2058
Ind. Eng. Chem. Res. 1993,32, 2058-2064
Amino Acid Separation in a Multistage Fluidized Ion Exchanger Bed Magdiel AgostoJ N.-H. Linda Wang, and Phillip C. Wankat' School of Chemical Engineering, Purdue University, West Lafayette, Indiana 47907
Theoretical and experimental investigations of the use of a Cloete-Streat contactor for ion exchange of amino acids are presented. A new step p H gradient approach was developed to concentrate and recover dilute components. For short cycles, the models have to include the supernatant fluid in between stages. Agreement between theory and experiment for phenylalanine separation was excellent for isocratic experiments but only fair with pH step gradients. Slow changes in internal pH were not included in the model and appear to be the reason for the lack of agreement between theory and experiment. For lysine separation, experimental concentration factors greater than 2.5 were obtained.
Introduction and Literature Review Downstream processing is almost always done in steps. A typical order is solids removal, initial recovery and concentration, purification, and final concentration. In a plant, the first step, solids removal, often creates major operating problems and often results in considerableyield losses. The opportunity to combine the solids removal step with the initial recovery and concentration step by use of fluidized bed ion exchange or affinity sorption has recently been the topic of several investigations. Single fluidized bed ion exchange has been studied by Buijs and Wesselingh (1980) for K+-Na+ exchange; by Wells et al. (1987),Bascouletal. (1989),andDraegerandChase(1990) for protein recovery;and by Gaillot et al. (1990) for recovery of immunomycin. Roe (1987) studied the adsorption of amylase on a nonionic resin, while Somers et al. (1989) studied the affinity separation of endo-polygalacturonase using alginate beads as the affinity sorbent. Gordon and Cooney (1990)used a stirred tank for both affinity and ion exchange separations of proteins. The experimental results with single fluidized beds typically show early breakthrough of the adsorbate. Bascoul et al. (1989) and Gaillot et al. (1990) found that, as expected, putting beds in series helps to reduce this problem significantly. Fluidized beds in series have been used commercially (Belter et al., 1973) for antibiotic recovery. Sharper separations can also be obtained with magnetically stabilized fluidized beds. Evans and Burns (1992) used this method with pH focusing to concentrate and separate myoglobin. An alternate type of fluidized system is the multistage contactor, which has appeared in several designs (Wankat, 1986). The most extensively studied is the Cloete-Streat contactor, which is a sieve plate system with no downcomers. The operation of the Cloete-Streat contactor is illustrated in Figure 1. During upflow, the solids are fluidizedon each stage, but there is very little mixing from stage to stage because of the sieve plates. To transfer solids down, a short settling period is followed by a downward fluid pulse. After a start-up period, a cyclic steady state is reached. Operation is not steady state but does approximate continuoussteadyatate countercurrent contact. Operation of the contactor is discussed in detail by van der Wiel and Wesselingh (1989), Wankat (1986), and Wesselingh and van der Meer (1986). The Cloete-Streat contactor was selected because it has several advantages. The pulsed operation essentially To whom correspondence should be sent. + Current address: Exxon Chemical, Baton Rouge, LA.
0888-5885193/2632-2058$04.0010
0
.. .. ..
I-?Figure 1. Schematic of operation of Cloete-Streat contractor: (a) fluid flow with fluidized solid, (b) settling step, (c) downward pulse of fluid and solid.
decouples the solid and liquid flow rates so that practically any desired LIS ratio can be used. Suspended solids will pass through the system. It is easy and inexpensive to add additional stages. In the mining industry, the system has been scaled up to very large sizes (Cloete, 1984). The patent on the system has expired, and extensive design information is available in the open literature (van der Wiel and Wesselingh, 1989;Wesselingh and van der Meer, 1986). Cloete-Streat contactors have been used for biotechnology, and various applications are reviewed by van der Wiel and Wesselingh (1989) and van der Wiel (1989). Amino acid separations are not reported in their literature review. A disadvantage of the design is that either relatively large or dense particles must be used for hydrodynamic reasons. In this paper, we report on the use of a Cloete-Streat contactor for amino acid recovery and concentration by ion exchange. The novel features include (1)the development of a new pH step gradient approach to simultaneously concentrate and purify the amino acid and (2) development of a staged model which includes the supernatant fluid.
