Production of Fluoroelastomers

4 Production of Fluoroelastomers � 4.1 � Introduction This chapter covers various aspects of the production of fluoroelastomer copolymers. After a general description of commercial production processes, free radical copolymerization and emulsion polymerization kinetics are described. Process variations designed to attain desired characteristics of major product families are summarized, with more detail covered in Chs. 5 and 6 on cure systems and processing of the various fluoroelastomer composition families. Other process steps, such as monomer recovery and polymer isolation, are described, along with process safety considerations. Finally, commercial processes are covered in detail.

4.2 � General Process Description Most commercial fluoroelastomers are copolymers of two or more monomers made by free-radical emulsion polymerization. Figure 4.1 is a schematic of the general process. The polymerization operation may be carried out in continuous or semibatch mode. Numerous process variations are used to produce different products. Molecular structures of fluoroelastomers are determined by polymerization and isolation process conditions, so product and process

Figure 4.1 General fluoroelastomer production process. ­

development are usually carried out simultaneously in laboratory semiworks units designed to emulate commercial operation. As indicated in Fig. 4.1, water and other liquid ingredients are added to the polymerization reactor. These include an initiator and soap as aqueous solutions, and an optional chain-transfer agent and curesite monomer. Two or three major monomers are fed as gases by a compressor. The reactor is maintained at the temperature, pressure, and holdup time required for the particular product. Air and other impurities are carefully excluded from the feed and reactor systems. Polymer is formed in the reactor as a dispersion containing 15%–30% solids, with particle size generally in the range 100–1000 nm diameter. At reactor conditions, much of the monomer present is dissolved in the particles at concentrations of 3%–30%, depending on polymer and monomer compositions, and on prevailing temperature and pressure. The polymer dispersion is discharged from the reactor to a degassing vessel maintained at low pressure to allow removal of residual gaseous monomer. In continuous reactor operation, the reaction vessel is maintained liquid-full and the dispersion is let down through a back-pressure control valve to the degasser. Recovered monomer is recycled continuously to the reactor through the monomer feed compressor. In semibatch reactor operation, the dispersion is let down to the degasser at the end of the

38 polymerization, and recovered monomer is held for subsequent recharging of the reactor for succeeding batches of the same composition. Additional vessels may be provided for final monomer removal and dispersion blending prior to isolation. Polymer isolation is effected by chemical coagulation of the dispersion, followed by separation of polymer crumb from the aqueous phase, removal of soluble soap and salt residues, and dewatering and drying of the polymer. Usual coagulants are soluble salts of aluminum, calcium, or magnesium. Various means of separating polymer from the coagulated slurry are used commercially, including continuous centrifuges, filters, and dewatering extruders. Methods used for salt removal include washing by repeated reslurrying in fresh water and separation of polymer; washing on a batch filter or continuous filter belt; or expelling most of the aqueous phase in a dewatering extruder. The purified polymer is dried in a batch oven or continuous conveyor dryer, or in a drying extruder. The isolated fluoroelastomer is generally formed into pellets or sheet for packaging and sale as gum polymer. Alternatively, the polymer may be precompounded by adding curatives and processing aids before forming and packaging.

FLUOROELASTOMERS HANDBOOK Chain radical reactivity is assumed independent of radical size, and depends only on the reactivity of the last unit added to the chain. Chain length is long, so monomer consumption is assumed to occur only by propagation. A stationary state is assumed with respect to radical concentrations. That is, the rate of change of radical concentration is negligible compared to the rate of polymerization. The following sections discuss the general free radical reaction scheme, followed by some aspects of relative monomer reactivity important in copolymerization of fluoroelastomers.

4.3.1

General Reaction Scheme

The steps in free radical polymerization are depicted in simple form below, together with individual rate expressions involved. Initiation: � Eq. (4.1) I → 2R· R· + M → R1·

2f k d[I] �

Propagation: � Eq. (4.2) Rn · + M → Rn+1·

kp[M][R·]

Transfer: � Eq. (4.3) Rr· + T → Pr + T·

ktr[T][R·]

4.3 � Free Radical Copolymerization

Termination: �

Free radical polymerization involves four types of reactions: initiation, propagation, transfer, and termination. Initiation includes generation of free radicals, moieties with free valences which are highly reactive, followed by addition of vinyl monomer units. The double bonds of the monomer open to form growing radical ends. Propagation is a relatively rapid process, with continued addition of monomer units to growing radical chains. Growth of a chain may be stopped by a transfer reaction in which the radical is capped by a reactive atom such as hydrogen or halogen and the radical activity is transferred to the residue of the transfer agent. This radical may add monomer to continue growth of the kinetic chain. Termination occurs by reaction of radicals to form dead chains that do not propagate further. Kinetic analysis is considerably simplified by making a number of assumptions that are good approximations in practical polymerization systems.

In the scheme above, initiation takes place by thermal decomposition of an initiator I, followed by addition of the first monomer unit at efficiency f. Propagation takes place quickly with addition of many monomer units. The rate coefficient kp is independent of radical chain length, but dependent on the nature of the radical end and the monomer (more detail is contained in the following section (Sec. 4.3.2). Transfer involves the transfer of a reactive atom such as hydrogen or halogen (usually from the transfer agent to cap the radical end) and transfer of the radical activity to the transfer agent residue. The new radical usually adds monomer to continue propagation of the kinetic chain. If the transfer radical has low reactivity toward propagation, it slows down the polymerization, acting as a retarder or inhibitor. Termination by combination of radicals is assumed, valid for most fluorocarbon polymer systems.

Eq. (4.4) Rr· + Rs· → Pr+s

2kt[R·]2 �

4 PRODUCTION OF FLUOROELASTOMERS

39

Kinetics relationships based on the reactions above must take into account the nature of the polymerization system. In suspension polymerization of fluoroelastomers, all the reactions occur in relatively large particles swollen with monomer. Emulsion systems are more complicated; reactions in both the aqueous phase and relatively small monomer-swollen particles must be considered. Initiation and propagation of short radical chains take place in the aqueous phase, along with termination and transfer reactions. A fraction of the short-chain radicals enter particles, where they undergo propagation to high molecular weight and also undergo transfer and termination reactions to form dead polymer chains. Rates in particles (either suspension or emulsion cases) may be reduced by the high-viscosity environment that reduces reactant mobility. Termination rates may be drastically reduced, since diffusion of chain radicals is greatly hindered.

Mayo,[3] Simha,[4] and Wall,[5] by making the further assumption that steady state applies to each radical type separately. This means that the rate of conversion of radical M1· to M2· (Eq. 4.5b) is balanced by the reverse conversion (Eq. 4.5d):

4.3.2 �

The ratio of monomers incorporated in the polymer is obtained by dividing Eqs. (4.8) by (4.9) and substituting for the ratio of radical types in Eq. (4.7). The resulting copolymer composition relationship is simplified by denoting the ratio of monomers in the polymer rp1/rp2 as Y and the ratio of monomer concentrations [M1]/[M2] as X, and defining the monomer reactivity ratios r1 = k11/k12 and r2 = k22/k21:

Copolymer Composition Relationships

Since fluoroelastomers are copolymers of two or more monomers, an understanding of the relationship between polymer composition and monomer ratios in the polymerization system is necessary for successful control. Composition relationships for dipolymers were derived by early workers in the polymer field. Using the assumption that the rate of monomer addition to a radical chain depends only on the nature of the last unit on the chain, Dostal,[1] in 1936, showed that only four propagation reactions and corresponding rates would describe copolymerization of two monomers: Eq. (4.5a)

M1· + M1 → M1·

k11[M1·][M1]

Eq. (4.5b)

M1· + M2 → M2·

k12[M1·][M2]

Eq. (4.5c)

M2· + M2 → M2·

k22[M2·][M2]

Eq. (4.5d)

M2· + M1 → M1·

k21[M2·][M1]

In 1944, a useful copolymer composition relationship was derived independently by Alfrey,[2]

Eq. (4.6)

k12[M1·][M2] = k21[M2·][M1]

Then the ratio of the radical types is: Eq. (4.7)

[M1 ·] = k21 [M1 ] [M 2 ·] k ­12 [M 2 ]

The rates of incorporation of each monomer into polymer are given by: Eq. (4.8)

rp1 = k11[M1·][M1] + k21[M2·][M1]

Eq. (4.9)

rp2 = k12[M1·][M2] + k22[M2·][M2]

Eq. (4.10)

Y=X

(r1 X +1) (r2 + X )

This copolymer relationship can be applied directly to fluoroelastomers containing two major monomers. Under polymerization conditions used for commercial production, whether continuous or semibatch operation, compositions of polymer and unreacted monomer (thus Y and X ) are held constant by continuous feed of monomer to the reactor. This allows estimation of reactivity ratios from carefully designed experiments. Composition relationships for systems of more than two monomers are much more complex.[6] However, the compositions of terpolymers and tetrapolymers were shown to be functions of the reactivity ratios of the various pairs of monomers in the system. This is helpful in determining characteristics such as monomer sequencing in such polymers.

40

FLUOROELASTOMERS HANDBOOK

4.3.3

Monomer Reactivity Ratios

To facilitate evaluation of reactivity ratios for copolymers with corresponding Y and X ratios determined over a range of compositions, Eq. (4.10) can be rearranged into several forms. Reactivity ratios r1 and r2 can be determined as the slope and intercept of a linear plot of the composition relationship in one of the following forms: Eq. (4.11)

Eq. (4.13)

X

r1 =

X

X (Y −1) X2 = (r1 ) − r2 Y Y Each polymerization experiment with reliable values of Y and X gives a point on the line, and regression analysis (least squares) can be applied to obtain the reactivity ratios. Alternatively, Eq. (4.10) can be solved for one of the reactivity ratios to obtain the following relationships: r1 = (r2 )

(Y −1) Y + 2 X X

r2 = (r1 )

X 2 X (Y −1) − Y Y

Y = r1X + 1

or

(Y −1) = r − (r ) Y 1 2 2

or

Eq. (4.12)

simplified further. In VDF/HFP copolymerization, the reactivity ratio r2 for HFP can be taken as zero. To a close approximation, HFP does not add to HFP· radical ends in the presence of VDF. The copolymer relationship for r2 = 0 becomes:

or

Each experiment gives a straight line in a plot of r1 versus r2. The reactivity ratios can then be estimated from the intersections of a number of lines from individual experiments. The extent of the region covered by intersections gives a visual idea of the errors in the reactivity ratio estimates. For important dipolymer fluoroelastomer families, the copolymer composition in Eq. (4.10) can be

(Y −1) X

Little polymerization data over a range of VDF/ HFP compositions have been published to allow good estimates of reactivity ratios. Some patent examples provide useful information, such as Example 1 of Moore and Tang, U.S. Patent 3,929,934.[7] In this example, the first of two reactors in series was operated in continuous mode to make VDF/HFP copolymer at high conversion (93%). At this conversion and the prevailing conditions of temperature (110°C) and pressure (6.2 MPa), all the unreacted monomer can be assumed to be dissolved in the polymer particles. In this continuous emulsion polymerization, a 2-liter reactor was fed with 8.0 L/h water (nominal residence time 0.25 hour) containing 16 g/h ammonium persulfate initiator and 3.0 g/h NaOH for pH control. The effluent polymer dispersion contained about 19% solids; the polymer composition was about 58% VDF and 42% HFP. A mass balance on monomer allows calculation of the ratios Y and X as shown in Table 4.1. From unreacted monomer composition, X = 0.36, and from the corresponding polymer composition, Y = 3.3, so that r1 = 6 from Eq. (4.13). This is a reasonable value, but should be considered as only a rough estimate (say, within ± 50%). From composition ranges noted in patents, perfluoroelastomers with major monomers TFE and PMVE must contain about 33 mole % PMVE to be

Table 4.1 Mass Balance on VDF/HFP Monomer

Monomer: VDF (1)

Feed g/h 1100

Monomer: HFP (2)

900

130

0.87

770

5.13

2000

150

1.18

1850

22.01

Monomer

Total:

Unreacted g/h mol/h 20 0.31

g/h 1080

Polymer mol/h 16.88

wt % 58 42

4 PRODUCTION OF FLUOROELASTOMERS

41

well in the amorphous composition range with negligible crystallinity from long runs of TFE units. This is in contrast to the VDF/HFP system, where only about 20 mole % HFP is sufficient to avoid crystallinity. Thus, it appears that there is significant clumping of adjacent PMVE units, allowing formation of relatively long runs of TFE units in TFE/PMVE copolymers. Then the propagation rate coefficients k11 and k22 and the reactivity ratios r1 and r2 for TFE and PMVE, respectively, must be greater than zero. Published data on TFE/PMVE polymerization are insufficient for calculation of reactivity ratios. However, in a subsequent section on monomer sequencing, it will be shown that a reactivity ratio product r1r2 ~ 0.5 is reasonable for this system. Substitution of r1/0.5 for r2 in the copolymer composition Eq. (4.10) yields a quadratic equation that can be solved for the reactivity ratio r1: Eq. (4.14)

(

Y −1+ 1 +Y 2 r1 = 2X

)

½

Example 1 of Apotheker and Krusic, U.S. Patent 4,035,565,[8] provides polymerization data on a terpolymer of TFE and PMVE with a small amount of cure-site monomer bromotrifluoroethylene (BTFE). Continuous emulsion polymerization was carried out in a liquid-full 3.8-liter reactor at 70°C and 4.1 MPa with about 2.7 hours residence time based on latex flow. Redox initiator components ammonium persulfate (6.38 g/h) and sodium sulfite (5.25 g/h), buffer dibasic sodium phosphate heptahydrate (4.5 g/h), and soap ammonium perfluorooctanoate (12.0 g/h) were fed in 1.2 L/h total water. A mass balance on monomer was obtained at steady state; the results are shown in Table 4.2. Ignoring BTFE in the calculation of the TFE/ PMVE monomer ratios, X = 0.20 and Y = 2.24. From

Eq. (4.14), reactivity ratios are approximately r1 ~ 9 and r2 ~ 0.06 under the assumption r1r2 ~ 0.5. For TFE/propylene copolymerization, both reactivity ratios are near zero, so the copolymer composition relationship reduces to Y = 1 at all values of X. Thus TFE and propylene units would alternate, no matter what the monomer ratio in the reactor. This assumption is not quite correct, since typical commercial TFE/P elastomers contain some 52–54 mole % TFE. It appears that with TFE in great excess, say X > 10, Y is ~ 1.1, indicating r1 ~ 0.01 for TFE. The monomer pair VDF and TFE appears to approximate the case r1r2 = 1. Substitution of 1/r1 for r2 in the copolymer composition equation (Eq. 4.10) leads to Y = r1 X, indicating that the monomer ratio in the polymer is directly proportional to the ratio of unreacted monomer, so-called random copolymerization. VDF/TFE copolymers are crystalline plastics with melting points varying with composition. Terpolymers with relatively high TFE content (45%–65%) and low HFP content (15%–20%) are sold by Dyneon as flexible thermoplastics with melting points 120°C–180°C. In elastomeric terpolymers of VDF and TFE with HFP or PMVE, the reactivity ratios of HFP or PMVE with respect to either VDF or TFE can be considered as near zero. Thus units of HFP or PMVE are usually isolated between mixed sequences of VDF and TFE units.

