Gas Phase Ethylene Polymerization-Production Processes, Polymer Properties, And Reactor Modeling

Znd. Eng. C h e m . Res . 1994,33, 449-479

REVIEWS

449

Gas Phase Ethylene Polymerization: Production Processes, Polymer Properties, and Reactor Modeling

Tuyu Xie, Kim B. McAuley,' James C. C. HSU, and David W. Bacon Department of Chemical Engineering, Queen's University, Kingston, Ontario, Canada K7L 3N6

A review of relevant macroscopic and microscopic processes of gas phase ethylene polymerization, both chemical and physical, is given. The commercial technology development of gas-phase ethylene polymerization processes is illustrated through a selective survey of the patent literature. Both advantages and disadvantages of gas phase polymerization processes are addressed, and the challenges of laboratory studies of gas phase polymerization are also outlined. Physicochemical phenomena of ethylene polymerization using heterogeneous catalysts are discussed, including examination of catalyst preparation, polymer morphological development, and elementary chemical reactions. Metallocene-based catalysts and their kinetic performance for olefin polymerizations are also discussed. The current state of the ar t for reactor modeling of polymerization rate, molecular weight development, reactor dynamics, and resin grade transition strategies is illustrated on the basis of the most recent academic studies. Finally, relationships between resin properties and polymer microstructures as well as characterization methods are described briefly. In particular, temperature-rising elution fractionation technology is emphasized for characterization of ethylene copolymers. The fundamental issues involved in gas phase ethylene polymerization and the& interrelationships are also discussed in some detail.

Contents 1. Introduction 2. Gas Phase Polymerization Processes

2.1. Commercial Gas Phase Polymerization Processes

2.2. Experimental Methods 3. Physicochemical Phenomena

3.1. Catalysts for Gas Phase Polymerization

3.2. Polymer Particle Morphology Developments

3.3. Chemical Reactions

4.1. Kinetic Modeling 4.2. Dynamic Process Modeling and

4. Reactor Modeling

Control 5. Polymer Properties and

Characterization

Properties 5.1. Physical and Mechanical

5.2. Polyethylene Characterization 6. Summary 7. References

449 450 450

455 456 456

460

461 463 463 467

469

469

470 474 474

1. Introduction Polyethylene (PE) is the largest synthetic commodity

polymer in terms of annual production and is widely used throughout the world due to its versatile physical and chemical properties. The American Society for Testing and Materials (ASTM) has classified PE into four groups: I (low density) at 0.910-0.925 g/cm3, I1 (also low density) at 0.926-0.940 g/cm3, I11 (high-density copolymers) at

0SSS-5SS5/94/2633-0449$04.50/0

0.941-0.959 g/cm3, and IV (high-density homopolymer) at 0.960 g/cm3 and above (Redman, 1991). In the literature, however, PE is normally classified as low density (LDPE) at 0.910-0.930 g/cm3 and high density (HDPE) a t 0.931- 0.970 g/cm3. Low-density polyethylene is further classified as low-density polyethylene (LDPE) and linear low-density polyethylene (LLDPE) based on polymer chain microstructure and synthesis processes. According to the figures reported in Mod. Plast. (19931, polyethylene production in the United States alone was over 10 million tons in 1992. The annual production of PE in Europe is about 9 million tons (Redman, 1991). The current annual worldwide capacity for PE production is over 30 million tons. Figure 1 shows the US. PE production profile over the past-decade. Although the annual rate of increase slowed down slightly at the end of the 1980s, the average annual increase rate is about 8 94 for HDPE and about 5 % for LDPE and LLDPE for the past decade. Consumption of PE is still rising through the 1990s with development of synthesis and processing technology. The main markets and applications of LLDPE and HDPE are summarized in Figure 2 after James (1986) and Foster (1991).

Polyethylene is commercially produced exclusively by continuous processes. On the basis of polymerization mechanisms and reactor operating conditions, PE production processes can be classified into at least five process categories as shown in Table 1. Among them, the gas phase polymerization process is the most recently developed and also the most versatile. Since its emergence, this process has been challenging other existing processes for market share, particularly, for production of LLDPE, due to its economic and technological advantages. Many excellent reviews of ethylene polymerization processes have been published (Vandenberg and Perka, 1977; Short, 1981; Choi and Ray, 1985a; Nowlin, 1985; James, 1986; Beach and Kissin, 1986). However, the fundamental issues involved

0 1994 American Chemical Society

450 Ind. Eng. Chem. Res., Vol. 33, No. 3, 1994

0.97

0.96

- 0.95 rn

0.94 v) z W 0 0.93

0.92

0.91

'Ooo L

-

-

-

-

-

-

-

z 6000 I- e o

LDPE & LLDPE

f LDPE & LLDPE z 6000 - e

HDPE

vj

5 2ooo y HDPE

.""" 80 82 84 86 88 90 92 94 96

TIME, year

Figure 1. U S . polyethylene production profile (data from Mod. Plast. (198C-1993)).

I

SHEET

0.90 1 I 0.01 0.10 1 .oo 10.0 100

MELT INDEX, 9/10 min

Figure 2. Markets and applications for low-pressure polyethylene. in gas phase polymerization processes have not been comprehensively discussed. With development of commercial gas phase polymerization processes, the importance of more complete understanding of gas phase polymerization processes has been recognized by the academic community. Significant experimental and reactor modeling work has been carried out in recent years. In the present review, the authors focus on the main issues involved in gas phase ethylene polymerization processes, including commercial production technology development, physicochemical phenomena, kinetictdynamic reactor modeling and control; polymer properties and characterization, as well as their interrelationships.

2. Gas Phase Polymerization Processes The distinguishing characteristic of gas phase polym-

erization is that the system does not involve any liquid phase in the polymerization zone. Polymerization does occur a t the interface between the solid catalyst and the polymer matrix, which is swollen with monomers during polymerization. The gas phase plays a role in the supply of monomers, mixing of polymer particles, and removal of reaction heat. Hence, gas phase polymerization is also called dry polymerization in some patents (Dormenval et al., 1975; Havas and Mangin, 1976). In this section, commercial gas phase polymerization processes and experimental studies of gas phase polymerization are discussed.

2.1. Commercial Gas Phase Polymerization Pro- cesses. The invention of gas phase ethylene polymeri-

zation can be traced back to the 1950s. Dye (1962), whose patent was filed in 1957, was perhaps the first to adopt a fluidized bed reactor for gas phase olefin polymerization. The original reactor consisted of three concentric super- imposed vertical sections. Polymer particles were discharged through an extruder, which was connected to the bottom section of the reactor. The reactor was operated at a pressure of 30 atm and a temperature of 100 "C for ethylene polymerization. Goins (1960) carried out ethylene copolymerization in a countercurrent fluidized bed reactor in the presence of inert diluent gas. In this process, polymer particles are passed downward in the reactor and monomer mixed with diluent gas is passed countercurrently upward in the reactor and monomer mixed with diluent gas is passed countercurrently upward through a series of vertical fluidized bed reaction zones. The reaction zones can be controlled independently by taking off-gas from the last reaction zone, cooling it, and recycling portions of such off-gas to each of the reaction zones. Both patents (Dye, 1962; Goins, 1960) were assigned to Phillips Petroleum Company. Schmid et al. (1967) carried out ethylene polymerization in a stirred fluidized bed reactor. In this configuration, polymer particles are moved in the direction of monomer flow by stirring, and reaction heat is removed by cooling the walls of the reactor, by the gas stream, and by introduction of liquified monomers. The patent of Schmid et al. was assigned to BASF. The benefits of gas phase polymerization were recognized by these pioneer inventors. Although these inventions have not been directly applied in commercial gas phase polymerization processes, the fundamental ideas demonstrated by these inventors provided the foundation for the later commercial gas phase process development.

The first commercial gas phase polymerization plant using a fluidized bed reactor was constructed by Union Carbide in 1968 at Seadrift, TX (Rasmussen, 1972; Batleman, 1975; Burdett, 1988). This process was developed initially for HDPE production. The success of this novel technology led to the extension of the process to LLDPE, which was produced initially on a commercial reactor in 1975 (Davis, 1978; Burdett, 1988). The Union Carbide gas phase process, commonly called UNIPOL, has been licensed worldwide with more than 25 licenses operating in 14 different countries (Burdett, 1988). Pro- duction of LLDPE using gas phase processes is more difficult than production of HDPE because the difference between the melting point and polymerization temperature is much narrower for LLDPE. The catalyst types and equipment design developed for HDPE cannot be used to produce LLDPE because of the potential for agglomeration of polymer particles. Hence, significant engineering and chemistry research has been required to assure the success of gas phase LLDPE production. According to Karol (19831, the keys to the success of the UNIPOL technology for LLDPE production are the proprietary catalysts that operate at low pressure and low temperature and which are suitable for use in a gas phase fluidized bed reactor. Union Carbide Corporation received the 1979 Kirkpatrick Chemical Engineering Achievement Award in recognition of the innovation of the UNIPOL process (Chem. Eng., 1979).

The UNIPOL process has been described in several US. Patents (Miller, 1977; Levine and Karol, 1977; Karol and Wu, 1978; Wagner et al., 1981; Jorgensen et al., 1982). A simplified flow diagram is shown in Figure 3. The fluidized bed reactor consists of a reaction zone and a disengagement zone. The reaction zone has a height to diameter ratio of about 6-7.5. The disengagement zone has a diameter to

Ind. Eng. Chem. Res., Vol. 33, No. 3, 1994 451 Table 1. Polymerization Processes and Reactor Operating Conditions

conventional high-press. solution gas phase

reactor type tubular or autoclave autoclave CSTR loop or CSTR fluidized or

reactor press., atm 1200-3000 600-800 - 100 30-35 30-35 temp, "C 130-350 200-300 140-200 85-110 80-100 polymn mech free radical coordination coordination coordination coordination loci of polymn monomer phase monomer phase solvent solid solid density, g/cm3 0.910-0.930 0.910-0.955 melt index, g/10 min 0.10-100 0.80-100 0.50-105


w

CATALYST FEEDER

INERT GAS

GAS FEED

a COMPRESSOR l h ACTIVATOR

FEEDER

Figure 4. Gas-phase ethylene polymerization process with internal heat exchanger (Brown et al., 1981).

this process. Compared with the process shown in Figure 3, this modified process has the following advantages:

1. The formation of sheets on the sloped walls of the transition zone is eliminated so that the frequency of reactor stoppages is reduced.