Experimental Section The glass laboratory apparatus is shown in Figure 2. The stages were constructed from glass O-ringjoints with a plate held between two O-rings. Experiments were done with three or five stages. The laboratory system is 3.8 cm in diameter, and each stage is about 10 cm tall with about 5.5-6.0 cm of fluidized resin on each stage. Each Teflon sieve plate except for the bottom plate has seven holes 0.4 cm in diameter, which gives a free hole area of 7.7 9%. To prevent slow loss of region from the bottom plate, one of 1993 American Chemical Society
Ind. Eng. Chem. Res., Vol. 32, NO.9, 1993 2059
ACID (OPTIONAL)
I
I
I
Figure 2. Sketch of laboratory fluidized bed apparatus.
the holes was plugged. Complete details of the apparatus and operation are given by Agosto (1991). The resin used was Rohm and Haas Duolite C-20 cation exchange resin from 20 to 50 mesh (0.297 to 0.840 mm). The resin was initially in the H+ form. Liquid flow rate was typically in the range 230-290 mL/min which is a superficial linear velocity from 20 to 25 cm/min. In this range of liquid flow rates, neither weeping nor entrainment was observed. In experiments, a 5-min upflow period, an 11-s period for settling, and 22 s of downward flow were used. We observed that about 60% of the resin in each stage was transferred (freak= 0.6),which gives a solid flow rate of S = 3.45 mL/min. These values were not optimized, but they resulted in reasonable values of LIS, and operation was simpler than with shorter cycles. The experiments have used phenylalanine or lysine as the solutes. Analysis of the Phe concentration was done with a UV spectrophotometer at 254 nm. Preliminary analysis of Lys was done at 233 nm, while final analysis used orthophthalaldehyde (OPA) derivatization and then injection into an HPLC with detection in a fluorimeter (Agosto,1991).Feed concentrations of 1 g/L amino acid were used. Isocratic experiments were done with a pH of approximately 2. Step gradient runs had resin in the H+ form added at the top and a feed containing NaOH at the bottom. Additional acid (1-2 M HC1) was sometimes added between stages to control the pH. With proper balancing the NaOH in the feed would neutralize the resin in the H+ form, and no extra acid was needed. Typically, the bottom two stages were at high pH and the top two at low pH. The third stage was a swing stage. After considerable initial experimentation, we were successful in doing pH step gradient runs with Phe and Lys. In the initial runs, the pH of the fluid in the bottom stages was high, but insufficient NaOH was used to increase the pH inside the resin. These experiments essentially repeated the isocratic experiments with good reproducibility. Obtaining the desired focusing required starting up in a manner which converted the bottom stages to Na+ form while the top stages stayed in H+form. After the column was loaded with resin, a feed solution containing
Figure 3. McCabe-Thiele diagram showing concentrationwith step pH gradient.
0.2 M NaOH and 1.0 g/L Phe or Lys was pumped into the column until the bottom three stages were converted to the dark brick-red color which indicated conversion to the Na+ form. During this conversion period, some 1.0 M HC1 was added between stages 3 and 4 to keep stages 4 and 5 in the H+ form (a light rust color). No downward transfers were done until resin in all stages was converted to the desired form. Because of resin shrinkage when converted to the Na+form, additional resin had to be added to the bottom stages. Then the feed was switched to 0.03 M NaOH with 1.0 g/L Phe or 0.042 M NaOH with 1.0 g/L Lys and cycles were started. For the first 11 cycles, small amounts of acid were added to keep the top two stages in H+ form. After cycle 11, no acid addition was required since the resin added in H+ form neutralized the NaOH. Cyclic steady state was reached after about 15 cycles.
