Limits of Operation for the Integration of Water Removal by

Water removal and crystallization were first studied separately, and models were used to define the limits of operation for the coupled system so that...
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Ind. Eng. Chem. Res. 2009, 48, 1566–1573

Limits of Operation for the Integration of Water Removal by Membranes and Crystallization of L-Phenylalanine Maria C. Cuellar, Simone N. Herreilers, Adrie J. J. Straathof,* Joseph J. Heijnen, and Luuk A. M. van der Wielen Department of Biotechnology, Delft UniVersity of Technology, Julianalaan 67 2628BC, Delft, The Netherlands

Integration of crystallization and water removal using membrane technology is being used as a means to improve control over supersaturation generation and, consequently, to improve crystal quality. However, it is not clear yet how water removal, which has a limited window of operation, can be combined with crystallization in case of occurrence of several crystal forms (and, therefore, also a limited window of operation). In this work we evaluated the use of nanofiltration and reverse osmosis for water removal to achieve selective crystallization of L-phenylalanine anhydrate from aqueous solution. Water removal and crystallization were first studied separately, and models were used to define the limits of operation for the coupled system so that the target crystal form and production rate could be achieved. A narrow window of operation was obtained, and it was shown that practical implementation can be difficult due to operation near the limits of the window of operation. Possibilities for widening the window of operation are briefly discussed. Introduction Industrial fermentation processes demand high product concentrations in order to be commercially interesting. However, factors such as product toxicity to the microorganism or product degradation result in a low product concentration, which in turn leads to a more complex or inefficient product recovery. One way of circumventing these issues consists in removing the product from the fermentation medium as soon as it is being formed. This approach, known as in situ product removal or in situ product recovery (ISPR), has been described in the literature with techniques such as adsorption, extraction, and membrane technology.1 Recently, product removal by crystallization has received attention for products that are commercialized in crystal form.2 In such cases, product removal by crystallization might lead not only to higher fermentation productivities (by reducing product toxicity and/or degradation) but also to simpler downstream operations (by recovering the product already in crystal form), without requiring the use of auxiliary materials. We are studying the applicability of product removal by crystallization during the fermentative production of L-phenylalanine (Phe). Phe is one of the most important commercially produced amino acids, being used as a feed and food supplement and as precursor for the production of the artificial sweetener aspartame. The fermentative production of Phe, however, suffers from product inhibition. Since Phe is undersaturated at fermentation conditions, product removal by crystallization can be done by means of an external crystallization loop. A water removal step is needed to generate the required supersaturation for crystallization; next to this, in order to maintain a low Phe concentration in the fermenter (and hence minimize product inhibition), the crystallization rate should equal the production rate by the microorganism. Phe is known to crystallize in two forms: a flake-like anhydrate and a needle-like monohydrate. Anhydrate is the commercial form, and it is easier to handle. The monohydrate, however, is the initial form obtained by spontaneous nucleation in water.3 The monohydrate may transform into the anhydrate at temperatures where the latter is thermodynamically more * To whom correspondence should be addressed. Tel.: +31-15-278 2330. Fax: +31-15-278 2355. E-mail: [email protected].

stable (>37 °C).4 This transformation process can take several hours depending on the process conditions. To overcome this problem, seeding anhydrate crystals within the metastable zone at temperatures above 37 °C has been recommended.3 In crystallization, the rate and the extent of supersaturation generation have a direct effect on the quality of the obtained crystals. It is well-known that uncontrolled supersaturation can lead to undesirable crystal size distributions (CSDs) which in turn will negatively affect the process performance after crystallization (for example during filtration, washing, drying, etc).5 Regarding the impurity content, crystallization at low supersaturation usually results in better crystal quality. Furthermore, for compounds exhibiting polymorphism, formation of solvates, and/or formation of hydrates, the conditions for obtaining the desired form spontaneously or by seeding might result in a limited supersaturation range for operation.6 In the case of Phe, the metastable zone width (MZW) was found to be rather narrow3 and, therefore, control over the extent of supersaturation generation is of great importance for obtaining the target anhydrate crystals. When designing a crystallization process, one way to increase the control over supersaturation generation is to uncouple this task from other crystallization tasks like nucleation and growth. Within this frame, the use of membranes has received particular interest. Lee et al.7 and Zeelen and Wierenga8 described the use of respectively reverse osmosis (RO) and dialysis membranes for the crystallization of biological molecules at the microscale. In a larger scale, the use of membranes linked to crystallization has been exploited for water treatment applications. Sluys et al.9 described the use of microfiltration (MF) for increasing the concentration of the solution up to the level required for seeded crystallization in the softening of drinking water. Drioli et al.10 used microporous hydrophobic membranes for membrane distillation (MD) during seawater desalination. In MD, the solvent evaporation occurs inside the membrane module (where the flowing solution is below the supersaturation condition) and the crystallization stage is performed in a separate tank on the retentate line.11 This concept has been further described for the crystallization of inorganic salts,11,12 for the removal of unreacted fumaric acid during an enzymatic reac-

