RU2814920C1 - Plant and method of producing flow-mixed polymer - Google Patents

Abstract

FIELD: processing of plastics.

SUBSTANCE: group of inventions can be used in the manufacture of polymer materials. Plant for producing a flow-mixed polymer comprises first reactor line for production of first polymer, second reactor line for production of second reactor and mixing unit for mixing of first and second polymers. First reactor line includes the first polymerisation reactor and the first purification separator located after the reactor. Second reactor line includes the second reactor and the second purification separator located after the reactor. First and second separators are connected to the mixing unit located after the separators. Ratio of internal volumes of the first and second reactors is in range from 95:5 to 55:45. Invention also discloses a version of the plant and a method of producing a flow-mixed polymer.

EFFECT: group of inventions increases melting and glass transition temperatures of the polymer while maintaining its density and melt flow rate.

15 cl, 2 dwg, 1 tbl, 3 ex

Description

The present invention relates to a plant for producing flow-mixed polymer and a method for producing flow-mixed polymer. In particular, the invention relates to an in-line mixing plant and process for two reactors operating in a parallel configuration.

In traditional polymer production, reactors operate in series. This principle of operation allows the use of different process conditions, as a result of which the properties of polymers obtained in a separate reactor can be modified. Within certain limits, it is possible to modify, for example, the molecular weight distribution of the total material, i.e. material obtained in all reactors.

Two or more reactors can also be operated in parallel. WO 2017/108951 describes an in-line polymer mixing process. Two or more low pressure reactor-separator units operate in a parallel configuration.

US 2011/0172375 A1 describes an apparatus for the continuous polymerization of one or more monomers in solution. The installation contains a primary reactor and a secondary reactor, configured to operate in parallel, where the ratio of the volume of the primary reactor to the volume of the secondary reactor is in the range from 60:40 to 95:5.

The monomer/solvent separation and recycling process is described in US 2009/0259005. Proposed are one or more serial or parallel homogeneous polymerization reactors and a downstream gravimetric separator connected in fluid transfer to one or more reactors.

WO2009/082468 discloses polypropylene ethylene-propylene copolymer blends in an in-line process. This in-line process includes two or more reactor lines arranged in parallel and a high pressure separator.

However, the use of two reactors with the same volume makes the installation unprofitable when mixing two different polymer fractions, especially if the ratio of materials in the mixture is high, for example 90:10. The reason is that in this case it is necessary to significantly reduce the productivity in one reactor.

In addition, some modern classes of polymers have some good properties, but have the disadvantage of slow extrusion before or during extrusion, resulting in long cooling times. This behavior has a negative impact on process economics for several applications. Thus, it would be preferable to have improved polymers targeting the same or similar class, however, with faster ejection and therefore shorter cooling time. In other words, there is a need in the art to improve the ejection and cooling time of such target polymers.

In view of the above, there is a need in the art to provide a plant and method for producing polymers by which the properties of the resulting polymer can be adapted or modified.

Thus, the purpose of the invention is to create a plant and method for producing polymers with improved properties compared to the target polymers.

A further object of the invention is to provide a plant and method for producing a polymer having improved ejection and shorter cooling time compared to target polymers.

A specific object of the invention is to provide a plant and method for producing a polymer having improved ejection and shorter cooling time compared to target polymers, while maintaining the density and MFR ₂ of these target polymers.

In the present invention, it has surprisingly been discovered that the apparatus and method of mixing two different polymers results in a polymer having a significantly higher melting temperature T _m and an improved glass transition temperature T _g while maintaining the density and melt flow rate of the target polymer. The significantly higher melting temperature _Tm and improved glass transition temperature _Tg advantageously result in faster ejection and shorter cooling times.

Therefore, in a first aspect, the present invention relates to a plant for producing a flow-mixed polymer, comprising:

the first reactor line to produce the first polymer,

a second reactor line for producing a second polymer and

a mixing unit for in-line mixing the first polymer with the second polymer to produce an in-line mixed polymer,

wherein the first reactor line includes a first polymerization reactor for producing a first polymer and a first separator located downstream of the first polymerization reactor,

the second reactor line includes a second polymerization reactor for producing a second polymer and a second separator located after the second polymerization reactor,

where both the first separator and the second separator are connected to a mixing unit located after the first separator and the second separator,

where the first polymerization reactor has a first internal volume, and the second polymerization reactor has a second internal volume,

and characterized in that the ratio of the first internal volume and the second internal volume is in the range from 95:5 to 55:45.

In a second aspect, the invention relates to a plant for producing in-line mixed polymer, comprising:

the first polymerization reactor for producing the first polymer,

a second polymerization reactor for producing a second polymer and

a mixing unit for in-line mixing the first polymer with the second polymer to produce an in-line copolymer,

where the first polymerization reactor and the second polymerization reactor are connected to a mixing unit located after the first polymerization reactor and the second polymerization reactor,

where after the first polymerization reactor and before the mixing unit there is a first heater, and/or where after the second polymerization reactor and before the mixing unit there is a second heater,

where the first polymerization reactor has a first internal volume, and the second polymerization reactor has a second internal volume,

and characterized in that the ratio of the first internal volume and the second internal volume is in the range from 95:5 to 55:45.

The present invention has several unexpected advantages. First, the use of this plant configuration and the different internal volumes of the two polymerization reactors makes the process more versatile in terms of producing different classes of polymers in each reactor and mixing them at the molecular level, which is not possible with conventional mixing methods. Therefore, the apparatus and method of the invention can be used to produce different classes of polymers in each reactor and produce an in-line blended polymer with improved properties compared to the target polymer produced in a single reactor.

In general, the method according to the invention allows the molecular weight and/or comonomer distribution to be adapted in different polymers produced in parallel reactors having different internal volumes. Such adaptation is advantageous for producing specific polymers suitable for specific applications. Specific applications require properties such as, for example, seal strength and hot seal, higher thermal stability of the seal under shrink film compression, higher thermal stability of soft composite applications such as dashboard coverings.

Secondly, the plant and method according to the invention can also be used to produce the same grade of polymer in both parallel polymerization reactors having different internal volumes, in case a higher throughput is required or required. This is especially important when using rectors with complex internal components that are typically difficult to scale.

In addition, the plant and method according to the invention allows the mixing ratio between the two polymers to be significantly varied, for example up to a mixing ratio of 95:5 wt.%, to allow precise control of the in-line polymer.

Finally, the installation and method according to the invention make it possible to obtain polymers that maintain the density and MFR ₂ of their target polymers. In particular, they can maintain the weight average molecular weight M _w as well as the comonomer content of their target ethylene-alpha olefin copolymers. In addition, by satisfying the requirements in terms of density, MFR ₂ , comonomer content and M _w , the polymers produced by the apparatus and process of the invention have an improved melting point T _m and an improved glass transition temperature T _g . The latter properties allow for faster ejection and shorter cooling times, which in turn improves process economics for multiple applications.

Typically, an in-line blending process for polymers as described herein involves continuously mixing two, or even more, intermediate polymers, optionally of different natures, to produce a final in-line blended polymer. In-line mixing processes are the opposite of traditional batch mixing processes. In batch mixing processes, the final polymer is created by combining various intermediate polymers from storage tanks in a mixing plant.

Both the first and the second polymerization reactor according to the present invention can be any reactor suitable for polymerization that can be operated in a continuous mode. Such reactors are well known in the art. Suitable examples are, inter alia, autoclave reactors or continuous stirred tank reactors or tubular reactors. The first and/or second polymerization reactor is preferably selected from a tubular reactor, a mixing autoclave, a pressure vessel reactor, or a loop reactor. The first polymerization reactor may be the same as or different from the second polymerization reactor.

In addition, both the first and the second polymerization reactor according to the invention have an internal volume. By "internal volume" is meant the volume or space inside the respective reactor. Within this internal volume or space, a polymerization reaction can take place.

The connecting line, as described herein, is typically a pipe. The pipe may optionally be provided with means for controlling, for example, flow and/or heat through the pipe.

The recycle line described herein connecting the separator and its associated polymerization reactor provides the separated monomer(s) back to the polymerization reactor. The recycle line preferably connects only the separator to its corresponding reactor. In other words, a recycle line between the separator and another parallel reactor is eliminated.

The terms “polymer-depleted vapor phase” and “polymer-rich liquid phase” as used herein indicate that the polymer-depleted phase is a vapor phase and the polymer-rich phase is a liquid phase.

DETAILED DESCRIPTION OF THE INSTALLATION ACCORDING TO THE FIRST ASPECT OF THE INVENTION

An apparatus according to a first aspect of the invention includes a first reactor line for producing a first polymer, the first reactor line including a first polymerization reactor for producing a first polymer and a first separator located downstream of the first polymerization reactor. Preferably, the first polymerization reactor includes a first reactor inlet for introducing a first reactor feed stream and a first reactor outlet for withdrawing a first reactor effluent stream containing the first polymer. The first outlet of the reactor is preferably fluid-transferably connected via a connecting line to the inlet of the first separator.

Preferably, the first separator includes an upper outlet for recovering a first polymer-depleted vapor stream and a lower outlet for recovering a first polymer-rich liquid stream.

Preferably, the lower outlet of the first separator is connected via a connecting line to the mixing unit.

