Process for the polymerization of olefins

Abstract

The present invention relates to a process for the continuous preparation of a polyolefin in a reactor from one or more -olefin monomers of which at least one is ethylene or propylene, wherein the reactor comprises a fluidized bed, an expanded section located at or near the top of the reactor, a distribution plate located at the lower part of the reactor and an inlet for a recycle stream located under the distribution plate wherein the process comprisesfeeding a polymerization catalyst to the fluidized bed in the area above the distribution platefeeding the one or more -olefin monomers to the reactorwithdrawing the polyolefin from the reactorcirculating fluids from the top of the reactor to the bottom of the reactor, wherein the circulating fluids are compressed using a compressor and subsequently cooled using a heat exchanger to form a cooled recycle stream comprising liquid, and wherein the cooled recycle stream is introduced into the reactor using the inlet for the recycle stream wherein a part of the cooled recycle stream is drawn to form a liquid comprising stream, wherein the liquid comprising stream is introduced into the expanded section during at least part of the polymerization process, and wherein the liquid comprising stream is brought into contact with at least part of the interior surface of the expanded section.

Claims

1. A process for the continuous preparation of a polyolefin in a reactor from one or more -olefin monomers of which at least one is ethylene or propylene, wherein the reactor comprises a fluidized bed, an expanded section located at or near the top of the reactor, a distribution plate located at the lower part of the reactor and an inlet for a recycle stream located under the distribution plate; wherein the process comprises feeding a polymerization catalyst to the fluidized bed in the area above the distribution plate feeding the one or more -olefin monomers to the reactor withdrawing the polyolefin from the reactor circulating fluids from the top of the reactor to the bottom of the reactor, wherein the circulating fluids are compressed using a compressor and subsequently cooled using a heat exchanger to form a cooled recycle stream comprising liquid, and wherein the cooled recycle stream is introduced into the reactor using the inlet for the recycle stream wherein a part of the cooled recycle stream is drawn to form a liquid comprising stream, wherein the liquid comprising stream is introduced into the expanded section during at least part of the polymerization process, and wherein the liquid comprising stream is brought into contact with at least part of the interior surface of the expanded section.

2. The process according to claim 1, wherein a gaseous stream is drawn from a point downstream of the compressor and upstream of the heat exchanger and mixed with part of the cooled recycle stream to form the liquid comprising stream.

3. The process according to claim 1, wherein an alkane containing stream comprising an inert alkane is added to a reaction system comprising the reactor and the cooled recycle stream.

4. The process according to claim 1, wherein the inert alkane is chosen from alkanes having 3 to 6 carbon atoms.

5. The process according to claim 1, wherein a thermal run away reducing agent is added to a reaction system comprising the reactor and the cooled recycle stream.

6. The process according to claim 5, wherein the thermal run away reducing agent is chosen from esters, amines, nitriles, amides and mixtures thereof.

7. The process according to claim 1, wherein the liquid comprising stream is added to the expanded section in a circumferential manner.

8. The process according to claim 1, wherein the liquid comprising stream is added to the expanded section via a pipe ring with a plurality of nozzles directed to the interior surface of the expanded section.

9. The process according to claim 1, wherein the alkane is added such that the molar composition of the alkane in the reactor is at least 1 mol %.

10. The process according to claim 1, wherein the reactor is a multi-zone reactor suitable for the continuous fluidized bed polymerization of one or more -olefin monomers of which at least one is ethylene or propylene, which multi-zone reactor is operable in condensed mode, which multi-zone reactor comprises a first zone, a second zone, a third zone, a fourth zone and a distribution plate, wherein the first zone is separated from the second zone by the distribution plate, wherein the multi-zone reactor is extended in the vertical direction, wherein the second zone of the multi-zone reactor is located above the first zone and wherein the third zone of the multi-zone reactor is located above the second zone, and wherein the fourth zone of the multi-zone reactor is located above the third zone, wherein the second zone contains an inner wall, wherein at least part of the inner wall of the second zone is either in the form of a gradually increasing inner diameter or a continuously opening cone, wherein the diameter or the opening increases in the vertical direction towards the top of the multi-zone reactor, wherein the third zone contains an inner wall, wherein at least part of the inner wall of the third zone is either in the form of a gradually increasing inner diameter or a continuously opening cone, wherein the diameter or the opening increases in the vertical direction towards the top of the multi-zone reactor, wherein the largest diameter of the inner wall of the third zone is larger than the largest diameter of the inner wall of the second zone.

11. The process according to claim 1, wherein the polyolefin is polypropylene.

12. The process according to claim 4, wherein the inert alkane is i-butane.

13. The process according to claim 5, wherein a thermal run away reducing agent is added to the liquid comprising stream.

14. The process according to claim 6, wherein the thermal run away reducing agent is p-ethoxy ethyl benzoate (PEEB).

15. The process according to claim 9, wherein the alkane is added such that the molar composition of the alkane in the reactor is at least 2 mol %.

16. The process according to claim 9, wherein the alkane is added such that the molar composition of the alkane in the reactor is at most 5 mol %.

17. The process according to claim 1, wherein the liquid comprising stream is introduced to the expanded section through a nozzle oriented tangential to the interior surface.

18. The process according to claim 1, wherein a superficial gas velocity in the reactor is at least 1.5 m/s and at most 5 m/s.

