WO2025219499A1 - Processes for polymerising olefins - Google Patents

Abstract

The present disclosure relates to a process for polymerising olefins in multistage polymerisation process configuration. The process comprises polymerising, in a reactor, first olefin monomer, optionally in the presence of at least one alpha olefin comonomer, in the presence of a polymerisation catalyst to produce a first polymer component (A); polymerising, in a downstream gas phase reactor, second olefin monomer at a first operating pressure in the presence of the first polymer component (A) and optionally in the presence of at least one alpha olefin comonomer to produce a second polymer component (B) at a production rate that provides a weight ratio of first polymer component (A) to second polymer component (B) that is below a target threshold value; and increasing the first operating pressure to a second operating pressure in the downstream gas phase reactor when the weight ratio of first polymer component (A) to second polymer component (B) increases above the target threshold value.

Description

PROCESSES FOR POLYMERISING OLEFINS

FIELD OF THE DISCLOSURE

The present disclosure relates to a process for polymerising olefins in multistage polymerisation process configuration.

BACKGROUND OF THE DISCLOSURE

Multistage polymerisation processes (e.g. Borstar PE, PP and Spheripol PP) employ multistage reactor configurations to provide the multimodal capability for achieving easy-to- process resins with desirable mechanical properties. In such processes, a combination of e.g. slurry loop reactors in series followed by a gas phase reactor may be employed to produce a range of polyolefins.

To be able to achieve a wide range of polyolefin specifications using such multi-stage olefin polymerisation processes, it is desirable to be able to achieve a broad range of production splits between the upstream reactor(s) and downstream gas phase reactor, while maintaining a desirable level of overall production throughput. In general, the product portfolio can be enhanced if the gas phase reactor production split can be increased for a given production throughput.

It can be particularly challenging to produce polyolefins with higher proportions of the downstream gas phaser reactor. For instance, a fast-decaying catalyst may be less active by the time polymerisation occurs in the gas phase reactor, particularly if the reaction mixture is exposed to significant residence times in the upstream e.g., slurry loop reactors. Even with catalysts that exhibit slower decay activity (i.e. , relatively flat catalyst activity profile), it can be a challenge to increase the gas phase reactor split in multi-stage reactor configurations because polymerisation conditions in the gas phase reactor are generally milder than those employed upstream.

WO 2022/200537 describes a process for the preparation of a heterophasic polypropylene resin in a multistage polymerisation process. The process employs a particular metallocene catalyst to produce an ethylene propylene copolymer rubber phase dispersed in a polypropylene matrix. The process comprises polymerising propylene in a first step to produce a matrix phase. Propylene and ethylene are then polymerised in a second step in the presence of the polymer from the first step to obtain the dispersed rubber phase. The second step is carried out at a pressure of at least 26 barg, most preferably 30 to 38 barg. The high pressure is used primarily to influence the nature of the ethylene propylene rubber produced and, in particular, to increase the reactivity of ethylene relative to propylene in the second step. The document does not address the problem of catalyst deactivation, and does not describe operating the second step initially at a lower pressure and subsequently increasing the pressure during the course of production.

DESCRIPTION OF THE DISCLOSURE

According to a first aspect, there is provided a process for polymerising olefins in multistage polymerisation process configuration. The process comprises polymerising, in a reactor, first olefin monomer, optionally in the presence of at least one alpha olefin comonomer, in the presence of a polymerisation catalyst to produce a first polymer component (A); and polymerising, in a downstream gas phase reactor, second olefin monomer at a first operating pressure, optionally in the presence of at least one alpha olefin comonomer, in the presence of the first polymer component (A) to produce a second polymer component (B) at a production rate that provides a weight ratio of first polymer component (A) to second polymer component (B) that is below a target threshold value. When the weight ratio of first polymer component (A) to second polymer component (B) increases above the target threshold value, the first operating pressure is increased to a second operating pressure.

In the process of the present disclosure, a first olefin monomer is polymerised in a reactor (e.g., a slurry loop reactor(s)) in the presence of a polymerisation catalyst to produce a first polymer component (A). The first polymer component (A) and, optionally, the polymerisation catalyst, are transferred to a downstream gas phase reactor, in which a second olefin monomer is polymerised at a first operating pressure in the presence of the first polymer component (A) to produce a second polymer component (B). The second polymer component (B) is produced at a production rate, whereby the weight ratio of the first polymer component (A) to second polymer component (B) is below a target threshold value.

In many instances, it can be challenging to produce polymer component (B) at a production rate sufficient to maintain the weight ratio of first polymer component (A) to second polymer component (B) below the target threshold value, while maintaining overall rates of polymer production at desirable levels. This may be because of deactivation of the polymerisation catalyst during the course of the reaction. Because the gas phase reactor in which the second polymer component (B) is produced is downstream of the reactor in which the first polymer component (A) is produced, the polymerisation catalyst may be less active during the downstream gas phase reaction, particularly if the residence times in preceding reactor(s) are significant. Furthermore, the downstream gas phase reactor may be operated under milder conditions, making it more challenging to maintain sufficiently high rates of production of polymer component (B) to keep the weight ratio below the target threshold. Nevertheless, it is important to maintain the weight ratio at desired levels so that a multimodal polymer with the required characteristics can be produced at desired production rates. The ability to boost the proportion of the second polymer component (B) produced in the gas phase reactor may also allow a broader range of multimodal polyolefin products to be produced in multistage configuration, as polymers with broader ranges of gas phase polymer content can be produced.

It has been found that, when the weight ratio of first polymer component (A) to second polymer component (B) increases above the target threshold value, the weight ratio may be reduced by increasing the first operating pressure. This is unexpected because increases in operating pressure generally result in more densely packed beds, which tend to have a negative impact on gas phase polymerisation. Surprisingly, it has been found that any negatives, perceived or otherwise, can be at least partially ameliorated because higher operating pressures increase the sorption of olefin in the polymer phase. This improves catalyst activity and the rate of polymerisation. Accordingly, it has been found that production rates of polymer (B) can be increased by increasing the operating pressure in the gas phase.

Preferably, the operating pressure of the downstream gas phase reactor may be used to control the weight ratio, such that any increases in the weight ratio between first polymer component (A) and second polymer component (B) can be limited or even at least partially reversed by increasing the operating pressure of the downstream gas phase reactor. Preferably, the second operating pressure is at least 10%, more preferably at least 15%, yet more preferably at least 20% greater than the first operating pressure. The second operating pressure may be 10 to 40% greater than the first operating pressure, for example 20 to 40% greater, 30 to 40% greater, 20 to 30% greater, or 10 to 20% greater.

The second operating pressure may be at least 22 barg and, preferably less than 27 barg. Preferably, the second operating pressure may be 22 to less than 26.5 barg, more preferably 22 to 26 barg, even more preferably 22 to 25.5 barg, for example, 22 to 25 barg. If desired, the superficial gas velocity in the downstream gas phase reactor may be increased when the first operating pressure is increased to the second operating temperature. Increasing the superficial gas velocity may reduce the risk of packing in the downstream gas phase reactor and/or improve fluidisation in the downstream gas phase reactor. On the other hand, higher superficial gas velocities can increase the risk of solids (e.g. catalyst particles and/or product) being loss as entrained solids. To ameliorate the risk of solid loss through entrainment, the process may further comprise recycling at least a portion of any entrained solids from the top of the downstream gas phase reactor to the process.

As discussed above, it may be desirable to produce polymer component (B) at a production rate sufficient to maintain the weight ratio of first polymer component (A) to second polymer component (B) that is below a target threshold value, while maintaining overall rates of polymer production.

