An Air-Lift Crystallizer Can Suppress Secondary Nucleation at a

Jun 5, 2014 - nucleation can be suppressed in both the air-lift crystallizer and the stirred ... Air-lift crystallizers can potentially operate with a...
0 downloads 0 Views 2MB Size
Download PDF
Article pubs.acs.org/crystal

An Air-Lift Crystallizer Can Suppress Secondary Nucleation at a Higher Supersaturation Compared to a Stirred Crystallizer Richard Lakerveld,* Jeroen J. H. van Krochten, and Herman J. M. Kramer Department of Process & Energy, Delft University of Technology, Delft, The Netherlands S Supporting Information *

ABSTRACT: Secondary nucleation is suppressed in an air-lift crystallizer at levels of supersaturation where in a stirred crystallizer a clear contribution of secondary nucleation is visible. A comparison of batch crystallization of Lascorbic acid in an air-lift crystallizer and in a stirred crystallizer is presented. The results demonstrate that at low supersaturation, secondary nucleation can be suppressed in both the air-lift crystallizer and the stirred crystallizer. At higher supersaturation, nucleation starts to dominate in the air-lift crystallizer. At an intermediate level of supersaturation, a clear contribution of secondary nucleation in the final product obtained from the stirred crystallizer is visible. However, experiments with similar conditions in the air-lift crystallizer show a significantly smaller contribution of secondary nucleation. The observed enlargement of the operating window in terms of supersaturation where secondary nucleation is suppressed in an air-lift crystallizer may have important practical consequences. Air-lift crystallizers can potentially operate with a higher crystal growth rate and the operating window for design and automated control can be extended.

■

INTRODUCTION The design and operation of industrial crystallization processes are complicated by stringent demands for product quality to enhance the efficacy of the final product and to reduce the costs of downstream operations. Crucial product specifications cover a range of characteristics including the crystal size distribution (CSD) and crystal shape, which are the result of several kinetic processes occurring within a crystallizer. In seeded batch crystallization, nucleation is often undesired as nucleation creates fines that broaden the CSD and hamper an accurate prediction of the CSD. In addition, small fluctuations in process conditions can trigger nucleation events with a large impact on the final CSD, which can lead, for example, to large batch-tobatch variations in industry. Therefore, there has been a high interest to understand the mechanisms of nucleation and to develop new design and control strategies to minimize nucleation. The various strategies to improve the control over nucleation during industrial crystallization processes can be understood from the operating conditions that are targeted. Most of the work has been dedicated to optimizing the supersaturation, which is the driving force for crystallization. Secondary nucleation refers to the birth of crystals in the presence of other crystals.1 Supersaturation has an influence on secondary nucleation, for example, via the survival chance of nuclei that are produced through crystal collisions. In general, damaged crystals grow with a different rate compared to undamaged crystals.2 Secondary nuclei have a large specific surface area, and the crystal lattice may contain strain,3 which may result in their dissolution at low supersaturation and minimize the effective secondary nucleation rate.4 Alternative mechanisms for © 2014 American Chemical Society

secondary nucleation rely on the presence of a crystal surface to induce nucleation via so-called activated nucleation, which depends strongly on supersaturation and requires a free-energy barrier to be passed.5 In either case, the strategy for design and operation boils down to preventing gradients in the supersaturation within the crystallizer and matching the consumption of supersaturation by the crystalline surface area with the production of supersaturation to prevent high levels of supersaturation. The former requirement explains why many industrial crystallizers still resemble well-mixed vessels, while the latter requirement has been the subject of intense scientific investigations for several decades. Consequently, significant progress has been achieved on effective strategies to control supersaturation via design and automated control.6−22 A complementary strategy to reduce secondary nucleation is to eliminate high-impact crystal collisions. Such strategy may require evolution of the design of the crystallizer hardware and has received less attention in the literature. The main source of high-impact crystal collisions is the presence of an impeller or a slurry circulating pump. Detailed kinetic models exist that describe the influence of an impeller on the secondary nucleation rate.4,23−26 An Oslo crystallizer, also called growth crystallizer, is an example of an industrial crystallizer in which collisions of crystals with moving internal parts are absent. These crystallizers are operated routinely in industry. The design and operation of such crystallizer are complicated by the need to maintain a fluidized bed of crystals and by the Received: January 16, 2014 Revised: May 20, 2014 Published: June 5, 2014 3264

dx.doi.org/10.1021/cg500090g | Cryst. Growth Des. 2014, 14, 3264−3275

Crystal Growth & Design

Article

circulation of supersaturated liquid, which may cause internal fouling. Especially when process development time is short and the same equipment has to be utilized for various purposes due to short production campaigns, such as in the pharmaceutical industry, stirred vessels offer advantages in terms of simplicity and flexibility. However, the impact of crystal-impeller collisions on the final CSD and crystal shape may be significant and unavoidable. Air-mixed devices are an interesting alternative to stirred crystallizers. An air-lift crystallizer combines the flexibility and simplicity of a stirred vessel with the absence of moving internal parts such as in an Oslo crystallizer. As such, they would also fit well in multipurpose production facilities and may reduce undesired secondary nucleation. An additional advantage is the extensive experience that exists in the design and operation of air-mixed devices obtained from applications in biotechnology. Crystallization in air-mixed systems have been studied for several applications, for example, involving reactions and salting-out of minerals.27−30 However, few studies exist in which the crystallization performance of air-mixed devices is compared systematically to stirred crystallizers when the objective is to minimize secondary nucleation. Bao et al.31 studied the crystallization of calcium gluconate in an external loop air-lift column and in a small-scale stirred tank. No significant differences in crystallization kinetics were found, and the growth of seed crystals proceeded without secondary nucleation in both vessels. In prior work, we investigated seeded crystallization of ammonium sulfate from water in a bubble column in batch and continuous flow mode. A comparison with results from a conventional stirred crystallizer did show a reduction in nucleation in the bubble column but also revealed the need for improved mixing.32,33 Therefore, we extended the bubble column to an air-lift system, which resulted in improved mixing and a further reduction of nucleation compared to a stirred vessel.34 In addition, the importance of obtaining favorable hydrodynamic conditions via proper design of a sparger and a gas disengagement zone was demonstrated. Despite these case studies that exist in the literature, the question remains under which operating conditions switching from a conventional stirred crystallizer to an air-lift crystallizer offers the most advantages. The objective of the present work is to investigate the influence of the operating conditions on the differences in performance of a conventional stirred draft-tube crystallizer and an air-lift crystallizer. An L-ascorbic acid−water system is used as a model system. The focus is on the attainable CSD and crystal shape. Compared to prior work, several extensions to the air-lift crystallizer have been implemented. First, an image probe is inserted in the crystallizer to capture the development of crystal size and shape in situ. Second, supersaturation data are available, which improves insights into the crystallization behavior. Third, a novel sparger to introduce nitrogen is used to improve the hydrodynamics. Fourth, pressure differences are measured between top and bottom of the crystallizer, which allow for quantification of liquid circulation velocities. LAscorbic acid is produced industrially via crystallization-based processes and is known to be sensitive for secondary nucleation.35