4.4

Emulsion Polymerization

Essentially all fluorocarbon elastomers are produced commercially by emulsion polymerization, depicted schematically in Fig. 4.2. As previously described, polymerization occurs in monomer-swollen polymer particles some 100 to 1000 nanometers (nm) in diameter, not in a liquid-liquid emulsion as implied by the name. Particles are stabilized by surfactant,

Table 4.2 Mass Balance on TFE/PMVE Monomer with a Small Amount of BTFE Cure-site Monomer ­

Monomer Monomer: TFE (1)

Feed g/h 260

Unreacted g/h mol/h 14 0.14

g/h 246 183

1.10

41.8

30.4

2.0

1.6

0.71

Polymer mol/h wt % 2.46 56.2

Monomer: PMVE (2)

300

117

Monomer: BTFE

10

1

9

0.06

Total

570

132

438

3.62

mol % 68.0

42

FLUOROELASTOMERS HANDBOOK Eq. (4.15)–O3SO– OSO3– → 2 •OSO3– The rate of decomposition is determined mainly by temperature, but is also somewhat dependent on pH. Fluoroelastomer polymerization is usually carried out at relatively low pH (~ 3–6), and the first order thermal decomposition rate coefficient kd (min-1) is given in the Arrhenius form as[10] Eq. (4.16)

Figure 4.2 Emulsion polymerization.

either added or made in situ by polymerization in the aqueous phase. A water-soluble initiator system generates free radicals, some of which grow and form or enter particles. In most fluoroelastomer polymerization systems, there is no sizeable reservoir of liquid monomer present. Much of the monomer is dissolved in the polymer particles, and is replenished by a continuous feed during the polymerization. Even in semibatch polymerization, the amount of monomer in the reactor vapor space is relatively small. The segregation of growing radicals in small particles under conditions of limited termination by incoming radicals allows attainment of the high molecular weights desired for good elastomeric properties. Especially for VDF-based fluoroelastomers, emulsion systems allow very high productivity in reactors of modest size.

E  k d = Aexp  a   RT   −17070  = 5.62 ×1018 exp    T 

In this equation, the factor A is in units min-1, activation energy Ea is in cal/mole, gas constant R = 1.987 cal/mole K, and absolute temperature T is in kelvin. With the high activation energy Ea = –33,900 cal/ mol, the rate of decomposition of persulfate is quite sensitive to temperature (Table 4.3). At temperatures below about 80°C, persulfate decomposition is slow, so relatively high initiator concentrations would be necessary to get reasonable radical generation rates. Instead various redox initiator systems may be used. Sulfite is a typical reducing agent that reacts rapidly with persulfate to generate two types of radicals: Eq. (4.17) �

–O

3SO–OSO3

–

+ SO32–

→ SO42– + •SO3– + •OSO3–

4.4.1 �

Emulsion Polymerization Kinetics

Table 4.3 Thermal Decomposition of Persulfate[10]

103 kd, min-1

Recent work by R. G. Gilbert[9] and coworkers at Sydney University and DuPont has greatly clarified the various complicated steps involved in emulsion polymerization, allowing development of improved kinetics models and setting of conditions to get polymer structures desired for commercial applications.

Temperature, °C

Polymerization mechanism. In emulsion polymerization, a water-soluble initiator system forms free radicals in the aqueous phase. Typically, thermal decomposition of persulfate is used to generate radicals by symmetrical scission of the O– O bond of the anion:

90

21.2

33

100

4.9

9

Half life, min

50

0.063

11,000 (184 h)

60

0.307

2,260 (38 h)

70

1.37

507 (8.4 h)

80

5.60

124 (2.0 h)

110

247

3

120

769

1

4 PRODUCTION OF FLUOROELASTOMERS At temperatures below 60°C, a small amount of a catalyst such as a copper salt may be added to increase the redox rate. In continuous polymerization, the components of the redox initiator system are fed to the reactor in separate streams. In semibatch polymerization at low temperature, persulfate is usually charged initially, and the reducing agent is added at a controlled rate to get the desired radical flux, twice the molar addition rate of sulfite times an efficiency factor. A considerable fraction of primary radicals may be lost by recombination with each other before monomer addition occurs to complete the initiation process. For VDF/TFE/HFP or PMVE elastomers, VDF and TFE are the most likely monomers to add to primary initiator radicals. For initiation with sulfate ion radicals, the following reactions would be typical: Eq. (4.18)

CH2=CF2 + •OSO3– → •CF2– CH2– OSO3–

Eq. (4.19)

CF2=CF2 + •OSO3– → •CF2– CF2–OSO3–

Perfluorinated sulfate end groups are likely to hydrolyze to carboxylate ends at polymerization conditions: Eq. (4.20)

•~CF2– CF2– OSO3– + 2 H2O → •~CF2– COO– + H2SO4 + 2 HF

The generation of sulfuric and hydrofluoric acids by hydrolysis reactions usually necessitates addition of a base or buffer to keep the pH above 3. Initiation with •SO3– leads to formation of sulfonate end groups, e.g.: Eq. (4.21)

CF2=CF2 + •SO3– → •CF2– CF2– SO3–

Unlike perfluorinated sulfate end groups, perfluorinated sulfonate ends are resistant to hydrolysis. These small radicals propagate further in the aqueous phase, reacting with the small amount of dissolved monomer present. Because polymer particles are stabilized by adsorbed anionic surfactant

43 and also carry surface charge from ionic end groups of polymer, growing radicals in the aqueous phase must add several monomer units (say, 3 to 5) to become surface active and hydrophobic enough to overcome the electrostatic surface barriers and enter particles. With this delay in entry, small radicals may undergo other reactions (e.g., termination reactions) such as: Eq. (4.22) •CF2– CH2– CF2– CH2– OSO3– + •CF2– CH2– CF2–CH2–CF2–CH2–OSO3– → –O3SO– (CH2– CF2)2– (CF2CH2)3– OSO3– The resulting combination products may serve as effective surfactants to stabilize particles. Depending on initiator level, little or no added surfactant may be necessary for adequate stabilization of dispersions of VDF copolymers. Note that if, instead of VDF or TFE, other less reactive monomer units such as HFP or PMVE add to these short radicals, termination becomes more likely than further propagation. Transfer to an active water-soluble species may result in a nonionic radical (e.g., transfer with isopropyl alcohol): Eq. (4.23) •CF2– CH2– OSO3– + (CH3)2CHOH → HCF2– CH2– OSO3– + (CH3)2C •OH Presumably, such a polar, uncharged radical would need to add only one or two monomer units to become hydrophobic enough to enter a particle. With all these possible aqueous-phase reactions, the fraction of primary radicals generated that grow and enter particles to continue growth to high polymer may be rather low (0.2–0.6), especially at high initiator levels. Once a radical enters a particle, it propagates rapidly by addition of monomer dissolved in the particle at much higher concentration than present in the aqueous phase. In some systems (e.g., TFE/ PMVE), high added soap levels give a large number of small particles (say, 200 nm in diameter) highly swollen with 20%–30% monomer. It is likely that the ideal emulsion 0,1 polymerization case prevails

44 here. That is, each particle contains only one or no growing radical at any time. No more than one radical at a time propagates in such a small particle; entry of a second radical leads to rapid termination, with no radical activity until another radical enters to restart polymerization. At the other extreme are most VDF copolymer systems. Here, very low added soap levels give a small population of large particles (500–1000 nm diameter) with relatively low concentration of dissolved monomer (about 10%). In these highly viscous particles, the mobility of long-chain radicals is so low that termination rates are drastically reduced. Only very small radicals entering from the aqueous phase or formed by transfer reactions are effective in terminating long chains. A large number of growing radicals (10 or more) may coexist in such large particles. Radical lifetimes may be quite high in such systems, so polymerization rates and molecular weights may be high. Typically, transfer agents are used to control molecular weights. Soaps used in the emulsion polymerization of fluoroelastomers are usually fully or partially fluorinated anionic surfactants. Efficacy at low concentrations and high water solubility are desirable to get low residual soap levels in isolated polymer. Soaps should be unreactive to radicals at polymerization conditions, to avoid excessive transfer and attachment of ionic soap moieties to polymer chain ends. Perfluoroalkyl carboxylates or sulfonates with 8- or 9-carbon alkyl chain lengths are inert and effective dispersion stabilizing agents. Ammonium perfluorooctanoate has been preferred for many fluoroelastomer emulsion systems. However, this stable soap is persistent in the environment and is not readily eliminated from the body after exposure. A major supplier (3M Co.) has stopped its production, and its use in fluoropolymer production is being phased out or reduced. A number of partially fluorinated soaps are effective, especially for VDF copolymers. These are usually of the structure F– (CF2– CF2)n–CH2– CH2– X–M+, with n = 2–8 (mostly 3–4); – X– may be sulfate, phosphate, or sulfonate, and M+ is H+, NH4+, or an alkali metal ion. The sulfate and phosphate forms are highly effective,[11] but may participate in unwanted transfer reactions. Recently, a particular partially fluorinated alkyl sulfonate form, F–(CF2– CF2)3– CH2– CH2– SO3–Na+, has been found to be a good replacement for ammonium perfluorooctanoate in many fluoroelastomer emulsion polymerization systems, both semibatch and con-

FLUOROELASTOMERS HANDBOOK tinuous.[12] This soap is effective as a dispersion stabilizer, inert to radical attack by transfer, and readily removed during polymer isolation. Polymerization rate Rp in an emulsion system can be represented as: Eq. (4.24)

Rp =

k p [M ]N p nr M o NA

In this relationship, kp is the overall propagation rate coefficient in the particles; [M] is the molar monomer concentration in a particle; Np is the total number of particles; nr is the average number of radicals per particle; Mo is the average monomer molecular weight; and NA is Avogadro’s number (6.022 × 1023). Note that kp[M ]Mo may be applied to a copolymer using average values for a particular composition. For most systems of interest, available data are insufficient to evaluate key parameters in the rate expression. The number of particles and the average number of radicals per particle are particularly difficult to determine. Factors that affect the number of particles are considered in the next section on particle nucleation. Particle formation mechanisms. Most studies of emulsion polymerization are based on batch polymerization of a liquid monomer, so particle formation and growth are treated as occurring in three distinct intervals:[13] I. Particle nucleation period: characterized by presence of monomer droplets and soap micelles, with formation of particles that grow in number and size. Polymerization rate Rp increases. II. Particle growth period: monomer droplets � are present, but no micelles; particle number is constant, particle size grows. Rp is � steady or increases. � III. Final stage: monomer is consumed, so Rp decreases with decreasing monomer concentration in particles. This is not an adequate picture of the fluoroelastomer emulsion polymerization processes. These are semibatch or continuous polymerizations in which monomer composition and concentration in particles are kept essentially constant by continuous feed of monomer to the reactor. No large reservoir of reactive monomer is present. Thus, only modified forms of

4 PRODUCTION OF FLUOROELASTOMERS intervals I and II exist in semibatch polymerization, and nucleation and growth stages coexist in continuous polymerization. For most fluoroelastomer systems, polymerization ceases when fresh monomer feed is stopped. The composition of unreacted monomers in the reactor, while necessary to set the copolymer composition, is not usually reactive enough toward propagation to support appreciable polymerization rates. In his analysis of particle formation, Gilbert[14] notes that a short radical formed from initiator with monomer addition must meet one of three fates: aqueous-phase termination, entry into a particle, or forming a new particle. Entry into a particle can occur when a sufficient degree of polymerization (number of monomer units added to an initiator fragment) denoted as z is reached so that the radical becomes surface active. Particle formation can occur when such a z-mer enters a micelle or when the radical grows to a sufficient longer degree of polymerization, jcrit , to homogeneously precipitate and nucleate to a precursor particle. Particle formation ceases when the number and size of particles reach levels such that all z-mer radicals are captured. Thus, two mechanisms of particle formation may occur: homogeneous nucleation in systems with soap levels below the critical micelle concentration (cmc), and micellar entry in systems above the cmc. Both particle formation mechanisms must be considered in fluoroelastomer polymerization systems. For VDF copolymers, the polymerizations are

45 characterized by low surfactant levels and high propagation rates at low monomer concentrations. In these systems, particle formation is by homogeneous nucleation. For most TFE copolymers containing no VDF (e.g., TFE/PMVE, TFE/P), soap levels are very high and propagation rates are low. Micellar entry may prevail as the major mode particle formation in these systems. Particle formation by homogeneous nucleation. Figure 4.3 illustrates steps in particle formation by homogeneous nucleation and coagulation as described by Gilbert.[15] After initiation by addition of a monomer unit to a primary ionic radical, the small radical may propagate in the aqueous phase. With ionic head groups such as sulfate, sulfonate, or carboxylate, these radicals are soluble in water when only a few monomer units (say, 1–3) have been added. Many are lost by mutual termination, becoming dead chains with one or two ionic head groups. Depending on their size, these may serve as surfactants to stabilize particles. When a sufficient number of units, z, have been added, a growing z-mer becomes surface active so that it can overcome the electrostatic surface barrier and enter a particle. A radical that propagates further to a critical length, jcrit , becomes insoluble in water, so it may coil and precipitate to form a precursor particle. Monomers enter such precursor particles, so the radicals may continue to grow. Precursor particles grow both by propagation and by co-

Figure 4.3 Particle formation by homogeneous-coagulative nucleation.[15]

46 agulation with other precursor particles, eventually becoming mature particles. Particles are stabilized by ionic end groups on their surfaces. These are from a combination of added surfactant, oligomeric surfactant formed in situ by aqueous phase termination of short radicals, and polymeric chains anchored in the particles. Ordinarily, the total of such ionic end groups is relatively low, since little or no soap is added and initiator levels may be low. Such dispersions usually have low particle number and large particle size. The first quantitative model of homogeneous nucleation was developed by Fitch and Tsai[16] and augmented by Hansen and Ugelstad[17] as HUFT theory. Coagulation of small particles was taken into account by Richards, Congalidis, and Gilbert,[18] using an extension of the standard DLVO model of colloid science.[19] This describes the coagulation of small particles stabilized by surface charge. Later versions of the model take better account of the variation of the number of particles with ionic strength.[20] Richards and Congalidis have developed proprietary models applicable to various DuPont products, including fluoroelastomers. Gilbert gives the general approach to formulation of such complicated models.[21] Particle formation by micellar entry. Figure 4.4 illustrates the micellar entry mechanism for particle formation, as described by Gilbert.[22] This mechanism is likely to prevail in systems with levels of added surfactant significantly higher than the critical micelle concentration. The initial steps are simi-

Figure 4.4 Particle formation by micellar entry.[22]

FLUOROELASTOMERS HANDBOOK lar to those for homogeneous nucleation, with formation of small radicals from initiator and monomer addition in the aqueous phase. When radicals reach the z-mer stage, they are readily incorporated into micelles. Even though a micelle is a dynamic moiety, with individual surfactant molecules residing in the micelle for only a short time, the micelle protects a growing radical from rapid termination with radicals in the aqueous phase. Also, the micelles solubilize monomers to facilitate rapid chain growth. Few radicals in the aqueous phase grow much longer than z-mer length before capture when micelles are present. Thus, homogeneous nucleation is unlikely when soap concentrations are well above the cmc. Micelles grow into mature polymer particles by radical propagation and coagulation with other particles. As the particle population and size grow, the surface area may become large enough to adsorb enough surfactant that the aqueous concentration falls below the cmc and micelles disappear. No new particles then can form by micellar entry, but homogeneous nucleation could occur when surfactant concentration falls to or below the cmc in the aqueous phase. Models that allow for both homogeneous nucleation and micellar entry generally predict a low particle number for soap concentrations below the cmc, then a large increase in the particle number as soap concentration is raised to and just above the cmc. However, most experimental data indicate a more gradual change in particle as the soap concentration is increased through the cmc. This suggests that,