Preferably, the predetermined threshold value is 1.2, preferably 1.1, more preferably 1.0, yet more preferably 0.95, yet more preferably 0.90, even more preferably 0.85, still more preferably 0.80, yet even more preferably 0.75, for example 0.70 or 0.60.

The first olefin monomer and second olefin monomer may be the same olefin monomer. For example, the first olefin monomer and the second olefin monomer may be propylene. Alternatively, the first olefin monomer and the second olefin monomer may be ethylene.

In some embodiments, the second olefin monomer may be polymerised in the downstream gas phase reactor in the presence of an induced swelling agent. The second olefin monomer may be polymerised in the downstream gas phase reactor in the presence of an induced swelling agent while the downstream gas phase is operated in condensed mode. The swelling agent may be introduced when the weight ratio of first polymer component (A) to second polymer component (B) increases above the target threshold value. For example, the swelling agent may be introduced when the operating temperature of the downstream gas phase reactor is increased from the first operating pressure to the second operating pressure. In some embodiments, polymerisation in the downstream gas phase occurs in the absence of swelling agent at the first operating pressure, but in the presence of swelling agent at the second operating pressure. Suitable swelling agents include a C4 to C10 alkane or a C4 to C10 alkene. Where a C4 to C10 alkene is used, the C4 to C10 alkene may be different from any alpha olefin comonomer used in the polymerisation in the downstream gas phase reactor. For example, where the second olefin monomer is polymerised with 1 -hexene in the downstream gas phase reactor, the swelling agent is preferably not 1 -hexene. The swelling agent may be selected from the group consisting of butane, pentane, hexane, heptane and octane.

Preferably, the induced swelling agent is present in an amount of 0.1 to 15 weight % of the polymerisation mixture in the downstream gas phase reactor. More preferably, the induced swelling agent may be present in an amount of 1 to 12 weight %, for example, 3 to 11 weight % of the polymerisation mixture in the downstream gas reactor.

The induced swelling agent can enhance the solubility of monomers in the polymer phase. For example, in the case of propylene polymerisation, the induced swelling agent can enhance the solubility of propylene in the polypropylene phase. Catalyst activity increases with increasing concentrations of reactants in the polymer phase. Accordingly, it has been found that production rates of polymer (B) can be increased by the introduction of the induced swelling agent.

It has also been found that, in combination, the addition of induced swelling agent and increase in operating pressure in the downstream gas phase reactor have a synergistic effect on catalyst activity and/or productivity in the downstream gas phase reactor. Preferably, therefore, the operating pressure and the concentration of induced swelling agent in the downstream gas phase reactor can be raised to increase the catalyst activity and/or productivity in the downstream gas phase reactor. This can help to maintain the weight ratio of first polymer component (A) to second polymer component (B) below the target threshold value and/or limit any increase in the weight ratio above the target threshold value.

In some embodiments, the residence time in the downstream gas phase reactor in which the second olefin monomer is polymerised may be longer than the residence time in the reactor in which the first olefin monomer is polymerised by at least 30%, preferably at least 40% or at least 50%.

Multistage Polymerisation As mentioned above, the process of the present disclosure is a process for polymerising olefins in multistage polymerisation process configuration. The process comprises polymerising, in a reactor, first olefin monomer, in the presence of a polymerisation catalyst to produce a first polymer component (A); and polymerising, in a downstream gas phase reactor, second olefin monomer at a first operating pressure in the presence of the first polymer component (A) to produce a second polymer component (B).

To distinguish the reactor(s) in which the first polymer component (A) is produced from the downstream gas phase reactor in which the second polymer component (B) is produced, the reactor(s) in which the first polymer component (A) is produced can be referred to as the upstream reactor(s). Polymerisation in the upstream reactor(s) may be referred to upstream polymerisation, while polymerisation in the downstream gas phase reactor(s) may be referred to as downstream gas phase polymerisation.

The first polymer component (A) may be transferred to the downstream gas phase reactor from the upstream reactor(s), such that the second olefin monomer is produced in the presence of first polymer component (A). Additionally, the polymerisation catalyst may be transferred from the upstream reactor(s) to the downstream gas phase reactor so that the second olefin monomer is produced in the presence of both the first polymer component (A) and the polymerisation catalyst.

Preferably, the same catalyst is used in each step and ideally, it is transferred from prepolymerisation to subsequent polymerisation steps in sequence.

Prepolymerisation

The purpose of the prepolymerisation is to polymerise a small amount of polymer onto the catalyst at a low temperature and/or a low monomer concentration. By prepolymerisation, it may be possible to improve the performance of the catalyst in the subsequent polymerization process stages..

The catalyst components are preferably all introduced to the prepolymerisation step when a prepolymerisation step is present. Preferably, the reaction product of the prepolymerisation step is introduced to the polymerisation in which the first polymer component (A) is produced. In some instances, the solid catalyst component and the cocatalyst can be fed separately. Here, it is possible that only a part of the cocatalyst is introduced into the prepolymerisation stage and the remaining part into subsequent polymerisation stages. Also in such cases it may not be necessary to introduce as much cocatalyst into the prepolymerisation stage to achieve sufficient polymerisation therein.

It is understood within the scope of the invention, that the amount or polymer produced in the prepolymerisation lies within 1 to 7 wt% in respect to the final multimodal (co)polymer. This is not counted as part of the first polymer component (A) produced.

Producing the first polymer component (A)

As discussed above, the process of the present disclosure comprises polymerising, in a reactor, first olefin monomer, optionally in the presence of at least one alpha olefin comonomer, in the presence of a polymerisation catalyst to produce a first polymer component (A).

Preferably, the first olefin monomer is propylene or ethylene.

The first olefin monomer may be polymerised to form a homopolymer as the first polymer component (A). Alternatively, the first olefin monomer may be polymerised to form a copolymer of the first olefin monomer and the at least one alpha olefin comonomer as the first polymer component (A).

For example, propylene may be polymerised in the upstream reactor to produce propylene homopolymer as the first polymer component (A). Alternatively, propylene may be copolymerised with at least one alpha olefin comonomer to produce a propylene copolymer as the first polymer component (A).

In one embodiment, ethylene is polymerised in the upstream reactor to produce ethylene homopolymer as the first polymer component (A). Alternatively, ethylene may be copolymerised with at least one alpha olefin comonomer to produce ethylene copolymer as the first polymer component (A).

Where an alpha olefin comonomer is used, the alpha olefin comonomer may be selected from a C2 to C10 alkene. For the avoidance of doubt, the alpha olefin comonomer is different from the first olefin monomer polymerised in the upstream reactor. Suitable alpha olefin comonomers include ethylene, propylene, 1-butene, 1-pentene, 1-hexene, 1-heptene, 1- hexene, 1-octene, 1-nonene and 1-decene. Preferably, the alpha olefin comonomer is selected from at least one of ethylene, propylene, 1-butene, 1-hexene, and 1-octene. Where the first olefin monomer is propylene, the alpha olefin comonomer may be ethylene, 1- butene, 1-hexene and/or 1-octene. Where the first olefin monomer is ethylene, the alpha olefin comonomer may be propylene, 1-butene and/or 1-hexene. Where an alpha olefin comonomer is used, one or more alpha comonomers may be employed. Preferably, one or two alpha comonomers are employed.

The alpha olefin comonomer content of the first polymer component (A) may be 0 to 10 mol%, preferably 0 to 8 mol%, more preferably 0 to 6 mol%. Where the first polymer component (A) is a polypropylene, the alpha olefin comonomer content may be 0 to 10 mol%, preferably 0 to 8 mol%, more preferably 0 to 6 mol%. Where the first polymer component (A) is a polyethylene, the alpha olefin comonomer content may be 0 to 10 mol%, preferably 0 to 8 mol%, more preferably 0 to 6 mol%.