10.1021/ie8012659 CCC: $40.75  2009 American Chemical Society Published on Web 01/08/2009

Ind. Eng. Chem. Res., Vol. 48, No. 3, 2009 1567

modeled as one well-mixed unit (dashed box in the figure). From this point on this unit will be referred to as the buffer vessel. The accumulation of total mass in the buffer vessel (under negligible evaporation) is dMb ) Φf - Φp - Φc,in dt The accumulation of Phe in the buffer vessel is

(1)

dMPhe,b aq aq aq ) Φf CPhe,f - ΦpCPhe,p - ΦcCPhe,b (2) dt aq is defined as The Phe mass fraction in the buffer vessel CPhe,b aq CPhe,b )

MPhe,b Mb

(3)

In the crystallizer, Phe crystals are produced and retained in the crystallizer (there is no flow of Phe crystals to and from the crystallizer). At this point, the crystals are assumed to have 100% purity and evaporation is assumed negligible. The accumulation of total mass in the crystallizer is Figure 1. Schematic representation of the coupled system for water removal and crystallization. 13

14

tion, and for the crystallization of lysozyme. A similar concept using a sweeping gas instead of water in the distillate line was suggested by Konig and Weckesser15 for NaCl crystallization. Zarkadas and Sirkar16 reported the use of a hollow fiber module containing a nonporous polymeric membrane for the cooling crystallization of KNO3 and salicylic acid. The applicability of a hollow fiber module for antisolvent crystallization was evaluated with a porous membrane for the production of L-asparagine monohydrate crystals.17 Van der Gun and Bruinsma18 used a nanofiltration (NF) membrane for concentrating a stream prior to crystallization during mineral processing. In many of these examples, the coupling of membrane technology and crystallization resulted in improved control over supersaturation generation, and consequently, over the product quality (e.g., CSD). However, it is not clear yet how water removal, which has a limited window of operation, can be combined with crystallization in case of occurrence of several crystal forms (and, therefore, also a limited window of operation). The aim of this paper is to provide a systematic approach for the selection of the conditions of operation for the crystallization of Phe anhydrate from aqueous solution using membrane technology for supersaturation generation. Two membranes, NF and RO, are compared in performance; next to this, a basic model is developed and used for designing the system. We provide an example in which the objective is to achieve a desired crystal production rate, as would be the case for its application within a concept for product removal by crystallization during fermentation. This example is also experimentally evaluated in order to assess the practical feasibility of the concept. Nevertheless, the approach developed here is also applicable to other compounds and in crystallization processes in which supersaturation generation and crystal growth need to be uncoupled. Theory and Model Development 1. Mass Balances. Figure 1 gives a schematic representation of a water removal unit coupled to a crystallizer. Several recycle configurations are possible but will not be considered in this work. The flow to and from the membrane is very large (kg min-1) compared to the feed and the flow to the crystallizer (g min-1). Therefore, the buffer vessel and the membrane are

dMc ) Φc,in - Φc,out dt The accumulation of aqueous mass in the crystallizer is

(4)

dMcaq s ) Φc,in - Φc,out - pPhe,c (5) dt where psPhe,c is the crystal production rate in the crystallizer. The accumulation of aqueous Phe in the crystallizer is aq dMPhe,c aq aq s ) Φc,in CPhe,b - Φc,out CPhe,c - pPhe,c dt