■

Figure 1. 5-L stirred draft-tube crystallizer. The draft tube is indicated by the orange surface. just above the bottom of the draft tube. The stirrer speed was set to 550 rpm by an overhead engine (IKA, RW20) to generate an upward flow in the draft tube. The temperature was measured inside the vessel with a resistance thermometer (Pt-100). Water was circulated through a jacket by a thermostatic bath (Lauda, Ecoline Star Edition RE304), which was used to control the temperature inside the vessel according to a time-varying set point. An image probe (MTS, PIA 524) was inserted diagonally from the top to capture in situ images of the slurry, which could be used to evaluate the qualitative characteristics of the CSD and the shape of crystals. The vessel was closed with a lid equipped with holes for the stirrer, thermometer, and image probe. The slurry was removed from the vessel via an outlet in the bottom. A typical experiment started with the preparation of a saturated solution in a separate jacketed vessel. An excess of L-ascorbic acid (DSM, Universal grade) was added to demineralized water. The demineralized water was stripped with nitrogen prior to usage to remove oxygen, which aimed to prevent oxidation of L-ascorbic acid. A resistance thermometer was used to measure the temperature inside the vessel. A jacket and a thermostatic bath (Huber, CC-231) were used to control the temperature inside the closed vessel at a fixed set point (Tsat). The vessel was well stirred with a marine type impeller with overhead engine (IKA, RW28) and kept overnight to ensure that the liquid became saturated with L-ascorbic acid at Tsat. At the start of an experiment, the stirrer was switched off, and crystals were allowed to settle such that clear liquid could be transferred to the 5-L drafttube crystallizer. Initially, the temperature inside the draft-tube crystallizer was controlled typically 2 °C above Tsat to ensure that any crystals that may have formed during transition or tiny crystals that did not settle sufficiently before transition would dissolve. The same thermometer that was used to prepare the saturated solution was also used for temperature control of the crystallizers to avoid any measurement bias. The actual concentration of L-ascorbic acid in water was measured at the start of each experiment via titration (TitraLAB, 865). Subsequently, the temperature was gradually brought to Tsat. The cooling profile during an experiment consisted of two segments separated by a seeding point. First, a linear cooling profile of 0.2 °C/min was executed until the seeding point (Ts) was reached. At the seeding point, seed crystals were added to provide surface area for crystal growth. Finally, a third-order cooling profile was executed described by the following equation:36

EXPERIMENTAL SECTION

Stirred Crystallizer. A scaled drawing of the stirred 5-L draft-tube crystallizer is given in Figure 1. A 4-blade marine-type propeller was placed inside a draft tube (indicated by the orange surface in Figure 1), 3265

dx.doi.org/10.1021/cg500090g | Cryst. Growth Des. 2014, 14, 3264−3275

Crystal Growth & Design

Article

⎛ t ⎞3 T (t ) = Ts − (Ts − Tf )⎜ ⎟ ⎝ tf ⎠

Several degrees of freedom were varied during the experiments. Essentially, all these degrees of freedom will change the supersaturation profile. The unseeded experiments were used to estimate the metastable zone width. Air-Lift Crystallizer. A scaled drawing of the 18-L air-lift crystallizer is given in Figure 2. A detailed description of the design is presented elsewhere.34 At the bottom of the column, 100% nitrogen was introduced into the riser by a ring-type sparger (Figure 3). The design of the sparger is a critical part of an air-lift crystallizer.34 In particular, the velocity in the holes of the sparger should be sufficiently large to prevent weeping of the liquid. An empirical correlation from

(1)

where Tf is the temperature at the end of the batch (tf). The in-line image probe was used to confirm for each experiment that the solution was clear at the start of the experiment and that the added seeds did not dissolve after seeding. A small fraction of mother liquor (10 mL) was removed from the vessel at the start of each experiment and, subsequently, every hour during the experiment to measure the concentration of the mother liquor via titration (TitraLAB, 865). For titration, a sample holder was placed on a scale (Mettler Toledo, AE200). The mother liquor was injected in the sample holder with a syringe or pipet (around 20 g for low concentrations and around 0.5 g for high concentrations). The measured mass of the sample was an input for the titration procedure with sodium hydroxide. At the start of each measurement, the sample was diluted with demineralized water (until around 10 mm of solution in sample holder), and the electrodes on the titration device were rinsed with demineralized water. During titration, a magnetic stirrer was used for mixing in the sample holder. The titration procedure itself was fully automated. At the end of the experiment in the crystallizer, the crystals were collected through the sampling point. The slurry was directly filtered with a Büchner funnel and filtration paper (Whatman, 597 grade) placed on a flask that was kept under a vacuum with a vacuum pump (KNF Neuberger Laboport) for rapid filtration. The slurry on the filter bed was shortly washed with demineralized water of 15 °C, which aimed to prevent caking during drying. The crystals were further dried in an oven (50 °C, 15 min). Note that some dissolution may still occur during drying, but no significant caking was observed on light microscope or SEM images. The CSDs of the crystals were measured in triplicate using a laser diffraction instrument (Microtrac, S3500). In general, the quantity of material needed for the CSD measurements was relatively small compared to the total yield of a batch, which complicated the measurement of the CSD due to possible particle segregation before sampling. Therefore, the standard deviations of the measurements in triplicate have been reported to quantify the reliability of the measured mean value for the volume density of each size class. Finally, light microscope and scanning electron microscope (SEM) images were made to verify the CSD measurements and to analyze the crystal morphology. The temperature profile given by eq 1 aimed to match the consumption of supersaturation with the production of supersaturation by increasing the cooling rate toward the end of the batch when more crystal surface was available. Detailed modeling or automated feedback control via in-line supersaturation measurements could in principle result in an optimized supersaturation profile. However, the objective of this work is to compare the growth of the seed crystals in the stirred vessel with the growth of seeds crystals in the air-lift crystallizer rather than optimizing the supersaturation profile. Therefore, a simple expression for the cooling rate was utilized. Note that the supersaturation profile was a strong function of selected temperatures (Tsat, Ts, Tf), batch time, and seed properties (mass, CSD). An overview of the conducted experiments is given in Table 1.

Table 1. Description of Experiments in the 5-L Draft-Tube Crystallizera no.

time [h]

Ms [g]

Cs [wt %]

Tsat [°C]

1 2 3 4 5 6

2 2 6 6 8 8

0.0 0.0 14.4 9.0 9.0 9.0

0.0 0.0 1.9 1.9 2.1 2.0

35.0 45.0 40.0 40.0 40.0 40.0

TS [°C]

Tf [°C]

31.6 37.0 37.0 37.0

10.0 10.0 25.0 31.0 31.0 31.0

a The seed load Cs is defined as the ratio of the seed mass (Ms) over the theoretical mass of the final product based on solubility and the measured initial concentration.

