Continuous Separation of Isomers in Fluidized Bed Crystallizers

Article pubs.acs.org/crystal

Continuous Separation of Isomers in Fluidized Bed Crystallizers D. Binev,* A. Seidel-Morgenstern, and H. Lorenz Max Planck Institute for Dynamics of Complex Technical Systems, Sandtorstrasse 1, 39106, Magdeburg, Germany S Supporting Information *

ABSTRACT: A fluidized bed process exploiting two coupled crystallizers was studied for continuous preferential crystallization in two ternary systems, namely, ortho-aminobenzoic acid (OABA)/para-aminobenzoic acid (PABA)/water and L-asparagine monohydrate (L-asn·H2O), D-asparagine monohydrate (D-asn·H2O)/water. A continuous supply of seeds was realized by ultrasonic comminution of crystals, withdrawn from the bottom of the partially conical-shaped fluidized bed crystallizers. Product removal occurred via outlets placed at specific positions in the conical sections of the crystallizers. The crystal size distribution (CSD) as well as the purity of the product crystals reached steady states with purities exceeding 97%. A simplified dynamic model was developed and applied to predict solution concentrations and product CSDs. Results of primary solid−liquid equilibria and metastable zone width studies are also reported.

1. INTRODUCTION Crystallization is an important technique for the production of pure substances. The production of crystals with high purity, similar size, and a desired morphology requires a specific design and a reliable control of crystallization processes. From an industrial point of view also the robustness of the process and its trouble-free control are of particular interest.1,2 Preferential crystallization permits substances that form a simple eutectic to be resolved from their mixtures. The resolution of enantiomers can be achieved by various methods, like microbiological methods, kinetic enzymatic resolution, and chromatography.3 Some of the recent studies show that of growing importance are methods that allow continuous production of enantiopure substances, such as the combination of continuous chromatographic processes and subsequent crystallization,4,5 chiralmembrane-based separation techniques,6,7 and polymers imprinted with chiral templates.8 Although all these techniques show high chiral discrimination, they are not yet applicable for large-scale resolution. Therefore, large-scale separations are still typically achieved by classical crystallization methods. In the case of an enantioseparation process, preferential crystallization can be performed in two main ways, i.e., entrainment and simultaneous crystallization.3 Galan et al. have recently demonstrated a continuous simultaneous crystallization realized in two coupled mixed-suspension mixed-product-removal crystallizers, both connected to a feed tank. The results have shown that this method of production is promising and further investigation is needed.9 Tung et al. have shown that fluidized bed crystallizers (FBCs) can be applied in the production of organic fine chemicals on an example of a continuous kinetic separation of enantiomers in FBCs.10 Midler has also investigated the application of fluidized bed crystallization by coupling two tubular crystallizers in serial and in parallel mode. He © 2016 American Chemical Society

successfully conducted a preferential crystallization of several conglomerate forming systems.11,12 An example of a continuous preferential crystallization process in two parallel connected FBCs is illustrated in a ternary phase diagram, shown in Figure 1 left, while a generic scheme of the setup is depicted in Figure 1, right. A solution, consisting of the two substances A and B present in eutectic composition, is saturated at a certain equilibrium temperature (saturation temperature, T) in a feed tank (point E, Figure 1 left). The solution is simultaneously pumped into the bottom of two FBCs, where it is cooled to a crystallization temperature (Tcryst), and thus supersaturation is achieved (point E′). To initiate the crystallization process, seed crystals of both pure substances (A and B) are introduced in the two FBCs respectively, while fresh feed solution is constantly circulated through the crystallizers and back to the feed tank. If the crystallization driving force for both substances is equal, then their ratio in the solution remains eutectic with time. In comparison with batch crystallizers, the fluidized bed crystallization process possesses a couple of advantages. The constant supersaturation, generated in the crystallizers during the whole process, ensures a robust control, stable crystal growth, and production of uniform product crystal sizes. Moreover, the risk of contamination through nucleation of the counter-enantiomer is very low due to the constant upward flow dragging the nuclei formed out of the crystallizer. Also as a result of the upward flow, the crystal mixing and heat transfer in the fluidized bed process are more uniform in comparison with the classical batch processes.10 Received: October 26, 2015 Revised: January 20, 2016 Published: February 3, 2016 1409

DOI: 10.1021/acs.cgd.5b01513 Cryst. Growth Des. 2016, 16, 1409−1419

Crystal Growth & Design

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Figure 1. Left: Ternary phase diagram example of a continuous preferential crystallization process of a simple eutectic system (here two enantiomers A and B), realized isothermally in two parallel coupled fluidized bed crystallizers (A and B − substances in the system; S − solvent; TN, Tcryst, and T − temperature of nucleation, crystallization and saturation, respectively; ΔT − difference of the crystallization temperature from saturation temperature; ΔTmax − the difference between T and TN, defining the metastable zone (shaded area); Right: Scheme of two coupled conical fluidized bed crystallizers (adapted from ref 10.).

Some of the major drawbacks of the fluidized bed process are the need of an increased reactor vessel volume and the use of pumping power to ensure stable crystal fluidization. Moreover, the current lack of understanding the process makes the eventual scale-up from laboratory pilot-plant very difficult. This can lead to higher initial capital costs including longer investigation times.2 The general feasibility of the fluidized bed crystallization process was already described in a previous work.13 It includes studies of the crystal size distribution along the crystallizer height during the crystallization process, along with a systematic study of the application of ultrasonic comminution for continuous generation of seed crystals with regard to its effects on the product crystal size distribution (CSD). In this work, two isomers of aminobenzoic acid and the two enantiomers of asparagine monohydrate were used as model substances. The article is divided into three main sections. In the first section, a simplified mathematical model is proposed, which describes the main features of the process. In the second, preliminary work concerning solid−liquid equilibria in the ternary OABA/PABA/water and L-asn·H2O/D- asn·H2O/water systems (system 1 and 2) is reported. Additionally, metastable zone width data for both systems are presented. On that basis, the third section addresses a continuous preferential crystallization process of OABA and PABA from a 50:50 aqueous solution, as well as L-asn·H2O and D-asn·H2O from their racemic mixture in water, conducted in a fluidized bed regime. It will be shown that pure product crystals with narrow crystal size distribution and high purity can be continuously produced in the steady state operated fluidized bed crystallizers from both ternary systems. Moreover, the solution concentrations, predicted by the model are in both cases in a good agreement with the experimental observations.

