Laboratory Simulation of Industrial Crystallizer Cycling - Industrial

Ind. Eng. Chem. Res. 1996, 35, 1163-1172

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Laboratory Simulation of Industrial Crystallizer Cycling Ru-Ying Qian and Gregory D. Botsaris* Department of Chemical Engineering, Tufts University, Medford, Massachusetts 02155

The crystal size distribution (CSD) of the product in a large continuous industrial crystallizer usually fluctuates periodically (cycles) even under well-controlled operating conditions. This long-period cycling was simulated in the laboratory in a bench-scale crystallizer (5 dm3) of ammonium sulfate, for the first time. A two-region crystallizer model consisting of a supersaturation production region (SPR) and a supersaturation consumption region (SCR) was proposed. Based on this model, a mechanism for the sustained CSD cycling was developed. This mechanism was used as a guideline in the construction of the laboratory crystallizer, in which a region of high and cycling local supersaturation was created, simulating the evaporative foam layer in the vacuum industrial crystallizers, as well as a bulk suspension with low supersaturation simulating the bulk suspension in the same crystallizers. The success of the crystallizer in achieving cyclic performance reinforces the proposed mechanism. 1. Introduction Crystallization is a widespread separation process in the chemical industry. The large continuous industrial crystallizers which are used for the process, however, are afflicted in many cases by a serious problem. The crystal size distribution (CSD) of the product of such a crystallizer usually fluctuates periodically even under well-controlled operating conditions of flow rates, feed concentration, temperatures, and vacuum pressure. This difficult problem has been studied for many decades, but the cause of CSD cycling is still elusive and as a result no effective way has been found yet for eliminating or reducing such a cycling. One serious difficulty in solving this problem was that the sustained CSD cycling in industry could not be simulated in traditional laboratory crystallizers. Thus, proposed industrial crystallizer models and CSD cycling mechanisms could not be checked by experiments. CSD in a laboratory mixed suspension mixed product removal (MSMPR) crystallizer always approaches a steady-state value after an initial transient state. In the same crystallizers the logarithm of the population density, n, decreases linearly with respect to crystal size L. However, in pilot- or industrial-scale crystallizers ln n increases with respect to L at an intermediate crystal size range and the so-called S-shaped CSD curve results (Bennett et al., 1973; Qian and Chen, 1986). A very small number of experimental investigations on CSD cycling have been performed, mainly because of the long duration of such an experiment; it requires an uninterrupted operation for a few days and a number of well-trained personnel to run it around the clock. Randolph et al. (1977) reported the results from a KClbrine system in an 18 dm3 classified crystallizer with fines dissolving (FDS). Two to four temporary cycles were observed before the experiments were terminated because of severe scaling of the cooling coil surface. However, in this case a product classifier was found necessary in order to get even these temporary cycles. Industrial crystallizers, however, usually have no product classifiers. Jager et al. (1991) used for ammonium sulfate a 970 dm3 crystallizer with an internal heater and FDS, that attained steady state, and found during the transient state more CSD fluctuation than in the * Author to whom correspondence should be addressed. Fax: (617) 627-3991. E-mail: [email protected].

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case of a 20 dm3 crystallizer. They assumed a lowsupersaturation threshold mechanism (i.e., fluctuation between no nucleation and slow secondary nucleation) as the mechanism for the fluctuations. Then they extrapolated their results and predicted a sustained cycling for a 50 000 dm3 crystallizer. However, most industrial crystallizers with severe CSD cycling (e.g., ammonium sulfate and potassium chloride crystallizations) have no internal heater, and even small pilotplant crystallizers, e.g., 250 dm3 for urotropine (Qian, 1986) and 520 dm3 for KCl (Qian and Chen, 1986), can generate sustained cycling. Recently, Eek et al. (1995) proposed another model for the oscillatory or badly dampened CSD dynamics in the above-mentioned 970 dm3 crystallizer. From the correlation of the transient nucleation rate with the total surface area of larger than 600 µm crystals, they suggested a surface breeding secondary nucleation mechanism to describe the CSD oscillations during the transient period. Their finding, however, that the nucleation rate is independent of the supersaturation is not expected for the case of surface breeding. Song and Douglas (1975) crystallized NaCl from aqueous solutions by salting out with ethanol. They reported short operational periods of 1-1.5 cycles in 6-9 retention times. It has been found, however, that the same system under similar conditions attains steady state after a longer operation of about 25 retention times (Timm and Larson, 1968). This paper describes a successful experimental effort to simulate in the laboratory the operation and the cycling of industrial-scale crystallizers; the latter usually have no product classifiers or internal heaters to account for their CSD cycling. To proceed with the design, construction, and operation of a cycling laboratory crystallizer, one needs to develop a CSD cycling mechanism and use it as a guideline. Such a mechanism was developed and is described in qualitative terms in this paper. A quantitative analysis of the mechanism requires a study of the nucleation in the evaporative foam layer and is the subject of an ongoing work. It should be noted that the cycling mechanism is supported by previously reported observations and phenomena in industrial crystallization. In brief the proposed cycling mechanism is based on the existence of a small region of high supersaturation (supersaturation production regionsfor instance, the foam layer in an evaporative crystallizer), in which most © 1996 American Chemical Society

