Ind. Eng. Chem. Res. 1996,34, 1779-1788
1779
Optimal Separation of Jojoba Protein Using Membrane Processes Hiroshi Nabetani,* Thomas P. Abbott, and Robert Kleiman USDA, Agricultural Research Service, National Center for Agricultural Utilization Research, New Crops Research, 1815 North University Street, Peoria, Illinois 61604
The efficiency of a pilot-scale membrane system for purifying and concentrating jojoba protein was estimated. In this system, a jojoba extract was first clarified with a microfiltration membrane. The clarified extract was diafdtrated and the protein was purified with a n ultrafiltration membrane. Then the protein solution was concentrated with the ultrafiltration membrane. Permeate flux during microfiltration was essentially independent of solids concentration in the feed, in contrast with the permeate flux during ultrafiltration which was a function of protein concentration. Based on these results, a mathematical model which describes the batchwise concentration process with ultrafiltration membranes was developed. Using this model, the combination of batchwise concentration with diafiltration was optimized, and a n industrial-scale process was designed. The effect of ethylenediaminetetraacetic acid (EDTA) on the performance of the membrane system was also investigated. The addition of EDTA increased the concentration of protein in the extract and improved the recovery of protein in the final products. The quality of the final product (color and solubility) was also improved. However, EDTA decreased permeate flux during ultrafiltration.
Introduction
Experimental Section
Jojoba [Simmondsia chinensis (Link) Schneiderl is endogenous to the Southwestern U.S.and Mexico, but is being grown commercially in many other countries. The cultivated acreage in the U.S.is now about 12 000 acres. Jojoba oil consists of unique wax esters which constitute 50-60% of the seed weight (Iwasaki, 1985; Spencer and Plattner, 1984). Jojoba oil is used in cosmetics and lubricants (Gillis, 1988). Although the meal, after screw pressing, typically contains 25% crude protein (Abbott et al., 1991; Verbiscar et al., 1979), it is underutilized. Other components of the meal are sugars (mainly pinetol) and unique antifeedants (simmondsin and its derivatives) which prevent the use of the meal as feed for domestic animals (Elliger et al., 1973, 1974a,b; Manos et al., 1986; Cokelaere et al., 1992a,b, 1993a,b; Verbiscar and Banigan, 1978; Weber et al., 1983). If the soluble protein in the meal could be purified, it could be tested in cosmetics formulations. Membrane technology (microfiltration or ultrafiltration) is used for separating components in a liquid according to their particle size or molecular weight. For example, insoluble materials such as suspended solids can be removed from a liquid using microfiltration (MF), and large molecules such as proteins can be separated from small molecules such as sugars using ultrafiltration (UF). A process as follows might be useful for jojoba meal. At first, proteins, sugars, and antifeedants are extracted from jojoba meal using water. Then, using MF and UF, soluble proteins in the extract are isolated by removing suspended particles, sugars, and antifeedants. The solid after extraction could be used as animal feed because antifeedants are removed. In this study, pilot-scale membrane processing was evaluated for purification and concentration of jojoba protein. Based on the experimental results, an industrial-scale process was designed. Effect of the addition of ethylenediaminetetraacetic acid (EDTA) on the performance of the membrane process was also investigated.
Materials. First pressed jojoba meal (12%oil, 27.4% crude protein) was obtained from International Flora Technologies (Apache Junction, AZ) as Lot No. 92336. It was commercially cold pressed from seed harvested in 1992. Crude protein in first pressed meal is 40% soluble at pH 6, and 20% of the total nitrogen is from nonprotein sources (Wolf et al., 1994). The first pressed jojoba meal was preliminarily tested in the laboratory for protein extractability and composition. Seventy percent of the total nitrogen was solubilized at pH 9 with 0.1% EDTA in the water. This was reduced to 62.5% when insoluble nitrogenous materials were removed at 13000g in a high-speed centrifuge. Using a nitrogen solubility of 62.5% and a nonprotein nitrogen content of 19.6%, the maximum soluble protein that could be isolated from first pressed meal was 11.8%of the starting meal weight. Membrane Apparatus. A membrane unit equipped with a variable-speed motor, PSK-1-MF, Prostak Microfltration System (Millipore Corp., Marlborough, MA) is shown schematically in Figure 1. Plate and frame type modules were used in this study. HVPP modules (Millipore Corp., Marlborough, MA) were used as the microfiltration (MF) membrane. The pore size was 0.45 pm and the membrane material was hydrophilic poly(viny1idene difluoride) (PVDF). The effective membrane area of each module was 0.93 m2 (10 R2). Three modules were connected in series. PTGC modules (Millipore Corp., Marlborough, MA) were used as the ultrafiltration (UF) membrane. The nominal membrane molecular-weightcutoff was 10 000, and the membrane material was polysulfone. The effective membrane area of each module was 0.93 m2 (10 ft2). Five modules were connected in series. Jojoba Extraction. Reverse osmosis purified water was used in all the experiments. Jojoba meal (33.4 kg, particle size 2 mm) was mixed with 0.357 m3 of water, and the pH was adjusted to 8.0. The mixture was filtered through cheesecloth and the solids were extracted again with 0.272 m3 of water at pH 8.0. The combined extracts were centrifuged with a continuous centrifuge (3-in. bowl), at 12780g and a flow rate of 2 Umin (Sharples Co., Philadelphia, PA). The superna-
* To whom correspondence should be addressed.