Just one external coordinate, x, associated with the axial position in the crystallizer, is applied in a one-dimensional model. In a population balance equation (PBE), the crystals are characterized by a single size coordinate, L. This results in a two-dimensional PBE for the particle size distribution n(x, L). The main effects considered in the PBE are crystal growth, segregation due to different particle sizes, particle comminution in an external ultrasonic bath, and reflux of generated seed particles.14,15 The following assumptions are made for the model formulation: • constant volumetric flow rate V̇ ; • no back-mixing (plug flow) along the crystallizer axis; • spherical crystals of diameter L; • size-independent growth rate of growth order one; • no nucleation, breakage and agglomeration; • isothermal conditions in the crystallizer (T = Tcryst). In Figure 2 is shown a schematic representation of the crystallizer with a total height H and characterized by a partially varying cross-section A. The solution enters with a concentration, cin, at the bottom of the crystallizer (x = 0) and leaves the crystallizer at the top (x = H) with a reduced concentration, cout. The flow rate uf varies in the conical section with x. Seed crystals with a mean diameter, L̅ s,0, are initially introduced at the bottom of the crystallizer (x = 0). A suspension containing product (p) crystals with a mean diameter L̅ p (>L̅s,0) is withdrawn continuously at the product collection position x = xp. The larger crystals positioned at the bottom of the crystallizer (x = 0) and characterized by a mean diameter L̅us are continuously withdrawn, “crashed” externally in an ultrasonic bath (US-bath), and pumped back at the same axial position as seeds of mean size L̅s ( x > xm ⎩

(1)

The solid phase can be described using the particle number distribution of a population of N crystals, n(L, x, t), in a fluidfree differential volume of size (1 − ε)A(x) dx: 1410



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⎛ ∂c ⎞ ⎜ ⎟ ⎝ ∂t ⎠ withdrawal

⎧ 0, for x ≠ xp ⎪ ⎪ ̇ = ⎨ Vp,liq ∂c(x) ⎪ ⎪ εA(x) ∂x , for x = xp ⎩

(6)

The volumetric liquid phase flow leaving the crystallizer at position xp, V̇ p,liq, is considered in eq 6 as a specific fraction α of the inlet flow, i.e.,

̇ = αViṅ Vp,liq

Initially (at t = 0) zero concentrations in the whole crystallizer were assumed. At the entrance the solution concentration is assumed to be the saturation concentration at the feed temperature Tfeed, i.e., cin(x = 0,t) = c*(Tfeed). In the experiments reported below, a specified amount of the substance of interest is additionally provided in the feed tank. After this excess amount is depleted, the feed concentration starts to drop in the arrangement applied (see section 5). Because of the conical section of the crystallizer (d(x) ≤ d(xm) for x ≤ xm, see Figure 2), also the cross section area A(x) used in eq 3 varies with the crystallizer length coordinate. Furthermore, the ratio between solid and fluid phases can vary due to local differences in the mass transfer between both phases. There are clearly two crystallizer sections divided by the product removal point:

Figure 2. Schematic representation of the tubular fluidized bed crystallizer with a conical lower section and continuous product removal.

∂n(L , x , t ) ∂(n) = nseg − nṗ − n→ ̇ −G ̇ us + n us ̇ → ∂t ∂L

⎧ Viṅ , for x < xp ⎪ V (̇ x) = ⎨ ⎪ Vout ̇ = (1 − α)Viṅ , for x > xp ⎩ ̇ = Viṅ − Vp,liq

(2)

where ε is the local fluid (void) fraction, A(x) is the corresponding local cross section area (see below), ṅseg is the flux due to particle segregation along the x-axis, G is the local growth rate, ṅp is the particle flux due to product removal at position x = xp, ṅ→us and ṅus→ are the particle fluxes to and from the US-bath at position x = 0 related to withdrawal of larger crystals and generation of seeds. . The local liquid phase concentration, c, in the differential fluid phase volume of size εA(x) dx can be described by the following partial differential equation πρp ∂(L̅ 3) ∂u ∂c(x , t ) ∂c = c f + uf −N ∂t ∂x ∂x 6εA dx ∂t ⎛ ∂c ⎞ − ⎜ ⎟ withdrawal ⎝ ∂t ⎠

(8)

The total local cross section can be seen as the sum of the contributions of the two phases: A(x) = (π /4)d 2(x) = Aliq (x , t ) + A solid (x , t ) 2

= Aliq (x , t ) + N (x , t)(π /4)L ̅

(3)

ε(x , t ) =

(4)

where N is the total number of crystals:

∫0

A (x )

2

=

A(x) − N (x , t )(π /4)L ̅ A (x )

u f = u pε 2.4

∞

n(L, x , t ) dL

Aliq (x, t )

(10)

(11)

In order to derive the particle flux due to segregation, ṅseg, required in eq 2, in our work the classical empirical RichardsonZaki model16 is used to evaluate quasi-stationary particle positions. This model is based on a force balance considering drag, gravity and buoyancy. It provides for the fluid velocity

∞

N (x , t ) =

V ̇ (x ) Aliq (x , t )

u f (x , t ) =

∫0 n(L , x , t )L dL N (x , t )

(9)

eq 9 and the PBE (eq 2) allow the calculation of the cross section available for the local fluid flow, Aliq(x), the local fluid velocity, uf(x), and the corresponding void fraction, ε(x), according to

with ρp being the solid particle density and L̅ the mean particle size: L̅ (x , t ) =

(7)

(12) 16

where up is the so-called particle terminal velocity. Hereby, the void fraction ε in eq 12 corresponds to the one given in eq 11. The power of 2.4 (Richardson-Zaki constant) holds for the laminar flow conditions present in the crystallizer. If up > uf a particle will sink, and if up > uf a particle will rise. In the laminar flow regime, based on Stokes’ law, can be further stated for up:16

(5)

In eq 3, the first two terms account for convection due to fluid flow. The third term evaluates the decrease of concentration due to particle growth. The last term considers the removal of liquid phase at the product collection position. This term can be formulated as follows: 1411


Crystal Growth & Design u p(L , x) =

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(ρp − ρf )gL2 18μf

(13)

where ρf and μf are the fluid phase density and viscosity. Using the single particle terminal velocity, the particle segregation flux term in eq 2 can be approximated by n seg ̇ (x , t ) =

∂uP(L , x)n(L , x , t ) ∂x

Figure 3. Structural formulas of ortho-aminobenzoic acid, paraaminobenzoic acid, L-(+)-asparagine, and D-(−)-asparagine. Dashed line represents mirror plane.