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Ind. Eng. Chem. Res., Vol. 35, No. 4, 1996

of the nucleation of new crystals occurs, and a larger region of low supersaturation (bulk region), in which most of the growth of the crystals takes place. The cycling is the result of the fluctuation of the supersaturation in the former region above and below a threshold; below the threshold the nucleation rate is low and above it it is very high. This fluctuation is created, in simplified terms, by the interaction of the creation of supersaturation by evaporation with the feedback effect of FDS, the growth of the fine crystals in the foam layer, and the crystal classification in the crystallizer. The proposed model differs from all previous models. The latter assumed a well-mixed suspension, leading to a coupling of nucleation and growth. Consequently, the nucleation and growth rates are both high or both low. The two-region model decouples them, creating the possibility of matching a high nucleation rate (in the foam layer) with a very low crystal growth rate (in the bulk of the crystallizer). This is an important element for CSD cycling. 2. Guiding New Mechanism of CSD Cycling The mechanism of CSD cycling in industry is related to the model of industrial crystallizers one formulates. There has been a tendency to approximate the large industrial crystallizers with a well-mixed suspension model and to assume that contact secondary nucleation (especially collisions between crystals and the impeller) is the predominant nucleation mechanism. These concepts, however, fail to explain the S-shaped CSD plot mentioned above as well as the observed CSD cycling in pilot plants or the industry. The proposed mechanism has the following elements: Two-Region Model for Industrial Crystallizers. There is a small region, the supersaturation production region (SPR), and a larger one, the supersaturation consumption region (SCR). The suspension in both regions is well mixed. The transfer of supersaturation from the former to the latter region is hindered by the incomplete hydrodynamic interaction between the two regions. There is also crystal classification during the transfer of crystals from the one region to the other. Difference in Supersaturation in the Two Regions. The supersaturation is high at the point of production, i.e., in the SPR. Consequently, most of the nuclei are produced there. The supersaturation is transferred to the SCR where it is consumed mostly by growth and partly by contact nucleation. The produced nuclei are also quickly transferred to the SCR, increasing the crystal surface area there available for growth. The other reason for most of the growth occurring in SCR is the much longer retention time in that region compared to the SPR. Supersaturation Fluctuation in SPR. When the supersaturation in SPR exceeds a certain threshold value, a shower of nuclei is produced and consumption of the supersaturation begins. This creates a peak of fines which are partly remaining in the SPR, consuming supersaturation, and partly injected into the SCR; their fate there is as follows: they (a) grow to produce a peak of large crystals and/or (b) are removed by the fines dissolving system and/or (c) are retransferred, preferentially relative to large crystals (classification), to the SPR where by their growth will adversely affect the build up of the supersaturation in that region. Eventually, after they grow above a certain critical size they settle into the SCR. Their removal will favorably affect the build up of the supersaturation in the SPR. The

time at which the supersaturation threshold value will be exceeded again is determined not only by the rate of supersaturation production but also by the feedback effect resulting from the interaction of the above processes. There are a number of previously reported experimental facts and observations that strongly support the above model: Existence of Two Regions in Industrial Crystallizers. Large industrial crystallizers are usually operated under vacuum. The foam layers formed on the evaporating surface in Oslo (fluidized suspension), drafttube baffle (DTB), standard messo (SM), forced-circulation (FC), and other industrial crystallizers (a detailed description of these crystallizers is found in Bennett (1984)) constitute a supersaturation production region. The growth chamber in the Oslo and the bulk of the suspension in the other crystallizers constitute a supersaturation consumption region. Obviously, the liquid volume of a foam layer is much less than, e.g., onetenth of, the bulk suspension volume. Presence of High Supersaturation in the Foam Layer. High supersaturation buildup in foam is confirmed by the formation of crystal scaling on the vessel wall and on the agitator shaft (in pilot plants) near the liquid level. In a previous investigation (Qian and Chen, 1986) it was observed that the growth rate of KCl crystals near the above areas in 520 and 250 dm3 pilot DTB crystallizers was found in several-days runs to be 3 × 10-7 m/s, while in the bulk it was only (7-17) × 10-9 m/s. Sudden Increase of Nucleation Rate at High Supersaturations. It is well-known that in a stirred batch cooling crystallization no nucleation is observed until the supercooling reaches the so-called metastable limit and a shower of nuclei appears suddenly. In other words, the nucleation is a catastrophic phenomenon. In the presence of crystals, the nucleation is slow at low supercooling (contact secondary nucleation) and a shower of nuclei appears when a metastable limit is reached (“catastrophic” nucleation). Only the metastable limit now is lower than that in the absence of seeds. This means that if we want to express the dependence of nucleation rate, B°, on supersaturation, S, by a power-form empirical equation:

B° ) KNSm

(1)

the exponent m has a small value at low supersaturations but jumps at a certain supersaturation to a much higher value. This type of behavior has been observed also in continuous crystallizers. For instance, in a continuous crystallization of KCl (Qian et al., 1989) m had a value of 2.5 up to a temperature difference of 4 K between suspension and cooling water, at which point it jumped to a high value (about 11). This high nucleation rate was observed in the foam layer, and it was much higher than the contact nucleation in the bulk where the nucleation was much lower. In summary, the previous experiments confirm that while the bulk in a crystallizer may have a low supersaturation, the foam layer is a region of high local supersaturation, capable of producing periodically showers of nuclei. The proposed mechanism attributes, thus, the cycling to supersaturation fluctuations about a high threshold (i.e., transition between secondary nucleation and catastrophic nucleation) in contrast to the hypothesis advanced by Jager et al. (1991) mentioned above,