19. A process for the continuous preparation of a polyolefin in a reactor from one or more -olefin monomers of which at least one is ethylene or propylene, wherein the reactor comprises a fluidized bed, a distribution plate located at the lower part of the reactor and an inlet for a recycle stream; wherein the process comprises feeding the one or more -olefin monomers to the reactor withdrawing the polyolefin from the reactor circulating fluids from the top of the reactor to the bottom of the reactor, wherein the circulating fluids are compressed using a compressor and subsequently cooled using a heat exchanger to form a cooled recycle stream comprising liquid, and wherein the cooled recycle stream is introduced into the reactor using the inlet for the recycle stream wherein a part of the cooled recycle stream is drawn to form a liquid comprising stream, wherein the liquid comprising stream is introduced into the expanded section during at least part of the polymerization process.

20. The process according to claim 19, wherein the liquid comprising stream is introduced to the expanded section through a nozzle oriented tangential to the interior surface.

Description

BRIEF DESCRIPTION OF THE FIGURES

(1) FIG. 1 illustrates a system suitable for the continuous preparation of a polyolefin according to the process of the invention.

(2) FIG. 2 illustrates a special embodiment of the system of FIG. 1.

(3) FIG. 3 illustrates a special embodiment of the system of FIG. 1, wherein the reactor is a multi-zone reactor.

(4) FIG. 4 illustrates a special embodiment of the flow of part of the cooled recycle stream in the expanded section of the reactor.

(5) FIG. 5 is a TREF profile of the polymer produced in example 1.

(6) FIG. 6 is a TREF profile of the polymer produced in example 2.

(7) FIG. 1 illustrates a system suitable for the continuous preparation of a polyolefin in a reactor from one or more -olefin monomers of which at least one is ethylene or propylene, comprising a reactor (8), a compressor (400), a heat exchanger (5), an expanded section (4) (which expanded section is located at or near the top of the reactor) wherein the reactor comprises a distribution plate (6) (which distribution plate is located at the lower part of the reactor) and an inlet for a cooled recycle stream (10) located under the distribution plate an inlet for providing the catalyst (20) an outlet for providing the polyolefin (30) an outlet for a recycle stream (40),

(8) wherein the outlet for the recycle stream (40) of the expanded section is connected to an inlet of the compressor (400) via a first connection means (AA), for instance pipes

(9) wherein the compressor (400) comprises an outlet for compressed fluids (50),

(10) wherein the outlet of the compressor (400) is connected to an inlet for compressed fluids of the heat exchanger (5) via a second connection means (BB)

(11) wherein optionally the second connection means (BB), for instance pipes, comprises an inlet for receiving a feed (70),

(12) wherein the heat exchanger (5) comprises an outlet for providing the cooled recycle stream (10) which outlet of the heat exchanger (5) is connected to the inlet of the reactor (8) for receiving the cooled recycle stream (10),

(13) wherein the first connection means (AA) may comprise an inlet for receiving a feed (60)

(14) wherein the reactor further comprises a means for receiving the liquid comprising stream and for introducing it into the expanded section. (d) represents the optional TRRA. (a) represents the optional alkane containing stream and/or TRRA.

(15) FIG. 2 (FIG. 2) illustrates a preferred embodiment of the system of FIG. 1 and shows that a gaseous stream (b4) and the liquid-containing stream (b3) are connected to a mixing unit (3) which mixing unit is connected to the means for receiving the liquid comprising stream (b5) and wherein the system further comprises one or more inlets for receiving the alkane containing stream and/or TRRA (a).

(16) FIG. 3 (FIG. 3) illustrates a special embodiment of the system of FIG. 3, wherein the reactor is a multi-zone reactor,

(17) which multi-zone reactor is suitable for the continuous fluidized bed polymerization of one or more -olefin monomers of which at least one is ethylene or propylene, which multi-zone reactor is operable in condensed mode, which multi-zone reactor comprises a first zone, a second zone, a third zone, a fourth zone and a distribution plate,

(18) wherein the first zone is separated from the second zone by the distribution plate,

(19) wherein the multi-zone reactor is extended in the vertical direction

(20) wherein the second zone of the multi-zone reactor is located above the first zone and wherein the third zone of the multi-zone reactor is located above the second zone,

(21) and wherein the fourth zone of the multi-zone reactor is located above the third zone

(22) wherein the second zone contains an inner wall, wherein at least part of the inner wall of the second zone is either in the form of a gradually increasing inner diameter or a continuously opening cone, wherein the diameter or the opening increases in the vertical direction towards the top of the multi-zone reactor

(23) wherein the third zone contains an inner wall, wherein at least part of the inner wall of the third zone is either in the form of a gradually increasing inner diameter or a continuously opening cone, wherein the diameter or the opening increases in the vertical direction towards the top of the multi-zone reactor

(24) wherein the largest diameter of the inner wall of the third zone is larger than the largest diameter of the inner wall of the second zone,

(25) wherein the zone (2) in the area directly above the distribution plate is either in the form of a gradually increasing inner diameter or a continuously opening cone (2A), wherein the diameter or the opening increases in the vertical direction towards the top of the multi-zone reactor and wherein the top part of the second zone has an inner wall having a cylindrical shape (2B) and wherein the top part of the second zone is connected to a bottom part of the third zone (3A),

(26) wherein the bottom part of the third zone is either in the form of a gradually increasing inner diameter or a continuously opening cone, wherein the diameter or the opening increases in the vertical direction towards the top of the multi-zone reactor and wherein the top part of the third zone has an inner wall having a cylindrical shape (3B) and wherein the top part of the third zone is connected to the top zone, for example to the fourth zone.

