Use of Continuous MSMPR Crystallization with Integrated

Article pubs.acs.org/crystal

Use of Continuous MSMPR Crystallization with Integrated Nanofiltration Membrane Recycle for Enhanced Yield and Purity in API Crystallization Steven Ferguson,† Franziska Ortner,†,§ Justin Quon,† Ludmila Peeva,‡ Andrew Livingston,‡ Bernhardt L. Trout,† and Allan S. Myerson*,† †

Novartis-MIT Center for Continuous Manufacturing & Department of Chemical Engineering, Massachusetts Institute of Technology, 77 Massachusetts Avenue, 66-568, Cambridge, 02139 Massachusetts, United States ‡ Department of Chemical Engineering, Imperial College London, Exhibition Road, London SW7 2AZ, U.K. § Centre for Life and Food Sciences Weihenstephan, TU Munich, Alte Akademie 8, 85354 Freising, Bavaria, Germany ABSTRACT: If continuous processing is to be employed in pharmaceutical production, it is essential that continuous crystallization techniques can meet the purity and yield achievable in current batch crystallization processes. Recycling of mother liquor in steady state MSMPR crystallizations allows the yield in the equivalent equilibrium batch process to be met or exceeded. However, the extent to which yield can be increased is limited by the buildup of impurities within the system. In this study, an organic solvent nanofiltration membrane was used to preferentially concentrate an API (deferasirox, M.W. = 373 Da) and purge the limiting impurity 4-hydrazinobenzoic acid (MW = 152 Da) from the mother liquor recycle stream in a mixed solvent (THF:ethanol) antisolvent (water) system. Incorporation of the membrane recycle allowed yields of 98.0% and 98.7% to be achieved. This compares to the following: a control MSMPR run without a membrane (70.3%), an equivalent batch process (89.2%), and the current commercial batch process (92%). Comparable product impurity levels were measured for the following: the MSMPR membrane recycle experiments (0.15 ppm and 0.22 ppm), the MSMPR control (0.13 ppm), and batch (0.32 ppm) control experiments. All processes met the regulatory specifications of a maximum of 3 ppm of the impurity 4-hydrainobenzoic acid.

1. INTRODUCTION Continuous manufacturing has long been used to produce a wide variety of commodity chemicals in a highly cost-effective and reproducible manner. In recent years there has been a concerted effort to both develop new continuous manufacturing technologies and to apply existing ones to pharmaceutical production. This movement is primarily driven by the potential to reduce production costs, improve product quality, and increase process safety. This renewed interest in continuous manufacturing has resulted in significant advances in many continuous operations aimed at pharmaceutical scales of operation including flow chemistry,1−5 extractions,6 and formulation.7 The first end-to-end continuous production run of an API (Aliskerin) from an intermediate right through formulation to final tabletting was demonstrated in 2012. Crystallization is no exception in this regard. 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 on pharmaceutical scales of operation. Crystallization is a key unit operation in the pharmaceutical industry with approximately 90% of small molecule active pharmaceutical ingredient (API) manufacturing processes containing a crystallization step.8 Crystallization serves as a cost-effective purification/separation operation, ensuring API purity specifi© 2013 American Chemical Society

cations are met as well as providing suitable API solid phase properties for formulation. Under normal conditions API losses following the final crystallization step are small. As such the yield for the final crystallization impacts the efficiency of the entire API processes. Therefore it is critical that the final crystallization operation in any continuous production scheme be able to meet that of the equilibrium batch crystallization. This is not trivial because continuous processes must operate with some supersaturation in the outlet at steady state and hence have a lower single pass yield.8 Several strategies can be used to ensure that continuous crystallizations meet the yield of the equivalent batch process. Similar to batch crystallizations, tubular crystallizations pass through a continuum of supersaturation values and approach equilibrium. The dynamics are identical except the crystallization progresses with respect to spatial position (residence time) within a tubular crystallizer rather than batch time.9 This can be seen in model organic and API systems antisolvent/ reactive crystallizations in standard tubular reactors with premixing.10−12 More novel configurations have been developed including segmented flow tubular reactors,13,14 continuReceived: October 8, 2013 Revised: November 26, 2013 Published: December 5, 2013 617

dx.doi.org/10.1021/cg401491y | Cryst. Growth Des. 2014, 14, 617−627

Crystal Growth & Design

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ous oscillatory baffled tubular crystallizers,15 Couette-Taylor reactors,16 and tubular crystallizers with static mixers.17 Each of these configuration aims to enable good mixing and narrow residence time distributions at low flow rates by imposing forced mixing through mechanical agitation or use of baffles. This provides additional residence time to reach equilibrium, enabling their use with compounds with slower kinetics. MSMPR crystallizers and other continuous stirred tank crystallizers operate at a fixed point on the phase diagram with the same conditions with respect to both spatial position and time.9 By definition these configurations must operate at a steady state with at least some remaining supersaturation in the product stream and hence some reduction in yield.8 This reduction in yield can be eliminated by recycling of concentrated mother liquor back into the crystallizer.18 In fact, the only constraint on the attainable yield with this method is the purification requirement of the system. A purge stream must be used to limit the steady state concentration of impurities that would otherwise build up inside the vessel or crystallize out in the solid product resulting in no purification (with 100% yield). The relative flow rates of the purge stream to the recycle stream (Recycle Ratio) determine the yield purity relationship for a given system and set of conditions within the crystallizer. As the recycle ratio is increased, a larger vessel volume is required to avoid reduction in productivity. In many cases, this strategy enables the equivalent yield and purity of a batch crystallization to be met, particularly where the steady state conditions (concentration/supersaturation/temperature, etc.) are within a kinetically favorable region for a given crystallization.18 Traditionally mother liquor concentration has been performed using vacuum evaporation. This method was demonstrated for the purification of cyclosporine (an immunosuppressant API) in a single stage MSMPR with mother liquor recycle.18 Concentration via evaporation has two major weaknesses. First, the rate of evaporation may be insufficient due to limits on evaporator temperature for heat sensitive APIs, especially with high boiling point solvents. Second, evaporative concentration of mother liquors cannot be used in mixed solvent systems or antisolvent systems without changing the makeup of the solvent or solvent/antisolvent ratio significantly. These problems can be addressed by the incorporation of an organic solvent nanofiltration (OSN) membrane into the process. OSN is a pressure-driven membrane based technique with separation capabilities in the molecular weight range between 200 and 2000 Da.19,20 Since OSN is not a temperature driven separation, there is no risk of API degradation, and for many membranes there should be limited or no effect on solvent/antisolvent concentrations. In systems with large differences in molecular weight between the API and the main impurity of interest, it may also be possible to utilize OSN membranes to preferentially purge impurities out of the system while simultaneously concentrating the mother liquor. For systems in which the impurity has a lower molecular weight than the API, membrane separation can be used to preferentially retain the API relative to the impurity. The purified and concentrated retentate stream can then be recycled back to the crystallizer. In cases where the impurity has a larger molecular weight than the API, the permeate stream can be recycled instead of the retentate. The purification of the membrane only has to be sufficient to prevent the buildup of impurities within the crystallizer. Therefore, a relatively modest difference in rejection factors