(6)

aq is defined as where the aqueous Phe mass fraction CPhe,c

aq CPhe,c )

aq MPhe,c

(7)

Mcaq

2. Membrane Equations. The mass transfer across an NF or an RO membrane can be described by the solution/diffusion model.19,20 In our system the solvent is water. The water flux Jw is dependent on the water permeability of the membrane, Aw, and the net transmembrane pressure: Jw ) Aw(∆PTM - ∆ΠTM)

(8)

In the system used for this study, the permeate is always at atmospheric pressure. Furthermore, the pressure drop within the module is negligible; therefore, ∆PTM is equal to the applied pressure, which is measured at the retentate line (Pr). The transmembrane osmotic pressure ∆ΠTM is calculated as the difference between the osmotic pressure at the membrane, Πm, and the osmotic pressure of the permeate, Πp. The osmotic pressure is calculated from the van’t Hoff equation: Π ) RTF

aq CPhe WPhe

(9)

where R is the universal gas constant, T is the solution temperature, F is the solution density (in our case comparable aq to that of water), CPhe is the mass fraction of Phe (at feed or permeate side of the membrane), and WPhe is the molecular mass of Phe. Equation 8 can be rewritten as

(

Jw ) Aw Pr -

)

RTF aq aq (C - CPhe,p ) WPhe Phe,m

(10)

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In our system other solutes are negligible; hence, in addition to water, only the transport of zwitterionic Phe is considered. The flux of Phe is controlled by diffusion only: aq aq - CPhe,p ) JPhe ) BPhe(CPhe,m

(11)

where BPhe is the mass transfer coefficient across the membrane. Due to polarization effects there is an additional mass transfer resistance in the boundary layer at the concentrate side of the membrane. The mass transfer of Phe across the polarized layer can be described by20

( )

aq aq aq aq CPhe,m - CPhe,p ) (CPhe,b - CPhe,p ) exp

Jw kPhe

(12)

where kPhe is the mass transfer coefficient in the polarized layer. Equations 10 and 11 thus become aq aq - CPhe,p ) JPhe ) B Phe(CPhe,b

(

Jw ) Aw Pr -

( ) ( ))

Jw exp kPhe

Jw RTF aq aq (CPhe,b - CPhe,p ) exp WPhe kPhe

(13) (14)

dimensions, L, using the aspect ratio of the other dimensions s can be defined as with respect to L. nPhe,c,0 and mPhe,c nPhe,c,0 )

s MPhe,c,0 s s VPhe,0 FPhe

)

s MPhe,c,0

(20)

s L03(AR1)(AR2)FPhe

s s s s mPhe ) VPhe FPhe ) L3(AR1)(AR2)FPhe

(21)

s s is the initial mass fraction of crystals and FPhe is where MPhe,c,0 the density of the Phe crystals. Equation 19 thus becomes

s pPhe,c )

s 3MPhe,c,0 L2 dL dt L03

(22)

From eq 20 it also follows that

() L L0

2

)

( ) s MPhe,c

2/3

(23)

s MPhe,c,0

The growth rate dL/dt can be expressed as a function of the supersaturation S:

(

) (

)

(15)

aq aq g g CPhe,c MPhe,c dL g ) kg(S - 1) ) kg aq - 1 ) kg aq aq -1 dt CPhe,c,eq Mc CPhe,c,eq (24)

The selectivity of the membrane is conventionally measured as the retention (also called rejection) coefficient:20

aq where CPhe,c,eq is the mass fraction of Phe at equilibrium (i.e., the solubility). Substitution of eqs 23 and 24 in eq 22 yields

The flux of Phe is calculated from the water flux and the permeate composition: aq JPhe ) JwCPhe,p

RPhe ) 1 -

aq CPhe,p aq CPhe,b

(16)

Combining eqs 13, 15, and 16 and solving for RPhe give RPhe )

Jw Jw BPhe exp + Jw kPhe

( )

(17)

