h (total production rate was 2.5 kg / gcat). The density of the produced polymer was 955 kg / m 3, MFR2 of 75 and crystallinity of 67 wt%. The material is then flashed off in a high pressure separator, the operating pressure being chosen to be equal to 2 barg and the residence time being estimated to be equal to 5 minutes. Subsequently, the polymer particles are transferred to a container with a total volume equal to 350 m 3 A gas phase reactor (including the separation zone) was installed, the reactor being operated at a total pressure of 20 barg, a temperature of 75° C., and having a gas phase composition of 52.5 mol% propane, 10 mol% nitrogen, 32.5 mol% ethylene, 5 mol% C₆, and H₂ / C₂=0.5 mol / kmol. The total residence time in the GPR was 2.8 hours. A superficial gas velocity in the gas phase reactor was selected to be 0.45 m / s.

[0197] A cyclone separator is placed at the outlet of the separation zone (recycle gas duct) to collect entrained particles (estimate particle transport) and to prevent small sized particles from passing through the gas compressor and heat exchanger.

[0198] The total catalyst productivity in the GPR was 3.5 kg / gcat (averaged over two days). The production cutoff was 55%. Solids transfer was measured at 280 kg / h. Operation of the GPR was interrupted and ultimately stopped after two days of operation due to sheeting and caking issues associated with the solids transfer phenomenon.

[0199] Example 6 (Comparative Example)

[0200] The procedure of Example 5 was repeated, except that the operating pressure in the high-pressure separator was chosen equal to 10 barg and the residence time was estimated to be equal to 5 minutes.

[0201] The total catalyst productivity in the GPR was 3.5 kg / gcat (averaged over 10 days). The production cutoff was 55%. Solids transfer was measured at 190 kg / h. Operation of the GPR was interrupted and ultimately stopped after 10 days of operation due to sheeting and caking issues associated with solids transfer.

[0202] Example 7 (Comparative Example)

[0203] The procedure of Example 5 was repeated, except that the operating pressure in the high-pressure separator was chosen equal to 15 barg and the residence time was estimated to be equal to 5 minutes.

[0204] The total catalyst productivity in the GPR was 3.5 kg / gcat. The production cutoff was 55%. Solids transfer was measured at 80 kg / h. GPR operation was stable during 15 days of operation and was interrupted after 18 days of operation due to sheeting and caking issues associated with solids transfer.

[0205] Example 8 (embodiment of the present invention)

[0206] The procedure of Example 1 was repeated, except that the operating pressure in the high-pressure separator was chosen to be equal to 15 barg and the residence time was estimated to be equal to 12 minutes.

[0207] The total catalyst productivity in the GPR was 3.5 kg / gcat. The production cutoff was 55%. The solids transfer rate was measured at 6 kg / h. The GPR operation was stable during the 20-day operation period.

[0208] Example 9 (Comparative Example)

[0209] The initial size was 25 μm, the span was 1.6 and the apparent density was 1.8 kg / m 3 The LLDPE film was prepared using a single-site catalyst. The catalyst was first prepolymerized in a prepolymerization reactor at T = 50 ° C and P = 65 barg. More specifically, 900 kg / h of ethylene, 95 kg of 1-butene / tn ethylene, 0.27 kg of hydrogen / tn propane and 6.50 tn propane / h (diluent) were fed into the prepolymer reactor with an average residence time of 30 minutes. The product was transferred to a volume equal to 80m 3 A split loop reactor was prepared. Ethylene (C2), propane (diluent), 1-butene (C4) and hydrogen (H2) were fed to the reactor. The polymerization conditions were T = 85°C, P = 64 barg, and an average residence time of 1.0 hour. The molar ratios of H2 / C2 and C4 / C2 were 6 mol / kmol and 20 mol / kmol, respectively, and the total production rate in the loop reactor was 25 tn / h (total production rate was 2.5 kg / gcat). The density of the polymer produced was 965 kg / m 3 , MFR2 of 85 and crystallinity of 75 wt%. The material is then flashed off in a high pressure separator, the operating pressure being chosen to be equal to 2 barg and the residence time being estimated to be equal to 5 minutes. Subsequently, the polymer particles are transferred to a container with a total volume equal to 350 m 3 A gas phase reactor (including the separation zone) was installed, the reactor being operated at a total pressure of 20 barg, a temperature of 75° C., and having a gas phase composition of 52.5 mol% propane, 10 mol% nitrogen, 32.5 mol% ethylene, 5 mol% C₆, and H₂ / C₂=0.5 mol / kmol. The total residence time in the GPR was 2.8 hours. A superficial gas velocity in the gas phase reactor was selected to be 0.45 m / s.

[0210] A cyclone separator is placed at the outlet of the separation zone (recycle gas duct) to collect entrained particles (estimate particle transport) and to prevent small sized particles from passing through the gas compressor and heat exchanger.

[0211] The total catalyst productivity in the GPR was 3.5 kg / gcat (averaged over two days). The production cutoff was 55%. Solids transfer was measured at 330 kg / h. Operation of the GPR was interrupted and ultimately stopped after one day of operation due to sheeting and caking issues associated with the solids transfer phenomenon.