(27) In FIG. 1 and FIG. 2, the fluidized bed (80) is indicated with dots.

(28) FIG. 4 (FIG. 4) illustrates the flow of the liquid-containing stream and of optional stream (d) in the expanded section, when this stream is added to the expanded section in a circumferential manner. In this figure, it is shown how the use of a nozzle that is tangential to the interior wall of the expanded section forces the (centrifugal/circumferential) movement of the liquid-containing stream and of optional stream (d) along the interior walls.

(29) In another aspect, therefore, the invention relates to a reaction system for the continuous preparation of a polyolefin in a reactor from one or more -olefin monomers of which at least one is ethylene or propylene, wherein the reaction system comprises a reactor (8), wherein the reactor comprises a fluidized bed (80), an expanded section (4) located at or near the top of the reactor, a distribution plate (6) located at the lower part of the reactor and an inlet for a recycle stream (10) located under the distribution plate

(30) wherein the system is arranged such that a polymerization catalyst is fed to the fluidized bed in the area above the distribution plate the one or more -olefin monomers are fed to the reactor the polyolefin is withdrawn from the reactor fluids are circulated from the top of the reactor to the bottom of the reactor, wherein the circulating fluids are compressed using a compressor (400) and subsequently cooled using a heat exchanger (5) to form a cooled recycle stream comprising liquid, and wherein the cooled recycle stream is introduced into the reactor using the inlet for the recycle stream

(31) and a part of the cooled recycle stream (b1) is drawn to form a liquid comprising stream, wherein the liquid comprising stream is introduced into the expanded section during at least part of the polymerization process,

(32) and wherein the liquid comprising stream is brought into contact with at least part of the interior surface of the expanded section.

(33) In the process and system of the invention, the one or more -olefin monomers and other fluids, such as hydrogen, an inert gas or liquid, for example a condensable inert component, may be added to the recycle stream (40) to make up for reacted fluids before cooling the fluids to form the cooled recycle stream.

(34) The feed (60) comprises a chain transfer agent, for example hydrogen and may further comprise gaseous -olefin monomers and insert gaseous components, for example nitrogen. A chain transfer agent, such as hydrogen may for instance be used to adjust the molecular weight of the polyolefin (30) produced.

(35) The feed (70) comprises condensable inert components, for example a condensable inert component selected from the group of alkanes having 4 to 20 carbon atoms, preferably 4 to 8 carbon atoms, and mixtures thereof, for example propane, n-butane, isobutene, n-pentane, isopentane, neopentane, n-hexane, isohexane or other saturated hydrocarbons having 6 C-atoms, n-heptane, n-octane and other saturated hydrocarbons having 7 or 8 C-atoms and any mixtures thereof; and may further comprise condensable -olefin monomers, -olefin comonomers and/or mixtures thereof.

(36) The condensable inert component is preferably selected from the group of isopentane, n-hexane, n-butane, i-butane and mixtures thereof. Because of their more attractive pricing, preferably isopentane and/or n-hexane are/is used as condensable inert component(s) in the feed (70)

(37) When copolymers are produced, the process of the invention further comprises supplying a comonomer using feed (60) or (70) in case of a non-condensable comonomer and using feed (70) in case of a condensable comonomer.

(38) It is apparent to the skilled person that the process of the present invention may also be applied when using multiple reactor. For example, for the purpose of the present invention, it is to be understood that if multiple reactors are employed, the liquid-containing stream may be added to the expanded section of any one of the reactorsand in case of two reactors, either reactorand that the liquid-containing stream need not be added to all reactors of the multiple reactor train.

(39) With condensed mode is meant that a liquid comprising stream is used to cool the multi-zone reactor (8).

(40) Hydrogen may for instance be used as a chain transfer agent to adjust the molecular weight of the polyolefin (30) produced.

(41) It is apparent to the skilled person that recycle streams may be present in the reaction system of the invention, for example there may be a recycle stream that is vented back from a polymer discharge system to the fluid bed reactor aiming at an efficient discharge of the product while at the same time recycling a large portion of unreacted gasses back to the reactor.

(42) In another aspect, the invention relates to a reaction system for the continuous preparation of a polyolefin in a reactor from one or more -olefin monomers of which at least one is ethylene or propylene, wherein the reaction system comprises a reactor (8), wherein the reactor comprises a fluidized bed (80), an expanded section (4) located at or near the top of the reactor, a distribution plate (6) located at the lower part of the reactor and an inlet for a recycle stream (10) located under the distribution plate

(43) wherein the system is arranged such that a polymerization catalyst is fed to the fluidized bed in the area above the distribution plate the one or more -olefin monomers are fed to the reactor the polyolefin is withdrawn from the reactor fluids are circulated from the top of the reactor to the bottom of the reactor, wherein the circulating fluids are compressed using a compressor (400) and subsequently cooled using a heat exchanger (5) to form a cooled recycle stream comprising liquid, and wherein the cooled recycle stream is introduced into the reactor using the inlet for the recycle stream

(44) and a part of the cooled recycle stream (b1) is drawn to form a liquid comprising stream, wherein the liquid comprising stream is introduced into the expanded section during at least part of the polymerization process,

(45) and wherein the liquid comprising stream is brought into contact with at least part of the interior surface of the expanded section.