In eq 17 RPhe is only dependent on the water flux Jw and the mass transfer coefficients BPhe and kPhe. For a given membrane BPhe is a constant. Furthermore, in our system the cross-flow velocity is approximately constant for all experiments. Thereaq fore, kPhe is assumed constant. For a set of experiments CPhe,b aq and CPhe,p, and hence RPhe, are determined at various values of Jw. Equation 17 is thus used for data fitting, from which the values of BPhe and kPhe are obtained. Data fitting was done by nonlinear regression with the Marquardt method21 using Matlab 7.3 (Mathworks, USA). The water permeability Aw can be derived from clean water flux measurements. Equation 14 becomes Jw ) AwPr

(18)

3. Crystallization Equations. For a seeded, batch crystallization, under the assumptions of negligible (on mass basis) nucleation, breakage, agglomeration, size-independent growth, and narrow size distribution of the seeds, the mass balance for the crystals becomes aq s dMPhe,c dmPhe,c ) nPhe,c,0 (19) dt dt s where nPhe,c,0 is the number of seed crystals and mPhe,c is the mass of a single crystal. The Phe anhydrate crystals are flakes. Assuming that the growth rate is proportional to each of the three dimensions, the aspect ratios remain constant. In this way, the crystal volume can be expressed in terms of one of the s )pPhe,c

s s pPhe,c ) 3(MPhe,c )2/3

(

)

s aq (MPhe,c,0 )1/3 MPhe,c kg aq aq -1 L0 Mc CPhe,c,eq

g

(25)

The kinetic parameters kg and g can be obtained from batch experiments by following the Phe mass in the solution in time and fitting the data to eqs 6 and 25. Data fitting was done by nonlinear regression with the Marquardt method21 using Matlab 7.3 (Mathworks, USA). 4. Steady-State Equations. A constant crystal production aq rate can be achieved by keeping CPhe,c constant at steady state: aq dCPhe,c )0 dt s Differentiating eq 7 and solving for pPhe,c give

s pPhe,c ) Φc,in

aq aq CPhe,b - CPhe,c aq 1 - CPhe,c

(26)

(27)

aq In the buffer vessel CPhe,b and Mb should also remain constant:

dMb )0 dt dMPhe,b )0 dt Equations 1 and 2 thus become

(28) (29)

Φf ) Φp + Φc

(30)

aq aq aq ) ΦpCPhe,p + ΦcCPhe,b ΦfCPhe,f

(31)

The permeate flow can also be expressed as Φp ) JwA

(32)

where A is the membrane area. Substituting eqs 32 and 30 in eq 31 and solving for Φc give

Ind. Eng. Chem. Res., Vol. 48, No. 3, 2009 1569 aq aq CPhe,f - CPhe,p Φc ) JwA aq aq CPhe,b - CPhe,f

(33)

The concentration factor in the buffer vessel is defined as F)

aq CPhe,b aq CPhe,f

(34)

Substituting eqs 16 and 34 in eq 33 yields Φc ) Jw A

1 - (1 - RPhe)F F-1

(35)

