Emulsion Polymerizations in a Pilot-Scale Loop Reactor with Inline

Conversions and particle-size distributions of the pilot-scale loop reactor ... line static mixers can help to maintain emulsion stability and provide...
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Ind. Eng. Chem. Res. 2005, 44, 5483-5490

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Emulsion Polymerizations in a Pilot-Scale Loop Reactor with Inline Static Mixers S. Fan,† J. H. G. Steinke,‡ and E. Alpay*,† Departments of Chemical Engineering and Chemical Technology and of Chemistry, Imperial College London, London SW7 2AZ, U.K.

Emulsion polymerizations of methyl methacrylate were carried out in a pilot-scale tubular reactor configured in a batch-loop mode. The tubular sections of the reactor were fitted with in-line static mixers to incite low-shear mixing. The reactor was used to investigate the influence of different recipes and operating conditions on reaction, particularly on the monomer conversion and the polymer particle-size distribution. Experimental data were compared to equivalent benchscale studies using a conventional stirred flask. A mathematical model was also developed for predicting the temperature dynamics and the conversion of methyl methacrylate polymerization in the pilot-scale reactor. Conversions and particle-size distributions of the pilot-scale loop reactor were found to be very similar to that of the bench-scale studies. The results indicate that inline static mixers can help to maintain emulsion stability and provide a means for good temperature control, without unduly influencing polymer particle-size distribution. 1. Introduction Compared with bulk or solution polymerizations, emulsion polymerization processes can simultaneously achieve relatively high reaction rates and polymer molecular weights. Typically, water is employed as the reaction medium, which provides a cheap, low-viscosity carrier, with the subsequent advantages for effective heat removal, as well as safety and environmental benefits. Although many industrial polymerizations are carried out in the emulsion phase, the majority of emulsion polymerizations take place in batch or semibatch reactors.1-3 Continuous processes can, of course, eliminate batch-to-batch variations and improve process productivity, but these have led to some fundamental limitations for emulsion-based processes. In particular, industrial experiences and theoretical studies have indicated complex dynamic phenomena, such as multiple steady states and sustained oscillations, within continuous-stirred tank reactor (CSTR) and CSTR-inseries processes.2,4,5 Nevertheless, emulsion polymerizations in a tubular reactor are expected to be relatively stable because of the lower extent of backmixing. The large surface area to reactor volume ratio of a tubular reactor also facilitates heat transfer, providing even further control and uniformity of temperature. However, tubular reactors may suffer from emulsion destabilization, particularly at low shear rates, as well as fouling and plugging of the tubes.6,7 Two approaches have been adopted in the past to accommodate such problems: the use of high flow rates (recycle rates) within the tubes8,9 or the introduction of some kind of pulse operation to incite mixing and create periodic pressure pulses for the alleviation of plugging.10,11 However, there may be a rather complex relation between the hydrodynamics of flow and the extent and nature of the emulsion reaction. For example, Bataille and co-workers have found opti* To whom correspondence should be addressed. E-mail: [email protected]. † Department of Chemical Engineering and Chemical Technology. ‡ Department of Chemistry.

mal conversions of emulsion homopolymerization and copolymerization in open-loop and closed-loop tubular reactors.6,7,12 In particular, maximum conversions were found to be achieved in the laminar-turbulent transition flow region. Lower conversions under turbulent flow were attributed to the reduction in the monomer droplet size and, subsequently, an increase in the droplet surface area. This was then postulated by the authors to lead to greater surfactant consumption and thus a reduction in the micelle number and subsequent conversion of monomer. Under laminar flow, however, lower conversions were explained by monomer mass transport limitations, as well as phase separations (emulsion destabilization) due to poor mixing. In other studies, experimental results in tubular reactors did not show a strong dependence of polymerization on the flow velocity when working within the laminar flow regime.9-11,13 Recently, attention has been given to the use of tubular reactor systems fitted with in-line static mixers.14,15 These have the advantage of inducing the good radial mixing of viscous mixtures under relatively low linear flow velocities, thus potentially narrowing the residence time distribution of the reactants and providing relatively good heat transfer to the heat-exchange surface. However, as far as the authors are aware, there is little published work in the use of static mixers for emulsion polymerization systems.16 The use of static mixers for emulsion polymerization may provide controlled and uniform mixing of the reaction medium under relatively low linear flow velocities and, subsequently, an energy-efficient means of temperature and emulsion stability control. Furthermore, if nearideal plug flow can be approached through the use of static mixers,15 then the complex dynamic phenomena associated with well-mixed flow systems (e.g., CSTRs) may be avoided. Batch-loop-mode operation of the tubular system is also possible, which may provide a large-scale analogue for an ideally (consistently) mixed batch vessel but in which there are relatively good heattransfer characteristics. In the following sections, pilotscale studies are presented for the emulsion polymeri-