Preferably, the connecting line connects the upper outlet of the first separator to either additional processing units downstream of the first separator or the first polymerization reactor, more preferably the connecting line connects the upper outlet of the first separator to the first polymerization reactor. In the latter case, the connecting line may also be referred to as a first recycle line, since the first polymer-depleted vapor stream of the first separator is recycled back to the first polymerization reactor. Thus, the first polymer-depleted vapor stream can be recycled back from the first separator to the first polymerization reactor.

Similar to the first reactor line, the second reactor line according to the first aspect of the invention produces a second polymer and includes a second polymerization reactor for producing the second polymer and a second separator downstream of the second polymerization reactor. Preferably, the second polymerization reactor includes a second reactor inlet for introducing a second feed stream into the reactor and a second reactor outlet for extracting a second reactor effluent stream containing the second polymer. The second outlet of the reactor is preferably fluidly connected through a connecting line to the inlet of the second separator.

Preferably, the second separator includes an upper outlet for recovering a second polymer-depleted vapor stream and/or a lower outlet for recovering a second polymer-rich liquid stream.

Preferably, the lower outlet of the second separator is connected via a connecting line to the mixing unit.

Preferably, the connecting line connects the upper outlet of the second separator to either additional processing units downstream of the second separator or the second polymerization reactor, more preferably the connecting line connects the upper outlet of the second separator to the second polymerization reactor. In the latter case, the connecting line may also be referred to as a second recycle line because the second polymer-depleted vapor stream of the second separator is recycled back to the second polymerization reactor. Thus, the second polymer-depleted vapor stream can be recycled back from the second separator to the second polymerization reactor.

In the first aspect of the invention, preferably the first heater is located after the first polymerization reactor and before the first separator, and/or the second heater is located after the second polymerization reactor and before the second separator, more preferably the first heater is located after the first polymerization reactor and before the first separator, and the second heater is located after the second polymerization reactor and before the second separator.

If present, the first heater heats the first reactor effluent to provide a heated first reactor effluent. The inlet of the first heater is fluidly connected through a connecting line to the outlet of the first polymerization reactor. The outlet of the first heater is connected with the possibility of transferring fluid through a connecting line with the inlet of the first separator.

If present, the second heater heats the second reactor effluent to provide a heated second reactor effluent. The inlet of the second heater is connected to transfer fluid through a connecting line to the outlet of the second polymerization reactor. The outlet of the second heater is connected with the possibility of transferring fluid through a connecting line with the inlet of the second separator.

In one embodiment, the mixing unit is coupled via a connecting line to a first separator, wherein the connecting line carries a first polymer-rich liquid stream from the first separator to the mixing unit, and via a connecting line to a second separator, wherein the connecting line carries a second polymer-enriched liquid stream from the second separator. into the mixing block.

In another embodiment, a connecting line carrying a first polymer-enriched liquid stream from the first separator is combined with a connecting line carrying a second polymer-enriched liquid stream from the second separator at a connecting assembly to produce a combined polymer-enriched liquid stream. The combined polymer-enriched liquid stream contains a first polymer and a second polymer. The connecting unit is located before the mixing block. The mixing unit is connected via a connecting line to the connecting assembly, wherein the connecting line carries the combined polymer-enriched liquid stream to the mixing unit.

In a mixing unit, the first polymer of the first polymer-enriched liquid stream and the second polymer of the second polymer-enriched liquid stream, or the first and second polymers of the combined polymer-enriched liquid stream, are in-line mixed to produce an in-line blended polymer.

Preferably, the mixing unit further includes an outlet for extracting the in-line polymer stream containing the in-line polymer. Preferably the outlet is a lower outlet.

DETAILED DESCRIPTION OF THE INSTALLATION ACCORDING TO THE SECOND ASPECT OF THE INVENTION

The apparatus according to the second aspect of the invention comprises a first polymerization reactor for producing a first polymer and a second polymerization reactor for producing a second polymer. After the first polymerization reactor and before the mixing unit, a first heater is located, and after the second polymerization reactor and before the mixing unit, a second heater is located.

Preferably, the first polymerization reactor includes a first reactor inlet for introducing a first reactor feed stream and a first reactor outlet for extracting a first reactor effluent stream containing the first polymer. Preferably, the first outlet of the reactor is fluidly connected through a connecting line to the inlet of the first heater.

The first heater heats the first reactor effluent to produce a heated first reactor effluent. The first heater preferably includes an outlet for extracting the heated first reactor effluent.

Preferably, the second polymerization reactor includes a second reactor inlet for introducing a second feed stream into the reactor and a second reactor outlet for extracting a second reactor effluent stream containing the second polymer.

Preferably, the second outlet of the reactor is fluidly connected through a connecting line to the inlet of the second heater.

The second heater heats the second reactor effluent to produce a heated second reactor effluent. The second heater preferably includes an outlet for extracting the heated second reactor effluent.

In one embodiment, the outlet of the first heater is connected via a connecting line to the inlet of the mixing unit, and the outlet of the second heater is connected via a connecting line to the inlet of the mixing unit.

In another embodiment, the outlet of the first heater is connected via a connecting line to the connecting node, and the outlet of the second heater is connected via a connecting line to the connecting node. At the connection assembly, the heated first reactor effluent and the heated second reactor effluent form a heated combined reactor effluent. The heated combined reactor effluent contains both the first polymer and the second polymer. The connecting unit is located before the mixing block. The mixing unit is connected via a connecting line to the connecting unit, wherein the connecting line carries the heated combined reactor effluent to the mixing unit.

In the mixing unit, the first polymer of the heated first reactor effluent and the second polymer of the heated second reactor effluent, or the first polymer and the second polymer of the heated combined reactor effluent, are in-line mixed to produce an in-line mixed polymer.

Preferably, the mixing unit further includes an outlet for extracting the in-line polymer stream containing the in-line polymer. Preferably the outlet is a lower outlet.

The following discloses preferred embodiments of installation in accordance with both the first and second aspects of the invention, as applicable.

Preferably, the first heater and/or the second heater is a heat exchanger, more preferably the first heater and the second heater is a heat exchanger.

Preferably, the first separator and/or the second separator are low pressure separators, more preferably the first separator and the second separator are low pressure separators. A low pressure separator refers to a unit for separating volatile components from a relatively dilute polymer solution. Volatile components are typically present in amounts ranging from about 10 to about 90 weight percent of the solution. In a low-pressure separator, a liquid phase containing a polymer dissolved in it and a vapor phase coexist. Preferably the low pressure separator operates at a pressure of up to 20 bar, for example from 1 to 15 bar (absolute pressure) and preferably from 2 to 12 bar (absolute pressure). Low pressure separators have been well known for several decades. The formation of two phases results in the formation of a polymer-rich liquid phase and a polymer-depleted vapor phase. The polymer-lean vapor phase, which contains unreacted monomer, solvent and traces of polymer, and the heavy comonomer, if present, has a density of 100 kg/m³ or less.

A low-pressure separator is the opposite of lower critical solution temperature (LCST) separation, where both the polymer-rich phase and the polymer-depleted phase are either essentially liquids or supercritical fluids.

Preferably, the ratio of the first internal volume to the second internal volume is from 85:15 to 60:40, more preferably from 80:20 to 65:35, and most preferably from 75:25 to 70:30.

Preferably, the mixing unit contains a mixer or separator, more preferably contains one or more mixers, for example two, three or four mixers, or one or more separators, for example two, three or four separators. Preferably the separator is a flash separator or a low pressure separator. The mixer is preferably a static mixer.

DETAILED DESCRIPTION OF THE METHOD ACCORDING TO THE INVENTION

In a third aspect of the invention, a method for producing a flow-mixed polymer is provided. The method is preferably carried out in a plant according to the invention, as described in all embodiments herein. However, the method according to the invention can also be carried out in any other suitable installation.

All preferred embodiments of the plant according to the first aspect and according to the second aspect of the invention are also preferred embodiments of the process for producing a flow-mixed polymer.

Thus, the invention provides a method for producing a flow-blended polymer, the method preferably being carried out in a plant in accordance with the first or second aspect as described herein, the method comprising the steps of:

a1) introducing a first feed stream containing a first monomer into a first polymerization reactor,

a2) polymerizing the first monomer in the presence of a first catalyst in a first polymerization reactor to produce a first polymer,

a3) removing the first reactor effluent containing the first polymer from the first polymerization reactor,

b1) introducing a second feed stream containing a second monomer into a second polymerization reactor,

b2) polymerizing a second monomer in the presence of a second catalyst in a second polymerization reactor to produce a second polymer,

b3) removing the second reactor effluent containing the second polymer from the second polymerization reactor,

c1) mixing the first polymer and the second polymer in a mixing unit to obtain a flow-mixed polymer.

In the process of the invention, a first monomer and a second monomer, a first catalyst and a second catalyst, optionally one or more comonomers, optionally one or more chain transfer agents, and optionally one or more solvents are used to initiate liquid polymerization. Each of the first polymer and/or the second polymer is produced by a high temperature solution polymerization process, preferably at temperatures above 100°C, as described in more detail below.

It goes without saying that the reaction conditions used in different reactors do not have to be the same, but may differ from each other. This allows for the production of different polymers that are ultimately mixed together in-line. Since the reaction conditions can be adjusted completely independently for additional reactor(s), the options for changing the microstructure of the polymers to be mixed are very wide. For example, catalyst, pressure, temperature, monomer feedstock, comonomer/monomer ratio, chain transfer agent feedstock, etc. may be different.