Figure 2. 18-L air-lift crystallizer. TT = temperature transmitter, dPT = differential pressure transmitter. 3266

dx.doi.org/10.1021/cg500090g | Cryst. Growth Des. 2014, 14, 3264−3275

Crystal Growth & Design

Article

Table 2. Description of Experiments in the 18-L Air-Lift Crystallizera #

time [h]

Ms [g]

Cs [wt %]

Ts [°C]

Tf [°C]

7 8 9 10 11 12 13 14

6 6 6 6 6 6 6 8

52.0 11.6 11.6 11.6 3.5 0.0 32.5 32.5

2.0 0.6 0.6 0.6 0.2 0.0 1.6 1.8

31.6 31.6 31.6 31.6 31.6 31.6 37.0 37.0

25.0 31.0 31.0 31.0 31.0 31.0 31.0 31.0

Tsat = 40 °C for all experiments. The seed load (Cs) is defined as the ratio of the seed mass (Ms) over the theoretical mass of the final product based on solubility and the measured initial concentration. a

thermostatic bath (Lauda, MS/M3) was used to control the temperature of the solution at Tsat. The concentration of the mixture corresponded to a saturated solution at 0.2 °C below Tsat, which aimed to dissolve crystalline dust that may be adhering to the crystals. The rotational speed of the magnetic stirrer was controlled at 500 rpm. A lower rotational speed resulted in clearly observable segregation and a higher rotational speed would produce excessive attrition. Figure 4

Figure 3. Sparger installed at the bottom of the 18-L air-lift crystallizer to distribute nitrogen. literature was used to estimate the critical weep velocity, and the gas flow rate was adjusted accordingly to operate well above this critical weep velocity.37 The nominal operating velocity of the nitrogen was 1.6 × 10−4 m3/s, which corresponded to a superficial gas velocity of 3.7 × 10−2 m/s in the riser. The gas left the system at the top of the column, and the liquid with crystals flowed into the downcomer. A gas disengagement zone has been installed to reduce the liquid circulation velocity at the top, which aimed to prevent gas bubbles from being dragged down into the downcomer. A heating medium flowed through the jacket of the downcomer via a thermostatic bath (Lauda, Ecoline Star Edition RE310), which was used together with an in situ resistance thermometer (Pt-100) to control the temperature inside the crystallizer according to a time-varying set point. The gas disengagement zone was not jacketed but was well insulated. Temperature was measured at several locations within the crystallizer. The thermometer for temperature feedback control was inserted in the top of the riser (indicated by TT5 in Figure 2). Furthermore, resistance thermometers (Pt-100) were installed at the bottom of the riser and downcomer and once more at 0.8 m from the bottom (see Figure 2). These temperature measurements at several locations could be used to detect any nonuniformity. The driving force for circulation for air-lift systems is determined by the density difference between the material in the riser and the downcomer. Two differential pressure transmitters (Siemens, Sitran P Series, D-76181) were installed to measure the density in-line in the riser and downcomer by comparison with a reference fluid. Furthermore, an image probe (MTS, PIA 524) was used for in situ observation of the crystals. The image probe was inserted horizontally with the gap positioned in the direction of the flow in the downcomer to obtain a representative fraction of the crystal slurry at minimal flow disturbance. The general procedure of a typical experiment in the air-lift crystallizer was identical to the procedure of the experiments in the 5-L draft-tube crystallizer as described previously. An overview of the conducted experiments in the air-lift crystallizer is given in Table 2. Note that for a number of experiments the temperature range in which the thirdorder cooling profile is executed is small, which essentially means that for those experiments the seed crystals grow at constant temperature. Such seeding at a low temperature aims to investigate the importance of nucleation at high supersaturation in the air-lift crystallizer. Seeding Procedure. Commercial L-ascorbic acid crystals (DSM, Universal grade) were introduced in a shaker machine (Retch, AS200 Basic) with sieves. The crystals were distributed on the top sieve until a small layer of crystals covered the complete mesh. Subsequently, the shaker machine was operated for 15 min after which the crystals that remained on the top sieve were replaced by new crystals. Rubber balls were added to each sieve to expedite the process. The 112−212 μm size fraction was collected and introduced into a flask with 100 mL of a mixture of L-ascorbic acid and water. The flask was placed on a magnetic stirring plate (IKA, RCT basic) with magnetic stirrer and a

Figure 4. CSD of seed crystals in suspension as a function of residence time in the seeding vessel. illustrates the CSD of the seed crystals as a function of the residence time in the seeding vessel, which shows the occurrence of a left shoulder after several minutes. Therefore, a residence time of only 5 min in the seeding vessel was selected to dissolve crystalline dust before introduction of the seed slurry in the crystallizer, which still produced a reproducible CSD of the seeds with limited influence of attrition (Figure 4). The growth of the seeds could be observed from the in-line image probe. For comparison between both crystallizers, the so-called seed load for each experiment was calculated, which was demonstrated to be a reliable scale-up parameter.38 Solubility and Metastable Zone. The solubility of L-ascorbic acid was measured using the same analytical techniques as were used to measure the concentration during an experiment. The solubility of L-ascorbic acid in water as a function of temperature was measured in a 100-mL jacketed vessel. Demineralized water was stripped with nitrogen to reduce oxidation of dissolved L-ascorbic acid and, subsequently, transferred to the vessel. An excess amount of L-ascorbic acid (DSM, Universal grade) was added, and temperature was controlled with a thermostatic bath (Lauda, MS/M3) at a certain temperature between 20 and 60 °C. Subsequently, the mixture was stirred with a magnetic stirrer (IKA, RCT basic) for 3 h. Finally, the composition of the liquid was measured with titration (TitraLAB, 865). The solubility at the standard saturated temperature was 3267

dx.doi.org/10.1021/cg500090g | Cryst. Growth Des. 2014, 14, 3264−3275

Crystal Growth & Design

Article

measured in triplicate to assess the accuracy of the titration device, which resulted in an average saturated concentration of 0.566 g/g at 40 °C with a standard deviation of 0.00920 g/g (1.6%). The obtained solubility curve is depicted in Figure 5 and compares reasonably well with values reported in the literature.41 Figure 5

Figure 6. Gas fraction in the riser and downcomer during Experiment #14. The difference in gas fraction has been used to estimate the average superficial liquid circulation velocity in the direction of flow below the gas disengagement zone.