4 PRODUCTION OF FLUOROELASTOMERS with homogeneous nucleation below the cmc, particles are more effectively stabilized against coagulation as surfactant concentration increases. Thus the particle number becomes relatively high as soap concentration is raised, well before the cmc is reached. Secondary nucleation. So-called secondary nucleation involves the formation of new particles in the presence of an established population of seed or previously formed particles. Gilbert[23] gives criteria for this situation. The number and size (thus the surface area) of established particles must be low enough that aqueous-phase radicals have sufficient probability of growing beyond the z-mer stage to jcrit size for nucleation. This case is important in continuous emulsion polymerization systems, which require continuing formation of new particles in the presence of a large number of particles present in the dispersion. Kinetics relationships. Complex models, such as those described by Richards, Congalidis, and Gilbert in Reference 20, can represent emulsion polymerization systems well, allowing extrapolation to conditions outside the range of available experimental data and providing insight into the effects of changing reaction variables. However, such models require considerable physical and kinetic data (e.g., solubilities of monomers in copolymers of varying composition over a range of temperatures and pressures, individual propagation rate coefficients and values of z and jcrit for oligomer entry and coagulation into particles). Such information has been obtained for only a few copolymers. Some parameter adjustment is usually necessary to fit experimental data. A particular difficulty seems to be prediction of both number-average molecular weight and particle number in a given system. It should be noted that both of these parameters are difficult to measure accurately so experimental error contributes considerably to differences from model predictions. In spite of the limited scope of the models developed so far, the models have helped workers understand the behavior of more complex terpolymer and tetrapolymer systems. Bonardelli, Moggi, and Russo[24] studied particle formation in the soapless emulsion polymerization of vinylidene fluoride (VDF) and hexafluoropropylene (HFP). Semibatch polymerizations were carried out in a five-liter reactor charged with 3.5 liters of water, using ammonium persulfate as initiator at

47 85°C with no added soap. Copolymer composition was held constant at a molar ratio VDF/HFP = 79/ 21, the same as most commercial dipolymers, by feeding this monomer mixture during polymerization. Reaction was stopped at 400 grams of polymer per liter of water (29% solids). Monomer concentration and initiator levels were varied in the study. Dispersion samples were taken during the polymerization for measurement of particle size by laser light scattering; the number of particles was calculated from average particle volume and total polymer formed. Experimental results were interpreted using Eq. (4.24), treating the constant copolymer composition as if it were a homopolymer. Monomer concentration [M] was expressed in terms of the product of average monomer fugacity fM and Henry’s Law constant H, so the rate equation becomes Eq. (4.25)

Rp =

k p Hf M M o N p nr NA

It should be noted that the concentration of monomer in the aqueous phase as well as that in the polymer particles varies with monomer fugacity (or total monomer pressure). Thus, the aqueous oligomeric radical growth and polymerization rate in the particles are both affected by varying monomer fugacity. For this copolymerization system, Bonardelli and coworkers[24] observed very long nucleation periods, with the number of particles Np increasing up to about 200 grams polymer/liter (17% solids). The polymerization rate also increased during this period, corresponding to interval I. Nucleation periods were longer in experiments with lower monomer fugacity or higher initiator level. Even during interval II, when Np and Rp are essentially constant, further particle formation may occur if balanced by particle agglomeration. The polymerization rate Rp and final particle number Np in interval II varied about as expected with initiator concentration [I], with Rp ∝ [I]0.6 and Np ∝ [I]0.4. Variation with monomer fugacity was somewhat more difficult to explain, with Rp approximately second order in fugacity and final Np varying inversely with monomer fugacity. Bonardelli and coworkers account for this by noting that monomer concentration affects particle size and the number of radicals per particle nr. They rearrange Eq. (4.25) into the form

48

Eq. (4.26)

FLUOROELASTOMERS HANDBOOK Rp N A Np fM

= k p HM o nr

This states that the polymerization rate per mole of particles divided by monomer fugacity is proportional to the number of radicals per particle. A plot of the left-hand side of Eq. (4.26) versus particle size indicates that nr is low and nearly constant at small particle size, but increases greatly with size at particle diameters above about 260 nm. At small sizes, particles would be expected to have either one or no radicals present because of rapid termination by incoming radicals, giving a 0,1 system with nr = 0.5. The more usual case for commercial VDF/HFP copolymerization with no soap (or low soap) is to have relative large particle sizes in the range 400 to 900 nm, and thus many radicals per particle. Consequently, we would expect a strong dependence of polymerization rate on monomer concentration in such systems. Also, initiation levels play a considerable role in determining rate and particle number. The hindered termination in large particles leads to significant broadening of molecular weight distribution in the absence of transfer agents. For production of commercial fluoroelastomers, empirical relationships are usually applied to estimate polymerization rates and to set and control polymer viscosities since most parameters in fundamental kinetics models are not known for most compositions. Polymerization rates may be correlated by equations of the form Eq. (4.27)

Rp = kp fMq ρ r(1 + S s)

Such equations may be applicable to VDF copolymerization with soap added at low concentration S. Monomer concentration may be represented by partial pressure or fugacity fM. An overall radical generation rate at 100% efficiency ρ is used, and an overall polymerization rate coefficient kp for the particular copolymer composition and reaction temperature. The usual ranges for the exponents are: q ~ 1–2, r ~ 0.5–0.7, and s ~ 0.4. It may also be necessary to incorporate additional factors to account for the effects of polymer concentration in the dispersion, or alternatively, for reaction time in a semibatch reactor or residence time in a continuous reactor. Ordinarily, Rp is known from experience and a commercial reactor is run at the same rate and

other conditions for a given product. Operating rates are often not set by kinetics, but are limited at lower levels because of other plant design constraints considered in Secs. 4.4.2 and 4.4.3. Relationships showing the dependence of polymer molecular weight or viscosity on reaction variables are of more use in setting and controlling fluoroelastomer properties. Number-average molecular weight Mn can be expressed as the ratio of polymerization rate (Rp g/h) to rate of chain formation (mol/h). For most fluoroelastomer emulsion systems, long chains are started and stopped by radical entry into particles (rate ρe), or by reactions with an added transfer agent (rate rtr). Ordinarily, transfer reactions with monomer, polymer, initiator, or adventitious impurities are negligible. Eq. (4.28)

Mn =

Rp  ρe   + rtr   2 

A more convenient measure of molecular weight than Mn is desirable for routine monitoring of product. In most situations, molecular weight distribution is reasonably constant, and thus Mv/Mn can be assumed constant. For a given polymer composition and solvent, the limiting viscosity number or intrinsic viscosity [η] is related to viscosity-average molecular weight Mv by the Mark-Houwink equation:[25] Eq. (4.29)

[η] = K´ Mvα

For commercial VDF copolymers in a good solvent such as methyl ethyl ketone, the exponent α is in the range 0.55–0.75. A good approximation to [η] is the inherent viscosity or logarithmic viscosity number: Eq. (4.30)

ηinh = (ln ηr)/c

The relative viscosity ηr is measured as the ratio of solvent to solution efflux times in a capillary viscometer, with solution concentration c = 0.1 g/dL. The overall radical generation rate at 100% efficiency ρ is used instead of radical entry rate, so an empirical relationship for inherent viscosity then becomes

Eq. (4.31)

ηinh

Rp   =   K (ρ + 2rtr )

a

4 PRODUCTION OF FLUOROELASTOMERS

49

The parameters K and a can be determined for a given polymer composition by making a number of experimental polymerization runs at varying initiator levels without any transfer agents present. Equation (4.31) simplifies to

Eq. (4.32)

ηinh

 Rp   =   Kρ 

a

For analysis of experimental data, this can be put in the form Eq. (4.33)

 Rp  log ηinh = a log   − a log K  ρ 

A plot of log ηinh versus log (Rp/ρ) has slope a, and K can be calculated from the intercept. With these parameters evaluated for a given copolymer composition, Eq. (4.32) can be used over a wide range of polymerization conditions with no transfer agent present. For the usual case of initiation by thermal decomposition of persulfate, ρ can be calculated from values of kd estimated from Eq. 4.16. To extend the correlation to include effects of transfer reactions, one must decide on an appropriate form to express the transfer rate rtr. The usual preference is for highly reactive transfer agents, so that transfer rate is proportional to feed rate of transfer agent Ftr. For VDF copolymers, transfer agents with active hydrogen are often used (e.g., low molecular weight alcohols, esters, or ketones). For such agents used at moderate levels at relatively high reaction temperatures (>100°C), Eq. (4.31) may be modified to the form

Eq. (4.34)

Rp   ηinh =    K (ρ + 2K tr Ftr )

a

With K and a already evaluated as above for a system, experiments may be run with varying transfer agent feed rates to determine the transfer coefficient ktr . For evaluation of ktr from experimental data, Eq. 4.24 may be put in the form: Eq. (4.35)

ηinh

−1 a

−

F Kρ = 2Kk tr tr Rp Rp

The left-hand side of Eq. (4.35) is plotted versus Ftr /Rp to get a straight line (if the assumptions above hold), and ktr can be calculated from the slope. Polymerization conditions for which a given value of ktr applies are quite restricted. Transfer rates are not highly sensitive to temperature, so ktr may be approximately constant over a range of 10°C–20°C. However, transfer agents such as those listed above for VDF copolymers are soluble in both polymer particles and the aqueous phase. Thus, these agents distribute between phases, and the fraction in the polymer particles increases with increasing solids. Thus, Eq. (4.34) may apply to only a narrow range of dispersion solids. Alternatively, ktr will appear to vary with reaction time in a semibatch reactor or with residence time in a continuous reactor. For less reactive transfer agents that might be used in a semibatch reactor at low temperature, a correlation in the classical form[26] based on the ratio of transfer agent to monomer in polymer particles may be used to obtain transfer coefficients: Eq. (4.36)

ηinh

−1 a

−

[T ] Kρ = C tr [M ] Rp

The applicability of these relationships for controlling product characteristics varies with the type of reactor system employed. Design, operation, and control of continuous and semibatch emulsion polymerization systems are considered in Secs. 4.4.2 and 4.4.3.

4.4.2 �

Continuous Emulsion Polymerization

DuPont pioneered VDF/HFP/(TFE) polymerization in continuous stirred tank reactors (CSTRs) in the late 1950s. An early version of a continuous fluoroelastomer production process, including isolation, is described by Bailor and Cooper.[27] Recent versions of the continuous emulsion polymerization process, as run by DuPont Dow Elastomers, feature more feed components, monomer recovery with continuous recycle of unreacted monomers, and considerably more monitoring and control systems. A schematic diagram of such a continuous polymerization system, including monomer recovery and recycle, is shown in Fig. 4.5. Details of monomer recovery are discussed in Sec. 4.7.

50 Continuous polymerization has the advantage of allowing sustained production at steady state. High rates are attained at moderately high dispersion solids (15%–30%). Most or all of the heat of polymerization is removed by the temperature rise of chilled feed water, so polymerization rates are not limited by relatively low rates of heat removal through a reactor cooling jacket. Continuous polymerization is particularly advantageous for production of a few high-volume types, especially if individual product campaigns are two days or more in length. After initial adjustments are made, uniform polymer can be produced at the same conditions for a considerable period. Continuous polymerization is less attractive for a product line comprising many types, requiring short campaigns with frequent reactor startups and shutdowns. Modern control systems allow rapid attainment of goal polymer characteristics and thus good quality even in this situation. However, semibatch systems are better suited to making product lines with many low-volume specialty types. The range of products suitable for a continuous emulsion polymerization process is somewhat restricted. Monomer compositions must allow aqueous-phase oligomerization rates high enough so that continuous generation of new particles occurs, and thus steady polymerization rates can be attained.

Figure 4.5 Continuous emulsion polymerization system. ­

FLUOROELASTOMERS HANDBOOK Reasonably high radical generation rates are required, with dispersion stabilization by ionic oligomers and added soap. Suitable compositions include most vinylidene fluoride copolymers, especially the commercially important VDF/HFP/(TFE) and VDF/ PMVE/TFE products. For continuous emulsion polymerization of these VDF copolymers, P. L. Tang has found that low levels of highly water-soluble short-chain hydrocarbon alkyl sulfonates (e.g., sodium octyl sulfonate) are effective in place of fluorinated soaps.[28] TFE/PMVE perfluoroelastomers and ethylene/TFE/PMVE base-resistant elastomers can also be made in continuous reactors, though at much lower rates. Sustained particle nucleation is difficult to attain for TFE/propylene compositions; these do not appear suitable for production in a continuous polymerization. Certain polymer designs that require initial formation of particles with little or no further initiation must be made in semibatch reactors. An example is the Daikin family of polymers with almost all chain ends capped with iodine, made in a living radical polymerization. Continuous reactor design and operation. Continuous stirred tank reactors used for emulsion polymerization of fluoroelastomers are run essentially liquid-full at pressures high enough to keep unreacted monomers dissolved in polymer particles.[29] Operating pressures are in the range 2–7

4 PRODUCTION OF FLUOROELASTOMERS MPa at temperatures 60°C–130°C. Most VDF copolymers are made at 5–7 MPa and 100°C–120°C with residence times of 10–60 minutes. Slower polymerizing specialties (e.g., TFE/PMVE and E/TFE/ PMVE copolymers) are made at low pressures and temperatures, with longer residence time (2–4 hours). Potential corrosion from dispersions with pH’s in the range 2–6 at elevated temperatures is avoided by stainless steel reactor construction. To facilitate rapid dissolution of feed monomers and good mixing of the dispersion, fairly intense agitation is necessary, usually with baffled turbine impellers. A number of inlets must be provided for various components. Gaseous monomers are usually introduced into regions of high shear near impeller tips. Especially for operation at short residence times, it is necessary for agitation systems to be designed with high impeller flow, so that liquid turnover times are much shorter than residence times. VDF copolymer dispersions are usually not highly stable, since it is desirable to minimize added soap which must be readily removed during isolation. Thus, the maximum shear rate or impeller tip speed must be limited to avoid shear coagulation. Dispersion exits the reactor through a back-pressure control valve. Stability of the dispersion must be high enough to withstand the high shear involved in the letdown to much lower pressure in the degasser. Removal of heat of polymerization is a major consideration in reactor design for fluoroelastomers. For specialty types such as perfluoroelastomers made at low rates and low temperatures in relatively small reactors, heat removal through a cooling jacket is feasible. VDF copolymers are generally made at much higher rates per unit volume and high overall rates that require larger reactors. Cooling jackets are inadequate in this situation, so such reactors are usually operated adiabatically, with the heat of polymerization taken up by the temperature rise of water fed to the reactor. Heat of polymerization calculated from bond energies is in the range 300–350 kcal/kg for commercial VDF copolymer compositions. For adiabatic operation, dispersion solids must be limited so that the ratio of water fed to polymer made is high enough to allow a reasonable temperature rise from a practical water feed temperature to the reaction temperature. For example, a practical VDF/HFP/TFE polymerization may be run at about 20% solids, with 4 kg water per kg polymer in the reactor dispersion. In this case, if the heat of poly-

