Novel Technique for Filtration Avoidance in Continuous Crystallization

Article pubs.acs.org/crystal

Novel Technique for Filtration Avoidance in Continuous Crystallization Nima Yazdanpanah, Steven T. Ferguson, Allan S. Myerson, and Bernhardt L. Trout* Department of Chemical Engineering, Massachusetts Institute of Technology, Cambridge, Massachusetts 02139, United States Novartis-MIT Center for Continuous Manufacturing, Cambridge, Massachusetts 02139, United States S Supporting Information *

ABSTRACT: Fouling and clogging of lines in both continuous crystallization and flow chemistry still present a challenge in designing continuous processes. This paper presents a proof of concept for a novel unit operation designed for purification and to eliminate the need for slurry handling in API manufacturing by utilizing a dynamic (falling) film-based solution layer crystallization technique called “filtrationavoidance”. Fenofibrate (with fenofibric acid as an impurity) and acetaminophen (with metacetamol as an impurity) in different solvents were tested as crystallization systems, and the final purities and yields are reported. Sixteen different combinations of materials and process parameters were investigated. The results show significant improvement in the purity of the main compounds and a high yield for the process, which indicates a potential industrial application for a filtration-free continuous separation process, e.g., final purity of 99.4% was achieved from the experiment with a solution of 98% fenofibrate and 2% fenofibric acid (impurity) in pure ethanol (at a flow rate of 15 mL/min and a column temperature of 0 °C). somewhat compartmentalized. Continuous filtration is seldom utilized in the literature or at small scales of operation. In the continuous manufacture of Aliskiren, a continuous disk-based filter with an auger for the conveying of wet cake was utilized.13 Such a design or similar belt-type designs are simple and inherently scalable. Continuous drying is more challenging. Screw-based dryers, e.g., fluidized beds and cyclones, have been used in industries where continuous or semicontinuous processing is common. It is not yet clear if such truly continuous isolation and drying operations will hold significant benefit compared to the parallel use of batch-filter dryers at pharmaceutical scales of operation. Several powder-free formulation techniques that could enable powder-/slurry-free production of orally dosed tablets have been developed recently: crystallization on the surface of polymers,11 cast within polymers,29 crystallization within codissolved polymers via spray drying or electro-spraying or -spinning,8 and crystallization within hydrogels.10,30 Decoupling the purification and isolation in filtration avoidance supplies in a single operation may be beneficial when specific isolation techniques that require impurity-free solutions are utilized to make composite or specialized solid phases that must be dried to completion for isolation.12,31 The falling film solution layer crystallization demonstrated here is analogous to falling film melt crystallization. Industrial melt crystallization would meet the requirements outlined for

1. INTRODUCTION Over the past decade, significant advances in many continuous operations aimed at pharmaceutical scales of operation have been made, e.g., flow chemistry, 1−5 extractions, 6 and formulations.7−12 The first end-to-end continuous production run of an API (Aliskiren) from an intermediate to the formulation of the final tableting was demonstrated in 2012.13 Crystallization is no exception in this regard, with standard tank and plug flow designs implemented for API and model organic systems on several occasions.14−20 Several alternate designs that can approximate plug flow conditions in tubular crystallizers at lower flow rates and hence longer residence times have also been demonstrated.21−26 In flow chemistry, the presence of solids is generally avoided where possible to eliminate the possibility of blocking or fouling of the process equipment. Because the goal of continuous crystallization is to generate solids in flow, concern is often attributed to the operation of continuous crystallizers for long periods of time on pharmaceutical scales of operation. Several methodologies have recently been outlined in the literature demonstrating a variety of designs and strategies that should enable continuous crystallizations to be successfully implemented.27 These include a novel pneumatic transfer mechanism for stirred-tank crystallization to prevent the blockage or fouling of transfer lines.14,15,20 Corrective action to counteract fouling can also be implemented without shutting down continuous crystallizers by adding solvent in segmented flows or by using segmented jackets with the periodic heating of individual sections.28 Despite some notable exceptions, the development of continuous operations for pharmaceutical purposes can be © XXXX American Chemical Society

Received: August 25, 2015 Revised: November 16, 2015

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DOI: 10.1021/acs.cgd.5b01231 Cryst. Growth Des. XXXX, XXX, XXX−XXX

Crystal Growth & Design

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the filtration avoidance concept detailed here. Industrial melt crystallization is not suitable for operation with most APIs because it requires the compound to be in the molten state for relatively long periods of time. Despite this, it is useful to study the literature on melt crystallization, as many of the operations outlined here also have direct comparisons that can be considered by analogy. Furthermore, the equipment and operating strategies used for both melt and solution layer crystallization would be very similar. The mechanisms of the two techniques are different: most strikingly, in melt, there is no solvent, and the compound itself is in essence the solvent for impurities, which are forced out by an advancing solid-phase layer. Additionally, the rate of mass and heat transfer is orders of magnitude lower than in a melt due to the high viscosity of melts. This has implications for mass transfer effects in purification in melt crystallization, but it is unlikely to have the same impact in solution layer crystallization. A good overview of melt crystallization can be found in the handbook of industrial crystallization.32 This paper highlights the development of alternate approaches to continuous crystallization referred to as filtration avoidance (via crystallization) to eliminate powder handling/ slurry transfer operations, which is achieved via the implementation of a dynamic film-based solution layer crystallization technique here.