This article not subject to U.S.Copyright. Published 1995 by the American Chemical Society
1780 Ind. Eng. Chem. Res., Vol. 34, No. 5, 1995
Strainer
Pump equipped with a variable speed motor Figure 1. Schematic flow diagram of the membrane apparatus.
tant was filtered through filter cloth for soy sauce processing (Nagata Brewing Machinery Co., Ltd., Sagamihara, Kanagawa, Japan) and used as the feed solution (0.418 m3) for the MF process. Clarification by MF. Recycling water in the unit, water permeate flux from the MF membrane, J,, was measured and water permeability, L,, was calculated using eq 1: Lp
=JvPav
(1)
where Pa, is the average pressure difference across the membrane and defined as:
Pa, = (Pin+ P0,,Y2
(2)
where Pi, and Poutare the pressures at the inlet and outlet of the membrane module, respectively. The water was then replaced with the jojoba extract, and total recirculation was performed by returning the permeate t o the feed tank and keeping the composition of the feed constant. Permeate flux from the membrane was measured under different operating conditions to determine optimum operating conditions. Under the optimum conditions, the permeate was collected outside the unit as a clarified extract. After the clarification treatment, water permeability was measured again. Protein Purification by UF. To remove sugar, nonprotein nitrogen, and toxicants from the clarified extract and purify the jojoba protein, constant volume diafiltration was performed with the UF membrane. Recycling water in the unit, water permeate flux from the UF membrane, J,, was measured and the water permeability, L,, was calculated using eq 1. The water was replaced with the clarified MF permeate and, using total recirculation, permeate flux was measured under different operating parameters, and the optimum operating conditions were identified. Under these conditions, the permeate was taken out of the unit, and constant volume diafiltration treatment was performed by adding water t o the feed tank t o keep the volume of the feed solution constant. The constant volume diafiltration treatment was continued until the total volume of the permeate exceeded three times the volume of the feed solution. Concentrating by UF. The constant volume diafiltration treatment was followed by batchwise concentration. To increase the concentration of jojoba protein in the feed, supplementation of water to the feed tank
was stopped, and the permeate was removed from the unit until a minimum feed volume was reached. After batchwise concentration, water permeability was measured again. The changes in permeate flux and total permeate volume were measured over time throughout the MF and UF experiments. In the preliminary experiment, a decrease in pH caused a sharp decline in permeate flux, and the extract became dark at a pH value higher than 8.5. The pH was, therefore, kept at 8.0 throughout the MF and UF experiments by adding NaOH solution. High temperature causes denaturation of protein, and low temperature usually results in a decrease in permeate flux (Nomura et al., 1987). The feed solution temperature was, therefore, kept at 20-25 "C during experiments with MF and UF. Cleaning Procedure. After the experiment, the membranes were cleaned to recover the original water permeability. The MF membrane was cleaned with 400 ppm of NaOCl for 40 min at 40 "C, and the UF membrane was cleaned with 0.1 N NaOH for 40 min at 40 "C and then with 400 ppm of NaOCl for 40 min at 40 "C. Freeze-Drying. After batchwise concentration, the retentate was freeze-dried with a Model 100-SRCfreezedrier (VIRTIS, Gardiner, NY). Isoelectric Point Precipitation. To compare the efficiency of isoelectric point precipitation with that of freeze-drying, a part of the retentate in the batchwise concentration treatment (500 mL) was acidified to pH 3.5, which is the isoelectric point of jojoba protein (Cardoso and Price, 1982; Wolf et al., 1988), using 1.0 N HC1, and then centrifuged for 60 min at 2100g, and the precipitate was freeze-dried. Analyses. Samples taken during the MF and UF experiments were tested for total solids, non-protein nitrogen, total nitrogen, and total sugars. Non-protein nitrogen was quantified as that nitrogen not precipitated with trichloroacetic acid. Total nitrogen was measured by using the Kjeldahl method (AOAC, 1984). The concentration of protein, Cpro, was calculated as in eq 3:
Cpro= 6.25(Ctn- Cnpn)
(3)
where Ct, and Cn,, are concentrations of total nitrogen and non-protein nitrogen, respectively. Total sugar was expressed as the concentration of dextrose (Dubois carbohydrate). Color of the protein was measured as absorbance at 400 nm of a 1%protein solution (0.16% nitrogen). If absorbance was greater than 1.0, the sample was diluted until the solution absorbance was 0.5-1.0, and the equivalent absorbance (in absorbance units, AU) for a 1%protein solution was calculated. Insoluble components were determined as the percentage of the sample not soluble in water at 5 g/lOO mL concentration.
Results and Discussion Optimization of MF and UF Operating Conditions. To operate the membrane unit stably, the Pin value should be lower than 345 kPa (50 psi) and higher than 103 kPa (15 psi) to avoid dry running or cavitation. The conditions which can be employed with the MF unit are, therefore, inside the hatched area in Figure 2a. Four different conditions inside the area (points 1-4 in Figure 2a) were tested to optimize the condition for MF treatment. Permeate flux rates through the MF mem-
Ind. Eng. Chem. Res., Vol. 34, No. 5, 1995 1781
a) MF 350 300
1
Pressure Control Valve is Fully Open
k. ... ...
7 250-
-
t
Ph =450kPa
*
:
200-
"0
50 100 150 200 250 300 350
"0
50
100 150 200 250 300 350
Pav [kpal
Pav
WI
Figure 2. Conditions tested in this study. Table 1. Effects of the Average Pressure Difference across the Membrane, Pay,and the Feed Flow Rate on Permeate Flux during MF' (permeate flux) point Pa, Pin- Pout flow rate no. (Wa) (US) x 106(m/S) (Wa) 1 2 3 4
94.5 241 276 310
86.9 207 138 69.0
1.72 2.14 1.59
5.87 6.00 6.32 5.83
brane under different conditions are listed in Table 1, where (Pi,- Pout)regulates the flow rate of feed solution in the module. This table indicates that the average pressure difference across the membrane, Pa,, had no effect on the value of permeate flux, if the value was larger than 94.5 kPa (13.7 psi). The flow rate of feed in the module also had no effect, if the value was higher than 1.59 U S[Pin - Pout= 69.0 kPa (10 psi)]. The pressure independent flux value implies that an increase in pressure develops a deposit layer on the membrane surface. Development of a deposit layer on an MF membrane sometimes causes formation of a dynamic membrane which can reject soluble proteins and other smaller solutes (Matsumoto et al., 1987; Nakao and Kimura, 1981; Nakao et al., 1982; Ohtani et al., 1985, 1987, 1988; Shoji et al., 1988; Watanabe et al., 1986,1988). Formation of a dynamic membrane on the MF membrane seems to decrease the recovery of soluble protein in the MF permeate. Based on these results, MF was carried out at Pa, = 94.5 kPa (13.7 psi) and Pi, - Pout= 86.9 kPa (12.6 psi). The conditions tested with the UF unit are shown in Figure 2b. To get a high permeate flux during UF, a large pressure difference across the membrane and high feed flow velocity in the module are usually required (Nabetani et al., 1990). The value of Pi,,therefore, was fixed at 345 kPa (50 psi), and Poutwas varied to optimize the feed flow rate in the membrane module (points 1-7 on the solid line in Figure 2b). Some other conditions (points 8-13 in Figure 2b) were also tested to confirm that the optimum condition is on the solid line in Figure 2b. Results in the total recirculation test are listed in Table 2. The maximum permeate flux value was obtained when (Pi,- Pout)was 138 kPa (20 psi). The UF treatments (the diafiltration and the batchwise concentration) were, therefore, performed under this condition.