51 merization is 320 kcal/kg, a water temperature rise of 80 degrees is necessary, so a reaction temperature of 110°C requires a water feed temperature of 30°C. Similar conditions are described in DuPont patent examples:[30] VDF/HFP (60/40 wt%) copolymer made at 107°C, 10–12 minutes residence time, 18% solids, with polymerization rate 1.1 to 1.3 kg/h·L. Operation of a continuous reactor is quite different from semibatch polymerization. CSTR startup procedures are crucial to proper operation. Aqueous feeds are first established to fill the reactor at the desired operating pressure and temperature. These feeds include the main water flow, initiator components, soap, and buffering agents. Other liquid feeds—cure-site monomer, chain-transfer agent—that may retard polymerization are usually withheld until reaction has been established. Polymerization is started by commencing monomer feed at full rate and calculated overall composition suitable for the copolymer being made. When the reaction starts, a considerable exotherm (“heat kick”) occurs, tending to increase the reactor temperature. Water feed temperature is reduced and jacket cooling may be applied to bring the reactor temperature back to goal. Polymer particles are formed quickly, but some oscillation of particle number occurs in the first few reactor turnovers. High monomer conversion (80%–95%) is attained within 1–2 turnovers when feeds and other operating conditions are properly set up. Steady-state operation at full dispersion solids concentration is established after about six reactor turnovers. Unreacted monomer recovered from degassing vessels is recycled back to the feed compressor and fresh feeds are adjusted to maintain the desired polymer composition and production rate. Monomer feeds to a CSTR are illustrated in Table 4.4 taken from Ex. 4 of Ref. 30 describing VDF/HFP copolymer production in a 10-gallon (38liter) reactor at 89% conversion. At steady state, with recycle set equal to unreacted offgas rate, the fresh feed rate and composition equals polymer rate and composition. The total monomer feed to the reactor remains constant throughout the operation. At startup, before any recycle is established, fresh feed must equal total feed. CSTR shutdown is accomplished by shutting off the monomer feeds. This immediately stops the polymerization, since the unreacted monomer mixture

52

FLUOROELASTOMERS HANDBOOK

Table 4.4 CSTR Monomer Mass Balance (from Ref. 30)

Monomer

Fresh feed (polymer)

Recycle (offgas)

Total feed

kg/h

%

kg/h

%

kg/h

%

VDF

24

60

1.25

25

25.25

56

HFP

16

40

3.75

75

19.75

44

Total

40

held up in the reactor is quite unreactive toward propagation. Initiator and chain-transfer agent flows are then stopped. The main water feed, including soap, is maintained long enough to displace remaining polymer dispersion from the reactor to the degasser and blend tank. Continuous emulsion polymerization control. Control of a continuous emulsion polymerization reactor involves a number of aspects including temperature, conversion stability, radical generation rate, polymerization rate, polymer composition, and polymer viscosity. These control issues are discussed below. Temperature control systems must be sufficiently robust to overcome the inherent instability of this type of CSTR. Polymerization rate Rp and monomer conversion are sustained by radical generation rate, which must be adequate to form new particles continuously, thus maintaining the particle population. Radical generation rate ρ, especially if based on persulfate thermal decomposition, is sensitive to temperature. Thus, a decrease in temperature decreases ρ, which in turn decreases Rp and heat generation, tending to further decrease temperature. The control system must be able to respond fast enough to overcome this sequence of events that could lead to loss of reaction. For an adiabatic reactor, the heat exchanger on the main water feed must be able to switch quickly from heating the feed to goal reactor temperature at startup to cooling the feed well below the reactor temperature to take up the heat of polymerization. Partially bypassing the exchanger may be a means to make such a rapid transition. The system must also respond rapidly to prevent decreases in reactor temperature. CSTR polymerization systems have two possible steady states—the desired high conversion state and a very low conversion state, with an unstable inter-

5.00

45.00

mediate region. Upsets such as loss of initiator or excessive feed of a retarder may cause a flip from high to very low conversion. The low conversion situation means unreacted monomer builds up in the reactor, causing poor agitation and excessive monomer flow to the degasser, leading to potential safety hazards. Recovery from such a low conversion state is accomplished by stopping monomer feeds, continuing aqueous feeds to displace unreacted monomer and refill the reactor, followed by correcting the problem that caused loss of reaction, and restarting the polymerization. As noted above, radical generation rate sustains the polymerization rate in a CSTR, supplying radicals to existing polymer particles, and renewing the particle population by supporting aqueous oligomerization for particle nucleation and stabilization. The ratio ρ/Rp determines the ionic end-group level in the polymer and is a major factor in setting polymer molecular weight. For the usual case of persulfate thermal decomposition, the overall radical generation rate ρ can be calculated for a CSTR with water volume Vr and total water volumetric feed rate Fw with initiator concentration [I]o by making a mass balance on initiator to get its concentration [I] in the reactor. Eq. (4.37)

Fw[I]o = Fw[I] + Vrkd[I] or

[I] = [I]o (1+ kd θ) The first order decomposition rate coefficient kd for persulfate can be estimated from Eq. (4.16) or Table 4.3. Reactor residence time θ is the ratio Vr /Fw of water volume in the reactor to water flow.

4 PRODUCTION OF FLUOROELASTOMERS The water volume Vr is less than the total reactor volume because of the presence of polymer and unreacted monomer. The total radical generation rate at 100% efficiency is then Eq. (4.38)

ρ = 2k dVr [I] = =

2k dVr [I]o (1 + k d θ)

2k d θFI (1 + k d θ)

The molar feed rate of initiator FI is equal to Fw[I]o and kdθ/(1 + kdθ) is the fraction of initiator decomposed in a CSTR with residence time θ, operating at a temperature giving an initiator decomposition rate coefficient kd. The radical entry rate ρe is lower than ρ by an efficiency factor f: Eq. (4.39)

ρe =

2 f k d θFI (1+ kd θ)

Radical entry efficiency is usually low in these systems, about 0.2 to 0.6, but is ordinarily not known, so overall ρ is used for practical correlations applied to reactor control. Polymerization rate and in turn monomer feed rate goals are set by estimates from kinetics models (see “Kinetics Relationships” in Sec. 4.4.1), empirical correlations (e.g., Eq. 4.27), or plant experience. Monomer feed adjustments may be necessary to get goal Rp and polymer composition. Reactor effluent samples may be analyzed to determine composition and dispersion solids. Several estimates of Rp can be made from CSTR monitoring the following: • Calculation from water feed rate and dispersion solids • Monomer mass balance from flow meters and GC analysis of fresh feed, recycle, total feed, and offgas monomer streams • Heat balance on an adiabatic reactor Monomer mass balances also provide estimates of polymer composition. The total monomer feed composition may be adjusted to obtain goal polymer composition. Cure-site monomer feed is usually set in ratio to Rp or total gaseous monomer feed, and cure-site level is monitored by analysis of effluent polymer.

53 Control actions are facilitated if Rp is set below the maximum possible for the goal polymer composition and molecular weight at the prevailing reactor temperature and pressure. Then the monomer conversion is high enough that the monomer concentration in the particles is below the solubility limit. In this situation, changing a variable in a direction that tends to reduce Rp and conversion results in an increase in monomer concentration that tends to offset the change in Rp. Then most individual control actions can be accomplished without significant changes in Rp or conversion. Polymer viscosity control is facilitated by the use of relationships such as Eq. (4.34), which can be used to set ρ and transfer agent feed rate Ftr in ratio to Rp to get the desired ηinh. The ratio ρ/Rp sets the ionic end-group level in the polymer; this in turn affects bulk viscosity and bisphenol curing characteristics. Usually, ηinh is monitored from analyses of effluent samples. Then Ftr can be adjusted from correlations like Eq. (4.34) to get the desired polymer viscosity. Reactor dynamics must be taken into account in managing such control actions, since a change in a reactor input variable takes about six turnovers (6θ) to be fully reflected in dispersion effluent analyses. Dispersion stability is affected by soap feed rate, ρ, and pH. Base or buffer feed is set as a ratio to ρ, with adjustments made in response to pH measurements on effluent dispersion. With proper setup of polymerization conditions, control actions taken after startup should be only small adjustments. Redundant measurements of reactor variables (e.g., monomer flow and composition) are desirable to allow checking for instrument errors. Besides the measurements taken around the reactor used for direct polymerization control, monitoring of many other systems—monomer feed and recycle compressors, degassers, agitators, impurities in feeds, and leaks—is necessary for safe, smooth operation.

4.4.3 �

Semibatch Emulsion Polymerization

All fluoroelastomer producers use semibatch emulsion polymerization systems. Detailed descriptions of commercial fluoroelastomer semibatch systems are not available in the open literature, but

54 smaller scale reactors are described in a number of patents. Figure 4.6 is a schematic representation of a fluoroelastomer semibatch reactor with associated charging and feed systems, and monomer recovery system. Shown are components usually charged initially, and those that may be fed during the course of the polymerization. Semibatch polymerization is suitable for a wide range of compositions, including those having very slow polymerization rates. Semibatch reactors are more versatile than continuous reactors for making specially designed polymers. Feeds of initiator, transfer agents, and cure-site monomers can be varied during the course of a batch to make polymers with different molecular weights and molecular weight distributions, end groups, and curesite distribution along chains. This allows control of rheology, processing, and curing behavior to an extent not attainable in CSTRs. Polymer composition and polymerization rate are readily controlled by setting monomer feeds during the reaction. Commercial semibatch reactors are capable of making a considerable number of low volume specialty products. However, the necessity of keeping different products separate in downstream handling equipment limits the versatility of the reactor system.

Figure 4.6 Semibatch emulsion polymerization system. ­

FLUOROELASTOMERS HANDBOOK Semibatch reactors have limitations compared to continuous reactors in the production of high-volume, fast-polymerizing types. Heat of polymerization must be removed by means of a cooling jacket. With this limited cooling capability, polymerization rates must be limited well below those possible in adiabatic CSTRs for many important high-volume products (e.g., VDF copolymers containing 60-80 mole % VDF). In campaigns of high-volume types, many batches with attendant shutdowns and startups are required, and batch-to-batch variability may be significant. For many types, reaction times may be too short to allow monitoring of product characteristics, feedback, and adjustments within each batch. Adjustments can be made on subsequent batches, but large blend tanks may be required to reduce final product variability. Holdup of gaseous monomer mixtures in semibatch reactors and feed systems is greater than that in CSTR systems. Considerable volumes of monomer mixtures under pressure in semibatch reactor vapor spaces and in accumulators after compressors may present potential explosion hazards. The lower operating pressures of semibatch reactors somewhat offsets this hazard, compared to

4 PRODUCTION OF FLUOROELASTOMERS CSTRs. However, barricades around semibatch reactors and feed facilities may be necessary to protect personnel. Semibatch reactor design and operation. Semibatch reactors are generally run at lower temperature and pressure than CSTRs. Usual ranges for semibatch operation are 60°C–100°C and 1–3 MPa (150–450 psi). Reaction times required to get dispersion solids of 25%–35% are quite variable, depending on composition and other variables related to polymer design, and may range from 2 hours to as much as 40 hours. Usually, the volume of aqueous dispersion is 60%–85% of total reactor volume. Dispersion volume increases significantly during the course of polymerization because of the increasing volume of polymer swollen with monomer. The general procedure for operation of a semibatch reactor is as follows: The reactor is charged with water and soap solution, and with monomers of the composition necessary to be in equilibrium with the desired polymer composition. Usually the unreacted monomer mixture from a previous batch makes up the bulk of the monomer charge. Ordinarily, this initial feed would bypass the accumulator after the compressor. The reactor is brought to the desired operating temperature and pressure. Reaction is started by adding persulfate initiator and chain-transfer agent. During polymerization, monomers are fed at the desired polymer composition to maintain reactor pressure. Additional initiator, transfer agent, cure-site monomer, and buffer may be fed during the polymerization as necessary to make the desired polymer. In semibatch emulsion polymerization of fluoroelastomers, particle formation occurs during the early part of the polymerization, but may be prolonged to rather high solids concentrations. With the increase in number of particles and growing radicals, the polymerization rate may also increase over a considerable fraction of the reaction time. To handle the large differences in monomer feed rate necessary to match the varying polymerization rate, an accumulator may be used between the feed compressor and reactor. The accumulator is maintained in a pressure range above the reactor pressure. A monomer mixture of the desired polymer composition is fed from the accumulator to maintain constant reactor pressure. Monomers may be fed periodically through the compressor to keep the accumulator in a set pressure range. This arrangement

55 allows metering of the monomers at convenient rates for accuracy in setting composition. The use of an accumulator does add a significant volume of highpressure monomer mixture to the feed system. This may be a potential explosion hazard for some monomer mixtures. Careful investigation is necessary to determine the extent of such hazards and provide means of avoiding damage or injury from possible deflagration of monomer mixtures. When the desired amount of polymer has been made, as estimated from the cumulative amount of monomer fed during the polymerization, shutdown is accomplished by stopping the feeds of monomer, initiator, and other minor components. Monomer may be removed by venting directly from the reactor. However, with the limited head space in this vessel, foaming and dispersion carryover into vapor lines can be a severe problem. It is more feasible to transfer the dispersion from the reactor to a larger degassing and blend tank. If the same composition is to be made in the next batch, it is convenient to leave a heel of dispersion in the reactor, along with the remaining unreacted monomer in the vapor space. This facilitates recharging and startup of the next batch. Commercial semibatch reactors used for manufacture of fluoroelastomers are generally 1,000 to 12,000 liters in size, larger than the CSTRs described in the previous section. Relatively small sizes are used for fast-polymerizing high-volume VDF copolymers, while larger reactors may be used for specialty types with lower polymerization rates. For copolymers of 60–80 mole % VDF with HFP and TFE or PMVE and TFE, rates may be limited by heat transfer capability. This situation can be analyzed by reference to Fig. 4.7, which shows a jacketed cylindrical reactor with diameter D, total height HT, and dispersion depth HL to get dispersion volume VL and heat exchange area A: Eq. (4.40)

VL = ðD 2 H L 4

Eq. (4.41)

A = π DHL

Consider the case of a reactor with total heightto-diameter ratio HT/D = 1.85, 83% full of dispersion, and thus with liquid height HL = 1.5D. Then, from Eq. 4.40, the liquid volume is given by VL =

56

FLUOROELASTOMERS HANDBOOK

Figure 4.7 Semibatch reactor: heat exchange area and liquid volume.