2. The depth of the bed in the polymerization zone can be varied over a wide range allowing a greatly increased rate of reactor output with good operation.

3. A smaller initial charge of powdered material is required to start up successfully.

4. The fabrication cost of a fluidized bed reactor is reduced.

5. Gas velocity can be much lower using an internal heat exchanger, thus reducing operation energy.

6. Since heat removal is independent of gas mass flow rate, the reactor pressure can also be decreased to a limit defined by the polymerization kinetics.

7. The internal heat exchanger also acts as a sets of baffles, and tends to deter the migration of bubbles to the center of the bed and to increase mixing near the reactor walls.

8. Recycle gas is at the same temperature as the entire fluidized bed.

Despite these advantages, there are no published reports on the application and performance of this type reactor.

About 1964 BASF first developed a gas phase ethylene polymerization process using Phillips catalysts (Wisseroth, 1969; Sailors and Hogan, 1981), and a commercial gas phase HDPE plant was built in Germany in 1976 (Choi and Ray, 1985a). BASF uses a continuous stirred-bed reactor for gas phase ethylene polymerization as shown in Figure 5 (Trieschmann et al., 1977; Choi and Ray, 1985a). The reactor is equipped with an anchor agitator and is operated at higher pressure and temperature (35 atm, 100-110 OC) than those employed in UNIPOL. Polymer is discharged through a cyclone. Unreacted monomer and the heat- transfer agent, such as cyclohexane, are removed through the top of the reactor and liquefied by a condenser. The liquefied recycle monomer and cyclohexane are fed into the reactor through a valve located at the bottom of the reactor, which is controlled by a thermostat. Liquefied monomer and cyclohexane can be fed into the reactor by a metering pump or by provision of an adequate static pressure difference between condenser and reactor. Fresh monomer is also fed through the bottom of the reactor.

The temperature in the polymerization zone can be maintained at a constant level by controlling the rate of feed of liquefied monomer and cyclohexane. Liquefied

P RECYCLE

Figure 5. BASF gas phase ethylene polymerization process (Trieschmann et al., 1977; Choi and Ray, 1985a).

RECYCLE MONOMER-HYDROGEN-QUENCH VAPOR

CATALYG SPYY NOZZLES n

- h - -

- I! 2 Y MONOMER AND HYDROGEN INLETS POLYMER DISCHARGE

Figure 6. AMOCO gas phase polymerization process (Shepard et al., 1976; Jezl et al., 1976; Peters et al., 1976; Jezl and Peters, 1978a,b).

monomer and cyclohexane vaporize in the reaction zone and rise through the bed at a very slow velocity of about 2 cm/s so that the polymer particles are not fluidized. The bed is mixed by agitation with a mean circulation time of 30-100 s. Fresh monomer is added at a rate such that the pressure in the system remains constant. The pressure and temperature in the polymerization zone are maintained at levels such that the vaporized monomer and heat- transfer agent remain in the gaseous state. Magnesium- supported titanium halides and aluminum alkyl compounds and silica-supported modified chromium oxide catalysts are used for HDPE production.

The AMOCO (Standard Oil Company) gas phase polymerization process was developed in the mid-1970s (Shepard et al., 1976; Jezl et al., 1976; Peters et al., 1976; Jezl and Peters 1978a,b). A compartmented horizontal stirred bed reactor is used in this process as shown in Figure 6. The reactor is divided into several stirred polymerization zones which are separated by weirs. Make- up and recycle monomer and hydrogen are fed into each reaction zone along the bottom of the reactor. Catalyst and quench liquid are introduced into each polymerization zone through spray nozzles at the top of the reactor. Unreacted monomer, hydrogen, and quench vapor, normally isobutane or isopentane, are removed through the top of the reactor to the recycle and separation sections. The weirs extend upward to somewhat above the middle of the reactor so that the polymer bed fills about half the volume of the reactor. As the polymer powder exceeds the weir height, it falls into the adjacent reaction zone in the direction of the polymer discharge. The polymer

discharge vessel contains a postpolymerization zone, wherein a controllable amount of adiabatic polymerization takes place, producing heat which, together with externally added heat, melts the polymer to form easily transferable liquid polymer. Following treatment with water and suitable additives, the melt polymer is finally converted to pellets.

Each polymerization zone is equipped with flat paddles extending transversely from the shaft and making close clearance with the inside wall of reactor to ensure adequate bed mixing. The paddles are so constructed to minimize any foward or backward movement of the bed. The agitation speed is about 5-30 rpm to provide the desired heat and mass transfer, to avoid polymer particle being thrown up into the space above the bed, and to provide a slow and regular turnover of the entire polymer bed contained in the reactor. The distinguishing feature of the AMOCO process, compared with previous processes, is that it can be operated like a set of CSTRs in series. The residence time distribution (RTD) of this process almost approaches that of plug flow. Tracer studies on the commercial reactor indicate that the RTD in the AMOCO reactor is roughly equivalent to that of three to five well- stirred tanks in series. This unique feature of the horizontal reactor permits production of ethylene-propylene copolymers with high impact strength and good stiffness (Brockmeier, 1991). AMOCO and Chisso have been collaborating for commercialization of ethylene- propylene gas phase polymerization processes (Krieger, 1992).

Soon after the Union Carbide process was introduced, Naphtachimie developed a gas phase polymerization process using a fluidized bed reactor (Dormenval et al., 1975; Havas and Magin, 1976). A commercial-scale fluidized bed plant with an initial capacity of 25 000 t/y was commissioned in 1975 (Delgrange, 1987). The gas phase polymerization technology developed by Naph- tachimie was later transferred to BP Chemicals (Delgrange, 1987; Brockmeier, 1987). After the amalgamation of Napthachimie into BP Chemicals, high priority was given to further development of the novel gas phase polymerization process, as can be confirmed by recent patents assigned to BP Chemicals (Dumain and Raufast, 1985; Collomb-Ceccarini and Crouzet, 1986; Bailly and Speak- man, 1986; Raufast, 1987; Bailly and Collomb, 1988; Chinh and Dumain, 1990; Speakman, 1991; Matens and Moterol, 1992; Havas and Lalanne-Magne, 1992). The BP Chem- icals gas phase process now has a capacity of 170 kt/y and can produce both LLDPE and HDPE with a density range from 0.91 to 0.96 (Delgrange, 1987; Redman, 1991; Speak- man, 1991). The major distinction of the BP Chemicals process from the UNIPOL process is that prepolymer particles rather than catalyst particles are fed into the fluidized bed reactor. The BP Chemicals gas phase process is a continuous two-stage polymerization process, as shown in Figure 7 (Chinh and Dumain, 1990; Redman, 1991; Matens and Moterol, 1992; Havas and Lalanne-Magne, 1992). This process uses a combination of a stirred tank reactor and a fluidized bed reactor in series.

During conventional gas phase reactor operation, the catalyst and the cocatalyst may be brought into contact either prior to their introduction into the fluidized bed, or in the interior of the reactor. Whichever method is employed, the polymerization reaction always starts up very abruptly and attains a maximum rate soon after the catalyst system is introduced into the fluidized bed. I t is therefore in the initial phase of polymerization that the risks of hot spots forming and grains bursting into fine

Q 9 Ind. Eng. Chem. Res., Vol. 33, No. 3, 1994 463

MONOMER RECYCLE

U I

- P 4- ' COMPRESSOR MONOMER AND GAS FEED Figure 7. BP Chemicals two-stage gas phase ethylene polymerization process (Chinh and Dumain, 1990; Redman, 1991; Maters and Moterol, 1992; Havas and Lalanne-Magne, 1992).

particles are greatest. Hot spots may lead to formation of agglomerates and to settling of polymer inside the fluidized bed (Collomb-Ceccarini and Crouzet, 1986). Furthermore, during polymerization small variations in the feed rates of catalyst, monomer, and comonomer or in the withdrawal rate of polymer will also cause an unex- pected increase in the quantity of heat evolved by the polymerization. If the heat cannot be removed efficiently, these small variations can cause hot spots in the reaction bed and formation of agglomerates by melting polymer. Such variations can therefore make it difficult to obtain a polymer of consistent quality, in particular, of constant molecular weight and particle size (Chinh and Dumain, 1990; Matens and Moterol, 1992; Havas and Lalanne- Magne, 1992). These problems can be eliminated by adopting a prepolymerization stage, as shown in Figure 7.

Prepolymerization gives advantages in polymer particle size control and control of the catalyst activity in the fluidized bed reactor. Prepolymerization can be carried out in a liquid hydrocarbon medium or in a gas phase stirred reactor a t temperatures from 40 to 115 "C. Catalyst is introduced into the prereactor in the form of dry powder or in suspension in a liquid hydrocarbon. Prepolymer- ization is carried out to a conversion wherein the prepolymer contains 0.002-10 millimol of transition metal/g of polymer. The diameter of the prepolymer is in the range from 200 to 250 pm (Bailly and Speakman, 1986; Chinh and Dumain, 1990; Martens and Moterol, 1992; Havas and Lalanne-Magne, 1992). Prepolymer is fed into the fluidized bed reactor through a metering feed device. The fluidized bed reactor operates a t a superficial velocity of approximately 0.5 m/s, 2-8 times the minimum fluid- ization velocity (Dumain and Raufast, 1985). To avoid induction time a t startup, the prepolymer is treated with triethylaluminum for polymerization with chromium oxide as catalyst (Durand and Morterol, 1986a,b; Bailly and Speakman, 1986). To create more porosity, the prepolymer is also treated with n-hexane to remove low molecular weight polymer (wax) (Bailly and Speakman, 1986). Monomer, comonomer, hydrogen, and inert gas are fed through the bottom of the fluidized bed. Some comonomer and inert volatile hydrocarbon, such as isopentane, are introduced into the inlet line of the heat exchanger to avoid fine particles depositing on heat exchanger surfaces and compressor blades. The ratio of comonomer to monomer partial pressure is kept constant (normally 0.1- 0.2)) in the reactor (Chinh and Dumain, 1990).