Theory Several different models are being used to explore the operation of the Cloete-Streat contactor. The simplest are steady-state approximations which assume that the device is a truly continuous, countercurrent contactor. The device shown in Figure 2 is now very similar to a countercurrent absorber, extractor, or distillation column, and the well-developed theories used for these systems are applicable. Specifically, for an equilibrium staged model or for a staged model with stage efficiencies, one can use the Kremser equation, a McCabe-Thiele diagram, a computer stage-by-stagecalculation,or amatrix inversion method (e.g., Henley and Seader, 1981; Wankat, 1988). A simple example of the insight which can be obtained from a continuous staged model is to look at the ratio L/(rniS)where mi = Ei/ci is the equilibrium parameter. If L/(miS)> 1, then solute i tends to exit with the liquid at the top of the column, and if L/(miS)< 1, then solute i tends to exit with the solid at the bottom of the column. This is useful for setting the flow rates to fractionate two solutes i andj. Qualitatively,the equilibrium staged model can be used to explain the step pH gradient operation. If we use a high pH at the bottom of the column, then mi will be low and L/(miS)> 1. A t the top of the column, we force the pH to be low, and then mi is high and L/(rniS)< 1. The exact pH values are not critical as long as the inequalities are satisfied. The net result is that amino acid i travels from both ends of the column toward the middle, where it concentrates. For an equilibrium staged model, this is illustrated in Figure 3. With very dilute feeds, it should
2060 Ind. Eng. Chem. Res., Vol. 32, No. 9, 1993
rium data can be fitted by a linear isotherm. For linear systems, the lumped parameter mass-transfer coefficient can be estimated from r
Figure 4. Model for fluidized bed system with supernatantsection.
be possible to achieve very high concentration factors. The concentrated fluid can be removed with a side stream or intermittent withdrawal (neither of which is shown in Figure 3). Although tremendously oversimplified for the Cloete-Streat system shown in Figure 2, this model does give a qualitative feel for the concentration. For amino acids, this flexibility in changing L / ( m i S ) comes from mi)s susceptibility to a change in pH. Amino acids are zwitterions whose ion exchange equilibrium affinities are easily influenced by their pH environment (Yu et al., 1987):
-
NH,+CHRCOO- + H+ NHCCHRCOOH Protonated and unprotonated amino acid forms have greatly different affinities. As a result of maintaining two different pH regions in the column, one can effectively establish "stripping and enriching sections" which can concentrate the trace feed component between these two regions. The next degree of sophistication after an equilibrium stage model would be to assume continuous steady-state operation but solve the mass-transfer expression for each stage. In effect, this approach allows one to calculate the stage efficiency from the mass transfer-coefficients. A major increase in physical accuracy and model complexity occurs when the periodic operation of the Cloete-Streat contactor is included. With these models, one will get the start-up and the limit cycle behavior of the contactor, and one will see differences between the Cloete-Streat contactor and a truly continuous system. The first model developed was a lumped parameter masstransfer model (Slater, 1974). This model treated each stage as a singlefluidized bed which was completely mixed. We have since observed that these models are inadequate at short cycle times. Observation of the experimental apparatus showed that there was a reproducible supernatant region in each stage which was not included in the model. A model including these regions is shown schematically in Figure 4. The fluidized bed region uses a lumped parameter mass-transfer expression. The resulting equations are as follows. fluidized bed's bulk liquid mass balance:
fluidized bed's resin mass balance: dCj dt
3kL
R
supernatant liquid's mass balance: (3) The prime on j designates the supernatant liquid stage above stage j, and mj is the Henry's law equilibrium parameter, e,* = Ej/rnh For all the systems tested, the solute concentrations are sufficiently low, so the equilib-
kLaP
=
1
.
-I
1 +
2
1-1
J
(4)
60Dmpmj up
where for nonlinear systems m = (dq/dc),. part ic1es
For spherical
ap = 6/dp (5) The film mass-transfer coefficient can be estimated from a suitable correlation such as the modified SnowdonTurner correlation (Rahmanand Streat, 1981),
and the bed porosity in the fluidized state qj can be estimated from the Richardson and Zaki relationship (Wesselingh and van der Meer, 1986). The boundary conditions for startup are
cj = c ~ , E j~= ~ ~ ~t =~0, ,j = 1 , 2 , i N (7a) co = C f d , EN+1= E,&, t >0 (7b) After the upflow step, resin is transferred downward. Observations of the experiments showed that there is considerablemixing of the fluidized and supernatant liquid regions between each part of the sieve plates. Thus, to update concentrations for the next cycle we assumed that regions j and 'j were perfectly mixed. c
cj =
'v**es + cj€fjv8*ej
'Vstageg + 'fjV,*ej
(8)
cy = cj (9) The resin update formula is unaffected by the presence of the supernatant liquid regions.