1.5πD 3/4, and the diameter can be expressed in terms of the liquid volume by D = (4VL/1.5π )1/3. The heat exchange area, from Eq. 4.41, can be expressed as A = 1.5πD 2, or related to liquid volume as A = 1.5π(4VL/1.5π)2/3 = 4.23VL2/3. For the situation with the maximum rate of polymerization limited by heat transfer capability, the following relationships apply: Eq. (4.42)

Rp max =

=

UA∆t ∆hp

4.23U∆tVL ∆hp

2/ 3

In Eq. 4.42, Rp max is the maximum polymerization rate for a polymer with heat of polymerization ∆hp in a reactor with heat exchange area A, overall heat transfer coefficient U, maximum average temperature difference between dispersion and jacket coolant ∆t, and dispersion volume VL. The maximum rate is proportional to the heat exchange area, thus to the two-thirds power of dispersion volume (or reactor volume). These relationships can be used to approximate the scale-up situation for semibatch reactors making polymer compositions for which rates are limited by heat exchange capabilities. A reasonable base case is that of a 1,500-liter reactor, charged with 1,000 liters of water, with capability of making 400

kg of VDF copolymer (28.6% solids in the dispersion after degassing) in two hours reaction time, thus Rp max = 200 kg/h. Assuming the monomer-swollen polymer has density 1.6 kg/liter, polymer volume is 250 liters and total dispersion volume, VL, is 1,250 liters or 1.25 m3, corresponding to 83% full. With HL/D = 1.5, D = 1.02 m from Eq. (4.40) and A = 4.89 m2 from Eq. 4.41, ∆hp = 320 kcal/kg or 1.34 MJ/kg and maximum ∆t = 50 K, which corresponds to a reaction temperature of 80°C and average coolant temperature of 30°C. The overall heat transfer coefficient U is 260 kcal/m2·h·K or 1.1 MJ/ m2·h·K, a reasonable value. Now consider scaling up this polymerization to a reactor eight times the size (12,000 liters) to make 3,200 kg polymer per batch. For the large reactor, VL = 10 m3, D = 2.04 m, and A = 19.6 m2, even with the optimistic assumption that U will be the same for the large reactor, the fourfold increase in heat exchange area limits the maximum rate to 800 kg/h, so that the reaction time increases twofold to four hours. This is probably a good tradeoff for scaling up, since total batch cycle time for the large reactor to make 3,200 kg would be much less than the total time required for eight batches in the small reactor. Offsetting the advantage of the larger reactor would be the cost of scaling up feed equipment, downstream blending, and isolation capacity. The larger monomer volumes under pressure may also introduce severe explosion hazards for the larger reactor. Design considerations for semibatch reactors differ significantly from those for continuous reactors. Since polymerization rates per unit volume are lower in semibatch reactors, these are usually much larger than continuous reactors. Pressures and temperatures are usually lower in semibatch reactors. Intensity of agitation is ordinarily lower, since high shear regions are not necessary to disperse monomers and other feeds to a semibatch reactor. Agitation systems should be designed for reasonable liquid turnover, with minimal baffling to avoid elastomer agglomeration and fouling. The presence of a sizeable volume of monomers under pressure in the head space of the reactor creates the potential for explosion hazards. Care must be taken to preclude possible sources of ignition, such as rubbing of moving metal parts, presence of air or other initiators, and electrical arcs or sparks. Proper relief area must be provided, which is capable of relieving the overpressure from a deflagration in the vapor space. For some

4 PRODUCTION OF FLUOROELASTOMERS monomer mixtures, this consideration may limit the reactor pressure or size. Feed systems for semibatch operation involve a combination of initial charging of some components and of feeding components at variable rates during polymerization. These requirements will be discussed in the next section. Semibatch emulsion polymerization control. Basic control of semibatch systems for making older VDF-based copolymer and terpolymer products is somewhat simpler than control of continuous reactor systems. However, new products require complex schemes for operation of the semibatch polymerization reactor to get the desired processing and curing characteristics. Major requirements of a semibatch reactor control system include: accurate initial charging of ingredients, good control of feeds of major monomers and minor components during polymerization, and maintaining the reactor at goal temperature and pressure. Before charging, it is necessary to clear air from the reactor system by flushing with inert gas, evacuation, or displacement of vapor with water. The proper amount of water (usually 50%–70% of total reactor volume) is charged along with dispersant (usually soap and buffer), and the reactor contents are heated to the desired polymerization temperature. Monomer is then charged with the appropriate composition to be in equilibrium with the desired polymer composition, and in an amount to bring the reactor to a chosen pressure at goal temperature. Major monomer composition is checked by gas chromatography. Polymerization is started by adding initiator. Then a monomer mixture with a composition essentially the same as the desired polymer composition is fed at a rate to maintain goal reactor pressure. Jacket coolant temperature is adjusted to keep the reactor at goal temperature as polymerization proceeds. Additional initiator is usually fed to hold the radical generation rate in a range that will maintain the polymerization rate and attain the desired polymer molecular weight and ionic end-group level. A transfer agent may be added to control the polymer’s molecular weight and molecular weight distribution. A cure-site monomer may be fed in ratio to the main monomer feed. Depending on reactor size and the polymerization rate, feeds of minor components may require special metering equipment to deliver low flows or small incremental shots accurately. Both instantaneous rates and cumulative amounts of ma-

57 jor monomers and minor components need to be monitored. The polymerization rate and total polymer formed are estimated from major monomer feeds. Note that the polymerization rate and the monomer feed rate may vary considerably over the course of a semibatch polymerization. An accumulator between the feed compressor and reactor may be necessary to facilitate delivery of a controlled monomer composition, especially during the early stages of the reaction when rates may be low. Polymerization is stopped by shutting off the monomer feeds when a desired dispersion-solids level is reached or a desired polymer viscosity is attained, both estimated from cumulative monomer feed. Reactor sampling may be feasible for relatively slow polymerizations with long enough reaction time to allow adjustments of feeds. The vapor space may be monitored by gas chromatography. Dispersion sampling may be difficult, especially in the usual situation of a barricaded reactor system precluding operating entry. Often the polymer characteristics must be inferred from monitoring of feed components during the polymerization. A number of strategies may be used for addition of initiator, transfer agent, and cure-site monomer components during a semibatch emulsion polymerization. The simplest initiator feed method is to add all of it at the start of polymerization. In this case, the total moles of persulfate initiator, I, in the reactor decrease with time, t, according to first order thermal decomposition kinetics: Eq. (4.43)

dI/dt = -kdI

Eq. (4.44)

It = I0 exp(-kdt)

Total radical generation rate, ρ t, at time, t, is given by Eq. (4.45)

ρ t = 2kdIt = 2kd I0 exp(-kdt)

In a semibatch emulsion polymerization, radical entry efficiency varies considerably as the particle population builds up. Since the efficiency is not readily estimated, it is easier to use total generation rate for correlation and monitoring purposes. The cumulative number of moles of radicals generated from time 0 to time t is Eq. (4.46)

Σρt = 2I0 [1- exp(-kdt)]

58

FLUOROELASTOMERS HANDBOOK

This method of adding initiator all at once may be usable for some semibatch polymerizations carried out at relatively low temperatures, say 80°C or below, with a persulfate initiator half life of two hours or more. The relatively high radical generation rate at the start facilitates particle formation, and the slowly decreasing radical generation rate may adequately sustain polymerization in later stages. However, this method is not versatile enough to control the polymerization rate, molecular weight, and end groups for most products of interest. A second method, often used for small reactors, is to add incremental shots of initiator to keep the initiator level between chosen levels, I0 and It, at intervals of time, t. The increment size, ∆I = I0 – It, is readily calculated from Eq. 4.44. Corresponding radical generation rates and cumulative radicals are estimated from Eqs. 4.45 and 4.46. For larger reactors, initiator may be fed continuously to get a desired profile of radical generation rate versus time, thus optimizing polymer viscosity and ionic end-group level, taking into account changes in polymerization rate over the course of the reaction. Initiator feed, FI, can be chosen to obtain constant, increasing, or decreasing ρ. For constant initiation rate, FI, is set equal to the initiator decomposition rate so that It = I0; then FI = kd I0 and ρ = 2kd I0. For the more general case, the following relationships apply: Eq. (4.47)

dI = FI − k d I dt

Eq.(4.48)

I t = I 0 exp(− k d t ) + FI

[1− exp(− kdt )] kd

Eq. (4.49) ρ t = 2k d I t

= 2kd I 0 exp(− kd t ) + 2FI [1− exp(− k d t )]

Eq. (4.50)

 F  Σρt = 2 I 0 − I [1− exp(− k d t )] + 2FIt kd  

Note that Eq. 4.50 can be applied to a number of intervals with varying initiator feed rates to get an overall summation of radicals generated over the course of the reaction. Chain-transfer agents are often used to control the molecular weight of fluoroelastomers. However, chain-transfer correlations and predictions are less readily obtained for semibatch systems than for continuous polymerization systems. Basic relationships like Eq. 4.31 are difficult to apply to semibatch systems. For transfer agents with low reactivity at the relatively low temperatures normally used in semibatch polymerization, Eq. 4.37 may be usable. This relationship is based on the ratio of transfer agent to monomer in particles. The transfer agent level in particles may not be readily estimated for agents that have substantial solubility in water because they are distributed between the aqueous and polymer phases. For some fluoroelastomer compositions, hydrocarbons may be used as transfer agents. These may be volatile enough to monitor by gas chromatography analysis of the vapor phase in the reactor. Such transfer agents are not usually used for VDF-containing polymers, because the hydrocarbon radicals formed by transfer are much less reactive toward propagation than the fluorocarbon radicals, thus retarding polymerization. Highly reactive transfer agents (e.g., lower alcohols or esters) may be fed continuously in set ratio to the monomer feed or the radical generation rate to get the desired polymer viscosity. However, while most of the chaintransfer agent reacts immediately in continuous systems operating at higher temperatures, that assumption can not be made for semibatch systems. Thus, considerable small scale polymerization work is often necessary to establish how to charge and/or feed transfer agents to obtain the desired polymer viscosity and molecular weight distribution for each composition. In a special case of transfer in semibatch emulsion polymerization, perfluorocarbon diiodides are used to make fluoroelastomers that have narrow molecular weight distribution and iodine at most chain ends for curing. As originally developed by Daikin workers,[31] a “living radical” polymerization is set up with very low levels of initiation and termination, so that propagation and transfer predominate. Soap levels are set high enough to obtain a large population of small particles containing no more than one growing radical each. Ordinarily, all of the diiodide

4 PRODUCTION OF FLUOROELASTOMERS transfer agent is added soon after polymerization starts, so that very few chains form without iodine end groups. Iodide ends, whether on polymer chains or the original perfluorocarbon iodide, continue to undergo transfer. Individual chains grow until they undergo transfer; the resulting iodide may transfer subsequently to allow further propagation and an increase in molecular weight. Since most chains start near the beginning of the polymerization, and very little radical-radical termination occurs, the chains have equal opportunity to grow. The result is a polymer with a very narrow molecular weight distribution and with iodine on most chain ends. Molecular weight continues to increase as polymerization proceeds; it can be estimated from the ratio of cumulative monomer feed to moles of iodide charged. Polymerization is stopped by shutting off the monomer feed when the estimated molecular weight goal is attained. In these polymerizations, adventitious impurities that may transfer to form unreactive radicals must be minimized. Even so, small amounts of initiator must be added from time to time to sustain the radical population and desired polymerization rate. It is crucial that a known amount of iodine is charged initially to allow an adequate estimation of molecular weight. Since these polymerizations are slow, it may be possible to take dispersion samples for measurement of the polymer’s inherent viscosity. A plot of inherent viscosity versus cumulative monomer feed may then be used to estimate the cumulative monomer level that will give the desired final viscosity. Most cure-site monomers are incorporated into the polymer at low levels. For many of these monomers, conversion is high, so they are not charged initially, but are fed in controlled ratio to the major monomers fed during the reaction. In a few cases, cure-site monomers may also be charged along with the initial monomer charge. Cure-site monomers, with active groups such as iodine or bromine, used for free radical curing present special problems in semibatch polymerization. Unlike the situation in continuous systems with continuous removal of polymer from the reactor, all chains formed in semibatch systems stay in the reactor until shutdown. This means that incorporated monomer units with reactive cure sites are exposed to radicals for considerable periods of time. The resulting chain transferto-polymer reactions may lead to excessive branching and gel formation, which may be detrimental to

59 processing characteristics. This situation has been circumvented in recent developments (e.g., by Ausimont workers)[32] by using small amounts of iodine-containing olefin monomer in conjunction with perfluorocarbon diiodide transfer agent. This allows production of fluoroelastomers with iodine units incorporated along the chains as well as at chain ends. Chain branching can then be controlled to allow reasonable polymer rheology and compound processing characteristics. A preferred Ausimont process variant for attaining reasonably high polymerization rates is a microemulsion process. A stable emulsion of perfluoropolyoxyalkalene solvent stabilized with a perfluoropolyoxyalkalene carboxylate surfactant is charged initially with monomers to obtain a large number of small particles and a subsequently high polymerization rate.

4.5 � Suspension Polymerization Suspension polymerization is used to make a number of thermoplastic polymers. In suspension polymerization, all reactions are carried out in relatively large droplets or in polymer particles stabilized by a small amount of water-soluble gum. Organic peroxide initiators are used to generate radicals within the droplets. A solvent may be used to dissolve a monomer at relatively high concentration. The main advantages of suspension polymerization over emulsion systems are that no surfactants, which are difficult to remove from the product, are used, and no ionic end groups are present which may be unstable during processing at high temperatures. What follows is a general introduction of suspension polymerization; S. Ebnesajjad[33] has presented an extensive review of suspension polymerization of vinylidene fluoride. In one semibatch suspension process for making VDF homopolymer,[34] the reactor is charged with water containing a cellulose gum (about 0.03%) as the suspending agent, an initiator solution, and a VDF monomer. The initiator of choice is diisopropyl peroxydicarbonate, which has a half life of about two hours at 50°C. The jacketed reactor is heated with agitation to a temperature in the range 40°C to 60°C, with a pressure in the range 6.5 to 7.0 MPa maintained by adding additional water or monomer

60 during the polymerization period of about 3.5 hours. Chain-transfer agents may also be fed. Average particle diameter is typically about 0.1 mm for the dispersion obtained in suspension polymerization. At the end of polymerization, the reactor is cooled, the dispersion is degassed by letting off pressure from the reactor, the polymer is separated by filtering or centrifuging the dispersion, and washed to remove residual dispersion stabilizer. Major features of this process were adapted by workers at Asahi Chemical Industry Co., Ltd. to make VDF/HFP/(TFE) fluoroelastomers. In the initial version of the Asahi Chem suspension polymerization process,[35] a relatively large amount of an inert solvent, trichlorotrifluoroethane (CFC-113, CCl2F–CClF2), is dispersed in water containing 0.01%–0.1% methyl cellulose suspending agent. The mixture is heated under agitation to the desired polymerization temperature (usually 50°C) and the proper composition of VDF/HFP/(TFE) monomer mixture to make the desired copolymer is charged in the amount necessary to get the goal concentration in the monomer-solvent droplets. With the solvent used, the pressure is usually relatively low, about 1.2–1.6 MPa. Reaction is started by adding diisopropylperoxydicarbonate initiator solution and a monomer mixture, with composition essentially that of the polymer being made, is fed to maintain the reactor pressure constant. Polymerization starts in the monomer-solvent droplets, with initial formation of a low molecular weight fraction. As polymerization proceeds, viscosity of the particles increases, long-lived radicals form, and both polymerization rate and molecular weight increase with reaction time. The resulting polymer has a bimodal molecular weight distribution, with the minor low molecular weight fraction acting as a plasticizer for the bulk high molecular weight polymer. Normally no chain-transfer agents are used for polymers cured with bisphenol. Polymer viscosity is set from the ratio of total polymer formed to initiator charged. Since reaction times are fairly long (six hours or more) to attain high dispersion solids (30%–40%), dispersion samples can be taken from the reactor during polymerization to monitor inherent viscosity and predict when to stop polymerization for goal viscosity. After polymerization is stopped by turning off the monomer feed, monomers are removed by venting the reactor. Considerable care is necessary during this operation to reduce pressure in stages so