hydrodynamic model42 has been used to correlate the difference in gas fraction to velocity, which was described for the present setup in prior work.34 The obtained liquid circulation velocity (see Figure 6) should be sufficiently high to keep all crystals in suspension during the experiments. Crystallization Kinetics and Final Product Quality. The investigated cooling profiles and seed properties will influence the supersaturation profiles in both the stirred crystallizer and air-lift crystallizer. The central question in this section is whether the difference in mixing mechanism, for a given supersaturation profile, has a significant influence on the growth of a suspended crystal population. Therefore, a comparison of the supersaturation during a batch and the CSD of the final product has been made in the case of the stirred and air-mixed crystallizer. Videos illustrating a series of in-line images for each experiment are available as Supporting Information. In general, three qualitatively different supersaturation profiles can be distinguished in the results: (1) a profile with seeding at a low temperature (high supersaturation) with sufficient seed surface area to reduce significantly the supersaturation in the first 2 h of a batch, (2) a profile with seeding at a high temperature (low supersaturation), and (3) a profile with seeding at a low temperature, but with insufficient seed surface area to reduce the supersaturation in the beginning of the batch. Figure 7 illustrates the development of the supersaturation in the stirred crystallizer (Experiment #3) and air-mixed crystallizer (Experiment #7) for a case with seeding at a low temperature and a sufficient seed load such that supersaturation drops significantly in the beginning of the batch. The development of the supersaturation in the first hour of the batch suggests that initially the seed population grows similarly in both vessels. Note that the identical uptake of supersaturation for an identical seed population and cooling profile in both crystallizers suggests that any difference in hydrodynamics has a limited influence on the growth kinetics of the crystals when assuming that no nucleation or agglomeration occurs in the first hour. The supersaturation in the stirred crystallizer drops more rapidly compared to the supersaturation in the air-lift crystallizer, which can either mean an increased crystal growth rate or an increase in crystal surface area driven by nucleation. The supersaturation profile in the air-lift crystallizer drops several hours later as well, which may suggest

Figure 5. Solubility of L-ascorbic acid in water as a function of temperature expressed in grams of L-ascorbic acid per gram of water. The black solid line is a fitted curve described by an exponential function. The center of the vertical black line at the saturated concentration at 40 °C corresponds to the mean value from a measurement in triplicate, and the upper and lower limit of the line are determined by a single standard deviation in both directions. The red diamonds correspond to the points at which primary nucleation was observed during Experiments #1 and #2. The orange triangle corresponds to the point at which primary nucleation was observed for Experiment #12. contains also the temperatures at which nucleation was observed with the in-line image probe for the unseeded Experiments (#1, #2, and #12), which provides an estimate of the metastable zone width. The unseeded experiments in the 5-L draft-tube crystallizer were conducted with a linear cooling profile until nucleation was observed. In the case of the air-lift crystallizer, the solution was cooled until the lowest seeding point used (T = 31.6 °C) was reached after which the thirdorder cooling profile as described by eq 1 was executed. The first nuclei were visible in the air-lift crystallizer approximately 1 h after the start of the third-order cooling profile. In general, the results indicate that operation is expected to stay within the metastable zone for primary nucleation at least until the seeding point. Note that the very limited number of unseeded experiments call for a cautious approach toward the width of a metastable zone as additional experiments would be required to get a complete understanding of the metastable zone.

■

RESULTS AND DISCUSSION Liquid Circulation Velocity in Air-Lift Crystallizer. In general, several hydrodynamic regimes can be distinguished in air-lift systems. A key distinction is whether bubbles are dragged into the downcomer and, possibly, re-enter the riser at the bottom.42 The highest driving force for mixing for a given inlet flow rate of the gas is obtained when no gas is present in the downcomer. The gas fraction in the riser and downcomer as calculated from the pressure difference measurements during Experiment #14 is illustrated in Figure 6. The results demonstrate that the gas fraction in the downcomer was very small, which was confirmed visually by in situ images of the slurry. The difference in gas fraction between the downcomer and riser can be used to estimate the average liquid circulation velocity in the part below the gas disengagement zone. A simple 3268

dx.doi.org/10.1021/cg500090g | Cryst. Growth Des. 2014, 14, 3264−3275

Crystal Growth & Design

Article

nucleation was minimal before and after the nucleation event. In the case of the stirred crystallizer, a larger peak at the small size classes can be seen, which demonstrates more nucleation compared to the air-mixed case toward the end of the batch. Furthermore, crystals are present with a broader range of sizes, which indicates a more sustained production of nuclei via secondary nucleation during the batch. Finally, the peak associated with the grown seeds has a peak value at a lower size class, which is the result of the larger competition for growth between the seed crystals and the produced nuclei in the case of the stirred crystallizer. Images obtained from a light microscope and SEM (Figure 9) show the significant presence of fines in the product obtained from the stirred crystallizer. In addition, the improved crystal facets suggest a reduction of attrition in the air-lift crystallizer. Note that the precise location of the main peak of the CSD of the product obtained from the stirred crystallizer could not be measured with high certainty as illustrated by the large standard deviation in this area of the graph. However, the significant difference in the volume of fines between the two distributions was measured with high certainty, which is the key property when assessing the influence of secondary nucleation. Additional in-line images taken at various point within the batch are available as Supporting Information. The development of the relative supersaturation during comparable batch experiments in the stirred and air-lift crystallizer where the seeding point is closer to saturation shows a different trend (Figure 10) compared to the previous case. The general trend in supersaturation is comparable for both experiments with a gradual increase in relative supersaturation throughout the batch, which indicates that the crystal population is not well capable to consume all generated supersaturation. Note that the relative supersaturation in the air-lift crystallizer remains significantly higher compared to the batch experiment in the stirred crystallizer. The reason behind the difference in initial concentration is not well understood but is likely to be related to some unintended variation in the startup procedure (e.g., solvent evaporation during start-up in the air-lift crystallizer or fine crystals being transferred accidentally to the crystallizers). Toward the end of the batch, some differences can be observed as the supersaturation in the air-lift crystallizer decreases slightly whereas the supersaturation in the stirred crystallizer keeps increasing until the end of the batch. The decrease in supersaturation in the air-lift crystallizer suggests that more crystal surface area is available at the end of the batch compared to the stirred crystallizer, which can be caused by increased crystal growth or by an increase in nucleation. Figure 11 shows the CSD of the final products obtained at the end of both batch experiments. Again for both cases a bimodal distribution can be observed with distinct differences between both CSDs. The peak associated with the grown seeds is located more to the left in case of the stirred crystallizer, which can be explained by two differences between both batches. First, significantly more fines are present, which compete for growth with seeds. Second, the supersaturation throughout the batch was lower, which reduces the crystal growth rate. Note that also the yield in the air-lift crystallizer is higher as a result of the higher initial concentration. A striking observation is that the height of the peak associated with the grown nuclei is still significantly smaller in the air-lift crystallizer compared to the stirred crystallizer despite the higher supersaturation that was present in the air-lift crystallizer.