(46) It is further noted that the term comprising does not exclude the presence of other elements. However, it is also to be understood that a description on a product comprising certain components also discloses a product consisting of these components. Similarly, it is also to be understood that a description on a process comprising certain steps also discloses a process consisting of these steps.

(47) The invention will now be elucidated by way of the following examples, without however being limited thereto.

EXAMPLES

(48) The following examples show that controlling the catalyst activity in the expanded section (by addition of a thermal run away reducing agent) leads to a more homogenous product.

(49) The inventors believe that the reason for obtaining a more homogeneous product is that the temperature is controlled in the expanded section due to the control of the catalyst activity.

(50) This invention concerns the cooling of the expanded section (by way of taking part of the cooled recycle stream and adding this to the expanded section). This is expected to also control the temperature in the expanded section. Consequently, the inventors believe that by drawing a liquid comprising stream from part of the cooled recycle stream, introducing the liquid comprising stream into the expanded section during at least part of the polymerization process, and bringing the liquid comprising stream into contact with at least part of the interior surface of the expanded section will therefore also increase the homogeneity of the polyolefin obtained.

(51) In addition, due to the contact of the liquid comprising stream with at least part of the interior surface of the expanded section, sheet formation in the expanded section will also be reduced as the liquid will physically wash the expanded section.

(52) In a special embodiment of the invention, the temperature reducing effect in the expanded section of both the liquid comprising stream and the TRRA is combined by adding the TRRA to the reaction system.

(53) The effect of the TRRA is demonstrated by the following examples;

(54) These examples show that TRRAs, for example esters, amines, nitriles, amides and mixtures thereof, in particular paraethoxyethylbenzoate (PEEB), trimethylacetonitrile (TA) and n,n-dimethyl benzamide (DB) are capable of reducing the catalyst activity (and hence controlling the temperature in the expanded section) without however affecting the properties of the produced polyolefin.

Example 1. TRRA is PEEB

(55) The polymerization catalyst was prepared as follows:

Example 1: Preparation of a Procatalyst on an Activated Butyl-Grignard Support

(56) Preparation of Grignard Reagent (Step o))Phase A

(57) This step o) constitutes the first part of phase A of the process for preparation of the procatalyst.

(58) A stirred flask, fitted with a reflux condenser and a funnel, was filled with magnesium powder (24.3 g). The flask was brought under nitrogen. The magnesium was heated at 80 C. for 1 hour, after which dibutyl ether (150 ml), iodine (0.03 g) and n-chlorobutane (4 ml) were successively added. After the colour of the iodine had disappeared, the temperature was raised to 80 C. and a mixture of n-chlorobutane (110 ml) and dibutyl ether (750 ml) was slowly added for 2.5 hours. The reaction mixture was stirred for another 3 hours at 80 C. Then the stirring and heating were stopped and the small amount of solid material was allowed to settle for 24 hours. By decanting the colourless solution above the precipitate, a solution of butylmagnesiumchloride with a concentration of 1.0 mol Mg/l was obtained.

(59) Preparation of Solid Magnesium Compound (Step i))Phase A

(60) This step i) constitutes the second part of phase A of the process for preparation of the procatalyst.

(61) This step is carried out as described in Example XX of EP 1 222 214 B1, except that the dosing temperature of the reactor is 35 C., the dosing time is 360 min and the propeller stirrer w is as used. An amount of 250 ml of dibutyl ether is introduced to a 1 liter reactor. The reactor is fitted by propeller stirrer and two baffles. The reactor is thermostated at 35 C.

(62) The solution of reaction product of step A (360 ml, 0.468 mol Mg) and 180 ml of a solution of tetraethoxysilane (TES) in dibutyl ether (DBE), (55 ml of TES and 125 ml of DBE), are cooled to 10 C., and then are dosed simultaneously to a mixing device of 0.45 ml volume supplied with a stirrer and jacket. Dosing time is 360 min. Thereafter the premixed reaction product A and the TES-solution are introduced to a reactor. The mixing device (minimixer) is cooled to 10 C. by means of cold water circulating in the minimixer's jacket. The stirring speed in the minimixer is 1000 rpm. The stirring speed in reactor is 350 rpm at the beginning of dosing and is gradually increased up to 600 rpm at the end of dosing stage.

(63) On the dosing completion the reaction mixture is heated up to 60 C. and kept at this temperature for 1 hour. Then the stirring is stopped and the solid substance is allowed to settle. The supernatant is removed by decanting. The solid substance is washed three times using 500 ml of heptane. As a result, a pale yellow solid substance, reaction product B (the solid first intermediate reaction product; the support), is obtained, suspended in 200 ml of heptane. The average particle size of support is 22 m and span value (d.sub.90d.sub.10)/d.sub.50=0.5.