Preferably, the first monomer and/or the second monomer is/is an alpha-olefin monomer. The first monomer and/or the second monomer has two or more carbon atoms, preferably from 2 to 10 carbon atoms. Suitable alpha-olefin monomers are selected from the group consisting of ethylene, propylene, 1-butene, 1-hexene, 1-octene, 1-decene and styrene. More preferably, the first monomer and/or the second monomer is selected from the group consisting of ethylene, propylene and 1-butene. More preferably, the first monomer and/or the second monomer is ethylene, and most preferably the first monomer and the second monomer are ethylene.

Preferably, the first feed stream further comprises a comonomer, preferably octene, and/or the second feed stream further contains a comonomer, preferably octene. More preferably, the first feed stream further comprises a comonomer, preferably octene, and the second feed stream further contains a comonomer, preferably octene.

Preferably, a cocatalyst and/or an activator and/or a catalyst support and/or an external donor is introduced into the first polymerization reactor and/or the second polymerization reactor.

Preferably, the first catalyst contains a metallocene complex, and/or the second catalyst contains a metallocene complex. More preferably, the first catalyst and/or the second catalyst contains or consists of:

(i) at least one metallocene complex,

(ii) an alumoxane co-catalyst and/or a boron-containing co-catalyst, and

(iii) an optional alkylaluminum compound Al(R ⁷ ) ₃ , where R ⁷ represents a linear or branched C ₂ -C ₈ alkyl group.

Preferably at least one metallocene complex (i) contains or consists of a metallocene complex of formula (I)

where M represents Hf,

X represents a sigma ligand,

R are the same or different from each other and may be saturated linear or branched C ₁ -C ₁₀ alkyl, C ₆ -C ₁₀ aryl, C ₄ -C ₁₀ heteroaryl, C ₆ -C ₂₀ alkylaryl or C ₆ -C ₂₀ arylalkyl groups , which may optionally contain up to 2 heteroatoms or silicon atoms,

R ¹ represents a C ₆ -C ₁₀ aryl or C ₆ -C ₂₀ alkylaryl group, optionally containing up to 2 heteroatoms or silicon atoms, or a C ₄ -C ₁₀ heteroaryl group,

R ² represents a C ₄ -C ₂₀ cycloalkyl group, optionally bearing alkyl substituents in the beta positions, formula (II)

where R' may be the same or different from each other and may be hydrogen or is defined as R, and n is from 1 to 17,

and/or

metallocene complex (i) contains or consists of a metallocene complex of formula (III)

where M represents Hf,

X represents a sigma ligand,

R are the same or different from each other and may be saturated linear or branched C ₁ -C ₁₀ alkyl, C ₅ -C ₁₀ aryl, C ₆ -C ₂₀ alkylaryl or C ₆ -C ₂₀ arylalkyl groups, which may optionally contain up to 2 heteroatoms or silicon atoms,

R ¹ is C ₆ -C ₂₀ -aryl, which may be unsubstituted or substituted with one or up to 5 linear or branched C ₁ -C ₁₀ alkyl groups,

R ² represents an unsaturated linear or cyclic C ₃ -C ₂₀ alkyl group or a branched CR ³ R ⁴ R ⁵ group, where R ³ represents hydrogen or a C ₁ -C ₂₀ alkyl group, and R ⁴ and R ⁵ are the same or different and may represent a C ₁ -C ₂₀ alkyl group.

Preferably, the at least one metallocene complex of formula (I) is a metallocene complex of formula (Ia)

Preferably, the at least one metallocene complex of formula (III) is a metallocene complex of formula (IIIa)

((Phenyl)(3-buten-1-yl)methylene(cyclopentadienyl)(2,7-di-tert-butylfluoren-9-yl)hafnium dimethyl).

Most preferably, (phenyl)(cyclohexyl)methylene(cyclopentadienyl)(2,7-di-tert-butylfluoren-9-yl)hafnium dimethyl) is used as the metallocene complex (i).

The above metallocene complexes of formulas (I) and (III) and their preparation are described in more detail in WO2018108917 and WO2018108918.

As cocatalyst (ii), alumoxane or a cocatalyst containing boron, or mixtures thereof, can be used.

The aluminoxane cocatalyst may be a cocatalyst of formula (IV)

where n is from 6 to 20 and R has the meaning given below.

Alumoxanes are formed by partial hydrolysis of organoaluminum compounds, for example, compounds of the formula AlR ₃ , AlR ₂ Y and Al ₂ R ₃ Y ₃ , where R can be, for example, C ₁ -C ₁₀ -alkyl, preferably C ₁ -C ₅ -alkyl , or C ₃ -C ₁₀ -cycloalkyl, C ₇ -C ₁₂ -arylalkyl or -alkylaryl, and/or phenyl, or naphthyl, and where Y may be hydrogen, halogen, preferably chlorine or bromine, or C ₁ -C ₁₀ -alkoxy, preferably methoxy or ethoxy. The resulting oxygen-containing alumoxanes are generally not pure compounds, but are mixtures of oligomers of formula (IV).

A preferred alumoxane is methylaluminoxane (MAO).

Since the alumoxanes used in the present invention as cocatalysts are not pure compounds due to the method of their preparation, the molarity of the alumoxane solutions is hereinafter based on their aluminum content.

Since the alumoxanes used in the present invention as cocatalysts are not pure compounds due to the method of their preparation, the molarity of the alumoxane solutions is hereinafter based on their aluminum content.

The molar ratio of Al in the alumoxane and the metallocene transition metal may range from 1:1 to 2000:1 mol/mol, preferably from 10:1 to 1000:1, more preferably from 50:1 to 500:1 mol/mol.

Suitable amounts of cocatalyst will be well known to one skilled in the art.

In one embodiment of the present invention, the cocatalyst used is alumoxane (ii), preferably methylaluminoxane, and an alkylaluminum compound of the formula Al(R⁷)₃, Where R⁷ represents a linear or branched C₂-C₈-alkyl group (iii).

In this case, the cocatalyst is preferably the reaction product of (ii) an alumoxane, preferably methylaluminoxane, with (iii) an alkyl aluminum compound such as tri-iso-butyl aluminum, tri-iso-hexyl aluminum, tri-n-octyl aluminum, tri-iso-octyl-aluminum and the like. The ratio between methylaluminoxane and alkylaluminum compound can be from 10:1 to 1:10, preferably from 5:1 to 1:5, most preferably from 3:1 to 1:3 moles of Al in methylaluminoxane relative to moles of aluminum in the alkylaluminum compound. The reaction between methylaluminoxane and an alkylaluminum compound is carried out by mixing the two components in a suitable solvent, which may be aromatic or aliphatic, at a temperature of minus 50 °C to plus 80 °C, preferably 10 to 50 °C, more preferably 20 to 40 °C WITH.

Boron-based cocatalysts of interest include boron compounds containing borate 3+ ion, that is, borate compounds. These compounds typically contain an anion of the formula:

(Z) ₄ B ^- (V),

where Z represents an optionally substituted phenyl derivative, said substituent being a halogenC _1-6 alkyl or halogen group. Preferred options are fluoro or trifluoromethyl. Most preferably, the phenyl group is perfluorinated.

Such ionic cocatalysts preferably contain a non-coordinating anion such as tetrakis(pentafluorophenyl)borate.

Suitable counterions are protonated amine or aniline derivatives or phosphonium ions. They may have the general formula (VI) or (VII):

NQ₄ ⁺ (VI) or PQ₄ ⁺ (VII),

wherein Q is independently H, C _1-6 alkyl, C _3-8 cycloalkyl, phenyl C _1-6 alkylene, or optionally substituted Ph. Optional substituents may be C _1-6 alkyl, halogen or nitro. One or more such substituents may be present. Thus, preferred substituted Ph groups include para-substituted phenyl, preferably tolyl or dimethylphenyl.

Preferably, if at least one group Q is H, preferred compounds are thus those of the formula:

NHQ ₃₊ ⁽ VIII) or ^PHQ ₃₊ (IX).

Preferred phenylC _1-6 alkyl groups include benzyl.

Thus, suitable counterions include: methylammonium, anilinium, dimethylammonium, diethylammonium, N-methylanilinium, diphenylammonium, N,N-dimethylanilinium, trimethylammonium, triethylammonium, tri-n-butylammonium, methyldiphenylammonium, p-bromo-N,N-dimethylanilinium or p -nitro-N,N-dimethylanilinium, especially dimethylammonium or N,N-dimethylanilinium. Using pyridinium as the ion is another option.

Phosphonium ions of interest include triphenylphosphonium, triethylphosphonium, diphenylphosphonium, tri(methylphenyl)phosphonium, and tri(dimethylphenyl)phosphonium.

A more preferred counterion is trityl (CPh ₃ ⁺ ) or analogs thereof in which the Ph group is functionalized to carry one or more alkyl groups. Therefore, particularly preferred borates used in the invention contain the tetrakis(pentafluorophenyl)borate ion.

Preferred ionic compounds that can be used in accordance with the present invention include:

tributylammonium tetra(pentafluorophenyl)borate, tributylammonium tetra(trifluoromethylphenyl)borate,

tributylammonium tetra-(4-fluorophenyl)borate,

N,N-dimethylcyclohexylammonium tetrakis-(pentafluorophenyl) borate,

N,N-dimethylbenzylammonium tetrakis(pentafluorophenyl) borate,

N,N-dimethylanilinium tetrakis(pentafluorophenyl) borate,

N,N-di(propyl)ammonium tetrakis(pentafluorophenyl)borate, di(cyclohexyl)ammonium tetrakis(pentafluorophenyl)borate,

triphenylcarbenium tetrakis(pentafluorophenyl) borate, ferrocenium tetrakis(pentafluorophenyl) borate.