(based on the solubility and saturation curve at the specific temperatures) and 5 mL of solvent to a 10 mL glass vial. The concentrations of the APIs were selected to be slightly above the saturation point at certain temperatures. The vials were then placed in the water bath at the control temperature. The samples were stirred with a controlled stirring speed of 300 rpm, and left in the bath for 48 h. Each experiment was performed twice at the different temperatures of 0, 10, 20, and 30 °C. Because the concentrations were very close to the saturation point, the slurry system was close to equilibrium, and the crystals were assumed to form in equilibrium conditions over the long residence time. At the end of the experiments, the crystals were filtered out and washed with 0.5 mL of cold solvent under vacuum to wash off the residual solvent and the impurities on the surface of the crystals. The harvested crystals were dissolved in the relevant mobile phase for HPLC, and the amount of impurity in the crystals was measured by the HPLC technique. The equilibrium distribution coefficient was calculated by the formula and reported as a function of temperature. 2.5. Thermogravimetric Analysis. Thermogravimetric analysis (TGA) of the samples was performed on a thermal gravimetric analyzer (TA Instruments, TGA Q5000), and the samples (approximately 5 mg) were heated from room temperature (30 °C) to 300 °C at a heating rate of 10 °C/min under a nitrogen flow. The results were analyzed by Universal Analysis 2000 V4.7A software, and the weight loss of the samples was reported on a mass basis as the amount of solvent inclusion. 2.6. Crystallization Solution Systems. Three different solution systems were prepared for the experiments, which are listed in Table 1.

2. MATERIALS AND METHODS

Table 1. Crystallization Solution Systems for the Falling Film Experiments

2.1. Materials. Acetaminophen (4-acetamidophenol) and metacetamol (3-acetamidophenol) with >98% purity were purchased from Sigma-Aldrich. Fenofibrate and fenofibric acid (fenofibrate impurity B) with >99% purity were purchased from the Xian Shunyi Bio-Chemical Technology Co., Ltd. Materials were used without any further purification. Deionized water was produced from a Barnstead Easypure II (Thermo Scientific). Ethyl acetate and ethanol (200 proof) were purchased from VWR International. 2.2. High Performance Liquid Chromatography (HPLC). The HPLC instrument (Agilent 1100) was equipped with a UV diode array detector. The column used was a YMC-Pack ODS-A (150 mm × 4.6 mm i.d.) packed with 3 μm particles with a 12 nm pore size (YMC America Inc.). The detection wavelength was set at 254 nm for acetaminophen and 230 nm for metacetamol. The samples were analyzed using an isocratic method with a 30:70 methanol:water mobile phase containing 0.3% trifluoroacetic acid for 10 min. The detection wavelength was set at 280 nm for fenofibrate and fenofibric acid. The samples were analyzed using an isocratic method with a 70:30 acetonitrile:water mobile phase containing 0.1% trifluoroacetic acid for 15 min. 2.3. Solubility Measurement. The saturation temperatures of APIs were measured in the solvents at different concentrations by adding a known amount of APIs and 1 mL of solvents, respectively, to a 1.5 mL glass vial. The vials were then placed in the Crystal16 (Aventium Technology BV, Netherlands), and the heating and cooling rates were set to 0.3 °C/min. The samples were stirred with a controlled stirring speed of 300 rpm using magnetic stirring bars. The samples were heated with a heating rate of 0.3 °C/min from −10 to 60 °C. The CrystalClear software (Aventium Technology BV, Netherlands) was used to analyze the results from the machine. The temperature at which the suspension turned into a clear solution was recorded and assumed to be the saturation temperature. After a waiting time of 30 min at 60 °C, the clear solution was cooled to −10 °C with a cooling rate of 0.3 °C/min to recrystallize the APIs. Then, the same temperature profile was repeated three times for each sample. 2.4. Equilibrium Distribution Coefficient Measurement. The equilibrium distribution coefficients of the APIs and relevant impurities were measured in the solvents at different initial concentrations of impurity by adding a known amount of APIs

systems System I System II System III

main compound (initial concentration, W %)

impurity (initial concentration, W %)

acetaminophen (95%) metacetamol (5%) fenofibrate (98%) fenofibric acid (2%) fenofibrate (98%)

fenofibric acid (2%)

solvent ethanol and water (50:50 vol%) ethyl acetate and ethanol (70:30 vol%) ethanol (100%)

Mixtures of 50:50 (volume basis) ethanol:water, 70:30 (volume basis) ethyl acetate:ethanol, and 100% pure ethanol at room temperature were prepared as solvents. From the solubility results, saturated solutions of the compounds were prepared at 60 °C by adding a known amount of APIs and 150 mL of solvents, respectively, to a 250 mL glass bottle. The solutions were stirred with a controlled stirring speed of 300 rpm using magnetic stirring bars and maintained at a constant temperature of 60 °C for the entirety of the experiments. 2.7. Sampling Method for the Deposited Layer of Crystal. As will be described in the Results section, the concentration of impurities in the deposited crystal layer is not constant and has an axial and radial profile. Therefore, solid samples were taken from the top, middle, and bottom sections (axial direction) of the deposited layer. Three small (1 × 1 cm) pieces of the crystal layer from the aforementioned positions were taken. The pieces were ground and mixed together to make a homogeneous mixture of powder, representing the overall quality of the deposited layer. The powder samples were maintained under vacuum at 60 °C for approximately 2 h to evaporate the residual solvents and obtain an accurate dry mass. This mixture was used for the HPLC sample preparation for measuring the amount of impurities. For TGA analysis, the pieces were gently crushed separately and marked as samples for the relevant positions. These samples were left at ambient conditions in an open bottle to evaporate the residual solvent on the surface of the crystals and were not vacuum-dried to retain the amount of occluded solvent between the diffused crystals. B