between the API and impurity should make a big difference in terms of attainable process yield. The combined approach should yield better results than with either the crystallization or membrane purification individually. Purification of a pharmaceutical intermediate using an OSN membrane was previously demonstrated by Sereewatthanawut et al.,21 for separation of an intermediate of a new drug candidate (MW = 675 Da) from its oligomeric impurities (MW ≥ 1000 Da). A model purification was also conducted using a blue (826.0 g mol−1)/yellow (198.2 g mol−1) dye based test system. In these cases solution cycling through a membrane module was the sole purification technique. In this study the API deferasirox (M.W. 373.37 Da) was used as a test system for the combined MSMPR with integrated OSN membrane recycle. The main impurity was 4-hydrazinobenzoic acid (4HBA, M.W. = 152.15 g mol−1), an unreacted reagent from API synthesis. 4-HBA must be reduced to 3 ppm to meet production specifications. The chemical structures for the two compounds are shown in Figure 1. These compounds provide a

Figure 1. Chemical structures of deferasirox (A) and 4-HBA (B).

good test for the crystallization system due to the 221.22 Da difference in molecular weights between API and impurity, which would not be uncommon in the case of unreacted species from API synthesis or degradation products. Furthermore, this process provides a proof of concept for use of concentrated mother liquor recycle in a mixed solvent/ antisolvent system. In this study a batch membrane screen was performed to assess the performance of a variety of OSN membranes for the deferasirox, 4-HBA system. The most promising membranes were then used in an integrated MSMPR with mother liquor recycle. Once the membrane retention factors and permeabilities were obtained a combined mass and population balance model of the whole system was used to generate the design space for the process in order to illustrate a methodology to quickly identify optimum operating regions, with a reduced number of experiments and material required.

2. EXPERIMENTAL SECTION 2.1. Materials. All of the experiments were carried out with the chemicals listed in Table 1. The crude refers to deferasirox obtained from a crude commercial crystallization which contains 4-HBA at a concentration 19.2 ppm. The “pure” deferasirox was obtained from a subsequent commercial purification process involving charcoal treatment followed by an antisolvent recrystallization that was recreated in this study. This commercial process had a yield of 92% with 0.3 ppm of 4-HBA. 2.2. Organic Solvent Nanofiltration. There are several criteria that describe membrane performance. Permeability (P) is a function of permeate volumetric flow rate (QP), membrane area (AM), and pressure (Δp): 618



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Table 1. Chemicalsa

a

chemical

description

manufacturer

ethanol water tetrahydrofuran charcoal 4-hydrazinobenzoic acid deferasirox deferasirox deferasirox

200 proof, anhydrous Chromasolv Plus, for HPLC anhydrous, inhibitor-free Norit SX-2

Koptec Sigma-Aldrich Sigma-Aldrich EMD Millipore Alfa Aesar Novartis Novartis Prepared

pure crude feed

purity 99.50% ≥99.9% 97% 0.3 ppma 19.2 ppma 5.9 ppma

Parts per million of 4-HBA.

Table 2. Properties of Screened OSN Membranes name

MWCO, Da

polymer

cross-linked

24-P84-1:1-PP 24-P84-1:2-PP 24-P84-1:3-PP 24-P84-1:1-PP-X 24-P84-1:2-PP-X 24-P84-1:3-PP-X 26-PBI-PP-X 23-PBI-PP 26-PBI-PP 24-PBI-PE 12-PEEK4000-PP

∼450 ∼400 ∼300 ∼350 ∼300 ∼200 250−300 >1500 >1500 300−400 ∼390

polyimide

no

moderately hydrophobic

yes

less hydrophobic (semihydrophilic)