10.1021/ie049783d CCC: $30.25 © 2005 American Chemical Society Published on Web 08/07/2004

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Table 1. Conditions of the MMA Emulsion Polymerizations variable monomer volume fraction emulsifier (mol/cm3 water) initiator (mol/ cm3 water) temperature (°C) stirring/recycle rate (rpm) pre-emulsification time (min)

bench 1 0.2

bench 2 0.3

pilot 1 0.2

pilot 2 0.3

2.0 × 10-5 3.0 × 10-5 2.0 × 10-5 3.0 × 10-5 1.5 × 10-5 1.5 × 10-5 1.5 × 10-5 1.5 × 10-5 60 350

60 420

60 500

60 1000

70

40

15

0

zation of methyl methacrylate (MMA) using a loop-mode arrangement of tubular reactor sections containing inline static mixers. Attention is given to the effects of static mixing on monomer conversion and product particle-size distribution (PSD), and comparisons are made to equivalent bench-scale studies in well-stirred flasks. As a means of guiding further process design, optimization, and scale-up, a mathematical model of emulsion polymerization within the pilot-scale reactor is presented. 2. Experimental Section 2.1. Reactants. The following reagents were used in the experiments: MMA (99%, Lancaster Ltd.), sodium dodecyl sulfate (99%, Sigma-Aldrich), and sodium persulfate (98%, Sigma-Aldrich). All materials were used as received. Distilled water was used in all of the experiments. 2.2. Bench-Scale Experiments. As a reference to the pilot-scale experiments, MMA emulsion polymerizations were first performed in a 250-mL three-necked, round-bottom flask fitted with a condenser. Agitation was achieved using a four-blade stainless stirrer (Janke & Kunkel RW20 DZM). Temperature control was achieved by a constant-temperature oil bath. The reaction vessel was first charged with the required amount of water, emulsifier (sodium dodecyl sulfate), and monomer (MMA), with the recipes given in Table 1. To remove traces of dissolved oxygen, the reactant mixture was purged with argon for at least 1 h at room temperature under agitation. The reactor was then heated in the oil bath to the desired temperature. Once the temperature was reached, the emulsion polymerization was started by injecting the initiator (sodium persulfate). During the polymerization, aliquots of approximately 6-mL samples were taken at specified times via an airtight poly(tetrafluoroethylene)-lined syringe. The aliquots were collected and stored in glass vials prior to offline analysis (conversion and PSD). The reaction was quenched by bringing the sample into contact with air. Conversions were determined gravimetrically (see, e.g., ref 17). Approximately 2 mL of the sample was transferred onto a preweighed aluminum foil plate (about 0.4 g). The thin emulsion film was dried at 70 °C in a vacuum oven to remove water and residual monomer. The samples were dried until no further reduction in the sample weight was observed, which took about 2 h. The relative errors of conversion measurements were less than 1%. Several independent bench-scale runs indicated good experimental reproducibility, i.e., deviations in measured conversion of less than 3%.

Figure 1. Pilot-scale polymerization reactor system.