Materials and Methods 1. Chemicals. L-Phenylalanine solutions were prepared by dissolving L-phenylalanine of purity >98% (batch no. 114K0140, Sigma-Aldrich, and lot no. 1357967, Fluka) in demineralized water. The pH of the solutions was adjusted to 6.5 by addition of 6% NH4OH. The seeds for crystallization were prepared by sieving and collecting the fraction between 100 and 202 µm. 2. Membrane Characterization. Two commercially available thin-film composite membranes were tested: a nanofiltration membrane (reference HL) and a reverse osmosis membrane (reference SE; both from GE Osmonics Inc., Minnetonka, MN). Both membranes were characterized with demineralized water and Phe solutions of mass fractions between 15 and 45 g kg-1. The solutions were kept in a 2 L jacketed vessel, and the solution temperature was maintained at 50 °C with a Lauda RE 307 water bath equipped with a Pt 100 sensor (Lauda-Ko¨nigshofen, Germany). The flat-sheet membranes, with an effective area of 140 cm2, were tested in a Sepa CF II Membrane Cell System (GE Osmonics Inc., Minnetonka, MN). The solution was pumped from the feed vessel to the membrane cell with a stainless-steel plunger pump (Model 3CP1231 direct drive, Cat Pumps, Minneapolis, MN). The retentate flow was measured by an Endress-Hausser electromagnetic flow meter (Reinach, Switzerland). The system was operated under full recycle to ensure a constant Phe mass fraction in the feed vessel. Pressure was built up in the system by closing a valve in the retentate line. After about 10-30 min of operation under a given pressure setting, samples from the feed vessel and the permeate were taken for mass fraction analysis. The permeate flow was measured gravimetrically. 3. Crystallization Experiments. Batch crystallization experiments were carried out using a 1.5 L jacketed vessel with a three-bladed propeller stirrer of 4.5 cm diameter operating at 180 rpm. The solution temperature was controlled by a Lauda RE 307 water bath equipped with a Pt 100 sensor (LaudaKo¨nigshofen, Germany). Approximately 0.5 kg of solution was brought to 45 °C. The seeds (about 0.35 g) were added as slurry prepared with about 4.5 mL of water per gram of seeds. Samples were taken at regular intervals using a syringe with a 0.2 µm filter attached to it. For each sample, a new syringe and filter were used. A continuous crystallization experiment was carried out in the same vessel, this time equipped with inlet and outlet ports. A stainless-steel microfilter was placed at the outlet to ensure that crystals would remain in the crystallizer. The feed solution (about 45 g kg-1) was kept in a 2 L jacketed vessel with three baffles and two four-bladed turbine stirrers of 4.5 cm diameter operating at about 200 rpm. The solution temperature was maintained at 50 °C by a Lauda RE 307 water bath equipped with a Pt 100 sensor (Lauda-Ko¨nigshofen, Germany). The flow to and from the crystallizer was fixed at 3 g min-1 using a

Figure 2. Solvent flux Jw as a function of pressure Pr for (a) nanofiltration and (b) reverse osmosis membranes. Symbols are experimental results at aq different CPhe,b values (in g kg-1).

Masterflex peristaltic pump. The tubing from the feed vessel to the crystallizer and from the crystallizer to the mother liquor storage vessel was jacketed and heated by means of a Lauda 003 water bath (Lauda-Ko¨nigshofen, Germany). Heating of the tubing was needed to prevent crystallization in the lines. The initial mass fraction in the feed vessel was the same as in the crystallizer. After the lines had been filled with solution, 0.5 g of seeds was added in the same way as described above. Samples were taken at regular intervals from the crystallizer outlet line. 4. Coupled Water Removal and Crystallization. The feed solution (about 20 g kg-1) was kept in the same vessel as for the continuous crystallization experiments described above. The vessel used for the membrane characterization experiments was used as the buffer vessel. The buffer vessel and the lines to and from the membrane module were prefilled with about 1 kg of the feed solution. The temperature in the buffer vessel was maintained at 50 °C. The crystallizer described in the crystallization experiments was prefilled with about 0.5 kg of solution (about 43 g kg-1), and the temperature was maintained at 45 °C. The experiment was started by setting the retentate pressure at 2600 kPa under full recycle for about 1 h to ensure that the permeate flow was stable and that all lines were filled. The feed solution was then pumped to the buffer vessel at about 6 g min-1, and the permeate was collected in a separate vessel. When the target mass fraction in the buffer vessel (about 45 g kg-1) was reached, the feed rate was increased to about 9 g min-1 and the flow to and from the crystallizer was started (about 5 g min-1). The mother liquor was collected in a separate vessel. In the same way as for the continuous crystallization experiments, the tubing to and from the crystallizer was heated to prevent crystallization in the line. After the lines had been filled with solution, seeds were added in the same way as described above. During the experiment, samples for mass

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Figure 5. Mass fraction/temperature window of operation for seeded crystallization of Phe anhydrate in water in the anhydrate stability region.