yes no

rather hydrophilic

P=

polybenzimidazole

moderately hydrophilic polyether ether ketone

QP AM Δp

C P,i C M,i

no

hydrophobic

in this study. Properties of the respective membranes are summarized in Table 2. Membranes code: the numbers and letters in the membrane code indicate the following: wt % of polymer in the membrane dope composition − type of polymer − solvent/cosolvent ratio − type of backing material − cross-linked (X)/noncross-linked (-). The MWCO was determined using standard solution of polystyrenes22 in acetone at 30 bar pressure and 30 °C, except for the PEEK membrane (12PEEK4000-PP) where MWCO was determined in DMF at 80 °C The filtration unit integrated into this study has been described in detail in two previous studies.24,25 To summarize the description, the filter module functions as a cross-flow filter with feed entering parallel to the membrane surface and retentate also exiting parallel to the membrane surface. It is slightly atypical of most cross-flow filtration devices in that the retentate channel is broad and is able to accommodate a magnetic stir bar. 2.2.1. Membrane Conditioning. Before each run, membranes were conditioned with solvent (THF:ethanol: H2O = 2.25:1:3.25 (w/w/w)) at 44 °C for a minimum of 12 h. The flow rate at the membrane inlet was adjusted to match that of the subsequent experiment. 2.2.2. Membrane Screening. The membranes listed in Table 2 were screened for their permeability and rejection properties. Charcoal treated solutions of 0.7% (w/w) crude deferasirox in THF:ethanol:H2O = 2.25:1:3.25 (w/w/w) (based on solution mass) were used. The solutions were spiked with additional 4-HBA to bring the concentration to 30 ppm in order to raise the 4-HBA concentration to detectable levels in the liquid phase samples. Temperature was maintained at 44 °C using a temperature controlled hot plate, and a stirring speed of 600 rpm was used. Each membrane was examined at three pressures ranging between 10 and 60 bar. Pressure levels were chosen individually for each membrane in order to cover a maximum operating range of pressures. API solution was recycled for 2.5 h at a constant pressure before measuring the permeate and retentate volumetric flow rates and taking samples. Deferasirox and 4-HBA concentrations were determined by HPLC. Permeability and rejection factors for deferasirox and 4-HBA were calculated from the flow rate and concentration data using eqs 1 and 2. 2.3. Crystallization Experiments. 2.3.1. Batch Crystallization Experiment. In order to provide a basis for comparison for the

(1)

Rejection is a measure of membrane separation performance and describes the percentage of a solute which is unable to pass the membrane. The rejection factor RRi is the rejection of solute i (i = A in for the API deferasirox, i = I in for the impurity 4-HBA) and is calculated according to eq 2, where CM.i is the concentration of solute i in the feed entering the membrane, and CP,i is the concentration of solute i in the permeate.19 RR i = 1 −

further properties

(2)

Another performance criterion is the molecular weight cutoff (MWCO), defined as the molecular weight at which 90% of the respective solute is retained by the membrane. This criterion is usually provided by the manufacturer to give an indication about the operating range of the membrane.19 The membranes used in the batch screening process (Section 3.1) vary in their polymer composition, manufacturing treatment (cross-linked/noncross-linked), and pore size. Six of the tested membranes (P84 series) are composed of polyimide (Lenzing P84 from HP Polymer GmbH, Austria). Their MWCO increases with decreasing solvent: cosolvent ratio (dimethylformamide: 1,4-dioxane). The asymmetric membranes are manufactured on a polypropylene (PP) backing. Three of them are noncrosslinked (PP series), whereas three are cross-linked with 1,6-hexanediamine (PP-X series).22 Four membranes consist of polybenzimidazole (PBI, CelazoleS26 polybenzimidazole (PBI) solution from PBI Performance Products Inc.), which is slightly more hydrophilic than polyimide. The 26-PBI-PP-X membrane is cross-linked with α,α′dibromo-p-xylene, whereas the PBI(-PE/-PP) series is noncrosslinked. 23-PBI-PP and 26-PBI-PP are manufactured on a polypropylene backing, whereas the backing of 24-PBI-PE is polyester. The PBI membranes were prepared using the procedure described elsewhere.23 The last membrane consists of polyether ether ketone (PEEK, Vestakeep 4000P, Evonik Degussa GmbH, Germany). It is noncross-linked and has a MWCO of approximately 390 Da.24 All membranes were expected to display high thermal and chemical stability concerning the experimental conditions and solvents required 619



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mL/min and thus QM = QR in eq 3). The apparatus is depicted in Figure 3. This experiment was operated without concentration of the recycle stream and improves yield over a single pass MSMPR by increasing the effective liquid phase residence time.

MSMPR with integrated OSN recycle, the equivalent batch experiment was performed in the same 155 mL vessel. A solution with identical composition to the MSMPR feed of 18.6% (w/w) of charcoal treated solution of crude deferasirox in THF:ethanol = 2.25:1 (w/w) at a temperature of 20 °C was added to the vessel. A 1 h linear addition of water (antisolvent) was conducted so that the final vessel volume of 155 mL and a solvent:antisolvent ratio of 1:1 (w/w) were obtained. Focused beam reflectance measurement (FBRM) was used in order to determine when the equilibrium chord length distribution was reached. Solid and liquid phase samples were taken at the end of the process. The solid phase (crystals) was washed with 60 mL of water and vacuum-dried. Concentrations of deferasirox and 4-HBA of both the liquid and solid samples were analyzed by HPLC. Concentration of deferasirox in the crystallizer after 125 min was determined to be 1.1% (w/w), whereas solubility of deferasirox at 20 °C and a solvent:antisolvent ratio of 1:1 (w/w) is 0.94% (w/w). This resulted in a yield calculation of 89.2%. This result is slightly below the expectations given in the Novartis process (92%); the discrepancy is most likely due to differences in the final isolation temperature between the commercial process and the batch process here (measured concentration exceeded solubility). Concentration of 4HBA in the crystals was 0.32 ppm. 2.3.2. Continuous Crystallization Experiments with/without Membrane. All continuous crystallization experiments using this combined MSMPR with integrated OSN for concentrated mother liquor recycle can be described in terms of the general flow sheet depicted in Figure 2.

Figure 3. Flowchart of experiment C1: single-step MSMPR with recycle, without membrane. Experiments C2 and C3 differed only in the OSN membrane types which were put into the membrane cell (24-P84-1:2-PP-X in C2 and 24-P84-1:3-PP-X in C3). The only waste stream of this configuration was the permeate stream, with split S = 0. Figure 4 shows the corresponding flowchart for these experiments.

Figure 4. Flowchart of experiments C2 and C3: single-step MSMPR with recycle and integrated membrane.

Figure 2. General setup, basis for crystallization experiments, and kinetic modeling.