To avoid pressure fluctuations during polymer discharge, BP Chemicals developed a continuous discharge device

4M Ind. Eng. Chem. Res., Vol. 33, No. 3, 1994

Table 2. Ethylene Gas Phaw Polymerization Prowsws and Reactor Operating Conditions Union Carbide BASF AMOCO BP Chemic&

nvletor type fluidized bed

catdlyst supported Ti. V and

caialyat size, pm 30-250 press., atm 20-30 temp. O C 75-110 comonomer 1-butene or 1-hexene

MWD. Mw/Mm 4-30 density, g/cmJ 0.91-0.87 palymer particle size. pm 500-1300 reference Rasmuassn (1972):

CrCOa catalyat

MWcontrol H a

N i d e t t i et al. (1988)

stirred bsd

supported CrCOS catalyat

-35 100-110 1-butene

narrow tu broad Ha

Trieaehmann et al. (1977): Choi and Ray (1985a)

consisting of continuously rotating plug valves shown as A and B in Figure 7. These valves are connected with each other in such a way that the two valves do not open at the same time, rotating a t a speed of about 0.5-1.0 rpm. The volume of the vessel between the valves is between 0.2% and 1% ofthevolumeofthefluidizedsolidcontained in the reactor (Raufast, 1987).

One of the disadvantages of gas phase polymerization is that the reactor operating temperature must be lower than the melting point ofthe polymer produced. Recently, Bailly and Collomb (1988; 1990) found that agglomeration ofthe polymer particles commencesat higher temperatures when pulverulent inorganic particles are employed. Thus, higher polymerization temperaturescan be used to produce LLDPE. The pulverulent inorganic particles, suchas silica oralumina,arechemicallyinerttocatalystsandmonomers. The inert particles have porosity 0.2-2.5 mL/g, surface area 20-90 mz, and diameter 0.5-20 fim. The particle size is 5(t500 times smaller than the mean diameter of prepolymer particles. According to the patent (Bailly and Collomb, 1988, LLDPE with density 0.860-0.910 g/cm3 can be produced with 0.005-0.2 wt % inert particles based on polymer in the reaction zone. The inert particles are mixed with prepolymer particles and then fed into the reactor. Rhee et al. (1991) also used inert particles, e.g., carbon black, silica, and clays, in the production of ethylene-propylene rubbers and found that the reactor can be operated a t temperatures higher than softening points of copolymers when amounts of about 0.3-50% by weight of inert material are employed. These inventions suggest that the soft polymer particles can he stabilized using inert particles. However, the morphology development of polymer polymerized at or near melting point has not been reported.

Typical reactor operating conditions and production capabilities of the gas phase processes discussed above are summarized in Table 2.

Compared with slurry and solution processes, gas phase processes have many distinct advantages. Slurry-based processes producePEcoveringa broad rangeof melt index; however, the range of densities attainable is limited. As density decreases, resin solubility in the dilute increases. At a density of about 0.930 g/cm3 sufficient dissolution occurs to foul the reactor. Hence, slurry processes are not suitable for LLDPE production. On the other hand, solution processes can produce PE over a broad range of densities, but only a limited range of molecular weights. As molecular weight increases, solution viscosity increases also. At some point, the increased viscosity limits reactor operability and productivity. Thus, solution processes are not suitable for making high molecular weight PE. Production capabilities of slurry and solution processes are shown schematically in Figure 8, after James (1986).

horizontal stirred bsd

~~ppor t sd Ti and crcos catalyst

20-40 70-110 propylene or 1-butene H, and temperature 5-17 0.95-0.98

Shepard et al. (1976); Jezl and Petern (1978a,b)

stirred reactor and f l u i d d bed

supported Ti and crco2 catalyat

-60 for prepolymn 15-25 70-115 I-butene or I-hexene Ha 6 2 0 0.91-0.96 300-1200 Speakman (1991): Bailly

and SpsaLman (1986)

0.970

0.980

0.950

a

5 0.93 0 . 0 -

In

0 0.920

0.910 0.01 0.10 1.00 10.0 100

MELT INDEX, 9/10 min Figure 8. Production capabilities of aolution and slurry polymerization processes.

A " I" I U" I "" 0910 L " "1 MELT INDEX, g/10 min

FiKure 9. Production capability of gas phase polymerization processes (James. 1986; Speakman, 1991).

Gas phase processes, which do not involve any liquid phase in the reaction zone, are not constrained hy solubility and viscosity. Hence, a gas phase process can produce a complete range of PE with densities of 0.91-0.97 g/cm$ and melt indexes of

Ind. Eng. Chem. Res., Vol. 33, No. 3, 1994 455 Table 3. Comparison of High-pressure Free-Radical Process and Low-Pressure Gas Phase Process

high-pressure process low-pressure process tubular or autoclave reactor

high operating pressure,

high reaction temperature,

high energy requirement high capital cost limited to low-density PE

no comonomer required (homopolymer) for LDPE production

catalyst is less sensitive to impurities in the feedstock

low raw material cost can produce ethylene and

vinyl acetate copolymer polymer with long-chain

branching

- 3400 atm -300 OC

fluidized bed or stirred

low operating pressure, bed reactor

-20 atm low reaction temperature,

80-110 OC low energy requirement low capital cost can produce both high- and

low-density PE up to l0-15% comonomer

required (copolymer)

catalyst is very sensitive

high raw material cost currently a limited range

of comonomers mostly linear polymer

chains with short-chain branching

to impurities

Compared with slurry and solution processes, however, gas phase processes do have the following disadvantages:

1. Reactor operating temperature is limited to the resin softening point. Thus the productivity of catalyst is also limited.

2. The poor heat-transfer efficiency of the gas phase is a disadvantage. Hence, additional inert heat transfer agent is required to maintain stable reactor operating conditions at high production rates.

3. There is a possibility of sintering and agglomeration of the polymer particles due to formation of local hot spots when a high-activity catalyst is used.

4. As polymerization progresses, fine particles will deposit on heat-transfer surfaces, compressor blades, and sloped walls of the reactor.

5. Sensitivity to variation in operating conditions, such as flow rates of catalyst and monomer feed and polymer discharge, is a disadvantage.

Gas phase processes offer many advantages over conventional high-pressure processes. The conventional high- pressure free-radical process can produce only low-density polyethylene due to the nature of its polymerization mechanisms. Table 3 compares some features of the high- pressure LDPE process and the low-pressure LLDPE gas phase process (Imhausen et al., 1981; Karol, 1983).

Significant progress has been made in overcoming the disadgantages of gas phase polymerization and in im- proving heat removal capacity, as discussed above. Overall, gas phase polymerization processes have had a great impact on polyethylene production technologydevelopment.

2.2. Experimental Methods. Despite the commercial success of gas phase ethylene polymerization technology, the public literature contains no accounts of fundamental scale-up studies of gas phase processes. There is a need, therefore, for a comprehensive understanding of detailed polymerization behavior in gas phase polymerization. A challenge for academic researchers studying gas phase polymerization of ethylene is how to scale down commercial processes for experimental laboratory studies.

Kinetic studies of ethylene gas phase polymerization were first carried out by Edgecombe (1963) and Lipman and Norrish (1963) using conventional Ziegler-Natta catalysts. Wisseroth (1969) reported on gas phase polymerization of ethylene using Phillips catalysts. With the development of commercial gas phase polymerization processes, more university-based researchers have at-

1 2

rt TO WATER BATH

/1 FROM WATER BATH

3 Figure 10. Laboratory reactors for gas phase ethylene polymerization. (1) Choi et al. (1983, 1985b); (2) Lynch et al. (1991); (3) Dusseault (1991).

tempted to extend experimental studies of gas phase polymerization in recent years. Choi et al. (1983,1985b) designed a 1-L gas phase stirred bed reactor equipped with an anchor-type agitator. Mabilon and Spitz (1985) and Spitz et al. (1988) used a 675-cm3 stainless steel reactor stirred by vertical oscillations for gas phase polymerization. However, the detailed reactor configuration used in this investigation was not described. More recently, Jejelowo et al. (1990) and Lynch et al. (1991) constructed a 1-L stainless steel reactor with a specially designed paddle agitator. The reactor temperature is controlled using a 26-L oil bath. To improve reactor mixing and heat transfer, Dusseault (1991) designed an 800-cm3 horizontal stirred reactor with both internal cooling coil and external cooling/ heating jacket. The agitator is designed in such a way so that the entire polymer bed can be turned over. Weist et al. (1989) used a 20-mm-diameter fluidized bed reactor to study polymer morphology development a t the initial polymerization stage. However, the reactor configuration and performance were not described. Figure 10 shows typical reactors used for ethylene gas phase polymerization studies reported in the literature.

Because the reactivity of the catalyst is very high, only a very small amount of catalyst, normally in the order of milligrams, is required for laboratory-scale polymerization using a 0.5-1.0-L reactor. Therefore, inert particles are usually added into the reactor to assist the initial mixing of the catalyst. Glass beads with diameters of 500-5000 pm are often used (Keii, 1972; Doi et al., 1982; Choi et al., 1983,1985b; Dusseault, 1991). Mabilon and Spitz (1985) and Spitz et al. (1988) used salt (KC1) or polyethylene powder to assist initial mixing. Lynch and Wanke (1991) used 450-pm Teflon powder. Glass beads may further grind the catalyst during the initial stage of polymerization. Salt is sensitive to moisture; hence careful drying procedure is required, but the separation of polymer from salt is very convenient. Although polyethylene powder is effective for initial mixing, it is not good for polymer property analysis because newly produced polymer can not be separated from the initial polymer powder. However, it may be useful for polymerization rate measurements without requiring polymer separation.