(64) Activation of First Intermediate Reaction Product (Step ii))Phase B

(65) This step ii) constitutes phase B of the process for preparation of the procatalyst as discussed above.

(66) Support activation was carried out as described in Example IV of WO/2007/134851 to obtain the second intermediate reaction product.

(67) In inert nitrogen atmosphere at 20 C. a 250 ml glass flask equipped with a mechanical agitator is filled with slurry of 5 g of reaction product B dispersed in 60 ml of heptane. Subsequently a solution of 0.22 ml ethanol (EtOH/Mg=0.1) in 20 ml heptane is dosed under stirring during 1 hour. After keeping the reaction mixture at 20 C. for 30 minutes, a solution of 0.79 ml titanium tetraethoxide (TET/Mg=0.1) in 20 ml of heptane was added for 1 hour.

(68) The slurry was slowly allowed to warm up to 30 C. for 90 min and kept at that temperature for another 2 hours. Finally the supernatant liquid is decanted from the solid reaction product (the second intermediate reaction product; activated support) which was washed once with 90 ml of heptane at 30 C.

(69) The activated support, according to chemical analysis, comprises a magnesium content of 17.3 wt. %, a titanium content of 2.85 wt. %, and a chloride content of 27.1 wt. % corresponding to a molar ratio of Cl/Mg of 1.07 and Ti/Mg of 0.084.

(70) C. Preparation of the Procatalyst

(71) A reactor was brought under nitrogen and 125 ml of titanium tetrachloride was added to it. The reactor was heated to 90 C. and a suspension, containing about 5.5 g of the support obtained in step C in 15 ml of heptane, was added to it under stirring. The reaction mixture was kept at 90 C. for 10 min. Then ethyl benzoate was added (EB/Mg=0.15 molar ratio). The reaction mixture was kept for 60 min. Then the stirring was stopped and the solid substance was allowed to settle. The supernatant was removed by decanting, after which the solid product was washed with chlorobenzene (125 ml) at 90 C. for 20 min. The washing solution was removed by decanting, after which a mixture of titanium tetrachloride (62.5 ml) and chlorobenzene (62.5 ml) was added. The reaction mixture was kept at 90 C. for 30 min. After which the stirring was stopped and the solid substance was allowed to settle. The supernatant was removed by decanting, after which a mixture of titanium tetrachloride (62.5 ml) and chlorobenzene (62.5 ml) was added. Then di-n-butyl phthalate (DBP) (DBP/Mg=0.15 molar ratio) in 3 ml of chlorobenzene was added to reactor and the temperature of reaction mixture was increased to 115 C. The reaction mixture was kept at 115 C. for 30 min. After which the stirring was stopped and the solid substance was allowed to settle. The supernatant was removed by decanting, after which a mixture of titanium tetrachloride (62.5 ml) and chlorobenzene (62.5 ml) was added. The reaction mixture was kept at 115 C. for 30 min, after which the solid substance was allowed to settle. The supernatant was removed by decanting and the solid was washed five times using 150 ml of heptane at 600 C., after which the procatalyst III, suspended in heptane, was obtained.

(72) The Polymerization was Conducted as Follows:

(73) Propylene polymerization experiments (Table 1) were performed using procatalysts I, II and III described above. Triethylaluminium (TEAL) was used as co-catalyst, and cyclohexylmethyldimethoxysilane (C-donor) or n-propyltrimethoxysilane (N-donor) was used as external donor (Si). Experiments were performed at different H2/C3 molar ratios.

(74) The polymerization of propylene was carried out in a stainless steel gas phase reactor with a volume of 1800 mL. Under a nitrogen atmosphere, the co-catalyst (TEAL) and procatalyst synthesized according to the procedure described above and the external electron donor were dosed to the reactor as heptane solutions or slurries. 10-15 mg (2% wt Ti) of procatalyst were employed. The molar ratio of co-catalyst TEAL to titanium (from the procatalyst) was set to 130, and the Si/Ti ratio was set to 8. During this dosing, the reactor temperature was maintained below 30 C. Subsequently, the reactor was pressurized using a set ratio of propylene and hydrogen, and the temperature and pressure were raised to its setpoint (67 or 82 C. and 20 barg). After the pressure setpoint has been reached, the polymerization was continued for 60 minutes. During the polymerization reaction the gas cap composition of propylene and hydrogen was controlled using mass flow meters and online-GC control. After reaching the polymerization time the reactor was depressurized and cooled to ambient conditions. The propylene polymer so obtained was removed from the reactor and stored in aluminium bags.

(75) The polymerization conditions are summarized in Table 1 below. In these experiments paraethoxyethylbenzoate (PEEB) was used as TRRA. The temperature of 67 C. reflects the temperature of a polymerization within the fluidized bed; the temperature of 82 C. reflects the temperature within the expanded section.