Start Up Methodology. For experiments C2 and C3, the respective membrane was initiated overnight as described in Section 2.2.1. Each crystallization was started as a fed batch, as per Section 2.3.1. Before starting continuous operation, the membrane cell (∼100 mL, ∼51 cm2 effective membrane area)24 was prefilled with charcoal treated solution of 1% (w/w) crude deferasirox in THF:ethanol:water = 2.25:1:3.25 (w/w/w). As sampling over the whole duration of the experiment lead to a significant reduction in volume in the system (approximately 60 mL), the filtrate hold vessel was prefilled with solution of the same concentration as the membrane cell. Continuous Operation. All continuous crystallization experiments were conducted as a single-stage MSMPR with recycle of the mother liquor. Each crystallization was performed in a 155 mL stirred, water jacketed reaction vessel with a residence time of 1 h, temperature of 20 °C, and a stir speed of 300 rpm. The feed solution of 18.6% charcoal treated deferasirox in THF:ethanol = 2.25: 1 (w/w) was pumped into the crystallizer at a flow rate of 0.16 mL/min. Water (antisolvent) at room temperature was added such that the MSMPR was operated at a solvent:antisolvent ratio of 1:1 (w/w). To improve the accuracy of the addition, both streams were pumped intermittently for 10 s per minute. Slurry was removed intermittently so that 10% of the vessel volume was removed rapidly, and the outlet tube was kept clear by pumping remaining slurry out using air transferred once the solution level had dropped below the level of the outlet dip tube. This is similar to a method previously demonstrated by Ferguson et al.,9 except in this case a peristaltic pump was used to transfer the fluid and air in the

In the general setup in Figure 2 the OSN membrane cell was included after the filter, which separated the solids from the mother liquor. The membrane then divided the mother liquor into a retentate and permeate stream in which the retentate of the API was concentrated and the impurity was depleted. The retentate was then recycled back into the crystallizer, while the permeate was discarded. In cases where the API preferential permeates the membrane, permeate would be recycled with the retentate discarded.21 In cases where the membrane purification is not sufficient to prevent the buildup of impurities in the system, a purge stream can be utilized to allow purification specifications to be met at the expense of yield. The volumetric flow rate of the purge stream (QW) is determined by the split (S), being defined as

S=

QW (Q M + Q W )

(3)

.where QM is the volumetric flow rate entering the membrane device, thus QM+QW is the volumetric flow rate of mother liquor obtained as filtrate in the filter unit. For a given crystallization, the split also sets the permeate volumetric flow rate. In total three continuous crystallization experiments were conducted with two different configurations. Experiment C1 was a control experiment conducted without a membrane. It was derived from the general configuration by setting the split S = 0.1 without membrane (i.e., permeate flow QP = 0 620



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place of vacuum/air lines. The slurry was sent to an inline filter unit to extract crystals from the mother liquor. The filtrate (mother liquor) was gathered in a holding vessel maintained at 40 °C. To avoid clogging of the lines, the feed line as well as the crystallizer outlet line leading to the filter unit were heated to 40 °C. In experiment C1, 90% (v/v) of the mother liquor was directly recycled into the crystallizer with remaining 10% (v/v) purged. In experiments C2 and C3, mother liquor entered the membrane cell containing the respective membranes as per Figure 4. 90% (v/v) of the inlet stream was retained and recycled into the crystallizer, whereas 10% (v/v) was purged through the permeate stream. The permeate flux was regulated with a backpressure regulator. The membrane cell was heated at 44 °C. In order to reach steady state, all experiments were run for 10 residence times (10 h). Chord length distribution of the crystals in the crystallizer was measured in-line by FBRM (Focused Beam Reflectance Measurement) over the entire duration of the experiment. Samples of the permeate, the retentate, and the liquid phase of the crystallizer (separated from solid phase by filtration) were taken after each residence time (1 h), and their concentrations were measured by HPLC. The filter cake (crystal phase) was removed after each residence time, washed with 60 mL of water, and vacuum-dried. The yield (Y) for each process was calculated based on the loss of API in the permeate and purge streams (where present) compared to the amount of API entering the system from the feed.

Ṁ PC P,A + Ṁ W C W,A Y=1− Ṁ FC F,A

Q IN = Q F + Q AS + Q R

C IN,A =

⎛ CSS,A − Csat,A ⎞b ⎟⎟ B = kB⎜⎜ Csat,A ⎝ ⎠

(12)

⎛ CSS,A − Csat,A ⎞ g ⎟⎟ G = k G⎜⎜ Csat,A ⎝ ⎠

(13)

mC,I

p=

mC,A CIN,I CIN,A

(14)

mC,I and mC,A represent the mass fractions of impurity and API in the crystal, and CIN,I and CIN,A are the concentrations of dissolved impurity and API at the crystallizer inlet (in % w/w). This assumption was used because the impurity concentration in the liquid phase was below the detection limit of the HPLC. More accurate prediction can be obtained by developing an equilibrium distribution coefficient model as per Alvarez et al.27 Measurement of the equilibrium distribution coefficient is straightforward and provides a method of calculating the impurity content in the solid phase. This assumes purity decreases with yield due to the relative increase in impurity concentration in the liquid phase but does not account for crystallization rate, which often significantly impacts on process purity. Hence this method gives an idealized estimate of the purification potential for a process. Dynamic equilibrium coefficients for organic solution crystallizations are rarely implemented but would be of great use in predicting process performance of continuous crystallizations. Given the kinetic parameters for crystallization, the distribution coefficient, membrane rejection factors and permeabilities, the use of the MSMPR model, and a mass balance for the combined system allows the development of a design space for the system from which optimal operating conditions can then be chosen.