Figure 7. Dynamic development of the relative supersaturation for Experiments #3 and #7. The lines are to guide the eye. The relative supersaturation is defined as the ratio of the actual concentration over the saturated concentration (from Figure 5).

a nucleation event. Toward the end of the batch, when the temperature decreases more rapidly, the supersaturation in the air-lift crystallizer goes up again, whereas the supersaturation in the stirred crystallizer remains constant, which indicates that more crystal surface area has developed in the stirred crystallizer. The measured CSD of the final product (Figure 8) demonstrates clearly that the larger crystal surface area in the

Figure 8. Crystal size distribution of the initial crystal population (seeds) and of the final product of Experiments #3 and #7. The centers of the vertical lines at each size class correspond to the mean value from a measurement in triplicate and the upper and lower limits of the vertical lines are determined by a single standard deviation in both directions.

stirred crystallizer is the result of significantly more secondary nucleation. The CSD of the final product from both crystallizers has a bimodal shape, but distinct differences can be seen. In the case of the air-lift crystallizer, the height of the peak at the small size classes is smaller. It is likely that the drop in supersaturation around t = 3 h was caused by a nucleation event, which initiated the occurrence of this small peak at the smaller size classes. Crystals with a size in between the small and large peak are practically absent, which indicates that 3269

dx.doi.org/10.1021/cg500090g | Cryst. Growth Des. 2014, 14, 3264−3275

Crystal Growth & Design

Article

Figure 9. Images of produced crystals at the end of Experiment #3 and #7. (A) Experiment #3 (light microscope, 5× magnification), (B) Experiment #3 (SEM), (C) Experiment #7 (light microscope, 5× magnification), (D) Experiment #7 (SEM).

Figure 10. Dynamic development of the relative supersaturation for Experiments #4 and #13. The lines are to guide the eye. The relative supersaturation is defined as the ratio of the actual concentration over the saturated concentration (from Figure 5).

Figure 11. Crystal size distribution of the initial crystal population (seeds) and of the final product of Experiments #4 and #13. The centers of the vertical lines at each size class correspond to the mean value from a measurement in triplicate and the upper and lower limits of the vertical lines are determined by a single standard deviation in both directions.

Similarly as in Figure 8, the characteristics of both CSDs suggest that in the case of the stirred crystallizer a more gradual production of nuclei occurs yielding a broader range of size classes in the corresponding CSD. Although the location of the nucleation peak is similar as in the previous case (Figure 8), the nucleation event that initiated the peak is likely to have occurred later in the batch as the higher supersaturation toward the end of the batch allows for faster growth of the nuclei. The change in direction of the supersaturation development toward

the end of the batch experiment in the air-lift crystallizer may be caused by the higher yield or by a nucleation event late in the batch. Finally, images of the final and intermediate product obtained via in-line imaging and SEM (Figure 12) show nicely faceted crystals obtained from the air-lift crystallizer and some fines and damaged crystals obtained from the stirred crystallizer, although qualitative differences are less pronounced compared to the previous case (Figure 9). Note that the in-line images reveal clearly the bimodal shape of the CSD. A number 3270

dx.doi.org/10.1021/cg500090g | Cryst. Growth Des. 2014, 14, 3264−3275

Crystal Growth & Design

Article

Figure 12. Images of produced crystals of Experiment #4 and #13. (A) Experiment #4 (SEM), (B) Experiment #4 (in-line photo taken 4 h after the start of the batch), (C) Experiment #4 (in-line photo taken 6 h after the start of the batch), (D) Experiment #13 (SEM), (E) Experiment #13 (inline photo taken 4 h after the start of the batch), (F) Experiment #13 (in-line photo taken 6 h after the start of the batch).

whereas the supersaturation during the shorter experiment drops at the end of the batch (Experiment #13 in Figure 10). The former observation suggests that less surface area was available at the end of Experiment #14 to reduce the supersaturation. When comparing the supersaturation for the experiments in the stirred crystallizer (Experiments #5 and #6 in Figure 13), some differences can be observed. First, the supersaturation in general is higher during Experiment #6 compared to Experiment #5. Second, the increase in supersaturation toward the end of the batch levels off during Experiment #6 and keeps on increasing during Experiment #5, which would suggest an increased importance of secondary nucleation during Experiment #6. The CSDs of the products obtained from Experiments #5, #6, and #14 show a difference between the experiments on the impact of nucleation (Figure 14). Nucleation had only a minor impact during the batch experiment in the stirred crystallizer

of additional in-line images obtained at different time points are available as Supporting Information. In the next case, supersaturation is even further reduced by seeding at a low supersaturation and extending the batch time, which allows for increased consumption of supersaturation. Figure 13 illustrates the development of the relative supersaturation during the batch Experiments #5, #6, and #14. Note that also for this case some differences exist in the initial concentration of the batch with the highest supersaturation in the air-lift crystallizer. The general trend of the development of the supersaturation during the batch is comparable for all cases with a rather flat profile of the supersaturation in the first half of the batch and a gradual increase in the second half of the batch. Such trend compares to the previous case (Figure 10) with some subtle differences. In the case of the air-lift crystallizer, the supersaturation during the batch experiment (Experiment #14 in Figure 13) keeps on increasing until the end of the batch, 3271

dx.doi.org/10.1021/cg500090g | Cryst. Growth Des. 2014, 14, 3264−3275

Crystal Growth & Design

Article

that limit and the actual occurrence of nucleation becomes dependent on, for example, small fluctuations in process conditions or impurities. The CSD of the product obtained from the air-lift crystallizer shows also for this case the smallest impact of secondary nucleation on the final CSD (Experiment #14 in Figure 14) despite having the highest supersaturation throughout the batch (Experiment #14 in Figure 13). Finally, several cases have been analyzed in the air-lift crystallizer where seeding with a small seed load was done at a high supersaturation, which moves the operating point of the system closer to the metastable zone limit. The objective is to investigate whether a high growth rate can be sustained for a long period of time in the air-lift crystallizer. The development of the supersaturation shows a similar trend for all those experiments (Figure 15). The supersaturation remains constant

Figure 13. Dynamic development of the relative supersaturation for Experiments #5, #6, and #14. The lines are to guide the eye. The relative supersaturation is defined as the ratio of the actual concentration over the saturated concentration (from Figure 5).

Figure 15. Dynamic development of the relative supersaturation for Experiments #8, #9, #10, #11, and #12. The lines are to guide the eye. The relative supersaturation is defined as the ratio of the actual concentration over the saturated concentration (from Figure 5). Figure 14. Crystal size distribution of the initial crystal population (seeds) and of the final product of Experiments #5, #6, and #14. The centers of the vertical lines at each size class correspond to the mean value from a measurement in triplicate, and the upper and lower limits of the vertical lines are determined by a single standard deviation in both directions.