Figure 3. Retention RPhe as a function of solvent flux Jw for (a) nanofiltration and (b) reverse osmosis membranes. Symbols are experimental results at aq different CPhe,b values (in g kg-1); the solid line was calculated from eq 17. Table 1. Fitted Parameters for the Membrane Model BPhe [kg m-2 s-1] kPhe [kg m-2 s-1]

NF membrane

RO membrane

(3.26 ( 0.14) × 10-3 0.099 ( 0.025

(3.32 ( 0.47) × 10-5 0.0099 ( 0.0041

fraction analysis were taken from the buffer vessel, the permeate, and the crystallizer outlet line. The permeate flow was measured gravimetrically at regular intervals. Clean water flux was measured before and after the experiment using demineralized water at 50 °C. 5. Analyses. The mass fraction of Phe in solution was determined by spectroscopy at 258 nm. After each crystallization run the crystals were harvested by vacuum filtration using a

Figure 6. Calculated steady-state profiles for the RO membrane: (a) required aq Pr as a function of Jw at 50 °C; (b) required CPhe,c as a function of Jw; metastability limit at 45 and 40 °C (lines 1 and 2, respectively) and anhydrate solubility at 45 and 40 °C (lines 3 and 4, respectively).

0.2 µm filter. The crystals were dried for about 48 h at 70 °C. Crystal form was confirmed by powder X-ray diffraction (XRD). Results and Discussion

Figure 4. Desupersaturation profile for a continuous crystallization with crystal retention at 45 °C. Symbols are experimental data, and the solid line is the best model prediction (kg ) 2.0 × 10-6 m s-1 and g ) 2.1). See text for further details.

1. Membrane Characterization. Figure 2 shows the permeaq ate flux as a function of (retentate) pressure for several CPhe,b values. Figure 3 shows the Phe retention as a function of permeate flux used for deriving the parameters for the membrane model (see Table 1). As expected, the RO membrane has a much lower permeate flux and Phe mass transfer coefficient BPhe than

Ind. Eng. Chem. Res., Vol. 48, No. 3, 2009 1571 Table 2. Limits of Operation (Jw and Pr) and Percentage of Phe Lost in Permeate at Steady State for Both Membranes at Two Crystallization Temperaturesa

NF membrane RO membrane

temperature [°C]

min Jw [kg m-2 s-1]

min Pr [kPa]

Phe loss [%]

max Jw [kg m-2 s-1]

max Pr [kPa]

Phe loss [%]

40 45 40 45

0.009 0.011 0.0035 0.005

710 782 1640 2078

49.6 41.4 1.7 1.4

0.013 0.030 0.007 0.011

848 1333 2703 4000

35.7 18.4 1.2 1.1

a Based on the water permeability Aw calculated from clean water flux measurements (6.3 × 10-8 and 5.6 × 10-9 s m-1 for the NF membrane and RO membrane, respectively) and the parameters given in Table 1.

Figure 8. XRD spectrum for the crystals obtained during coupled operation. The five major peaks correspond to the five peaks suggested by Mohan et al.4 for the differentiation between Phe anhydrate and monohydrate (2θ ) 5.7, 17.1, 22.8, 28.6, 34.5).

Figure 7. (a) Phe mass fraction profiles in the buffer vessel and the crystallizer and (b) Jw during coupled operation. Symbols are experimental data, and lines are best simulation results (parameter values within the ranges shown earlier: BPhe ) 3.79 × 10-5 kg m-2 s-1, kPhe ) 0.0065 kg m-2 s-1, kg ) 1.0 × 10-6 m s-1, and g ) 1.8). Temperature in the crystallizer ) 45 °C.

the NF membrane. Accordingly, the RO membrane has a higher RPhe with lower Jw compared to the NF membrane. 2. Crystallization Experiments. Two batch experiments with initial supersaturation values of 1.17 and 1.21 resulted in kg ) (2.37 ( 1.38) × 10-6 m s-1 and g ) 1.8 ( 0.3 (results not shown). Parameter values within these ranges were used in eq 25 together with eqs 5-7 to evaluate the mass fraction profile in the crystallizer during an experiment with continuous inflow and outflow of solution. Figure 4 shows the Caq Phe,c profile during the experiment. For the first part of the experiment (t < 100 min) the best model prediction (see parameter values in the aq , but by the end of figure legend) slightly overestimates CPhe,c the experiment a supersaturation equal to 1.06 is correctly predicted. Therefore, it was decided to use the model for further process evaluation. 3. Selection of Conditions of Operation for Coupled Operation. A typical objective is to crystallize Phe at the production rate of the fermentation. Let us assume a target Phe s of 3.8 × 10-7 kg s-1 (which is typical for production rate pPhe,c aq a 1.2 kg fermenter containing 30 g kg-1 cell mass) with a CPhe,f