For the boundary condition n(0) = n this can be solved analytically resulting in eq 6 (6)

where L is the characteristic size of the crystal, n(L) is the number density of crystals per unit volume of size between L and L + dL, G is the size-independent growth rate, and τ is the residence time in the crystallizer.26 n0 describes the number density of zero sized particles n(0) and is linked to the nucleation rate B: (7)

A mass balance on the crystallizer is as per eq 8

Q OUTM T = Q INC IN,A − Q OUTCSS,A

(11)

The kinetic parameters kB, b, kG, and g can be estimated by fitting the model to the CSD from an MSMPR type experiment. The parameters are fitted as per the membrane crystallization experiments. However, they could be determined from a standard MSMPR experiment to assess whether an OSN recycle is a good strategy in conjunction with permeability and rejection factor data from membrane batch screening. In this study cube weighted FBRM chord length distributions normalized using measured solid phase suspension density were used to fit the kinetic parameters. CLDs do not give an accurate measure of full CSD and were utilized here to illustrate an approach for design rather than to accurately assess the kinetics for this industrial process. For prediction of the impurity level (4-HBA) in the crystals, a constant purification ratio p based on the experimental results was assumed:

(5)

B G

(10)

where ρc is the solid density, and kv is the volume shape factor. Crystal nucleation and growth rates were expressed as a function of supersaturation employing empirical power-law equations. CSS,A and Csat,A are the steady state and saturation concentrations in the crystallizer, kB, b (eq 12), kG, and g (eq 13) are kinetic parameters for nucleation.

0

n0 =

Q F + Q AS + Q R

M T = 6ρc k vn0(Gτ )4

Ṁ P, Ṁ W, and Ṁ F and CP,A, CW,A, and CF,A are the mass flow rates and the corresponding concentrations of deferasirox in % w/w of permeate, purge, and feed, respectively. Impurity levels of 4-HBA in the crystals were measured via HPLC from the solid phase samples. 2.4. Process Modeling. A simple population balance model combined with a mass balance for the system was utilized in order to generate a design space to illustrate their usefulness in the assessment and design of continuous MSPMPR crystallizations with concentrated mother liquor recycle. Such an approach can lead to a huge reduction in the required number of experiments and is an excellent way to assess the viability of the integrated recycle strategy for a given compound. The crystallization is modeled using the MSMPR model. The MSMPR model makes the assumption that the vessel is perfectly mixed; withdrawal is representative. Assuming size independent growth, the population balance equation is represented by the ODE in eq 4.

n(L) = n0e−L /(Gτ)

Q FC F,A + Q R C R,A

The suspension density can be calculated from the third moment of the population density25

(4)

dn n =− dL Gτ

(9)

(8)

with QOUT being the volumetric stream leaving the crystallizer, CSS,A being the outlet (= steady state) concentration of deferasirox, and MT being the suspension density of crystals in the outlet stream. QIN and CIN,A in this context are the volumetric inlet stream and the inlet concentration of deferasirox resulting from lumping all inlet streams (feed stream QF, antisolvent stream QAS, and recycle stream QR) together:

3. RESULTS AND DISCUSSION 3.1. Membrane Screen. There are two objectives when performing OSN membrane screening. The first is to find a membrane with high API rejection and sufficient permeability 621



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to allow the required permeate flow rate to be obtained (given a specific membrane surface area). This in itself is significant, as it enables concentration of mother liquor for recycle in mixed solvent and antisolvent systems for the first time. As outlined within the introduction the only limitation that constrains yield in MSMPRs with mother liquor recycle is the purification specifications. The second objective of the screen is to find a membrane that has a significantly lower rejection factor with respect to the impurity than the API. The impurity will be preferentially purged from the system allowing greater concentration and higher yield at the same level of purification. As outlined in Section 2.4 the data from this screening process combined with crystallization kinetics for the system can quickly give a prediction of system behavior. A batch screening methodology may be preferable in pharmaceuticals as it minimizes the amount of material consumed. Long-term continuous membrane stability experiments would also be required before putting such systems into production to determine the membrane lifetime and the frequency of required back flushing/cleaning. In the screening experiments, permeability slightly decreases with increasing pressure for all membranes. No trends concerning the rejection of deferasirox or 4-HBA were recognized with changing pressure. Since performance is largely unaffected by pressure, mean permeabilities and rejection factors shown in Figure 5 were averaged across the full pressure range investigated for each membrane. The standard deviation error bars describe both experimental variability and variation between pressure levels for a given experiment. It has been shown previously that the pressure affects rejection, more considerably for the low molecular weight species22,23 which may explain much larger standard deviations for the 4-HBA rejection. The permeability of the cross-linked (PPX) and noncrosslinked (PP) polyimide membranes tested in Figure 5A correlate with molecular weight cut off (MWCO) for each subset of membranes (Table 2). A lower molecular weight cut off indicates a smaller nominal pore size, and so this result is expected. Decreasing MWCO is correlated with an increased ratio of solvent:cosolvent (used in membrane production) for both the cross-linked and noncross-linked polyimide membranes. At the same solvent:cosolvent ratio, the cross-linked membranes have a lower MWCO and also a lower permeability than the equivalent noncross-linked membrane. The 24-P841:2-PP and 24-P84-1:1-PP-X membranes had similar MWCO with the cross-linked one having lower permeability. This is not unexpected as separations with OSN membranes are primarily through steric effects, as such MWCO is not a definitive measure and emphasizes the need for screening. Rejection of 4HBA varied but was lower than rejection of deferasirox in most of the polyimide membranes (Figure 5B). Deferasirox rejection increased with decreasing MWCO (see Table 2). Since crosslinked membranes of the same solvent:cosolvent ratio had a higher MWCO than noncross-linked membranes, higher rejection of deferasirox with cross-linked membranes occurred as expected. The cross-linked polybenzimidazole membrane (26-PBI-PPX) displayed a lower permeability compared to equivalent cross-linked polyimide membrane (24-P84-1:3-PP-X). Rejection of deferasirox was in the same range (approximately 95%), whereas the rejection factor of 4-HBA was higher (72% compared to 27% for 24-P84-1:3-PP-X). The noncross-linked polybenzimidazole membranes (23-PBI-PP, 24-PBI-PE, and

Figure 5. Performance criteria of screened membranes (means with standard deviations), A: permeability, B: rejection factors of deferasirox and 4-HBA. Names of the first six membranes displayed were shortened (24-P84 polyimide series).