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in which the supersaturation fluctuates about a low threshold (Randolph and Larson, 1988), leading to a transition between no nucleation and secondary nucleation. On the basis of the elements on the new mechanism for CSD cycling, the experimental approach described below was devised. 3. Experimental Approach Scaling down on industrial crystallizer in the laboratory has been proven to be a very difficult task. An important factor is that industry employs mainly vacuum evaporative crystallizers, while investigators in the laboratory in view of operational problems prefer to work with cooling crystallizers. The hydrodynamic conditions and the supersaturation gradients in a large vacuum crystallizer differ markedly from those in a laboratory cooling crystallizer. Attempts of vacuum crystallization have been made in laboratories in the United States, China, and The Netherlands. The reproducibility, however, was found to be poor. A cooling crystallization technique was used by Randolph et al. (1977) in his above-mentioned work on CSD cycling. Serious experimental difficulties arose because of crystal scaling on the cooling surface and clogging by crystals in the FDS loop, and the experiments were terminated after 2-4 cycles. In this work unique techniques were used to simulate industrial crystallizers in the laboratory. These techniques were developed using the cycling model described in section 2 as a guideline. Their main features are described next. (a) Water Evaporation and Cooling by Air Bubbling. Air bubbling through a hot supersaturated solution at atmospheric pressure was used to simulate a vacuum evaporative crystallizer. Bubbling produces supersaturation by evaporation of water and by cooling of the solution. Bubbling creates also a foam layer similar to the foam layer in industrial crystallizers, in which layer the supersaturation is generated. The foam layer has a high heat-transfer coefficient and avoids the scaling on the cooling surface. In cooling crystallizers a high-temperature difference between cooling water and solution has to be used, leading to severe scaling on the cooling surface. (b) Region of High Supersaturation. The positioning of a funnel baffle between the draft tube and the air bubbling cups (see Figure 1) divides essentially the crystallizer into two parts. The lower part simulates the growth chamber in an Oslo crystallizer or the bulk region in other industrial crystallizers. The upper part is the region in which the supersaturation is generated. It simulates the evaporative head in an Oslo crystallizer or the evaporative foam layer in other crystallizers. The funnel baffle also simulates the partial mixing between the foam layer and the bulk suspension which is encountered in industrial crystallizers. The bottom diameter of the funnel and thus the flow rate from bulk suspension to the head can be varied; in this way in the foam layer, a retention time which is closer to that in industry, e.g., 1 min, can be obtained. (c) Fines Dissolving. Under our model it is necessary, for cycling to occur in the laboratory, to build up the supersaturation in the foam to a value exceeding the metastable limit. Evaporation is not sufficient to accomplish this in a laboratory crystallizer; the evaporation rate there is many times lower than the rate in industry. To obtain supercoolings in the foam layer similar to those in industry, part of the stream from the

Figure 1. Basic crystallizer.

fines dissolver was diverted to the foam layer in our laboratory crystallizer. For reasons associated with the lower suspension density (150-300 kg/m3 compared to 350-550 kg/m3 in industry) and the shorter retention time in the laboratory crystallizer (1-1.5 h compared to 6-8 h for ammonium sulfate crystallization in industry), a longer residence time in the fines dissolver and a higher ratio of cut size for FDS Lf to product size LD than in industry had to be used. (d) Metastable Limit. The concept of metastable limit, that is, the existence of a supersaturation value below which the nucleation rate is zero and above which it rises to a very high value, was first proposed by Miers and Isaac (1906). With the subsequent development of the classical nucleation theory, this concept lost its original physical meaning. Its use, however, remains as a practical way of expressing the supersaturation point at which a shower of nuclei will appear in the solution. Since the metastable limit is not only a property of the system, it will also depend on the conditions of the experiment: cooling rate, seeding, agitation, etc. In this work the metastable limit for the system ammonium sulfate-water was measured at a saturation temperature of 315.2 K, which is the crystallization temperature for the cycling experiments reported in the next sections. Analytical reagent ammonium sulfate and HPLC-grade water were used. The limit was measured in batch cooling crystallization at constant cooling rate, with and without seeding. The procedure followed was similar to that used by Nyvlt (1968) and Mullin et al. (1970). The limit was also measured in a batch crystallization in a jacketed 250 mL beaker, in which air was bubbled through a sintered-glass dispersor and the solution was cooled at a constant cooling rate. In the latter experiments the supercooling at nucleation was determined by solution density measurements with a precise densimeter (Paar, Model DMA602) after the agitation and bubbling were stopped following the nucleation.

