Separation of Enantiomers by Continuous Preferential Crystallization

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Separation of enantiomers by continuous preferential crystallization: Experimental realization using a coupled crystallizer configuration Joussef Hussein Chaaban, Kim Dam-Johansen, Tommy Skovby, and Søren Kiil Org. Process Res. Dev., Just Accepted Manuscript • DOI: 10.1021/op400087g • Publication Date (Web): 02 Jul 2013 Downloaded from http://pubs.acs.org on July 11, 2013

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Organic Process Research & Development

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Separation of enantiomers by continuous preferential

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crystallization: Experimental realization using a coupled

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crystallizer configuration

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Joussef Hussein Chaaban1, Kim Dam-Johansen2, Tommy Skovby3 and Søren Kiil4*

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1

Department of Chemical and Biochemical Engineering, Technical University of Denmark, DTU,

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Søltofts plads , Building 229, DK−2800 Kgs. Lyngby, Denmark

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E-mail: [email protected]

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Department of Chemical and Biochemical Engineering, Technical University of Denmark, DTU,

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Søltofts plads , Building 229, DK−2800 Kgs. Lyngby, Denmark

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E-mail: [email protected] 3

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H. Lundbeck A/S, Chemical Technology and Implementation,

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Oddenvej 182, DK−4500 Nykøbing Sjælland, Denmark

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Department of Chemical and Biochemical Engineering, Technical University of Denmark, DTU, Søltofts plads , Building 229, DK−2800 Kgs. Lyngby, Denmark

ACS Paragon Plus Environment

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* Corresponding author: Tel.: +45 4525 2827. Fax: +45 4588 2258. E-mail: [email protected]

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ABSTRACT

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The experimental realization of a continuous preferential crystallization process, consisting of two

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mixed flow crystallizers coupled via crystal-free liquid exchange streams and with only the liquid

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phases operated continuously, is addressed. Experiments in triplicate, using the conglomerate forming

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system of DL-asparagine monohydrate in water, were conducted and the achievement of nearly racemic

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composition of the liquid phase in the crystallizers was verified. An experiment was also carried out

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using seed crystals of a smaller average particle size than used in the reference experiments. Successful

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enantioseparation by crystal growth, with the repeatability being within ±10% deviation, was obtained.

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However, slow crystal growth, due to a low surface integration rate, led to a negligible consumption of

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the desired enantiomer in the feed solution added resulting in low productivities. Productivities, yields,

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and purities of solid products were influenced by the morphological differences in the seed crystals. Due

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to irregularly shaped seed crystals, increase in the productivities and yields were achieved in the L-

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Tank. Lower purities of solid products from the L-Tank compared to solid products from the D-Tank

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were obtained. This could be due to surface nucleation of D-asparagine monohydrate, ascribed to the

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surface structure of the seeds of L-asparagine monohydrate supplied. Improvements in productivity,

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yield, and purity in the L-Tank, for the same process duration, were realized using seed crystals of lower

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average particle size having a smoother surface structure. The main advantages compared to other

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separation processes are low capital cost, high crystal purity and yield, ease of upscaling, increased

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safety, and reduced environmental impact due to reduction in amount of solvent used. The application is

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currently limited to conglomerate forming systems, but the separation concept may open new

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possibilities for process improvements regarding enantioseparation of racemic compound forming

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systems as well.

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KEYWORDS: Conglomerate; Crystal growth; Continuous pharmaceutical production; Process


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intensification

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INTRODUCTION

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Continuous pharmaceutical production and separation processes in minireactors are receiving increasing

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attention as a mean of efficient and safe manufacturing1,2. Continuous crystallization also has a great

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potential in this respect. A substantial number of active ingredients in the pharmaceutical, agrochemical,

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food, and cosmetic industries are racemic mixtures, i.e. conglomerates and racemic compounds.

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Usually, only one enantiomer is biologically effective and the other enantiomer can be considered as an

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impurity comprising 50% of the mixture. Human and environmental side effects are in many cases

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attributed to the presence of counter enantiomers. A clear evidence of the potential consequences of

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administering racemic drugs is e.g. the dramatic incident of the drug thalidomide in the last century,

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which caused fetal abnormalities, such as severely underdeveloped limbs3,4. During the last decade, the

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commercialization of chiral active ingredients attracted great attention, which has resulted in new

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guidelines for the commercialization of chiral ingredients by the Food and Drug Administration (FDA),

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the Japanese authorities, and the European Community5-7. Strict regulatory requirements and increasing

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public demand for the production of chiral active ingredients in optically pure form have sparked

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intensive research devoted to the development of reliable and cost-efficient production processes. In this

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context, two options are available: 1) Enantioselective synthesis (chemical or enzymatic asymmetric

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synthesis) or 2) Enantioselective separation of racemic mixtures.

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As a result of intensified efforts put into the development of enantioselective synthesis, remarkable

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progress has been achieved, chemically as well as enzymatic5,8. However, an obstacle to the wide

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commercial application is the cost of such selective synthesis processes. Among other reasons, the use

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of expensive chemical auxiliaries and reagents as well as enzymes must be considered. Therefore,

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separation of enantiomers from racemic mixtures comprises an attractive alternative for the production

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of optically pure enantiomers5.