26-PBI-PP) displayed an extremely low permeability. Permeability was high when pressure was increased but decreased rapidly at the new constant pressures. Rejection factors for these membranes were mostly negative values. Since the negative rejections were observed only for the deferasirox this could also be due to a higher affinity of the solute to the membrane than the solvent. The negative rejection phenomenon is still not well understood and has been attributed to the high interaction of solute with the membrane material and consequently preferential transport when solvents show poor permeation properties. 28 The PEEK membrane (12PEEK4000-PP) also showed low permeability and was deemed unsuitable for use with the compound in this study. As stated, requirements of high rejection of deferasirox, low rejection of 4-HBA, and suitable permeability were sought. For the planned experimental setups (Figure 4) a permeate flow rate of 10% of the membrane inlet (0.26 mL/min) was set. For permeabilities between 0.1 and 0.4 L/(h m2 bar) with the membrane disc area of 51 cm2 this flow rate was reachable with pressures ranging from 30 to 5 bar. From inspection of the results the most suitable membranes were 24-P84-1:2-PP-X, 24-P84-1:3-PP-X, and 26-PBI-PP-X. The best purification performance was achieved with 24-P841:3-PP-X; however, variability of data for the rejection factor of 4-HBA was very high. The 24-P84-1:2-PP-X membrane showed 622



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Figure 4. In this section the performance of this system is compared in terms of yield and level of 4-HBA (impurity) to the equivalent batch process (Section 2.3.1), the commercial batch process (Section 2.1), and the equivalent MSMPR process without membrane. The results for the MSMPR recycle without membrane concentration experiment are shown in Figure 7. Since an

lower variability of rejection factors as well as higher permeability. Compared to the two polyimide membranes, retention of 4-HBA was higher with the polybenzimidazole membrane 26-PBI-PP-X. The two cross-linked polyimide membranes (24-P84-1:2-PP-X and 24-P84-1:2-PP-X) were thus selected for integration into the continuous crystallization process. 3.1.1. Stability Test. The membrane used in the continuous crystallization C2 (24-P84-1:2-PP-X) was tested over 24 h. It was conditioned as per section 2.2.1. After conditioning, screening solution (as per Section 2.2.2, was recycled at 44 °C for 24 h with a flow rate of 2.6 mL/min. Pressure was adjusted to that of the operating range of the continuous crystallization setup (7 bar for membrane 24-P84-1:2-PP-X) using the back pressure regulator on the membrane cell. Figure 6 shows the permeability and rejection factor of deferasirox and 4-HBA of membrane 24-P84-1:2-PP-X over the 24 h run.

Figure 7. Concentration of deferasirox in the liquid phase of the crystallizer, the recycle, and the purge stream over time for a continuous crystallization process without integrated membrane.

MSMPR operates with some supersaturation in the outlet line, recycling of mother liquor boosts yield by increasing the effective liquid phase residence time. The process parameters for this run were laid out in Figure 3; a 1 h residence time was utilized with a recycle ratio of 90% (v/v) of the mother liquor and hence 10% (v/v) purge. Figure 7 shows the liquid phase concentration of deferasirox in the crystallizer, the recycle, and the purge over 10 residence times. A final concentration of 3.4% (w/w) was reached in the crystallizer after approximately 7 residence times, while steady state was reached for the CSD after 8 residence times as observed via FBRM CLD mean (μm) and total counts (#/s). All three streams should reach the same concentration. The delay in the recycle and purge streams reaching the steady state concentration was due to the buffering effect of the volume of conditioning solution in the filtrate reservoir. The steady state yield of the process without membrane recycle was 70.3% with a 4-HBA concentration of 0.13 ppm in the solid product. 3.2.1. Continuous Crystallization with Integrated Membrane. As per Figure 4 the continuous crystallization processes were performed with the integrated membrane unit, and 90% v/v of the total mother liquor flow was recycled into the crystallizer with the remaining 10% purged from the system via the membrane permeate stream. Concentration profiles of deferasirox in the mother liquor, retentate, and permeate streams for both experiments are depicted in Figure 8. Both experiments were started from a batch crystallization and show an increase in concentration in the mother liquor and the retentate from the initial 1% w/w (batch equilibrium and prefill concentration) to approximately 3% w/w. Due to the high recycle ratio, the concentration of deferasirox in the retentate stream was marginally higher than that of the crystallizer. As with experiment C1 the buffer tank is responsible for the delay of the retentate reaching the same steady state concentration as the crystallizer. For both experiments, steady state was reached after 10 residence times. Attainment of steady state with respect to crystal size

Figure 6. Permeability and rejection of deferasirox and 4-HBA for 24P84-1:2-PP-X over 24 h.

The permeability and rejection of deferasirox were slightly lower in the long term study than in the batch membrane screen at 0.4 L/(h m2 bar) and 0.9, respectively, compared to the membrane screen (0.47 L/(h m2 bar) and 0.96 from Section 3.1). This level of variation is normal in OSN membranes and is thought to occur via small discrepancies in the preparation and manufacturing processes. Both performance characteristics did not show a significant trend over the 24 h duration of the run. Rejection of 4-HBA decreased considerably from initial 95% to 40% after 24 h. This is an interesting result suggesting that there could be some chemical interactions between the membrane and 4-HBA. One possible hypothesis is that since the P84 material is susceptible to reactions with amines there still may be some free reactive groups (due to incomplete cross-linking) which are able to react with 4-HBA at the beginning of the experiment and thus increase the rejection (chemical sorption). Once these groups were saturated the rejection would mainly be due to the sieving mechanism and is expected to decrease. This hypothesis could also explain the higher rejection of 4-HBA than that of deferasirox by the noncross-linked P84 membranes (Figure 5). This result indicated that better performance could have been achieved with prolonged steady state operation in the crystallization experiments. No degradation or fouling was observed over the course of the experiment. 3.2. Continuous Crystallization. The most promising membranes from the screening process (24-P84-1:2-PP-X and 24-P84-1:3-PP-X) were integrated into the continuous singlestage MSMPR process of deferasirox with recycle as shown in 623