Table 2. Effects of the Average Pressure Difference across the Membrane, Pa",and the Feed Flow Rate on Permeate Flux during TJF (permeate flux) P,n - Pout flow rate point Pa, no. (Pa) (US) x 106(m/S) (kPa) 1 2 3 4 5 6 7 8 9 10 11 12 13
172 190 207 224 241 276 310 155 207 259 103 155 51.7
345 310 276 241 207 138 69.0 310 207 103 207 103 103
0.49 0.48 0.47 0.41 0.39 0.29 0.18
8.39 8.99 9.75 10.29 10.81 11.69 11.57 8.03 9.43 10.43 5.49 7.10 2.47
Clarification with MF. The changes in permeate flux, volume reduction factor, and rejection rates for protein and sugar with time are shown in Figure 3. The volume reduction factor, VRF, was defined as
VRF = V,flf
= V,d(V,o- Vp)
(4)
where Vf,o,Vf and V, are the volumes of initial feed, feed, and total permeate, respectively. Rejection for the ith solute, R,, was calculated as
R,= 1 - Ci(in permeate)/Ci(in feed)
(5)
where Ci is concentration of the ith solute. In the middle of this experiment, the membrane was cleaned once. The jojoba extract was treated until the VRF became 9. About 90% of the jojoba extract volume (0.373 m3) was recovered in the MF permeate. The permeate flux values were 4 x 10-6-6 x m f s and had little dependence on the VRF, or concentration of solids in the feed. In the first half of the MF treatment, rejection for protein, Rpm, was almost constant and rejection for sugar, Rsug,was very low. However, in the latter half, not only Rpro,but also Rsug,gradually increased. This increased rejection could be caused by decreases in the mass of soluble proteins and sugars in the feed with time, that is, increased ratio of insoluble proteins and polysaccharides to total solids, or protein clogging of the membrane. A deposit layer of protein formed on an MF membrane sometimes rejects soluble protein and other
--
Ind. Eng. Chem. Res., Vol. 34,No. 5, 1995 Cleaning
20
*OF--I+- J"
-
DilutioyRq
+ 0.8 '
I
0.6 0.6
I
1
1 -f'
0.4 l... . . . ..........
I .
...........
!.................~.........................l
, o
'
o
2
~
I
4
! ~
, "
6
"
, i
, 8
'
,
,
t [hl Figure 3. Changes in permeate flux,J,, volume reduction factor, VRF, and rejection of protein, R, and sugar, Rsug, with time during microfiltration.
smaller solutes (Matsumoto et al., 1987; Nakao and Kimura, 1981; Nakao et al., 1982; Ohtani et al., 1985, 1987, 1988; Shoji et al., 1988; Watanabe et al., 1986, 1988). The effect of protein clogging was not negligible because cleaning in the middle of the experiment decreased the protein rejection. This implies that, a t the end of the MF treatment, not only insoluble protein but also soluble protein was rejected by a deposit layer formed on the MF membrane. In order to keep the value of R,, low and increase the recovery of soluble protein, the thickness of the deposit layer should be reduced. Although a higher feed flow rate in the module and a smaller pressure difference across the membrane should decrease the thickness of the deposit layer and reduce the rejection of solutes (Ohtani et al., 1988; Shoji et al., 1988;Watanabe et al., 1986),increasing the feed flow rate in this system caused increased pressure difference across the membrane. A design improvement in the system, which would combine high feed flow rate with small pressure difference across the membrane, would be required to keep the value of R , low and increase the recovery of soluble protein. A diafiltration process using MF might increase the recovery of soluble protein to some degree. However, it does not seem probable because the high rejection for protein (0.5-0.9) would require too much time and too large of a permeate volume for the MF process, resulting in a longer process time for the UF separation. Diafiltration. Changes in permeate flux, dilution ratio, rejection for protein, purity of protein in the feed, and concentrations of protein, sugar and non-protein nitrogen in the feed during constant volume diafiltration are shown in Figure 4. Dilution ratio, DR, was defined as
DR = VJVf
(6)
and the purity of protein was calculated as purity of protein = CprJCb where Ct, is the total solids concentration.
I
10
(7)
I
-A- R pro
I*
CnDn
Purity of Protein 9Cpro
- ~ tsug
I
Figure 4. Changes in permeate flux,J,, dilution ratio, rejection for protein, Rp,, purity of protein in the feed, and concentrations of protein, C,, sugar, C,,, and non-protein nitrogen, Cnpn, during the constant volume diafiltration treatment.
The R,, value was always unity, and C,,(in feed) was almost constant. The J, value was also constant throughout the experiment. Therefore, the dilution ratio increased linearly with time. The C,,,(in feed) value decreased from 0.59 to 0.16 kg/m3, and purity of protein increased from 0.20 to 0.66. Batchwise Concentration. Changes in permeate flux, volume reduction factor, rejection for protein, purity of protein in the feed, and concentrations of protein, sugar, and non-protein nitrogen in the feed during batchwise concentration treatment are shown in Figure 5. During the treatment, the purity of protein increased from 0.66 to 0.83, because sugar and non-protein nitrogen were removed with permeate. The rejection of the ultrafiltration membrane for protein was unity throughout the treatment. While the permeate flux decreased with time, the feed solution could be processed up to a VRF value of 14, and the concentration of protein in the feed increased from 3.28 to 52.9 kg/ m3. However, the concentration of sugar in the feed also increased slightly (from 0.4 to 2.4 kg/m3). This means that part of the sugar was rejected by the membrane. Rejection of sugars or polysaccharides by UF membranes is apt to be increased by protein clogging of the membrane: i.e., formation of a deposit layer on the UF membrane (Nakao and Kimura, 1981; Nakao et al., 1982) or adsorption of solutes to the membrane (Nabetani et al., 1988). There is a possibility that the purity of protein in the product could be increased by eliminating these fouling phenomena and decreasing the rejection of sugars. Addition of EDTA or some kind of salt might be effective. Cleaning. Water permeability of the MF membrane before the experiment was 5600 x m / ( s Pa). After the experiment it had decreased to 258 x m / ( s Pa). However, cleaning with 400 ppm of NaOCl restored the water permeability to its original value.