The laboratory-scale reactors shown in Figure 10 can provide satisfactory temperature control within a restricted operating temperature range. With high-activity catalysts,


Table 4. Main Requirements of Industrial Catalysts for Ethylene Gas Phase Polymerization

polymerization process requirements polymer property control

high productivity molecular weight proper kinetic behavior catalyst morphology density control of polymer morphology good comonomer incorporation easy feed to reactor low cost and reproducible

catalyst preparation

molecular weight distribution

chain-branching distribution particle size and bulk density polymer chain unsaturation

the initial polymerization rate is very high, and so the initial heat is difficult to remove with an oil or water bath cooling system. Therefore, reactor temperature control systems shown in Figure 10 are still not adequate for a broad range of reactor operating conditions. To carry out gas phase ethylene polymerization under commercial reactor operating conditions, the reactor must be able to provide fast heat removal to avoid hot spots in the polymerization zone and adequate mixing conditions to disperse a small amount of catalyst powder.

3. Physicochemical Phenomena In light of the preceding observations, it is evident that

gas phase polymerization involves a complicated physicochemical transition from gaseous monomer to solid polymer, although the synthesis process itself is relatively simple compared with liquid phase processes. In gas phase polymerization, catalyst preparation and polymer particle morphology development play important roles in reactor operation. Hence, it is crucial to understand the relevant microscopic processes of polymerization and their relationships.

3.1. Catalysts for Gas-Phase Polymerization. Ca- talysis is the most active area of research in ethylene polymerization with worldwide participation in both industrial and academic laboratories. In recent years, hundreds of publications on polyethylene synthesis and manufacture have been recorded annually in Chemical Abstracts. Most are related to catalysis studies. The extent and variety of research work dealing with the same problem reflects not only the great interest and extensive commercial applications in this area, but also the complexity of catalysts for ethylene polymerizations. Research progress concerning catalysts for ethylene polymerization has been a focus of several recent international conferences (Seymour and Cheng, 1987; Kaminsky and Sinn, 1988; Quirk et al., 1988; and Keii and Soga, 1990; Vandenberg and Salamone, 1992), and excellent reviews have appeared in the literature (Zakaharov and Yermakov, 1979; Karol, 1984; Hsieh, 1984; McDaniel, 1985; Nowlin, 1985; Hsieh et al., 1987; Barbe et al., 1987; Tait, 1989; Kryzhanovskii and Pranchev, 1990; Dusseault and Hsu, 1993). Catalyst development plays a vital role in the polymerization of ethylene, particularly in successful gas phase polymerization processes (Karol, 1983). For gas phase polymerization, although high productivity of the catalyst is a necessary condition, other features, such as copolymerization kinetic behavior and catalyst particle size and shape also play important roles in the production of high-quality LLDPE and HDPE. The main requirements of industrial catalysts are summarized in Table 4 (Karol et al., 1987; Dall'Occo et al., 1988; Burdett, 1988).

Two major catalyst families have been commonly used in commercial ethylene production: titanium1 vanadium based catalysts (Ziegler-Natta catalysts) and chromium oxide based catalysts (Phillips catalysts). In the past four

Table 5. Properties of Silica Support Used for Gas Phase Polymerization (Joraensen et al.. 1982: Nicoletti et al.. 1988)

silica range preferable range particle size, pm 10-250 30-100 surface area, m2/g at least 3 50

concn of catalyst metal, 0.054.5 0.2-0.3 pore size, A at least 80 100

mmol/g of support decades, research has been focused mainly on titanium and chromium oxide based catalysts (Karol et al., 1988). A typical industrial catalyst consists of active metal, modifiers, and inert support. Porous silica and magnesium halide are common supports. Table 5 shows typical properties of silica support used for commercial gas phase polymerization (Jorgensen et al., 1982; Nicoletti et al., 1988). Internal pore size distribution is in the range of 80-1000 A, and with the distribution peak in the range of 100-200 A, depending on preparation (Weist et al., 1987). A comparison of the performance of titanium-, vanadium-, and chromium-based catalysts is summarized in Table 6 (McDaniel, 1985; Karol, 1988). The significant difference between Ti- and Cr-based catalysts is that the molecular weight distribution (MWD) of polymers made using Ti catalysts in narrower than that using Cr-based catalysts.

High activity or productivity of catalyst means that the catalyst leaves no more than 10 ppm catalyst metal and 30 ppm chloride in the polymer formed. Titanium levels higher than this level lead to polymer color problems, and higher chloride contents result in excessive corrosion problems in processing equipment (Hsieh, 1980, 1984). According to recent patents, the metal content in polymer is very low for the catalyst used in gas phase polymerizations, e.g.,

Ind. Eng. Chem. Res., Vol. 33, No. 3, 1994 467 Table 6. Performance of Titanium, Vanadium, and Chromium Oxide Based Catalysts (McDaniel, 1986: Karol, 1988) - .

Mg/Ti/EDa VC13(THF)3 CrOs R&l + halocarbon promoter high high

activator R3A1 productivity high Hz response moderate high low MWD narrow, 3-6 intermediate-broad intermediate-broad, 6-30 a-olefin incorporation moderate high high activity decay rate moderate low-moderate very low polymer chain unsaturation low very low one C=C per polymer molecule

thermal, 300-1000 OC

long-chain branching no no

ED is an electron donor.

Table 7. Operational Methods to Produce High-Activity Catalysts (Karol, 1984) 1. chemical anchoring to surface of support

(a) TiCWMg(0H)Cl (b) CrOa/SiOz (c) (CKHdoCr/SiO, . . . 1 I,_

2. formation of bimetallic complexes (a) MgClz + 2TiC4 + 8Poc13 - [TiCl~~lZ[Mg(POC13)sl~+~2POC13 (b) 2MgC1~ + Tic4 + 7THF - [TiCls(THF)l-[Mg2Cl~(THF)61+ (c) MgC12 + Tic4 + 4CH3COzCzHs - TiMgCl&H&OzCzHs)4

3. insertion into defects of support (a) MgClz + Tic14 + ethyl p-toluate - ball-milling (b) MgClz + TiQedioxane - ball-milling (a) Tic4 + EtzAlCl - ,9-TiC1,.xEtAlC12 (b) Mg(0Et)z + Tic4 - [MgClz.Mg(OEt)z.Mg(TiC&)l (c) Mg(0Et)Z + Ti(O-nC,H& + EtAlC12 - trimetallic sponge

5. formation of solid solutions by cocrystallization (coprecipitation) (a) EtMgCl + Tic4 - TiCls-MgClz + organic fragments

4. formation of high surface area sponge

Tic4

bwmyl ether - treated solid - [TiCl~~(EtA1C12)o,03(ether)o,oJ

The patent literature usually gives only the state of the art for preparation of catalysts. The reasons for the choices of different compounds are not explained. To better understand the effect of each component used on the properties of a catalyst, Chien and his co-workers have fully characterized a typical MgC12 supported catalyst: MgC12/EB/PC/A1Et3/TiC14-A1Et3/MPT, where EB, PC, and MPT are abbreviations for ethyl benzoate, p-cresol, and methyl p-toluate, respectively (Chien, 1987, 1988, 1992). The major findings of academic studies on MgC12 supported catalysts have been comprehensively reviewed by Dusseault and Hsu (1993). The main reasons that MgCl2 is the most widely used as a catalyst support can be summarized as follows:

1. MgClz has crystalline forms similar to those of Tic13 (conventional Ziegler-Natta catalyst) (Chien, 1987; Barbe, 1987; Dusseault and Hsu, 1993). This distinct feature suggests that MgClz should have the ability to mimic the structure of active Tic13 and the ability to efficiently incorporate TiC14.

2. MgCl2 possesses desirable morphology as a support. Mercury porosimetry showed that MgClz has a large number of pores with radii > 200 nm (Chien et al., 1983). The structure of MgClz is sufficiently resistant to fracture by operational manipulation, but weak enough so that its internal structure will be broken down during polymerization.

3. MgC12 has a lower electronegativity compared with those of other metal halides. A metal halide with lower electronegativity than Tic13 will increase productivity of ethylene polymerization (Soga et al., 1987). MgClz also enhances chain-transfer reactions because Mn decreases with an increase in Mg/Ti ratio (Karol, 1984).

4. MgC12 is inert to the chemicals used for polymerization. It can be left in the final polymer product without deashing at the levels used in commercial production.

Because of the above features of MgClz, Tic14 can be incorporated with MgClz in different ways. Co-milling of

low concn of long-chain branching

Table 8. Approximate Dimensions of MgCll Supported Catalyst (Chien et al.. 1983: Chien. 1987: 1988)

~~ ~~ ~

pore surface volume, area,

composition diameter,A cm3/g m2/g MgC12 (HCl treated) -4000 0.41 MgCldEB 700-1500 0.41 2.0-7.0 MgClz/EB/PC/AlEts 300-500 1.2 50-70 MgC12/EB/PC/AlEt3/TiC& 25-170 1.3 100-150

MgClz with an electron donor (internal Lewis base), such as ethyl benzoate, can produce a disordered crystal structure and reduce crystalline size to 20-40 A. The surface area of the support increases to a maximum value of ca. 6.5 m2/g. The surface area of the support further increases through each step of the modification. Table 8 shows approximate dimensions of a MgClz supported catalyst obtained by ball-milling (Chien et al., 1983; Chien, 1987, 1988). Although the composition of commercial catalysts may not be the same as that shown in Table 8, the main features of MgC12 supported catalysts are the same. From Table 8 it can be seen that MgClz supported catalysts have large surface areas and internal pore volumes. Both surface area and pore volume change with modifier composition. Incorporation of T i c 4 with such a support results in dispersion or isolation of the titanium centers so that the population of active centers which are accessible for polymerization increases dramatically compared with that form unsupported conventional Ti catalysts. This is the main reason why supported Ti catalysts have productivities so much higher than those of conventional catalysts. In addition, magnesium ions can stabilize active titanium from deactivation, enhance chain- transfer processes (& of PE decreases when Mg/Ti ratio increases), and lead to narrow MWD for polyethylene production (Karol, 1984; Nowlin, 1985).