FLUOROELASTOMERS HANDBOOK that rapid release of monomer from particles does not occur, and carryover of particles into vapor lines is avoided. Particle sizes after degassing are 0.1 to 1 mm in diameter, and are readily separated by filtering or centrifuging the dispersion. Fluoroelastomers made by the suspension process have no ionic end groups and contain a significantly low molecular weight fraction. These copolymers can be made with high inherent viscosities for enhanced vulcanizate properties, while they still retain good processibility because their compounds have relatively low viscosity at processing temperatures. Compared to emulsion products of similar composition, bisphenol-curable suspension products exhibit better compression set resistance, faster cure, and better mold release characteristics. Asahi Chem also developed peroxide-curable VDF/HFP/TFE fluoroelastomers by charging methylene iodide along with the initiator to the suspension polymerization reactor. The resulting chaintransfer reactions allow incorporation of iodine on more than half the chain ends. Final polymer molecular weight is determined mainly by the ratio of total monomer fed during the polymerization to iodine incorporated. The suspension process has been adapted to make bimodal VDF/HFP/TFE polymers for extrusion applications, such as automobile fuel hoses, to get smooth extrudates with minimal die swell at high shear rates.[36] These polymers contain 50%–70% very high molecular weight fractions (ηinh about 2.5 dL/g, Mn about 106 daltons) and 30%– 50% very low molecular weight fraction (ηinh about 0.15 dL/g, Mn about 17,000 daltons), with polymer bulk viscosity determined by the relative amounts of the two fractions. The low viscosity fraction has a molecular weight below the critical chain length for entanglement (Me about 20,000 to 25,000), so it acts as a plasticizer to facilitate extrusion with low die swell. Similar bimodal polymers with low viscosity fractions having molecular weights greater than Me would exhibit very high die swells. Synthesis of these polymers is carried out in two stages of suspension polymerization. A very small amount of initiator is used in the first stage to make the high molecular weight fraction. Then additional initiator and a relatively large amount of methylene iodide are charged to make the low viscosity fraction. The relative amounts of each fraction are estimated from the cumulative monomer feed in each stage. The amount of methylene iodide charged is that required to in-

4 PRODUCTION OF FLUOROELASTOMERS corporate 1.5%–2% iodine in the low viscosity fraction. Polymerization rate in the second stage is very low, so the total reaction time required for the bimodal polymer synthesis is some 40–45 hours. These bimodal polymers are ordinarily cured with bisphenol, but the iodine ends on the low viscosity fraction allow a mixed cure system with both bisphenol and radical components. The radical system links very short chains into longer moieties that can be incorporated into the bisphenol crosslinked network. Similar bimodal polymers made by emulsion polymerization with conventional chain-transfer agents are cured only with bisphenol. The resulting vulcanizates contain sizeable fractions of short chains that are not incorporated into the network and are thus susceptible to extraction when exposed to solvents. The suspension process described above was used by Asahi Chem for commercial production of Miraflon fluoroelastomers during the early 1990s. However, it was recognized that the use of large amounts of the ozone-depleting solvent CFC-113 would need to be phased out. A second version of the suspension process uses a small amount of a hydrogen-containing solvent such as HCFC-141b, CH3-CFCl2. Since only enough solvent is used to dissolve the initiator, the reactor operating pressure must be increased to 1.5–3.0 MPa so that a fraction (10–30%) of the initial monomer charge condenses to form an adequate volume of droplets to serve as the polymerization medium. In a further improvement, the hydrochlorofluorocarbon solvent is replaced with a small amount of a water-soluble hydrocarbon ester, preferably methyl acetate or t-butyl acetate.[37] These polar hydrocarbon solvents are used mainly to feed the initiator to the reactor. The methyl or tbutyl groups are relatively inactive toward transfer, and these solvents are so soluble in water that little is in the polymer phase. After the Asahi Chem suspension polymerization technology was acquired by DuPont in 1994, additional development was carried out to extend the technology to VDF/PMVE/ TFE fluoroelastomers with cure-site monomers incorporated along the chains.[38] Cure-site monomers can be incorporated evenly along chains by careful feed in controlled ratio to polymerization rate of major monomers. In this way, bromine- or iodine-containing monomers can be incorporated, in addition to iodine on chain ends from methylene iodide transfer agent, to get polymers with improved characteristics in free radical cures. It should be noted that simi-

61 lar polymers can be made more readily by continuous emulsion polymerization.[39] Of more interest are bisphenol-curable VDF/PMVE/TFE compositions with 2H-pentafluoropropylene, CF2=CH–CF3, as cure-site monomer. Bisphenol-cured parts from such polymers have better thermal stability than products made by radical curing.

4.5.1 �

Polymer Compositions

The suspension polymerization process works well for VDF/HFP/TFE and VDF/PMVE/TFE compositions. These monomer mixtures exhibit high propagation rates at relatively low temperatures (45%–60°C) and low monomer concentrations (less than 15% in monomer/polymer particles). Reasonably high polymerization rates are possible at temperatures below 60°C, so elastomer particle agglomeration is minimized. The amorphous polymers are insoluble in the monomer/solvent mixtures and also the monomer and solvent have low solubility in the polymer-rich phases. The high viscosity of the polymer-rich phase gives hindered termination, so that long-lived radicals can grow to high molecular weights. The initial monomer mixtures charged to the reactor can be partially condensed at about 50°C and moderate pressure to form droplets as the initial locus of polymerization, without the need for charging large amounts of solvent or for charging polymer seed particles. Slower propagating compositions like TFE/ PMVE give a lower molecular weight and a less useful polymer when made by suspension polymerization than polymer that can be obtained by emulsion polymerization. For these perfluoroelastomers, monomer solubility in the polymer is high, so particle viscosity remains too low for hindered termination and the formation of long-lived radicals. Considerable initiator must be fed during the polymerization to sustain reasonable reaction rates. Several other TFE copolymer compositions give similar results.

4.5.2 �

Polymerization Mechanism and Kinetics

In all versions of the suspension-polymerization process, an initial dispersion of low-viscosity droplets is present, either from solvent containing dissolved monomer or from liquid monomer partially

62

FLUOROELASTOMERS HANDBOOK

condensed from the initial monomer charge. With the low viscosity of the monomer-solvent phase and the relatively high initial radical flux, both initiation and radical-radical termination rates are high, so the polymer formed in the early stages of the reaction is low in molecular weight. Solution kinetics apply in this early stage. The general reaction scheme outlined in Sec. 4.3.1, describing initiation, propagation, and termination reactions (Eqs. 4.1, 4.2, and 4.4) can be used in this situation. In the mobile droplets, rates of radical generation and termination are equal: Eq. (4.51)

2f kd[I] = 2kt[R·]2

Radical concentration in the droplets can be expressed as

Eq. (4.52)

  [R·] =  f k d [I ]  kt 

12

The polymerization rate, Rp, and the numberaverage molecular weight Mn (assuming termination by radical combination) are then given by: 12

Eq. (4.53)

 f k [I ] Rp = k p [M ][R·] = k p �[M ] d   kt 

Eq. (4.54)

Mn =

k p [M ]

( f k d k t [I])1 2

The high termination rate coefficient, kt, leads to a low rate and molecular weight in this initial stage of suspension polymerization. As the reaction proceeds, the insoluble polymer formed builds up as a second high-viscosity phase in the droplets. Radical mobility is limited in this viscous polymer-rich phase, so termination rate decreases and both molecular weight and polymerization rate increase with time. In the later stages of polymerization, kt approaches zero. Long-lived radicals persist in the dominant viscous polymer phase, so that growth of these chains continues even though most of the initiator has decomposed and the new radical formation rate is low. Ordinarily, the initial low molecular weight polymer is a small fraction of the total polymer formed.

When methylene iodide is used to form iodine end groups for radical curing, initiator levels are minimized so that transfer reactions predominate. Depending on iodide level, polymerization rates may be quite low, even in the later stages of the reaction. Ordinarily, all the methylene iodide is charged with the initiator. Since the iodide is somewhat soluble in water, its level in the droplets is initially low enough so the polymerization can be started at a reasonable rate. As the reaction proceeds, all the iodide enters the droplets and undergoes transfer. As in the semibatch emulsion case with perfluorinated iodide as the transfer agent, chains undergo alternating periods of propagation interrupted with iodide transfer from other chain ends. Usually the suspension polymers made with methylene iodide contain no more than about 1.5 iodine ends per chain and have somewhat broader molecular weight distribution than semibatch emulsion polymers made with perfluorinated iodides and very low initiation levels. The suspending agent, usually a water-soluble gum such as methyl cellulose with moderate molecular weight, prevents agglomeration of droplets and monomer-swollen polymer particles by forming a water-swollen coating on them. These gums are effective at low concentrations, typically less than 0.1% concentration in the water charged. Also polymerization temperature must be less than about 70°C so the swollen elastomer particles are not too sticky. Cellulose derivatives contain structures that normally would participate in chain-transfer reactions. However, these materials are so water-soluble that essentially none is in the droplets or polymer particles, thus do not reduce polymer molecular weight. The initiator of choice for fluoroelastomer suspension polymerization is diisopropyl peroxydicarbonate (or isopropyl percarbonate, IPP), R–O– C(:O)–O–O–C(:O)–O–R, where R is isopropyl. Under polymerization conditions, the IPP added to the reactor is dissolved in the fluorinated monomer/ polymer droplets, and its half-life is about 2.5 hours (kd = 0.27/h) at 50°C. IPP decomposition by thermal homolysis gives isopropyl carbonate radicals, R– O–C(:O)–O·, which react readily with fluorinated monomers to initiate polymerization. In the absence of a reactive monomer, the isopropyl carbonate radicals may undergo further decomposition to isopropoxy radicals, R–O·, and carbon dioxide. Isopropoxy radicals may react with IPP to induce

4 PRODUCTION OF FLUOROELASTOMERS further decomposition. The IPP decomposition rate varies with the medium, and increases significantly in polar solvents. Thus solutions must be kept cold and used soon after makeup. IPP is supplied as a solid (m.p. 8°C–10°C) which must be stored in a dedicated freezer at temperatures below –20°C. Above –10°C, IPP decomposes slowly, but generates heat internally so that the temperature may increase rapidly and the decomposition autoaccelerates. Decomposition products include flammable vapors which may be ignited. Proper storage and handling procedures are necessary to avoid these problems.

4.5.3

Reactor Design and Operation

A reactor used for the suspension polymerization of fluoroelastomers must be designed to minimize agglomeration of swollen particles and fouling of vessel surfaces. Agitation must be sufficient to disperse the initial condensed monomer-solvent phase into small droplets and to keep polymer particles from settling. Standard turbine agitators may be used with minimal baffling that is sufficient to avoid vortex formation without producing regions of high turbulence. Reactor fouling must be monitored and removed periodically. This maintains heat removal capacity through the cooling jacket and allows adequate temperature control. Removal of polymer deposits is facilitated by ports for water jets. As with emulsion semibatch reactors, vessel size must accommodate a considerable increase in volume of the liquid phase as reaction proceeds to high solids. Typically, the initial aqueous solution charge occupies about 60% of the vessel volume. The final dispersion, containing up to about 40% polymer, may occupy some 80%–85% of total volume. If degassing is carried out by letting down reactor pressure after completion of a batch, enough vapor space must be allowed to minimize entrainment of particles in the vapor stream vented from the reactor. This ordinarily requires a ratio of length-to-diameter of about two for the vessel. Adequate relief area should be provided for the vapor space to avoid damage from potential monomer deflagration. As with semibatch emulsion polymerization, monomer feed rates for a suspension reactor may vary over a wide range during the course of each batch operation. An accumulator may be necessary between the feed compressor and reactor to facili-

63 tate metering, as discussed in semibatch reactor design and operation. Careful measurement of initiator, modifier, and cure-site monomer feeds is also necessary. Special design considerations apply to storage and handling of peroxydicarbonate initiator, as noted in Sec. 4.5.2.

4.5.4

Polymerization Control

Similar polymerization control considerations apply to semibatch suspension systems as those described in semibatch emulsion polymerization control. Reference 38 describes suspension polymerization system operating and control procedures for making two fluoroelastomers of different compositions: one is a VDF/HFP/TFE polymer with a bromine-containing cure-site monomer and iodine end groups for peroxide curing, and the other is a VDF/ PMVE/TFE polymer with a cure-site monomer for bisphenol curing. In both cases, a 40-liter reactor was configured for carrying out semibatch polymerizations. The gaseous monomer feed system consisted of a source line for each gaseous monomer, a compressor, an accumulator, and a pressure controller between the accumulator and reactor vessel. At the beginning of the polymerization, monomers were consumed in the reactor at a low rate. The monomer supply rate to the compressor was considerably higher to maintain an accurate monomer composition. The difference in the amount of monomer fed to the compressor and the amount consumed in the reactor was stored in the accumulator. The storage in the accumulator was controlled by a pressure controller, which was cascaded to several flow controllers metering the monomer mixture to the compressor. As monomers flowed into the accumulator, the pressure increased to a high preset limit. When the high limit was reached, the flow controllers closed the gaseous monomer feed valves. As monomers flowed into the reactor, the accumulator pressure dropped to a low limit. At the low limit, the monomer supply valves opened and compressed gases were fed to the accumulator until pressure reached the high set limit, which shut off the monomer feed. This cycle continued until the polymerization was terminated. An exponential digital filter was used to calculate the average flow rate of gaseous monomers during each period that the supply valves were in the open position. The calculated average gaseous monomer flow rates were used to adjust

64

FLUOROELASTOMERS HANDBOOK

the flow rate of the metering pump delivering the liquid cure-site monomer to the reactor during the same time periods. For the peroxide-curable VDF/HFP/TFE elastomer, the 40-liter reactor was charged with 20 liters of water containing 14 g (0.07%) methyl cellulose (Mn about 17,000 daltons) and was heated to 50°C. Gaseous monomers were charged as listed to bring the reactor pressure to 2.56 MPa: Monomer

Amount, g

Wt %

TFE

183

6.3

VDF

872

29.8

HFP

1,870

63.9

Total

2,925

Part of the monomer charged condensed under these conditions to form liquid droplets. The polymerization was initiated by adding a solution of 20 g diisopropyl peroxydicarbonate (IPP) in 80 g methyl acetate. A solution of 36 g methylene iodide in 44 g methyl acetate was also charged to the reactor; about a third was added at the start and the rest during the feed of the first 1,800 g of incremental monomer. A gaseous incremental major monomer mixture was fed to maintain constant reactor pressure at the controlled temperature of 50°C. The liquid cure-site monomer, 4-bromo-3,3,4,4-tetrafluorobutene-1 (BTFB), was fed in a controlled ratio into the incremental gaseous monomer feed. BTFB was initially fed at a ratio of 0.35% to the digitally filtered value of monomer flow. The ratio was gradually increased to 0.75% to attain an overall average of 0.60% BTFB, based on the total incremental monomer fed. The polymerization rate was approximately equal to the incremental monomer feed rate, and increased from approximately 100 g/h initially to 1,000 g/h after 10 hours. A total of 14,278 g incremental monomer was fed over a 20-hour period in the amounts shown: Monomer