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Table 1. Range of Operating Parameters parameter

range

crystallizn temp in bulk (K) air flow rate (std. dm3/s) agitation speed (rpm) saturation temp for feed (K) dissolver bath temp for FDS (K) air inlet temp (K) suspension density (kg/m3) seeds (kg/m3) inlet temp of feed and FDS

308-323 2.1-2.5 700-900 336-345 350-360 312-319 150-350 110-220 slightly higher than their saturation temp

The metastable limit in cooling crystallization, at a cooling rate of 6-7 K/h, was 4.60-5.55 K without seeding and 2.15-2.80 K with seeding. Previously reported values (Mullin et al., 1970) were 3.8 and 2.0 K, respectively, for the same cooling rate. The higher values obtained in this work could be attributed to the fewer particulates the HPLC grade of water may contain. In the air bubbling experiments, with bubbling rates of 2.1 × 10-5 and 11 × 10-5 std. m3/s and cooling rates of the supersaturated solution of 6 and 60 K/h, respectively, the corresponding metastable limits were 1.7 and 3.1 K. There are no previous data with these types of experiments to compare with. Mullin’s work (1970) with cooling crystallization showed at the highest cooling rate of 30 K/h a metastable limit of 4.8 K for unseeded and 3.2 K for seeded solutions. It can be concluded, therefore, that the metastable limit for the ammonium sulfate-water system in the CSD cycling experiments, which will be presented below and involve both air bubbling and the presence of crystals, was somewhere between 2 and 3 K. An effort was made to measure the supersaturation in the foam layer and in the bulk suspension directly using the above-mentioned Anton Paar densimeter. Unfortunately, the experimental setup did not permit the introduction of a large settling tube to separate the suspended fines from the liquid. A very small amount of fines interferes seriously with the density measurement of the solution. 4. Experimental Apparatus and Conditions 4.1. General Approach. The purpose of the experiments was to demonstrate that a laboratory crystallizer can be constructed whose operation can be shifted from a steady state to a sustained CSD cycling, simulating the behavior of industrial crystallizers, and that this shifting can be accomplished by condition changes consistent with the above criteria. In this way a successful outcome will provide an indirect support for the proposed model. The values of a number of variables, whose effect on the dynamics of the system is not important according to our model, were fixed within certain narrow limits, as shown in Table 1. 4.2. Basic Crystallizer. This is a funnel-baffled combined crystallizer shown schematically in Figure 1. It consists of an 148 mm i.d. Pyrex glass vessel with flange, a plexiglass cover plate, a hollow stainless steel draft tube of 70 mm i.d. and 89 mm o.d. with four baffle plates and cooling water inlet and outlet tubes, a speed adjustable agitator with a 50 mm diameter stainless steel three-blade propeller, three 30 mm diameter Teflon bubble caps with 12 2.4 mm × 25 mm slots, 8-10 mm o.d. tubes for feed and product withdrawal, and a funnel

Figure 2. Flow sheet of a basic crystallizer system.

baffle made of polyethylene, whose cone apex angle was 70° and bottom diameter 43 or 60 mm. The bottom of the crystallizer was contoured by a Teflon truncated cone section to help the suspension of the crystals. The crystallization temperature was controlled in the bulk in the range of 308-323 K with a precision of (0.1 K by the constant-temperature bath in which the crystallizer was immersed. In some runs cooling water from another constant-temperature bath was circulated through the hollow draft tube. The immersion of the bubble caps during operation was 75 mm from the liquid level. The active volume of bulk suspension in the crystallizer was 4.5 dm3, while that of the foam layer above the funnel was 0.35 dm3. The product was withdrawn by vacuum into a bottle every 1, 2.5, or 5 min, and after venting it was sent into a 5 dm3 dissolver immersed in a 341 or 351 K bath. The dissolved solution overflowed to a 5 dm3 saturator by a periodical siphon. After saturation at a temperature of 336 or 345 K, the clear solution was sent back into the crystallizer as the feed by a peristaltic pump (see Figure 2). Suspension samples were taken from the product withdrawal bottle at time intervals equal to the retention time. The samples were then filtered and washed by 77.5% and 100% ethyl alcohol successively. The filtered cake was dried and sieved by a shaker on 15 standard sieves with size ranges of 125-1400 µm. Few crystals were larger than 1000 µm, and many of them were agglomerated. It usually took several hours to saturate the solutions and attain the set temperatures in the crystallizer and the saturator. After the flow rates were calibrated and seed crystals were put into the crystallizer under controlled agitation, the feed, the FDS, and the air flows were started. In most of the runs, dry large seeds (>700 µm) were used. They provided a strong first shower of nuclei, and the transient period was found to be 4-6 retention times. The exception was run D in which the bulk suspension from run C was used, and as a result, because of the closeness of the CSD in both runs, the transient period was less than 1 retention time. The product withdrawal was adjusted to maintain the constant level in the crystallizer. The Teflon bubble caps were removed and washed at intervals of 1.5-2 h, to get rid of crystals plugging them. Samples were


to prevent any nucleation or clogging in the FDS system. Most of the pipelines carrying the solution were jacketed with hot water to avoid crystal clogging during the run. An air heater was used to keep the air inlet temperature near the crystallization temperature. 5. Experimental Runs and Their Results

Figure 3. Separated crystallizer.