PSDs were measured using a Malvern Mastersizer 2000 with a Hydro 2000SM sample dispersion unit. The instrument could measure latex particle sizes in the range of 20 nm to 2 mm. While a micelle with a diameter of about 5 nm could not be captured by this instrument, both the monomer droplets and the larger polymer particles could be measured properly. A high surfactant concentration ensured sample stability over the period of analysis of approximately 48 h. 2.3. Pilot-Scale Batch-Loop Experiments. The pilot-scale polymerization reactor is shown schematically in Figure 1. It consists of 316 stainless steel tubular elements of 32 mm inside diameter and a total length of 5.4 m. The tubes are fitted with Sulzer SMXL static mixers, which are also of stainless steel construction. Three independent heat exchangers control the different sections of the reactor jackets. A gear pump in the recycle loop provides the flexibility of several operation modes. The loop section has an effective volume of 1.76 L. For the purpose of batch-loop operation, the loop and straight-pass sections of the reactor (see Figure 1) were separated through the insertion of a blank plate at a suitable flange point. The reactor additionally has three microdosing pumps and a data acquisition system. Online data monitoring and control software (HP VEE version 5.01) is used to store the reactor temperature and pressure readings every 30 s to a PC. Further details about the polymerization reactor system and its operation can be found in work by Fan et al.15 Before being charged to a reactor feed tank, the aqueous emulsifier and monomer mixture was degassed in a glass tank. After 2 h of degassing, transfer was carried out under an argon atmosphere. Initiator powder was then added into the solution and degassed for another 0.5 h. At the start of the run, the reactor was empty and the gear pump was shut down. The far end of the reactor tube (the devolatilization section) was closed by a pressure control valve. The positive dosing pump was set to full stroke and 1000 rpm, with an inflow rate of 200 mL/min. About 3 L of the reactant mixture was pumped into the reactor, using a recipe identical with that for the corresponding bench-scale experiment, given in Table 1. After all reactants were pumped in, the recycle gear pump was started to achieve proper mixing inside the reactor. At a recycle rate of 500 rpm (4.7 cm3/revolution), the Reynolds number was about

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2000. Thus, the process was operated as a batch-loop reactor, with the rest of the reactor as a reactant reservoir to ensure the loop was filled with liquid the entire time during the experiment. The positive pressure of 1.5 bar inside the reactor facilitated the removal of sample aliquots. Two different operation procedures were investigated: In the experiment pilot 1, the reactants were first recycled for 15 min, during which an emulsion was formed, the reactor was then heated from room temperature to the specified reaction temperature. In the experiment pilot 2, the reactants were directly heated to the reaction temperature without allowing time for emulsification. The beginning of the heating program was set as time zero in both of the experiments. The reactor reached the specified polymerization temperature in less than 10 min (see Figure 5). Samples were taken from the sampling point located at 0.89 m from the inlet. Monomer conversions (gravimetrically) and PSDs (by a Malvern Mastersizer 2000) were analyzed as described for the bench-scale experiments. Finally, reactor cleanup was achieved by releasing the vessel contents through a discharge valve located at the bottom of the loop section and purging the loop with distilled water at 70 °C. Occasionally, a xylene wash was carried out to dissolve any polymer deposits. However, constant drop readings across the loop section indicated no plugging of, or significant polymer buildup on, the static mixers. 3. Mathematical Modeling There exist many uncertainties in the modeling of emulsion polymerization processes.18-20 In this paper, the model was developed mainly to help in the analysis of the performance of the pilot-scale reactor. Key model assumptions are summarized below: (a) The polymer particles are monodisperse when the reactor is operated in batch mode. (b) Polymer particles are generated by both micellar and homogeneous nucleation. (c) No particle coagulation occurs under the high emulsifier concentrations used in this research. (d) The quasi-steady-state assumption is valid for describing radical concentrations in the aqueous phase. (e) Monomer concentrations in the different phases are in thermodynamic equilibrium. An emulsion polymerization process can exhibit three stages.21 In stage I, the emulsifier molecules aggregate into micelles, with the typical diameter of 5 nm for a monomer-swollen micelle. Beyond about 5% conversion (stage II), the micelles disappear while the monomer droplets are still present. The diameters of monomer droplets are on the order of 10 µm and decrease in size with the progress of polymerization. Stage II continues until the monomer droplets disappear. Then (stage III), with a scarce amount dissolved in the aqueous phase, monomer is only present within swollen polymer particles. The polymer particles grow in size until nearly complete monomer conversion. The diameters of the final polymer particles are typically in the range of 100 nm. Further details on the emulsion polymerization mechanism can be found in work by Dube et al.22 and Coen et al.23 In this work, the usual free-radical polymerization mechanism is assumed. The kinetic models proposed by Ray and co-workers are employed, which

have been found to be consistent with various experimental results and are relatively simple to implement within a mathematical model.24-26 For the MMA emulsion polymerization in the pilotscale batch-loop reactor, the model consists of six state variables: the polymer particle concentration, Np, the concentrations of initiator, Ci, total monomer, Cm, and polymer, Mp, and the jacket, Tj, and reaction, T, temperatures. Such a lumped parameter system can be described by a set of ordinary differential equations based on mass and energy balances. 3.1. Mass Balances. The balance equation for the polymer particle concentration, Np, is

dNp/dt ) Rmic + Rhom

(1)

where Rmic and Rhom are the micellar and homogeneous particle nucleation rates, respectively. The mass balance equations for initiator, overall monomer, and polymer concentrations can be written as

dCi/dt ) -kdCi

(2)

dCm/dt ) -kpCmpNpjı

(3)

dMp/dt ) kpCmpNpjı MWm

(4)

where kd and kp are kinetic constants, Cmp is the monomer concentration in polymer particles, jı is the average number of radicals per polymer particle, and MWm is the molecular weight of monomer. The monomer conversion is defined as the weight fraction of polymer