where fminS 1 (10) = (1 - fmin)Ej + freainEj+l The model is applicable to both isocratic and pH step gradient operation. In the latter case, we assumed that the pH change is very rapid after resin transfer. Thus, a stage is either at mL or mH. Unfortunately, the resin in the experimental apparatus had not read our model. The resin took about 2 min to change color, signifyinga change from H+ to OH-form. Internal changes probably took longer. Thus, the model could not accurately predict the behavior of the very important swing stage. Cj
Results Isocraticresults for Phe for both theory and experiment are shown in Figure 5. The experimental results show 86 5% recovery of Phe with only three stages. The values of kL and m were obtained by fitting breakthrough data in a two-stage fluidized bed system with no solids flow. Because of experimental problems with the breakthrough experiment, the value of kL used may be too low. A higher kL value would give better agreement between theory and experiment. For phenylalanine, kL = 0.13 cm/min and m = 250 if the supernatant sections are excluded. If the supernatant sections are included, kL = 0.11 cm/min and m = 249. The closeness of these results indicates that the supernatant section could be excluded from the model for breakthrough experiments which are essentially long cycle time experiments. The value of m obtained from batch equilibrum studies was m = 220.
Ind. Eng. Chem. Res., Vol. 32, No. 9, 1993 2061 "
n
n 0I
b
0
" 6 ,
1
I
I
X
W
n
E
W
2
2
1
0.
Stage n u m b e r
Figure 5. Cyclic steady-state results for theory and experiment for Phe recovery in three stage systems: Phe isocratic three-stage CloeteStreat run,pH 2. c f d = 0.006 053 M, L = 238 mL/min, cyclic time = 5min. Data represent samples taken at the end of a cycle. Particle diameter = 0.056 85 cm, H(1) = 6.1 cm, H(2) = 6.0 cm, H(3) = 5.5 cm, m = 262.8, kL = 0.246 cm/min, capacity = 1.26 mol/L Phe of wet isocratic experiment. bed: (-) LMT model, (0) n
nI
0
$ 0 1 490
3
I 3
I
I
'
"
491
"
492
'
493
'
"
494
I
495
T i m e (minutes)
Figure 7. Cycle transient period: effluent histories at cycle 99. Model results for isocratic three-stage recovery. Cycle time = 5 mix (-) stage 1, (- -) stage 2, stage 3. (-e)
Table I. Theoretical Results for Three-Stage Ieocratic Modeling for Phenylalanine' cycle L recovery time(min) (mL/min) S LIS CN(M) (75) 1.0 280.0 18.81 14.89 8.4155 X 1od 98.6 350.2 4.70 74.47 4.4988 X lo-' 92.6 4.0 94.7 3.76 74.42 3.1953 X lo-' 5.0 280.0 3.135 74.42 2.3889 X lo-' 6.0 233.3 96.1 87.4 280.0 1.881 148.86 7.6058 X lo-' 10.0 a
Details are given by Agosto (1991).
v
c
s0m .012. 2
I'
I
1 $' oo 10
20
30
Cycles
Figure6. Overallcolumn starbup period: model results for isocratic three-stage recovery. Data a t end of each cycle. Cycle time = 5 min, f = 0.6 (-) stage 1, (-3 stage 2, (--) stage 3.