Ind. Eng. Chem. Res., Vol. 34, No. 5, 1995 1783 20
40
Jojoba Meal 33.4 kg Protein 7.35 kg &
I
Slurry and coarse filter
I-.
17 kg solids 2.75 kg protein
1
16.2 kg solubles 48% passes through 4.73 kg protein
centrifugation
1
6.22 kg solids 1.86 kg protein
8.76 kg solubles 2.33 kg protein
1
Microfiltration
1 I 1
I-.
Retentate 1.76 kg solids 0.64kg protein
Permeate 6.68 kg solids 1.32 kg protein
UF-Diafiltration
0
0.5
1.0
1.5
2.0
2.5
3.0
t [hl
I-.
Permeate 5.14 kg solids 0.08 kg protein
Retentate 1.50 kg solids 0.99 kg protein 0.0034 kg protein
Figure 5. Changes in permeate flux,Jv,volume reduction factor, VRF, rejection for protein, R,,, purity of protein in feed, and concentrations of protein, C ,, sugar, Csug,and non-protein nitrogen, Cnpn, during the batchwise concentration treatment.
Water permeability of the UF membrane decreased to 88 x m/(s Pa) during the from 412 x experiment. Cleaning with 0.1 N NaOH alone did not completely recover the water permeability. The membrane was, therefore, cleaned again with 400 ppm of NaOC1. The combination of NaOH and NaOCl cleaned the membrane well, and the original water permeability was recovered. The decreases in water permeabilities of the membranes were attributed to fouling phenomena, not to degradation of the membrane itself. Mass Balance. The mass balance obtained in this experiment is shown in Figure 6. In this experiment, 4.1% of the starting meal was recovered as product (1.38 kg of protein powder was obtained from 33.4 kg of jojoba meal.). In addition, about 75% of the starting meal was dried and recovered as water-extracted meal for animal feed or other uses. Sugars and simmondsin were separated in the UF permeate and constituted 16% of the starting meal. Extraction and recovery of proteins needs to be improved. Quality of Products. The final retentate protein had a color value of 2.20 AU. This could be described qualitatively as a light golden color a t 1%protein concentration and a dilute iced-tea brown a t 5% concentration. Although this color value is acceptable for some cosmetic uses, further research is now underway to improve the color value. At 5% weight per volume in distilled water adjusted to pH 6.6, a small amount of insoluble matter (2.4%)was centfiged solids at 22000g. The insoluble matter might be an aggregate of protein which was denatured during freeze-drying. The freezedried protein (final retentate) was 86.8% protein, 5.05% moisture, and 2.4% insoluble. Isoelectric Point Precipitation. Although isoelectric point precipitation gave some precipitate, the mass of the freeze-driedprecipitate was only 7.3% of the total solids contained in the solution. Purity of protein in the freeze-dried precipitate was 0.85, and the isoelectric
1
Retentate 1.30 kg solids 1.08 kg protein
Figure 6. Mass balance obtained in this experiment.
point precipitation did not increase this value and was judged to be ineffective for recovering solid phase protein from solution after the membrane process. Process Design for Industrial Use. To design an industrial-scale process, based on this experiment, the permeate flux values were first quantified. The permeate flux during MF treatment was assumed to be m / s . The permeate flux during constant a t 5.0 x UF treatment (batchwise concentration and diafiltration) was assumed to be a function of protein concentration in the feed. This assumption was based on the fact that the permeate flux during diafiltration did not change in spite of a change in concentration of sugar in the feed (Figure 4). In UF, a linear relationship between permeate flux and the log of the macromolecular concentration in the feed, C, is usually observed. The relationship can be expressed as
J , = k ln(Clim/C) where Clim is the concentration of macromolecules at the intercept of the linear relationship to Jv= 0 axis. This linear relationship can be explained theoreticallyby two different models. One is the gel-polarization model (Baker and Strathman, 1970; Blatt et al., 1970; Porter, 1972), and the other is the osmotic pressure model (Clifton et al., 1984; Goldsmith, 1971; Kozinski and Lightfoot, 1971,1972; Leung and Probstein, 1979;Mitra and Lundblad, 1978; Nabetani et al., 1990; Trettin and Doshi, 1981; Vilker et al., 1981; Wijmans et al., 1984, 1985). In the gel-polarization model, the permeate flux is thought to be limited by the hydraulic resistance of a gel layer formed on the membrane surface, and Clim accounts for the solute concentration in the gel-layer which does not depend on operating conditions such as pressure difference across the membrane. On the other hand, the osmotic pressure model indicates that the permeate flux is limited by the increase in osmotic
1784 Ind. Eng. Chem. Res., Vol. 34, No. 5, 1995 20
time required for the diafiltration treatment, t(for DF), could be calculated with the following equation:
i\
t(for DF) = DRVdfor DF)/{AJ,(for DF)} = DRC,ro,oVf,~{AC,ra(for DF)J,(for DF)} (12)
5' 0
:\
: : . : : : : ; /
1
10
100 Cpro
;::;:;I 1000
IWm31
Figure 7. Relationship between concentration of protein in the feed, Cpro,and permeate flux, J,.