taken at time intervals equal to the retention time. Usually it took more than 40 h for a run and for some runs as long as 3 days. For the last run (run E) a revised crystallizer was used that is described in section 5.5 and shown in Figure 3. 4.3. Fines Dissolving System. For simulation on the fines dissolving system (FDS) of an industrial crystallizer, a system was used to dissolve the fines in the solution withdrawn from a fines trap in the crystallizer and to send the resulting solution to the foam layer and/or the bulk of the crystallizer. The system consisted of two fines traps, a preheater, one or two fines dissolvers, a heater, a peristaltic pump with one or two pump heads, and corresponding coolers. The flow diagram of the crystallization system with fines dissolving is shown as Figure 2. The bath temperature of the fines dissolver was between 345 and 370 K, depending on the fines suspension density and the fines dissolving flow rate. Two pump heads with two pairs of coolers were used for high FDS flow rate. One dissolver of 0.8 or 2 dm3 or both was used depending on the FDS flow rate and its suspension density. The fines dissolver was operated under reduced pressure of the pump inlet; thus, a magnetic stirrer was used in the 2 dm3 dissolver instead of a regular mechanical agitator. Two sizes of the fines-trap set were used to match the size cut of fines at different FDS flow rates. Set I included a 23 mm i.d. tube and a 28 mm i.d. tube. Set II consisted of a pair of tubes with 48 mm i.d. for its upper part and 38 mm i.d. for its lower part. The upper part was above the funnel, while the lower part, which had a 35 mm diameter side hole facing the crystallizer wall, was below the funnel in order to get better separation of larger crystals from the withdrawn FDS flow. It was possible to switch from one trap to the other or to both during the operation without interrupting the operation. The inlet temperatures of FDS flows were usually 7-12 K higher than the bulk suspension temperature

Five representative runs will be presented here. The variables to be tested were the position of the funnel and its bottom diameter size, the position of the fines dissolving trap and its cross-sectional area, and the FDS flow rates. The variation of these factors gave a wide range of temperature differences between bulk and foam (from 0.1 to 5 K) and nucleation rates (from 6 × 105 to 2 × 107 no./m3‚s), which covered the whole range of dynamic behavior from transient to steady state and to cycling. All solutions sent to the crystallizer were saturated or nearly saturated. The very low supersaturation in the bulk was estimated from the growth rates {(0.7-4.0) × 10-8 m/s}. Experimental results are summarized in Table 2. This table contains also the values of the important operating variables. Details of each run follow. 5.1. Run A: Steady-State Operation at Very Low Supersaturation in Foam. This is a run designed in such a way so that the difference in temperature between the foam layer and bulk is very small and thus the supersaturation in foam is very low. This is achieved by utilizing the basic crystallizer with the funnel baffle removed. In addition, no fines removal was employed. The absence of the funnel baffle led, as expected, to a well-mixed crystallizer. The measured temperature difference ∆T between the bulk and foam was negligible. The supersaturation, estimated from low growth rates, was also very low, and the dominant product size LD, which is equal to 3Gτ, was about 500 µm. A shower of nuclei appeared within a few minutes after the air bubbling started. This phenomenon was observed in all runs of this type with air flow rate higher than 1.67 std. dm3/s. The nucleation rates for the last two samples were estimated at 6.4 × 105 and 8.8 × 105 no./m3‚s. These values are not significantly different from the nucleation rate of 5.0 × 105 no./m3‚s from a run with an air rate of 0.83 std. dm3/s, in which no shower of nuclei appeared. In run A the nucleation in the foam was obviously insignificant. According to the proposed mechanism, the above experimental conditions should lead eventually to a steady-state operation. The obtained results indicate that the system did reach steady state. The CSD dynamics is shown in Figure 4. The figure shows a transient state consisting of a strong first peak and a weak second peak. The steady state was attained in about 30 h. The correlation coefficient of linear regression of the last product samples on the ln n vs L plot was greater than 0.99. Incidentally, the growth rates were estimated during the transient period by following a fractional crystal weight peak and/or valley of a certain size as it advances in size with time. This average growth rate was found to be (0.8-1.0) × 10-8 m/s in the later period of the transient state. The growth rate at steady state was calculated from the conventional ln n vs L plot from the last two samples at 32 and 34 h; the growth rates were (1.05-1.10) × 10-8 m/s. These values are very close to the growth rate toward the end of the transient period.

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Table 2. Experimental Results of CSD Dynamics Simulation in the Laboratory crystallizer run (funnel size) (mm) A B C D E

MSMPR basic (φ 148/φ 60) basic (φ 148/φ 43) basic (φ 148/φ 43) separated

FDS overflow split feedback split feedback split feedback

temp retention time difference a ∆T (K) τ (103 s) τ/τf 0-0.1 0.4-0.7 1.8-2.8 2.7-3.7 4.0-4.9

15.3 11.8 4.54 4.64 4.68

CSD dynamics

transient & steady 2.95 damping oscillation 13.6 sustained cycling 22.7 sustained cycling 13.0 transient & steady

crystallizn kinetics cycling period τc/τ B° (106 no./m3‚s) G (10-8 m/s) 1.9 3.5-4.0 2.9-3.8

0.64-0.88 4.7-6.6 cycling cycling 8.3-21b

0.8-1.1 0.72-0.78 2.4-5.1 3.0-4.1 1.25-1.68b

a Temperature difference between bulk and head suspensions. b In head, B ° 108-109 no./m3‚s, G (3.0-3.5) × 10-7 m/s. In bulk, the h h median product size, L50, is small, 188-248 µm, compared to 320-580 µm in runs A-D.

Figure 4. CSD dynamics of run A.