(14)

Further details regarding the standard submodel related to the terms in eq 2 related to crystal growth, product removal and seed generation are summarized in the appendix. Finally, to complete the model, as a simple (artificial) initial condition it was assumed in the calculations, that at t = 0 a specified number of initial seed crystals, N0, of identical size, Ls,0, is present at position x = 0. It was further assumed, that a particle free liquid phase coming from the feed tank enters the crystallizer at position x = 0. The dynamic model introduced above was solved numerically using simple explicit finite differences to approximate the time and spatial derivatives in Matlab. Axial discretization was bases on the total height of the crystallizer applied (see section 3.2) in mm. Thus, the number of nodes was Nseg = 1100. To achieve sufficient numerical precision, the number of grid points in time was varied to find a compromise between accuracy and calculation time. The model was partly tested in this work to predict transients of the overall solution concentration in the crystallizer as well as corresponding recovered crystal product masses. Furthermore, the model could be applied to evaluate consequences of changes in the feed concentration in case of depletion of the excess of solid mass initially provided. The required basic physical properties of the specific substances (e.g., solid crystal density) can be typically taken from literature (see Table 1). In order to use the proposed model, there are in addition several thermodynamic and kinetic parameters required, which should be determined experimentally. Hereby, parameters like substance solubility are more complex and depend on the physical and chemical properties of the solute and the solvent as well as on temperature and the pH of the solution. It should be mentioned that Mangold et al. described recently an approach for reducing the model type described above using proper orthogonal decomposition (POD) together with a corresponding error estimator.15 This approach allows for faster solution and planned application in optimization and control studies.

substances were above 98% and therefore used without further treatment. The water used for preparing the solutions was deionized. 3.2. Experimental Setup. A pilot plant setup was constructed to study continuous preferential crystallization in conjunction with the fluidization process. A schematic diagram of the experimental facility is shown in Figure 4. Its main component parts are a double jacketed solution reservoir with a volume of 6 l (1); two gear pumps (2); two double jacketed homemade glass filters for product collection (5); and two double jacketed homemade tubular FBCs (6), having a volume of 0.8 L. The conical shape of the FBC ranges from d(x = 0) = 15 mm to d(xm = 515 mm) = 30 mm. Specific values for x correspond to product withdrawal at height xp = 365 mm and height of the crystallizer H = 1100 mm. This position was selected based on preliminary results.13 From the top of one of the crystallizers, part of the solution flow was diverged and send through an ultraviolet detector or polarimeter (in the case of system 1 or 2) (7) and a densitometer (8) for online tracking of the solution concentration and composition. From the bottom of the crystallizers, part of the suspension was constantly transported through a tempered US bath (3), where crystals and crystal agglomerates were comminuted. The resulting smaller crystals and crystal fragments were then mixed with the main inlet flow and again introduced into the bottom of the crystallizers. Homogenization of the feed solution in the feed tank was performed with a 3-blade marine type propeller at a constant rate of 250 min−1 (10). All connecting pipes are thermo-isolated, except those, connecting the top of the crystallizer and the feed tank, which are electrically heated. Water bath thermostats (Lauda RC6 CP, not shown on figure) provide the necessary conditions for the process to be isothermally conducted with temperature control through thermocouples Pt100. 3.3. Methods and Procedures for Solubility Studies. Solubility equilibria of aqueous mixtures of OABA/PABA and L-asn· H2O/D-asn·H2O were determined using an isothermal method. Temperatures of 20, 35, and 50 °C were used in the solubility measurements of aminobenzoic acid isomers and mixtures, while 20 and 40 °C were used in the case of asparagine monohydrate specimens. The reproducibility of the solubility measurements was studied carrying out three experiments under same conditions. Predefined samples of the pure substances and mixtures were inserted in vials, and subsequently 10 g of distilled water was added. The vials were heated well above the expected saturation temperature to completely dissolve the solids and then tempered at the selected temperature under stirring from 24 to 48 h to guarantee a saturated solution with an excess of crystals. After equilibration, 1 mL was taken out from each sample with a syringe through a filter (0.45 μm), diluted in 10 mL distilled water, and analyzed for composition by HPLC (1200 series equipment, Agilent Technologies, Germany). For the aminobenzoic acid system, the following analytical method was applied: column Kinetex reverse C18, 5 μm, 250 mm × 4.6 mm, Phenomenex, Germany, and eluent: 80/20 v/v, acetonitrile/water. HPLC measurements of the asparagine monohydrate system were done with column Astec Chirobiotic T, 5 μm, 250 mm × 4.6 mm, Astec, USA, and eluent: 30/70 v/v, ethanol/water. Solid state phase analysis of the substances used was performed via XRPD measurements. Additionally, samples for XRPD studies were prepared by separation through filtration of aqueous suspensions of OABA, PABA and a 50:50 mixture equilibrated at 20 and 50 °C. The XRPD patterns were measured with an X’Pert Pro diffractometer (PANanalytical

3. EXPERIMENTAL SECTION 3.1. Materials. The materials used in this work are orthoaminobenzoic acid (also anthranilic acid) (OABA), para-aminobenzoic acid (PABA), L-asparagine monohydrate (L-asn·H2O), and Dasparagine monohydrate (D-asn·H2O). Their structural formulas are shown in Figure 3. Aminobenzoic acid is a stereomeric system, showing positional isomerism. It consists of three isomers, which depending on their position on the benzene ring are denoted with prefixes ortho-, meta-, and para-. The optically active molecules (or enantiomers) of asparagine monohydrate are part of the conglomerate-forming D,Lasparagine monohydrate system. For the investigations in this work, OABA and PABA were obtained from VWR and Merck respectively, while L-asn·H2O, D-asn·H2O, and D,L-asparagine monohydrate (D,Lasn·H2O) were purchased from Alfa-Aesar. The purities of the 1412