Preference is given to triphenylcarbenium tetrakis(pentafluorophenyl)borate, N,N-dimethylcyclohexylammonium tetrakis(pentafluorophenyl)borate, N,N-dimethylbenzylammonium tetrakis(pentafluorophenyl)borate or N,N-dimethylanilinium tetrakis(pentafluorophenyl)borate.

More preferred borates are triphenylcarbenium tetrakis(pentafluorophenyl)borate and N,N-dimethylanilinium tetrakis(pentafluorophenyl)borate.

N,N-dimethylanilinium tetrakis(pentafluorophenyl)borate is most preferred.

In addition, it is possible to add an aluminum alkyl compound. Suitable alkylaluminum compounds are those of formula (VIII) AlR ₃ where R is a straight or branched C ₂ -C ₈ alkyl group.

Preferred alkylaluminum compounds are triethylaluminum, tri-isobutylaluminum, tri-isohexylaluminum, tri-n-octylaluminum and tri-isooctylaluminum.

Suitable amounts of cocatalyst will be well known to one skilled in the art.

The molar ratio of boron to metallocene metal ion may be in the range of 0.5:1 to 10:1 mol/mol, preferably 1:1 to 10:1, especially 1:1 to 5:1 mol/mol.

Even more preferably, the molar ratio of boron to metallocene metal ion is from 1:1 to less than 2:1 mol/mol, for example from 1:1 to 1.8:1 or from 1:1 to 1.5:1.

According to the present invention, it is preferable to use a cocatalyst containing boron as described above.

Preferably, the first catalyst is the same as the second catalyst, or the first catalyst is different from the second catalyst, more preferably, the first catalyst is the same as the second catalyst.

The reaction temperature in the first polymerization reactor and the second polymerization reactor is preferably such that the first and second polymers, respectively, formed during the polymerization reaction are completely dissolved in the reaction mixture containing the solvent, comonomer(s), chain transfer agent and polymer. The reaction temperature is sufficiently higher than the melting point of the polymer.

Preferably, the polymerization step a2) is carried out at a first reaction temperature, and the polymerization step b2) is carried out at a second reaction temperature, wherein the first reaction temperature is the same as or different from the second reaction temperature.

Thus, when the polymer is an ethylene homo- or copolymer, the first reaction temperature in step a2) and the second reaction temperature in step b2) are preferably 120 to 240 °C, for example 140 to 220 °C, most preferably 150 up to 200 °C, depending on the content of comonomer units in the polymer. When the polymer is a propylene homo- or copolymer, the first reaction temperature in step a2) and the second reaction temperature in step b2) are preferably 120 to 250 °C, for example 140 to 235 °C, most preferably 150 to 225 °C C, depending on the content of comonomer units in the polymer.

The reactor pressure in the first and second polymerization reactors depends on the temperature, on the one hand, and the type and amount of comonomer, on the other hand.

Preferably, the polymerization step a2) is carried out at a first reactor pressure, and the polymerization step b2) is carried out at a second reactor pressure, wherein the first reactor pressure is the same as or different from the second reactor pressure.

The first reactor pressure and the second reactor pressure are preferably 50 to 300 bar, more preferably 60 to 250 bar, and most preferably 70 to 200 bar.

Preferably, the first polymer has a first density and the second polymer has a second density, the first density being different from the second density.

The solvent is preferably introduced into the first polymerization reactor and the second polymerization reactor. The solvent may be any suitable straight-chain or branched alkyl containing from 3 to 20 carbon atoms, a cyclic alkyl optionally containing alkyl substituents containing from 5 to 20 carbon atoms, or an aryl optionally containing alkyl substituents containing from 6 to 20 carbon atoms , or a mixture of two or more of the above compounds. The solvent must be inert towards the catalyst and monomers and comonomers. In addition, it must be stable under polymerization conditions. Additionally, it must be capable of dissolving the monomer, optional comonomers, optional chain transfer agent and polymer under polymerization conditions.

Preferably, the first feed stream further contains a solvent and/or a chain transfer agent, and/or the second feed stream further contains a solvent and/or a chain transfer agent. Alternatively, the solvent and/or chain transfer agent may be introduced at a different stream from the first feed stream into the first polymerization reactor, and/or the solvent and/or chain transfer agent may be introduced at a stream different from the second feed stream into the second polymerization reactor.

The solvent and/or chain transfer agent introduced into the first polymerization reactor may be different from the solvent and/or chain transfer agent introduced into the second polymerization reactor. However, preferably, the solvent and/or chain transfer agent introduced into the first polymerization reactor is the same as the solvent and/or chain transfer agent introduced into the first polymerization reactor.

A chain transfer agent can be used in one or both polymerization reactors to control the molecular weight of the polymer, as is known in the art. A suitable chain transfer agent is, for example, hydrogen. A chain transfer agent is a substance capable of reacting with a growing polymer chain, thereby transferring activity to the growing polymer chain. By maintaining different concentrations of the chain transfer agent in the two reactors, a polymer mixture with a broad molecular weight distribution can be obtained.

The method according to the invention preferably operates continuously.

In step a3), the first reactor effluent containing the first polymer is removed from the first polymerization reactor. The first reactor effluent may further comprise an unreacted first monomer, an optional unreacted comonomer, an optional chain transfer agent, and an optional solvent.

The first reactor effluent preferably contains from 10 to 35 wt% of the first polymer, more preferably from 12 to 30 wt% of the first polymer, and most preferably from 15 to 25 wt% of the first polymer, based on the total weight content of the first reactor effluent.

In step b3), a second reactor effluent containing a second polymer is removed from the second polymerization reactor. The second reactor effluent may further comprise an unreacted second monomer, an optional unreacted comonomer, an optional chain transfer agent, and an optional solvent.

The second reactor effluent preferably contains from 10 to 35 wt% of the second polymer, more preferably from 12 to 30 wt% of the second polymer, and most preferably from 15 to 25 wt% of the second polymer based on the total weight content of the second reactor effluent.

Heat

Preferably, the method further includes a heating step a4) between step a3) and step c1), wherein in the heating step a4), the first reactor effluent extracted from the first polymerization reactor in step a3) is heated to obtain a heated first reactor effluent prior to the mixing step c1), and/or further includes a heating step b4) between step b3) and step c1), where the second reactor effluent stream recovered from the second polymerization reactor in step b3) is heated to obtain a heated second reactor effluent stream before mixing step c1) . More preferably, the method further includes a heating step a4) between step a3) and step c1), wherein in the heating step a4), the first reactor effluent extracted from the first polymerization reactor in step a3) is heated to obtain a heated first reactor effluent to the step mixing step c1), and further includes a heating step b4) between step b3) and step c1), where the second reactor effluent stream recovered from the second polymerization reactor in step b3) is heated to obtain a heated second reactor effluent stream before mixing step c1).

The purpose of heating steps a4) and b4) is to heat the first reactor effluent and the second reactor effluent, respectively, before they enter the mixing step c1) or the separation step before step c1).

Heating is preferably carried out in a first heater and/or a second heater as described above, such as a heat exchanger.

Typically, the temperature of the first reactor effluent before entering the first heater is 110 to 250°C, preferably 120 to 240°C, most preferably 130 to 230°C if the first polymer is an ethylene homo- or copolymer.

Typically, the temperature of the second reactor effluent before entering the second heater is 120 to 240°C, preferably 140 to 220°C, most preferably 150 to 200°C if the second polymer is an ethylene homo- or copolymer. Typically, the temperature of the second reactor effluent before entering the first heater is 120 to 250 °C, preferably 140 to 235 °C, most preferably 150 to 225 °C if the second polymer is a propylene homo- or copolymer.

Preferably, the temperature of the heated first reactor effluent, after the first heater, is usually from 200 to 300 °C, preferably from 210 to 260 °C and more preferably from 210 to 230 °C if the first polymer is an ethylene homo- or copolymer, or usually 200 to 300 °C, preferably 210 to 270 °C and more preferably 220 to 250 °C if the first polymer is a propylene homo- or copolymer.

Preferably, the temperature of the heated second reactor effluent, after the second heater, is typically 200 to 300 °C, preferably 210 to 260 °C, and more preferably 210 to 230 °C if the second polymer is an ethylene homo- or copolymer, or usually 200 to 300 °C, preferably 210 to 270 °C and more preferably 220 to 250 °C if the second polymer is a propylene homo- or copolymer.

It is preferable that the pressure of the heated first reactor effluent stream and the heated second reactor effluent stream does not significantly influence heating steps a4) and b4), respectively. The pressure of the heated first reactor effluent and the heated second reactor effluent is suitably 50 to 300 bar, preferably 50 to 250 bar, and more preferably 70 to 200 bar.

Separation

Preferably, the method further includes a separation step a5) between step a3) and step c1) or, if a first heater is present, between step a4) and step c1). In separation step a5), the first reactor effluent recovered from the first polymerization reactor in step a3), or the heated first reactor effluent recovered from the first heater in step a4), is separated in a first separator to produce a first polymer-lean vapor stream and a first polymer-rich vapor stream. liquid flow polymer.