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Figure 8. Concentrations of deferasirox in the liquid phase of the crystallizer, retentate, and permeate over 10 residence times, A: experiment C2, membrane 24-P84-1:2-PP-X integrated, B: experiment C3, membrane 24-P84-1:3-PP-X integrated.

without integrated membrane. A slightly higher yield combined with a higher level of impurity in experiment C3 was a result of higher retention of deferasirox and 4-HBA by 24-P84-1:3-PP-X compared to 24-P84-1:2-PP-X membranes. From these results it is clear that a significant improvement in process performance can be achieved by utilizing this combined crystallization nanofiltration purification and isolation strategy in terms of process yield. The impurity level in the crystals was slightly increased in the membrane runs compared to the MSMPR without membrane, due to some retention of 4-HBA by the membranes. The concentration of 4-HBA in the processes with membranes was still lower than the level achieved in the equivalent batch process. All processes comfortably met the production specification of 3 ppm of 4-HBA. These results are significant as meeting the yield/purity relationship of a well-designed batch crystallization that reaches thermodynamic equilibrium with a continuous steady state process can be quite challenging. Use of this combined MSMPR with integrated OSN membrane for mother liquor recycle could negate this issue where impurities have significant differences in molecular weight compared to the API/ compound of interest. The system highlighted in this study is reasonably typical of a degradation product or unreacted species from reactive synthesis and could be of use in many pharmaceutical purifications but will not be suitable for cases where structurally similar compounds have to be separated. The population balance model was fitted to the steady state chord length distributions obtained from the continuous crystallization experiments (C1−C3) after 10 residence times as described in Section 2.4. Experimental and model-predicted size distributions (volume based), as well as kinetic parameters, are depicted in Figure 10. The model distributions agree reasonably with the experimental data. Experimental distributions were broader and biased toward larger particle sizes. As outlined in the description of the population balance model (Section 2.4) the reason for deviations between experimental data and model prediction are likely due to the detection method (FBRM), which records chord lengths instead of particle sizes. This is

distribution in the crystallizer was observed from the FBRM square weighted mean and total counts as per Figure 9.

Figure 9. Total counts per second and mean chord length (square weighted) recorded by FBRM for a continuous crystallization with integrated membrane 24-P84-1:2-PP-X (experiment C2).

The permeate stream concentrations reached a final level of 0.23% and 0.15% w/w in experiments C2 and C3, respectively. Based on these permeate concentrations and the permeate streamflow rate of 0.26 mL/min, the yields for experiments C2 and C3 were calculated to be 98.0% and 98.7% w/w with 4HBA concentration in the crystals of 0.15 ppm and 0.22 ppm, respectively. From inspection of the results in Figure 7 and the membrane runs in Figure 8 it is easy to see that the high retention of API via the membrane drastically reduces the amount of API purged from the system thus increasing the yield. The concentration of 4-HBA in the liquid phase was below the detection limit of the HPLC but was detectable from the solid phase product samples. The purity results for the continuous crystallization runs with integrated membranes were 0.15 ppm and 0.22 ppm for the 24-P84-1:2-PP-X and 24P84-1:3-PP-X membranes, respectively. These results are shown in Table 3 alongside the purity values in the equivalent batch, commercial batch, and MSMPR with recycle with and

Table 3. Results of Crystallization Experiments (Batch, Continuous with/without Membrane, Novartis Process)

deferasirox, % w/w equilibrium/steady state purge/permeate yield 4-HBA in crystals, ppm

batch

MSMPR, no membrane

MSMPR with 24-P84-1:2-PP-X

MSMPR with 24-P84-1:3-PP-X

commercial batch

1.11 89.22% 0.32

3.39 3.40 70.29% 0.13

3.00 0.23 98.03% 0.15

3.03 0.15 98.71% 0.22

92% 0.3

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Figure 10. Experimental chord length distributions and model predicted crystal size distributions (volume-weighted) of experiments C1 (no membrane), C2 (24-P84-1:2-PP-X), and C3 (24-P84-1:3-PP-X), model predictions based on given kinetic parameters.

Figure 11. Performance of the continuous crystallization process for varying operating conditions. A: yield as a function of split and retention factor of deferasirox, B: impurity level (4-HBA) as a function split and rejection factor of 4-HBA, assuming a retention factor of 0.9 for deferasirox.

additional membrane purge stream. This may not be the case for all systems, and an additional purge stream may be required. Figure 11 shows the design space for the deferasirox system. Figure 11A shows yield as a function of both the split (S) between the purge stream and flow rate to the membrane and of the rejection factor of deferasirox. In Figure 11B the impurity level as a function of the split and of the rejection factor of 4HBA for a membrane with a rejection factor of 0.9 for deferasirox is shown. When taken together these plots allow determination of the attainable yield and purity via the integration of all membranes. This can also be solved as an optimization problem. Yield increases with rejection of deferasirox because less deferasirox is lost by permeation through the membrane. When a higher percentage is purged directly to further reduce impurity concentration, additional deferasirox is also removed, thus reducing the yield. If a membrane with a rejection of 0.9 is used as in Figure 11B, the purity at each split for all of the rejection factors for 4-HBA can be seen. Again as would be expected, an increasing split also increases the amount of purged 4-HBA and thus reduces 4-HBA concentration in the crystallizer, which leads to a lower impurity level in the crystal phase. 4-HBA concentration in the crystallizer and thus in the crystals is increased by higher retention of 4-HBA by the membrane. This can be useful in assessing what membrane performance is required before membrane screening in addition to predicting process performance using information identified by screening. More generally the results show that in cases where the purification of the membrane is more limited (i.e.