I
2a
t 3
The model was used to extensively study the effect of operating parameters on the isocratic operation of the Cloete-Streat device. The start-up results for a threestage system are shown in Figure 6. At the cyclic steady state, each stage has transient behavior during each cycle. This is illustrated in Figure 7. As expected, increases in m,kL, and N increased the recovery. The effect of cycle time is more complex since cycle time can affect LIS and the residence time. The results are shown in Table I. If LIS is kept constant, then increasing the cycle time increases the recovery. However, if L is kept constant, increasing the cycle time increases LIS and the recovery drops. Thus, the slope of the operating line is more important than the increasing residence time. The experiments for pH step gradient operation were not optimized but do show that significant concentration can be obtained. The results for a five-stage pH step gradient run with Phe are shown in Figures 8and 9. Figure 8 shows the column profile at the end of a cycle, while Figure 9 shows the transient history for stage 3. Both of these results were obtained after the system had reached
Stage number
Figure 8. Comparison between theory and experiment. Column profile: Phe pH gradient five-stage Cloete-Streat run. Stages 1-3 in Na+ form. Stages 4 and 5 in H+ form. Feed had a 0.03 M NaOH concentration. No acid added, c f d = 0.006 053 M, L = 287 mL/min, cycle time = 5 min (data taken a t the end of the cycle). Theory: kL = 0.11 cm/min, m(1-3) = 27.0 < L/S< m(4-5) = 200.4. At 17 cycles: (0) experiment, (-) LMT model.
cyclic steady state. The theoretical results were obtained by fitting parameters and will be discussed later. Note that there is a modest concentration effect on stages 3 and 4. Since the Phe has significantnonionic interactionswith the resin, the change in m as pH changes is not as dramatic as with other amino acids. The results obtained for Lys are shown in Figures 10 and 11. Figure 10 shows the column profile at the end of the cycle, while Figure 11 shows the transient behavior of stage 4 at the cyclic steady state. Note the maximum concentration obtained is more than 2.5 times the feed concentration. Extensive theoretical studies were done for the pH step gradient recovery of Phe. For a 5-min cycle,the theoretical
2062 Ind. Eng. Chem. Res., Vol. 32, No. 9, 1993
-
h
h I
\
-
3
"
1
1 31
6
h
E 1.6
v
0 0 0
0
u
2
.4
v:
x 4
a r
0 3
1
5
Time ( m i n . )
T i m e (mln.)
Figure 9. Comparison between theory and experiment for Phe recovery. Effluent history for stage 3. Stages 1-3 in Na+ form, stages 4 and 5 in H+ form. c f d = 0.006 053 M, L = 287 mL/min, cycle time = 5 min. Stage height = 4 cm, d, = 0.056 85 cm. Feed had a 0.03 M NaOH concentration. No acid added. Theory: kL = 0.11 cm/ min, m(1-3) = 27.0 < L/S < m(4-5) = 200.4. At 17 cycles: (0) experiment, (-) LMT model.
Figure 11. Effluent history for stage 4. Stages 1-3 in Na+ form, stages 4 and 5 in H+ form. Lys-HC1 c f d = 1 g/L, NaOH feed concentration = 0.042 M, pH 11.9. Fluidized resin stage height = 3 cm, L =316 mL/min, cycle time = 5 min. No acid added. h N
2 -
2
h
v
7 M
1.01
1 0
1
I d
v
0
.St 0
L
0
g 3
.3
J
2
0
d
c, al
u e
2 '
0
u
'a h
0 0
'?
i-
b 3
I
t 0 ' 1
3
e 5
Stage n u m b e r
Figure 10. Lysine recovery for five-stagepH gradient Cloete-Streat run. Feed is 0.042 M NaOH. Stages 1-3 in Na+ form, stages 4 and 5 in H+ form. Lys-HC1 feed concentration = 1 g/L, NaOH feed concentration = 0.042 M, L = 316 mL/min, cycle time = 5 min (data taken at the end of the cycle). pH gradient for Cloete-Streat contractor. No acid added.
results are shown in Figures 12-14. The value of m = 249.4 in the top stages is that obtained at pH 2.0. The value of m = 40 in stages 1-3 was set arbitrarily at a value less than LIS.Figure 12 shows the column profile at the end of a cycle after the cyclic steady state has been reached. Note that the fluid concentrations in the fluidized and in the supernatant sections are equal. The start-up behavior over many cycles is shown in Figure 13 for stage 3. On the scale of this figure, the difference between fluid concentrations within the resin and supernatant regions cannot be seen. About 45 cycles were required to reach cyclic steady state. In Figure 14, we see that the two fluid concentrations are different, but only at the start of the cycle. Thus, for long cycle periods the simpler model which ignores the supernatant sections is certainly adequate. Figure 14 indicates that the more complicated model may be important for short cycles.