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Figure 1. Schematic of slurry- and powder-free manufacturing scheme utilizing processes capable of filtration avoidance, such as falling film solution layer crystallization, to avoid the transfer and isolation of solids.

employed in this study is a falling film. A schematic diagram of the falling film operation, an image of the experimental test rig and a schematic drawing of the column can be seen in the Supporting Information. A superheated feed is distributed as a film over a cooled tube. The solution rapidly cools as it flows down the cold surface and becomes supersaturated. Nucleation occurs on the surface of the tube, which, in this case, is sandblasted to aid nucleation in this process, and later the subsequent growth on this nucleated layer occurs. The solution falls from the cold element into a superheated well where it can then be recycled to be distributed on the film again; a percentage of this will be sent on to another stage or waste stream in continuous operation. The column is sealed, and the second outside interface is composed of air or a gas phase of some type in which the vapor pressure can be controlled. This can also be used for evaporation from the surface (vacuum) or for the stripping of solvent (dry gas flow) to boost the process yield or even to drive the crystallization process itself. The stripping off of solvent with dry gas as the primary driving force for crystallization will enable far thicker crystal layers to be generated because heat transfer limitations are eliminated. In theory, an antisolvent or a reagent for a reactive crystallization could also be brought in through the gas phase. Once the crystal layer has been built up to the desired thickness for a given continuous processes or the mother liquor phase reaches equilibrium, the crude solution is removed from the column. At this point, a solid crystal layer is left on the column that can undergo secondary purification operations, if desired (washing), to remove entrained solvents and impurities. Clean solvent can be brought in to dissolve or partially dissolve the layer, bringing purified product on to the next operation, whether that be another reaction, a powder-free formulation, another falling film stage, or for further purification. Figure 2 shows the process flow diagram for the crystallization process in recirculation mode. The solution is dissolved in a hot feed tank, maintained at a constant temperature of 60 °C, and stirred with a magnetic bar for the entire experiment. This solution is pumped from the feed tank through heated lines, and the temperature in these lines is maintained above the solubility point of the solution using a tube-in-tube heat exchanger configuration controlled by an

3. PROCESS AND EQUIPMENT DESCRIPTION; FALLING FILM CRYSTALLIZER SETUP 3.1. Concept. Figure 1 shows a general schematic of the overall filtration avoidance concept. This approach aims to eliminate all powder or slurry transport operations required in current API manufacturing. This technique involves crystallization on a surface from a thin liquid film of crude mother liquor flowing over its surface; the crude solution is then removed, and the layer is redissolved, allowing the purified solution to be sent to the next operation without the generation of a slurry. As per Figure 1, this allows crystallization, isolation, drying, resuspension, and redissolution, and the solid phase transport mechanisms between them to be condensed into a single-unit operation. The elimination of all slurry promises to make this operation far more robust than continuous suspension crystallization in terms of avoiding blockages and fouling during continuous operation. Using the clean solution generated from the layer crystallization in conjunction with a powder-free formulation technique avoids the need to transfer slurries or powders at any point in the manufacturing process. In developing a flow synthesis, multiple steps are telescoped together to minimize the need to conduct continuous isolation and reconstitutions, which can be extremely arduous to implement. In trying to avoid isolation, care must be taken to minimize dilution as an increasing amount of streams are added and also to avoid propagation of impurities that may negatively impact downstream reactions. Alternative approaches that utilize crystallization to deliver clean liquid phase streams for downstream operations use standard equipment, but will not be capable of completely eliminating the generation of slurries, as per the solution-layer crystallization proposed here. A conventional continuous suspension crystallization, with continuous or batch inline filtration followed by washing and then dissolving from the filter, would provide a purified stream of any solvent for the next operation. However, this will still face the challenges of conveying a slurry, fouling, and slow filtration that are eliminated via layer crystallization. A work-around would be to implement the broader concept of the slurry/powder-free production of APIs outlined in Figure 1. 3.2. Falling Film Solution Crystallization and Operating Set-Up. The dynamic film-based solution crystallization C



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surface of the cold column, which has been sandblasted and includes the micropitches.

Figure 3. Column surface (sandblasted) with the Nylon rings (micropitches) installed.