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The technique most utilized commercially is chiral chromatography with its wide selection of chiral

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stationary phases. This allows the separation of many racemates9-11. A version of chiral chromatography


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successfully applied in industrial scale is continuous countercurrent Simulated Moving-Bed (SMB)

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chromatography enabling an increase in productivity and realizing the goal of achieving optically pure

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enantiomers12. Nevertheless, highly expensive chiral stationary phases and a dilute separation product

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have initiated the search for and development of alternative technologies which, among others, include

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1) diastereomeric (or classical) resolution13-15, 2) chiral membranes16-18, 3) dynamic resolution19-22, 4)

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liquid-liquid extraction23-26, and 5) preferential crystallization27-34. Preferential crystallization, in

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particular, has attracted great attention, the reason being that the majority of chiral pharmaceutical,

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agrochemical, food, and cosmetic products, and their intermediates are produced as highly pure solids.

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Coupled Batch Preferential Crystallization

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A thoroughly studied preferential crystallization process is the Coupled Batch Preferential

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Crystallization (CB-PC) process27-28, 33-34. It consists of two batch crystallizers coupled via crystal-free

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liquid exchange streams between the crystallizers. The process can be operated either isothermally

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throughout the duration of separation27-28 or with a transient decrease in crystallization temperature27.

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Both operations are performed with external supply of enantiopure seeds to facilitate selective crystal

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growth. Both operation modes for the CB-PC process have proven to be feasible for the simultaneous

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separation of enantiomers by crystallization27-28, 34 with improved productivity and yield compared to

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preferential crystallization in a single batch crystallizer. Because of the crystal-free liquid exchange

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streams, the delay or even the suppression of the nucleation of excess counter enantiomer is possible

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due to the nearly racemic composition of the mother liquors. However, a characteristic drawback is the

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decrease in the crystallization driving force, i.e. supersaturation of desired enantiomer, with time in the

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crystallizing solutions.

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The CB-PC process is applicable for conglomerate forming systems27-28 and racemic compound forming

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systems as well34. In case of conglomerate forming systems, the enantiomers are separated

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simultaneously (one in each crystallizer) by seeding. Enrichment of initial racemic solution is not


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necessary. An essential requirement allowing for separation is a sufficiently higher solubility of the racemate compared to that of the pure enantiomers. For a racemic compound forming system the process is feasible only if 1) sufficient enrichment of the desired enantiomer beyond the eutectic composition is present in the initial racemic solution, and 2) the solubility of the racemic compound is sufficiently higher than that of pure enantiomers13,30. In this case, the simultaneous separation of the racemic compound and the desired enantiomer in excess can be performed. A loss in the yield of pure solid product of the desired enantiomer is unavoidable in this process due to the consumption of that enantiomer in the formation of solid racemic compound. A detailed theoretical study followed by experimental investigation of the parameters influencing the enantioseparation performance of the CB-PC process was performed by Elsner et al.27,33. The study was conducted on the conglomerate forming system of the amino acid threonine using water as solvent. The study highlighted that an increase in productivity can be achieved by manipulating critical process parameters such as flow rate of crystal-free liquid exchange between crystallizers, initial enantiomeric excess in solutions, mass of seeds supplied, time for initiating liquid exchange (switching time), process duration time, and crystallization temperature.

Coupled Continuous Preferential Crystallization In the present study, the continuous preferential crystallization apparatus called Coupled Continuous Preferential Crystallization (CC-PC) process is investigated. This newly developed process is based on the thoroughly investigated CB-PC apparatus developed by Elsner et al.27, 34, and a proof-of-concept investigation of a continuous isothermal version of the process is conducted. The objectives are to assess the potential achievement of simultaneous separation of the enantiomers by crystal growth and the achievement of a nearly racemic composition of the liquid phase in each crystallizer. Due to the nearly racemic composition of the mother liquor, the crystal-free liquid exchange could ensure the suppression of crystallization of the counter enantiomer.


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The newly developed continuous process is evaluated with respect to productivity, yield, and purity of the solid product produced. To evaluate the robustness of the process, experiments in triplicates with almost identical experimental conditions were conducted. Furthermore, to investigate the flexibility of the process, an experiment where seed particles with smaller size were used was also conducted.

Chemical model system The investigation is performed as a case study of enantioseparation of the enantiomers of the amino acid asparagine monohydrate from an aqueous racemic mixture by preferential crystallization. The selection of the asparagine monohydrate system as model compound for the investigation in this study was based mainly upon two reasons: 1. Aqueous solution of the compound belongs to the conglomerate forming systems36. 2. Solubility data and data for crystallization kinetics are available in the litterature28. The solubility of the asparagine monohydrate compound in water decreases with decreasing temperature28. This facilitates the application of cooling crystallization as separation method in the continuous preferential crystallization process considered. Enantioseparation of the enantiomers of the selected model compound in the isothermal CB-PC process was also investigated by PetruševskaSeebach et. al 28. It was concluded that the crystallization process is controlled predominantly by surface integration at the given experimental conditions. In general, surface integration rate exhibited by many organic molecules is known to be rather low35 leading to a slow consumption of the enantiomer from the liquid phase. It was also found that the growth rate of the desired enantiomer is dependent on the concentration of the counter enantiomer which can inhibit the growth of the desired enantiomer28. The low rate of crystal growth means that the consumption of the desired enantiomer from the continuously added feed solution in the CC-PC process considered in this study is, for any practical uses, expected to be negligible. Therefore, the purpose of this work is not to demonstrate a quantitative separation of the two monohydrate enantiomers in the feed streams, but, as an initial step, to show that it is possible to


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grow only one of the enantiomers in each of the two crystallizers. Thus, the estimated productivities of the solid product produced are expected to be low. However, it is believed that process parameters influencing the growth rate of the seed crystals, such as crystallization temperatures, can be manipulated so that feasible separation of the enantiomers of quantitative scale can be realized.