According to Chien (19871, both internal Lewis base and internal alcohol increase surface area and pore volume, but they do not compete for the same surface sites of MgC12 crystal because the content of EB on the support is the same after addition of internal alcohol. Therefore, they may adsorb on different planes of MgClz crystal. Alu- minum alkyls such as triethylaluminum (AlEt3) are usually used as cocatalysts to activate the catalyst. The activating action of AlEt3 is believed to reduce Ti4+ to Ti3+ and alkylation. In addition to the activation requirement, AlEt3 also acts as a scavenger of impurities of the polymerization system. The optimum ratio of Al/Ti is in the range of 10-30 (Dusseault and Hsu, 1993). AlEt3 can also reduce MgClz support particle size and increase surface area 2-3-fold if AlEt3 is used as an internal modifier during support preparation (Chien, 19871, as shown in Table 8.

Compared with Ti-based catalyst preparation, chromium-based catalyst preparation seems to be simpler. Chromium-based catalyst (Phillips catalyst) is usually prepared by impregnating a chromium compound onto a wide-pore silica and then calcining an oxygen environment to activate the catalyst. This leaves the chromium in the


hexavalent state, monodispersed on the silica surface. Chromium trioxide (CrOa) has been impregnated most commonly (McDaniel, 1985). Cr03 chemically anchors to the silica surface to form a new surface composition of silica (Karol, 1984; Hsieh, 1984; McDaniel, 1985; Nowlin, 1985). However, whether the chromium atoms are isolated as chromate, or exist in pairs as dichromates, is still debated in chromium catalyst studies. It is believed that chromium trioxide probably binds to the silica as chromate initially, and during calcining some chromate may convert to dichromate (McDaniel, 1985).

At least six methods have been described in preparation of Cr03 catalysts for ethylene gas phase polymerization (Bailly and Speakman, 1986). The main distinction among them is due to incorporation of modifiers. After CrOs is impregnated with porous silica, the catalyst is further impregnated with a titanium compound such as titanium tetraisopropoxide. Then, the catalyst is activated by a heat treatment a t 350-1000 "C under air for 30 min-5 h. The final catalyst contains 0.05-10% chromium and 0.50- 20% titanium (Bailly and Speakman, 1986; Speakman, 1991).

Catalyst prepared following the above procedure is not suitable for direct use in gas phase polymerization because of its long induction time. Cr(VI)/silica catalyst develops polymerization activity gradually when exposed to ethylene monomer. It is believed that Cr(V1) is reduced by ethylene to the Cr(I1) active site. Below 100 "C the induction time becomes longer until a t 60 "C there is almost no activity. However, a t 150 "C, the catalyst exhibits an immediate and constant activity in solution polymerization (McDaniel, 1985). For gas phase polymerization, the normal operating temperature is between 80 and 100 "C. The Cr-based catalyst has a significant induction period which causes difficulty not only in starting polymerization, but also in maintaining constant polymerization conditions in the reaction zone. When the polymerization is started, the catalyst activity rapidly accelerates, which also makes polymerization conditions difficult to control. Therefore, the catalyst has to be prereduced to eliminate the induction time. CO and alkyls of aluminum, boron, and zinc can be used as reducing agents to eliminate induction time (McDaniel, 1985). However, for gas phase polymerization, catalyst with fully activated chromium could lead to the formation of hot spots and cause agglomeration due to poor heat transfer in the reaction zone. According to Bailly and Speakman (1986) and Speakman (1991), the best way to overcome these problems is to convert the catalyst to prepolymer particles through a prepolymerization stage as described above.

It should be pointed out that the activation temperature is a very important parameter in controlling catalyst activity and polymer molecular weight for ethylene polymerization using Cr-based catalyst. Figure 11 shows the effects of activation temperature on the reactivity and molecular weight of polymer (McDaniel, 1985,1988). Here, activity is defined as the inverse of the time needed to produce 5000 g of polymer/g of catalyst. The activity increases with an increase in activation temperature up to maximum at around 925 "C, and then declines as sintering destroys the surface area and porosity of the silica base. The weight-average molecular weight decreases with an increase in calcining temperature up to 925 "C; then the trend reverses as the silica begins to sinter. The melt index shows the same type of relationship. The relationships shown in Figure 11 suggest that the chain termination reaction is very sensitive to the activation temperature for Cr-based catalyst systems.

lo 200 _I_I 150 600 600 700 800 900 lo00

ACTIVATION TEMPERATURE ('C)

Figure 11. Effecta of activation temperature on reactivity and polymer properties (McDaniel, 1985, 1988). Reprinted with per- mission from McDaniel (1985). Copyright 1985 Academic Press.

The silica supported catalysts described above have been successfully used for PE production. Deashing processes have been eliminated for many years. However, silica does remain in the final polymer product as an impurity. This impurity causes the polyethylene to have a poor film appearance. Hence, it is desirable to eliminate silica by use of polymer particles as a support. Research in this area has been reported in recent patent publications (Cana et al., 1992). Polypropylene and polyethylene or their copolymers have been used as supports. The catalyst precursor is bonded on the polymer particles. However, detailed information about how the catalyst is bonded on the polymer particles is not given. The results using polymer supported catalyst are sufficiently promising that polymer supported catalysts may replace nonpolymer supported catalysts for the manufacture of polyethylene film product in the near future.

The term "supported catalyst" is used in a very wide sense and includes not only systems in which the transition metal compound is linked to the substrate by means of a chemical bond, but also systems in which the transition metal atom may occupy a position in a lattice structure, or where complexation, adsorption, or even occlusion may take place (Tait, 1989). In commercial applications, MgC12 supported catalyst is further supported on silica or polymer through prepolymerization. All of these supports play a role in polymerization and in polymer morphology development.

Besides the heterogeneous catalysts discussed above, another catalyst family that may soon be commercialized in gas phase olefin polymerizations is metallocene-based catalysts. The metallocene-based catalysts are also called homogeneous Ziegler-Natta catalysts in the literature. The first homogeneous Ziegler-Natta catalyst, CpzTiClz/AlEt~- C1 (Cp = cyclopentadienyl) was independently discovered by Natta and Pino (1957) and by Breslow and Newburg (1957). The kinetic behavior of this catalyst was studied by Breslow and Newburg (19591, Chien (19591, and others (Reichert, 1983). The productivity of this catalyst is very low, less than 300 g of PE/(g of Tiehaatm) (Chien, 1959; Chien et al., 1990). However, if cocatalysts (AlEt, or AlEtC12) are treated with water, both productivity and molecular weight (MW) of PE increase dramatically (Reichert and Meyer, 1973; Andresen et al., 1976; Chihlar

Ind. Eng. Chem. Res., Vol. 33, No. 3, 1994 469 to be well on the way to commercialization of metallocene technologies for polyolefins (SRI, 1993; Kuber, 1993; Payn, 1993).

While polyolefin companies worldwide jostle in the race to commercialize metallocene-based technologies, the fundamental studies of metallocene catalysts carried out in companies are highly confidential. In academic insti- tutes the kinetic studies of metallocene catalyzed olefin polymerization were mainly carried out in solution processes. The following metallocene-based catalysts have been studied extensively: CpZMClZ/MAO (M = Ti, Zr, Hf) (Herwig and Kaminsky, 1983; Kaminsky et al., 1983; Kaminsky and Luker, 1984; Kaminsky and Hahnsen, 1987; Chien and Wang, 1988; Mallin et al., 1988; Ahlers and Kaminsky, 1988; Tsutaui and Kashiwa, 1988; Tait, 1988; Chien and Wang, 1989; Chien and Wang, 1990; Kaminaka and Soga, 1992; Uozumi and Soga, 1992); [IndlzZrCld MA0 (Ahlers and Kaminsky, 1988; Chien and He, 1991a- d); chiral metallocenes, e.g., Et[IndIzZrClz (ethylene bis(indeny1)zirconium dichloride), Et[IndIzZrMez, and Et[IndH4]zZrCl2 (Ewen, 1984; Kaminsky et al., 1985,1988; Kaminsky, 1986; Ahlers and Kaminsky, 1988; Drogemuller et al., 1988; Ewen et al., 1988; Tait, 1988; Antberg et al., 1990; Chien and He, 1991a; Chien and Sugimoto, 1991; Kioka et al., 1990; 1992; Heiland and Kaminsky, 1992; Soga and Kaminaka, 1992; Uozumi and Soga, 1992; Soga et al., 1993; Chien and Nozaki, 1993; Chien and Gong, 1993); and other metallocene derivatives (Kaminsky et al., 1986,1991; Ewen et al., 1988; Tait et al., 1992; Chien et al., 1992). Compared with heterogeneous Ziegler-Natta catalysts, metallocene-based catalysts have distinct kinetic characteristics for olefin polymerizations, which are summarized as follows.

Superhigh productivity and good polymerization rate profile. Polymerization rate reaches a maximum within 5 min, decays slightly, and then remains almost constant. For the Cp2MRz (M = Ti, Zr, Hf; R = C1, CH3) catalyst family, zirconium catalysts are more active than titanium or hafnium systems, a t temperatures over 50 OC. The cocatalyst methylaluminoxane is better suited than ethylaluminoxane or isobutylaluminoxane (Sinn and Ka- minsky, 1980; Kaminsky et al., 1983, 1986; Herwig and Kaminsky, 1983; Kaminsky and Hahnsen, 1987; Mallin et al., 1988; Tsutsui and Kashiwa, 1988; Chien and Razavi, 1988; Chien and Wang, 1989,1990; Chien and Sugimoto, 1991; Tait et al., 1992). For Et[Indl2MC12 (M = Zr, Hf), Hf catalysts show higher activities (Heiland and Kaminsky, 1992). Productivity of metallocene catalysts can be up to 3 X lo7 g of PE/(g of Zr*h) (Kaminsky et al., 1986).