.a
c@
a
O
0 '
"
1
2
"
3
"
4
"
5
'
t
Stage n u m b e r
Figure 12. Five-stage pH gradient recovery-model results. 200 cycles, cycle time = 5 min. Flow rate = 280 mL/min, f = 0.60. Feed concentration = 0.006 053 M. Data points represent samples taken at the end of a cycle. kL = 0.11 cmfmin, m(1-3) = 40 < LIS = 74.5 < m(4-5) = 249.4. Height of fluidized resin section = 5.5 cm, supernatant section = 4.5 cm, bed porosity = 0.50 (0) fluidized resin, (0) supernatant liquid.
The theoretical results for a pH step gradient run with Phe in a five-stage system are shown in Figures 15 and 16 for a 1-min cycle. Because of the faster cycles, the ratio LIS drops to 14.94. The value of m = 1.004 was estimated from breakthrough runs at high pH (Agosto,1991). With these short cycle times, the fluid concentrations in the fluidized and supernatant regions are markedly different. Figure 15 shows the column profile a t the end of the cycle. The detailed transient for stage 3 is shown in Figure 16. Cyclic steady state was reached in less than 70 cycles. Immediately following the transfer step, the Phe desorbs from the resin and the fluid concentration in the fluidized region shoots up to a high value. The fluid concentration in the supernatant region lags behind this since it takes time to transfer fluid to the Supernatant region. The concentration in the fluidized region drops as less concentrated fluid is transferred in from stage 2. Note that very high concentrations are obtained a t very short times, indicating that considerable further optimization can be done with the experimental separations.
Ind. Eng. Chem. Res., Vol. 32, No. 9,1993 2063 n
h
uI
I
2
2.0r
2
I
W
X
8
n
h
E 1.6.
Y
L
I
5 1.0
I
0
1.21
0
E 1.2.
v
3
I
I
1.4.
0
0 I
4
cu O h a
'
5'0
'
100
'
150
'
200
'
"
za 1
2:o
Time (minutes)
Figure 13. Start-up history for stage 3 (including individual cycle transients). Model results for fluidized resin bed and supernatant sections. Cycle time = 5 min. (-) stage 3 fluidized resin, (- -) stage 3 supernatant. CI I
0
2
3
4
5
Stage number
Figure 16. Five-stage pH gradientrecovery. Model results for Phe. 300 cycles, cycle time = 1 min, flow rate = 280 mL/min, f = 0.60.Feed concentration = 0.006 053 M. Data points represent samples taken at the end of a cycle. kL = 0.11 cm/min,m(1-3) = 1.004 < L / S = 14.94 < m(4-5) = 249.4. Height of fluidized resin asction = 5.5 cm, supernatant section = 4.5 cm. Bed porosity = 0.50 ( 0 )fluidized resin, ( 0 )supernatant liquid.
'
n I
h
2 1.6
v
e
.2,o
1.2
1
8.0
X
t
v h
E 6.4
v
e
0
$ 4.8
E
4
i!0 3.2 fi
0 U
2 1.6
2
za Although both the experimental and theoretical results were reasonable, we were at first unable to predict the experimentalresults for the pH step gradient runs. When , and mL obtained from breakthrough the values for k ~m~ curves were used, the theoretical results showed much less concentration than the experimental results. Adjustment of the parameters allowed a reasonable fit of the experiments to be obtained. The fit obtained by raising m in stages 1-3 (the high pH stages) and lowering m in stages 4 and 5 (the low pH stages) is shown in Figures 8 and 9. Thus, Figures 8 and 9 must be interpreted as fits to the experiments,not predictions. A reasonable fit could also be obtained by drastically reducing kL. What is apparently happening in the experiments is that the pH change inside the particles is quite slow, and therefore the Henry's law constant decreases slowly. The particle then serves as a source for Phe as the Phe is slowly desorbed. This slow desorption is crudely modeled by reducing kL or increasing mL. This view agrees with the experimental observation that it took the resin about 2 min to change color after the resin had been transferred into a region of lower pH. Note that even with this fitting of variables the
" 99
99.5
100.0
Time (minutes)
Figure 16. Cycle transient period for pH gradient recovery of Phe. Stage 3's effluent history at cycle 99. Model results for fluidized resin bed and supernatant sections. Cycle time = 1 min: (-1 stage 3 fluidized resin, (- -1 stage 3 supernatant.
model predicts an incorrect slope for the transient in Figure 9. Clearly, a more detailed model is needed for the pH change and desorption inside the particles.