EXPERIMENTAL SECTION Materials The chiral compounds used as enantiopure seeds in the present study were D-(+)-Asparagine monohydrate (D-AsnH2O) (≥99%) and L-(−)-Asparagine monohydrate (L-AsnH2O) (≥99%) purchased from Sigma-Aldrich (Copenhagen, Denmark). DL-(±)-Asparagine monohydrate (DL-AsnH2O) (≥99.0%) was also obtained from Sigma-Aldrich (Copenhagen, Denmark) and dissolved in demineralized water (solvent) and used as racemic feed for the experiments.

Experimental setup and procedures Coupled Continuous Preferential Crystallization (CC-PC) process. The schematic diagram of the experimental setup is shown in Figure 1 and a picture of the setup is shown in Figure 2.


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Figure 1. Schematic diagram of the CC-PC process for the preferential crystallization experiments. TF symbolizes the temperature in the feed tank, TF1 is the temperature of feed entering the D-Tank, T1 is the temperature in the D-Tank, T2 is the temperature in the L-Tank, TE1 is the temperature of mother liquor effluent from the D-Tank, T1-2 is the temperature of exchange stream from the D-Tank to the LTank and T2-1 is the temperature of exchange stream from the L-Tank to the D-Tank. PF1 and PF2 are the feed pumps for the D- and L-Tank, respectively. P1-2 is the liquid exchange pump from the D-Tank to the L-Tank. Similarly, P2-1 is the liquid exchange pump from the L-Tank to the D-Tank. PP1 and PP2 are the purge pumps for the D- and L-Tank, respectively. Blue and red filled circles symbolize growing enantiopure seeds of D-AsnH2O and L-AsnH2O in the D- and L-Tank, respectively. The grey boxes in the D- and L-Tank represent filters for mother liquor exchange. The CC-PC process in Figure 1 is a modification of the CB-PC process. The modification consists of the addition of two feed streams consisting of undersaturated racemic solution, one to each crystallizer, and


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two purge streams from the crystal-free liquid exchange to maintain constant volumes in the crystallizers. It should be emphasized that the process is operated continuously with respect to the liquid phase only.

Figure 2. Photograph of experimental setup for the preferential crystallization experiments. The small insert (lower left corner) shows the instrumentation in the crystallization vessels. Six individual pumps are necessary to run the setup in a continuous mode. Each of the crystallizers (D-Tank and L-Tank) consisted of jacketed round-bottom glass vessels of 250 mL in total volume with 75 mm ID flange and lid with five necks. The stirrer used in each crystallizer was a 45 mm ringed propeller-type impeller (Heidolph PR 32 Ringed Propeller, Schwabach, Germany) mounted on a motor (IKA Eurostar Digital, Staufen, Germany). Two separate thermostated circulation water baths (LAUDA Alpha RA 8 for the D-Tank and LAUDA Ecoline Star Edition RE 104 for the LTank, Lauda-Königshofen, Germany) were used for cooling the crystallizers. The liquid solutions from and into the crystallizers were transported by six separate peristaltic pumps (Longer Pump LP-BT-1001F, pump head YZ1515x, Baoding Longer Precision Pump Co. Ltd., Hebei, China).


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Preparation of undersaturated racemic feed solution. The racemic feed solution was prepared by dissolving 100 g of solid DL-AsnH2O in 1.90 kg of demineralized water in a jacketed cylindrical glass vessel with total volume of 2.5 liters (Feed tank). An in-house made stirrer shaft fitted with a half-moon impeller of PTFE mounted on a motor (IKA Eurostar Power Control-Visc, Staufen, Germany) was placed in the center of the feed vessel. The feed slurry was heated up to about 31 °C by a thermostated circulation water bath (LAUDA M3, Lauda-Königshofen, Germany) and maintained for 1 hour for complete dissolution. The solution was maintained at that temperature throughout the duration of the experiment. The temperature of feed solution was recorded via the Vernier Software package Logger Lite 1.6.1 using the thermocouple sensor connected to Go! Link interface (Vernier Software and Technology LLC, Beaverton, OR, USA).

Continuous preferential crystallization experiments. The first experiments were performed with almost identical experimental conditions according to Table 1. An additional experiment was conducted with the experimental conditions given in Table 1, but using enantiopure seeds with a smaller average particle size (i.e. a sieve fraction of 150-250 µm).