pressure at the membrane surface, and C1im gives the solute concentration for which osmotic pressure of the solution equals the pressure difference across the membrane. In both models, k represents the mass transfer coefficient in the boundary layer on the membrane surface (Wijmans et al., 1984). The relationship between Cpro and J, obtained in this experiment is shown in Figure 7. From this result, the J, value was quantified as
J, = 3.97 x
lob6ln(120/Cpro)
(9)
The straight line in Figure 7 is described by eq 9. In our experiment, the intercept of the linear relationship to J, = 0 axis resulted in 120 kg/m3. A solution containing 120 kg/m3 jojoba protein did not give gel. Permeate flux during UF of jojoba protein might be, therefore, controlled by osmotic pressure at the membrane surface. Using eq 9, the batchwise concentration process could be expressed as dVddt = -AJ, = -A[3.97 x d(C,,Vf)/dt
ln(120/C)I =0
(10)
(11)
where A is membrane area. The curves in Figure 5 were obtained by solving eqs 10 and 11, where initial conditions were as follows
at t = 0: Vf = 0.362, C,,, = 3.28 The batchwise concentration process was fit well with eqs 10 and 11. Then, using eqs 9-11, a combination of a diafiltration process with a batchwise concentration process was optimized. In this experiment, the diafiltration process was followed by the batchwise concentration process. However, the batchwise concentration process could be followed by diafiltration. In order to resolve which operation should be done first, the time required for the UF process was calculated. If the initial volume of the feed solution, Vf,o,and the concentrations of protein at the beginning, CPro,o,and at the end of the are given, the time required for UF treatment, Cpro,~, the batchwise concentration could be calculated solving eqs 10 and 11. This time should not be changed by the existence of the diafiltration treatment, because the diafiltration treatment never changes the concentration of protein in the feed. If the diafiltration is performed in the middle of the batchwise concentration when the concentration of protein in the feed is C,,,(for DF), the
where Vdfor DF) and DR are the volume of the feed solution during the diafiltration treatment and the dilution ratio required for the treatment. The flux value during the diafiltration treatment, J,(for DF), could be obtained by substituting C,,(for DF) into eq 9. Ng et al. (1976) indicated that if there is a linear relationship between J,(for DF) and ln(C,,,(for DF)), eq 12 gives minimum time when C,,,(for DF) = Cpro,lim/e To minimize the time required for the UF treatment, the batchwise concentration treatment should be performed until C , becomes 44.1 kg/m3(=Cpm,lim/e).Then, the diafiltration treatment should be done. These treatments should be followed by further batchwise concentration. On the basis of the results described above, an industrial membrane process which can produce 100 kg of protein powder in 1 day can be described. Both the MF and UF units should be cleaned once a day. It takes 2 h to clean each unit. The MF and UF units are, therefore, operated 6 h/day. The final concentration of protein, Cpro,E, is 100 kg/m3, because, batchwise concentration beyond this point did not increase Cp, significantly (Figure 5). So, Vf,E in the ultrafiltration process is 1.0 m3. On the basis of the results obtained in this experiment, Cpro,oin the UF process was assumed to be 3.28 kg/m3, and Vf,ois 30.5 m3. The DR value is 3. To get 30.5 m3 of MF permeate within 6 h, the area of MF membrane should be 282 m2 mk 21 600 s)]. [=30.5 mY(5.0 x Equations 10-12 indicate that 234 m2 of UF membrane are required to concentrate the protein from 3.28 to 100 kg/m3 within 6 h. In the middle of the concentration treatment, at 44.1 kg/m3protein concentration, the solution should be diafiltrated to a DR value of 3. A schematic flow diagram of this UF process is shown in Figure 8. Effect of EDTA on the Performance of the Membrane System. Although the purities and qualities of the protein fractions obtained with the membrane processes were good enough for use in cosmetics, the recovery of jojoba protein was not optimized. To improve recovery, protein extractability from jojoba meal should be increased. Less rejection of protein by the MF membrane would also improve recovery. Higher permeate flux values for both MF and UF membranes would shorten the operating time and increase productivity in the process. Lower rejection of non-protein solutes by the UF membrane would increase productivity because it decreases the dilution ratio required for diafiltration. Abbott et al. (1991) reported that EDTA helped protein extraction from hexane-extracted jojoba meal. A large amount of residual EDTA might create a problem for cosmetics formulators, but EDTA is most probably removed in the UF permeate. EDTA solution was, therefore, tested in the extraction process to improve the recovery of protein from an optimized membrane process, and the effects of EDTA, not only on the extractability, but also on membrane perfor-
Ind. Eng. Chem. Res., Vol. 34, No. 5,1995 1785 Vf : 30.5 m3 C pro : 3.28 kg/m3
*
Initial batch-wise concentration
-
- _ _
10
20
Permeate 28.2 m 3
V f : 2.271~13 C pro : 44.1 kg/m3
Diatiltration Operation tlme: 2.03 h
Permeate 6.81 m 3
V f : 2.27m3 44.1 kg/m3
Cpro:
*
Final batch-wise concentration
" 00
Permeate 1.27 m 3
V f : l.00m3 C pro: 100 kg/m3 Figure 8. Schematic flow diagram of an ultrafiltration membrane process which can process 30.5m3 of microfiltration permeate and produce 100 kg of protein powder in a day. (Membrane area 234 m2, total operation time 6 Wd.)
mance (permeate flux and rejection ability for solutes) were investigated. Jojoba meal (27.0 kg, particle size < 2 mm) was mixed with 0.277 m3 of 0.1% EDTA (Sigma Chemical Company, St. Louis, MO) solution, and the pH adjusted to 8.0. The mixture was filtered through cheesecloth, and the solids were extracted again with 0.138 m3 of EDTA solution (0.1%)at pH 8.0. The combined extracts were centrifuged with the continuous centrifuge, and the supernatant was filtered through filter cloth for soy sauce processing as described previously. Then the filtrate was used as the feed solution (0.329 m3) for the membrane processes. In the UF process, the clarified solution was concentrated as much as possible, and then, to remove sugars and simmondsin and purify the protein, the concentrated solution was diaflitrated. The diafiltration treatment was performed until a dilution ratio value of 3 was achieved. The pH was kept at 8.0 throughout the MF and UF experiments using NaOH solution. 1. Clarification with MF' Process. Composition of the feed solution for the MF process is shown in Table 3 together with that of the solution without EDTA. Concentrations of all components analyzed were increased by addition of EDTA. However, the purity of protein, (C,,JC@),decreased from 0.263 t o 0.221. The changes in permeate flux, volume reduction factor, and rejection rates for protein and sugar with time are shown in Figure 9. The permeate flux values were about 5 x lov6d s and almost independent of the VRF, or concentration of solids in the feed. This value was almost same as the value observed with the solution without EDTA. Rejection values for both protein, Rpro,and sugar, Rsug, increased with time. These increases were also observed in the process without EDTA and are thought
1
L
1
2
3
4
5
6
t [hl
Figure 9. Changes in permeate flux,Jv,volume reduction factor, VRF, and rejection rates for protein, R,, and for sugar, Itsug, with time, t , during the MF treatment. (Experiment with EDTA.) Table 3. Composition of the Feed Solutions for the MF' ProcessO Cte Ctn Csug Cnpn Cpm Cprd (kg/m3) (kg/m3) (kg/m3) (kg/m3) (kg/m3) Ct, withEDTA 28.2 1.42 2.21 0.429 6.22 0.221 without EDTA 20.5 1.00 1.67 0.138 5.40 0.263 a Cb, Ct,, Caug,Cnpn,and C, are concentrations of total solid, total nitrogen, total sugar, non-protein nitrogen, and protein, respectively.