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Figure 4. Left: Scheme of the fluidized bed crystallization setup: (1) double jacketed feed tank (solution reservoir); (2) gear pump Ismatec MCP-Z; (3) tempered ultrasonic bath Bandelin Sonorex Digital 10P; (4) peristaltic pump Heidolph Pump Drive 5201; (5) double jacketed glass filter; (6) double jacketed tubular fluidized bed crystallizer; (7) ultraviolet detector Knauer, K-2501 or polarimeter IBZ Messtechnik P3002; (8) densitometer Mettler Toledo DE40; (9) vacuum pump ILMVAC MP 601 E; (10) stirrer Heidolph RZR 2021; (11) glass filter ROBU glass, porosity 3; Right: Photograph of the fluidized bed crystallization setup. GmbH, Germany) using Cu Kα radiation and an X’Celerator detector. The patterns were recorded in a 2Θ range of 3−40°, a step size of 0.017°, and a counting time of 50 s per step. 3.4. Determination of Metastable Zone Widths. It is important to determine the width of the metastable zone (MSZW) in order to efficiently avoid undesired nucleation within the crystallization process. Metastable zone widths of aqueous 1:1 mixtures of OABA/PABA and L-asn·H2O/D-asn·H2O were determined using a polythermal method. Specific amounts of the substances were weighted and filled together with 1 g of distilled water and a magnetic stir bar in small vials. The vials were inserted in a Crystal16 multiplereactor system (Avantium, The Netherlands). The turbidity was recorded per individual reactor to detect the appearance of crystals (“cloud point”) in multiple heating runs. Additionally, two series of experiments were conducted in the experimental setup. A 50:50 mechanical mixture of OABA and PABA (System 1) and D,Lasparagine monohydrate (System 2) were dissolved in water and the respective solutions were saturated at 35 °C. After starting the circulation of the solutions in the crystallization setup, the temperature of both crystallizers was set to specific values, lower than 35 °C, while the ultrasonic bath was operated at the saturation temperature. Thus, the influence of the cavitation effect on the formation of nuclei is greatly suppressed. Nucleation was detected via measuring the density of the solutions, while considering the time elapsed since starting the experiment. Nucleation is recognized by a negative deviation of the density measurements. The time of the event is recorded as nucleation time, tN. 3.5. Continuous Preferential Crystallization Runs. Studies of preferential crystallization of the two ternary systems were conducted in conjunction with the dynamic behavior of the fluidization process. Two experiments for each ternary system were conducted. In Table 1 are given selected parameters, used for performing the experiments and for model calculations. Crystal density and crystal growth rates of the pure substances were taken from the literature.17−22 The parameter values for D-asn·H2O are not shown in the table as they are assumed to be identical to the respective values for L-asn·H2O. The ultrasonic bath was operated at 35 kHz, 10% of its maximum output power (i.e., 48 W), and the saturation temperature of the respective ternary system.

Table 1. Parameters Used in the Experiments and Model for the Ternary Systems 1 and 2 experimental conditions

System 1

System 2

saturation temperature, Tfeed, [°C] crystallization temperature, Tcryst, [°C] initial solution concentration, c, [wt %] supersaturation value, k [−] seed amount per crystallizer, ms,0, [kg] seed number per crystallizer, N0, [−] initial median seed size (d50), Ls,0, [m] feed flow rate, V̇ in, [L/h] seed generation flow rate, V̇ us, [L/h] splitting factor α used in eq 7, [−] mean void fraction, ε, [−]b mean residence time of crystals in USbath, τus, [s] solution density at 35 °C, ρf, [kg/m3] solution viscosity at 35 °C, μ, [Pa·s] OABA crystal density,17 ρp, [kg/m3] PABA crystal density,18 ρp, [kg/m3] OABA crystal growth rate,19 G, [m/s] PABA crystal growth rate,20 G, [m/s] 21 3 L-asn·H2O crystal density, ρp, [kg/m ] 22 L-asn·H2O crystal growth rate, G, [m/s]

34.9 26.8 1.52 1.43 8 × 10−3 1.16 × 1010 7.6 × 10−5 10.5 18 0.083 0.78 600

35 27 8.76 1.37 4 × 10−3 6 × 109 5.5 × 10−5 9 4 0.083 0.66 600

994.6 6.53 × 10−4a 1409 1393 1.75 × 10−7 1 × 10−7

1031 8.5 × 10−4 20

1568 1.117 × 10−6

The viscosity of water at 40 °C was used. bValues experimentally estimated in previous work.13

a

In the following, the experimental procedures applied will exemplarily be described for System 1. Experiments with System 2 were performed in an analogous manner. A solution of 50:50 mixture of OABA and PABA crystals saturated at 35 °C was created in the double jacketed, heated reservoir. Homogenization was provided through an electric stirrer at a constant stirring rate. Additional 10 g of each OABA and PABA crystals were added into the reservoir to keep the saturation of the solution over a certain time to maintain a continuous crystallization process. In case of System 2, an additional 1413



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80 g of solid D,L-asn·H2O were added. The solution was then pumped into the bottom of the FBC with the aid of a gear pump at a constant flow rate through a sintered glass filter. The porosity of the filter was chosen such a way that on one side it ensures a constant solution flow, and on the other, all solid particles (>16 μm) remain in the feed tank. The clear saturated solution was cooled in the FBC to generate supersaturation, while flowing from bottom to the top. After the top of the crystallizer was left on, the solution was fed back into the feed tank, where it was saturated again by dissolution of the excess solid. The crystallization process was initiated by introducing 8 g dry seed crystals of each pure substance (4 g in the case of system 2) in the respective crystallizer. During the crystallization process, crystals and crystal agglomerates were continuously taken from an outlet at the bottom side of each crystallizer and transported by a peristaltic pump into a heated ultrasonic bath, where crystal comminution took place through ultrasonic waves. The suspension was afterward pumped back into the FBC together with the feed flow. This continuous seed-generation loop ensures that the crystallization process runs continuously by providing the needed seed crystals and also avoiding potential clogging at the FBC bottom. Product crystals were collected as suspension and filtered via the thermostated filter units, while the liquid filtrate was transported back to the feed tank. For measuring the product CSD, offline measurements were performed using laser diffraction (CILAS 1180L, Quantachrome GmbH & Co., Germany). From the measured CSDs, d50 values were empirically estimated and additionally compared with results from offline light microscope image analyses (Axioscope 2, Carl Zeiss AG).