(76) TABLE-US-00001 TABLE 1 Experimental polymerization conditions. Parameter At T = 67 C. At T = 82 C. PC3 (barg) 26 28 H2 (mol %) 0.37 0.37 TEAL/Ti (mol/mol) 130 130 Si/Ti (mol/mol) 8 8 PC3 pressure of propylene H2 hydrogen TEAL tri-ethylaluminium

(77) The effect of different TRRA/Si ratios on the yield is shown in Table 2 below:

(78) TABLE-US-00002 TABLE 2 yield versus TRRA/Si molar ratio at 67 and 82 C. Molar ratio Yield at 67 C. Yield at 82 C. TRRA/Si (Kg-PP/g-Cat) (Kg-PP/g-Cat) TRRA/Si = 0 24.2 20.0 TRRA/Si = 0.68 21.6 16.8 TRRA/Si = 0.98 17.2 TRRA/Si = 1.12 15.4 TRRA/Si = 1.25 12.8 TRRA/Si = 2 12.5 TRRA/Si = 4 17.9 5.9 TRRA/Si = 4.5 19.9 7.6

(79) The polypropylene produced with TRRA/Si ratio 0; 4 and 4.5 and polymerization temperatures of 67 C. or 82 C. was characterized in terms of its molecular weight distribution (MWD) and crystallinity.

(80) Crystallinity was determined using analytical temperature rising elution fractionation (ATREF) analysis was conducted according to the method described in U.S. Pat. No. 4,798,081 and Wilde, L.; Ryle, T. R.; Knobeloch, D. C; Peat, L R.; Determination of Branching Distributions in Polyethylene and Ethylene Copolymers, J. Polym. ScL, 20, 441-455 (1982), which are incorporated by reference herein in their entirety. The composition to be analyzed was dissolved in 1,2-dichlorobenzene as solvent of analytical quality filtrated via 0.2 m filter and allowed to crystallize in a column containing an inert support (Column filled with 150 m stainless steel beans (volume 2500 L) by slowly reducing the temperature to 20 C. at a cooling rate of 0.1 C./min. 1 g/L Irgafos and BHT were used as stabilizers The column was equipped with an infrared detector. An ATREF chromatogram curve was then generated by eluting the crystallized polymer sample from the column by slowly increasing the temperature of the eluting solvent (1,2-dichlorobenzene) from 20 to 130 C. ata rate of 1 C./min.

(81) The instrument used was Polymer Char Crystaf-TREF 300.

(82) Stabilizers: 1 g/L Topanol+1 g/L Irgafos 168

(83) Sample: approx. 70 mg in 20 mL

(84) Sample volume: 0.3 mL

(85) Pump flow: 0.50 mL/min

(86) The software from the Polymer Char Crystaf-TREF-300 was used to generate the spectra.

(87) molecular weight distribution was determined using IAV Molecular Characterization method. The chromatography equipment used is Polymer Laboratories PL-GPC220 with Viscotek 220R viscometer and Refractive index detector. The column set consists of three columns of Polymer Laboratories 13 m PLgel Olexis, 3007.5 mm. Standard linear polyethylene was used for calibration and reference.

(88) The results of the aTREF analysis are represented in FIG. 5.

(89) Profile A in FIG. 5 shows the aTREF temperature profile of a polypropylene produced without a TRRA at T=67 C.

(90) Profile B in FIG. 5 also shows the aTREF temperature profile of a polypropylene produced without a TRRA at T=82 C.

(91) Profile C in FIG. 5 shows the aTREF temperature profile of a polypropylene produced with TRRA/Si ratio of 4 at T=67 C.

(92) As can be seen in FIG. 5 the addition of a TRRA does not significantly affect the crystallinity of the polymer produced at normal operating temperature (67 C.).

(93) The crystallinity of these polypropylenes shows an almost similar peak temperature with a slight shift of the polymer sample produced at 82 C. towards lower crystallinity.

(94) The results of the molecular weight distribution is given in Table 3 below:

(95) TABLE-US-00003 TABLE 3 Summary of the molecular weight averages and their corresponding viscosities. Polymerization Mn Mw Mz TRRA/Si temperature (10.sup.3 (10.sup.3 (10.sup.3 [] Run (mol/mol) ( C.) g/mol) g/mol) g/mol) Mw/Mn Mz/Mw (dL/g) 1 0 67 69 460 1500 6.7 3.2 2.10 2 0 82 68 340 1000 5.1 2.9 1.73 3 4 67 89 550 1700 6.2 3.1 2.42 4 4 82 82 460 1600 5.7 3.5 2.12 5 4.5 82 90 470 1300 5.2 2.8 2.17 6 4.5 67 95 560 1700 6.0 3.1 2.46

(96) The results of Table 3 show that a molecular weight distribution (Mw/Mn, MWD) closer to the desired MWD (run #1, MWD=6.7) is obtained when using a TRRA/Si in a preferred ratio of 4 mol/mol (runs #3 and 4) as compared to not using a TRRA at 82 C. (run #2). Moreover, since the activity of the catalyst is more reduced by the TRRA at a higher temperature, the contribution of the MWD of the polymer produced at 82 C. is less, thereby making an overall product that has a MWD closer to the desired MWD of 6.7 (combination of a bit of 5.7 of run #4 and 6.2 of run #2) than in the situation where a TRRA is not used; in the latter case the contribution of the MWD of 5.1 at 82 C. is more pronounced.

Example 2

(97) With C-donor is meant: cyclohexylmethyldimethoxysilane.

(98) With N-donor is meant: n-propyltrimethoxysilane

(99) The polymerization catalyst was prepared as described in Example-1. Similar procedures and equipment used in Example-1 were used in this example for the molecular weight distribution (MWD) using GPC and crystallinity using ATREF.