t o be caused by deposit layers formed on the MF membrane (Matsumoto et al., 1987;Nakao and Kimura, 1981;Nakao et al., 1982; Ohtani et al., 1985,1987,1988; Shoji et al., 1988; Watanabe et al., 1986, 1988). Therefore, EDTA was found to be ineffective in decreasing protein clogging of membranes (formation of a deposit layer). After the MF process, 56% of the protein contained in the original feed solution was recovered in the permeate solution. This value was the same as was observed in the previous experiment without EDTA and was not high enough. Some other method is needed to reduce the rejection for protein and improve the recovery. 2. Batchwise Concentration by UF Process. Changes in permeate flux, volume reduction factor, rejection for protein, purity of protein in the feed, and concentrations of protein, sugar, and non-protein nitrogen in the feed during batchwise concentration treatment are shown in Figure 10. During the treatment, the purity of protein, (C,,JCt,), increased from 0.18 to 0.55,because of the removal of sugar and non-protein nitrogen with permeate. The rejection of the UF membrane for protein was unity throughout the treatment. While the permeate flux decreased with time, the feed solution could be processed up t o a VRF value of 6.4, and the concentration of protein in the feed increased from 3.55 t o 25.9 kgl m3. However, the concentration of sugar in the feed also increased slightly (from 1.62 to 3.06 kg/m3). This means that part of the sugar was rejected by the membrane. The same results had been obtained in this process without EDTA. EDTA did not affect the rejection ability of the UF membrane.
1786 Ind. Eng. Chem. Res., Vol. 34,No. 5, 1995 10
10
c
5/5
a
* Dilution Ratio
1
4
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2
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0
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4
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.-0
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IO 0
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10
15
20
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0 25
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Figure 10. Changes in permeate flux, J,, volume reduction factor, VRF, rejection for protein, Rpm,purity of protein in the feed, (CpJ Ca),and concentrations of protein, C,,, sugar, Csug,and nonprotein nitrogen, Cnpn, in the feed with time, t , during the
Figure 12. Changes in permeate flux, J,, dilution ratio, DR, rejection for protein, Rpm,purity of protein, (C,,dCb), in the feed, and concentrations of protein, C,,, sugar, Csug,and non-protein nitrogen, Cnpn,in the feed with time, t, during the constant volume diafiltration treatment. (Experiment with EDTA.)
batchwise concentration treatment. (Experiment with EDTA.) 12 10
X
$
4 \
2 -
\ \
03 1
10 Cpro
100
~ m 3 I
Figure 11. Relationship between protein concentration in the feed, C,,, and permeate flux, Jv, during the batchwise concentration treatment.
The relationship between protein concentration in the feed, C,,,, and permeate flux, J,, is shown in Figure 11. The J , value decreased linearly with In Cpm,and the relationship could be expressed as
J , = 3.97 x
ln(44.5/CP,,)
(14)
The solid line in Figure 11 represents eq 14,and the dotted line indicates the relationship obtained with the solution without EDTA. Addition of EDTA decreased permeate flux during batchwise concentration. Figure 11 shows that addition of EDTA did not change the slope of the linear relationship; only the intercept at the J , = 0 axis was changed. This means that the mass transfer coefficient was not affected but osmotic pressure of the solution was increased by EDTA, because the permeate flux seems to be controlled by the osmotic pressure at the membrane surface as described above. This phenomenon could be described as follows. Addition of EDTA increased the solubility of jojoba protein and dimer proteins or larger proteins in the solution
changed to monomer proteins resulting in a marked increase in osmotic pressure. On the other hand, the mass transfer coefficient was not affected by the dissociation so much. For example, the diffusion coefficient of p-lactoglobulin increases 25% when it changes from dimer to monomer (Nippon Seikagakukai, 1981). According to the LBvCique equation (Blatt et al., 1970),a 25% increase in the diffusion coefficient results in only a 16% increase in the mass transfer coefficient. Using eq 14 and the model developed previously, changes in Cpro, VRF, and J , with time, t , were estimated. The curves shown in Figure 10 are calculated from the model. The calculated curves were in good agreement with the experimental values, and the model explained the batchwise concentration process well. The validity of the model was, therefore, confirmed again in the process with EDTA added. 3. Diafiltration by UF. Changes in permeate flux, dilution ratio, rejection for protein, purity of protein in the feed, and concentrations of protein, sugar, and nonprotein nitrogen in the feed during constant volume diafiltration are shown in Figure 12. The R, value was always unity, and Cpm(infeed) was almost constant. The J , value slightly decreased with time. In the experiment without EDTA, the J , value was constant during the diafiltration treatment. Decrease in concentration of EDTA during diafiltration treatment seemed to cause precipitation of the solutes which could be extracted only by EDTA solution, not by water. This precipitate might have clogged the UF membrane and decreased the J , value. The C,,(in feed) value decreased from 0.31 to 0.07 kg/m3 and purity of protein, (C,,JCt,), increased from 0.55 to 0.81 during the treatment. 4. Recovery and Quality of the Protein. In this experiment, 5.7% of the starting meal was recovered as product (1.54kg of protein powder was obtained from 27.0 kg of jojoba meal). In the experiment without EDTA, the product amounted to 4.1% of the starting meal. Addition of EDTA increased the recovery by 39%. The purity of the protein (0.81)was not changed by the addition of EDTA and increased recovery. The final
Ind. Eng. Chem. Res., Vol. 34, No. 5, 1995 1787 Table 4. Water Permeability of the Membranes before the Exueriment,L,. and after the Experiment, L,’ L, x 1012 L,’ x 1012 [m/(s Pall I d ( s Pall L,’IL, MF membrane 5600 258 0.0461 without EDTA 5550 181 0.0326 with EDTA UF membrane 0.214 without EDTA 412 88 398 56.5 0.142 with EDTA
retentate protein had a color value of 1.06; the color of the protein from the process without EDTA was 2.20. The amount of insoluble matter which was precipitated a t 22000g decreased from 2.4% to 1.1%by the addition of EDTA. 5. Membrane Fouling. The water permeabilities of the membranes before and after the experiments are shown in Table 4, together with the ratios of water permeability after the experiment to that before the experiment, (Lpl/Lp). The addition of EDTA decreased (Lpl/Lp)values of both MF and UF membranes. In a membrane process for protein solutions, decrease in water permeability, which is one of the primary factors decreasing permeate flux, is caused mainly by protein clogging of membranes, i.e., formation of a deposit layer on the membrane surface or adsorption to the membrane (Fane et al., 1983; Koutake et al., 1987, 1992, 1993; Matsumoto et al., 1987; Matthiasson, 1983; Nabetani et al., 1988, 1990). The process with EDTA appears t o cause increased clogging of membranes. The water permeabilities of the membranes were recovered with the cleaning procedure described previously.