This validates the peak/valley method for obtaining average growth rates at unsteady state. 5.2. Run B: Damping Cycling at Low Supersaturation in the Foam. The objective here was to have a run in which the supersaturation in the foam is higher than that in the bulk but still lower than the metastable limit of 2-3 K for catastrophic nucleation. This was achieved by using the basic crystallizer with a funnel baffle of 148 mm top diameter and 60 mm bottom diameter in order to reduce the amount of the circulating suspension transferred from the bulk to the foam layer. The build up of the supersaturation in the foam layer was helped also by sending the feed to that layer. However, no FDS stream from fines traps with 23 and 28 mm diameter was sent back to the foam layer. The fines overflow with a cut size of 90 µm. The operating conditions are indicated in Table 2. Under these conditions the temperature difference between the bulk and foam was found to be between 0.4 and 0.7 K as opposed to negligible in run A. The CSD dynamics presented in Figure 5 clearly showed a damping oscillation of CSD. After 21 h of operation the system approached steady state. The sample obtained after 29 h give a straight line in the ln n vs L plot. For example, the nucleation and growth rates obtained from the line at 29 h were 5.6 × 106 no./ m3‚s and 0.75 × 10-8 m/s, respectively. The nucleation rate is about an order of magnitude higher than the (6.4-8.8) × 105 no./m3‚s rate in run A, in which the nucleation in the foam was negligible. This means that the dominant site for secondary nucleation is now the foam. The bulk growth rate of (0.72-0.78) × 10-8 m/s was even lower than that in run A, due probably to the higher nucleation rate in the foam layer and thus larger crystal surface in the bulk. The temperature difference of 0.4-0.7 K indicated a supercooling in the foam much

Figure 5. CSD dynamics of run B.

lower than the metastable limit of 2-3 K. In conclusion, run B was designed to demonstrate a nonsustained (damping) cyclical behavior and it did. 5.3. Run C: Sustained Cycling at a Fines Retention Time, τf, of 5.6 min. This laboratory experiment was designed to simulate the cycling in DTB, SM, and similar industrial crystallizers. This was accomplished through the following characteristics. A fines stream was withdrawn from a settling zone in the bulk crystallizer, sent to the dissolver, and then fed back to the crystallizer. The return stream was divided into two parts. The major part was sent directly into the bulk suspension. This simulates the fact that, in a large crystallizer, the circulating suspension is mixed with the FDS returning stream and the feed, thus bypassing the foam layer. The minor part of the returning FDS flow was mixed with the feed and sent to the foam layer. This assists a buildup of the local supersaturation in the foam layer to a value necessary for the production of showers of nuclei from catastrophic nucleation and thus the production of a large number of fines. The FDS flow fed back to the bulk suspension supplied there the supersaturation necessary for the growth of crystals and the conversion of a size peak of fines to a peak of large crystals. This results in fluctuations in the fines suspension density, the total crystal surface area, and the supersaturation in the bulk. In turn, the feedback effect of these fluctuations in the bulk to the foam exaggerated the CSD cycling. The FDS cut size in this run was about 330 µm. The internal interaction between bulk and foam is controlled by the opening of the funnel baffle. The use in this run of a small opening of 43 mm diameter led to reduced internal flow from the bulk to the foam, greater temperature difference between them, and higher supercooling in the foam layer. The temperature difference between the foam and bulk varied between 1.8 and


Figure 7. Growth rate in run C.

Figure 8. CSD dynamics of run D.

Figure 6. CSD dynamics of run C.

2.8 K. The variation was due to fluctuations in the foam temperature at different times and places. The CSD dynamics presented in Figure 6 shows four cycles with no damping. Due to the crowding of the large number of points, separate figures for each individual crystal size were necessary. It should be noted that during the operational time of 13.5-19 h the whole FDS return flow was sent to the foam layer, i.e., no splitting. This short disturbance changed somewhat the shape of the curves mainly in the period of the first 2 h; however, the cyclic performance of the crystallizer was retained. The advance of the crystal weight fraction peaks and valleys as a function of time is shown in Figure 7 for various crystal sizes, directly from sieve analysis. As stated above, average growth rates can be calculated from these peak or valleys advances as in Figure 7. The estimated growth rates G of (2.4-5.1) × 10-8 m/s were higher than rates of (1.05-1.10) × 10-8 and (0.72-0.78) × 10-8 m/s in runs A and B, in which the supersaturation in the foam was lower. They were also higher than that of about 1.7 × 10-8 m/s in industrial crystallizers, evidently due to the short retention time τ of 1.26 h and low suspension density of 117-175 kg/m3. In the crystallization of ammonium sulfate and potassium

chloride in industry, those values are 6-8 h and 350550 kg/m3, respectively. The average period of cycling τc in run C was 4.7 ( 0.3 h, which corresponds to about 3.7 retention times (3.7τ). In industry the τc value of about 24 h corresponds also to 3τ to 4τ (Qian and Chen, 1986; Randolph et al., 1977; Bennett, 1983). In conclusion, this run exemplifies a laboratory crystallizer which exhibits cycling similar to that in large industrial crystallizers. 5.4. Run D: Sustained Cycling at Fines Retention Time, τf, of 3.4 min. The operation conditions in this run (see Table 2) were about the same as in run C with the exception of the FDS flow rates, whose effect was actually tested in this run. The total FDS flow rate was 23.7 cm3/s as opposed to 14.5 cm3/s for run C. Both fines traps in the crystallizer were used instead of the single one in run C. The FDS cut size was 300 µm. Also both fines dissolvers were used as shown in the flowsheet in Figure 2. A temperature difference between the foam and the bulk of 2.7-3.7 K was measured. This was higher by about 1 K from the difference in run C, and consequently in this run a higher supersaturation in the foam was also expected. The crystallizer was loaded with the bulk suspension from run C, resulting in a very short transient period. In view of this, one can conclude that the CSD dynamics in Figure 8 indicates a sustained CSD cycling, although only two nondamping cycles are