The upstream polymerisation may take place in any suitable reactor or series of reactors. The upstream polymerisation step may take place in one or more slurry polymerisation reactor(s) or in a gas phase polymerisation reactor, or a combination thereof. In other words, the first polymer component (A) may be formed by slurry polymerisation, preferably in slurry loop reactors.

Preferably the upstream polymerisation step may takes place in one or more slurry polymerisation reactor(s), more preferably in at least two (e.g., two) or at least three (e.g. three) slurry phase reactors. Where, for example, at least three (e.g., three) slurry reactors are used, this does not include a slurry reactor for prepolymerisation.

Slurry ethylene polymerisation can take place in an inert diluent, typically a hydrocarbon diluent such as methane, ethane, propane, n-butane, isobutane, pentanes, hexanes, heptanes, octanes etc., or their mixtures. Preferably the diluent is a low-boiling hydrocarbon having from 1 to 4 carbon atoms or a mixture of such hydrocarbons. An especially preferred diluent is propane, possibly containing minor amount of methane, ethane and/or butane. In bulk propylene polymerization, propylene acts as monomer and diluent at the same time.

Polymer particles of the first polymer component (A) formed by polymerisation, together with the catalyst fragmented and dispersed within the particles, may be suspended in the inert diluent. The slurry is agitated to enable the transfer of reactants from the fluid into the particles.

The olefin content in the fluid phase of the slurry may be from 2 to about 99 % by mole, preferably from about 3 to about 96% by mole and in particular from about 5 to about 90 % by mole. Where the first olefin monomer is propylene, the propylene content in the fluid phase of the slurry may be from 2 to about 99 % by mole, preferably from about 3 to about 96 % by mole and in particular from about 5 to about 90 % by mole. Where the first olefin monomer is ethylene, the ethylene content in the fluid phase of the slurry may be from 2 to about 50 % by mole, preferably from about 3 to about 20 % by mole and in particular from about 5 to about 15 % by mole. The benefit of using higher olefin concentration is that the productivity of the catalyst can be increased.

The temperature in the upstream polymerisation may be from 20 to 115 °C, preferably from 25 to 110 °C and in particular from 30 to 100 °C. The pressure is from 1 to 150 bar, preferably from 10 to 100 bar.

The residence time in the upstream reactor (including any prepolymerisation reactor employed) may be typically from 0.15 h to 3.0 h, preferably from 0.20 h to 2.0 h and in particular from 0.30 to 1.5 h.

It is sometimes advantageous to conduct the slurry polymerisation above the critical temperature and pressure of the fluid mixture. Such operation is described in US-A- 5391654. In such operation, the temperature is typically from 70 to 110 °C, preferably from 80 to 105 °C and the pressure is from 40 to 150 bar, preferably from 50 to 100 bar.

Where the upstream polymerisation is carried out by slurry polymerisation, the slurry polymerisation may be conducted in any known reactor used for slurry polymerisation. Such reactors include a continuous stirred tank reactor and a loop reactor. It is especially preferred to conduct the polymerisation in loop reactor. In such reactors the slurry is circulated with a high velocity along a closed pipe by using a circulation pump. Loop reactors are generally known in the art and examples are given, for instance, in US A-4582816, USA-3405109, US-A-3324093, EP-A-479186, and US-A-5391654. The slurry may be withdrawn from the reactor either continuously or intermittently. A preferred way of intermittent withdrawal is the use of settling legs where slurry is allowed to concentrate before withdrawing a batch of the concentrated slurry from the reactor. The use of settling legs is disclosed, among others, in US-A-3374211 , US-A-3242150 and EP- A-1310295.

Continuous withdrawal is disclosed, among others, in EP-A-891990, EP-A- 1415999, EP-A- 1591460 and WO-A-2007/025640. The continuous withdrawal is advantageously combined with a suitable concentration method, as disclosed in EP-A- 1310295, EP-A-1591460, and EP3178853B1.

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

Downstream gas phase reactor

Once formed in the upstream reactor(s), the first polymer component (A) is transferred to the downstream gas phase reactor.

As mentioned above, polymerisation in the downstream gas phase reactor involves polymerising a second olefin monomer and optionally at least one alpha olefin comonomer in the presence of the first polymer component (A).

As discussed above, the second olefin monomer is preferably the same as the first olefin monomer. For example, the first and second olefin monomers may be propylene. Alternatively, the first and second olefin monomers may be ethylene.

The second olefin monomer may be polymerised in the downstream gas phase reactor to form a homopolymer as the second polymer component (B). Alternatively, the second olefin monomer may be polymerised to form a copolymer of the second olefin monomer and the at least one alpha olefin comonomer as the second polymer component (B).

For example, propylene may be polymerised in the downstream gas phase reactor to produce propylene homopolymer as the second polymer component (B). Alternatively, propylene may be co-polymerised with at least one alpha olefin comonomer to produce a propylene copolymer as the second polymer component (B).

In one embodiment, ethylene is polymerised in the downstream gas phase reactor to produce ethylene homopolymer as the second polymer component (B). Alternatively, ethylene may be co-polymerised with at least one alpha olefin comonomer to produce ethylene copolymer as the second polymer component (B).

The second polymer component (B) may or may not be an ethylene propylene rubber.

Where an alpha olefin comonomer is used, the alpha olefin comonomer may be selected from a C2 to C10 alkene. For the avoidance of doubt, the alpha olefin comonomer is different from the first olefin monomer polymerised in the upstream reactor. Suitable alpha olefin comonomers include ethylene, propylene, 1-butene, 1-pentene, 1-hexene, 1-heptene, 1- hexene, 1-octene, 1-nonene and 1-decene. Preferably, the alpha olefin comonomer is selected from at least one of ethylene, propylene, 1-butene, 1-hexene and 1-octene. Where the first olefin monomer is propylene, the alpha olefin comonomer may be ethylene, 1- butene, 1-hexene and/or 1-octene. Where the first olefin monomer is ethylene, the alpha olefin comonomer may be propylene, 1-butene, 1-hexene, and/or 1-octene. Where an alpha olefin comonomer is used, one or more alpha comonomers may be employed. Preferably, one or two alpha comonomers are employed.

Where an alpha olefin comonomer is used, the alpha olefin comonomer used in the downstream gas phase reactor may be the same or different from any alpha olefin comonomer used in the upstream reactor(s).

The alpha olefin comonomer content of the second polymer component (B) may be 0 to 10 mol%, preferably 0 to 8 mol%, more preferably 0 to 6 mol%. Where the second polymer component (B) is a polypropylene, the alpha olefin comonomer content may be 0 to 10 mol%, preferably 0 to 8 mol%, more preferably 0 to 6 mol%. Where the second polymer component (B) is a polyethylene, the alpha olefin comonomer content may be 0 to 10 mol%, preferably 0 to 8 mol%, more preferably 0 to 6 mol%.

The downstream gas phase polymerisation may take place in one or more gas phase polymerisation reactor(s) to produce the second polymer component (B).

The gas phase polymerisation may be conducted in any known reactor used for gas phase polymerisation. Such reactors include a fluidized bed reactor and a fast fluidized bed reactor or in any combination of these. When a combination of reactors is used then the polymer is transferred from one polymerisation reactor to another. Furthermore, a part or whole of the polymer from a polymerisation stage may be returned into a prior polymerisation stage. The temperature in the gas phase polymerisation may be from 50 to 100 °C, preferably from 65 to 90 °C.