of 20 g kg-1 and maximal 5% Phe loss in the permeate (relative to Phe in the feed). Figure 5 shows the window of operation for the seeded crystallization of Phe anhydrate. Crystallization should be performed in the region of anhydrate thermodynamic stability. The lower limit in the mass fraction axis is obviously the anhydrate solubility curve. The upper limit corresponds to the metastability limit in order to prevent spontaneous monohydrate crystallization. In the temperature axis the upper limit corresponds to the temperature limit of the membranes used for water removal, while the lower limit corresponds to the crystal transition temperature. From eq 27 it can be seen that s aq aq to be large Φc,in should be large, CPhe,b - CPhe,c should for pPhe,c aq aq should be small (thus both CPhe,c and be large, or 1 - CPhe,c aq as high as possible). Based on this, it was decided to CPhe,b aq ) 45.0 g kg-1 operate the buffer vessel at 50 °C and CPhe,b (just below the metastability limit). For the crystallization we chose a temperature within the window of operation, say 45 aq ) 38.3 g kg-1 and F ) °C. Under these conditions, CPhe,c,eq 2.25. In selecting the conditions for water removal, the pressure limits of the membranes and equipment need to be considered. Even though the water removal setup described under Materials and Methods can operate up to 7000 kPa, the membranes have a limit of operation at 4000 kPa. Substitution of eqs 16 and 17 in eq 14 allows the evaluation of Pr as a function of Jw. Figure 6a shows that, for the RO membrane, Jw should remain below 0.011 kg m-2 s-1. Next, eq 35 is substituted in eq 27 to evaluate aq as a function of Jw. In Figure 6b it can be seen that, in CPhe,c order to ensure supersaturation, Jw should be higher than 0.005 kg m-2 s-1, which corresponds to 2078 kPa. Membrane operation at lower pressure can be favored by crystallizing at a lower temperature, say 40 °C (see Figure 6b). At this temperature, the minimal Jw becomes 0.0035 kg m-2 s-1 (1640 kPa) and the maximal Jw is now defined by the metastability limit in the crystallizer (0.007 kg m-2 s-1 or 2703 kPa). The same approach can be used to select the conditions with the NF

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membrane. Table 2 summarizes the results for both membranes at 40 and 45 °C. From this table it can be seen that it should be possible to reach the target crystallization rate with either membrane within the pressure and temperature limits of the system. However, as discussed above, the NF membrane has a much lower Phe retention than the RO membrane. This leads to a Phe loss in the permeate much larger than the constraint of 5% (see Table 1). Therefore, the RO membrane is chosen for the coupled experiment. Figure 7 shows the experimental profiles and the best simulation results (see parameter values in the figure legend) of the coupled experiment. The first 220 min of operation show the process of concentration in order to reach the required value aq of CPhe,b . It can be seen that the model agrees reasonably well with the experimental values. At 220 min the flow toward the crystallizer started and, as shown by the model line, it was expected that both Caq Phe,b and Jw would remain constant in order s to achieve the target pPhe,c . The results show that Jw did remain aq stable, but at a lower value than predicted. Next to this, CPhe,b aq continuously decreased to values comparable to CPhe,c. Hence, even though the right crystal form was obtained (as confirmed s by XRD; see Figure 8), pPhe,c was not reached (as confirmed from mass balances). These results illustrate that a narrow window of operation increases the chance of failure due to operation at conditions near the limits of the window of operation. At t < 220 min, aq CPhe,b did not increase as smoothly as predicted by the model. There were fluctuations, probably due to instability in the setup (pump rate, pressure, etc). Occurrence of such fluctuations during the coupled operation could easily result in crossing the metastability limit. Thus, failure might have originated on crystallization at the membrane surface; this would explain the reduction in Jw, which translates as loss of water removal capacity of the system and, consequently, dilution in the buffer vessel. This is also suggested by the clean water flux measurements performed after the experiment, which resulted in a 14% decrease in Aw (from 6.0 × 10-9 to 5.2 × 10-9 s m-1) compared to the values before the experiment. Further confirmation of crystallization at the membrane surface by means of microscopy, scanning electron microscopy, or XRD measurements was not performed due to practical reasons: after the experiment, the temperature of the membrane could not be kept at 50 °C. This means that it would not be possible to differentiate between crystals formed during the experiment and crystals formed upon cooling of the membrane. 4. Widening the Window of Operation. Clearly, a wider window of operation would allow for a more stable process. From the crystallization side, this could be done by working with a different solvent or with additives that result in solubility reduction or shifting of the transition temperature.4,22 Solubility reduction, however, might be followed by a reduction in the metastability limit. From the water removal side, the window of operation can be widened by several means: by using membranes with better temperature and pressure resistance than the ones used in this study; by screening for membranes with large Jw but still high RPhe, and/or by increasing the membrane aq area. These measures would also allow maintaining CPhe,b at undersaturated levels, hence minimizing the risk of crystallization at the membrane surface. Widening the window of operation from the water removal side might be easier to implement,