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Figure 9. Growth rate in run D. Figure 10. CSD dynamics of run E.

shown. The growth rates G obtained from the data in Figure 9 were in the range of (3.0-4.1) × 10-8 m/s. The average cycling period τc was 4.3 ( 0.6 h or about 3.4 retention times. The above values are the same as in run C. The high correlation coefficients of the growth rate regressions in Figures 7 and 9 for runs C and D, respectively, prove strongly the advance of CSD peaks and valleys. Thus, this method provides further support to the observation of a sustained CSD cycling in runs C and D. In conclusion, the increase in the FDS flow rate or the decrease in retention time from 5.6 to 3.4 min (keeping the split ratio of flow to foam and to bulk the same) did not change appreciably the CSD cycling characteristics. 5.5. Run E: Steady-State Operation at High Supersaturation in Foam. This run was designed with the purpose of obtaining a supersaturation in the foam layer higher than the metastable limit. The supersaturation in the foam could not be raised to high values operating the basic crystallizer or any variations of it. A heat balance in the foam layer showed that the suspension interflow rate from the bulk to the foam was much greater than either one of the flow rates of the other two streams bringing liquid to the foam, i.e., the returning FDS and the feed. The bulk suspension sent to the foam had a low supersaturation, and as a result it was preventing the buildup of high supersaturation in the foam by the other two streams and the evaporation. This means that another crystallizer with an additional bath to cool the foam layer was necessary. In this run the two parts, head and bulk, were physically separated. A twin crystallizer was constructed consisting of two interconnected mixed suspension crystallizers, with a suspension volume of 5.1 and 0.61 dm3, respectively, as shown in Figure 3. In a way similar to that of the two regions in the basic crystallizer presented above, the two crystallizers simulate the two regions, SPR and SCR, in industrial crystallizers. The small crystallizer (head) with air bubbling simulates an evaporative foam layer in an industrial crystallizer. The large one (bulk) with the fines trap used in run C simulates the growth chamber of Oslo crystallizers or the bulk suspension in DTB, SM, or FC crystallizers. Both head and bulk crystallizers have their own constanttemperature water bath, and their temperatures are controlled separately. The suspension exchange between the two crystallizers is controlled by two metering

pumps (peristaltic pumps). A funnel baffle was used in the head with a bottom opening of 25 mm. The top diameter was the same as the diameter of the vessel, 125 mm (see Figure 3). The funnel was on top of a 150 cm3 reservoir bottle. Suspension was withdrawn from the bottle by a metering pump and sent to the bulk crystallizer. Three bubble caps were used as in the basic crystallizer. Suspension from the bulk can be sent to the region above the funnel in the head using a metering pump. In run E, however, only a minor part of the returning FDS flow was sent to the head, which enters the head with the feed at a point above the funnel. The operation parameters are given in Table 2. The FDS cut size was about 300 µm. The air was bubbled through three caps in the head. At the high supersaturation created there, a high nucleation rate always occurred. The resulting suspension was collected in the reservoir bottle below the funnel, from where it was withdrawn and sent to the bulk. Samples were taken from this interflow, and CSDs were determined by sieve analysis. The suspension density in the head is low, 20-30 kg/m3. The high nucleation rate Bh° and growth rate Gh of 108-109 no./m3‚s and (3.0-3.5) × 10-7 m/s, respectively, obtained from the straight line in a ln n vs L plot confirmed the existence of high supersaturation in the head. The CSD dynamics is shown in Figure 10. Steady state was attained after about 12 h or 9.3 retention times. There was still some random CSD fluctuation, but no peak or valley advance was found. The correlation coefficients of regression for the lines on the ln n vs L plots were greater than 0.99 in the last samples. 6. Discussion The five experimental runs simulated the whole spectrum of performances of an industrial crystallizer. Runs A and B simulated the steady-state CSD performance of crystallizers in industry at the so-called underloading production. On the other hand, run E simulated the steady state but fine product size production. This represents the so-called overloading operation. For example, the product size of run E is 188248 µm, much less than 572-581 and 317-371 µm of runs A and B, respectively. The important point of the investigation, however, is that a bench-size laboratory crystallizer exhibiting cyclic performance with respect to product CSD can be con-