The first operating pressure in the gas phase polymerisation may be from 5 to 30 barg. For example, 5 to 29 barg. Preferably, the first operating pressure may be 10 to 27 barg, yet more preferably 15 to 22 barg. Where the second olefin monomer is propylene, the first operating pressure may be 5 to 30 barg, more preferably 10 to 27 barg, yet more preferably 15 to 22 barg.

The second operating pressure in the gas phase polymerisation reactor may be at least 10%, more preferably at least 15%, yet more preferably at least 20% greater than the first operating pressure. In some embodiments, the second operating pressure is 10 to 60%, preferably 10 to 50%, more preferably 10 to 45%, yet more preferably 10 to 40% greater, even more preferably 10 to 35%, still more preferably 10 to 20% greater than the first operating pressure. In some embodiments, the second operating pressure is 15 to 60%, preferably 15 to 50%, more preferably 15 to 45%, yet more preferably 15 to 40% greater, even more preferably 10 to 35%, still more preferably 10 to 20% greater than the first operating pressure. In some embodiments, the second operating pressure is 20 to 60%, preferably 20 to 50%, more preferably 20 to 45%, yet more preferably 20 to 40% greater, still more preferably 20 to 35% greater than the first operating pressure.

The second operating pressure may be 10 to 40 barg, preferably from 15 to 35 barg. Preferably, the second operating pressure is at least 22 barg. More preferably, the second operating pressure is less than 27 barg. The second operating pressure may be 22 to 27 barg. Preferably, the second operating pressure may be 22 to less than 26.5 barg, more preferably 22 to 26 barg, even more preferably 22 to 25.5 barg, for example, 22 to 25 barg. The residence time in the gas phase polymerisation is from 1.0 h to 4.5 h, preferably from 1.5 h to 4.0 h and in particular from 2.0 to 3.5 h. The residence time in gas phase polymerisation may be longer than the residence time in the upstream reactor(s) by at least 30%, preferably at least 40% or at least 50%.

The polymer production rate in the gas phase reactor may be from 10 tn/h to 65 tn/h, preferably from 12 tn/h to 58 tn/h and in particular from 13 tn/h to 52.0 tn/h, and thus the total polymer withdrawal rate from the gas phase reactor may be from 15 tn/h to 100 tn/h, preferably from 18 tn/h to 90 tn/h and in particular from 20 tn/h to 80.0 tn/h. The gas phase polymerisation is preferably conducted in a gas-solids fluidized bed(s).

An example of a suitable downstream gas phase polymerisation reactor may comprise three zones: a) a bottom zone where fluidization gas is introduced into the reactor; b) a middle zone, which may have a generally cylindrical shape, where olefin monomer(s) present in the fluidization gas are polymerised to form the polymer particles; and c) a top zone, where fluidization gas is withdrawn from the reactor. A fluidization grid (also named distribution plate) may be employed to separate the bottom zone from the middle zone. The top zone may also act as a disengaging or entrainment zone in which, due to its expanding diameter compared to the middle zone, the fluidization gas can expand and disengage from the polyolefin powder.

As polymer particles are formed during polymerisation, a “dense phase” having an increased bulk density will form in the area within the middle zone of the reactor with due to the formation of the polymer particles.

As discussed above, when the operating pressure in the downstream gas phase reactor is increased, the bulk density of the fluidised bed may increase. In some embodiments, to maintain desirable fluidisation characteristics within the fluidised bed, it may be desirable to increase the superficial gas velocity through the downstream gas phase reactor. For example, the superficial gas velocity of the fluidising gas introduced into the reactor may be increased.

In some embodiments, the bulk density of the dense phase during polymerization may be in the range of from 100 to 500 kg/m³, preferably of from 120 to 470 kg/m³, most preferably of from 150 to 450 kg/m³.

It may be desirable to maintain the bulk density within a target range even when the operating pressure of the downstream gas phase reactor is increased to the second operating pressure. To achieve desirable bulk densities and/or adequate fluidisation, it may be desirable to increase the superficial gas velocity through the downstream gas phase reactor. For example, the superficial gas velocity of the stream of fluidization gas introduced into the downstream gas phase reactor (e.g., in the bottom zone) may be in the range of from 0.1 to 1.3 m/s, more preferably of from 0.15 to 1.1 m/s, most preferably of from 0.2 to 1.0 m/s. It may be desirable to increase the superficial gas velocity through the downstream gas phase reactor by at least 5%, at least 10%, for example, by 5 to 30% when the operating pressure is increased from the first operating pressure to the second operating pressure.

While an increase in superficial gas velocity may improve fluidisation and/or help to maintain bulk densities within desired values, this can increase the risk of polyolefin powder being entrained towards the reactor exit, increasing the risk of product loss. To ameliorate the risk of product loss, it may be desirable to recover the entrained solids, for example, using a separator, such as cyclone(s).

For example, a fluidization gas stream may be withdrawn from the reactor exit (e.g., at the highest location). The withdrawn fluidization gas stream may then be cooled and reintroduced to the gas phase reactor (e.g., bottom zone), for example, as at least part of the fluidization gas. In order to remove entrained polyolefin powder from the withdrawn stream, a separator, for example, a cyclone(s) may be installed in the circulation gas line used to withdraw the fluidization gas stream. The cyclone can be used to remove any entrained polymer material from the withdrawn fluidization gas stream used as circulation gas. The polymer stream recovered from the cyclone can be directed to another polymerization stage, or it may be returned into the gas-solids olefin polymerization reactor or it may be withdrawn as the polymer product.

After separation of entrained solids and prior to being reintroduced into the gas phase reactor as or as part of the fluidisation gas, the withdrawn stream may be compressed in a compressor and, optionally, cooled. Additional olefin monomer(s), eventual comonomer(s), hydrogen and inert gas are suitably introduced into the circulation gas line. It is preferred to analyse the composition of the circulation gas, for instance, by using on-line gas chromatography and adjust the addition of the gas components so that their contents are maintained at desired levels.

Controlling the predetermined target weight ratio

In the present disclosure, the second polymer component (B) at a production rate that provides a weight ratio of first polymer component (A) to second polymer component (B) that is below a target threshold value.

It may be desirable to produce polymer component (B) at a production rate sufficient to maintain the weight ratio of first polymer component (A) to second polymer component (B) that is below a target threshold value, while maintaining overall rates of polymer production. Preferably, the predetermined threshold value is 1.2, preferably 1.1 , more preferably 1.0, yet more preferably 0.95, yet more preferably 0.90, even more preferably 0.85, still more preferably 0.80, yet even more preferably 0.75, for example 0.70 or 0.60.

The weight ratio of the first polymer component (A) to second polymer component (B) may be determined using any suitable method. For example, the weight ratio of the first polymer component (A) to the second polymer component (B) may be determined from catalyst productivity values. The catalyst productivity value for the first polymer component (A) may be determined in terms of the weight (e.g. kg) of first polymer component (A) produced per g of catalyst employed in the reactor(s) in which the first polymer component (A) is produced. Where prepolymerisation takes place, the catalyst productivity in the prepolymerisation reactor is not included for determining the weight amount of polymer component (A) produced.

The catalyst productivity value for the second polymer component (B) may be determined in terms of the weight (e.g. kg) of second polymer component (B) produced per g of catalyst employed in the gas phase reactor(s) in which the second polymer component (B) is produced. The ratios of these values can be determined to provide the weight ratio of the first polymer component (A) to second polymer component (B).