(100) The effect of different TRRA/Si ratios for both TA and DB on the yield are shown in Table 4 below:

(101) TABLE-US-00004 TABLE 4 Yield versus TRRA/Si molar ratio for TA and DB at 67 and 82 C. Polymerization TRRA/Si Temp. Yield TRRA (mol/mol) ( C.) (g-PP/mg-Cat .Math. hr) TA 0 67 22.6 TA 7 67 12.8 TA 7 82 7.47 DB 0 67 20.8 DB 4 67 17.6 DB 4 82 8.67

(102) Table 4 shows the yield with different molar ratio of TRRA to electron donor (TRRA/Si) for TA. The addition of TA shows almost similar deactivation degree at both temperatures where the drop in activity was 43% and 67% at 67 and 82 C., respectively, when bench marked with zero TA at 67 C. Therefore, TA is preferably used as one of the components in a mixture with another TRRA. Table 4 also shows the yield with different molar ratio of TRRA to electron donor (TRRA/Si) for DB. The TRRA/Si ratio of 4 is for DB, a ratio that satisfies the functionality of reducing catalyst activity by around 60% at elevated temperature, 82 C.; and maintaining the catalyst activity to be not lower than 15% at normal polymerization temperature, 67 C.

(103) A in FIG. 6 is the TREF of a polymer produced without TRRA at 67 C. B in FIG. 6 is the TREF of a polymer produced using TA as TRRA with a TRRA/Si molar ratio of 7 at a production temperature of 67 C. C in FIG. 6 is the TREF of a polymer produced using TA as TRRA with a TRRA/Si molar ratio of 4 at a production temperature of 82 C.

(104) As can be seen from FIG. 6, the crystallinity of the polymer remains similar with or without using a TRRA (as exemplified by the use of TA).

(105) Table 5 summarizes molecular weight averages and their corresponding viscosities.

(106) TABLE-US-00005 TABLE 5 Molecular structure parameters obtained from SEC-IR using convention calibration. Pol. M.sub.n M.sub.w M.sub.z Temp. (10.sup.3 (10.sup.3 (10.sup.3 [] Si TRRA TRRA/Si ( C.) g/mol) g/mol) g/mol) M.sub.w/M.sub.n M.sub.z/M.sub.w (dL/g) C 0 82 68 340 1000 5.1 2.9 1.73 C 0 67 85 450 1300 5.4 2.8 2.12 C TA 7 82 81 370 940 4.6 2.5 1.85 C TA 7 67 80 460 1300 5.7 2.8 2.14 C DB 4 67 84 480 1400 5.7 3.0 2.19 N 0 67 72 400 1100 5.5 2.8 1.91 N PEEB 4 67 87 510 1500 5.9 3.0 2.30

(107) Table 5 shows that the use of a TRRA does not significantly affect the polymer properties. It also shows that the functionality of the TRRA is not dependent on the type of external donor used.

(108) Table 6 shows the yield of the polymer versus the TRRA/Si molar ratio for PEEB with different silane donors at 67 C.

(109) TABLE-US-00006 TABLE 1 Yield versus TRRA/Si molar ratio for PEEB with different Si at 67 C. TRRA/Si Yield @ 67 C. Si (mol/mol) (kg-PP/g-cat) C-donor 0 24.2 C-donor 4 17.9 N-donor 0 14.5 N-donor 4 8.1

(110) In order to visualize the effect of the type of external donor used (Si), the PEEB as TRRA was tested when Si is N-donor (n-propyltrimethoxysilane) and C-donor. Table 1 shows that the TRRA, represented here by PEEB, functioning in similar pattern when the Si changed from C-donor to N-donor. The productivity drop when TRRA added at similar TRRA/Si ratio is the same with C and N-donors at both polymerization temperatures regardless of the effect of Si type on catalyst productivity. Therefore, this example shows that regardless of the type of external donor used, the TRRA is effective in reducing catalyst productivity.

(111) A TREF was performed on the polypropylene produced using N-donor and a PEEB/Si ratio of 4 and compared to the TREF of a polypropylene produced using the same C-donor but without using a TRRA. Also, in the case of N-donor and PEEB, it was seen that the crystallinity of the polymer remains similar with or without using a TRRA.

(112) Conclusion

(113) A TRRA is capable of reducing the catalyst activity at a temperature of 82 C. (which is the maximum desired temperature inside the expanded zone in a commercial polypropylene plant) or higher while maintaining very good overall catalyst activity at a normal polymerization temperature of e.g. in the range of 65 to 72 C. (representing the temperature of a fluidized bed in a commercial polypropylene plant). This allows a better control of the temperature in the expanded zone and consequently the production process will be more stable.

(114) Therefore, the process of the invention, wherein a stream comprising a thermal run away reducing agent (TRRA-containing stream) is introduced into the expanded section during at least part of the polymerization process will provide a more uniform temperature profile across the fluidized bed, leading to a stable process, which in addition will lead to the production of more uniform polyolefins.

(115) In addition, by bringing the TRRA-containing stream is brought into contact with at least part of the interior surface of the expanded section, fouling/sheeting will be significantly decreased if not eliminated.

(116) In a special embodiment of the invention, the composition of the recycle stream (from which part is drawn to form the liquid comprising stream which cools the expanded section) is improved by the addition of an inert alkane.