Table 1. Operational conditions for experiments conducted in triplicates. Parameter

Value

Unit

Feed temperature

304

K

Concentration

5

wt% DL-AsnH2O

Total mass of feed solution

2000

gram

Stirring rate

250

rpm

Heating rate

14

K/h

Feed tank

D-Tank and L-Tank ACS Paragon Plus Environment

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Inlet feed temperature

300

K

Saturation temperature

296

K

Crystallization temperature

292

K

Effective volume

220

mL

Residence time of liquid phase

15

min

Inlet feed flow rate

7.3333

mL/min

Liquid exchange flow rate

7.3333

mL/min

Effluent flow rate

14.6667

mL/min

Mass of seeds

0.4648

gram

Sieve fraction of seeds

250-355

µm

Cooling rate

14

K/h

Stirring rate

250

rpm

Initially, 220 mL of undersaturated racemic solution from the feed vessel was transferred into each crystallizer by two separate peristaltic feed pumps (Longer Pump LP-BT-100-1F, YZ1515x, Baoding Longer Precision Pump Co. Ltd., Hebei, China). Stirring of solutions was initiated and maintained at about 250 rpm. A liquid sample of about 4 mL was withdrawn from the feed solution for offline polarimetric analysis and density determination. The racemic solutions in the crystallizers were then cooled with constant cooling rate about 14 K/h to a crystallization temperature of about 19 °C and maintained throughout the duration of the experiments. The temperatures were recorded via the Vernier Software package Logger Lite 1.6.1 using Al-Cr thermocouples type K (0.1 K in resolution) connected to thermocouple sensors which in turn were connected to Go!Link interface (Vernier Software and Technology LLC, Beaverton, OR, USA). The experiments were initiated by starting the peristaltic pumps at a flow rate of 7.333 mL/min. The liquid phase was separated from the solids by in-situ sintered metal filters with pore size of 10 µm (IDEX Health & Science, Wertheim-Mondfeld, Germany) fitted in each crystallizer. When the temperature of the solutions in the crystallizers was stabilized, dry homochiral seeds of pure enantiomer ACS Paragon Plus Environment

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were externally supplied to each crystallizer. The mass of seeds was collected from a certain sieve fraction and analyzed offline for the particle size distribution with laser diffraction particle size analyzer (Malvern Mastersizer 2000, Malvern Instruments Ltd.). Liquid samples of about 4 mL were withdrawn from the respective purge stream from each crystallizer every 15 minutes after seeding. The samples were analyzed offline for optical rotation at about 23 °C and 589 nm in 1 mL glass cell (PerkinElmer 341 Polarimeter, PerkinElmer Life and Analytical Sciences, Shelton, CT, USA) with a reproducibility of ±0.002˚. Additionally, the liquid density of the liquid samples was measured offline at about 23 °C using Density Meter DMA 35 (Anton Paar GmbH, Graz, Austria) with a reproducibility of ±0.0005 g/mL. After each experiment, the pumps and the heating and cooling water baths were turned off and the slurries were filtered using ashless cellulose filter paper with 7.0 cm in diameter and pore size of 3 µm placed on a Büchner funnel with vacuum suction. The filtered solids were washed with 50 mL ice cold ethanol to remove mother liquor adhering to the crystal surface. A short time delay occurred between the filtration tasks. The solids were dried for 24 hours in a ventilated cupboard and weighed for yield and productivity estimation. The solids were also analyzed for optical purity using offline polarimetry and simultaneous thermogravimetric analysis (STA) consisting of differential scanning calorimetry (DSC) and thermogravimetric analysis (TGA) (STA 409, NETZSCH-Gerätebau GmbH, Selb, Germany) using Pt/Rh crucibles with lids. The STA analysis was conducted at a heating rate of 5 K/min and up to 300 ˚C with purge gas flow of 100 mL N2/min. Furthermore, samples of the seeds and harvested solids were coated with gold particles and analyzed by Scanning Electron Microscopy (SEM) imaging at the Center for Electron Nanoscopy (CEN) in DTU.

PROCESS METRICS To quantify the efficiency of the present preferential crystallization process, productivity, yield, and purity of the solid product have been considered. Additionally, the liquid phase concentration (mass


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fraction) change during the crystallization experiments has been assessed to observe the performance of the process.

Analysis Concentration of enantiomer in the liquid phase The concentration (mass fraction) profiles for each enantiomer in the mother liquor in each crystallizer during the crystallization process are calculated by a combination of the polarimetric analysis and the density analysis of samples of the mother liquor37. The polarimetric analysis determines the difference in the concentrations of D-Asn and L-Asn in the mother liquor according to the calibration curve obtained for different aqueous solutions with known concentrations of each enantiomer. The density analysis determines the total concentration of enantiomers in the mother liquor by using the calibration curve of liquid density for known concentrations in a series of aqueous solutions.

Supersaturation ratio The crystallization driving force, i.e. supersaturation, in terms of the supersaturation ratio of the desired enantiomer in each crystallizer is also estimated. The supersaturation ratio of the desired enantiomer is defined as the ratio between the mass fraction of the desired enantiomer and the equilibrium mass fraction of the desired enantiomer. The equilibrium mass fraction is in turn a function of the crystallization temperature and the concentration of the counter enantiomer in the crystallizing suspensions27,28.

Productivity and yield The productivity is defined as the ratio between the total mass of solid product produced of both enantiomers and the total volume of crystallizing liquid in both crystallizers per unit time. The total mass of solid product produced of both enantiomers is estimated as the sum of the difference between


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the mass of dry solids harvested and the initial mass of the supplied seeds multiplied by the estimated optical purity for each enantiomer. The total volume of crystallizing liquid in both crystallizers is defined as the sum of the volume of the mother liquors in the crystallizers. The solid product yield is in turn defined as the product of the optical purity and the ratio between the mass of the solid product of enantiomer produced and the total mass of enantiomer present in the feed solution. The total mass of enantiomer in the feed solution is in turn defined as the product of the volumetric flow rate of the feed entering the crystallizer, process time, the density of liquid feed solution, and the concentration (mass fraction) of the enantiomer in feed solution.