especially when the choice of solvents or additives for the crystallization is limited (e.g., in fermentation processes). Conclusions In this work, a coupled water removal and crystallization system was developed for the crystallization of Phe anhydrate. In this system, supersaturation generation and crystal growth are decoupled, giving the possibility of more control over the crystallization process. This is crucial for processes that require operation within a narrow supersaturation range, as is the case for Phe anhydrate. For the application concerning this study, the target is to achieve a certain crystal production rate. Both operations, water removal by membranes and crystallization, were evaluated independently in order to derive models that would allow the simulation of the coupled system. With these models it was possible to define the operating limits for the coupled system so that the target crystal form and production rate might be achieved. Experimental evaluation showed that conditions of operation near the limits of the system demand smooth setup performance, which was not achieved in the current case. In order to operate away from the limits of the system, a widening of the window of operation is required. This might be achieved by changing the solvent or using additives for crystallization, or by selecting better membranes in terms of temperature and pressure limits, Jw, RPhe, and area. The approach developed here should be applicable for other compounds and crystallization processes in which supersaturation generation and crystal growth are uncoupled. Moreover, it is expected that the model developed in this study can be extended to include CSD prediction; in this way the coupled system can be further optimized toward given product specifications. Implementation of this system with more complex solutions (e.g., from fermentation) requires studying the effect of impurities on both the water removal and the crystallization performance. Acknowledgment These investigations are supported in part by The Netherlands Research Council for Chemical Sciences (CW) and The Netherlands Technology Foundation (STW) in the NWO research program Separation Technology. We thank Sanaz Allaie for her contribution in the selection of equipment and the preliminary experimental design. Niek van der Pers at the Department of Materials Science and Engineering of the Delft University of Technology is acknowledged for the X-ray diffraction analysis. Nomenclature A ) membrane permeability [s m-1] AR ) aspect ratio B ) mass transfer coefficient [kg m-2 s-1] C ) mass fraction [g g-1] F ) concentration factor g ) order of crystal growth J ) flux [kg m-2 s-1] k ) mass transfer coefficient [kg m-2 s-1] kg ) constant [m s-1] L ) length [m] m ) mass per crystal [g] M ) mass [g] n ) number of crystals p ) productivity [kg s-1] P ) pressure [Pa] R ) universal gas constant [J mol-1 K-1]

Ind. Eng. Chem. Res., Vol. 48, No. 3, 2009 1573 RPhe ) retention S ) supersaturation T ) temperature [K] V ) volume [m3] W ) molecular weight [g mol-1] Greek Symbols F ) density [g L-1, g m-3] Φ ) mass flow rate [g s-1] Π ) osmotic pressure [Pa] Superscripts aq ) aqueous s ) solid, crystals Subscripts 0 ) initial b ) in the buffer vessel c ) in the crystallizer eq ) at equilibrium f ) in the feed g ) growth m ) at the membrane p ) in the permeate Phe ) L-phenylalanine r ) in the retentate TM ) transmembrane w ) water, solvent

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ReceiVed for reView August 19, 2008 ReVised manuscript receiVed November 4, 2008 Accepted November 13, 2008 IE8012659