Preferably, the method further includes a separation step b5) between step b3) and step c1) or, if a second heater is present, between step b4) and step c1). In separation step b5), the second reactor effluent recovered from the second polymerization reactor in step a3), or the heated second reactor effluent recovered from the second heater in step b4), is separated in a second separator to produce a second polymer-depleted vapor stream and a second enriched vapor stream. liquid flow polymer.

Preferably, the method includes both a separation step a5) and a separation step b5) as described above.

Preferably, the first separator and/or the second separator are low pressure separators. In low pressure separators, temperature and pressure are controlled to produce a liquid phase and a vapor phase. The polymer is dissolved in a liquid phase that contains a portion of an optional solvent and a portion of an optional unreacted comonomer, while the majority of the unreacted monomer, the optional unreacted chain transfer agent, an optional portion of the unreacted comonomer, and ultimately a portion of the solvent form the vapor phase.

The temperature in separation step a5) is preferably in the range from 120 to 240 °C, more preferably from 140 to 220 °C and more preferably from 150 to 200 °C if the polymer is an ethylene homo- or copolymer. The temperature in separation step a5) is preferably in the range from 120 to 240 °C, more preferably from 140 to 220 °C and more preferably from 150 to 200 °C if the polymer is a propylene homo- or copolymer.

The pressure in the separation step a5) is preferably 1 to 15 bar, more preferably 2 to 12 bar and more preferably 5 to 10 bar.

The temperature in separation step b5) is preferably in the range of 120 to 240 °C, more preferably 140 to 220 °C and more preferably 150 to 200 °C if the polymer is an ethylene homo- or copolymer. The temperature in separation step b5) is preferably in the range from 120 to 240 °C, preferably from 140 to 220 °C and more preferably from 150 to 200 °C if the polymer is a propylene homo- or copolymer.

The pressure in the separation step b5) is preferably 1 to 15 bar, more preferably 2 to 12 bar, and more preferably 5 to 10 bar.

The conditions in both separation steps a5) and b5) must be such that no undesired polymerization occurs after the reactors, which would require destruction of the catalysts, usually using polar substances. Thus, in a preferred aspect of the present invention, the catalyst destroying agent is not added to the first reactor effluent or the heated first reactor effluent and the second reactor effluent or the heated second reactor effluent before or during separation steps a5) and b5), respectively.

In another aspect of the present invention, which is however not preferred, a catalyst destroying agent is added to the first reactor effluent or heated first reactor effluent and the second reactor effluent or heated second reactor effluent before or during separation steps a5) and b5), respectively. The catalyst destroying agent is typically a polar component such as water, alcohols (such as methanol and ethanol), sodium/calcium stearate, CO, and combinations thereof.

As discussed above, the conditions in the separation step must be such that a vapor phase and a liquid phase are formed. Therefore, a polymer-rich phase and a polymer-depleted phase are prepared as described above. The temperature and pressure are adjusted such that vapor and liquid separate, resulting in a two-phase system containing a polymer-rich liquid phase and a polymer-depleted vapor phase. These two phases are then separated from each other.

The polymer-depleted vapor phase is then separated from the polymer-rich liquid phase in each of the low-pressure separators to form separated polymer-lean vapor streams and polymer-rich liquid streams.

The separation step can be carried out in accordance with any separation method known in the art, where a liquid phase and a vapor phase coexist. It is preferable to carry out the separation step as a flash evaporation step due to ease of implementation. As is well known in the art, the liquid feedstock is fed into a vessel operating at reduced pressure. In this way, part of the liquid phase evaporates and can be removed as an overhead stream (or vapor stream) from the flash evaporation. The portion remaining in the liquid phase is then recovered as an underflow (or liquid stream).

The advantage of having a vapor phase and a liquid phase present in the separation stage is, firstly, a simple device and therefore a low investment cost. In addition, polymer transfer with the vapor stream is relatively small. The polymer-enriched liquid stream is recovered from the liquid phase of the separation step.

In a preferred embodiment, the separation step is a flash evaporation step, as described below. The flash evaporation step is suitably carried out in a flash evaporation vessel as a separator, which is a vertical vessel, preferably having a generally cylindrical shape. Thus, the flash evaporation vessel has a section that has an approximately circular cross-section. Preferably, the flash evaporation vessel has a cylindrical section which has the shape of a circular cylinder. In addition to the cylindrical section, the flash vessel may have additional sections, such as a lower section, which may be conical, and an upper section, which may be hemispherical. Alternatively, the rapid evaporation vessel may also have a generally conical shape.

The temperature in the rapid evaporation vessel is typically between 120 and 240 °C. The temperature must be high enough to maintain the viscosity of the liquid stream at a suitable level, but below the temperature at which the polymer decomposes. The pressure in the flash vessel is usually between 15 bar and atmospheric or even below atmospheric.

The first reactor effluent or the heated first reactor effluent enters the flash vessel at the top. The liquid stream moves downward in the rapid evaporation vessel, while the gases evaporating from the liquid stream move upward. According to this preferred embodiment, the liquid stream forms a thin film which flows downward in the vessel for rapid evaporation. This facilitates the removal of hydrocarbons from the liquid stream. The first polymer-depleted vapor stream generated from the evaporated gases is typically withdrawn from the top of the flash vessel, while the first polymer-rich liquid stream is removed from the bottom. The above applies equally to the second reactor effluent stream or the heated second reactor effluent stream.

According to a particularly preferred embodiment, the first reactor effluent or the heated first reactor effluent is sprayed into the flash vessel. Atomization may be accomplished using one or more suitable nozzles that atomize the unreduced reactor effluent into droplets. Such nozzles are well known in the industry and include air atomization nozzles, flat spray nozzles, hollow cone nozzles and solid cone nozzles. Preferably, the nozzles break up the flow into droplets having a size of no more than about 1 mm. The above applies equally to the second reactor effluent or the heated second reactor effluent.

The nozzle creates a stream of droplets in the vessel for rapid evaporation. The stream of droplets then coagulates within the vessel for rapid evaporation and forms a downward film having a relatively large surface area. This improves the mass transfer of volatile components from the solution.

As described above, the flash vessel may have a vertical, generally cylindrical shape. The droplet flow is then directed tangentially to the vessel wall for rapid evaporation using a suitable nozzle position. Thus, the nozzle is suitably positioned relatively close to the wall so that its outlet is directed tangentially to the wall. When the stream of droplets leaves the nozzle, it moves towards the wall, forming a downward film. It is also possible that the rapid evaporation vessel has a vertical overall conical shape. In such an embodiment, it is possible to direct the flow of droplets tangentially to the wall of the vessel for rapid evaporation, as described above. However, it is also possible to direct the droplets axially towards the wall of the vessel for rapid evaporation. The nozzle or nozzles are then positioned asymmetrically within the vessel for rapid evaporation. In both arrangements, the liquid flow forms a cascading film inside the vessel for rapid evaporation.

The polymer content of the first polymer enriched liquid stream recovered from the first separator is typically from 40 to 90 wt%, preferably from 50 to 80 wt% and most preferably from 60 to 75 wt% based on the total weight content of the first enriched liquid flow polymer. In other words, the first polymer-enriched liquid stream recovered from the flash step typically contains from 10 to 60 wt.%, preferably from 20 to 50 wt.%, and most preferably from 25 to 40 wt.% residual hydrocarbons based on total mass content of the first polymer-enriched liquid stream.

Viewed from another angle, the first polymer-lean vapor stream recovered from the first separator represents from 35 to 80 wt.% of the total material streams recovered from the flash vessel. The polymer-lean vapor stream typically contains unreacted monomer as well as solvent and unreacted comonomer.

By using fast evaporation as described above, high separation efficiency can be achieved. For example, the separation efficiency for hydrocarbons containing 6 carbon atoms is at least 75% and preferably at least 80%. In addition, the separation efficiency for hydrocarbons containing 8 carbon atoms is at least 60% and preferably at least 65%. Separation efficiency is defined as the mass flow rate of the component recovered in the steam stream divided by the (theoretical) mass flow rate of the component in the steam stream under equilibrium conditions.

In the mixing step c1), the first polymer contained in the first polymer-enriched liquid stream and the second polymer contained in the second polymer-enriched liquid stream are mixed in a mixing unit to obtain a flow-mixed polymer.

In a preferred aspect, both the first and second polymer-enriched liquid streams are heated before, during or after said mixing in the mixing unit.

The mixing unit in step c1) is preferably a mixer or a separator.

When the mixer is used to mix polymer-enriched liquid streams, the mixer is preferably a static mixer. Static mixers are well known in the art, and one skilled in the art will be able to select a suitable mixer for the process. The use of a mixer improves the mixing of polymer-enriched liquid streams, on the one hand, and the mass transfer of volatile components from the liquid phase to the vapor phase, on the other hand, due to a significant increase in the mass transfer area.

In case the mixing unit is a separator, the separator is preferably a flash separator or a low pressure separator. The low pressure separator is typically only used when the separation efficiency of the low pressure separators in steps a5) and b5) is insufficient, ie the residual hydrocarbon content in the polymer rich liquid streams recovered from the low pressure separators is quite high.

When a low pressure separator is used to mix both polymer-rich liquid streams in step c1), i.e., essentially polymer-only streams, the thus obtained mixed polymer-lean vapor phase can be recycled back as a mixed polymer-lean vapor stream into any or all polymerization reactors.

To further illustrate the invention, two illustrative embodiments of the invention are described using FIGS. 1 and 2.

In fig. 1 shows an exemplary embodiment of an installation configuration in accordance with the first aspect of the invention. The installation includes a first reactor line and a second reactor line.