Very high initiation efficiency. Active center concentration is 75-100% of Zr for ethylene polymerization using zirconocene/MAO, increases monotonically with temperature, and is sensitive to Al/Zr ratio (Chien and Razavi, 1988; Tait, 1988; Chien and Wang, 1989; Chien and He, 1991b; Tait et al., 1992).

Narrow molecular weight distribution. In general, polydispersity of PE or its copolymers produced using metallocene catalysts is between 2 and 5; in most cases, polydispersity is about 2 (Tsutsui and Kashiwa, 1988; Chien and Wang, 1990; Chien and He, 1991a; Uozumi and Soga, 1992). Polydispersity can be controlled by mixing metallocene catalysts or by mixing metallocene with heterogeneous Ziegler-Natta catalysts (Ahlers and Ka- minsky, 1988; Kioka et al., 1990; Heiland and Kaminsky, 1992; Fries and Bowles, 1993).

Molecular weight sensitivity to hydrogen and temperature. Chain transfer to hydrogen and @-hydride elimination rate constants are 2-3 orders of magnitude greater than the corresponding values found for MgClz

et al., 1978). It was believed that aluminoxane plays an important role in the increase in productivity (Andresen et al., 1976; Chihlar et al., 1978; Sinnand Kaminsky, 1980). Significant breakthroughs however occurred in the early 1980s when Kaminsky and his co-workers found very high productivity for ethylene polymerization using CpzZrCl2 as catalyst and methylaluminoxane (MAO) as cocatalyst (Sinnet al., 1980; Kaminsky et al., 1983; Kaminsky, 1983; Herwig and Kaminsky, 1983; Kaminsky and Luker, 1984). Furthermore, Ewen (1984) and Kaminsky et al. (1985) demonstrated that chiral rac-Et[IndH412TiClz and rac- Et[Ind]zZrClz/MAO (Ind = indenyl) catalysts are capable of stereospecific polymerization of a-olefins. Since the mid-l980s, metallocene-based catalysts have become of academic and commercial interest.

The level of industrial research activities in metallocene catalysis has been tremendous. More than 50 companies worldwide have been involved in metallocene catalyst studies. Among them, Mitsui Petrochemical, Indemitsu, Exxon, Mitsui Toatsu, Hoechst, Asahi Chemical, and Dow Chemical are the leading companies in terms of the number of patents issued. To date, over 50 metallocene catalysts have been described based on Ti, Zr, and Hf metallocenes (SRI, 1993). Exxon Chemical was the first company to apply metallocene catalysis to polyolefin production on a commercial scale. The company built a 15 000 t/y high- pressure autoclave reactor unit a t the Baton Rouge site. This plant was started up on June 1991 and has produced specialty linear ethylene-based polymers called EXACT resins: about 20 grades with a range of densities from 0.865 to 0.935 g/cm3 (Mod. Plast., 1991; SRI, 1993; Montagna and Floyd, 1993; Rogers, 1993a). To further increase product range, Exxon and Mitsui Petrochemical have been collaborating for commercialization of gas phase process technology using metallocene catalysts (SRI, 1993; Rogers, 1993a). To further increase product range, Exxon and Mitsui Petrochemical have been collaborating for commercialization of gas phase process technology using metallocene catalysts (SRI, 1993; Rogers, 1993b). The gas phase process is expected on stream in the mid-1990s and will be operated at temperatures of 80-90 "C and will produce 100 000 t/y of ethylene copolymers with densities ranging from 0.86 to 0.975 g/cm3. Dow Chemical, in competition with Exxon, has developed a new family of polyolefins using Constrained Geometry Catalyst Tech- nology, This technology was commercialized under the trade name INSITE in mid-1993 (Rotman and Wood, 1993; Schwank, 1993; Lindsay, 1993; Mod. Plast., 1993; Story and Knight, 1993). The distinct characteristic of INSITE polyethylenes is improvement in physical properties without sacrificing processability as a result of a significant amount of long-chain branching formed in polymer chains (SRI, 1993; Stevens, 1993). Thus, for the first time, polyethylenes made through coordination polymerization have the melt rheological advantages currently obtainable only in free-radical-polymerized high-pressure LDPE. INSITE resins are now produced in solution processes (Rotman and Wood, 1993). However, Dow claims to have successfully carried out polymerizations using its metallocene catalysts in other processes as well, including high- pressure, gas phase, and slurry processes (SRI, 1993). Mobil Chemical has also carried out ethylene and l-hexene copolymerization using metallocene catalysts in a pilot- scale fluidized bed reactor. Commercialization of a gas phase process using a metallocene catalyst is on the way (Furtek, 1993). In addition to the companies above, there are others in Japan, Germany, and Belgium that appear


supported heterogeneous Ziegler-Natta catalysts (Ka- minsky and Luker, 1984; Chien and Wang, 1990). Hy- drogen also reduces polymerization rate (Kaminsky and Luker, 1984; Kioka et al., 1992).

Uniform incorporation of comonomers. Composi- tion of copolymer is almost the same as the monomer feed composition, and comonomer is uniformly distributed in copolymer chains regardless of chain length (Chien and He, 1991a,d; Uozumi and Soga, 1992). The effect of comonomer on ethylene polymerization rate depends on polymerization temperature. 1-Hexene or propylene enhances ethylene consumption rate a t 30 and 50 "C (Tsutsui and Kashiwa, 1988; Koivumaki and Seppala, 1993). However, 1-hexene reduces ethylene polymerization rate at 70 and 95 "C (Chien and Nozaki, 1993; Koivumaki and Seppala, 1993).

Stereochemical control in microstructures of olefin polymers. Metallocene catalysts by manipulating the structures of organocompounds can polymerize olefins with very high stereoregularity to give either isotactic or syndiotatic polymers (Ewen, 1984; Kaminsky et al., 1985, 1986; Kaminsky, 1986; Ewen et al., 1988; Antberg et al., 1990; Chien and Sugimoto, 1991; Antberg et al., 1991; Uozumi and Soga, 1992).

Very high cocatalyst ratio. To obtain high polymerization rate and high active center concentration, the Al/Zr ratio needs to be 1OL1O4or higher (Chien and Razavi, 1988; Sinn et al., 1988; Chien and Wang, 1989). If the AVZr ratio is small, the polymerization rate increases slowly and has an induction period (Chien and Sugimoto, 1991). The reason for high cocatalyst ratio is still unclear. The cost of cocatalyst could be more than 200-300 times the cost of the catalyst in commercial production (SRI, 1993). Hence, high cocatalyst ratio is one of the major barriers to commercialization of metallocene technologies.

Metallocene catalysts must be supported on a carrier for use in gas phase processes. Chien and He (1991~) found that Et[Ind12ZrClz supported on Si02 does not change polymerization behavior. Interestingly, a smaller amount of MA0 is required in the case of the supported catalyst than for the homogeneous system to achieve the same catalytic activity. Et[IndHrlzZrClz supported on A1203 or MgC12 can be activated by common trialkylaluminums; however, the catalyst cannot be activated by common alkylaluminums if it is supported on Si02 (Soga and Kaminaka, 1992; Soga et al., 1993). The MgClz supported metallocene catalysts produced a broader MWD (Soga et al., 1993). CpzZrClz supported on A1203 or MgClz produces atactic polypropylene, but is inactive to propylene using Si02 as a carrier (Kaminaka and Soga, 1992). The effect of a carrier on the performance of metallocene catalysts is unclear a t present.

For six decades, the polyolefin industry has experienced several great revolutions. Despite the fact that there are some technological difficulties, commercialization of metallocene-based catalyst technologies may be the greatest revolution of all. The impact of metallocene-based catalysts on polyolefin industry is perhaps beyond the current imagination. 3.2. Polymer Particle Morphology Developments.

One of the most important features of ethylene polymerization using heterogeneous catalysts is replication of polymer particles (Berger and Grieveson, 1965; Mackie et al., 1967). This suggests that the morphology development of polymer particles depends on the original morphology of the catalysts. Polymer growth on titanium crystal surfaces has been carefully examined using electron microscopy (Carradine and Rase, 1971; Baker et al., 1973),

and it was found that the polymer first accumulates on boundaries of catalyst fragments and at existing cracks. Carradine and Rase (1971) observed that the polymer appearing on lateral surfaces is nodular in character while that growing on basal surfaces is fibrous. Baker et al. (1973) observed nodular polymer growth around small catalyst particles (-0.1 pm in diameter) during gas phase polymerization. The shape of the polymer growth was usually governed by that of the active catalyst particle. Wristers (1973a,b) examined the morphology of classical Ziegler-Natta catalyst using electron scanning microscopy, and found that the gross catalyst particles are composed of primary particles whose size and shape depend on the preparation conditions of the catalyst. According to Hock (1966) and Graff et al. (1970), the size of a primary particle is in the range 10-1000 A in diameter. The shape and reactivity of primary catalyst crystals are responsible for the fiberlike structures that make up the final polymer particles (Wristers, 1973a,b). On the basis of electron microscopy and adsorption data, Bukatov et al. (1982) suggested that the 10-30-pm catalyst granules consist of three levels of particles: 200-1000-A, 1000-5000-A, and

Ind. Eng. Chem. Res., Vol. 33, No. 3, 1994 461

Waals force. There are interstices or pores between the primary particles, providing a large surface area.

2. At the beginning of polymerization, monomer diffuses through the interstices and adsorbs onto the surface of catalyst primary particles, where the active centers are located (Bohm, 1978a).