The predetermined weight ratio is controlled by:

(i) optionally determining the weight ratio of the first polymer component (A) to the second polymer component (B) exiting the downstream gas phase polymerisation reactor; and

(ii) increasing the operating pressure in the downstream gas phase polymerisation reactor from a first operating pressure to second operating pressure when the weight ratio of the first polymer component (A) to the second polymer component (B) is greater than the target threshold value.

The predetermined weight ratio may additionally be controlled by:

(iii) decreasing the operating pressure in the downstream gas phase polymerisation reactor from the second operating pressure to third operating pressure if the weight ratio of the first polymer component (A) to the second polymer component (B) is more than 30% below, preferably 25% below, more preferably more than 20% below, even more preferably more than 15% below, yet more preferably more than 10% below, still more preferably more than 5% less than the target threshold value. The third operating pressure may be the same or different from the first operating pressure.

In some embodiments, the operating pressure of the downstream gas phase polymerisation reactor may be varied (up or down) to control the weight ratio within ± 30%, preferably ± 25%, more preferably ± 20%, yet more preferably ± 15, still more preferably ± 10% of a specific target. The specific target for operating pressure in the downstream gas phase polymerization reactor may be 20 barg, more preferably 19 barg, most preferably 18 barg.

As discussed above, as well as varying the operating pressure of the downstream gas phase reactor, it may be desirable to perform the gas phase polymerisation in the presence of an induced swelling agent as described in WO 2022/268951. The induced swelling agent may be introduced into the downstream gas phase reactor. The concentration of induced swelling agent may be increased when the weight ratio increases above the target threshold value. The concentration of induced swelling agent may be increased when the operating pressure is increased to the second operating pressure.

The concentration of induced swelling agent may be increased from an initial value when the weight ratio increases above the target threshold value. The initial value may be zero. For example, the induced swelling agent is only introduced when the weight ratio increases above the target threshold value and/or the operating pressure is increased from the first operating pressure to the second operating pressure. Alternatively, the induced swelling agent may be present even when the weight ratio is below the target threshold value. However, the concentration of the induced swelling agent may be increased when the weight ratio rises above the target threshold value.

The term “induced swelling agent” used herein refers to a compound capable of permeating the shell and swelling the core of a polymer particle, in particular due to mass uptake. Thus, the induced swelling agent is capable of solubilizing into the polymer particles produced in the polymerisation process in the presence of the said polymer particles and monomers, in particular under the conditions of the specific process for which the swelling agent is used. The induced swelling agent increases the solubility of reactants, including the second olefin monomer, any alpha olefin comonomer and, optionally, other chain transfer agents (e.g., hydrogen, etc.) present in the polymer phase. This can increase catalytic activity in the gas phase polymerisation reaction, helping to maintain the weight ratio below the target threshold value and/or at least partially reversing any increase in the weight ratio above the target threshold value. As mentioned above, it has been found that, by increasing the operating pressure and using an induced swelling agent in combination, it is possible to achieve a synergistic increase in catalytic activity. For example, the combination may result in an improvement that is greater than the sum improvement achieved by increasing the operating pressure and using an induced swelling agent separately.

The term “induced” as used herein in particular refers to intentional aim to create a swelling effect and that the swelling effect is not merely caused because of a circumstantial presence of a component which is anyhow required for the process. Preferably, the induced swelling agent is used to create as high as possible degree of swelling.

The induced swelling agent may be an inert chemical compound that is part of the reaction medium. The induced swelling agent may be selected from alkanes, preferably C4- -alkanes (such as n-heptane, n-butane, n-pentane and any isomers thereof) and Cs-w-comonomer (such as 1 -hexene). Preferably, the induced swelling agent is butane, pentane, heptane, 1- pentene or 1 -hexene or a mixture thereof, more preferably n- butane, n-pentane, n-heptane, 1-pentene or 1-hexene or a mixture thereof.

Where the second olefin monomer is propylene and 1-hexene and/or 1 -butene is used as the alpha olefin comonomer, the induced swelling agent may be an alkane, preferably a C4-- 10 alkane (such as n-heptane, n-butane, n-pentane and any isomers thereof). Alternatively, the induced swelling agent may be Cs-w-comonomer other than 1-hexene, such as 1 - heptene and 1 -octene.

Where the second olefin monomer is ethylene and 1-hexene and/or 1 -butene is used as the alpha olefin comonomer, the induced swelling agent may be an alkane, preferably a C4-IO alkane (such as n-heptane, n-butane, n-pentane and any isomers thereof). Alternatively, the induced swelling agent may be Cs-w-comonomer other than 1-hexene, such as 1 -heptene and 1 -octene.

The concentration of the induced swelling agent in the downstream gas phase polymerisation can be controlled by the total concentration of oligomers (i.e. , expressed as C₆-Ci4 components) in the gas phase reactor, measured by on-line gas chromatographer.

The total concentration of oligomers, i.e. Cs-14 components, in the second polymerisation step is typically in the range 50 to 1200 ppm, preferably lower than 600 ppm, more preferably lower than 500 ppm, most preferably lower than 400 ppm of the total amount of the reaction mixture.

The induced swelling agent may be introduced to the reactor via an injection line that is placed at the bottom of the gas phase reactor and it is mixed with the recirculation gas stream that in turn is introduced into the gas phase reactor.

The induced swelling agent concentration in the downstream gas phase polymerisation reactor may vary from 0.1 weight % to 15 weight%, more preferably 0.5 weight% to 12 weight%, most preferably 1.0 weight % to 1.0 weight%.

Preferably, the predetermined weight ratio may be controlled by:

(i) optionally determining the weight ratio of the first polymer component (A) to the second polymer component (B) exiting the downstream gas phase polymerisation reactor; and either

(ii a) increasing the operating pressure in the downstream gas phase polymerisation reactor from a first operating pressure to second operating pressure and introducing an induced swelling agent into the downstream gas phase reactor if the weight ratio is greater than the target threshold value; or

(ii b) increasing the operating pressure in the downstream gas phase polymerisation reactor from a first operating pressure to second operating pressure and increasing the concentration of induced swelling agent present in the downstream gas phase reactor if the weight ratio is greater than the target threshold value.

It may additionally be possible to:

(iii) decrease the operating pressure in the downstream gas phase polymerisation reactor from the second operating pressure to third operating pressure if the determined weight ratio is more than 30% less, preferably 25% less, more preferably more than 20% less, even more preferably more than 15% less, yet more preferably more than 10% less, still more preferably more than 5% less than the target threshold value, and/or

(iv) decrease the concentration of induced swelling agent present in the downstream gas phase reactor if the determined weight ratio is more than 30% less, preferably 25% less, more preferably more than 20% less, even more preferably more than 15% less, yet more preferably more than 10% less, still more preferably more than 5% less than the target threshold value.

Polymerization catalyst

The polymerization in the multi-stage olefin polymerization reactors is conducted in the presence of an olefin polymerization catalyst. The catalyst may be any catalyst which is capable of producing the desired olefin polymer. Suitable catalysts are, among others, Ziegler - Natta catalysts based on a transition metal, such as titanium, zirconium and/or vanadium catalysts. Especially Ziegler - Natta catalysts are useful as they can produce olefin polymers within a wide range of molecular weight with a high productivity.

Suitable Ziegler-Natta catalysts preferably contain a magnesium compound, an aluminium compound and a titanium compound supported on a particulate support.

The particulate support can be an inorganic oxide support, such as silica, alumina, titania, silica-alumina and silica-titania. Preferably, the support is silica.