Xm ) Mp/MWmCm0

(5)

where Cm0 is the initial monomer concentration. 3.2. Energy Balances. The energy balance for the batch-loop reactor is given by15

(mixFmixCp + rFrCp,r)V

dT ) Qr - UA(T - Tj) + Qp dt (6)

where mix is the void fraction of the reactor in the presence of the static mixers, r is the volume fraction of the static mixers, Fmix and Fr are the densities of the reactant mixture and steel, Cp and Cp,r are the corresponding heat capacities, and V is the total reactor volume. On the right-hand side of the above equation, UA is the overall heat-transfer coefficient, and the reaction heat, Qr, is given by

Qr ) kpCmpNpjı mixV(-∆HP)

(7)

The heat generated by the recycling pump, Qp, can be estimated from the steady-state temperature measurements when the reaction almost ceases, and thus Qr and dT/dt are 0. The jacket temperature, Tj, is modeled by first-order dynamics

τj

dTj + Tj ) Tjsp dt

(8)

where τj is the jacket time constant and Tjsp is the setpoint of the thermostat.

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3.3. Kinetic Model. 3.3.1. Nucleation Equations. A collision model for radical entry is used to describe the micellar nucleation

Rmic ) (kmm4πrm2Na)CmicCRφw

(9)

and the homogeneous nucleation rate is given by

Rhom ) kpCmwCRjcrφw

(10)

where kmm is the constant of radical entry into micelles, rm is the radius of a micelle, Na is Avogadro’s number, Cmic is the micelle concentration, CR is the concentration of radicals in the aqueous phase that may enter particles, Cmw is the monomer concentration in the water phase, CRjcr is the concentration of radicals in the aqueous phase with a length of jcr, and φw is the volume fraction of water in the emulsion. 3.3.2. Emulsifier Partitions. The total emulsifier concentration, Ce, remains constant during the reaction, but its partition within water, micelles, and the particle surface changes. The micelle concentration, Cmic, is calculated from the partition equation of the emulsifier

Ce ) Cewφw + 4πrm2Cmicφw/ae + 4πrs2Np/ae (11) In the above equation, Cew is the emulsion concentration in the aqueous phase, whose maximum is the critical micelle concentration, Scmc, and ae is the area of a micelle or polymer particle covered by an emulsifier molecule; the surfactant surface coverage is assumed to be equal to unity. It is noted that when Cew < Scmc, the micelle concentration, Cmic, is zero. The average radius of polymer particles swollen with monomer, rs, is calculated from the polymer weight and particle number

4/3πrs3NpNa(1 - φm)Fp ) Mp

(12)

where Fp is the density of the polymer. 3.3.3. Monomer Partitions. The monomer partitions among the droplets, the water phase, and polymer particles. When monomer droplets are present, polymer particles are saturated, and the monomer volume fraction, φm, is at its maximum

φm ) φmsat

(13)

In stage III, when monomer droplets disappear, the volume fraction can be determined by the equilibrium relationship24,26

ln(Cmw/Cmwsat) ) 1 - φm + ln(φm) + ψ(1 - φm)2 (14) and

Cm ) φwCmw +

MpCmp Fp(1 - φm)

CR ) CRw/2

(17)

CRjcr ) CRw/jcr

(18)

In eq 18, it is thus assumed that equal concentrations of free radicals, of chain length 1 to jcr, exist in the water phase. Half of this total concentration of free radicals may be captured by micelles or polymer particles, and thus eq 17. The quasi-steady-state assumption is applied to the aqueous-phase radical concentration, CRw

2fkdCi + D(rs) Npjı /φw ) kt0CRw2 + (kmm4πrm2Na)CmicCR + (kmp4πrs2Na)NpCR (19) When chain transfers occur in polymer particles, the small radicals formed can traverse to the aqueous phase.22 The radical desorption rate, D(rs), is given by24

D(rs) )