structed and operated (runs C and D). The way this was achieved in this study provides also clues about the possible mechanism leading to the cycling of the large crystallizers. These clues are consistent with the proposed above two-region model. A significant element of this model and that resulting from its CSD cycling mechanism is the existence of a local high supersaturation in the foam layer (SPR) near the metastable limit that results in periodic catastrophic nucleation in the same layer. The emphasis is first on catastrophic, meaning that the nucleation follows the characteristic kinetics of the primary nucleation model. However, the nucleation in the foam layer is promoted by the presence of crystals; the data indicate that it occurs at a supercooling lower than the supercooling at which primary nuclei would appear. Therefore, it is by definition a type of secondary nucleation, which nevertheless is of a catastrophic type. The existence and the mechanism of such a type of secondary nucleation are the subject of a current investigation in our laboratory. The emphasis is also on periodic. The presented model states that nucleation fluctuates around an upper threshold separating a supercooling region of low rate secondary nucleation and that of catastrophic nucleation. The results of the five runs can be explained in terms of the above mechanism. The absence of the catastrophic type of nucleation in the foam layer in the case of a low supersaturation that remains always below the threshold explains the achievement of steady-state operation with relatively large product size in runs A and B. In the case of intermediate supersaturation in the foam (runs C and D), the supersaturation fluctuates across the upper threshold, the nucleation oscillates between secondary and catastrophic type, and CSD cycling results. In run E the supersaturation in the foam region is high enough so that, although it fluctuates, it remains always above the threshold; the nucleation is continuously of the catastrophic type and the product size is small. The results suggest a possible reason for the cycling of large industrial crystallizers. Do they also suggest how the crystallizers should be constructed or their operation controlled in order to prevent cycling? The laboratory simulation, as well as previous pilot-plant experiments, point to four main factors that initiate and exaggerate CSD cycling, namely, (1) low circulation flow per unit foam layer (SPR) volume, (2) high evaporation rate per unit area of liquid-gas interface, (3) excessive fines dissolving, i.e., high τ/τf ratio, and (4) settling (classification) of large crystals (>300 µm) from the foam layer. To devise a way to control cycling in an industrial crystallizer, knowledge of the relative importance of each factor in such a crystallizer is required. This can be accomplished only by a quantitative treatment of the proposed model. Such a treatment, a dynamic analysis of the proposed model, is the subject of an ongoing investigation. 7. Conclusions A bench-size crystallizer capable of simulating the long-period CSD cycling of an industrial crystallizer was constructed and operated, for the first time, in the laboratory. A model for large industrial crystallizers was formulated, and based on this model, a mechanism

for long-period CSD cycling was advanced. This mechanism was used as a guideline in the construction of the laboratory crystallizer. The success of this crystallizer in achieving cyclic performance supports the proposed mechanism. Acknowledgment Finding for this work has been provided by The National Science Foundation (CPS 9107442) and Allied Signal Fibers. Special thanks are in order to Mr. Brian Emgushov and Mr. Christos K. Dalianis for their invaluable assistance in the laboratory work. Nomenclature B° ) nucleation rate in bulk or whole crystallizer, no./m3‚s Bh° ) nucleation rate in head (foam layer), no./m3‚s G ) growth rate in bulk or whole crystallizer, m/s or nm/s Gh ) growth rate in head (foam layer), m/s or nm/s KN ) coefficient in general nucleation kinetic equation (1), (no./m3‚s)(kg/m3)-m L ) crystal size, µm LD ) dominant product size, µm Lf ) cut size of fines dissolving, µm L50 ) median product size (50 wt % undersize), µm m ) power index of supersaturation in general nucleation kinetic equation (1) n ) population density, no./m4 S ) supersaturation, kg/m3 ∆T ) temperature difference between bulk and head suspensions, K Greek Letters τ ) retention time of product crystals, h or s τc ) cycling period of crystal size distribution, h or s τf ) retention time of fine crystals, h or s

Literature Cited Bennett, R. C. Swenson Process Equipment, Inc., personal communication, 1983. Bennett, R. C. Crystallization from Solution. In Perry’s Chemical Engineer’s Handbook, 6th ed.; Perry, R. H., Green, D. W., Eds.; McGraw-Hill: New York, 1984; pp 19-34 and 19-39. Bennett, R. C.; Fiedelman, H.; Randolph, A. D. CrystallizerInfluenced Nucleation. Chem. Eng. Prog. 1973, 69 (7), 86. Eek, R. A.; Dijkstra, S.; van Rosmalen, G. M. Dynamic Modeling of Suspension Crystallizers, Using Experimental Data. AIChE J. 1995, 41 (3), 571. Jager, J.; Kramer, H. J. M.; Scarlett, B.; de Jong, E. J.; de Wolf, S. Effect of Scale of Operation on CSD Dynamics in Evaporative Crystallizers. AIChE J. 1991, 37 (2), 182. Miers, H. A.; Isaac, F. Refractive Indices of Crystallizing Solutions. J. Chem. Soc. 1906, 89, 413. Mullin, J. W.; Chakraborty, M.; Mehta, K. Nucleation and Growth of Ammonium Sulfate Crystals from Aqueous Solution. J. Appl. Chem. 1970, 20, 367. Nyvlt, J. Kinetics of Nucleation in Solutions. J. Cryst. Growth 1968, 3/4, 377. Qian, R.-Y. Shanghai Research Institute of Chemical Industry, unpublished results, 1986. Qian, R.-Y.; Chen, Z.-D. An Approximate Mathematical Model for Design of Industrial Crystallizers with Elutriators. J. Chem. Ind. Eng. (China) 1986, 1 (2), 80.

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Received for review July 6, 1995 Revised manuscript received December 4, 1995 Accepted January 8, 1996X

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IE950412U

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X Abstract published in Advance ACS Abstracts, March 1, 1996.

Laboratory Simulation of Industrial Crystallizer Cycling - Industrial

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