The residual solution is recirculated to the feed tank by a drain pump. However, it is possible to design the process in single-pass mode and have a separate drain tank. In this case, the returning solution, which is rich in impurities but has a lower concentration of the API, does not mix with the feed, and the concentration of the feed tank remains constant during the experiment. Both methods have their own advantages and disadvantages in terms of overall purity and yield of the process. At the start of the experiment, the concentration of the solution is high, and there is no deposited layer of crystals on the surface of the column. The starting point (red line) is at 60 °C for the high-concentration feed. Once the solution is introduced to the falling film crystallizer, the heat transfer takes place rapidly, and the temperature drops to an equilibrium temperature that is close to the column’s cold temperature (in this case, 0 °C). At this supersaturation level, the API precipitates on the surface of the column in the form of crystals, and an integrated layer of the crystals is formed on the column. Because the process is in recirculation mode, the concentration in the feed tank reduces over time by the dilution effect of returning low-concentration solution. Therefore, as the experiment progresses, the process path moves downward to a new starting point at a lower concentration. The deposition of the crystal layer during the experiment acts like an insulator over the cold column because the thermal conductivity of the API is significantly lower than that of steel. Therefore, the new equilibrium temperature is slightly higher, and the supersaturation ratio is therefore lower. This change in supersaturation ratio lowers the crystallization rate (and the deposition rate), and, at some point in time, the change in concentration in the feed tank (de-supersaturation profile) becomes minimal. Layer crystallization processes, whether in melt or in the proposed solution-layer crystallization process, result in a large reduction in the surface area of a layer compared to the equivalent suspension crystallization. The unit used in this study is for design purposes only; many tubes in a single unit supplying more surface area would be implemented commercially to boost the productivity of the device. However, the surface area will still be much lower than in suspension crystallization. This will impact the relative purification efficiency compared to suspension crystallization. However, it can be offset by utilizing multiple stages to meet the purification requirements needed in API manufacturing. The proposed solution-layer crystallization process can be operated at higher concentrations than are possible in suspension crystallizers (and can even be brought to dryness without disrupting the operation), which should in many cases mitigate or reverse the yield losses associated with a multiple stage operation. The

Figure 2. Process flow diagram.

internal PT100 thermocouple (in the heating medium). As Figure 2 shows, the temperature is monitored and controlled at many points to ensure that the solution remains above the solubility point and avoids bulk nucleation and subsequent blockage of the lines due to solid formation. Peristaltic pumps are used to transfer solution. Unheated sections are kept to a minimum and insulated. The solution is pumped into the top well of the crystallization column. The cooling temperature inside the cold column is controlled by an external chiller and has a range of −10 to 30 °C for the different experiments. The cooling media is a mixture of 30% ethylene glycol and 70% water in mass, and its flow rate is 24 L/min. As the solution is pumped into the top well, the level increases and spills over into the distributor, where it forms a film on the surface of the cold tubular element. The core is a 40-cm-long stainless steel 304 tube with an outer diameter of 127 mm and a thickness of 16 mm. The flow rate through the system is set by the flow rate of the peristaltic pump. The internal cold tube is not connected to the heated distribution top of the column, but is sealed using a standard vacuum fitting at the top of the column. The gap size between the distributor and the internal coldfinger is 1 mm. This gap size in the distributor may need to be adjusted depending on the wettability of the column for a given solution, and, in some cases, other construction materials may need to be considered. The column was leveled before running each experiment to ensure an even distribution of fluid. The surface tension of the solution, the flow rate of the falling liquid, and the wettability of the column govern the uniform film formation around the cold core. It is a challenge to maintain the uniformity of the thin falling film around the pipe at lower flow rates (less than 10 mL/min in some cases). Therefore, some “regional redistributors” are introduced on the surface of the pipe in the form of micropitches. The micropitches are Nylon rings with a 0.38 mm thickness that are placed at a 1 cm distance parallel to provide local redistribution of the flow and improve the mixing of the film in the initial stages of the experiments. Nylon rings are used because they are easy to install to mix and redistribute the flow. However, these micropitches can be composed of the same material as the pipe and formed by etching the surface. Figure 3 shows the D



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multiple stage operation to enhance the purification would be designed to meet a specific required purity, and the number of required stages and the optimum operating conditions could be examined experimentally or by mathematical modeling. A mathematical model, the results of which were verified by several experiments, has been developed for this purpose and was published elsewhere.33 There are several other falling film geometries (internal tube, plates, horizontal tubes) with advantages and disadvantages. Plates are stackable, which could make it easy to create multiunit operations. Tubes are more efficient for packing surface area into a given reactor volume. For these experiments, the column and the entire process are washed, cleaned, and dried thoroughly, and a fresh clean dry column was used every time. It is possible, and recommended, to not fully dissolve the deposited layer from previous experiments and use the partial layer as a “seed” for growing the new layer of crystal on top. However, because we were recording the desupersaturation profiles and comparing different process parameters, this time saving method could not be utilized.

Figure 4. Purification results for acetaminophen in ethanol:water with 5% initial impurity (system I) for different flow rates and cooling temperatures, compared with the equilibrium distribution.