Figure 5. Upper 10% section of the ternary solubility phase diagram of the system OABA/PABA/water. Isotherms at 20, 35, and 50 °C are shown with thick blue, green, and red lines, respectively. Additionally, the respective colored thin lines represent tie-lines, separating twophase from three-phase zones. Black dots represent initial concentration. Axes are in weight fractions.

The fitted results from solubility and metastable zone width measurements of a 50:50 mixture of OABA and PABA in water (System 1) are shown in Figure 6.

4. RESULTS AND DISCUSSION OF THE SOLUBILITY EQUILIBRIA AND MSZW 4.1. Solubility Equilibria and MSZW of OABA and PABA Mixtures. Solubility studies of pure OABA and PABA in water can be found in the literature. Additionally, it was shown that all isomers of aminobenzoic acid exhibit polymorphic transformations.23−25 The isomers exhibit generally very low solubility in water, which at 50 °C is between 1.2 and 1.4 wt %. The pure isomers, having a hydrophobic benzene core, cannot form strong interactions with water molecules, which could be a reason for their low solubility. On the contrary, the carboxyl- and amino- groups can form relatively strong hydrogen bonding interactions, thus promoting the solvation of the molecules in the water. This is confirmed by He et al., who has published theoretical and experimental studies of water complexes of OABA and PABA.26 On the basis of the isothermal solubility measurements of samples, containing defined amounts of OABA and PABA in water, the ternary solubility phase diagram was constructed with solubility isotherms at 20, 35, and 50 °C, (blue, green, and red lines respectively; see Figure 5). It can be clearly seen from Figure 5, that the system exhibits simple eutectic behavior with eutectic composition at 50%. Because of the nearly identical solubilities of OABA and PABA isomers, the isotherms are almost symmetric around the eutectic. This also means that aqueous solutions of OABA and PABA with 50:50 ratios possess the highest solubility at the eutectic composition. XRPD diffractograms of the samples from the solid phase with 1:1 ratio, taken at 20 and 50 °C, unambiguously show that a mechanical mixture of both substances was present. This means it is possible to preferentially crystallize one isomer in the presence of the other. Additionally, the results from XRPD measurements from the respected crystal samples have confirmed that only the stable polymorphs, type I of ortho-aminobenzoic acid and αpolymorph of para-aminobenzoic acid, have been formed.23,24

Figure 6. Solubility and MSZW of a 50:50 mixture of OABA and PABA in water. The solubility/supersolubility curves are depicted with solid/dashed lines respectively and correspond to the appropriate polynomial equations given.

It can be seen that the metastable zone has been found to be relatively wide, with a ΔT of about 18 K at lower temperatures and 15 K at 50 °C. On the basis of the solubility measurements and the respected metastable zone width, it can be concluded that preferential crystallization experiments could be planned in the temperature region 25−35 °C with a supersaturation ratio up to 1.5. Thus, a relatively high yield might be achieved, while attaining low nucleation probabilities, crystallization only of the stable polymorphs of OABA and PABA, and keeping the solutions stable with time. 4.2. Solubility Equilibria and MSZW of L-asn·H2O and D-asn·H2O Mixtures. Solubilities of pure D-asn·H2O, L-asn· H2O, and D,L-asn·H2O in water can be found in the literature.22,27−29 In this work, isothermal solubility measurements of samples containing defined amounts of D-asn·H2O and L-asn·H2O in water are reported, and the ternary solubility phase diagram of the system D-asn·H2O/L-asn·H2O/water is constructed with solubility isotherms at 20 and 40 °C (blue and red lines respectively; see Figure 7). It can be seen that solubilities of L-asn increase steadily with increasing content of D-asn reaching maximum solubility at the racemic composition that marks the eutectic composition in the system (conglomerate). Moreover, XRPD diffractograms of 1414



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Table 2. MSZW Measurements of Systems 1 and 2, Performed Isothermally in the Experimental Setup under the Influence of Ultrasonic Attenuation in the Seed Loop T (°C)

Tcryst of system 1/2 (°C)

supersaturation of system 1/2 c/c* (−)

tN of system 1/2 (h)

35 35 35

22/20 25/24 28/27

1.81/1.88 1.57/1.55 1.37/1.37

∼2.5/∼1 ∼6.5/∼3 >24

It can be seen from the table that in both systems at high supersaturation values (above 1.8), nucleation took place within a short time (after about 1 and 2.5 h). At moderate supersaturation levels (∼1.5) the solution in both systems remained stable for quite longer time (about 3 and 6.5 h), but still it is not sufficient for a nuclei-free operation of a continuous process. A comparison of the values for the supersaturation (Table 2) with values, recalculated from the laboratory measurements (Figure 6 and Figure 8), reveals that the MSZW data measured in the crystallization setup correspond well with the data from the laboratory small scale measurements. It should be noted, that the formation of nuclei possibly could have happened at earlier time than shown in the table. But due to the constant upflow of the solution in the crystallizer, the newly formed nuclei are taken away with the flow and then dissolved in the feed tank, thus not triggering a detectable concentration change. In our previous work,13the critical size, above which a crystal is not swept away with the flow was estimated. It was proven that at 9 L/h flow rate particles with sizes above 60 μm remain in the crystallizer. The results have shown the preferential crystallization experiments in both systems could be performed without triggering nucleation at a crystallization temperature of 27 °C, while maintaining a supersaturation of 1.43 and 1.37 for the Systems 1 and 2 respectively. Moreover, in the case of nuclei formation, the residence time is not sufficient for them to reach the critical size and remain in the crystallizer.