3. The propagation reaction occurs on the surface of the catalyst primary particles. When monomer is converted to polymer, polymer precipitates and accumulates around the catalyst primary particles. As polymerization proceeds, the interstices are gradually filled with polymers. Hydraulic pressure or stress between catalyst primary particles increases due to accumulation of polymer chains in the pores leading to separation of catalyst primary particles. Polymer around the catalyst primary particles cements these primary particles. The distance between the catalyst primary particles increases with an increase in polymer yield. Thus polymer particles grow isotropically during polymerization. When the polymer particles with high polymer yield are examined using electron microscopy, catalyst primary particles are found to be uniformly dispersed in the polymer matrix as though the catalyst were broken up during polymerization. In fact, this should be regarded as a separation process. Catalyst primary particles should be broken in the preparation stage. Interstices or pores must exist between catalyst primary particles, where active centers are located, so that polymerization occurs on the surfaces of the interstices, leading to separation of catalyst primary particles. 4. The polymer around the catalyst primary particles

is swollen with monomers (ethylene and a-olefin). Since olefin polymerization is a highly exothermic process, the temperature within the polymer primary particle could increase rapidly, and may exceed the softening or melting point of the resin (Karol et al., 1987). Due to the rapid growth of primary particles, the polymer chains between primary particles fuse together to form polymer primary particle aggregates of diameters in the range of 0.5-1.0 pm.

5. The final polymer particles are composed of aggregates of primary particles with average diameters in the range of 200-500 pm (Galli et al., 1981, 1984; Kakugo, 1988; Bohm et al., 1988). The morphology of the final particles depends on polymer yield and the balance between annealing and distorting forces during particle growth. The primary particles grow isotropically (Kakugo, 1988); therefore, in general, the polymer particles replicate their parent catalyst particles in shape and size.

The mechanism described above can be expressed schematically as in Figure 13. To avoid confusion between catalyst primary particle and polymer primary particle, the catalyst primary particle is called a domain in Figure 13. Polymer primary particles are formed with a catalyst domain as a core, which is the smallest resulting catalyst particle during polymerization.

3.3. Chemical Reactions. From the preceding observations, the chemical reactions of ethylene polymerization can be envisioned as occurring at the interface between the solid catalyst and the polymer matrix, where the active centers are located. From gas-state monomer to solid-state polymer, ethylene experiences a dramatic physicochemical transition within a very short time. The polymerization environment changes with the composition of catalyst, polymerization process, reactant composition, reactor operating conditions, and extent of polymerization. Although intensive research activity has been focused on Ziegler-Natta catalyst systems since their discovery in the early 1950s, no definitive, unequivocal chemical

100 1 P t 8o 70 L

CATALYST 7 i

/ I , , , I

0-38 45-75 100-250 500-1000 3845 75-100 250-500 >loo0

PARTICLE SIZE, micron Figure 12. Particle size replication of ethylene polymerization with silica-supported Ziegler-Natta catalyst (Munoz-Escalona et al., 1988).

scanning electron microscopy and found that the catalyst consists of small particles with a rough surface. However, the internal structure of the catalyst was not shown. If the catalyst is impregnated on silica support, the particle size is in the range of 30-250 pm (Bailly and Speakman, 1986; Munoz-Escalona, 1988). The average polymer particle size is about 20 times larger in diameter than the catalyst particle size (Galli et al., 1981, 1984; Karol et al., 1987).

Kakugo et al. (l988,1989a,b) recently studied internal morphology development of polypropylene particles using transmission electron microscopy. Their results indicate that each polymer primary particle has a nucleus 50-170 A in diameter, which is believed to be the catalyst primary particle. It is interesting to notice that the catalyst primary particle size identified in the polymer particle is in excellent agreement with catalyst crystal particle size estimated by Chien (19871, as shown in Table 8. This suggests that the minicrystal size of a catalyst achieved by ball-milling cannot be reduced further during polymerization, or that the Ti catalyst active center is fixed on the surface of this mini MgClz crystal and not inside the minicrystal. The diameter of a polymer primary particle is about 0.2 pm. The fine polymer globules, about 1 pm in diameter, are formed by fusion of several polymer primary particles. The gross polymer particles are composed of these primary particle aggregates. Weist et al. (1989) found that the pore distribution of silica supported catalyst does not change at the initial stage of polymerization. However, the pore size of a catalyst increases with an increase in polymer yield, indicating fragmentation during polymerization. The SEM photographs of polymer particles show that the catalyst particles are completely encapsu- lated by polymer in the early stage (


10-30 micron in diameter (30-100 micron for silica supported catalyst) Composed of minute crystal domains

Formation of polymer primary particles, -0 2 micron in diameter Separation of catalyst domain

CATALYST

'zH4 20-200 A"in diameter

LOW POLYMER

YIELD

Growth of primary particles Formetion of primary particle aggregates, -1 micron in diameter

MEDIATE

Figure 13. Scheme of polyethylene morphology development during gas phase polymerization.

reaction mechanisms have been developed to fully describe the kinetic behavior of ethylene homo/copolymerization, due to the complexity of the systems employed. Never- theless, the key elementary reactions have been estab- lished, which include formation of active centers, insertion of monomer into the growing polymer chains, chain- transfer reactions, and catalyst deactivation. Most of the proposed mechanisms are based on information about polymerization rate, molecular weight and its distribution, polymer chain microstructure, and active center concentrations. Detailed mechanisms have been discussed in monographs (Boor, 1979; Kissin, 1985) and in recent reviews (Tait and Watkins, 1989; Dusseault and Hsu, 1993). Among the proposed propagation mechanisms, the one proposed by Cossee (Cossee, 1964; Arlman and Cossee, 1964) together with its subsequent modifications has been widely adopted. With Cossee's mechanism in mind, Bohm (1978a) proposed perhaps the most comprehensive elementary reactions involved in the ethylene polymerization process.

Since commercial production of LLDPE and HDPE consists of a copolymerization process, obviously, copolymerization mechanisms are required to understand kinetic behavior and polymer properties. Therefore, Bohm's reaction model has been modified and extended to the ethylene copolymerization processes in recent modeling studies (Villermaux et al., 1989; de Carvalho et al., 1989,1990; McAuley et al., 1990; Lorenzini et al., 1991; Hutchinson et al., 1992). Furthermore, to illustrate multimodal distribution of polymer chain composition and broad moleclar weight distribution, the multiple active center concept has been adopted in recent modeling studies. Therefore, a comprehensive chemical reaction mechanism should include not only reactions proposed by Bohm (1978) but also reactions involving multicom- ponents at different active centers. Whether each type of active center has the same or different reaction mechanisms is still uncertain (multiple center formation itself has not been proved fundamentally). If one assumes that all of the active centers perform the same reaction mechanisms, but with different reaction rates for each elementary reaction, then elementary reactions which are commonly adopted in modeling studies can be summarized as in Table 9, where C, is the catalyst potential active

Table 9. Summary of Elementary Reactions for Ethylene and a-Olefin Copolymerization

reaction description Activation

spontaneous activation

activation by aluminum

activation by electron

activation by hydrogen

activation by monomer

activation by monomer

alkyl (A)

donor (E)

(Hz)

1 (Mi)

2 (Mz) Initiation

initiation of M 1 by normal

initiation of Mz by normal

initiation of MI by active

initiation of Mz by active

initiation of MI by active


initiation of M 1 by active


active center

active center

center with H

center with H

center with A

center with A

center with E

center with E Propagation

KPll propagation of chain

K ~ L Z propagation of chain

K ~ P I propagation of chain

KPln propagation of chain

P m , n , l + [ M I 3 - P m + l , n , I P*m,n,l + [ M 2 1 - p*m,nt1,2 Pm,n,z + [MI3 -* P m + l , n , *

P m , n , z + [Mz] - P m , n + l , z

type 1 with M 1

type 1 with M 2

type 2 with MI

type 2 with Mz Chain Transfer

Kbpl spontaneous chain transfer P*m.n.i - PO + 4'm.n or &elimination

KIN chain transfer to hydrogen P m n L + [HJ - PH.O + Q'm.n (Hz) I .

KfAl chaintransfer to aluminum

Kls, chain transfer to electron Pm,n ,L + [AI - P*A,O + Q'm,n P m , n , i + [El -+ P*E,o + P'm,n

alkyl (A)

donor (E) KMl3 chain transfer to M 1

P m , n i + [ M I ] - P*I,o,I + Q'm,n K m chain transfer to Mz

P m , n , i + [Mz] - *o,I,, + q'm.n Deactivation

KW spontaneous deactivation Pm,n,i -* cd ' qm,n

Kd7i deactivation by impurities or

KW deactivation by aluminum

KdW deactivation by electron

Pm,n,i + lZ1 -* cd + qm,n

plm,n,i + - 'd + qm,n alkyl (A) poison (Z) donor (E) Pm.n.i + - cd ' 4m.n . .