in which the roles of flow rate and temperature are notable in this set of experiments. As Figure 4 shows, the highest flow rate improved the purity. This improvement in purity is the result of a thinner boundary layer at the higher flow rate because the concentration of the impurity (the concentration profile of the impurity in the radial direction) is lower in the thinner boundary layer on the growing face of the crystal. Therefore, the rejection rate is higher, and less impurity is entrapped in the forming crystal lattice. Conversely, the lowest level of purity corresponds to the lowest flow rate of 5 mL/min, when the rejection rate of impurity on the boundary layer is less. With a starting amount of 5% impurity content at 0 °C, flow rates of 5, 20, and 30 mL/ min gave average final purities of 96.6%, 96.8%, and 97.1%, respectively. As described previously for the relative solubility in the system and the distribution coefficient, the increase in temperature improves the purification. At 10 °C, 5 mL/min gives an average final purity of 97%, and 20 mL/min gives 97.4% final purity. The experiment for 30 mL/min at 10 °C was unsuccessful due to the limited heat transfer rate when a uniformly integrated deposited layer could not be formed. The deposition at this temperature and flow rate is not even, and a few patches of “islands” of crystals formed on the column. The same pattern is observed at higher temperatures and lower flow rates; therefore, 10 °C is the temperature limit for this experiment. Figure 5 demonstrates the purification results for system II at flow rates of 10, 20, and 40 mL/min and a temperature range of 0 to 10 °C. The top blue line is for the ideal purification based on the equilibrium distribution coefficient. For this system, the experiment at a higher flow rate of 40 mL/min is successful because the supersaturation ratio for this system is relatively higher, which provides enough driving force to deposit a substantial amount of solid on the column. With a starting 2% impurity content at 0 °C, 10, 20, and 40 mL/min gave average final purities of 98.4%, 98.8%, and 98.9%, respectively. The overall purification results for this system are not significant, which is the result of the high growth rate for this system. The rapid crystal formation and dramatic growth and diffusion causes the inclusion of impurities in the layer (inside the crystal and in the cake) to be large. At 10 °C, the results are even worse, with average final purities of 98.3% and 98.4% at 10 mL/min and 20 mL/min, respectively. The

4. RESULTS 4.1. Purity for the Falling Film Crystallization Experiments. The systems described in the previous sections were used for the falling film crystallization experiments. Based on the solubility and distribution coefficient results (Appendix A), different final purities and overall yields were expected. The growth rates, based on supersaturation, ΔC, primarily govern the processing time and final impurity content; in recirculation mode, the system with the higher supersaturation (system II) has the potential to rapidly deposit the solid layer on the column in a shorter amount of time. The experiments were conducted at different flow rates of 5−40 mL/min and at temperatures of 0−10 °C. The temperature and flow rate ranges depend on the dimension of the column, and these results are for the bench-scale setup. The limits are associated with the following: (1) The surface tension of the solution (wettability), when the uniform film formation at a very low flow rate is challenging (although the micropitches perform satisfactorily for very low flow rates). (2) The size of the overhead reservoir and the distributor size, where higher flow rates could result in flooding. (3) A short length and small diameter of the column, which limit the heat transfer area in cooling the solution instantly and providing enough area for the mass deposition. (4) The size of the connecting tubes, which could cause issues in handling highly concentrated solutions and in clogging at the bottom of the column. Figure 4 represent the purification results for system I at flow rates of 5, 20, and 30 mL/min for a temperature range of 0 to 10 °C. The top blue line is for ideal purification based on the equilibrium distribution coefficient. The gap between this line and the experimental results is the result of cooling crystallization, which is far from the equilibrium condition, as well as solvent and impurity inclusion. The solvent and impurity inclusion is for the overall entrapped impurity in the crystal matrix, due to rapid crystal formation, and between diffused crystals in the deposited layer (crystal cake). As mentioned before, the ultimate goal is to reduce this gap and achieve further purification through the falling film technique, E



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increases in flow rate improved the final purity over the range of temperatures. The experiment at 15 mL/min and 10 °C was again unsuccessful due to the limited heat transfer rate when a uniform integrated deposited layer could not be formed. 4.2. Yields for the Falling Film Crystallization Experiments. Figures 7−9 show the average yields for the different

Figure 5. Purification results for fenofibrate in ethyl acetate:ethanol with 2% initial impurity (system II) for different flow rates and cooling temperatures, compared with the equilibrium distribution.

experiment at 40 mL/min and 10 °C was again unsuccessful due to the limited heat transfer rate when a uniform integrated deposited layer could not form. The changes in purity for the different temperatures at 10 and 20 mL/min are not similar. Hence, many stochastic parameters are involved in the impurity inclusion; the slope at 20 mL/min is relatively wider, resulting in a significant decrease in purity at 20 mL/min and 10 °C. Figure 6 represents the purification results for system III at flow rates of 5, 10, and 15 mL/min for a temperature range of 0

Figure 7. Yield results for acetaminophen in ethanol:water with 5% initial impurity (system I) for different flow rates and cooling temperatures.

Figure 8. Yield results for fenofibrate in ethyl acetate:ethanol with 2% initial impurity (system II) for different flow rates and cooling temperatures.

systems at different temperatures and flow rates, which are all within the range of 65−76%. The yields are calculated based on the theoretical concentration difference at the cooling temperature, assuming the falling film has the same temperature as the cold column in the late stage of the experiments. In fact, toward the end of the experiment, the thin film of the solution on the column has a different temperature than the cold temperature of the column due to the formation of a less conductive layer deposited on the column surface. This variation in temperature was described before. Due to the recirculation of the solution and the heating of the feed tank to maintain the 60 °C feed temperature, the solution entering the top of the column is undersaturated. Therefore, this unsaturated warm solution dissolves a portion of the deposited layer (on the top of the column), and, if the heat transfer rate is adequate enough in the axial direction, this dissolved portion can redeposit on the lower part of the column. The proper mitigating action for this issue

Figure 6. Purification results for fenofibrate in ethanol with 2% initial impurity (system III) for different flow rates and cooling temperatures compared with the equilibrium distribution.