3Dmkfm/kp 3Dm/kpCmp + rs2

(15)

(16)

(20)

where Dm is the effective diffusivity coefficient and kfm is the chain-transfer coefficient. 3.3.5. Average Number of Radicals per Particle. The Stockmayer-O’Toole method is used for the computation of the average number of radicals per polymer particle27

jı )

a Ib(a) 4 Ib-1(a)

(21)

where Ib(a) is the modified Bessel function of the first kind and the argument and order are

a ) 4(4/3πrs3Nakmp4πrs2NaCR/kt)0.5

(22)

b ) 2D(rs) 4/3πrs3Na/kt

(23)

3.3.6. Gel and Glass Effect Equations. In emulsion polymerizations, the viscosity in the polymer particles can be high from the onset of the reaction. To account for the diffusion-controlled rate constant change during the process, the following semiempirical relationships were adopted in this work. For the gel effect on the radical termination rate constant, the free-volume theory28 is used:

kt ) gtkt0

where Cmwsat is the saturated aqueous-phase monomer concentration, ψ is the Flory-Huggins interaction parameter, and Cmp is the monomer concentration in polymer particles

Cmp ) φmFm/MWm

3.3.4. Aqueous-Phase Radical Concentration. In the aqueous phase, radicals of chain length between jcr/2 and jcr can be captured by micelles or polymer particles, while those with a length of jcr + 1 undergo homogeneous nucleation to form new latex particles.21 Because the propagation probability of a free radical in the aqueous phase is close to unity, the calculations of CR and CRjcr can be simplified to

gt ) 1.0

for vf > 0.1731

gt ) exp[75(vf - 0.1731)]

for vf e 0.1731

(24) (25) (26)

where vf is the free volume and is given by

vf ) 0.025 + 0.001(T - 167)φm + 0.00048(T - 387)(1 - φm) (27)

Ind. Eng. Chem. Res., Vol. 44, No. 15, 2005 5487 Table 2. Kinetic and Physical Parameters parameter 10-6

reference

Scmc ) 6.4 × rm ) 2.5 × 10-7 cm ae ) 5.7 × 10-16 cm2 kd ) 1.8 × 1017 exp(-143220/RT) s-1 f ) 0.5 kp0 ) 4.92 × 108 exp(18283/RT) cm3/mol‚s kt0 ) 9.8 × 1010 exp(2944/RT) cm3/mol‚s kfm ) 2.0 × 10-5kp Dm ) 1.1 × 10-7 cm2/s Fm ) 0.919 g/cm3 Fr ) 7.83 g/cm3 Fp ) 1.19 g/cm3 Cp ) 4.2 J/g‚K Cp,r ) 0.5 J/g‚K kmm ) 28 cm/s kmp ) 28 cm/s φmsat ) 0.73 Cmwsat ) 1.56 × 10-4 mol/cm3 ψ ) 0.613 jcr ) 16 MWm ) 100.13 g/mol -∆HP ) 5.78 × 104 J/mol mol/cm3

26 24 this work 24 24 24 24 24 24 32 33 32 32 33 24 24 24 24 24 21 32 32

Table 3. Reactor Parameters τj ) 78 s mix ) 0.64

V ) 2891 cm3 UA ) 87.1 J/s‚K

r ) 0.36 Qp ) 150 J/s

For the glass effect on the propagation rate constant, the relationships given by Ballard et al.29 are adopted:

kp ) gpkp0 gp ) 1.0

(28)

for wp e 0.84

gp ) exp[-29.8(wp - 0.84)]

Figure 2. Evolution of the PSD during the emulsion polymerization experiment run pilot 2: (a) stages I and II when monomer droplets are present; (b) the transition from stage II to stage III.

for wp > 0.84

(29) (30)

where wp is the polymer weight fraction in a particle and is given by

wp )

(1 - φm)Fp (1 - φm)Fp + φmFm

(31)

The above model equations were solved in the Matlab Simulink environment (MathWorks, Inc.), using the method of ode15s with a variable-integration step length. The various parameters used in the model are summarized in Tables 2 and 3. 4. Results and Discussion The latex PSD is one of the most important product properties in emulsion polymerization. The evolution of the PSD in the loop reactor experiment pilot 2 is shown in Figure 2. During stages I and II of the MMA emulsion polymerization, the monomer droplets gradually decreased in size, whereas the polymer particles formed and grew in size. As shown in Figure 2a, the distribution curves shifted from micrometer to nanometer diameter ranges. After 10 min from the start of reaction, there were no detectable polymer particles. At 15 min, some polymer particles with diameters bigger than 20 nm formed. At 25 min, the monomer droplets shrank further, resulting in a bimodal curve: one peak representing the polymer particles (smaller) and the other the monomer droplets (larger). Figure 2b shows the transition from emulsion polymerization stage II to stage III. At 30 min, the monomer droplets disappeared completely. Then the polymer particles continued to