The average particle size of the silica support can be typically from 6 to 100 pm. However, it has turned out that special advantages can be obtained if the support has median particle size from 6 to 90 pm, preferably from 10 to 70 pm.

The magnesium compound is a reaction product of a magnesium dialkyl and an alcohol. The alcohol is a linear or branched aliphatic monoalcohol. Preferably, the alcohol has from 6 to 16 carbon atoms. Branched alcohols are especially preferred, and 2-ethyl-1-hexanol is one example of the preferred alcohols. The magnesium dialkyl may be any compound of magnesium bonding to two alkyl groups, which may be the same or different. Butyl-octyl magnesium is one example of the preferred magnesium dialkyls.

The aluminium compound is chlorine containing aluminium alkyl. Especially preferred compounds are aluminium alkyl dichlorides and aluminium alkyl sesquichlorides.

The titanium compound is a halogen containing titanium compound, preferably chlorine containing titanium compound. Especially preferred titanium compound is titanium tetrachloride.

The catalyst can be prepared by sequentially contacting the carrier with the above mentioned compounds, as described in EP-A-688794 or WO- A- 99/51646 . Alternatively, it can be prepared by first preparing a solution from the components and then contacting the solution with a carrier, as described in WO-A-01/55230 .

Another group of suitable Ziegler-Natta catalysts contains a titanium compound together with a magnesium halide compound acting as a support. Thus, the catalyst contains a titanium compound on a magnesium dihalide, like magnesium dichloride. Such catalysts are disclosed, for instance, in WO-A-2005/118655 and EP-A-810235 .

Still a further type of Ziegler-Natta catalysts are catalysts prepared by a method, wherein an emulsion is formed, wherein the active components form a dispersed, i.e. a discontinuous phase in the emulsion of at least two liquid phases. The dispersed phase, in the form of droplets, is solidified from the emulsion, wherein catalyst in the form of solid particles is formed. The principles of preparation of these types of catalysts are given in WO-A- 2003/106510 of Borealis.

The Ziegler-Natta catalyst is used together with an activator. Suitable activators are metal alkyl compounds and especially aluminium alkyl compounds. These compounds include alkyl aluminium halides, such as ethylaluminium dichloride, diethylaluminium chloride, ethylaluminium sesquichloride, dimethylaluminium chloride and the like. They also include trialkylaluminium compounds, such as trimethylaluminium, triethylaluminium, triisobutylaluminium, trihexylaluminium and tri-n-octylaluminium. Furthermore they include alkylaluminium oxy-compounds, such as methylaluminiumoxane (MAO), hexaisobutylaluminiumoxane (HIBAO) and tetraisobutylaluminiumoxane (TIBAO). Other aluminium alkyl compounds, such as isoprenylaluminium, may also be used. Especially preferred activators are trialkylaluminiums, of which triethylaluminium, trimethylaluminium and tri-isobutylaluminium are particularly used. If needed the activator may also include an external electron donor. Suitable electron donor compounds are disclosed in WO-A- 95/32994 , US-A-4107414 , US-A-4186107 , US-A-4226963 , US-A-4347160 , US-A- 4382019 , US-A-4435550 , US-A-4465782 , US 4472524 , US-A-4473660 , US-A-4522930 , US-A-4530912 , US-A-4532313 , US-A-4560671 and US-A-4657882 . Also electron donors consisting of organosilane compounds, containing Si-OCOR, Si-OR, and/or Si-NR2 bonds, having silicon as the central atom, and R is an alkyl, alkenyl, aryl, arylalkyl or cycloalkyl with 1-20 carbon atoms are known in the art. Such compounds are described in US-A-4472524 , US-A-4522930 , US-A-4560671 , US-A-4581342 , US-A-4657882 , EP-A-45976 , EP-A- 45977 and EP-A- 1538167 . The amount in which the activator is used depends on the specific catalyst and activator.

Typically triethylaluminium is used in such amount that the molar ratio of aluminium to the transition metal, like Al/Ti, is from 1 to 1000, preferably from 3 to 100 and in particular from about 5 to about 30 mol/mol.

Metallocene catalysts may also be used. Metallocene catalysts comprise a transition metal compound which contains a cyclopentadienyl, indenyl or fluorenyl ligand. Preferably the catalyst contains two cyclopentadienyl, indenyl or fluorenyl ligands, which may be bridged by a group preferably containing silicon and/or carbon atom(s). Further, the ligands may have substituents, such as alkyl groups, aryl groups, arylalkyl groups, alkylaryl groups, silyl groups, siloxy groups, alkoxy groups or other heteroatom groups or the like. Suitable metallocene catalysts are known in the art and are disclosed, among others, 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 .

EXAMPLES

Examples 1 to 4

In Examples 1 to 4, propylene was polymerised in the presence of a Ziegler Natta catalyst having an average particle size equal to 50 microns under pre-polymerization conditions at T = 23 °C and P = 10 bar for 6 mins. Subsequently, polymerization under loop conditions (i.e. , T = 80 °C and P = 38 bar) was carried out for 10 hours. Then, the material was flushed out so as to remove a significant amount of the propylene sorbed to the particles of the first polymer component (A) produced. The polymer particles and catalyst were then transferred to a downstream gas phase reactor. Polymerization in the gas-phase was carried out for a residence time of 2.0 hours, at a temperature of 80 °C to produce the second polymer component (B). In the case of Examples 3 and 4, during transition from the slurry to gasphase process, n-heptane was then added to the feed to the gas phase reactor. The operating pressure of the gas phase reactor was 20 barg in the case of Examples 1 and 3, and 25 barg in the case of Examples 2 and 4.

Example 1 (Comparative)

Multistage polymerisation (excluding the prepoly and transition time) as described above carried out for overall time of 3 h (1 h loop polymerization and 2 h gas phase polymerization). In this example, n-heptane was not added to the gas phase reaction and the gas phase operating pressure was 20 barg. Table 1 shows the average values of the catalyst activity and catalyst productivity in the relevant stages of the multi-stage process. The weight ratio of the first polymer component (A) to second polymer component (B) calculated using the catalyst productivity values was 1.16.

Table 1. Average catalyst activity values at different PP polymerization stages (no C7 in GPR, P= 20 barg).

Example 2 (Inventive)

The procedure of Example 1 was repeated with the exception that the operating pressure in the gas phase reactor was 25 barg. Table 2 shows the average values of the catalyst activity and catalyst productivity in the relevant stages of the multi-stage process. The weight ratio of the first polymer component (A) to second polymer component (B) calculated using the catalyst productivity values was 1.02.

By increasing the operating pressure in the gas phase reactor, catalytic activity in the gas phase reactor was increased. It can be seen that the weight ratio of first polymer component (A) to second polymer component (B) was reduced from 1.16 in Example 1 to 1.02 in Example 2.

Table 2. Average catalyst activity values at different PP polymerization stages (no 07 in GPR, P= 25 barg).

Example 3 (Comparative) The procedure of Example 1 was repeated with the exception that that 10.4 % wt of n- heptane was added to the gas phase reaction and the overall pressure in the GPR was equal to 20 barg. Table 3 shows the average values of the catalyst activity and catalyst productivity in of the relevant stages of the multi-stage process. The weight ratio of the first polymer component (A) to second polymer component (B) calculated using the catalyst productivity values was 0.95.

By using n-heptane as an induced swelling agent in the gas phase reactor, catalytic activity in the gas phase reactor was increased. It can be seen that the weight ratio of first polymer component (A) to second polymer component (B) was reduced from 1.16 in Example 1 to 0.95 in Example 3.

Table 3. Average catalyst activity values at different PP polymerization stages (10.4 % wt C7 in GPR, P= 20 barg).