Figure 7. Upper 20% section of the ternary solubility phase diagram of the system D-asn·H2O/L-asn·H2O/water. Isotherms at 20 and 40 °C are shown with thick blue and red lines, respectively. Additionally, the respective colored thin lines represent tie-lines, separating two-phase from three-phase zones. Green and black dots represent initial concentration at 20 and 40 °C respectively. Literature data iare represented with pink dots.22 Axes are in weight fractions.

samples from the solid phase with racemic ratio, taken at 20 and 40 °C, unambiguously verify that a mechanical mixture of both substances is present. The fitted results from polythermal measurements for determination of the MSZW of D,L-asn·H2O are shown in Figure 8.

5. RESULTS AND DISCUSSION OF THE CONTINUOUS PREFERENTIAL CRYSTALLIZATION RUNS 5.1. Continuous Preferential Crystallization of OABA and PABA (System 1). The discussion of the continuous preferential crystallization of OABA and PABA from their 50:50 aqueous solution will begin with the data, represented in Table 1 and Table 3. The latter summarizes the mass of products obtained at the end of the process, their median crystal size as well as purity achieved. The initial concentration of the 50:50 aqueous solution of OABA and PABA was 1.52 wt % at 34.9 °C, while the saturated solution concentration at 26.8 °C is about 1.00 wt %. Thus, the difference between the two concentrations c* and c is ∼0.52 wt

Figure 8. Solubility and MSZW of D,L-asn·H2O. The solubility/ supersolubility curves are depicted with solid/dashed lines respectively and correspond to the appropriate polynomial equation.

The two curves define an expanding metastable region, where ΔT at concentration of 5 wt % is about 7 K and rises up to 14 K at 23 wt %. This suggests that for the crystallization experiments, relatively high supersaturation levels can be used at higher temperatures. Nevertheless, in highly supersaturated solutions the possibility of spontaneous nucleation as a stochastic process cannot be neglected. For this purpose, a temperature difference of no more than 6 K (or 1/2 of the MSZW) appears to be suitable for eventual crystallization experiments conducted at 35 °C. 4.3. MSZW Determination in the Experimental Setup. Nuclei-free operation is desired for conducting a preferential crystallization providing product crystals with high purity. MSZW is a kinetic parameter and thus it strongly depends on the conditions in the experimental setup (e.g., US waves, upward flow, and long operation time). Hence, measurements of nucleation times were conducted additionally in the experimental setup. A summary of the results is shown in Table 2, where T is the solution saturation temperature, Tcryst the temperature of the crystallizer and tN - the nucleation time determined.

Table 3. Results from the Continuous Preferential Crystallization of OABA and PABA from Their 50:50 Aqueous Solution experimental result

value

excess of solid feed in feed tank (OABA + PABA)a, [g] total product recovered (OABA/PABA), [g] mean product size L̅ p, [μm] product purity OABA/PABA, [%]

10 + 10 11.15/11.59 175 97/97.9

a

Amount of solid feed excess above the solubility limit for maintaining constant feed concentration.

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Figure 9. Left: Experimental (solid line) and predicted (dashed line) steady state cumulative CSDs for initial OABA seed (blue color) and product (red color) crystals at position x = xp. Right: Microscopic photographs of the product crystals of OABA polymorph type I.

%, which corresponds to a theoretical yield of ∼29.4 g for each isomer at the end of the crystallization process. It is calculated by taking into account the amount of solvent used (7460 g) and the excess solid in the feed tank (10 g). From the experiment, the collected mass of each substance (∼11−12 g, Table 3) is less than the theoretical yield. This was expected, as the product outlet is 365 mm above the bottom of the fluidized bed crystallizer, thus forming a volume of about 125 mL filled with crystal suspension. Moreover, the loop connecting the bottom of the crystallizer with the ultrasonic bath has also a volume of 120 mL, again filled with suspension. Thus, this unrecoverable volume is filled with crystal suspension, which is lowering the product recovery. From Table 1 can be seen that the seed generation suspension flow rate is relatively high (18 L/h). This was done due to the fast formation of crystal agglomerates obtained in the tubes to/from US bath and to avoid their settling in the tubing and possibly blocking the suspension flow. According to the higher suspension flow rate, the residence time of the crystals and crystal agglomerates in the US bad is greatly reduced, thus affecting the generation of seeds. Nevertheless, crystals and agglomerates from the bottom part of the crystallizer are actually forced to loop several times in the US bath in order to be broken into smaller crystals. Important requirements of crystallization processes are the quality and the purity of the product crystals. In Figure 9, the experimentally determined crystal size distributions are shown for the initial OABA seeds and the product crystals. The figure also presents as dashed lines the theoretical CSDs calculated from the model described. Additionally for comparison in the figure, microscopic photographs of product crystals are shown. From the microscopic photograph it can be seen that the OABA product crystals are more or less uniform in size and show a relatively narrow CSD (median product size ∼175 μm, Table 3). On the contrary, PABA product crystals were needlelike with mean crystal length of about 350 μm. It should be mentioned that the measured experimental CSD of the product includes the seed crystals from the US loop, seen also in the photograph (Figure 9). In the model, the initially provided seed size distribution was used to initialize the calculations, which underestimates the seed size distribution width. Nevertheless, the resulted product size distribution is rather well located. Both mathematically derived CSDs are too steep in comparison with the experimental ones. Another reason for the discrepancy, seen in Figure 9, could be that the seeds generated in the US-bath are withdrawn theoretically from the crystallizer by the convective flow. This obviously did not occur to this extent in the experiments.