Figure 3. Comparisons of the PSDs of the bench- and pilot-scale experiments.

grow in size, with the peak diameter increasing from 112 nm at 30 min to 195 nm at 90 min. Figure 3 compares the PSDs of the pilot-scale loop reactor and the bench-scale experiments. With the same recipes, the corresponding curves of the two reactions are very similar. Because of the lower emulsifier concentrations in the experiments of pilot 1 and bench 1, there were fewer micelles, and thus the final particles were larger than those of data set 2. It is important to note, however, that the accuracy of the Malvern particlesize analyzer is low for sizes of less than 100 nm, and therefore the data provide a qualitative comparison of PSD. In future work, a capillary hydrodynamic fractionator will be employed for accurate particle-size analysis. Such high-fidelity data may then be used for further model verification, i.e., prediction of the average particle size. The monomer conversions for the pilot-scale loop reactor and the bench-scale experiments are compared in Figure 4. For the same recipes and similar operation procedures (pre-emulsification time and stirring/recycle rate), the final conversions were similar. As shown in Figure 4a, the final conversions in bench 1 and pilot 1 were about 70%, well below 100%. This may be due to the lack of stirring or recycle, and thus reaction which

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Figure 4. Comparisons of the monomer conversions of the benchand pilot-scale experiments: (a) gravimetric analysis of runs bench 1 and pilot 1 and model simulation of bench 1; (b) gravimetric analysis and model simulations of runs bench 2 and pilot 2.

is monomer-diffusion-controlled.30 For run pilot 1, the 15-min pre-emulsification could not compensate for the lack of mixing during the reaction, but some monomer polymerized at this stage, leading to a higher initial conversion after 5 min. The model simulation could satisfactorily predict the experimental conversion data. Note that, in this simulation, an effective monomer concentration of 80% of the actual value was used to account for the diffusion-controlled reaction. Under sufficient stirring or recycle conditions, the conversions of emulsion polymerization of MMA could go beyond 90% quickly, as shown in Figure 4b. For the bench 2 experiment, the bath oil temperature was set at 60 °C, but because of the highly exothermic reaction, the actual reaction temperature reached as high as 68 °C. The reaction finished in about 15 min. Compared with this, the pilot 2 experiment in the loop reactor exhibited an induction period of 20 min. Two possible reasons exist for this induction period. First, in the pilot 2 experiment, pre-emulsification was not carried out. The reactants needed some time in the loop to form an emulsion. This was validated by the phase separation of samples taken before 20 min. Second, it also took some time to heat the loop reactor from room temperature to the specified reaction temperature (see Figure 5b). Figure 5 shows the temperature dynamics of the loop emulsion polymerization process. Two sensors installed at different locations inside the tube gave the same reaction temperature, which meant the temperature distributions were uniform at the operation conditions. For the pilot 1 experiment with a 20% charge of monomer, the reaction temperature was very close to that of the jacket and near-isothermal operation was achieved (see Figure 5a). While for the pilot 2 experiment with a 30% charge of monomer, a temperature peak occurred at about 28 min, with a maximum temperature difference of 6 °C between the reaction and the cooling jacket (see Figure 5b). This corresponded to

Figure 5. Jacket and reaction temperatures of the emulsion polymerizations in the pilot-scale batch-loop reactor: (a) online measurements of run pilot 1; (b) online measurements and model simulations of run pilot 2.