Optical purity The optical purity of the solid product was determined quantitatively by polarimetry. Solid samples were dissolved in water and the solutions were analyzed for optical rotation and compared to the reference solutions of the pure enantiomer estimated from a linear fit of the calibration curve. The optical purity was then calculated as the ratio between the actual observed optical rotation of a solution of the solid product with a certain concentration and the expected observed optical rotation of a solution of the pure enantiomer with the same concentration.

RESULTS AND DISCUSSION Results of the crystallization process, namely crystallization temperature and optical rotation of the liquid phase are now presented and discussed. Estimations of the liquid phase compositions, supersaturation ratios, and growth rates of crystals are also given and discussed. Moreover, the results obtained of the solid phase analysis in terms of purity, productivity, and yield of solid products produced are evaluated.

Crystallization temperature and observed optical rotation


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Crystallization temperatures in the D- and L-Tank and the observed optical rotation of the mother liquors are shown in Figure 3 for all three initial experiments.

Figure 3. Temperatures and observed optical rotations of the mother liquors in the D- and L-Tank during preferential crystallization experiments. Error bars represent the standard deviation. Where error bars cannot be seen, a standard deviation of zero was found. The racemate has an optical rotation of zero. At time equal zero, enantiopure seeds are introduced. The experimental conditions are provided in Table 1. Minor fluctuations in the crystallizer conditions are due to heat transfer limitations, the efficiency of mixing in the crystallizers, and the placement of the thermocouple. Nevertheless, under isothermal conditions the process performed with a reasonable repeatability.


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The reason for the selection of 19 °C as crystallization temperature in the present investigation is that a lower crystallization temperature than the surrounding air temperature enables natural heating of the liquid exchange streams and thereby avoids the nucleation of the counter enantiomer in the tubing between the crystallizers. The liquid exchange temperature achieved was about 23 °C, which is the saturation temperature of the pure enantiomer28. Operating at crystallization temperature higher than the temperature of the surroundings requires a heating system to avoid cooling of the liquid exchange streams. Due to limited space, such a heating system was problematic to implement. The observed optical rotations over time indicate that the composition of the mother liquor is nearly racemic because of values very close to zero.

Liquid phase concentration profiles of enantiomers The estimated liquid phase concentrations of D-Asn and L-Asn in each tank as a function of time, calculated from the polarimetric and the liquid density measurements, are shown in Figure 4.


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1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32324 33 34325 35 36 37326 38 39 40327 41 42328 43 44 45329 46 47330 48 49 331 50 51 52332 53 54333 55 56 334 57 58 59 60


Figure 4. Liquid phase concentrations of enantiomers in each tank as a function of time. Error bars represent the standard deviation. At time equal zero, enantiopure seeds are introduced. As shown above by the observed optical rotation, the liquid phase composition is nearly racemic and this is due to the crystal-free liquid exchange streams between the two crystallizers. Furthermore, as shown in Figure 4, a nearly constant concentration of the preferred enantiomer in each crystallizer is achieved and that the concentration of enantiomers is the same as the concentration in the feed solution. Because of the slow rate of surface integration, negligibly low consumption of the desired enantiomer in each crystallizer occurs. It is thus expected that the crystallization driving force, i.e. supersaturation, and thereby the rate of crystal growth, is constant for the desired enantiomer in each crystallizer and can be determined only by knowing the concentration of the enantiomers in the feed solution.


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335

1 2 3 336 4 5 337 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38338 39 40339 41 42 340 43 44 45 46341 47 48342 49 50 343 51 52 53344 54 55345 56 57 58346 59 60

Supersaturation and crystal growth The estimated supersaturation ratio of the preferred enantiomer in each crystallizer is given in Figure 5.

Figure 5. Transient development in the supersaturation ratio for the preferred enantiomer in each tank. At time equal zero, enantiopure seeds are introduced. Indeed the crystallization driving force of the preferred enantiomer is almost constant in the respective crystallizers. Obviously, there exists a slight difference in the supersaturation driving force of the crystallizing enantiomer which could be explained by the difference in the morphology and the size of the enantiopure seed crystals supplied27,38 and the minor fluctuations in the temperature of the crystallizing solutions. An inherent consequence of the above mentioned constant driving force is a constant crystal growth rate


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of the enantiopure seeds in each tank throughout the duration of the crystallization experiments. This is desirable because of the possibility for improving the control of the purity and the particle size of the growing crystals.

Analysis of the solid product A summary of the results of the analysis of the solid product produced in terms of productivity, yield, and purity for the triplicate experiments (Exp 1-3) conducted and the additional experiment (Exp 4) are listed in Table 2. Some statistics are given in Table 3. With respect to the size and the mass of the seed crystals used in this study (see Table 1), the selection was mainly based upon the amount of the seed crystals of each enantiomer available for performing three identical experiments, taking into account the increase in the risk of agglomeration and secondary nucleation caused by attrition and breakage when using a large amount of seed crystals.