(117) As can be seen from the below examples, the use of an insert alkane, preferably n-butane or i-butane, more preferably i-butane has a positive effect on the production rate.

Example 3

(118) A computer-based mathematical model capable of generating mass and heat balances along a fluidized bed reactor was used to run simulation in condensed mode operation to show the advantage of adding an alkane having 3 to 6 carbon atoms in the process of the invention. Firstly, the model was run using actual data from commercial polypropylene production to validate the model. The results are shown in Table 7.

(119) TABLE-US-00007 TABLE 7 Commercial data validated versus the computer-based mathematical model Example Number Reactor Conditions a (Commercial) b (Model Results) Internal Reactor diameter (m) 4.8 4.8 Recycle Gas Superficial 0.3081 0.3081 Velocity (m/s) Recycle Gas Composition (mole fraction): Propylene 0.83893 0.83893 Propane 0.09398 0.09398 Hydrogen 0.00267 0.00267 Nitrogen 0.06443 0.06443 n-Butane Iso-Butane Recycle Gas Density (kg/m.sup.3.) 67.25 72.31 Reactor Temperature ( C.) 68.08 68.08 Reactor Inlet Temperature ( C.) 65.70 65.70 Reactor Pressure (kPag) 3214.85 3214.85 Reactor inlet Pressure (kPag) 3334.53 3334.53 Inlet Dew Point Temperature ( C.) 67.19 68.56 Condensed Liquid in Recycle 39.78 36.94 Stream (% weight) Production Rate (ton/h) 50.6 51.2

(120) As can be seen from Table 4, the actual data and the data from the model are very well comparable.

(121) Subsequently, the model was ran representing commercial data when an alkane having 3 to 6 carbon atoms was not added (propane present in the recycle stream is produced by hydrogenation of propylene in the reactor in case hydrogen is present and/or comes from the feed-stream comprising propylene) versus when different amounts of n-butane, iso-butane or cyclopropane were added to the recycle stream. The n-butane, i-butane or cyclopropane were added using feed (70)

(122) The results are shown in Table 8 below. Table 8 shows the effect of a molar composition in the reactor of 0% additional alkane having 3 to 6 carbon atoms versus 2.0 mole % n-butane (I), 2.3 mole % n-butane (II), 2.65 mol % i-butane (III) and 2.75 mole % i-butane (IV) and 3.89 mole % cyclopropane (V).

(123) As can be seen from Table 8, the feeding of an alkane having 3-6 carbon atoms increases the production rate from 51.2 tons/hour to for 2.0 mol % n-butane to 52.9 tons/hour, for 2.3 mol % n-butane to 54.9 tons/hour; for 2.65 mol % i-butane to 69.4 tons/hour, for 2.75 mole % i-butane to 71 tons/hour and for 3.89 mol % cyclopropane to 73.1 tons/hour.

(124) Therefore, it has been shown that in the process of the invention, it is advantageous to add an alkane having 3-6 carbon atoms, preferably an alkane having 3 to 6 carbon atoms chosen from the group consisting of i-butane, n-butane, propane and mixtures thereof, more preferably chosen from the group consisting of i-butane, more preferably liquid i-butane is added to the reactor such that the molar composition of the alkane having 3 to 6 carbon atoms in the reactor is at least 1 mol %, preferably at least 2 mol %, more preferably at least 2.4 mol % and/or at most 10 mol %, preferably at most 5 mol %.

(125) TABLE-US-00008 TABLE 8 Reactor Conditions b I II III IV V n-butane (mol %) 2.0 2.3 i-butane (mol %) 2.65 2.75 cyclopropane 3.89 (mol %) Internal Reactor 4.8 4.8 4.8 4.8 4.8 4.8 diameter (m) Recycle Gas 0.3081 0.3081 0.3081 0.3081 0.3081 0.3081 Superficial Velocity (m/s) Recycle Gas Composition (mole fraction): Propylene 0.83893 0.80893 0.80593 0.80593 0.80593 0.80000 Propane 0.09398 0.09398 0.09398 0.09398 0.09398 0.09398 Hydrogen 0.00267 0.00267 0.00267 0.00267 0.00267 0.00267 Nitrogen 0.06443 0.07443 0.07443 0.07093 0.06993 0.06443 n-Butane 0.02000 0.02300 Iso-Butane 0.0265 0.0275 cyclopropane 0.03893 Recycle Gas 72.31 72.17 72.41 73.12 73.38 73.17 Density (kg/m.sup.3.) Reactor 68.08 68.08 68.08 68.08 68.08 68.08 Temperature ( C.) Reactor Inlet 0.03893 0.03893 0.03893 0.03893 0.03893 65.70 Temperature ( C.) Reactor Pressure 3214.85 3214.85 3214.85 3214.85 3214.85 3214.85 (kPag) Reactor inlet 3334.53 3334.53 3334.53 3334.53 3334.53 3334.53 Pressure (kPag) Inlet Dew Point 68.56 69.44 69.73 71.03 71.03 70.69 Temperature ( C.) Condensed Liquid 36.94 38.6 40.18 51.10 52.25 53.40 in Recycle Stream (% weight) Production Rate 51.2 52.9 54.9 69.4 71.0 73.1 (ton/h)