Conclusions Permeate flux during microfiltration treatment was found to be almost independent of concentration of solids in the feed. On the other hand, permeate flux during ultrafiltration was a function of protein concentration in the feed. Based on the results, a model to express the efficiency of a batchwise concentration process with an ultrafiltration membrane was developed. Using this model, the combination of batchwise concentration with diafiltration was optimized and an industrial-scale process was designed. The addition of EDTA increased the concentration of protein in the extract and improved the recovery of protein in the final product. The permeate flux values for the ultrafiltration membrane were decreased by EDTA, although those for the microfiltration membrane were not changed. Solute rejection values of the ultrafiltration and microfiltration membranes were not changed by the addition of EDTA. Water permeabilities of the membranes were lowered by EDTA, implying that EDTA increased protein clogging of the membranes (formation of a deposit layer on the membrane surface or adsorption to the membrane). In both microfiltration and ultrafiltration processes, the protein clogging of membranes was found to have large effects on the performance of the membranes. However, the mechanism of protein clogging of membranes has not been established. Acknowledgment This research was supported in part by Cooperative Research and Development Agreement No. 58-3K95-2109 with the Biotechnology Research and Development Corporation. The use of brand or trade names may be
necessary to report factually on available data. The USDA neither guarantees nor warrants the standard of the product, and the use of the name by USDA implies no approval of the product t o the exclusion of others that may also be suitable. This work was presented, in part, a t the American Institute of Chemical Engineers Meeting, Atlanta, GA, April 17-21,1994.
Nomenclature A = membrane area [m21 C = concentration [kg/m31 DR = dilution ratio J, = volume flux through membrane [ d s l J, = volume flux of water [m/s] L, = water permeability [m/b Pall L,’ = water permeability decreased by fouling [m/b Pall P = pressure [Pal R = rejection t = time Esl V = volume [m31 VRF = volume reduction factor Subscripts E = at the end f = feed i = the ith solute in = inlet of the membrane module npn = non-protein nitrogen out = outlet of the membrane module p = permeate pro = protein sug = sugar tn = total nitrogen ts = total solid 0 = at the beginning
Literature Cited Abbott, T. P.; Nakamura L. K.; Buchholz, G.; Wolf, W. J.; Palmer, D. M.; Gasdorf, H. J.; Nelsen, T. C.; Kleiman, R. Processes for making animal feed and protein isolates from jojoba meal. J . Agric. Food Chem. 1991,39 (81,1488-1493. AOAC. Official Methods of Analysis, 14th ed.; Association of Official Analytical Chemists: Arlington, VA, 1984. Baker, R. W.;Strathman, H. Ultrafiltration of macromolecular solutions with high . Sei. 1970, - flux membranes. J . h - ~_ lPolrm. 14, 1197-1214. Blatt, W. F.; Dravid, A.; Michaels, A. S.; Nelson, L. Solute Dolarization and cake formation in membrane ultrafiltration: bauses, consequences and control techniques. In Membrane Science and Technology; Flinn, J. E., Ed.; Plenum Press: New York, 1970;p 47. Cardoso, F. A.; Price, R. L. Extraction, characterization and functional properties of jojoba proteins. In Proceedings of the Fourth International Conferenceon Jojoba, Hermosillo, Mexico; Consejo Nacional de Ciencia y Technologia: Mexico, D.F., 1982; p 305. Clifton, M. J.;Abidine, N.; Aptel, P.; Sanchez, V. Growth of the polarization layer in ultrafiltration with hollow-fiber membranes. J . Membr. Sei. 1984,21, 233-246. Cokelaere, M. M.; Dangrean, H. D.; Amouts, S.; Kiihn, E. R.; Decuyere, E. M.-P. Influence of pure simmondsin on the food intake in rats. J . Agric. Food Chem. 1992a, 40, 1839-1842. Cokelaere, M. M.;Dangrean, H. D.; Daenens, P.; Bruneel, N.; Amouts, S.; Decuyere, E. M.-P.; Kiihn, E. R. Investigation of possible toxicological influence of simmondsin after subacute administration in the rat. J . Agric. Food Chem. 199213, 40, 2443-2445. Cokelaere, M. M.;Buyse, J.; Daenens, P.; Decuypere, E.; Kiihn, E.; Boven, M. V. Influence of jojoba meal supplementation of growth and organ function in rats. J.Aerie. - Food Chem. 1993a, 41, 1444-1448. Cokelaere. M. M.: Buvse. J.: Daenens. P.; Decuvuere, E.; Kuhn. E.; Boven, M. V. Fertility in rats after long-term jojoba meai supplementation. J.Agric. Food Chem. 1993b,41,1449-1451.