Table 2. Summary of the results of the productivity, yield of solid product, purity of solid product produced, and the Seed Efficiency Number for experiments Exp1-3 in the D-Tank and L-Tank and the additional experiment (Exp 4). Note the units. Exp 1

Exp 2

Tank

D

Sieve fraction of seeds [µm]

250-355

Purity* [%] Productivity** [g/(L h)]

L

100.0

93.4

0.53

D

Exp 3 L

250-355 100.0

92.2

0.54

D

Exp 4 L

250-355 100.0

93.9

0.60

D

L

150-250 100

100

0.94

Yield*** [%]

0.88

1.54

1.12

1.41

1.25

1.62

1.64

2.08

Fraction of solids produced**** [%]

23.8

36.8

28.1

35.5

30.2

37.2

36.6

43.4

Seed Efficiency Number***** [%]

31.2

54.5

39.0

50.8

43.2

57.0

57.7

76.8

* The optical purity is defined as the ratio between the actual observed optical rotation of a solution of the solid product with a certain concentration and the expected observed optical rotation of a solution ACS Paragon Plus Environment

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1 2 366 3 4 367 5 6 368 7 369 8 9 370 10371 11372 12 13373 14 374 15 16 375 17 18 376 19 20 21377 22 23378 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39379 40 41380 42 43 381 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

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of the pure enantiomer with the same concentration. ** The productivity is defined as the ratio between the total mass of solid product produced of both enantiomers and the total volume of crystallizing liquid in both crystallizers per unit time. *** The solid product yield is defined as the product of the optical purity and the ratio between the mass of the solid product produced and the total mass of enantiomer in the feed solution. **** The fraction of solids produced is defined as the ratio of the mass of solids produced by growth of seed crystals to the total mass of solids harvested (seeds and solids produced) at the end of the experiment. ***** The Seed Efficiency Number is defined as the product of the optical purity and the ratio between the mass of the solid product produced and the initial mass of enantiopure seeds added.

Table 3. Summary of some statistical results of the productivity, yield of solid product, purity of solid product produced, and Seed Efficiency Number for experiments Exp 1-3 in the D- and L-Tank. Note the units. Tank Purity [%] Productivity [g/(L h)]

D

L

100.0 ±0.0

93.9 ±2.0

0.56 ±0.04

Yield [%]

1.08 ±0.19

1.52 ±0.11

Fraction of solids produced [%]

27.4 ±3.3

36.5 ±0.9

Seed Efficiency Number [%]

37.8 ±6.1

54.1 ±3.1

SEM photomicrographs in Figure 6 show the difference in the growth of the seed crystals in both the DTank and the L-Tank for the experiments Exp 1 and 3.


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1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18382 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37383 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55384 56 57385 58 59 60


Figure 6. SEM photomicrographs of the seeds of D-AsnH2O ((a)) and L-AsnH2O ((b)) and the final


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1 2 3 387 4 5 388 6 7 8 389 9 10 11390 12 13391 14 15392 16 17 18393 19 20394 21 22 395 23 24 25396 26 27397 28 29 30398 31 32399 33 34400 35 36 37401 38 39402 40 41 403 42 43 44404 45 46405 47 48 49 50 51 52 53 54 55 56 57 58 59 60

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solid products showing the growth of the seed crystals of D-AsnH2O in the D-Tank for experiments Exp 1 and 3 ((c) and (e)) and the growth of the seed crystals of L-AsnH2O in the L-Tank for experiments Exp 1 and 3 ((d) and (f)). Note the scale of the images. Productivity and yield As shown in Table 2, the productivity and yield of solid product are low due to the slow crystal growth controlled by low surface integration rates. However, slightly higher values are obtained for the product from the L-Tank, in which L-AsnH2O seed crystals were grown, compared to the D-Tank. This could be explained by the difference in the morphology and the size of enantiopure seed crystals acquired27,38. The enantiopure seed crystals of L-AsnH2O have a more random structure with many irregularities compared to seeds of D-AsnH2O. These irregularities could be a result of attrition and abrasion in the manufacturing process and the sieve analysis performed prior to introduction to the crystallizing solution in the L-Tank. The stirring of the crystallizing suspensions does not cause any attrition or breakage of the enantiopure seed crystals introduced during the experiments performed. The rougher and irregular morphology of the seed crystals of L-AsnH2O (Figure 6 (b)) allows for rapid and more random crystal growth (Figure 6 (d) and (f)) compared to the more systematic growth of seed crystals of D-AsnH2O (Figure 6 (a), (c) and (e)). The reason is the higher number of nearest neighboring molecules accompanied with the irregular grow sites such as kink sites on the seed crystals of L-AsnH2O, which promote rapid crystal growth40. The particle size distributions (PSDs) of the enantiopure seeds used in the experiments are shown in Figure 7.