[0212] Example 10 (Comparative Example)

[0213] The procedure of Example 9 was repeated, except that the operating pressure in the high-pressure separator was chosen equal to 10 barg and the residence time was estimated to be equal to 5 minutes.

[0214] The total catalyst productivity in the GPR was 3.5 kg / gcat (average over 10 days). The production cutoff was 55%. Solids transfer was measured at 240 kg / h. Operation of the GPR was interrupted and ultimately stopped after 7 days of operation due to sheeting and caking issues associated with solids transfer.

[0215] Example 11 (Comparative Example)

[0216] The procedure of Example 9 was repeated, except that the operating pressure in the high-pressure separator was chosen equal to 15 barg and the residence time was estimated to be equal to 5 minutes.

[0217] The total catalyst productivity in the GPR was 3.5 kg / gcat. The production cutoff was 55%. Solids transfer was measured at 120 kg / h. Operation of the GPR was stable during 10 days of operation and was interrupted after 11 days due to sheeting and caking issues associated with solids transfer.

[0218] Example 12 (Example of the present invention)

[0219] The procedure of Example 1 was repeated, except that the operating pressure in the high-pressure separator was chosen to be equal to 15 barg and the residence time was estimated to be equal to 12 minutes.

[0220] The total catalyst productivity in the GPR was 3.5 kg / gcat. The production cutoff was 55%. The solids transfer was measured at 7 kg / h. The GPR operation was stable during the 20-day operation period.

[0221] Table 2 summarizes the results of the examples.

[0222] Table 2: Summary of results.

[0223]

[0224]

Claims

1. A method for polymerizing ethylene in a multi-stage polymerization process configuration, the method comprising: (a) polymerizing ethylene in a first polymerization step, optionally in the presence of at least one other alpha-olefin comonomer, in the presence of a metallocene polymerization catalyst to form a first polymer component (A), said first polymer component (A) having a density of from 930 to 950 kg / m 3 within the scope of as well as (b) transferring the first polymer component (A) to a separation unit to remove low molecular weight permeants and obtain separated solid polyethylene particles (A*) of the first polymer component, wherein the separation unit is carried out at a pressure ranging from 1 bar to 25 bar, in a range of 40° C. to 90° C., in one or more flash separators, and the average residence time of the particulate matter in the separation unit is from 2 minutes to 30 minutes; and (c) polymerizing ethylene in a second polymerization step in gas phase, optionally in the presence of at least one other alpha-olefin comonomer, in the presence of the isolated solid polyethylene particles (A*) of step b) to form a second polymer component (B).

2. The process according to claim 1, wherein the separation unit is operated at 2 to 22 bar, in the range of 50°C to 85°C, and the average residence time of the particulate matter in the separation unit is in the range of 3 to 25 minutes.

3. The method according to claim 2, wherein the separation unit operates in the range of 3 bar to 20 bar.

4. The method of claim 2, wherein the separation unit operates in the range of 60°C to 80°C.

5. The method of claim 2, wherein the average residence time of the particulate matter in the separation unit is in the range of 5 minutes to 20 minutes. The method according to claim 1 or 2, wherein the low molecular weight penetrant is one or more selected from α-C2-6 olefins, C2-6 alkanes and chain transfer agents.

7. The method according to claim 6, wherein the low molecular weight permeant is selected from the group consisting of ethylene, 1-butene, 1-hexene, propane, hydrogen, and any mixture thereof.

8. The method according to claim 1 or 2, wherein the metallocene polymerization catalyst is a single-site catalyst.

9. The method of claim 1 or 2, wherein the metallocene polymerization catalyst comprises (i) a transition metal complex, (ii) a cocatalyst and optionally (iii) a support.

10. A method for improving the performance of metallocene polymerization catalysts in multi-stage ethylene polymerization, wherein the first polymer component (A) produced in the first polymerization step is transferred to a separation unit to remove low molecular weight permeants and obtain separated solid polyethylene particles (A*) of the first polymer component, and then the obtained separated solid polyethylene particles (A*) of the first polymer component are transferred to another polymerization step, wherein the separation unit is carried out at 1 bar to 25 bar, in the range of 40°C to 90°C, in one or more flash separators, and the average residence time of the particulate matter in the separation unit is 2 minutes to 30 minutes.

11. The process according to claim 10, wherein the operating pressure of the flash separator is in the range of 2 bar to 22 bar.

12. The process according to claim 11, wherein the operating pressure of the flash separator is in the range of 3 bar to 20 bar.

13. The process according to claim 10 or 11, wherein the operating temperature of the flash separator is in the range of 50°C to 85°C.

14. The process of claim 13, wherein the operating temperature of the flash separator is in the range of 60°C to 80°C.

15. The process of claim 10 or 11, wherein the average residence time of the particulate matter in the flash separator is in the range of 3 minutes to 25 minutes.

16. The process of claim 15, wherein the average residence time of the particulate matter in the flash separator is in the range of 5 minutes to 20 minutes.

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