Discussion We have shown both isocratic and pH step gradient ion exchange separation of amino acids in a Cloete-Streat contactor. Obviously, these methods can be extended to other separations such as proteins. In a complete system, the resin exiting from the bottom of the isocratic column would be eluted in a separate vessel. It would be of interest to compare this operation with an optimized pH step gradient system. The hydrodynamics of the CloeteStreat system are not straightforward. The reason for resin depletion in the bottom stage if the free area is not reduced is unclear, but this result is consistently reported (Wesselingh and van der Meer, 1986; van der Wiel and Wesselingh, 1989; van der Wiel, 1989). With pH gradients, the shrinkage of the
2064 Ind. Eng. Chem. Res., Vol. 32, No. 9, 1993
resin further complicates retaining resin in each stage. During our experiments, the amount of resin in the lower stages was slowly reduced. To control similar difficulties, Cloete-Streat contactors are provided with a method for introducing resin holdup on the bottom stage. These contactors have been scaled up to very large sizes in the mining industry (Cloete, 1984). Clearly, the theoretical model in this paper is still too simple. The transfer and neutralization of OH- and H+ when the pH is changed need to be included in the model. It is of interest to operate the Cloete-Streat system with fermentation broth. Gaillot et al. (1990) found a decrease in exchanger capacity in a fluidized bed. This probably occurred due to the sorption of trace components in the broth. One must also check for foulingof the resin. Gaillot et al. (1990) obtained excellent results with Mitsubishi Kasei resin SP-207.This brominated resin has a specific gravity of 1.18, which allows for high fluid velocities. It would also be interesting to operate the contactor in series with the fermenter recycling whole cells. This operation has been done with fluidized beds (Holst and Mattiasson, 1991).
Acknowledgment The support of NSF Grant BCS 8912150 and the donation of resin by Rohm and Haas, Inc., are gratefully acknowledged. The assistance of Mr. Joon-Ho Koh in checkingtheoretical predictions and obtaining equilibrium data is gratefully acknowledged.
Nomenclature up = area/volume of resin (cmZ/cm3) C j = solute concentration in bulk liquid or fluidized resin on stage j (M) cy = solute concentration in supernatant liquid (M) ci. = solute concentration in equilibrium with solid phase (MI cj = solute concentration in exchanger phase (mol/L of fluidized resin) d, = resin bead diameter (cm) D,, = effective mass diffusivity inside particle (cmZ/s) trmh= fraction of resin in stage transferred down to next stage during settling and downflow steps kf = film mass-transfer coefficient (cm/s) kL = overall mass-transfer coefficient (cm/s) L = total flow in liquid phase (cm3/s) m = linear equilibrium parameter = E,/c~* N = number of stages S = resin flow rate in column (cm3/s) R = resin radius (cm) t = time (s) u = fluid velocity (cm/s) Vstagej = volume of fluidized resin on stage j (cm3) VSaes = volume of supernatant fluid on stage j’ (cm3) Greek qj = pf p
porosity of fluidized bed
= fluid density (g/cm3) = liquid viscosity (P)