Amount, g

Wt %

TFE

2,736

19.2

VDF

7,056

49.4

HFP

4,486

31.4

Total

14,278

The polymerization was terminated after 20 hours by discontinuing the incremental monomer feed. After degassing, the resulting polymer slurry was filtered and washed. Total dry polymer recov-

ery was 15.4 kg, corresponding to 43% solids in the dispersion. Major monomer composition in the polymer was determined by FTIR, and bromine and iodine cure-site levels by x-ray fluorescence. Polymer composition was 22.1% TFE, 51.4% VDF, 25.7% HFP, 0.54% BTFB, and 0.20% I, close to the goal composition set by incremental monomer feeds. Polymer inherent viscosity was 0.73 dL/g, Mooney viscosity ML-10 (121°C) was 42, and glasstransition temperature, Tg, was -19°C. The addition of the cure-site monomer BTFB in a closely controlled ratio to the incremental monomer feed allowed the polymerization to proceed at a satisfactory rate to form a high molecular weight polymer having a homogeneous distribution of cure sites for good curing characteristics. For the bisphenol-curable VDF/PMVE/TFE elastomer, the 40-liter reactor was charged with 20 liters of water containing 14 g methyl cellulose and heated to 50°C. Gaseous monomers, including the cure-site monomer 2H-pentafluoropropylene (2HPFP), were charged in the amounts listed to bring the reactor pressure to 1.55 MPa: Monomer

Amount, g

Wt %

TFE

45

3.0

VDF

405

27.0

PMVE

600

40.0

2H-PFP

455

30.0

Total

1,505

The polymerization was started by adding an initiator solution of 40 g IPP in 160 g methyl acetate. A gaseous incremental monomer mixture was fed to maintain constant pressure at 50°C. The gaseous cure-site monomer, 2H-PFP, was fed along with the major monomers. The incremental feed rate, approximately equal to the polymerization rate, increased from about 176 g/h initially to about 1,956 g/h at the termination of the polymerization period of 10.7 hours. A total of 12,000 g incremental monomer was fed: Monomer

Amount, g Wt %

TFE

480

4.0

VDF

6,960

58.0

PMVE

4,320

36.0

240

2.0

2H-PFP Total

12,000

4 PRODUCTION OF FLUOROELASTOMERS

65

After termination of the polymerization by discontinuing the incremental monomer feed, the polymer slurry was degassed, filtered, and washed. Total polymer recovery was 12.0 kg, corresponding to 37% solids in the dispersion. Polymer composition and properties are listed, with the ratio of 2H-PFP to PMVE determined by 19F nmr: Inherent viscosity, dL/g Mooney viscosity, ML-10 (121°C)

0.81 43

Composition, wt % TFE

3

VDF �

59

PMVE �

36

2H-PFP �

2

Glass transition temperature, Tg, °C –31 Curing characteristics and physical properties of cured compounds were determined for the medium-viscosity bisphenol-curable polymer above and a high-viscosity peroxide-curable commercial polymer made by continuous emulsion polymerization. The commercial polymer, Viton GLT®, has a composition 10% TFE, 54% VDF, 35% PMVE, and 1.2% BTFB, and has inherent viscosity about 1.3 dL/g, and a Mooney viscosity ML-10 (121°C) about 90. As shown in Table 4.5, cure rates and physical properties are similar, but the bisphenol-cured compound gives much better mold release and better retention of properties after heat aging at 250°C. Peroxide Luperox 101XL is 2,5-dimethyl-2,5di(t-butyl peroxy)hexane, 45% on an inert filler. Tremin EST is an epoxysilane-treated wollastonite mineral filler. This formulation, with the special filler molecular sieve zeolite, and metal oxides (but no calcium hydroxide), is advantageous for bisphenol curing of VDF/PMVE/TFE elastomers containing the reactive 2H-PFP cure-site monomer.

4.6 � Process Conditions and Polymer Characteristics Processing behavior, curing characteristics, and vulcanizate physical properties of fluoroelastomers are largely set by polymerization process conditions. Molecular weight distribution is important for most polymer compositions, and can be varied consider-

ably by choice of polymerization process and operating conditions. The nature of chain end groups, determined by initiation and transfer reactions, may affect both processing and curing behavior. Polymer composition and monomer sequence distributions affect suitability for various end uses.

4.6.1

Molecular Weight Distribution

Little information on molecular weight distribution of commercial fluoroelastomers has been published. The usual method of size exclusion liquid chromatography (SELC; also known as gel permeation chromatography, GPC) is not easy to apply. SELC measures macromolecule size in solution, which varies with polymer composition as well as molecular weight. Reliable calibrations exist for only a few VDF copolymer compositions. Several TFE copolymers are so resistant to fluids that solvents suitable for SELC measurements are not available. However, some generalizations can be made, especially for VDF/HFP/TFE and VDF/PMVE/TFE fluoroelastomers, on the variation of molecular weight distribution with polymerization process condition. For these polymer families, average monomer unit weight is about 100 daltons, and polymers with low to medium-high bulk viscosities have a number-average molecular weight, Mn, in the range 60,000 to 120,000 daltons corresponding to 600 to 1,200 monomer units per chain. Bulk characteristics such as viscosity are related to weight-average molecular weight, Mw, which varies from 1.2 to 8 or more times Mn, depending on the distribution set by the polymerization process and operating conditions. Older products, such as Viton® A and B, made by the original DuPont continuous emulsion polymerization process with no added soap or transfer agents have relatively broad molecular weight distribution, with Mw/Mn about 4 to 8. The large particles (about 1 µm in diameter) contain many growing radicals. Termination is hindered, but may involve a combination of long-chain radicals as well as a combination of long-chain radicals with entering oligomeric radicals, leading to broad distribution. A similar semibatch operation, with low soap and use of an initiator level to set overall polymer viscosity, also results in broad molecular weight distribution. Such polymers and their compounds have high green strength and modulus, but poor extrusion characteristics.

66

FLUOROELASTOMERS HANDBOOK

Table 4.5 Comparison of Curing Characteristics and Physical Properties[38]

Formulation, phr Polymer Tremin 283600 EST filler

GLT

Suspension Polymer

100

100

45

45

MT Black, Thermax FF N990

2.5

2.5

Calcium oxide VG

–

6.0

MgO, Elastomag 170

–

1.0

Molecular sieve 13X

–

3.0

Bisphenol AF

–

2.0

Tetrabutyl ammonium hydrogen sulfate

–

0.5

Ca(OH)2, Rhenofit CF

5

–

Peroxide, Luperox 101XL 45

2

–

Triallyl isocyanurate, Diak 7

4

–

GLT

Suspension Polymer

Process aid, octadecyl amine, Armeen 18

0.5

–

Process aid, rice bran wax, VPA 2

1.0

1.0

ML, dN·m

3.9

2.2

MH, dN·m

22.9

23.8

Formulation, phr

Cure Characteristics (MDR, 180°C)

ts2, minutes

0.52

0.29

t´50, minutes

0.93

0.42

t´90, minutes

2.74

2.70

Tensile Properties M100, MPa

14.3

8.2

TB, MPa

18.5

12.0

EB, %

153

176

Hardness, Shore A

75

74

Compression set (disks), % (70 h @ 200°C)

32

37

4 PRODUCTION OF FLUOROELASTOMERS Newer types made by either continuous or semibatch emulsion polymerization use added soap to get smaller particle size and chain-transfer agents to control polymer viscosity. These have narrow molecular weight distribution, with Mw/M n about 2–3. Such polymers and their compounds exhibit relatively low green strength and modulus, but have good flow and extrusion characteristics. Perfluorocarbon diiodide modifiers in semibatch emulsion systems with very low initiator levels may attain “living radical” polymerizations, resulting in fluoroelastomers with very narrow molecular weight distributions, Mw/Mn about 1.2–1.5.[40] Other iodidemodified polymers made with higher initiator levels and optional cure-site monomers in continuous or semibatch emulsion systems have somewhat broader molecular weight distributions, with Mw/Mn about 1.8–2.5.[41] When bromine- or iodine-containing curesite monomers are incorporated in fluoroelastomers made by continuous emulsion polymerization with little or no added chain-transfer agents, these reactive sites may undergo transfer and branching reactions. The long chain branches give considerable high molecular weight fractions, and broad distributions,

67 with Mw/Mn about 4–8. Extensive branching and gel formation may occur in semibatch polymerization, since all polymer made stays in the reactor until polymerization is stopped. Such highly branched, broad distribution polymers give marginal to poor processing characteristics. Figure 4.8 illustrates characteristic molecular weight distributions produced by the three process variations described for polymers made with iodide transfer and/or bromine-containing curesite monomers. Operating conditions in continuous emulsion polymerization or semibatch emulsion or suspension systems can be manipulated to get tailored bimodal molecular weight distributions. To obtain a bisphenol-curable VDF copolymer with good processing characteristics, a blend of a major modified low-viscosity (LV) component with a high-viscosity (HV) component is made by cyclic operation of a single continuous emulsion polymerization reactor.[42] HV component is made with a low persulfate initiator level for a period of at least six reactor turnovers; then a chain-transfer agent is fed for a longer period of time to make the LV component. HV and LV periods alternate in a series of cycles of several hours

Figure 4.8 Fluroelastomer molecular-weight distributution. ­

68

FLUOROELASTOMERS HANDBOOK

each, with conditions otherwise set to maintain a nearly constant polymerization rate and polymer composition. Effluent dispersion from the reactor is blended in tanks downstream before isolation of the bimodal polymer. Operation of a semibatch reactor is readily adapted to making bimodal polymers (see Sec. 4.5 for an example involving suspension polymerization). The reactor is started up with a low initiator level to make the HV component; then a transfer agent is fed to make the LV fraction desired.

4.6.2 �

End Groups

Three kinds of end groups are important for fluoroelastomers: ionic, nonionic, and reactive ends. The types of chain ends may largely determine the product processing and curing characteristics. The process variations discussed in the previous section give varying molecular weight distributions and also result in different end groups. Ionic end groups form from the inorganic initiators used in emulsion polymerizations. Transfer reactions with anionic soaps may also contribute to ionic end groups. Persulfate initiation results in a mixture of sulfate and carboxylate end groups in VDF copolymers, or in carboxylate end groups in TFE/ PMVE perfluoroelastomers. Redox systems, such as persulfate-sulfite, give sulfonate end groups. These ionic end groups increase the bulk viscosity of polymers and compounds by forming ionic clusters that act as chain extenders or temporary crosslinks. The effects are larger for polymers with higher fluorine content. Perfluoroelastomers made with full redox initiation and no chain-transfer agents contain sulfonate ends, which form clusters that are stable at the usual processing temperatures. The compounds are very difficult to mix and form into parts. VDF copolymers made with high persulfate initiator levels may have enough ionic end groups to interfere with bisphenol curing. The ionic ends tend to tie up variable fractions of the quaternary ammonium or phosphonium accelerators used, leading to variable cure rates. Residual soap and oligomers with ionic ends may also affect bisphenol curing. Ionic ends have little effect on radical curing, but these acidic ends may cause some premature decomposition of the organic peroxides used for curing. Ionic end groups contribute to compression set of o-ring seals. Ionic ends may be labile enough to form clusters when the seal is under strain

at high temperature. When the seal is cooled, the secondary network of ionic clusters prevents full recovery of seal shape and sealing force. Nonionic end groups form from the use of organic chain-transfer agents in emulsion polymerization, or from organic peroxide initiators used in suspension polymerization. Fluoroelastomers with mostly nonionic end groups have lower bulk viscosity, lower green strength of uncured compounds, and lower modulus and tensile strength of vulcanizates compared to similar composition with predominately ionic end groups. The polymers with nonionic end groups exhibit better compound flow and bisphenol cure characteristics. Compression set resistance is improved, since the nonionic ends do not impede shape recovery on relief of strain. Reactive ends are mainly formed from use of iodide transfer agents. When enough iodide end groups are present, the chains can be linked by attachment of multifunctional crosslinking agent to chain ends. The resulting networks can attain very good compression set resistance in seals.

4.6.3 �

Composition and Monomer Sequence Distributions

In the usual operation of a continuous or semibatch reactor, the monomer feed composition is essentially constant, and the reactor contains a constant composition of unreacted monomer. Under these conditions, copolymer composition is constant, with a very narrow overall composition distribution. However, the same copolymerization kinetics (reactivity ratios) that determine overall polymer composition as described in Secs. 4.3.2 and 4.3.3 also allow for the presence of monomer sequences that may differ considerably from the overall average composition. The fraction and length of certain monomer sequences may affect polymer characteristics such as the tendency to crystallize. Reactor operation may also be manipulated to produce blends of different compositions or block copolymers containing segments of different compositions within the same chain. Both of these situations are discussed in this section. W. Ring[43] calculated monomer sequencing in dipolymers by considering the relative probabilities of each monomer adding to a given radical end. The probability P11 of Monomer 1 adding to a radical ending in a Monomer 1 unit is given by:

4 PRODUCTION OF FLUOROELASTOMERS

69

k11 [M 1 ] P11 = k11 [M 1 ] + k12 [M 2 ]

tially amorphous and where they have significant crystallinity. For VDF/HFP copolymers, Sec. 4.3.3 notes that, to a good approximation, the HFP monomer does not add to a radical ending in an HFP unit, so r2 = 0 and the copolymer composition relationship reduces to Y = r1X + 1. The monomer addition probabilities P11 and P12 given by Eqs. 4.55 and 4.56 can then be expressed in terms of polymer composition Y, the ratio of VDF to HFP units in the copolymer:

Eq. (4.55)

=

r1 X r1 X +1

Similarly, the probability P12 of Monomer 2 adding to a Monomer 1 radical end is: Eq. (4.56)

P12 = =

k12 [M 2 ] k11 [M 1 ] + k12 [M 2 ]

Y −1 Y

Eq. (4.60)

P12 =

1 Y

The bulky –CF3 of HFP is attached directly to the polymer chain, crowding adjacent groups to produce severe steric hindrance, and thus stiffens the chain to reduce segment mobility for 1–2 VDF units on either side of an HFP unit. Thus, a long sequence of some 12 VDF units seems to be the minimum length required for crystallization with other similar sequences. Using α = 12 in Eq. 4.58, crystallizable fractions for VDF/HFP copolymers of various compositions can be estimated, as shown in Table 4.6. The average VDF sequence length for each composition is Y. The last column gives an estimate of the maximum crystallizable fraction in each co-

P1(n) = P11n-1P12

The fraction Q1 of Monomer 1 units in sequences α or longer in length is: Eq. (4.58)

P11 =

1 = 1 −P11 r1 X +1

The probability P1(n) of a sequence containing n Monomer 1 units is then: Eq. (4.57)

Eq. (4.59)

Q1 = αP11α-1 – (α – 1)P11α

The weight fraction of copolymer in the form of Monomer 1 sequences α or longer is w1Q1, where w1 is the total weight fraction of Monomer 1 in the copolymer. These relationships can be applied to fluoroelastomer families of interest to determine composition ranges where the copolymers are essen-