for about 2 h after which a reduction in supersaturation can be seen both for the unseeded case (Experiment #12) and the seeded cases with small seed load (Experiments #8−11). The drop in supersaturation is caused by the onset of nucleation and, subsequently, the growth of the nuclei, which compete for growth with the seed crystals. Indeed, the CSD of all these seeded experiments show a bimodal shape of the distribution and the CSD of the unseeded experiment exhibits a significant shoulder on the left-hand side of the main peak (Figure 16). Two qualitatively different CSDs can be seen from the seeded experiments. The CSD of the final product from Experiment #8 strongly resembles the CSD of the unseeded case. The supersaturation profile of Experiment #8 (Figure 15) indeed confirms the importance of nucleation as the supersaturation drops from a high value in the first 2 h to a low value in the final hours of the batch. The remaining seeded experiments produced a final product with a CSD that has the main peak located at the higher size classes each with a smaller peak at the lower size classes attributed to grown nuclei. In the latter cases, the growth of the seeds dominates the formation of the final product with nucleation toward the end of the batch, whereas in the former case nucleation may have occurred earlier in the batch, which competes with the growth of the seeds. The qualitatively different CSDs cannot be well explained just by the

with the lowest supersaturation (Experiment #5). The supersaturation during Experiment #6 was slightly higher compared to Experiment #5 (see Figure 13). Although the CSD shows a relatively high standard deviation around the fines peak, nucleation did seem to have an impact during Experiment #6. Note that the occurrence of nucleation reduces the increase in supersaturation toward the end of Experiment #6 compared to Experiment #5 where nucleation did not occur significantly. It is unlikely that the small difference in supersaturation between the experiments in the stirred crystallizers provides a clear explanation for the qualitatively different CSDs. However, it is to be expected that at some point secondary nucleation is also suppressed in a stirred crystallizer upon further reduction of the supersaturation, because of the low survival rate of attrition fragments. The suppression of secondary nucleation at low supersaturation in stirred crystallizers is commonly observed for batch crystallization.38−40 It is possible that the supersaturation in Experiments #5 and #6 are getting close to 3272

dx.doi.org/10.1021/cg500090g | Cryst. Growth Des. 2014, 14, 3264−3275

Crystal Growth & Design

Article

the air-lift crystallizer. Although we have only conducted a limited number of experiments in which primary nucleation was dominant, results suggest that the width of the metastable zone in a stirred crystallizer is larger compared to a air-lift crystallizer for the studied case. This difference might be caused by the presence of a gas−liquid interface, which may induce heterogeneous primary nucleation.45 Comparison with Novel Continuous-Flow Crystallizers. A key motivation for developing the air-lift crystallizer is the absence of moving internal parts, which leads to significantly reduced secondary nucleation compared to a conventional stirred crystallizer under certain operating conditions. The strong recent interest in continuous-flow crystallization has led to novel types of crystallizers that share similarities with the air-lift crystallizer presented in our work. An important similarity is the absence of moving internal parts in several of those novel types of continuous crystallizers. Although a comprehensive discussion is outside the scope of the present work, a comparison between an air-lift crystallizer and novel types of crystallizers that are developed for continuous flow is of interest since hydrodynamic conditions are mostly independent of the mode of operation. Furthermore, the concept of an air-lift crystallizer could easily be extended to continuous flow simply by connecting a number of air-mixed compartments in series to approximate plug-flow behavior while maintaining the favorable properties leading to reduced secondary nucleation. In an oscillatory-baffled crystallizer (OBC), mixed compartments are separated by periodically spaced orifice baffles. Mixing is provided by an oscillatory flow on top of a net flow toward the outlet of the crystallizer.46 In this way, residence time and hydrodynamic conditions in the mixed compartments can be decoupled and plug-flow behavior can be approached with a net velocity that would otherwise result in laminar flow conditions. Several researchers have compared the nucleation behavior in an OBC with a conventional stirred vessel. Chew et al.47 concluded that smaller crystals were produced in an OBC compared to a stirred crystallizer at the same power density and residence time in both crystallizers. Furthermore, the final crystal size in the OBC appeared to be less sensitive for the initial supersaturation. CFD simulations revealed a high local shear rate in a stirred crystallizer around the impeller, but an average shear rate of an order of magnitude smaller compared to the OBC.48 Lawton et al.49 also found a higher nucleation rate in an OBC compared to a conventional crystallizer. However, they also demonstrated that different types of crystals can be obtained by changing the mixing conditions in the OBC. For all cases, narrow size distributions were obtained, which was attributed to the improved mixing conditions compared to a conventional industrial crystallization process.49 Recently, Callahan et al.50 compared the secondary nucleation rate of sodium chlorate in a stirred crystallizer and OBC. No real difference between the two systems was seen in terms of similarity with respect to the seed enantiomer, which indicated comparable nucleation behavior when no scraping was applied. In the present work, it is demonstrated that an air-lift crystallizer consistently produced crystals with a larger mean size for all tested conditions compared to a stirred crystallizer. Therefore, one may expect that an air-lift crystallizer can in principle also produce larger crystals compared to an OBC. A systematic comparison between these two systems would have to prove this hypothesis in future research.

Figure 16. Crystal size distribution of the initial crystal population (seeds) and of the final product of Experiments #8, #9, #10, #11, and #12. The centers of the vertical lines at each size class correspond to the mean value from a measurement in triplicate and the upper and lower limits of the vertical lines are determined by a single standard deviation in both directions.

differences in the experimental parameters (Table 2). Possibly, the variation in obtained product quality is the result of the stochastic nature of nucleation and the strong dependency on gradients in process conditions. The rate of nucleation within an experiment at high supersaturation can, therefore, be unpredictable and result in different shapes of the final CSD. In general, batch-to-batch variations are reported in literature for systems that are dominated by nucleation. For example, Kalbasenka et al.43 report numerous unseeded fed-batch experiments with identical operating conditions for an ammonium-sulfate water system in a pilot-plant crystallizer, which show a wide range of values for supersaturation at which primary nucleation occurs. Consequently, the outcomes of the batches differ significantly. It is likely that also for our case the values of supersaturation at which primary or activated secondary nucleation occur vary and batch-to-batch variations will result as a consequence of that. Therefore, strictly speaking, only qualitative differences in the performance between the stirred and air-lift crystallizer are meaningful when nucleation is present. Reproducibility between batches is improved when seeding is applied such that nucleation is suppressed during the batch.38,43,44 The investigated supersaturation profiles during batch crystallization of L-ascorbic acid in aqueous medium reveal various new insights about the behavior of an air-mixed crystallizer compared to a stirred crystallizer. In the case where the relative supersaturation is high initially and reduces rapidly due to crystal growth, a clear reduction in secondary nucleation could be observed in the air-lift crystallizer compared to the stirred crystallizer. Furthermore, secondary nucleation could be reduced in an air-lift crystallizer compared to a stirred crystallizer for the case where seeding occurs at a low supersaturation despite an overall higher supersaturation in the air-lift crystallizer. Only in the case of a very low supersaturation, it is possible that the performance of a stirred crystallizer starts to resemble the performance of the air-lift crystallizer operated at higher supersaturation. In both cases nucleation is suppressed. On the other hand, at a high supersaturation, we have observed that nucleation can have a significant impact on the final product quality obtained from 3273

dx.doi.org/10.1021/cg500090g | Cryst. Growth Des. 2014, 14, 3264−3275

Crystal Growth & Design

Article

application of state-of-art process control strategies, which are developed for conventional batch crystallization, to an air-lift crystallizer to reveal the full potential of the concept. Furthermore, a question that remains is how well the presented results translate to other systems. Secondary nucleation is very common in industrial crystallization and depends on the material and particle properties of the crystallization system (e.g., hardness, shape, size). For future research, it would be interesting to investigate the influence of these material and particle properties on the difference in performance between a stirred and air-mixed crystallizer. For example, air-mixed crystallizers are likely to demonstrate much different behavior compared to stirred crystallizers for brittle materials with needle-shaped crystals. Finally, from a more fundamental point of view, the availability of kinetic data from different vessels with similar process conditions except the high shear induced locally by an impeller offers opportunities to improve kinetic models for secondary nucleation by separating experimentally the role of supersaturation and shear without compromising on the degree of mixing.