The first reactor line (1) for producing the first polymer includes a first polymerization reactor (2) and a first separator (3). The first polymerization reactor (2) includes a first reactor inlet for introducing a first feed stream into the first reactor (2) and a first reactor outlet for extracting a first reactor effluent stream containing the first polymer. The first outlet of the reactor is connected with the ability to transfer fluid through the first connecting line (4) with the inlet of the first separator (3). The first separator (3) includes a lower outlet for recovering the first polymer-enriched liquid stream, the lower outlet being connected via a second connecting line (14) to a mixing unit (13). The first separator (3) further includes an upper outlet for extracting the first polymer-depleted vapor stream. The first recycle line (5) connects the upper outlet of the separator (3a) back to the first polymerization reactor (2) to recycle the first polymer-lean vapor stream back to the first polymerization reactor (2).

Similar to the first reactor line (1), the second reactor line (7) according to the first aspect of the invention produces a second polymer and includes a second polymerization reactor (8) and a second separator (9). The second polymerization reactor (8) includes a second reactor inlet for introducing a second feed stream into the reactor (8) and a second reactor outlet for extracting a second reactor effluent containing the second polymer. The second outlet of the reactor is connected with the possibility of transferring the fluid through the third connecting line (10) with the inlet of the second separator (9). The second separator (9) includes a lower outlet for recovering a second polymer-enriched liquid stream, the lower outlet being connected via a fourth connection line (15) to a mixing unit (13). The second separator (9) further includes an upper outlet for extracting a second polymer-depleted vapor stream. A recycle line (11) connects the upper outlet of the separator (9) back to the second polymerization reactor (8) to recycle the second polymer-lean vapor stream back to the second polymerization reactor (8).

The internal volume of the first polymerization reactor (2) is 1.5 times larger than the internal volume of the first polymerization reactor (8).

The first heater (6) and the second heater (12) are located after the first polymerization reactor (2) and before the mixing unit (13) and after the second polymerization reactor (9) and before the mixing unit (13), respectively, see FIG. 1.

The first heater (2a) heats the first reactor effluent to provide a heated first reactor effluent, wherein the heated first reactor effluent is introduced into the first separator (3). The second heater (12) heats the second reactor effluent to provide a heated second reactor effluent, wherein the heated second reactor effluent is introduced into the second separator (9).

The mixing unit (13) is connected through a second connecting line (14) to the first separator (3) and through a fourth connecting line (15) to the second separator (9). Line (14) transfers the first polymer-enriched liquid stream from the first separator (3) to the mixing unit (13), while line (15) transfers the second polymer-enriched liquid stream from the second separator (9) to the mixing unit (13). In the mixing unit (13), which in this illustrative embodiment is a static mixer, the first polymer from the first polymer-enriched liquid stream and the second polymer from the second polymer-enriched liquid stream are in-line mixed so as to obtain an in-line mixed polymer.

The mixing unit (13) further includes a lower outlet connected to a withdrawal line (16) for extracting the in-line polymer stream containing the in-line mixed polymer.

In fig. 2 shows an exemplary embodiment of an installation configuration in accordance with the second aspect of the invention.

The installation according to the second aspect includes a first polymerization reactor (17a) and a second polymerization reactor (17b). The first polymerization reactor (17a) includes a first reactor inlet (20a) for introducing a first feed stream into the reactor (17a) and a first reactor outlet for extracting a first reactor effluent stream containing the first polymer. The first outlet of the reactor is connected with the possibility of transferring fluid through a connecting line (21a) and a line (22a) with the inlet of the mixing unit (19). The mixing unit (19) in this illustrative embodiment is a mixer.

The second polymerization reactor (17b) includes a second reactor inlet (20b) for introducing a second feed stream into the reactor (17b) and a second reactor outlet for extracting a second reactor effluent containing the second polymer. The second outlet of the reactor is connected with the possibility of transferring fluid through a connecting line (21b) and a line (22b) with the inlet of the mixing unit (19).

The internal volume of the first polymerization reactor (17a) is 1.5 times larger than the internal volume of the second polymerization reactor (17b), as shown schematically in FIG. 2.

The first heater (18a) and the second heater (18b) are located after the first polymerization reactor (17a) and before the mixing unit (19) and after the second polymerization reactor (17b) and before the mixing unit (19), respectively, see FIG. 2.

In the mixing unit (19), the first polymer from the first reactor effluent and the second polymer from the second reactor effluent are in-line mixed so as to obtain an in-line mixed polymer.

The mixing unit (19) further includes a lower outlet connected to a withdrawal line (23) for extracting the in-line polymer stream containing the in-line mixed polymer. The flash separator (24) is optionally located after the mixing unit (19).

EXAMPLES

1.Measuring methods

a) Melt flow rate (MFR) and flow rate ratio (FRR)

Melt flow rate (MFR) was determined in accordance with ISO 1133 - Determination of mass melt flow rate (MFR) and melt volume flow rate (MVR) of thermoplastic plastics - Part 1: Standard method and reported in g/10 min. MFR is a measure of the flowability and therefore the processing performance of a polymer. The higher the melt flow rate, the lower the viscosity of the polymer.

MFR ₂ of polypropylene was determined at a temperature of 230 °C and a load of 2.16 kg.

MFR ₂ of polyethylene was determined at a temperature of 190 °C and a load of 2.16 kg.

The flow rate ratio (FRR) is MFR ₂₁ /MFR ₂ .

b) Density

Polymer density was measured according to ISO1183.

c) Comonomer content

Quantitative nuclear magnetic resonance (NMR) spectroscopy was used to quantify the comonomer content of the polymers.

Quantitative ^13C { ^1H } NMR spectra were recorded in the molten state using a Bruker Avance III 500 NMR spectrometer operating at 500.13 and 125.76 MHz for ^1H and ^13C , respectively. All spectra were acquired using a ^13C optimized 7mm magic angle spinning (MAS) probe head at 150°C using nitrogen gas for all pneumatic systems. Approximately 200 mg of material was packed into a 7 mm OD zirconium MAS rotor and centrifuged at 4 kHz. This setup was chosen primarily for the high sensitivity required for rapid identification and accurate quantification. Standard single-pulse excitation was used using transient NOE (nuclear overhauser effect) with short recovery latency periods of 3 s and an RS-HEPT (rotor-synchronized Hahn echo pulse train) decoupling circuit. A total of 1024 (1k) transients were obtained per spectrum.

Quantitative ^13C { ^1H } NMR spectra were processed, integrated, and quantified using special spectral analysis automation programs. All chemical shifts are internally related to the bulk methylene (d+) signal at 30.00 ^ppm .

Characteristic signals corresponding to the incorporation of 1-octene were observed, and all comonomer contents were calculated relative to all other monomers present in the polymer.

Characteristic signals resulting from the isolated inclusion of 1-octene, that is, the comonomer sequence EEOEE, were observed. Isolated 1-octene inclusion was quantified using the integral of the signal at 38.3 ^ppm . This integral is assigned to the unresolved signals corresponding to both the *B6 and *bB6B6 sites of isolated (EEOEE) and isolated double non-sequential (EEOEOEE) 1-octene sequences, respectively. To compensate for the influence of two *bB6B6 sites, the integral of the bbB6B6 site at 24.6 ^ppm was used:

O = I _*B6+*bB6B6 - 2×I _bbB6B6

Characteristic signals resulting from the sequential incorporation of 1-octene, that is, the comonomer sequence EEOOEE, were also observed. This sequential 1-octene incorporation was quantified using the integral of the signal at 40.4 ^ppm attributable to aaB6B6 sites, taking into account the number of reporter sites per comonomer:

OO = 2×I _aaB6B6

Characteristic signals resulting from the isolated non-sequential incorporation of 1-octene, that is, the comonomer sequence EEOEOEE, were also observed. This isolated, inconsistent 1-octene incorporation was quantified using the integral of the signal at 24.6 ^ppm attributable to the bbB6B6 sites, taking into account the number of reporter sites per comonomer:

OEO = 2×I _bbB6B6

Characteristic signals resulting from isolated triple sequential incorporation of 1-octene, that is, the comonomer sequence EEOOOEE, were also observed. This isolated triple sequential 1-octene incorporation was quantified using the integral of the signal at 41.2 ^ppm attributable to the aagB6B6B6 sites, taking into account the number of reporter sites per comonomer:

OOO = 3/2×I _aagB6B6B6

In the absence of other signals indicating other comonomer sequences, total 1-octene comonomer content was calculated solely based on the number of isolated (EEOEE), isolated double-sequential (EEOOEE), isolated non-sequential (EEOEOEE), and isolated triple-sequential (EEOOOEE) 1-octene comonomer sequences:

O _total = O + OO + OEO + OOO

Characteristic signals resulting from the saturated end groups were observed. These saturated end groups were quantified using the average integral of the two resolved signals at 22.9 and 32.23 ^ppm . The integral at 22.84 ^ppm is assigned to the unresolved signals corresponding to both the 2B6 and 2S sites of the 1-octene and saturated chain end, respectively. The integral at 32.2 ^ppm is assigned to the unresolved signals corresponding to both the 3B6 and 3S sites of the 1-octene and saturated chain end, respectively. To compensate for the influence of the 2B6 and 3B6 sites of 1-octene, the total content of 1-octene was used:

S = (1/2)×(I _2S+2B6 + I _3S+3B6 - 2×O _total ).