to 10 °C. For this system, the experiments at a higher flow rate of more than 15 mL/min were unsuccessful. The hypothesis for this observation involves the difference in the induction time for this system, although the induction time was not measured for the systems. With a starting 2% impurity content at 0 °C, 5, 10, and 15 mL/min gave average final purities of 99.2%, 99.3%, and 99.4%, respectively. The overall purification results for this system are significantly improved compared to system II due to “proper” solvent selection, which shows the effect of crystal growth rate in impurity inclusion. At 10 °C, 5, 10, and 15 mL/min gave average final purities of 99.1%, 99.3%, and 99.4%. The effect of flow rate on the purification is distinguishable here; the F



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Table 2. Average Yield and Purity Results for Falling Film Crystallization Experiments with System I column temperature yield (%)

purity (%)

flow rate

0 °C

10 °C

0 °C

10 °C

5 mL/min 20 mL/min 30 mL/min

71 69 68

66 65

96.6 96.8 97.1

97 97.4

Table 3. Average Yield and Purity Results for Falling Film Crystallization Experiments with System II column temperature yield (%)

Figure 9. Yield results for fenofibrate in ethanol with 2% initial impurity (system III) for different flow rates and cooling temperatures.

is to control the feed tank temperature and decrease the temperature slightly during the experiment. In this way, toward the end of the experiment, the cooler solution entering the column will be in a condition closer to the saturation line. The other possible reason for this range of yield is the metastable zone width for the systems, which were not measured for these experiments. The hypothesis is that the supersaturation ratio is so small at the end of the experiment, and the solution remains slightly above the saturation point at a specific axial position without forming any new crystals or significantly contributing to crystal growth. To maximize the yield of the process, the cold column temperature could be lowered. The experiments were performed in a range of 0 to 20 °C. The lowest level of 0 °C was low enough to maximize the yield without going beyond the capacity of the chiller, although a few runs were tested at −10 °C. Figure 7 represents the yields for system I; at 0 °C, 5, 20, and 30 mL/min gave average yields of 71%, 69%, and 68%, respectively. At 10 °C, 5 mL/min resulted a yield of 66%, and 20 mL/min had a 65% yield. As described for the purity results, the experiment at 30 mL/min and 10 °C was unsuccessful due to the limited heat transfer rate when a uniform integrated deposited layer could not be formed. Although the effect of flow rate on the overall yield is not significant, a maximum of 5% for this system, the overall trends show that the increases in the flow rates depress the yield over the range of temperatures studied. The same trend is observed for the increase in temperature; the yield is lower at higher temperatures. Similarly, a key reason for this behavior lies in the obvious temperature differences and the level of the supersaturation ratio. The change in solubility at a temperature slightly above 0 °C is relatively much less than the change in solubility at temperatures above 10 °C. Due to the exponential trend of the solubility curve, the slope of the solubility curve for the 0−5 °C section is much less than the slope of the solubility curve for the 10−15 °C section. Although the surface temperature profiles of the deposited layer and the thin falling film were not measured, these profiles were modeled by a mathematical simulation, the results of which will be published in a separate work.33 Tables 2−4 summarize the purification and yield results for the systems. Many possible changes in the process design and column design can improve the overall yield and purity. For

purity (%)

flow rate

0 °C

10 °C

0 °C

10 °C


76 74 70

71 68

98.4 98.8 98.9

98.3 98.4

Table 4. Average Yield and Purity Results for Falling Film Crystallization Experiments with System III column temperature yield (%)

purity (%)

flow rate

0 °C

10 °C

0 °C

10 °C


74 73 72

68 69 69

99.2 99.2 99.4

99.1 99.3 99.4

example, employing multiple columns in series will have a significant effect on the final purity, although the yield might be sacrificed for purity in multicolumn, countercurrent, cascade arrangement, or many other feasible designs. The solvent selection and crystallization solution system design have a more significant effect on the final purity, due to manipulating the crystal growth rate, where at higher growth rate the impurity inclusion is higher (i.e., comparing systems II and III). Figure 10 shows a final deposited layer of the API (in this case, acetaminophen), which was sleeved off of the column for imaging purposes. The overall view of the deposited layer (Figure 10A) shows the integrity and uniformity of the layer, where the crystals are uniform in size and distribution in the axial direction. The radial crystal size distribution profile consists of smaller crystals in the inner layer, where the fine crystals form on the surface of the cold column from nucleation on the rough surface, and larger crystals form in the outer layer. At the start of the experiments, when the supersaturation ratio is higher, crystal nucleation dominates crystal growth, ultimately resulting in smaller crystals. The outer large crystals form from the growing smaller crystals over the underlying layer, and diffused crystals are on the top. The 3 mm layer thickness is used for this specific example and is proportional to the feed concentration and total amount of the feed solution. The processing time is another key factor for developing the thickness; this layer is harvested at the end of the experiment when the supersaturation ratio is low and no further precipitation is possible. 4.3. Solvent Inclusion. The formation of inclusions is related to the rapid crystal growth and uneven levels of supersaturation across the crystal interface. Therefore, the control of the crystal growth influences the inclusion of solvent G



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Figure 10. (A) Overall look of the surface of the layer and (B) cross-sectional view of the deposited layer.