Tubular crystallizers are employed successfully to produce crystals with a narrow CSD in a flexible fashion.51−53 Such crystallizers can be further extended by incorporating a static mixer in the crystallizer to enhance mixing and by using multistage antisolvent addition.54 In a segmented flow tubular reactor, small volumes of slurry in which crystal growth occurs move through a tube. When plug flow is approached, each volume has the same residence time and a narrow crystal size distribution with tailored mean size can be achieved.55−58 In order to provide sufficient residence time for crystal growth, such systems typically operate in the laminar flow regime. Therefore, a nonmiscible phase is introduced to separate the small volumes that are moving through the tube to better approach plug-flow conditions and thus prevent a broad residence time distribution. Jiang et al.58 demonstrate how slugs can form spontaneously due to inherent hydrodynamic stability when choosing carefully the properties of a liquid and gas phase, which assures ease of operation for continuous crystallization. Such systems provide very mild conditions for crystal growth. Therefore, although differences in properties of model systems and operating conditions prevent a one-to-one comparison, one would expect that these systems perform at least as good as an air-lift crystallizer in terms of suppressing secondary nucleation and possibly even better. Experimental studies demonstrate that residence times when using OBCs or tubular crystallizers can be much smaller compared to the residence times that are used in the present work while still achieving sufficient yield, which is beneficial. However, in case needed, larger residence times would be more difficult to attain in a tubular crystallizer compared to an air-lift crystallizer, because of a potentially excessive length of such tubular crystallizer leading to, for example, a high pressure drop and high risks for scaling due to the large surface area. Larger residence times might be needed for systems with a slow integration-controlled growth rate or for products that require large crystals with high purity. In general, systematic comparisons of the performance of an air-lift crystallizer with novel types of continuous-flow crystallizers are an exciting topic for future research.

■

ASSOCIATED CONTENT

S Supporting Information *

Videos representing a series of in situ images obtained from an in-line image probe during the experiments. This material is available free of charge via the Internet at http://pubs.acs.org.

■

AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. Tel:+31152783852. Notes

The authors declare no competing financial interest.

■

ACKNOWLEDGMENTS This work was supported by the European Commissions Framework 7 program through the OPTICO consortium. The authors thank DSM for stimulating discussions and for providing L-ascorbic acid. Furthermore, the authors thank Michel van den Brink for assistance with SEM and titration.

■

CONCLUSION The application of an air-lift crystallizer for the controlled growth of a seed population is, at least for the tested model system, mainly advantageous compared to a stirred crystallizer in the operating region corresponding to an intermediate level of supersaturation. At such intermediate level of supersaturation, the air-lift crystallizer shows significantly reduced secondary nucleation compared to a conventional stirred crystallizer for all tested cases. At low supersaturation, nucleation could be suppressed in both cases, and only at high supersaturation nucleation becomes dominant in the airlift crystallizer as well. The enlargement of the operating window for seeded batch crystallization in an air-lift crystallizer can have important practical consequences. The robustness of the process can improve as more variation is allowed before significant nucleation is induced, which may improve process control. Furthermore, the batch time could potentially be decreased by operating the system at a higher supersaturation and thus at a higher crystal growth rate. The scope of the present work is to investigate differences in performance of a stirred and air-mixed crystallizer, and no attempt has been made to optimize the system for the current application. An interesting topic for future research is the

■

REFERENCES

(1) Myerson, A. S.; Ginde, R. In Handbook of Industrial Crystallization, 1st ed.; Myerson, A. S., Ed.; Butterworth-Heinemann: Stoneham, MA, 1993; Chapter 2. (2) Offermann, H.; Ulrich, J. In Proceedings of the 9th Symposium on Industrial Crystallization; Jancic, S. J.; De Jong, E. J., Eds.; Elsevier: Amsterdam, 1984, 31−36. (3) Van der Heijden, A. E. D. M. J. Cryst. Growth 1992, 118, 14−26. (4) Gahn, C.; Mersmann, A. Chem. Eng. Sci. 1999, 54, 1283−1292. (5) Mersmann, A.; Heyer, C.; Eble, A. In Crystallization Technology Handbook, 2nd ed.; Mersmann, A., Eds.; Marcel Dekker Inc.: New York, 2001; Chapter 2. (6) Jones, A. G.; Mullin, J. W. Chem. Eng. Sci. 1974, 29, 105−118. (7) Chung, S. H.; Ma, D. L.; Braatz, R. D. Can. J. Chem. Eng. 1999, 77, 590−596. (8) Chiu, T.; Christofidis, P. AIChE J. 1999, 45, 1279−1297. (9) Ma, D. L.; Tafti, D. K.; Braatz, R. D. Comput. Chem. Eng. 2002, 26, 1103−1116. (10) Fujiwara, M.; Chow, P. S.; Ma, D. L.; Braatz, R. D. Cryst. Growth Des. 2002, 2, 363−370. (11) Feng, L.; Berglund, K. A. Cryst. Growth Des. 2002, 2, 449−452. (12) Hu, Q.; Rohani, S.; Wang, D. X.; Jutan, A. Powder Technol. 2005, 156, 170−176. 3274