Ethylene comonomer content was quantified using the integral of the bulk methylene (v/v) signals at 30.00 ^ppm . This integral included the D and 4B6 sites from the 1-octene, as well as the D ^D sites. Total ethylene comonomer content was calculated based on volume integral and compensation for the observed 1-octene sequences and end groups:

E _total = (1/2)×[I _rev + 2×O + 1×OO + 3×OEO + 0×OOO + 3×S].

It should be noted that volume integral compensation for the presence of isolated triple inclusion (EEOOOEE) of 1-octene sequences is not required, since the number of undercounted and overcounted ethylene units is the same.

The total mole fraction of 1-octene in the polymer was then calculated as follows:

fO = _Otot /( _Etot + _Otot ).

The total inclusion of 1-octene comonomer in mass percent was calculated from the mole fraction in the standard way:

O [wt%] = 100×(fO×112.21)/((fO×112.21)+((1-fO)×28.05)).

Additional information can be found in the following sources:

Klimke, K., Parkinson, M., Piel, C., Kaminsky, W., Spiess, H.W., Wilhelm, M., Macromol. Chem. Phys. 2006; 207:382.

Parkinson, M., Klimke, K., Spiess, H.W., Wilhelm, M., Macromol. Chem. Phys. 2007; 208:2128.

NMR Spectroscopy of Polymers: Innovative Strategies for Complex Macromolecules, Chapter 24, 401 (2011)

Pollard, M., Klimke, K., Graf, R., Spiess, H.W., Wilhelm, M., Sperber, O., Piel, C., Kaminsky, W., Macromolecules 2004; 37:813.

Filip, X., Tripon, C., Filip, C., J. Mag. Resn. 2005, 176, 239

Griffin, J. M., Tripon, C., Samoson, A., Filip, C., and Brown, S. P., Mag. Res. in Chem. 2007 45, S1, S198

Castignolles, P., Graf, R., Parkinson, M., Wilhelm, M., Gaborieau, M., Polymer 50 (2009) 2373

Zhou, Z., Kuemmerle, R., Qiu, X., Redwine, D., Cong, R., Taha, A., Baugh, D. Winniford, B., J. Mag. Reason. 187 (2007) 225

Busico, V., Carbonniere, P., Cipullo, R., Pellecchia, R., Severn, J., Talarico, G., Macromol. Rapid Commun. 2007, 28, 1128

J. Randall, Macromol. Sci., Rev. Macromol. Chem. Phys. 1989, C29, 201.

Qiu, X., Redwine, D., Gobbi, G., Nuamthanom, A., Rinaldi, P., Macromolecules 2007, 40, 6879

Liu, W., Rinaldi, P., McIntosh, L., Quirk, P., Macromolecules 2001, 34, 4757

d) Unsaturation

Quantitative nuclear magnetic resonance (NMR) spectroscopy was used to quantify the content of unsaturated groups present in the polymers.

Quantitative ¹ H NMR spectra were recorded in solution using a Bruker Avance III 400 NMR spectrometer operating at 400.15 MHz. All spectra were recorded using a ¹³ C optimized 10 mm head for selective excitation at 125 °C using nitrogen gas for all pneumatic systems. Approximately 200 mg of material was dissolved in 1,2-tetrachloroethane-d ₂ (TCE-d ₂ ) using approximately 3 mg of Hostanox 03 (CAS 32509-66-3) as a stabilizer. Standard single-pulse excitation was applied using a 30-degree pulse, a 10-s recovery delay, and 10-Hz sample rotation. A total of 128 transitions per spectrum were obtained using 4 blank scans. This setup was chosen primarily for the high resolution needed to quantify the unsaturation and stability of vinylidene groups. All chemical shifts were indirectly referenced to TMS at 0.00 ^ppm using the signal obtained from the residual protonated solvent at 5.95 ^ppm .

Characteristic signals corresponding to the presence of terminal aliphatic vinyl groups (R-CH=CH ₂ ) were observed and the quantity was determined using the integral of the two bound nonequivalent terminal protons of CH ₂ (Va and Vb) at 4.95, 4.98 and 5.00 and 5.05 ^ppm taking into account the number of reporter sites per functional group:

Nvinyl = IVab/2.

When characteristic signals consistent with the presence of internal vinylidene groups (RR'C=CH ₂ ) were observed, the amount was quantified using the integral of two CH ₂ protons (D) at 4.74 ^ppm , taking into account the number of reporter sites per functional group:

Nvinylidene = ID/2.

When characteristic signals corresponding to the presence of internal cis-vinylene groups (E-RCH=CHR') or related structure were observed, the amount was quantified using the integral of the two CH(C) protons at 5.39 ^ppm , taking into account the number of reporter sites per functional group:

Ncis = IC/2.

When characteristic signals corresponding to the presence of internal trans-vinylene groups (Z-RCH=CHR') were observed, the amount was quantified using the integral of the two CH(T) protons at 5.45 ^ppm , taking into account the number of reporter sites per functional group:

Ntrans = IT/2.

When characteristic signals consistent with the presence of internal trisubstituted vinylene groups (RCH=CHR'R'') or related structure were observed, the amount was quantified using the integral of the CH proton (Tris) at 5.14 ^ppm , taking into account the number of reporter sites per functional group :

Ntris = ITris

Hostanox 03 stabilizer was quantified using the multiplet integral of aromatic protons (A) at 6.92, 6.91, 6.69 and 6.89 ^ppm and taking into account the number of reporter sites per molecule:

H = IA/4

As is typical for the quantification of unsaturation in polyolefins, the amount of unsaturation was determined relative to the total number of carbon atoms, even though quantified by ¹ H NMR spectroscopy. This allows direct comparison with other quantitative microstructure values obtained directly from ¹³ C NMR spectroscopy.

The total number of carbon atoms was calculated from the integral of the bulk aliphatic signal between 2.85 and minus 1.00 ^ppm , with compensation for methyl signals from the stabilizer and carbon atoms belonging to unsaturated functional groups not included in this scope:

NCtotal = (Irev – 42×H)/2 + 2×Nvinyl + 2×Nvinylidene + 2×Ncis + 2×Ntrans + 2×Ntris

The content of unsaturated groups (U) was calculated as the number of unsaturated groups in the polymer per thousand total carbon atoms (kCHn):

U = 1000×N/NCtot

The total number of unsaturated groups was calculated as the sum of the individual observed unsaturated groups and thus also reported as a ratio per thousand total carbon atoms:

Utotal = Uvinyl + Uvinylidene + Ucis + Utrans + Utris

The relative content of a particular unsaturated group (U) was expressed as the proportion or percentage of that unsaturated group relative to the total amount of unsaturated groups:

[U] = Ux/Utot

Additional information can be found in the following sources:

He, Y., Qiu, X., and Zhou, Z., Mag. Res. Chem. 2010, 48, 537-542.

Busico, V. et. al. Macromolecules, 2005, 38 (16), 6988-6996

e) Determination of average molecular weights, molecular weight distribution

The average molecular weights (Mz, Mw and Mn), molecular weight distribution (MWD) and its breadth, described by the polydispersity coefficient, PDI = Mw/Mn (where Mn is the number average molecular weight and Mw is the mass average molecular weight), were determined using gel permeation chromatography (GPC) in accordance with ISO 16014-1:2003, ISO 16014-2:2003, ISO 16014-4:2003 and ASTM D 6474-12 using the following formulas:

For a constant elution volume interval, ΔV _i , where A _i and M _i are the chromatographic peak cutoff area and polyolefin molecular weight (MW), respectively, associated with the elution volume, V _i , where N is equal to the number of data points obtained in the chromatogram between the limits integration.

A high-temperature GPC instrument equipped with a multi-band infrared (IR) detector model IR5 (PolymerChar, Valencia, Spain) equipped with 3 Agilent-PLgel Olexis columns and 1 Agilent-PLgel Olexis Guard column was used. 1,2,4-trichlorobenzene (TCB) stabilized with 250 mg/L 2,6-di-tert-butyl-4-methylphenol was used as a solvent and mobile phase. The chromatographic system was used at 160 °C at a constant flow rate of 1 mL/min. 200 μl of the test solution was injected for analysis. Data collection was performed using PolymerChar GPC-one software.

The column set was calibrated using a universal calibration (according to ISO 16014-2:2003) with 19 narrow MWD polystyrene (PS) standards ranging from 0.5 kg/mol to 11500 kg/mol. PS standards were dissolved at room temperature over several hours. Conversion from polystyrene peak molecular weight to polyolefin molecular weight was carried out using the Marc-Houwink equation and the following Marc-Houwink constants:

K _PS = 19×10 ^-3 ml/g, α _PS = 0.655

K _PE = 39×10 ^-3 ml/g, α _PE = 0.725

A third-order polynomial fit was used to fit the calibration data.