Figure 11. Solvent inclusion (A) in the crystal lattice and (B) between diffused crystals.

and impurity. The amount of impurity inclusion was previously reported as the form of final purity. Solvent inclusion is another incorporation present in macroscopic pockets inside the crystal matrix or between diffused crystals. Figure 11 shows these two types of inclusion, where A is the solvent inclusion inside the crystals of fenofibrate from system II and B is the trapped solvent in the crystal cake between the diffusing crystals of acetaminophen from system I. The included solvent is not pure, and it would attain an impurity concentration because the impurity concentration at the interface layer is high. The impurity portion in the solvent inclusion is included in the total impurity content of the crystals, which was measured by HPLC. TGA was used to measure the amount of solvent inclusion, and the results are reported in Table 5. The highest inclusion was in system II, which had the fastest growth rate, with an average amount of 1.9%. Systems I and III had similar solvent inclusion. The variation between the different experiments was significant at 30%, which could be

the result of heterogeneity in solvent inclusion in the cake and between the crystals. From TGA, the solvent inclusion in the crystal matrix and in the cake can be differentiated. During the TGA analysis, the “free” solvent, which is entrapped in the cavities between diffused crystals, evaporates first at the boiling point of the solvent, while the solvent incorporated inside the crystal matrix can be released upon melting of the crystal. Therefore, two inflections are expected in the TGA graph; the first inflection around the boiling point of the solvent is associated with the solvent in the cake, and the second is for the within-matrix inclusion. The melting point of fenofibrate (80.5 (°C)) is close to the boiling points of ethyl acetate and ethanol (77, 78 (°C)). Therefore, it is not possible to differentiate the different types of solvent inclusion for systems II and III. However, acetaminophen has a melting point of 169 °C, which is far beyond the boiling point of the solvent (ethanol:water 82 (°C)). Therefore, the amount of solvent inclusion in the cake (between diffused crystals) could be measured by TGA, and the value was approximately 60% of the total 1.4 ± 0.5% solvent inclusion. The lowest solvent inclusion depends on whether the operating conditions allow a low crystal growth rate, rather than on the solute−solvent system possessing low crystal growth kinetics.

Table 5. TGA Results for Solvent Inclusion in the Different Systems system

API

I II

acetaminophen fenofibrate

III

fenofibrate

solvent ethanol:water ethyl acetate:ethanol ethanol

avg solvent inclusion (% mass)

5. CONCLUSIONS This study presents a new concept for filtration-free purification via crystallization using a dynamic (falling) film solution layer crystallization approach. The technique was tested with three

1.7 ± 0.5 2.1 ± 0.7 1.6 ± 0.4 H



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kg.kg−1 for this system. The slopes of the solubility curves for systems I and III are relatively much smoother, but the supersaturation ratios are still high. The supersaturation ratio for system I is around 7.5 with a concentration difference, ΔC, of 0.52 kg.kg−1; and that for system III is aproximately 17 with a concentration difference, ΔC, of 0.66 kg.kg−1. However, the supersaturation ratios and the concentration differences can be controlled by the design of the crystallization solution systems and the selection of solvents and solvent mixtures. An important factor for the purification by crystallization is the solubility of impurities in the working solvent. Figure A.2

different crystallization systems with different growth rates and distribution coefficients. This falling film solution layer crystallization operation combines the functions of crystallization, isolation, washing, resuspension, dilution, and the slurry and powder transport operations between them into a single unit operation. Thus, this operation provides the possibility to eliminate the isolation of solids or transfer of slurries from continuous pharmaceutical manufacturing. If combined with powder-free formulation operations, there is potential to conduct fully continuous manufacture of drug products without the need to transfer or process slurries or powders within the manufacturing process. In continuous production, this could be extremely advantageous because the fouling and blocking of transfer lines is one of the biggest concerns in process robustness of continuous processes. Depending on the crystallization solution systems (growth rates at different supersaturation levels), the purification performance of the technique is significant. The comparison between systems II and III shows the effect of solvent selection, where the growth rate (and distribution coefficient) can be manipulated with a variety of solvents and mixtures. The results for the different temperatures and flow rates show no major effect on final purity, although the final purity is significantly improved compared with the initial amount of impurity in the solution. It can be concluded that the main parameters for purification performance of the system are the design of the system (geometry and size of the column) and growth rate of the particular crystallization solution system (supersturation degree). Although the final purity for a single stage/pass process would not fulfill required purity for drug formulation, utilizing further stages and serial columns can technically further improve the final purity. For instance, the results for one of the experiments (acetaminophen with metacetamol in a ethanol:water solvent at 30 mL/min flow rate and column temperature of 0 °C) with initial purity of 95% show the yield of 68% product recovery, and final purity of 97.1%.

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Figure A.2. Solubility of the impurities in the solvents of the different systems.

shows the solubility of the impurities in the solvents for the different systems. If the solubility of an impurity is low in the system, the cooling crystallization for purification will be challenging because, when reducing the solution temperature, the impurity starts to crystallize first or forms co-crystals with the main compound. The selection of proper solvent, working temperature, and maximum amount of impurity is crucial to avoid this issue. As Figure A.2 represents, the solubility of metacetamol in the solvent is very similar to that of acetaminophen. This similia similibus solvuntur, which is demonstrated in Figure A.3, mostly depends on the structure and molecular weight of the impurity.

APPENDIX A

A.1. Solubilities Results

Figure A.1 shows the solubility profiles for the systems. The solubility curve for system II (fenofibrate in ethyl acetate:ethanol) is very steep; however, the supersaturation ratio is approximately 8, with a concentration difference, ΔC, of 2.19

Figure A.1. Solubility curves for the APIs in the corresponding solvents (System 1, acetaminophen in ethanol:water; System II, fenofibrate in ethyl acetate:ethanol; System III, fenofibrate in ethanol).