24

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1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31406 32 33407 34 35 408 36 37 38409 39 40 41410 42 43 411 44 45 46412 47 48413 49 50 414 51 52 53415 54 55416 56 57 58417 59 60


Figure 7. Normalized frequency functions of the PSDs for the enantiopure seeds of D-AsnH2O and LAsnH2O supplied in experiments Exp1-3 (sieve fraction 250-355 µm) and Exp 4 (sieve fraction 150-250 µm). Note the logarithmic scale on the abscissa. The PSDs are plotted as dFv/dlog(dp) versus log dp in order to normalize the area under the differential curve to 1 which make the PSDs comparable41. The seed crystals of L-AsnH2O in both sieve fractions have a narrower PSD with larger average particle size compared to the seeds of D-AsnH2O. This could possibly be due to the manufacturing of the respective seeds and the subsequent sieve fraction analysis. The result is more uniform and larger seeds of L-AsnH2O. This means a larger surface area available for the crystal growth of L-AsnH2O, which could contribute to the higher yield and slightly higher productivity achieved in Exp 1-3. Additionally, the seeds used in Exp 4 have a smaller average particle size compared to the larger sieve fraction used in Exp 1-3 (Table 2). Thus, the higher yield and


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1 2 3 419 4 5 420 6 7 421 8 9 10422 11 12423 13 14 424 15 16 17425 18 19426 20 21 22427 23 24428 25 26429 27 28 29430 30 31431 32 33 432 34 35 36433 37 38434 39 40 41435 42 43436 44 45437 46 47 48438 49 50439 51 52440 53 54 55441 56 57442 58 59 60

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productivity in this case could be due to the larger surface area available for crystal growth in the crystallizing solutions. In the present study, the duration of the crystallization experiments of 90 minutes was limited by the volume of the feed solution and the mean residence time of the liquid phase in the crystallizers. A mean residence time of the liquid phase of 15 minutes was chosen to ensure that the flow rate of the crystalfree liquid exchange is sufficiently high to obtain a racemic composition in the mother liquor in both crystallizers, because identical crystallization kinetics of the monohydrate enantiomers are assumed. The supersaturation ratio of the preferred enantiomer in each crystallizer could be maintained constant (Figure 5) thereby obtaining a constant growth rate. In this context, a longer duration could be realized by a longer mean residence time of the liquid phase in order to improve the utilization of the mass of enantiomers in the feed solution thereby improving the productivities and the yields of the solid product in each crystallizer. In Table 2 it is shown that the mass of seed crystals supplied comprises the major fraction of the mass of solids harvested at the end of each experiment. Increasing the mean residence time of the liquid phase could increase the fraction of the solids produced of the total amount of the harvested solids.

Optical purity The optical purity estimated is lower than the desired 100% for the solid product from the L-Tank as listed in Table 2. The reason for the purity drop in the L-Tank could be explained by the surface structure and size of the seed crystals at which surface nucleation of the counter enantiomer could occur38. The rougher surface morphology of the seed crystals of L-AsnH2O and the larger surface area available for growth of the seeds of L-AsnH2O in the L-Tank could result in a relatively higher growth rate of the seeds of L-AsnH2O (Figure 5) and thereby the observed purity drop in the L-Tank38. The randomly oriented particles observed on the irregular locations on the surface of the grown seed crystals of L-AsnH2O (Figure 6 (b), (d) and (f)) suggest the surface nucleation of the counter enantiomer D-


26

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443

1 2 3 444 4 5 445 6 7 446 8 9 10447 11 12448 13 14 449 15 16 17450 18 19451 20 21 22452 23 24453 25 26 454 27 28 29455 30 31456 32 33 34457 35 36458 37 38459 39 40 41460 42 43461 44 45 462 46 47 48463 49 50464 51 52 465 53 54 55466 56 57467 58 59 60


AsnH2O in the L-Tank. In contrast, due to the more uniform and smoother surface structure of the smaller seed crystals of L-AsnH2O supplied to the L-Tank in Exp 4, 100% purity was achieved. This fact could explain the effect of the surface morphology of the seed crystals supplied on the purity of the solid product achieved in the preferential crystallization experiments27,38. However, it should be emphasized that in general the purity drop in preferential crystallization is unavoidable due to the changes of the surface structure of the seed crystals during the crystallization process. The purity drop can be delayed by controlling the surface morphology of the seed crystals and the process duration27,38. To qualitatively assess whether the nucleation of the counter enantiomer D-AsnH2O on the surface of the seed crystals of L-AsnH2O in the L-Tank occurred or not, STA measurements were conducted. The STA measurements showed that the decomposition of D-AsnH2O and L-AsnH2O occurs simultaneously at about 215 °C. Therefore, it was not possible to distinguish between the racemate and the pure enantiomer. However, a difference in the stage of the dehydration of crystal water occurred, but it was not possible to confirm the presence of the counter enantiomer D-AsnH2O. The reason is that the difference in the surface structure, crystallinity, crystal defects due to grinding, ageing etc. of the crystals in both pure form and the solid products from the crystallization experiments allow for different dehydration kinetics42-43.

Unutilized enantiomer in purge From a simple steady-state total mass balance for both enantiomers for a control volume of the CC-PC process, the fraction of unutilized enantiomer in the purge can be estimated. It was found that the growth rate of the desired enantiomer is very low at the given experimental conditions and therefore practically no consumption of the enantiomer in the feed occurs in the crystallizer. As mentioned, this is due to the nature of the controlling mechanism for crystal growth comprising slow surface integration of the asparagine monohydrate molecules onto the surfaces of the supplied seed crystals. Furthermore, rarely do all crystal faces have the same growth rate35 and therefore growth rate differences between the


27


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1 2 3 469 4 5 470 6 7 471 8 9 10472 11 12473 13 14 474 15 16 17475 18 19476 20 21 22477 23 24478 25 26479 27 28 29480 30 31481 32 33 482 34 35 36483 37 38484 39 40 41485 42 43486 44 45487 46 47 48488 49 50489 51 52490 53 54 55491 56 57492 58 59 60

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crystal faces of the growing seed crystals could slow down the overall growth kinetics. Additionally, the inhibitory effect of the presence of the counter enantiomer28, which is also supplied by the liquid feed solution, could also contribute to lower the rate of surface integration.