the same period of sharp increase of monomer conversions shown in Figure 4b. The bench-scale model assumes isothermal operation, while the pilot-scale model takes into account the jacket and reaction temperature dynamics. As shown in Figure 5b, the jacket temperatures can be described well by the model simulation. The reaction temperatures predicted by the simulation have the same trend and peak values as the experimental data, but it leads by about 6 min. As explained above, this may be due to the lack of pre-emulsification in this experiment. Interestingly, just by applying a simple 6 min shift, both the temperature and the monomer conversion can be accurately predicted. This is shown in Figure 4b, where both the original and shifted modeling results of the pilot-scale reactor are given. 5. Conclusions Emulsion polymerizations of MMA were successfully performed in a pilot-scale batch-loop reactor with inline static mixers. Through the use of a high recycle flow rate, the emulsion process was found to be stable and no plugging occurred. The conversions and PSDs of the final PMMA latex were comparable to that of small bench-scale experiments. Loop operation, in which radial mixing is further facilitated by the use of static mixers, was also found to lead to good temperature control by minimizing the reaction exotherm; further analysis of the temperature dynamics of the reactor can be found in work by Fan and Alpay.31 The model developed could satisfactorily predict the temperature dynamics and monomer conversions. The experimental results also indicate that pre-emulsification can significantly influence the reaction dynamics of the system under batch-loop mode, and hence proper attention to this is needed under continuous operating analogues (i.e., a single-pass tubular reactor containing in-line

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static mixers). The above results provide much impetus for the evaluation of the in-line static mixers under such single-pass, continuous operation, and such work is currently ongoing in our laboratories. Acknowledgment We thank Professor Dame J. S. Higgins for her helpful discussions. S. P. Gretton-Watson and D. L. Palmer for their help in the pilot-scale experiments, and P. Carry for her help in the PSD measurements. We are grateful to the EPSRC for the funding of the Polymers, Properties and Polymerisation Processes (P4) program at Imperial College London. Notation ae ) area occupied by a surfactant molecule, cm2 Ce ) total emulsion concentration, mol/cm3 Cew ) emulsion concentration in the aqueous phase, mol/cm3 water Ci ) initiator concentration, mol/cm3 water Cm ) total monomer concentration, mol/cm3 Cmic ) micelle concentration, mol/cm3 water Cmp ) monomer concentration in polymer particles, mol/cm3 Cmw ) monomer concentration in water, mol/cm3 Cp ) heat capacity of the reaction mixture, J/g‚K Cp,r ) heat capacity of steel, J/g‚K CR ) concentration of radicals in water that can enter micelles or particles, mol/cm3 water CRjcr ) concentration of radicals with critical length in water, mol/cm3 water CRw ) total concentration of radicals in water, mol/cm3 water D(rs) ) radical desorption rate, s-1 Dm ) effective diffusivity coefficient, cm2/s f ) initiator efficiency, dimensionless -∆HP ) heat of polymerization, J/mol jı ) average number of radicals per polymer particle jcr ) critical radical length for homogeneous nucleation kd ) kinetic constant for initiator decomposition, s-1 kfm ) kinetic constant for chain transfer to monomer, cm3/mol‚s kmm ) mass-transfer coefficient for radical entry into micelles, cm/s kmp ) mass-transfer coefficient for radical entry into particles, cm/s kp ) kinetic constant for propagation of free radical, cm3/mol‚s kt ) kinetic constant for termination of free radical, cm3/mol‚s MWm ) molecular weight of monomer, g/mol Mp ) total polymer concentration, g/cm3 Na ) Avogadro’s number, 6.023 × 1023/mol Np ) particle concentration in the reactor, mol/cm3 Qp ) heat generated by the recycle pump, J/s Qr ) heat generated by polymerization, J/s rm ) radius of a monomer-swollen micelle, cm rs ) radius of a swollen polymer particle, cm Rhom ) homogeneous particle nucleation rate, mol/cm3‚s Rmic ) micellar particle nucleation rate, mol/cm3‚s Scmc ) critical micelle concentration, mol/cm3 water T ) reaction temperature, K Tj ) jacket temperature, K Tjsp ) jacket temperature setpoint, K UA ) overall heat-transfer coefficient, J/s‚K V ) total volume of the pilot-scale loop reactor, cm3 vf ) free volume, dm3 Xm ) monomer conversion

Greek Letters mix ) void fraction of the reactor r ) fraction of the static mixers and reactor wall φm ) monomer volume fraction in polymer particles φw ) volume fraction of water in the reactor Fm ) density of the monomer, g/cm3 Fmix ) density of the reaction mixture, g/cm3 Fp ) density of the polymer, g/cm3 Fr ) density of steel, g/cm3 τj ) time constant of the jacket dynamics, s ψ ) Flory-Huggins interaction parameter Subscripts 0 ) initial value sat ) saturated value

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Received for review March 19, 2004 Revised manuscript received June 3, 2004 Accepted July 1, 2004 IE049783D