All samples were prepared at concentrations ranging from 0.5 to 1 mg/mL and dissolved at 160 °C for 3 hours with continuous gentle shaking.

f) Melting temperature ( _Tm ) and crystallization temperature ( _Tc )

Experiments were performed with TA Q200 instruments calibrated using indium, zinc, tin, and ISO 11357-3. About 5 mg of material was placed in a pan and tested at 10 °C/min during the experiments at a nitrogen flow rate of 50 ml/min with lower and higher temperatures of −30 °C and 180 °C, respectively. Only the second heating cycle was considered for analysis. The melting temperature T _m is defined as the temperature of the main peak of the thermogram, while the enthalpy of fusion (∆Hm) was calculated by integrating between 10 °C and the end of the thermogram, typically T _m plus 15 °C. The integral along the direction of movement in this range was also calculated.

g) Glass transition temperature ( _Tg )

The glass transition temperature Tg was determined by dynamic mechanical analysis in accordance with ISO 6721-7. Measurements were carried out in torsion mode on pressed samples (40×10×1 mm ³ ) at temperatures from minus 100 °C to plus 150 °C with a heating rate of 2 °C/min and a frequency of 1 Hz.

h) Vicat softening point (T _Vicat )

Vicat temperatures were measured in accordance with ISO 306, method A50. A flat-tipped needle loaded with a mass of 10 N was brought into direct contact with an injection molded test piece having dimensions of 80 x 10 x 4 mm ³ as described in EN ISO 1873-2. The sample and needle were heated at 50 °C/h. The temperature at which the needle penetrated to a depth of 1 mm was recorded as the Vicat softening point.

2. Materials

a) Comparative example 1 (CE1)

CE1 is an ethylene-based 1-octene plastomer (octene content 15.7 wt%) having an MFR ₂ of 1.1 g/10 min, a density of 902 kg/m ³ and a melting point T _{m of} 97 °C, commercially available from Borealis. CE1 was prepared by a solution polymerization process using a metallocene catalyst.

b) Copolymer A is an ethylene-based 1-octene plastomer (octene content 11.9 wt%) obtained by solution polymerization using a metallocene catalyst, having an MFR ₂ of 1.1 g/10 min, density 910 kg/ m ³ and melting temperature T _m 106 °C.

c) Copolymer B is an ethylene-based 1-octene plastomer (octene content 25.8 wt%) obtained by solution polymerization using a metallocene catalyst, having an MFR ₂ of 1.1 g/10 min, density 883 kg/ m ³ and melting temperature T _m 73 °C.

d) Copolymer C is an ethylene-based 1-octene elastomer (octene content 37.1 wt%) obtained by a solution polymerization process using a metallocene catalyst, having an MFR _{2 of} 1.0 g/10 min, a density of 862 kg/ m ³ and melting temperature T _m 35 °C.

e) Copolymer D is an ethylene-based 1-octene elastomer (octene content 31.5 wt%) obtained by solution polymerization using a metallocene catalyst, having an MFR _{2 of} 1.0 g/10 min, density 870 kg/ m ³ and melting temperature T _m 56 °C.

Copolymers A-D were prepared using Borceed ^™ solution polymerization technology from Borealis, the metallocene catalyst used in the present invention was (phenyl)(cyclohexyl)methylene(cyclopentadienyl)(2,7-di-tert-butylfluorenyl)hafnium dimethyl and N,N- dimethylanilinium tetrakis(pentafluorophenyl)borate (AB) (CAS 118612-00-3), commercially available from Boulder, as a cocatalyst.

The polymerization conditions were chosen in such a way that the reaction system represents a single liquid phase. (T 130 to 230 °C; 60 to 150 bar)

3. Results

Mixing of the appropriate material was carried out using a Prism TSE-16, 16 mm co-rotating twin screw extruder with an L/D of 25, with a throughput of approximately 1.4 kg/h. The temperature profile was set to 180–200 °C and the instrument was operated at 250 rpm. Samples were prepared by mixing a dry mixture of base resin granules and extruding the mixture. About 2.5 kg of the dry mixture was fed into the batch hopper and after stabilization, about 2.0 kg of the finished extruded mixture was collected.

Inventive examples IE1-1 to IE1-3 are mixtures of two copolymers in specific blend ratios. The results are presented in Table 1 below.

Table 1: Results

CE1 IE1-1 IE1-2 IE1-3 Mixture ratio - 83 wt.%
Copolymer A
17 wt.%
Copolymer C 80 wt.%
Copolymer A
20 wt.%
Copolymer D 71 wt.%
Copolymer A
29 wt.%
Copolymer B C8 content, wt.% 15.7 14.9 15.2 15.3 Density, kg/m ³ 902 902 903.1 902.1 _Mw , g/mol 81650 81350 82800 82250 _Mw / _Mn 2.6 2.72 2.64 2.72 MFR ₂ , g/10min 1.1 1.02 0.99 1.03 MFR ₂₁ , g/10 min 31.54 37.05 37.53 36.3 MFR ₂₁ /MFR ₂ 30.62 36.32 37.91 35.24 _Tm , °C 97 102.63 103.7 102.01 _Tc , °C 78.42 89.9 _Tg , °C minus 35.48 minus 41.55 minus 41.55 T _Vicat , °C 82 87.2 Vinylidene, 100 kCHn 12.3 11.5 12.1 12.1 Vinyl, 100 kCHn 5.6 5.0 6.9 5.8 Trisubst, 100 kCHn 19.2 17.7 17.40 21.2 Vinylene, 100 kCHn 8.8 12.7 14.40 14.0

The above results show that mixing two different copolymers targeting an existing product (CE1) results in copolymers (IE1-1 to IE1-3) with significantly better melting temperature _Tm , as well as improved _Tg , improved _Tc and improved T _Vicat at comparable density, melt flow rate, M _w and 1-octene comonomer content.

Claims (38)

1. Installation for producing flow-mixed polymer, including the first reactor line to produce the first polymer, a second reactor line to produce a second reactor and a mixing unit for in-line mixing the first polymer with the second polymer to produce an in-line mixed polymer, wherein the first reactor line includes a first polymerization reactor for producing a first polymer and a first separator located downstream of the first polymerization reactor, the second reactor line includes a second polymerization reactor for producing a second polymer and a second separator located after the second polymerization reactor, where both the first separator and the second separator are connected to a mixing unit located after the first separator and the second separator, where the first polymerization reactor has a first internal volume and the second polymerization reactor has a second internal volume, and characterized in that the ratio of the first internal volume and the second internal volume is in the range from 95:5 to 55:45, and in that the mixing unit contains a cleaning separator. 2. Installation according to claim 1, where after the first polymerization reactor and before the first separator the first heater is located and/or where after the second polymerization reactor and before the second separator the second heater is located. 3. Installation according to claim 1 or 2, where the first separator includes an upper outlet and a lower outlet and/or where the second separator contains an upper outlet and a lower outlet. 4. Installation for producing flow-mixed polymer, including the first polymerization reactor for producing the first polymer, a second polymerization reactor for producing a second polymer and a mixing unit for in-line mixing the first polymer with the second polymer to produce an in-line copolymer, where the first polymerization reactor and the second polymerization reactor are connected to a mixing unit located after the first polymerization reactor and the second polymerization reactor, where after the first polymerization reactor and before the mixing block the first heater is located and/or where after the second polymerization reactor and before the mixing block the second heater is located, where the first polymerization reactor has a first internal volume and the second polymerization reactor has a second internal volume, and characterized in that the ratio of the first internal volume and the second internal volume is in the range from 95:5 to 55:45, and in that the mixing unit contains a cleaning separator. 5. Installation according to any one of paragraphs. 1-4, where the first polymerization reactor includes a first reactor inlet for introducing a first feed stream into the first polymerization reactor and a first reactor outlet for extracting a first reactor effluent stream containing the first polymer, and/or where the second polymerization reactor includes a second reactor inlet for introducing a second feed stream into a second polymerization reactor; and a second reactor outlet for recovering a second reactor effluent stream containing the second polymer. 6. Installation according to any one of paragraphs. 1-5, where the ratio of the first internal volume and the second internal volume is from 85:15 to 60:40. 7. Installation according to any one of paragraphs. 1-6, where the mixing unit has an outlet for extracting the in-line mixed polymer. 8. A method for producing a flow-mixed polymer, carried out in an installation according to any one of paragraphs. 1-7 and including stages a1) introducing a first feed stream containing a first monomer into a first polymerization reactor, a2) polymerizing the first monomer in the presence of a first catalyst in a first polymerization reactor to produce a first polymer, a3) removing the first reactor effluent containing the first polymer from the first polymerization reactor, b1) introducing a second feed stream containing a second monomer into a second polymerization reactor, b2) polymerizing a second monomer in the presence of a second catalyst in a second polymerization reactor to produce a second polymer, b3) removing the second reactor effluent containing the second polymer from the second polymerization reactor, c1) mixing the first polymer and the second polymer in a mixing unit to obtain a flow-mixed polymer. 9. The method of claim 8, wherein the first catalyst contains a metallocene complex and/or the second catalyst contains a metallocene complex. 10. The method according to claim 8 or 9, wherein the polymerization step a2) is carried out at a first reaction temperature and the polymerization step b2) is carried out at a second reaction temperature, where the first reaction temperature is the same as or different from the second reaction temperature. 11. The method according to claim 9 or 10, wherein the polymerization step a2) is carried out at a first reactor pressure and the polymerization step b2) is carried out at a second reactor pressure, where the first reactor pressure is the same as or different from the second reactor pressure. 12. Method according to any one of paragraphs. 8-11, where the first monomer and/or the second monomer is ethylene. 13. Method according to any one of paragraphs. 8-12, wherein the first feed stream further comprises a comonomer, preferably octene, and/or where the second feed stream further comprises a comonomer, preferably octene. 14. Method according to any one of paragraphs. 8-13, where the first feed stream further contains a solvent and/or a chain transfer agent and/or where the second feed stream further contains a solvent and/or a chain transfer agent.