Figure A.3. Solubility ratio of metacetamol to acetaminophen. I



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A.2. Distribution Coefficients

The process of capture of impurities in crystals during their growth from the liquid phase could be affected by several factors, such as the chemical composition of the liquid phase, relative solubilities of the host and impurity phases, interactions between host and impurity molecules, relative dimensions of substituting (impurity phase) and substituted (host phase) ions/molecules, similarity in crystallographic structure of the two phases, and crystallization conditions. When impurities incorporate at small levels into the crystal, the incorporation can be characterized by a distribution coefficient, which is defined as the ratio of the impurity concentration to the host compound concentration in the solid phase divided by that ratio in the liquid phase. The following equation defines the distribution coefficient: Distribution Coefficient =

(C imp/C H)solid (C imp/C H)liquid

Figure A.4. Equilibrium distribution coefficients for the systems (System 1, acetaminophen in ethanol:water; System II, fenofibrate in ethyl acetate:ethanol; System III, fenofibrate in ethanol).

(1)

where Cimp is the concentration of the impurity and CH is the concentration of the host compound. Distribution coefficients for Cimp were calculated by applying eq 1 to the data obtained experimentally. There are mainly two types of distribution coefficient, “equilibrium” and “effective”. The equilibrium distribution coefficient of an impurity in a solute is a function of the solution thermodynamics and solid-state thermodynamics. The effective distribution coefficient is manipulated by the kinetics and process parameters, such as the cooling rate, mixing, local concentration of the impurity, and crystal growth rate. In the falling film crystallization, the system is not in equilibrium, nor is the distribution coefficient. The effective conditions that result in the inclusion of more impurities in the crystals in the form of surface impurities, lattice impurities, and solvent inclusion increase the distribution coefficient to the effective level. It is difficult to measure the effective distribution coefficient separately for a continuous large-scale system, and mimicking the actual process conditions in the small batch vials is almost unfeasible. It would be possible to measure the purity of the deposited layer from each experiment (at a different flow rate, temperature, and impurity level) and extrapolate the effective distribution coefficient from an equilibrium basis by the actual process parameters. A plot of the equilibrium distribution coefficients as a function of the purity of the initial solution is shown in Figure A.4. The equilibrium distribution coefficient is the ideal purification limit for the falling film process, where the crystallization process would occur in equilibrium conditions at which the growth rate is slow and impurity inclusion is minimal. The ultimate goal of any purification process, including this falling film, would be to maximize the purity to an extent that is very close to the equilibrium distribution coefficient. Therefore, as a benchmark for purification, the design of the system and process parameters should be optimized to narrow the gap. The difference in the order of the distribution coefficients for systems I, II, and III is due to the molecular structure, activity group, and molecular size and weight similarities between the main compound (host molecule) and the impurity. Thus, if the impurity atoms are similar in size and in chemical properties of the atoms of the host crystal lattice, the impurity atoms will replace the host atoms at the lattice sites without creating much disturbance or strain in the lattice. Metacetamol is a structurally

related impurity in acetaminophen and has the same molecular weight and functional groups. The only difference is the position of the −OH group, which is on the fourth position in acetaminophen (4-acetamidophenol) and on the third position in metacetamol (3-acetamidophenol). The difference between fenofibrate and fenofibric acid is more significant in the differences in functional groups (ether vs ester), acidic groups, the existence of H-donor/-acceptor molecules, and molecular weight. Therefore, the equilibrium distribution coefficient for system I is five times higher than those of systems II and III. The solubility of both impurity and solute can change enormously with temperature. Therefore, the relative solubility could be variable over a range of temperatures. Consequently, in growth from solutions, temperature can play an even greater role in impurity incorporation by changing the solution thermodynamics. These relationships generally show that the purity of a material is improved by crystallization from solutions composed of solvents in which impurities are more soluble relative to the product material in the temperature range of crystallization. Figure A.5 shows the equilibrium distribution coefficients (in the form of purity) for system I over a range of processing temperatures for different initial impurity contents

Figure A.5. Purity of acetaminophen from system I (acetaminophen and metacetamon in ethanol:water) based on the equilibrium distribution coefficient for different initial impurity contents for a range of temperatures. J



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of 7%, 5%, 3%, and 2%. The segregation ratio of the impurity and the main compound is the same, based on the equilibrium distribution coefficient, and is proportional to the different initial impurity contents. However, the effect of temperature and relative solubility on the slightly variable distribution coefficient is noticeable. For system I, the distribution coefficient over temperature profile has a convex shape (Figure A.5) with a maximum point at approximately 25 °C. This behavior is similar to the relative solubility profile (Figure A.3), in which the solubility of metacetamol compared to acetaminophen is at its maximum level of approximately 25 °C. The trend suggests that, if the crystallization experiments were performed at 25 °C, the theoretical purification level would be optimum. However, the other process variables, such as the cooling rate, mixing, and supersaturation ratio, also have significant effects on the final purity.

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ASSOCIATED CONTENT

S Supporting Information *

The Supporting Information is available free of charge on the ACS Publications website at DOI: 10.1021/acs.cgd.5b01231. Schematic diagram of the falling film methodology, experimental falling film setup, and schematic diagram of the liquid distributor (PDF)

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AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS We would like to acknowledge Novartis International AG for its generous funding of this research. REFERENCES

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Novel Technique for Filtration Avoidance in Continuous Crystallization

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