dx.doi.org/10.1021/cg500090g | Cryst. Growth Des. 2014, 14, 3264−3275

Crystal Growth & Design

Article

(49) Lawton, S.; Steele, G.; Shering, P. Org. Process Res. Dev. 2009, 13, 1357−1363. (50) Callahan, C. J.; Ni, X. CrystEngComm 2014, 16, 690−697. (51) Eder, R. J. P.; Radl, S.; Schmitt, E. K.; Innerhofer, S.; Meier, M.; Gruber-Woelfler, H.; Khinast, J. G. Cryst. Growth Des. 2010, 10, 2247− 2257. (52) Eder, R. J. P.; Schmitt, E. K.; Grill, J.; Radl, S.; Gruber-Woelfler, H.; Khinast, J. G. Cryst. Res. Technol. 2011, 46, 227−237. (53) Ferguson, S.; Morris, G.; Hao, H.; Barrett, M.; Glennon, B. Chem. Eng. Sci. 2012, 77, 105−111. (54) Alvarez, A. J.; Myerson, A. S. Cryst. Growth Des. 2010, 10, 2219−2228. (55) Vacassy, R.; Lemaître, J.; Hofmann, H.; Gerlings, J. H. AIChE J. 2000, 46, 1241−1252. (56) Eder, R. J. P.; Schrank, S.; Besenhard, M. O.; Roblegg, E.; Gruber-Woelfler, H.; Khinast, J. G. Cryst. Growth Des. 2012, 12, 4733− 4738. (57) Jongen, N.; Donnet, M.; Bowen, P.; Lemaître, J.; Hofmann, H.; Schenk, R.; Hofmann, C.; Aoun-Habbache, M.; Guillemet-Fritsch, S.; Sarrias, J.; Rousset, A.; Viviani, M.; Buscaglia, M. T.; Buscaglia, V.; Nanni, P.; Testino, A.; Herguijuela, J. R. Chem. Eng. Technol. 2003, 26, 1521−4125. (58) Jiang, M.; Zhu, Z.; Jimenez, E.; Papageorgiou, C. D.; Waetzig, J.; Hardy, A.; Langston, M.; Braatz, R. D. Cryst. Growth Des. 2014, 14, 851−860.

(13) Fujiwara, M.; Nagy, Z. K.; Chew, J. W.; Braatz, R. D. J. Process Control 2005, 15, 493−504. (14) Chew, J. W.; Black, S. N.; Chow, P. S.; Tan, R. B. H. Ind. Eng. Chem. Res. 2007, 46, 830−838. (15) Nowee, S. M.; Abbas, A.; Romagnoli, J. A. Cryst. Growth Des. 2008, 8, 2698−2706. (16) Nagy, Z. K.; Chew, J. W.; Fujiwara, M.; Braatz, R. D. J. Process Control 2008, 18, 399−407. (17) Nagy, Z. K. Comput. Chem. Eng. 2009, 33, 1685−1691. (18) Mesbah, A.; Huesman, A. E. M.; Kramer, H. J. M.; Nagy, Z. K.; Van den Hof, P. M. J. AIChE J. 2011, 57, 1557−1569. (19) Nagy, Z. K.; Aamir, E. Chem. Eng. Sci. 2012, 84, 656−670. (20) Aamir, E.; Rielly, C. D.; Nagy, Z. K. Ind. Eng. Chem. Res. 2012, 51, 16677−16687. (21) Nagy, Z. K.; Braatz, R. D. Annu. Rev. Chem. Biomol. Eng. 2012, 3, 55−75. (22) Saleemi, A.; Rielly, C.; Nagy, Z. K. CrystEngComm 2012, 14, 2196−2203. (23) Gahn, C.; Krey, J.; Mersmann, A. J. Cryst. Growth 1996, 166, 1058−1063. (24) Gahn, C.; Mersmann, A. Chem. Eng. Sci. 1999, 54, 1273−1282. (25) Neumann, A. M.; Bermingham, S. K.; Kramer, H. J. M.; Van Rosmalen, G. M. J. Cryst. Growth 1999, 198/199, 723−728. (26) Briesen, H. Chem. Eng. Sci. 2009, 64, 661−672. (27) Bao, J.; Koumatsu, K.; Furumoto, K.; Yoshimoto, M.; Fukunaga, K.; Nakao, K. Chem. Eng. Sci. 2001, 56, 6165−6170. (28) Arimatsu, Y.; Bao, J.; Furumoto, K.; Yoshimoto, M.; Fukunaga, K.; Nakao, K. J. Chem. Eng. Jpn. 2004, 37, 1035−1040. (29) Jenkins, D. H. Can. J. Chem. Eng. 2009, 87, 869−878. (30) Gonzalez-Contreras, P.; Weijma, J.; Buisman, C. J. N. Cryst. Growth Des. 2012, 12, 2699−2706. (31) Bao, J.; Koumatsu, K.; Arimatsu, Y.; Furumoto, K.; Yoshimoto, M.; Fukunaga, K.; Nakao, K. Biochem. Eng. J. 2003, 15, 177−184. (32) Lakerveld, R. Development of a task-based design approach for solution crystallization processes. Ph.D. Dissertation, Delft University of Technology, Delft, 2010. (33) Lakerveld, R.; Kramer, H. J. M.; Jansens, P. J.; Grievink, J. Comput. Chem. Eng. 2009, 33, 1692−1700. (34) Soare, A.; Lakerveld, R.; Van Royen, J.; Zocchi, G.; Stankiewicz, A. I.; Kramer, H. J. M. Ind. Eng. Chem. Res. 2012, 51, 10895−10909. (35) Wierzbowska, B.; Piotrowski, K.; Koralewska, J.; Matynia, A.; Hutnik, N.; Wawrzyniecki, K. Cryst. Res. Technol. 2008, 43, 381−389. (36) Nyvlt, J. Collection Czechoslov. Chem. Commun. 1976, 41, 342− 349. (37) Kulkarni, A. V.; Badgandi, S. V.; Joshi, J. B. Chem. Eng. Res. Des. 2009, 87, 1612−1630. (38) Doki, N.; Kubota, N.; Sato, A.; Yokota, M.; Hamada, O.; Masumi, F. AIChE J. 1999, 45, 2527−2533. (39) Jagadesh, D.; Yokota, M.; Doki, N.; Sato, A. J. Chem. Eng. Jpn. 1999, 32, 514−520. (40) Kubota, N.; Doki, N.; Yokota, M.; Sato, A. Powder Technol. 2001, 121, 31−38. (41) Apelblat, A.; Manzurola, E. J. Chem. Thermodyn. 1989, 21, 1005−1008. (42) Heijnen, J. J.; Hols, J.; Van der Lans, R. G. J. M.; Van Leeuwen, H. L. J. M.; Mulder, A.; Weltevrede, R. Chem. Eng. Sci. 1996, 52, 2527−2540. (43) Kalbasenka, A. N.; Spierings, L. C. P.; Huesman, A. E. M.; Kramer, H. J. M. Part. Part. Syst. Charact. 2007, 24, 40−48. (44) Kubota, N.; Onosawa, M. J. Cryst. Growth 2009, 311, 4525− 4529. (45) Wohlgemuth, K.; Kordylla, A.; Ruether, F.; Schembecker, G. Chem. Eng. Sci. 2009, 64, 4155−4163. (46) Ni, X.; Valentine, A.; Liao, A.; Sermage, S. B. C.; Thomson, G. B.; Robert, K. J. Cryst. Growth Des. 2004, 4, 1129−1135. (47) Chew, C. M.; Ristic, R. I. AIChE J. 2005, 51, 1576−1579. (48) Chew, C. M.; Ristic, R. I.; Reynolds, G. K.; Ooi, R. C. Chem. Eng. Sci. 2004, 59, 1557−1568. 3275

dx.doi.org/10.1021/cg500090g | Cryst. Growth Des. 2014, 14, 3264−3275