Example 4 (Inventive)

The procedure of Example 1 was repeated with the exception that 10.4 %wt of n-heptane was added to the gas phase reaction and the overall pressure in the GPR was equal to 25 barg. Table 4 shows the average values of the catalyst activity and catalyst productivity in the relevant stages of the multi-stage process. The weight ratio of the first polymer component (A) to second polymer component (B) calculated using the catalyst productivity values was 0.80.

By using n-heptane as an induced swelling agent and increasing the operating pressure in the gas phase reactor to 25 barg, catalytic activity in the gas phase reactor was increased. It can be seen that the weight ratio of first polymer component (A) to second polymer component (B) was reduced from 1.16 in Example 1 to 0.80 in this example.

As summarised in Table 5 below, the use of the induced swelling agent in Example 3 improved the GPR activity over that observed in Example 1 by 1.50 kg/g_cat/h. The 5 barg increase in pressure in Example 2 improved the GPR activity over that observed in Example 1 by 2.50 kg/gcat/h. The additive improvement would have been expected to be 4.00 kg/g_cat/h. However, it can be seen that a synergistic improvement of 5.00 kg/g_cat/h was observed. Table 4. Average catalyst activity values at different PP polymerization stages (10.4 % wt C7 in GPR, P= 25 barg).

Table 5. Results summary.

Example 5 (Comparative)

A single site catalyst, having an initial size of 25 microns, span (i.e. , (d90 - d10)/d50) of 1.6 was used to produce LLDPE film. The catalyst was first prepolymerized in a prepolymerisation reactor at T=50 °C and P = 65 barg. More specifically, 900 kg/h of ethylene, 95 kg of 1-butene per tn ethylene, 0.30 kg hydrogen per tn of propane and 6.50 tn propane/h (diluent) were fed to the prepolymerisation reactor and the mean residence time was 30 mins.

The product was transferred to a split loop reactor configuration having volume equal to 80 m³. Ethylene (C2), propane (diluent), 1-butene (C4) and hydrogen (H2) were fed to the reactors and the polymerisation conditions were T = 85 °C, P = 64 barg and the mean residence time was 1.0 h. The molar ratio of H2/C2 and C4/C2 were 0.4 mol/kmol and 40 mol/kmol, respectively, and the overall production rate in the loop reactor (without accounting for the prepolymerization reactor) was 13.5 tn/h. The overall catalyst productivity was 1.2 kg/gcat. C2 % mol equals to 7.5, The first polymer component (A) produced in the loop reactor had a density of 940 kg/m³, and an MFR of 6.0 g/10min.

Then, the material was flushed out in a high-pressure separator during the transition from the slurry to gas-phase process. The polymerisation process in the gas-phase reactor was carried out for residence time of 2.5 hours to produce the second polymer component (B). The overall pressure was 20 barg, and the temperature was 75 °C. The gas phase composition in the gas phase reactor was 59.6 % mol propane, 15 % mol nitrogen, 25 % mol ethylene, 0.37 % mol 1-hexene and 0.03 % mol hydrogen.

The size of the gas phase reactor was 3.5 m diameter. The fluidized bed height was 18 m and the superficial gas velocity (SGV) was 0.5 m/s. The overall mass flow rate of the recirculation gas was 520 tn/h. The final material had a density of 913 kg/m³ and MFI equal to 1.3 g/10min.

The overall catalyst productivity in the gas phase reactor was 2.5 kg/gcat and the production rate in the gas phase reactor was 14.50 tn/h production and 28.0 tn/h overall throughput. The weight ratio of the first polymer component (A) to the second polymer component (B) calculated based on the polymer production rates in the respective reactors was 0.93.

Example 6 (Inventive)

The procedure of Example 5 was repeated with the exception that 22 bar operating pressure was employed in the gas phase reactor (i.e. , increase of partial pressure of propane by 2 bar or 10%). The catalyst productivity in gas phase reactor was 2.8 kg/gcat. The production in the gas phase reactor was 16.80 tn/h. The overall throughput was increased to 30.3 tn/h. The weight ratio of the first polymer component (A) to the second polymer component (B) calculated based on the polymer production rates in the respective reactors was 0.80. This is a marked reduction from the value of 0.93 achieved in Example 5.

Example 7 (Comparative)

A single site catalyst, having an initial size of 25 microns, span (i.e., (d90 - d10)/d50) of 1.6 was used to produce LLDPE film. The catalyst was first prepolymerized in a prepolymerisation reactor at T=50 °C and P = 65 barg. More specifically, 900 kg/h of ethylene, 95 kg of 1-butene per tn ethylene, 0.30 Kg hydrogen per tn of propane and 6.50 tn propane/h (diluent) were fed to prepolymerisation reactor and the mean residence time was 30 mins.

The product was transferred to a split loop reactor configuration having volume equal to 80 m³. Ethylene (C2), propane (diluent), 1-butene (C4) and hydrogen (H2) were fed to the reactors and the polymerisation conditions were T = 85 °C, P = 64 barg and the mean residence time was 1.0 h. The molar ratio of H2/C2 and C4/C2 were 0.4 mol/kmol and 40 mol/kmol, respectively, and the overall production rate in the loop reactor (without the prepolymerization reactor) was 10.0 tn/h (the overall catalyst productivity was 1.0 kg/gcat, C2 %mol equals to 5. The first polymer component (A) produced in the loop reactors had a density of 940 kg/m³, and an MFR of 6.0 g/10min.

Then, the material flushed out in a high-pressure separator during the transition from the slurry to gas-phase process. The polymerisation process in the gas-phase reactor was continued for a residence time of 2.5 hours to produce the second polymer component (B). The overall pressure was 20 barg, and the temperature was 75 °C. The gas phase composition in the gas phase reactor was 44.6 % mol propane, 30 % mol nitrogen, 25 % mol ethylene, 0.37 % mol 1-hexene and 0.03 % mol hydrogen.

The size of the gas phase reactor was 3.5 m diameter, the fluidized bed height was 18 m and the superficial gas velocity (SGV) was 0.5 m/s. The overall mass flow rate of the recirculation gas was 520 tn/h. The final polymer blend had a density equal to 913 kg/m³ and an MFI of 1.3 g/10min.

The overall catalyst productivity in the gas phase reactor was 2.5 kg/gcat and the production rate in the GPR was 9.50 tn/h production. The overall throughput was 19.5 tn/h. The weight ratio of the first polymer component (A) to the second polymer component (B) calculated based on the polymer production rates in the respective reactors was 1 .05

Example 8 (Inventive)

The procedure of Example 7 was repeated with the exception that 22 bar operating pressure was employed in the gas phase reactor. The catalyst productivity in the gas phase reactor was 2.8 kg/gcat. The production in gas phase reactor was 16.80 tn/h and the overall throughput was increased to 26.8 tn/h. The weight ratio of the first polymer component (A) to the second polymer component (B) calculated based on the polymer production rates in the respective reactors was 0.60, a marked reduction compared to the value of 1.05 achieved in Example 7.