Process improvement options To increase the separation potential of this particular chemical system, a modification of the crystal surfaces of the seed crystals could alter the surface integration rates and thereby enhancement of the crystal growth can potentially be realized27. Additionally, a prohibitive long residence time (days), high crystallization temperature (evaporation of water), or large amount of seed crystals with smaller size (risk for agglomeration/breakage) would be required. Therefore, it seems more beneficial to continuously add seed crystals and remove solid product. In this respect, enantiopure seed crystals of could be produced in a separate MSMPR crystallizer or by a technique called the impinging jet39. The continuous removal of solid product from each crystallizer could take place by employing a continuously operated filtration system, e.g. continuous drum filters, or by periodic slurry discharge using e.g. a hydrocyclone. Additionally, recycle of the unutilized enantiomer in the purge streams to the feed solution could improve the process efficiency in terms of productivity and solid product yield. However, the combined purge streams accumulate a diluted nearly racemic solution and would require the addition of more solid material of the enantiomers to the feed solution in order not to dilute the feed solution entering the crystallizers. This approach could result in higher productivity and yield of solid product due to the larger surface area of the continuously supplied seed crystals available for growth. This cyclic fully continuous process, however, requires a substantial extension and modification of the present setup and will be explored in a future investigation. For other conglomerate forming chemical systems, with higher rates of crystallization, the setup has great promise as a continuous separation unit.


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CONCLUSIONS In the present work, an experimental realization of the newly suggested continuous preferential crystallization process has been performed using the conglomerate forming system DL-asparagine monohydrate in water. It can be concluded that enantioseparation by the continuous process developed is achievable and that the experiments performed are repeatable. This confirms that the process is robust. It has been proven that a nearly racemic composition is achieved throughout the duration of the experiments taking into account the effects of morphological differences in the initially supplied seed crystals. Furthermore, separate growth of the supplied seed crystals of the desired enantiomer in each crystallizer was verified. It can be concluded that reasonable productivities, yields, and purities are achieved, taking into account the short process duration of 90 minutes and the morphological differences in the enantiopure seeds supplied. The lower purity achieved for the solid product from the L-Tank is ascribed to the surface nucleation of counter enantiomer D-AsnH2O due to the morphology of the enantiopure seeds of LAsnH2O. An increase in the productivity, yield and purity can be achieved in the L-Tank for the same duration time of 90 minutes only by using smaller particle size of enantiopure seeds of both enantiomers with a smoother surface structure of the L-AsnH2O crystals. The present continuous process has potential for further development into a fully continuous process to increase the efficiency of enantioseparation for conglomerate forming systems having slow growth rates controlled by surface integration. Further work is therefore required to develop and implement proper solid feeding systems for the continuous supply of enantiopure seeds, efficient downstream units necessary for the continuous harvest of solid product, and effective recycle systems for the recycle of the purge streams.


29


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ACKNOWLEDGMENTS The authors are grateful for the financial support of H. Lundbeck A/S and to Mette Larsen and Emine Yüksel Coskun for their help in the laboratory and for performing STA and Malvern PSD measurements.

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35. Tung, H.-H.; Paul, E. L.; Midler, M.; McCauley, J. A, Kinetics, In: Crystallization of Organic Compounds – An Industrial Perspective. John Wiley and Sons: 2009. 36. Thiemann, W., Disproportionation of Enantiomers by Precipitation. Journal of Molecular Evolution 1974, 4, (1), 85-97. 37. Elsner, M. P.; Menéndez, D. F.; Muslera, E. A.; Seidel-Morgenstern, A., Experimental Study and Simplified Mathematical Description of Preferential Crystallization. CHIRALITY 2005, 17, (S1), S183-S195. 38. Profir, V. M.; Matsuoka, M., Processes and phenomena of purity decrease during the optical resolution of DL-threonine by preferential crystallization. Colloids and Surfaces A: Physicochemical and Engineering Aspects 2000, 164, (2-3), 315-324. 39. Jiang, M.; Wong, M. H.; Zhu, Z.; Zhang, J.; Zhou, L.; Wang, K.; Ford Versypt, A. N.; Si, T.; Hasenberg, L. M.; Li, Y.-E.; Braatz, R. D., Towards achieving a flattop crystal size distribution by continuous seeding and controlled growth. Chemical Engineering Science 2012, 77, 2-9. 40. Myerson, A. S.; Ginde, R., Crystals, Crystal growth, and Nucleation. In: Handbook of Industrial Crystallization. Elsevier Inc.: 2002. 41. Allen, T., Particle Size Measurement. Chapman and Hall: 1996. 42. Garnier, S.; Petit, S.; Mallet, F.; Petit, M.-N.; Lemarchand, D.; Coste, S.; Lefebvre, J.; Coquerel, G., Influence of ageing, grinding and preheating on the thermal behavior of α-lactose monohydrate. International Journal of Pharmaceutics 2008, 361, (1-2), 131-140. 43. Guarini, G. G. T.; Dei, L., Kinetics and mechanism of the thermal dehydration of l. asparagine monohydrate microcrystals and single crystals: Role of ageing. Thermochimica Acta 1998, 311, (1-2), 129-140.


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