KdHl deactivation by hydrogen (H2)

KdJ& deactivation by monomers Pm,n,+ + [ H 2 1 'd + qm,n

Pm,n,L + r M j ] - 'd + 4m,n Other Possible Reactions

K'Pl formation of short-chain

K'pz formation of short-chain

K'Pl formation of long-chain

K'PI formation of long-chain

P I , O J + q'ra - p r + 1 + , 1 p O , l , z + q'ra - P r + + 1 . 2 P*m,n,l + q f r a -+ ~ m t r , n t s , l

p m , n , 2 + Q'ra - P m + r , n t a , z

branches

branches

branches (rare)

branches (rare)

Ind. Eng. Chem. Res., Vol. 33, No. 3, 1994 463

center; P*o is the active center without polymer chain; P*m,n,i is the active center with m units of monomer 1 and n units of monomer 2, with the third subscript i denoting the chain terminal monomer type bonded to the active center; qm,n is a dead polymer chain with a terminal double bond; and q,,,, is a dead polymer chain without terminal double bond. For simplification of notation, no subscript is shown corresponding to the type of active center. The reactions shown in Table 9 should be considered to occur a t each type of active center. The terminal model is assumed to be valid for binary copolymerizations. How- ever, if the penultimate model is valid for some systems, the mechanisms in Table 9 can be modified accordingly. Hutchinson et al. (1992) considered transformation reactions, which are assumed to form different active centers. If an active center formed by transformation is independent of the original polymer chain, then the effect of a transformation reaction is on the initiation reactions which have been shown in Table 9. If an active center formed by transformation depends on the original polymer chain produced by this center, then the formation of each active center must be treated individually. The detailed mechanisms of transformation reactions have not been given (Hutchinson, 1992). Formation of short-chain branches was proposed by Bohm (1978). It was assumed that a polymer chain with a terminal double bond will participate in propagation at an active center with one monomer unit. Formation of long-chain branching is unlikely for Ziegler- Natta polymerization systems. However, for Phillips catalyst systems, polymer chains with long-chain branching are formed, although branching frequency is low (McDaniel, 1985). For ethylene polymerization with a Ziegler-Natta catalyst at high temperature and high pressure, polymer chains dissolve in the monomer phase. Hence, polymer chains with a terminal double bond are able to diffuse to active centers to participate in propagation and to form long-chain branching. However, detailed information regarding this reaction has not been published to date. The mechanisms in Table 9 are summarized based on Ti-based catalysts. Reaction mechanisms for chromium oxide based catalysts are a subset of those in Table 9. It should be pointed out that second- order deactivation of active centers is not included in Table 9.

4. Reactor Modeling Reactor modeling is the determination of a quantitative

relationship between reactor performance and reactor operating conditions. It requires comprehensive understanding of polymerization processes, physical phenomena, and chemical reaction mechanisms. The importance and benefits of reactor modeling have been widely recognized by both industrial and academic researchers. The application capability of a model depends on scope of the modeling effort. Ray (1986, 1991) defined a modeling hierarchy as microscale, mesoscale, and macroscale, according to the characteristics of the polymerization reactor systems. The relationships between modeling scale and polymerization systems are outlined in Figure 14 after Ray (1986,1991). The emphasis of each modeling level, in particular for ethylene polymerization systems, can be summarized as follows:

1. Microscale modeling: polymerization rate development, or polymer yield; reactant species distribution; molecular weight development and ita distribution; chemical composition of polymer chains and its distribution; microstructures of polymer chains, including chain branching, unsaturated groups, and sequence distribution.

MICROSCALE + 4, MESOSCALE I

I I 1 1 MACROSCALE Figure 14. Levels of polymerization reactor modeling.

2. Mesoscale modeling: interphase heat and mass transfer; intraphase heat and mass transfer; fluid me- chanics and micromixing; polymer particle morphology development; polymer particle size and distribution.

3. Macroscale modeling: macromixing and residence time distribution; overall material and energy balance; heat and mass transfer from the reactor; reactor dynamics and control; polymer grade transition and control.

There are no definitive boundaries between these modeling levels. In fact, they often overlap during modeling studies. For instance, knowledge of some microscale modeling is required for mesoscale modeling studies, and macroscale modeling depends upon understanding of microscale and mesoscale phenomena. A complete model for ethylene polymerization, including all three levels, has not been developed in the literature. However, significant modeling efforts with emphasis on specific levels have been published. Microscale and mesoscale modeling studies are normally reported as kinetic modeling in the literature. Macroscale modeling is often referred to as dynamic modeling. Major contri- butions to ethylene polymerization reactor modeling are now discussed.

4.1. Kinetic Modeling. Kinetic modeling has focused mainly on polymerization rate and polymer molecular weight development, which can be expressed as follows:

2 2

i=l j -1

MW f ~ ~ , ~ ~ * ~ l , ~ ~ j l , ~ ~ ~ l , ~ ~ l , ~ ~ l ~ .-.) (2) where binary copolymerization is assumed. With the long- chain approximation, polymerization rate is a function of the propagation rate constants, monomer concentrations, and active center concentration. Complexity can arise through the expression for active center concentration. Molecular weight development is a function of all the components which affect polymer chain length, as shown in Table 9. One of the important features of ethylene polymerization is the broad molecular weight distribution of the polymer. Two theories have been developed to explain broad MWD of polyethylene, namely, diffusion theory and multiple active center theory. The former emphasizes the effect of monomer concentration on molecular weight development. The latter is more con- cerned with the influence of active centers and associated kinetic parameters. Some of the experimental results in support of these theories are as follows.

Diffusional limitations: 1. Polymerization rate depends on stirring speed when

agitation speed is below a critical level (Berger and Grieveson, 1965; Bohm, 1978b).


2. Molecular weight of PE formed in the initial period of polymerization is much higher than that of polymer formed in later stages (Crabtree et al., 1973).

3. The polymerization rate of ethylene increases significantly in the presence of small amounts of comonomers, such as propylene and l-hexene (Finogenova et al., 1980; Soga et al., 1989). Soga et al. (1989) found that the crystallinity of the copolymer decreases with an increase in comonomer concentration. The reactivity ratio of propylene to ethylene is less than 1. Therefore, an increase in reaction rate in the presence of a small amount of propylene is due to a decrease in the monomer diffusion resistance. This suggests that diffusion resistance is very significant in highly crystalline polymers, such as HDPE (Soga et al., 1989).

The results of Finogenova et al. (1980) and Soga et al. (1989) are very convincing. In fact, the effect of crystallinity on monomer concentration in the polymer matrix has not been comprehensively studied. However, the effect of a-olefins on the ethylene polymerization rate depends on the a-olefin types. It is not certain whether physical or chemical effects dominate. Crabtrees result is in contradiction with data reported by Bohm (1978~) and others (Kissin, 1985). Bohm (1978~) showed that number- average molecular weight increases significantly with reaction time during the early period of polymerization and then levels off with an increase in reaction time.

Multiple active centers: The following phenomena cannot be explained by diffusion limitation theory (Zuc- chini and Cecchin, 1983; Karol, 1984).

1. Different transition metal catalysts can provide large changes in polymer MWD.

2. Chemical modification of catalyst changes the breadth of the MWD.

3. Heterogeneous catalyst can produce broad MWD, even when the polymer is in solution.

4. Homogeneous soluble catalysts provide narrow MWD, even when the polymer is insoluble in the reaction medium.

5. High-activity catalyst does not necessarily provide broad MWD.

6. Ti-based catalyst can provide polymer with narrow MWD, and vanadium-based catalyst produces resins of broad MWD. A combination of these two catalyst systems, however, gives MWD of intermediate breadth with the same reactor operating conditions (Nicoletti et al., 1988).

7. Composition of LLDPE separated using temperature- rising elution fractionation had two or three distinct peaks (Wild et al., 1982a; Usami et al., 1986; Kelusky et al., 1987; Mirabella et al., 1987; Hosoda, 1988). The reactivityratios, as measured by 13C NMR, were different for each peak (Usami et al., 1986).

Although the above phenomena cannot provide con- clusive proof of multiple active center formation mechanisms, they do suggest that molecular weight development of PE depends on the nature of the catalyst and on the polymerization mechanisms. Many mathematical models have been developed to quantitatively describe the kinetic phenomena of ethylene polymerization. Most of the early models focused on polymerization rate development (Keii, 1972). However, the emphasis has been placed on polymer quality development in more recent modeling studies. The progress of kinetic model development is now discussed.

Bohm (1978a) developed a model based on a comprehensive chemical reaction scheme, which is applicable to both homogeneous and heterogeneous catalyst systems. The model development is based on assumptions of steady state for all active species, kinetic rate constants inde-

pendent of chain length, and slow deactivation of the catalyst. However, the model is limited to polymerization rate and number-average molecular weight (M,) of ho- mopolymerization. Higher order molecular weight and MWD were not considered.

Dussealt (1991) and Dusseault and Hsu (1993) developed areaction rate model for gas phase ethylene polymerization based on the multiple active sites hypothesis. This model is able to incorporate different distribution functions of active centers to simulate polymerization rate behavior. Both first-order and second-order deactivation mechanisms are considered in the model calculations (Dussealt, 1991). However, the model treatment is limited to polymerization rate. The effect of multiple active site distribution on polymer molecular weight development was not considered.

de Carvalho et al. (1989) developed a comprehensive ethylene and a-olefin copolymerization model accounting or the formation, initiation, and deactivation of active centers, and for the spontaneous and chain transfer to hydrogen, monomer, and organometallics. The state of the art to model copolymerization on multiple active sites was described in detail. This general model was simplified by assuming a stationary state of growing chains and neglecting deactivation of active sites. The molecular weight distribution was developed on the basis of these assumptions. This model shows that the instantaneous polydispersity for each active site is equal to 2, and the instantaneous polydispersity for the entire amount of polymer produced would deviate from 2 by the ratio of the variance of the propagation rate constants to the mean propagation rate constant. The effect of the propagation rate constant distribution on molecular weight development was demonstrated by assuming unimodal and skewed bimodal distributions of propagation rate constants. The pseudo-kinetic rate constant method was used to evaluate the kinetic rate constants for each site. However, chain- transfer reactions were not considered in the active center fraction calculations. Effects of chain-transfer reactions on active center fraction calculations can be very important if chain-transfer reactions are significant (Xie and Hamielec, 1993a-c). de Carvalho et al. (1990) recently extended their model to calculate both the stereo and chemical sequence distributions of copolymer. The models developed by de Carvalho et al. (1989,1990) have not been evaluated using experimental data.

Villermauax et al. (1989) and Lorenzini et al. (1991) developed a model for ethylene and a-olefin copolymerization with Ziegler-Natta catalyst a t high temperature and high pressure. The model development is based on a single active site, which is then extended to two or three active sites using a mixing law. This model can fit molecular weight averages and distribution data over the range of the experimental conditions. The kinetic parameters were estimated by fitting the model to experimental data (Lorenzini et al., 1991).

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Gas Phase Ethylene Polymerization-Production Processes, Polymer Properties, And Reactor Modeling

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