Samples from both OABA and PABA product crystals were finally subjected to HPLC purity analysis. The results show that during the crystallization process crystals with an overall purity of over 97% were produced (see Table 3). In Figure 10 are shown the experimental and calculated overall concentration transients as a function of time for OABA

Figure 10. Experimental and calculated (using the model described in section 2) transient for the overall solution of the aqueous mixture of OABA/PABA in 50:50 ratio in the FBC (position x = 0). Conditions are shown in Table 1.

and PABA preferential crystallization. A run is shown, in which a feed concentration drop occurred. It can be seen from the experimental curve, that approximately after 3 h the overall concentration of the feed solution dropped as well as the supersaturation in the crystallizer due to depletion of the excess solid phase placed initially in the feed tank. After the solution concentration in the crystallizer, c, has equaled the saturation concentration of the substance at Tcryst, c*, the crystallization process stopped. It can be seen that the predicted overall concentration transient is in fairly good agreement with the experimental one. From the beginning of the process until 2.4 h both transients are constant, as the excess of solid in the feed tank is not yet dissolved. After this stationary process regime, the calculated concentration transient started to decrease, while the experimental one remained still constant for another 40 min. Possible reasons could be uncertainties in the solubility description and erroneous quantification of the seeds generated in the model. Moreover, the removal of product crystals was implemented continuously, thus making an impact on the solid concentration in the crystallizer and directly influencing the 1416



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Figure 11. Transients of the total solution concentration and relative optical rotation for the preferential crystallization experiment. Solution measurements, where the courses of L-asn·H2O and D-asn·H2O are shown in blue and red color, respectively. The time the seeds were added is designated as t = 0.

asparagine monohydrate in the feed tank is equal to the crystallization rate in the crystallizers. The two absolute values for the optical rotation are very close to zero, just with the opposite sign. This means that a very small excess of the counter-enantiomer is detected in the mother liquor and the solution composition remains close to racemic. This is an advantage over the classical batch crystallization process. The almost constant optical rotation transients confirm that the enantiomer ratio remains stable with time. Every hour a 150 mL fraction of the product suspension was collected from the product outlets of the crystallizers. After 220 min the excess racemic solid in the feed tank was completely depleted and the supersaturation started to decrease, thus lowering the crystallization rate. After further 60 min the crystallization process was stopped (at 4.7 h in Figure 11) and remaining crystals above the product outlet were collected. The crystallized product recovery was in total 95.2 g (46.2 g of Lasn·H2O and 49 g of D-asn·H2O). An overall productivity of 28 g/(L·h) (or 14 g/(L·h) per enantiomer) was calculated taking into account the steady state operation of the crystallization process. From Table 1 can be seen that the seed generation flow rate is relatively low compared to the one used for preferential crystallization of OABA and PABA (4 L/h in comparison with 18 L/h, respectively). In the asparagine case almost no agglomerates were formed and no crystal settling was detected in the tubing. Thus, the residence time of the bigger crystals and agglomerates in the US bath is sufficient for seed generation. Microscopic photographs of the collected L-asn·H2O product crystals as well as the experimentally determined crystal size distributions for L-asn·H2O seeds and both L-/D-asn·H2O product crystals are shown in Figure 12 together with the theoretically derived cumulative CSDs (model described in section 2). Obviously besides product crystals, a small portion of seed crystals is noticed, as already found in the case of System 1 (see Figure 9). The CSD of the product crystals shows that crystals with mean size of ∼180 μm are collected. Figure 13 presents the measured and the predicted solution concentration transients with time. Analogous to System 1, the point of the drop in feed concentration is not fully accurately predicted, although the kinetic trend is rather well represented. The predicted solution concentration transient is in good agreement with the experimentally measured one. From the beginning of the process until roughly 3.5 h the concentration is constant, as there is still an excess of solid feed available allowing the realization of a continuous crystallization process. After this time the concentration starts to decrease due to depletion of the excess solids in the feed tank. After approximately 5 h the crystallization process was stopped. The results from the HPLC purity analysis of the product enantiomer samples showed that the purity exceeded 97%.

After addition of 4 g of seed crystals of the pure enantiomers into the respective crystallizer, the crystallization process was initiated, indicated by a small change in the solution density (i.e., concentration), detected in both crystallizers as seen in Figure 11. Afterward, a steady state operation was observed confirmed by both analytical circles, which show nearly the same values for the solution concentration. The constancy of the solution concentration transients verifies that the dissolution rate of the excess solid D,L-

6. CONCLUSIONS AND OUTLOOK An attractive concept for continuously separating isomers by preferential crystallization was studied experimentally and theoretically. Additionally, a fundamental study of solubility equilibria and metastable zone widths in the ternary systems of OABA/PABA and L-asn·H 2 O/ D -asn·H 2O in water was conducted. The resulting solubility phase diagram of the OABA/PABA/water system indicates that it forms a simple eutectic based on the shape of the solubility isotherms, the

solution concentration. To reduce significantly the calculation times, it should be mentioned that a mean value for the void fraction was used in the calculation, thus leading to further discrepancies with the experiments. After depletion of the excess solids in the feed tank, both concentration trajectories (experimental and calculated) show almost the same behavior until a concentration of about 1.05 wt % indicating a certain reliability of kinetic parameters and models. At this point the shift to the lower concentration values from the experiment could be eventually due to clogged filter, thus contaminating the density measurement. After 6 h the crystallization process is finished with only fluidization running. The excess solid mass in the feed tank was dissolved in the first 3 h of the process, which recalculated per hour and per volume crystallizer gives a productivity of 4.27 g/(L·h) per isomer. 5.2. Continuous Preferential Crystallization of L-asn· H2O and D-asn·H2O (System 2). In Table 1 are given the start parameters of the continuous preferential crystallization of L-asn·H2O and D-asn·H2O from an aqueous solution of racemic D,L-asn·H2O. The results from the experiments are summarized in Table 4. The transients of the solution concentration and the Table 4. Results from the Preferential Crystallization Experiments excess of solid racemic feed D,L-asn·H2O in tank, [g] total product recovered L-asn·H2O/D-asn·H2O, [g] mean product size L̅ p (both enantiomers), [μm] product purity L-asn·H2O/D-asn·H2O, [%]

80 46.2/49 184 97/97.9

relative optical rotation, measured in both crystallizers, are presented in Figure 11. The letters L and D in the figure indicate the crystallizer in which L-asn·H2O and D-asn·H2O are crystallized, respectively.

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Figure 12. Left: Experimental (solid line) and predicted (dashed line) steady state cumulative CSDs for L-asn·H2O seed (blue color), L-/D-asn·H2O product crystals (red/black color respectively). Right: Microscopic photographs of the L-asn·H2O product crystals.

lization was successfully conducted for two exemplary systems forming simple eutectics in their solubility phase diagrams.

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ASSOCIATED CONTENT

S Supporting Information *

The Supporting Information is available free of charge on the ACS Publications website at DOI: 10.1021/acs.cgd.5b01513. An appendix which elaborates details and submodels concerning the mathematical model (PDF)

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AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. Notes

The authors declare no competing financial interest.