Literature Cited Agosto, M. Amino acid recovery using a multistaged fluidized bed contactor. Ph.D. Thesis, Purdue University, West Lafayette, IN, 1991. Bascoul, A.; Delmas, H.; Flaschel, E.; Renken, A.; Couderc, J. P. Improvementa of solid-liquid fluidized bed design: use of static mixers and fluidized distributors. In Fluidization Vfi Grace, J. R., Shemilt, L. W., Bergougnou, M. A., Eds.; Eng. Foundation: New York, 1989;pp 717-724. Belter, P. A.; Cunningham, F. L.; Chen, J. W. Development of a recovery process for Novobiocin. Biotechnol. Bioeng. 1973, 15, 533. Buijs, A.; Wesaelingh, J. A. Batch fluidized ion-exchangecolumn for streams containing suspended particles. J. Chromatogr. 1980, 201,319-327. Cloete, F. L. D. Comparative engineering and process features of operating continuous ion exchange planta in South Africa. In Zon Exchange Technology; Naden, D., Streat, M., Eds.; Ellis Horwood: Chichester, England, 1984,661. Draeger, N. M.; Chase, H. A. Protein adsorption in liquid fluidized beds. Advances Separation Rocesses; Institution of Chemical Engineers, Symp. Ser. #118,Rugby, UK, 1990; pp 161-172. Evans, L. L.; Burns, M. A. The focusing of biochemicals in countercurrent adsorption columns. Paper 15i,AIChE Meeting, Miami, Nov. 5, 1992. Gaillot, F. P.; Gleason, C.; Wilson, J. J.; Zwarick, J. Fluidized bed adsorption for whole broth extraction. Biotechnol. h o g . 1990,6, 370-375. Gordon, N. F.;Cooney, C. L. Impact of continuous affinity-recycle extraction (CARE) in downstream processing. In Protein Purification-From Molecular Mechanisms t o Large-Scale R o cesses; Ladisch, M. R., Willson, R. c.,Painton, C.-d.C., Builder, S. C., Eds.; ACS Symposium Series 427; American Chemical Society: Washington, DC, 1990;pp 118-138. Henley, E. J.; Seader, J. D. Equilibrium-Stage Separation Operations in Chemical Engineering; Wiley: New York, 1981. Holst, 0.;Mattiasson, B. Solid sorbenta used in extractive bioconversion processes. In Extractive Bioconversions;Mattiasson, B., Holst, O., Eds.; Marcel Dekker: New York, 1991;pp 189-205. Rahman, K.; Streat, M. Mass transfer in liquid fluidized beds of ion exchange particles. Chem. Eng. Sci. 1981,36,293. Roe,S. D. Whole broth extraction of enzymes from fermentation broths using commercially available adsorbenta. In Separations for Biotechnology; Verrall, M. S., Hudson, M. J., Ede.; Ellis Horwood: Chichester, England, 1987;pp 210-216. Slater, M. J. Continuous ion exchange in fluidized beds. Can. J. Chem. Eng. 1974,52,43. Somers, W.; Van’t Reit, K.; Rozie, H.; Rombouta, F. M.; Visser, J. Isolation and purification of endo-polygalacturonase by affiiity chromatography in a fluidized bed reactor. Chem Eng. J. (Lausanne) 1989,40,B7-B9. Van der Wiel, J. P. Continuousrecoveryof bio-producta by adsorption. Ph.D. Dissertation, Technische Universiteit Delft, Delft, The Netherlands, 1989. van der Wiel, J. P.; Wesaelingh, J. A. Continuous adsorption in biotechnology. In Adsorption: Science and Technology; Rodrigues, A. E., et al., Eds.; Kluwer Academic Publishers: The Netherlands, 1989. Wankat, P. C. Large Scale Adsorption and Chromatography; CRC Press: Boca Raton, FL, 1986;Vol. 11, Chapter 6. Wankat, P. C.Equilibrium-Staged Separations; Elsevier: New York, 1988. Wells, C. M.; Lyddiatt, A.; Patel, K. Liquid fluidized bed adsorption in biochemical recovery from biological suspensions. In Separations For Biotechnology; Verrall, M. S., Hudson, M. J., Eds.; Ellis Horwood Chichester, England, 1987;pp 217-224. Wesselingh,J. A.; van der Meer, A. P. Counter-current ion exchange. In Zon Exchange: Science and Technology;Rodrigues, A. E., Ed.; Martinus Nijhoff Publishers: B.V., Dordrecht, The Netherlands, 1986;pp 289-318. Yu, Q., Yang, J., and Wang, N.-H.L. Multicomponent Ion-Exchange Chromatography for Separating Amino Acid Mixtures. Reactive Polymers 1987,6,33-44. Received for review August 31, 1992 Revised manuscript received May 11, 1993 Accepted May 20,1993