particularly true for needlelike particles formed in the crystallization of deferasirox. Microscope images of the particles indicate that breakage and agglomeration were not dominant processes and can be neglected in the case of this crystallization. Nonetheless the kinetic parameters obtained were typical of organic systems that crystallize at comparable rates. This approach is sufficient for the proof of concept experiments conducted in this study. However, when conducting a more comprehensive design for an industrial process, a more detailed investigation of kinetic parameters with a wider range of process supersaturations should be conducted. This will enable a more robust prediction of crystallization performance over a range of process conditions. The design spaces produced by the combined population and mass balance process model (Figure 11) illustrate an efficient way of assessing the applicability of this combined MSMPR membrane approach given just crystallization kinetics, crystallization distribution coefficients, and permeability and rejection factor data from batch screening. The kinetic parameters fitted in Figure 10 were employed to investigate the effects of varying operating conditions on process performance based on the general setup displayed in Figure 2. For a given crystallization, the design space shows the performance that can be achieved by integrating all possible membranes. In particular, the influence of an additional purge stream and membrane performance on process yield and impurity level was examined. In the case of the deferasirox process, purification requirements were easily met without an 625



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higher impurity rejection) introduction and adjustment of purge rate (split) is a powerful tool in controlling purity. The experimental cases described earlier (C2 and C3) are characterized by the lines when the split is zero. Yields of experiments C2 and C3 (98.0% and 98.7%, respectively) met the model predictions for membranes with a rejection of deferasirox between 0.9 and 0.95. The product impurity levels in these experiments (C2 and C3) were extremely low (0.15 ppm and 0.22 ppm), corresponding to rejection factors of 4HBA of 0.2 (C2) and 0.4 (C3) compared to factors of 0.44− 0.633 (24-P84-1:2-PP-X) and 0.39−4.61 (24-P84-1:3-PP-X) obtained from the batch screen. Deviations from the prediction are most likely due to decreasing rejection in terms of 4-HBA over time, as indicated in the membrane stability experiment (Figure 6). These values are reasonably accurate given the screening experiments are designed with more focus on speed. Membrane to membrane variation is common in terms of lab scale OSN membranes and would also introduce significant variability.

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4. CONCLUSION Integration of an organic solvent nanofiltration membrane into continuous MSMPR crystallization with process mother liquor recycle can allow for a dramatic increase in crystallization yield without sacrificing purity as long as the membrane preferentially retains the API. This process can be used for systems with a large difference in molecular weights between API and impurity, as it is the case for deferasirox (373 Da) and 4-HBA (152 Da). Because of the crystallization step, the purification requirements of the membrane are less onerous than if the membrane were used on its own. The integrated membrane module only needs to prevent accumulation of impurities over time. Membrane filtration has the additional advantage that it can be used to concentrate the mother liquor in mixed solvent/ antisolvent systems. The fact that pure solid phase is produced (as opposed to purely liquid phase separations like extractions) will also be of benefit in producing a final crystal form for formulation of a drug product without any additional steps for isolation. In a continuous steady state processes, it can be difficult to attain the same yield purity relationship as a well-designed batch equilibrium one. It was demonstrated that the steady state yield and purity values obtained using the integrated MSMPR with membrane recycle matched or exceeded the values of the batch equilibrium crystallizations. The methodology outlined in this study can also be applied to other systems to obtain acceptable yield and purity in continuous systems. Finally, membrane screening and process modeling can be used to assess the impact of inclusion of an OSN membrane with a relatively low experimental burden.

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AUTHOR INFORMATION

Article

ABBREVIATIONS AND NOMENCLATURE 4-HBA = 4-hydrazinobenzoic acid API = active pharmaceutical ingredient CLD = chord length distribution CSD = crystal size distribution FBRM = focused beam reflectance measurement HPLC = high performance liquid chromatography MW = molecular weight MWCO = molecular weight cutoff OSN = organic solvent nanofiltration PBI = polybenzimidazole PEEK = polyether ether ketone PP = polypropylene THF = tetrahydrofuran NOMENCLATURE AM = membrane area, m2 B = birth rate, #/(m3 min) Csat,i = equilibrium concentration (solubility) of solute i*, kg/m3 CP,i; CM,i; CF,i; CR,i; CW,i = liquid concentration of solute i* in permeate/membrane inlet/feed/recycle/purge stream, % w/ w (kg/m3 in model calculations) CSS,i = steady state concentration of solute i*, kg/m3 G = growth rate, m/min kB, b = kinetic nucleation parameters, #/(m3 min), kG, g = kinetic growth parameters, m/min, kv = volumetric shape factor L = characteristic size of particle/crystal, m Ṁ P; Ṁ M; Ṁ F; Ṁ R; Ṁ W; Ṁ AS = mass flow rates of the permeate/membrane inlet/feed/recycle/purge/antisolvent stream, g/min MT = suspension density (crystals in slurry), kg/m3 n(L) = number density per unit volume and crystal size, #/m4 n0 = number density of zero sized crystals, #/m4 P = permeability, L/(m2 h bar) p = purification ratio QIN, QOUT = volumetric flow rate at crystallizer inlet and outlet, (mL/min) QP; QM; QF; QR; QW; QAS = volumetric flow rate of permeate/membrane inlet/feed/recycle/purge/antisolvent stream, mL/min RRi = rejection factor of solute i Y = yield Δp = pressure drop over membrane, bar ρc = solid density, kg/m3 τ = residence time, min *subscript i: i = A for API deferasirox; i = I for impurity 4HBA REFERENCES

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Corresponding Author

*E-mail: [email protected]. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS We acknowledge the Novartis-MIT Center for Continuous Manufacturing for funding and technical guidance. 626



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Use of Continuous MSMPR Crystallization with Integrated

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