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Figure 13. Experimental and calculated (model described in section 2) transients for the overall solution concentration of the DL-asparagine monohydrate solution in the FBC (position x = 0). Conditions are shown in Table 1.

ACKNOWLEDGMENTS The authors thank Dr. Stefan Palis and Prof. Achim Kienle from Institute for Automation Technique, Otto-von-Guericke University in Magdeburg for their support in the mathematical modeling. This project is supported by the German Research Foundation within the Priority Program “Dynamische Simulation vernetzter Feststoffprozesse” (DFG-SPP 1679).

course of the corresponding tie lines, and particularly on the measured XRPD diffractograms of the resultant solid phase. On the basis of the solubility and MSZW data measured, continuous preferential crystallization in two coupled fluidized bed crystallizers was conducted for both model systems. The resultant product crystals featured high purities (>97%) and sustained a stable and narrow crystal size distribution throughout the duration of the process. Product crystals were collected with median sizes from 175 μm (OABA) to 350 μm (PABA) and about 180 μm in the case of asparagine enantiomers. The productivity of the crystallization process was found under steady state conditions to be 4.27 g/(L·h) for the aminobenzoic acid isomers, and 14 g/(L·h) for the asparagine enantiomers. In the crystallization of aminobenzoic acid isomers, only polymorph I for OABA and α-form for PABA were produced, and no polymorph transitions were detected. A simplified dynamic mathematical model was developed. It was partially validated evaluating solution concentration transients and CSDs. The predicted product CSDs are in rather good agreement with the experimentally measured ones. Further systematic testing and validation is still pending. The constructed fluidized bed crystallization setup, consisting of two connected crystallizers, is capable of continuous production of uniform crystals with specific size and high purity. Continuous isomer separation by preferential crystal-

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REFERENCES

(1) Myerson, A. S.; Handbook of Industrial Crystallization, 2nd ed.; Elsevier: Oxford, 2001. (2) Hoffmann G. VDI-Berichte 1901 2005, Bd. 2, 759−772. (3) Jacques, J., Collet, A., Wilen, S. H.; Enantiomers, Racemates, and Resolutions; Krieger Publishing Company: Malabar, FL, 1994. (4) Kaemmerer, H.; Seidel-Morgenstern, A.; Lorenz, H. Chem. Eng. Process. 2013, 67, 71−79. (5) Swernath, S.; Kaspereit, M.; Kienle, A. Chem. Eng. Technol. 2013, 36, 1417−1429. (6) Xie, R.; Chu, L.−Y.; Deng, J.−G. Chem. Soc. Rev. 2008, 37, 1243− 1263. (7) Gou, L.; Robl, S.; Leonhard, K.; Lorenz, H.; Sordo, M.; Butka, A.; Kesselheim, S.; Wolff, M.; Seidel−Morgenstern, A.; Schaber, K. Chirality 2011, 23, 118−127. (8) Wattanakit, C.; Côme, Y. B. S.; Lapeyre, V.; Bopp, P. A.; Heim, M.; Yadnum, S.; Nokbin, S.; Warakulwit, C.; Limtrakul, J.; Kuhn, A. Nat. Commun. 2014, 5, 1−8. (9) Galan, K.; Eicke, M. J.; Elsner, M. P.; Lorenz, H.; SeidelMorgenstern, A. Cryst. Growth Des. 2015, 15, 1808−1818. (10) Tung, H.-H., Paul, E. L., Midler, M., McCauley, J. A. Crystallization of Organic Compounds: An Industrial Perspective; Wiley: Hoboken, NJ, 2009. (11) Midler M. United States Patent No. 3892539, 1975. (12) Midler M. United States Patent No. 3996018, 1976. 1418



Article

(13) Binev, D.; Lorenz, H.; Seidel-Morgenstern, A. Chem. Eng. Sci. 2015, 133, 116−124. (14) Palis S., Binev D., Lorenz H., Seidel−Morgenstern A., Kienle A. Proceedings of 20th International Workshop on Industrial Crystallization (BIWIC), Odense, Denmark, 2013; pp 252−259. (15) Mangold, M.; Feng, L.; Khlopov, D.; Palis, S.; Benner, P.; Binev, D.; Seidel-Morgenstern, A. J. Comput. Appl. Math. 2015, 289, 253− 266. (16) Richardson, J. F.; Zaki, W. N. Trans. Inst. Chem. Eng. 1954, 32, 35−53. (17) Brown, C. J.; Ehrenberg, M. Acta Crystallogr., Sect. C: Cryst. Struct. Commun. 1985, 41, 441−443. (18) Groth, P. Chem. Krystallogr. 1908, 4, 509. (19) Temmel E., Eisenschmidt H., Lorenz H., Seidel-Morgenstern A. Chem. Eng. Sci., to be submitted. (20) Lin, S. W.; Ng, K. M.; Wibowo, C. Ind. Eng. Chem. Res. 2007, 46, 518−529. (21) Lindenberg, C.; Mazzotti, M. AIChE J. 2011, 57, 942−950. (22) Petrusevska-Seebach, K.; Seidel-Morgenstern, A.; Elsner, M. P. Cryst. Growth Des. 2011, 11, 2149−2163. (23) Jiang, S.; Jansens, P. J.; ter Horst, J. H. Cryst. Growth Des. 2010, 10, 2541−2547. (24) Gracin, S.; Rasmuson, Å.C. Cryst. Growth Des. 2004, 4, 1013− 1023. (25) Svärd, M.; Nordström, F. L.; Jasnobulka, T.; Rasmuson, A. C. Cryst. Growth Des. 2010, 10, 195−204. (26) He, Y.; Wu, C.; Kong, W. J. Phys. Chem. A 2005, 109, 2809− 2815. (27) Elgarhy K. Ph.D. Thesis, McGill University: Canada, 2006; ISBN 9780494251386. (28) Orella, C. J.; Kirwan, D. J. Ind. Eng. Chem. Res. 1991, 30, 1040− 1045. (29) Dalton, J. B.; Schmidt, C. L. A. J. Biol. Chem. 1935, 109, 241− 248.

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