Example 9 (Comparative)

A single site catalyst, having an initial size of 25 microns, span (i.e. , (d90 - d10)/d50) of 1.6 was used to produce LLDPE film. The catalyst was first prepolymerized in a prepolymerisation reactor at T=50 °C and P = 65 barg. More specifically, 900 kg/h of ethylene, 95 kg of 1-butene per tn ethylene, 0.30 Kg hydrogen per tn of propane and 6.50 tn propane/h (diluent) were fed to prepolymerisation reactor and the mean residence time was 30 mins. 1

The product was transferred to a split loop reactor configuration having volume equal to 80 m³. Ethylene (C2), propane (diluent), 1 -butene (C4) and hydrogen (H2) were fed to the reactors and the polymerisation conditions were T = 85 °C, P = 64 barg and the mean residence time was 1.0 h. The molar ratio of H2/C2 and C4/C2 were 0.4 mol/kmol and 40 mol/kmol, respectively, and the overall production rate in the loop reactor (without the prepolymerization reactor) was 10.0 tn/h (the overall catalyst productivity was 1.0 kg/gcat, C2 %mol equals to 5). The first polymer component (A) produced in the loop reactor had a density of 940 kg/m³, and an MFR of 6.0 g/10min.

Then, the material flushed out in a high-pressure separator during the transition from the slurry to gas-phase process (operating conditions, T= 70 °C and P = 20 barg). The polymerisation process in the gas-phase reactor was continued for a residence time of 2.5 hours to produce the second polymer component (B). The pressure set up value has been set to 16 barg, and the temperature was 75 °C. The gas phase composition in the gas phase reactor was 40.0 % mol propane, 24.6 % mol nitrogen, 35 % mol ethylene, 0.35 % mol 1 -hexene and 0.05 % mol hydrogen.

The size of the gas phase reactor was 3.5 m diameter, the fluidized bed height was 18 m and the superficial gas velocity (SGV) was 0.55 m/s. The overall mass flow rate of the recirculation gas was 420 tn/h. The final polymer blend had a density equal to 913 kg/m³ and an MFI of 1.5 g/10min.

The overall catalyst productivity in the gas phase reactor was 1.6 kg/gcat and the production rate in the GPR was 6.50 tn/h production. The overall throughput was 16.5 tn/h. The weight ratio of the first polymer component (A) to the second polymer component (B) calculated based on the polymer production rates in the respective reactors was 1.54. The product design requires a ratio value of the first polymer component (A) to the second polymer component (B) equals to 0.6.

Example 10 (Inventive)

The process as described in Example 9 was operated for 2 days and continued without making any changes in the prepolymerization and loops reactors. Then, the pressure in the gas phase reactor was progressively increased from 16 barg to 20 barg by increasing the pressure set point of the gas compressor in the gas recycling stream, while the operating temperature was equal to 75 °C. The ramp up in operating pressure was realized within 8 hours and the overall mass flow rate of the recirculation gas was 500 tn/h and the superficial gas velocity (SGV) was 0.50 m/s. The gas phase composition in the gas phase reactor was 40.0 % mol propane, 24.6 % mol nitrogen, 35 % mol ethylene, 0.35 % mol 1-hexene and 0.05 % mol hydrogen. The final polymer blend had a density equal to 914 kg/m³ and an MFI of 1.5 g/10min.

The catalyst productivity in the gas phase reactor was 2.5 kg/gcat. The production in gas phase reactor was 12.00 tn/h and the overall throughput was increased to 22.00 tn/h. The weight ratio of the first polymer component (A) to the second polymer component (B) calculated based on the polymer production rates in the respective reactors was 0.83. After the gas phase process has been stabilized (10 hours of operation), the pressure in the gas phase reactor was progressively increased from 20 barg to 22 barg by increasing the pressure set point of the gas compressor in the gas recycling stream, while the operating temperature was equal to 75 °C. The ramp up in operating pressure was realized within 4 hours and the overall mass flow rate of the recirculation gas was 520 tn/h and the superficial gas velocity (SGV) was 0.48 m/s. The gas phase composition in the gas phase reactor was 40.0 % mol propane, 24.6 % mol nitrogen, 35 % mol ethylene, 0.35 % mol 1-hexene and 0.05 % mol hydrogen. The final polymer blend had a density equal to 914 kg/m³ and an MFI of 1.5 g/10min.

The catalyst productivity in the gas phase reactor was 2.8 kg/gcat. The production in gas phase reactor was 16.65 tn/h and the overall throughput was increased to 26.65 tn/h. The weight ratio of the first polymer component (A) to the second polymer component (B) calculated based on the polymer production rates in the respective reactors was 0.6 that represents the desired ratio value to meet the required product design features.

Examples 9 and 10 show that a higher throughput is attainable while also maintaining a high level of control of the final product e.g. in the A:B polymer ratio and/or avoiding known issues with high pressure operation such as increased costs and/or product losses.

Claims

Claims

1. A process for polymerising olefins in multistage polymerisation process configuration, said process comprising: polymerising, in a reactor, first olefin monomer, optionally in the presence of at least one alpha olefin comonomer, in the presence of a polymerisation catalyst to produce a first polymer component (A); and polymerising, in a downstream gas phase reactor, second olefin monomer at a first operating pressure in the presence of the first polymer component (A) and optionally in the presence of at least one alpha olefin comonomer to produce a second polymer component (B) at a production rate that provides a weight ratio of first polymer component (A) to second polymer component (B) that is below a target threshold value; and increasing the first operating pressure to a second operating pressure in the downstream gas phase reactor when the weight ratio of first polymer component (A) to second polymer component (B) increases above the target threshold value.

2. A process as claimed in claim 1 , wherein the second operating pressure is at least 10%, more preferably at least 15%, yet more preferably at least 20% greater than the first operating pressure.

3. A process as claimed in claim 2, wherein the second operating pressure is 10 to 40% greater than the first operating pressure.

4. A process as claimed in any one of the preceding claims, wherein the second operating pressure is at least 22 bar and, preferably less than 27 bar.

5. A process as claimed in any one of the preceding claims, wherein the predetermined threshold value is 1.2, preferably, 1.0, more preferably 0.90, yet more preferably 0.80, even more preferably 0.70, still more preferably 0.60.

6. A process as claimed in any one of the preceding claims, wherein the first olefin monomer and second olefin monomer are the same kind of olefin monomer.

7. A process as claimed in claim 6, wherein the first olefin monomer and the second olefin monomer are both propylene, or wherein the first olefin monomer and the second olefin monomer are both ethylene.

8. A process as claimed in any one of the preceding claims, wherein the second olefin monomer is polymerised in the gas phase reactor in the presence of an induced swelling agent, optionally in condensed mode, and, wherein the induced swelling agent is a C4 to C10 alkene or a C4 to C10 alkane, preferably selected from the group consisting of butane, pentane, hexane, heptane and octane.

9. A process as claimed in claim 8, wherein the induced swelling agent is present in an amount of 0.1 to 15 weight % of the polymerisation mixture in the further reactor.

10. A process as claimed in any one of the preceding claims, wherein the first olefin monomer is polymerised in a slurry polymerisation step.

11. A process as claimed in any one of the preceding claims, wherein the residence time in the gas phase reactor in which the second olefin monomer is polymerised is longer than the residence time in the reactor in which the first olefin monomer is polymerised by at least 50%.

12. A process as claimed in any one of the preceding claims, wherein the superficial gas velocity in the gas phase reactor is increased when the first operating pressure is increased to the second operating pressure, and optionally wherein the recovered portion of any entrained solids from the gas phase reactor are recycled to a polymerisation stage in the multistage polymerisation process.

13. A process as claimed in any one of the preceding claims, wherein the polymer production rate in the gas phase reactor is from 10 tn/h to 65 tn/h, preferably from 12 tn/h to 58 tn/h and in particular from 13 tn/h to 52.0 tn/h, and thus the total polymer withdrawal rate from the gas phase reactor is from 15 tn/h to 100 tn/h, preferably from 18 tn/h to 90 tn/h and in particular from 20 tn/h to 80.0 tn/h.