Table 4.6 Crystallizable Fractions of VDF/HFP Copolymers (α = 12) ­

Mol % VDF

100 w1, Wt % VDF

Y, VDF/HFP

P11

Q1

100 w1Q1, % cryst’n

70.1

50

2.34

0.573

0.013

0.6

74.1

55

2.86

0.651

0.043

2.4

77.9

60

3.52

0.716

0.104

6.2

71.3

65

4.35

0.770

0.200

13.0

84.5

70

5.47

0.817

0.327

22.9

87.5

75

7.03

0.858

0.474

35.6

90.4

80

9.38

0.893

0.628

50.3

93.0

85

13.28

0.925

0.773

65.7

70

FLUOROELASTOMERS HANDBOOK

polymer. Actual crystallinity would be less, and would depend on thermal history of the copolymer (e.g., rate of cooling from the melt, annealing time, and temperature). These calculations are in general accord with observations for various VDF/HFP copolymer compositions. Copolymers containing 60% or less VDF have little or no crystallinity, and are amorphous elastomers. Copolymers containing 65%– 70% VDF have significant crystallinity, with relatively low melting ranges (40°C to 80°C). At higher VDF contents, the copolymers behave as crystalline thermoplastics, with melting ranges increasing with VDF level (100°C to 140°C). Crystalline copolymers with high VDF contents have poor lowtemperature flexibility, even though the glass transition temperature of amorphous regions decreases with increasing VDF content. From Sec. 4.3.3 on reactivity ratios, composition relationships for TFE/PMVE perfluoroelastomers are approximated reasonably well by assuming r1r2 = 0.5. From Eq. 4.14, the following relationship for r1X can be substituted into Eq. 4.55 for estimating P11 values from polymer composition Y: Eq. (4.61)

(

Y −1+ 1 + Y 2 r1 X = 2

to a partially crystalline polymer that gives vulcanizates with a high modulus, but poor low-temperature flexibility. Special reactor operating conditions may be set up to obtain blends or block copolymers with amorphous elastomer and crystalline thermoplastic components. An elastomer in powder form has been made in a two-stage continuous emulsion polymerization system.[44] In one example, an elastomeric VDF/HFP copolymer (58% VDF) was made in the first reactor at high conversion to minimize the amount of unreacted HFP. The total effluent dispersion, including the unreacted monomer, was fed to a second reactor along with the VDF monomer and additional aqueous feed containing initiator. A thermoplastic VDF/HFP component (about 91% VDF) was made in an amount about 29% of the total blend. The dispersion from the second reactor was flocculated and spray-dried to a fine powder. Since the monomer in the second reactor was soluble in the elastomeric particles, the thermoplastic component formed as a separate phase within the elastomer matrix. The composite product, containing about 68% VDF, exhibited a glass-transition temperature of -20°C, characteristic of the elastomeric component and a crystalline melting point of about 140°C, characteristic of the thermoplastic component. The blend was cured with bisphenol to obtain a vulcanizate with a higher modulus and tensile strength than that of the first-stage elastomer alone; elongation-at-break and compression set were comparable. Compared to a physical blend of VDF/HFP elastomer with commercial poly-VDF, the cascade blend gave vulcanizates with a lower modulus, higher elongation, and better compression set resistance. In the cascade reactor operation, relatively few active radicals of the elastomeric component enter the second stage to add thermoplastic chain segments in the same macromolecule. The blend consists mostly of separate chains of the two components.

)

12

The –O–CF3 group of PMVE does not hinder chain segment mobility greatly, with the flexible –O– linkage separating the bulky –CF3 group from the chain. Thus, relatively short TFE sequences are able to crystallize, and a value of α = 8 appears reasonable for the lower limit of crystallizable segment lengths. Estimates of crystallizable fractions for a few TFE/PMVE compositions are shown in Table 4.7. The first two compositions, with TFE content up to about 55%, are nearly amorphous, with little crystallinity likely. The third composition corresponds

Table 4.7 Crystallizable Fractions of TFE/PMVE Copolymers (α = 8) ­

Mol % TFE

100 w1, Wt % TFE

Y, TFE/PMVE

P11

Q1

100 w1Q1, % cryst’n

62.4

50

1.66

0.565

0.074

3.7

67.0

55

2.03

0.622

0.131

7.2

73.9

63

2.83

0.707

0.269

17.0

4 PRODUCTION OF FLUOROELASTOMERS The Daikin “living radical” semibatch emulsion polymerization process can be used to make block copolymers with segments of different composition.[45] An iodine-terminated VDF/HFP/TFE elastomeric component is made with perfluorocarbon diiodide and a small amount of initiator in a first stage of operation. Unreacted monomer is removed, the dispersion is recharged to the reactor, and polymerization is continued with VDF or TFE/E added to make thermoplastic chain segments attached to central elastomeric segments of block copolymer macromolecules. The copolymer can be compounded and cured by the usual elastomer processing techniques. However, it can also be molded as a thermoplastic at a relatively low temperature, then can be removed from the mold and optionally cured by raising the temperature of the part.[46] Adapting a similar diiodo transfer process, Carlson developed A-B-A segmented thermoplastic elastomers with compositions more resistant to strong base and solvents.[47] In one embodiment, the central elastomeric B blocks have the base-resistant composition E/TFE/PMVE, and the outer thermoplastic blocks are E/TFE. Uncompounded molded parts have good properties without curing, but may be cured by ionizing radiation for enhanced properties.

4.7

Monomer Recovery

In the continuous emulsion polymerization process, as shown in Fig. 4.5, pressure is let down at the reactor exit so that the unreacted monomer flashes in the line leading to a degassing vessel. A small amount of defoamer may be added to avoid entrainment and carryover of polymer dispersion from the degasser. Effluent from the degasser goes to a second vessel at a lower pressure to remove most of the rest of the monomer by diffusion from the small particles. The vaporized monomers may be recycled directly back to the suction of the reactor feed compressor. However, it is usually more convenient to take the unreacted monomers through a recycle compressor to a monomer recovery tank. The monomer can then be fed at a controlled rate from the recovery tank to the reactor feed compressor. Less volatile components in reactor efflu-

71 ent dispersion are usually not recovered in this process; they are removed in polymer drying. In semibatch polymerization processes, as shown in Fig. 4.6, the monomer is recovered after completion of the polymerization batch. It is possible to vent unreacted monomer directly from the reactor to a monomer recovery compressor and hold tank. However, control of foaming and carryover is difficult with the limited head space in the reactor. Usually, the dispersion is let down at a controlled rate to a degassing vessel maintained at a low pressure. If a polymer of a similar composition is to be made in the next reactor batch, a heel of dispersion plus a considerable fraction of the unreacted monomer may be left in the reactor. Monomer vaporized from the degasser goes through a compressor to a recovery tank. The degassing vessel may be heated to allow final monomer removal by unsteady state diffusion at low pressure. The recovered monomer is fed to the reactor feed compressor to provide a proper composition for starting the next polymerization batch.

4.8

Isolation

In fluoroelastomer emulsion polymerization processes, dispersions are stabilized by anionic soaps, oligomers, and end groups. Salts of aluminum, calcium, or magnesium are usually used to cause coagulation into particles of convenient size (about 1 mm diameter) for washing and separation by filtering or centrifuging. The coagulant metal ion is chosen for its effectiveness at low concentration and also to keep the soap in solution to facilitate its removal. Coagulation conditions (temperature, holdup, concentration) are controlled to get reliable crumb size for washing, separation, and drying. In the original DuPont continuous emulsion polymerization process for VDF/HFP/(TFE) elastomers, described in Ref. 27 and shown in Fig. 4.9, isolation is also a continuous operation. Potassium aluminum sulfate solution is added to the dispersion in an agitated tank to produce a slurry of crumb that is fed to a continuous centrifuge for removal of most soap and salts. Crumb from the centrifuge is suspended in fresh water in a second wash tank. The slurry is again centrifuged and the wet crumb is fed to a continuous-belt conveyer air oven dryer. Dry crumb is

72

FLUOROELASTOMERS HANDBOOK

Figure 4.9 Isolation by crumb washing and drying.

taken to a crumb blender, then fed to an extruder to produce the final form of pellets or sheet for packaging. Isolation of a polymer from a semibatch emulsion process is similar. Batch coagulation may be carried out by adding coagulant to the dispersion in a stirred tank. Polymer crumb may be separated by filtration or centrifuging, and washed to remove residual soap and salts before drying in an air oven or extruder. In a different version of a continuous isolation process,[48] polymer dispersion is pumped through a coagulation section to a dewatering extruder, as shown in Fig. 4.10. Coagulant is added in-line, and conditions are set to produce large polymer agglomerates. The extruder is set up in a vertical configuration with the inlet in the top section with a large diameter screw. Water containing residual soap and salts is removed from the top. The system is maintained under enough pressure to collapse vapor bubbles that could otherwise cause some polymer crumb to rise to the top water exit. The screw picks

up the polymer and compresses it in a metering section with a smaller diameter to force almost all of the water out of the top of the machine. Polymer with less than 5% water content exits the bottom outlet of the extruder. Final drying is carried out in a vented drying extruder. Since the dewatering extruder removes some 99% of the water in the original dispersion, most water-soluble soap and salts are also removed. Such a system is adequate for continuous emulsion polymerization systems using modest soap levels for dispersion stabilization. Extrusion isolation systems can also be used in semibatch emulsion polymerization processes. However, the higher soap levels used may necessitate a separate coagulation and crumb-washing step before extrusion. A major fraction of the bisphenol-curable VDF/ HFP/TFE fluoroelastomers produced is sold as a precompound, rather than a gum polymer. The isolated gum polymer from a process described above is sent to a compounding facility for incorporation of bisphenol crosslinking agent, accelerator, and optional

4 PRODUCTION OF FLUOROELASTOMERS

73

Figure 4.10 Extruder isolation system.

processing aids. The precompound compositions are usually proprietary, and are designed for specific end uses and fabrication methods. The advantage of precompounds to customers is that the supplier assures good dispersion of curatives and reproducible processing characteristics. Conventional rubber compounding equipment is used, usually an internal mixer and a sheet extruder.

4.9

Process Safety

The major hazards in fluoroelastomer production processes involve handling of toxic or potentially explosive monomer mixtures. Hazards relating to individual monomers are discussed in Ch. 3. In some cases, mixtures may be less hazardous than one or more of the monomers present. For example, the explosive potential of TFE or VDF is reduced in mixtures containing HFP or PMVE. In other cases, mixtures may be more energetic than the individual monomers (e.g., TFE with olefins such as propylene or ethylene). Explosivity testing is necessary to establish the explosion potential of various mixtures at conditions encountered in production facilities. Such testing can be used to establish ratios of required pressure relief areas to volumes of monomer under pressure. Tests can also establish ranges of

monomer composition, pressure, and temperature that can be allowed in plant operation. Systems designed for VDF/HFP copolymerization may not be suitable for TFE/propylene polymerization, for example. In addition to proper design of relief systems, it may be necessary to provide additional protection to personnel by barricading some systems to avoid consequences of possible compromising of relief devices by polymer plugs. Volumes of monomers under pressure should be minimized in monomer feed and polymerization reactor systems. This may be difficult for semibatch systems, as discussed in semibatch reactor design and operation Sec. 4.4.3. Propagation of deflagration pressure pulses from one vessel to another should be prevented by proper placement of relief devices and by minimizing line sizes. Potential ignition sources for monomer deflagration should be minimized by proper design of the system and by proper operating procedures. Electrical systems should not produce arcs or sparks, and surface temperatures should be limited. Electrical energy in instruments should be lower than levels necessary for ignition of monomer mixtures. Metal parts of moving equipment such as agitators should be designed to avoid metal-to-metal contact that could produce sparks or hot surfaces. Trace oxygen levels in monomers should be monitored, and operating steps should be taken to remove oxygen

74 and air from the systems. Trace oxygen can lead to initiation of polymerization through formation and decomposition of monomer peroxides. Polymerization in high-pressure monomer feed systems can give plastic compositions that can cause plugging or local hot spots that might initiate deflagration. Monomer piping should be as direct as possible, avoiding sharp elbows and tees to closed pipe sections, with dead volumes that cannot be readily flushed. In addition to having operating procedures set up to avoid monomer hazards during normal production operation, adequate procedures for equipment maintenance and modification are needed. A large fraction of mishaps causing injury or equipment damage have occurred during mechanical maintenance rather than production operation. Special attention must be paid to assuring adequate flushing and clearing of equipment before mechanical work is started. Similar attention is necessary when putting equipment back into service, especially in careful removal of air before introducing monomers. A number of toxic materials are present in a fluoroelastomer production facility. Very low exposure limits have been established for several major fluoromonomers. Monitoring of work areas is necessary to detect leakage of gaseous monomers, so that steps can be taken to limit exposure of operators to potentially toxic levels. Handling of minor liquid components, such as bromine- or iodine-containing cure-site monomers and transfer agents, may require special procedures and personal protective equipment. For some materials, toxicity information may be limited; these should be handled by procedures adequate to protect personnel from exposure. Peroxide initiators should be handled by procedures recommended by their suppliers to avoid potential hazards caused by decomposition in storage or contact with readily oxidized materials. Considerations for storage and handling of peroxydicarbonates are discussed in Sec. 4.5.2. Monomer compressors should be carefully monitored and controlled to avoid condensation between stages because of the potential for equipment damage or release of excessive amounts of monomers from relief devices. Maximum temperatures should be kept well below those at which monomer decomposition can occur. Trace oxygen levels should be monitored and controlled to avoid possible polymerization in compressor systems which have the potential to plug or initiate monomer deflagration.

FLUOROELASTOMERS HANDBOOK In continuous emulsion polymerization at normally high conversion, loss of reaction (e.g., by interruption of initiator feed or by introduction of a retarder) may lead to the rapid buildup of monomer mixtures with an increased explosivity hazard and with volume flows above the handling capacity of the downstream degassing equipment. Quick action is required to shut off the monomer feeds and to clear the monomers from the reactor by continuing the water feed. The reason for the loss of reaction should be established and corrected before restarting with the normal operating procedure. Proper design and operating procedures are necessary to assure the safe operation of other equipment normally present in chemical process plants, including pumps, agitated vessels, conveyors, extruders, and the like. These will not be covered here.

4.10 Commercial Process Descriptions Commercial continuous fluoroelastomer production facilities used by DuPont Dow Elastomers have the general configuration depicted in Fig. 4.5 for continuous polymerization and monomer recovery, with isolation carried out either by crumb handling as shown in Fig. 4.9 or by extruder dewatering and drying as shown in Fig. 4.10. With the wide range of VDF/HFP/TFE and VDF/PMVE/TFE compositions made, polymerization rates per unit volume vary over a wide range. To keep the overall production rate in a reasonably narrow range for good operation and control, it is convenient to have more than one reactor size available in each facility. Then, a relatively small reactor can be used for products with high polymerization rates per unit volume, and a larger reactor is available for slower polymerizing types. This arrangement also allows for optimizing reaction conditions to get the desired polymer characteristics. Other considerations for design, operation, and control of continuous emulsion polymerization systems are discussed in Sec. 4.4.2 under the headings “Continuous reactor design and operation” and “Continuous emulsion polymerization control.” Monomer recovery and isolation systems are described in Secs. 4.7 and 4.8.

4 PRODUCTION OF FLUOROELASTOMERS Less information is available on the various semibatch process facilities operated by the other fluoroelastomer suppliers. Generally, polymerization system configurations are as shown in Fig. 4.6 and as described in Secs. 4.4.3, and 4.7 and 4.8. As in the continuous process, it may often be convenient to have reactors of different sizes available to accommodate the very wide ranges of polymerization rates per unit volume exhibited by different products. Previous discussion of major monomer handling has assumed that these monomers would be

75 fed as gases at temperatures well above critical temperatures for individual components or mixtures. It is also possible to keep the feed monomers in the liquid phase, as in some processes for making plastic TFE copolymers. Feed system pressures need to be high and temperatures low to keep monomers in liquid form. Such a system may be particularly useful for a product like the Daikin perfluoroelastomer, since the perfluoroalkyl vinyl ether used as the major comonomer is a liquid with a high critical temperature.

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