1788 Ind. Eng. Chem. Res., Vol. 34,No. 5 , 1995 Elliger, C. A,; Waiss, A. C.; Lundin, R. E. Simmondsin, an unusual 2-cyanomethylenecyclohexylglucoside from Simmondsia californica. J. Chem. Soc., Perkin Trans. 1 1973,19,2209-2212. Elliger, C. A.; Waiss, A. C.; Lundin, R. E. Structure and stereochemistry of simmondsin. J. Org. Chem. 1974a,39,2930-2931. Elliger, C. A.; Waiss, A. C.; Lundin, R. E. Cyanomethylenecyclohexyl glucosides from Simmondsia californica. Phytochemistry 1974b,13, 2319-2320. Fane, A. G.; Fell, C. D. J.; Water, A. G. Ultrafiltration of protein solutions through partially permeable membranes-the effect of adsorption and solution environment, J. Membr. Sei. 1983, 16, 211. Gillis, A. Developing new commercial crops. J.An. Oil Chem. Soc. 1988,65 (11,6-20. Goldsmith, R. L. Macromolecular ultrafiltration with microporous membranes. Znd. Eng. Chem. Fundam. 1971,10,113-120. Iwasaki, K. JOJOBA: Its economic feasibility and the possibility of its cultivation in developing - - countries. Sakyu-Kenkyu 1985, 28 (2),84-96. Koutake, M.; Uchida, Y.; Sato, T.; Shimoda, K.; Watanabe, A.; Nakao, S. Filtration membrane fouling in ultrafiltration of skim milk, I. Causes and cleaning. Nippon fiogeikagaku Kaishi 1987, 61 (6), 677-681. Koutake, M.; Matsuno, I.; Nabetani, H.; Nakajima, M.; Watanabe, A. Classification of resistance to permeation caused by fouling during ultrafiltration of whey and skim milk. Biosci. Bwtechnol. Biochem. 1992,56(5),697-700. Koutake, M.; Matsuno, I.; Nabetani, H.; Nakajima, M.; Watanabe, A. Osmotic pressure model of membrane fouling applied to the ultrafiltration of whey. J. Food Eng. 1993,18, 313-334. Kozinski, A. A,; Lightfoot, E. N. Ultrafiltration of proteins in stagnation flow. AZChE J. 1971,17, 81-85. Kozinski, A. A.; Lightfoot, E. N. Ultrafiltration: A general example of boundary layer filtration. AIChE J. 1972,18, 1030-1040. Leung, W. F.; Probstein, R. F. Low polarization in laminar ultrafiltration of macromolecular solution. Znd. Chem. Eng. Fundam. 1979,18,274-278. Manos, C. G.; Schrynemeeckers, P. J.; Hogue, D. E.; Telford, J. N.; Stoewsand, G. S.; Beerman, D. H.; Babish, J. G.; Blue, J. T.; Shane, B. S.; Link, D. J. Toxicologic studies with lambs fed jojoba meal supplemented rations. J. Agric. Food Chem. 1986, 34, 801-805. Matsumoto, Y.; Nakao, S.; Kimura, S. Cross-flow filtration of polymer solutions by ceramic microfiltration membranes. Kagaku Kogaku Ronbunshu 1987,13(l),100-106. Matthiasson, E. The role of macromolecular adsorption in fouling of ultrafiltration membranes. J. Membr. Sei. 1983,16,23-36. Mitra, G.; Lundblad, J. Ultrafiltration of immune serum globulin and human serum albumin: Regression analysis studies. Sep. Sei. Technol. 1978,13, 89-94. Nabetani, H.; Nakajima, M.; Watanabe, A.; Nakao, S.; Kimura, S. Change of permeate flux and solute rejection by ovalbumin adsorption on ultrafiltration membranes. Membrane 1988,13, (1)51-57. Nabetani, H.; Nakajima, M.; Watanabe, A.; Nakao, S.; Kimura, S. Effects of osmotic pressure and adsorption on ultrafiltration of ovalbumin. AZChE J. 1990,36(61, 907-915. Nakao, S.;Kimura, S. Effect of gel layer on rejection and fractionation of different-molecular-weight solutes by ultrafiltration. In Synthetic Membranes, Vol ZZ, Hyper- and Ultrafiltration Uses; Turbak, A. F., Ed.; ACS Symposium Series 154; American Chemical Society: Washington, DC, 1981;p 119. Nakao, S.; Yumoto, S.; Kimura, S. Analysis of rejection characteristics of macromolecular gel layer for low molecular weight solutes in ultrafiltration. J. Chem. Eng. Jpn. 1982,15(6), 463468. Nippon Seikagakukai, Ed. Seikagaku Data Book Z; Tokyo Kagaku Dojin Co.: Tokyo, 1981;p 126 (in Japanese). Nomura, T; Nakao, S.; Kimura, S. Influence of feed temperature on ultrafiltration performance. Kagaku Kogaku Ronbinshu 1987,13,811-817. Ohtani, T.; Watanabe, A.; Hoshino, C.; Kimura, S. Application of dynamic membrane to ultrafiltration. Kagaku Kogaku Ronbunshu 1985,ll(2),140-146.
Ohtani, T.; Ohi, T.; Horikita, H.; Nakajima, M.; Nabetani, H.; Watanabe, A. Recovery of B-amylase from sweet potato with self-rejection type of dynamic membrane. Nippon Shokuhin Kogyo Gakkaishi 1987,34(lo),640-646. Ohtani, T.;Nakajima, M.; Nabetani, H.; Nawa, Y.; Watanabe, A. Formation and nature of ovalbumin dynamic ultrafiltration membrane. Nippon Shokuhin Kogyo Gakkaishi 1988,35(12), 807-812. Porter, M. C. Concentration polarization with membrane ultrafiltration. Znd. Eng. Chem. Prod. Res. Dev. 1972,11,234-248. Shoji, T.; Nakajima, M.; Nabetani, H.; Ohtani, T.; Watanabe, A. Effect of pore size of ceramic support on the self-rejection characteristics of the dynamic membrane formed with water soluble proteins in waste water. Nippon Nogeikagaku Kaishi 1988,62 (7),1055-1060. Spencer, G. F.; Plattner, R. D. Compositional analysis of natural wax ester mixtures by tandem mass spectroscopy. J . Am. Oil Chem. Soc. 1984,61, 90-94. Trettin, D. R.;Doshi, M. R. Pressure-independent ultrafiltration-Is it gel limited or osmotic pressure limited? In Synthetic Membranes Volume ZZ, Hyper- and Ultrafiltration Uses, Turbak, A. F., Ed.; ACS Symposium Series 154; American Chemical Society: Washington, DC, 1981;p 373. Verbiscar, A. J.; Banigan, T. F. Composition of jojoba seeds and foliage. J. Agric. Food Chem. 1978,26, 1456-1459. Verbiscar, A. J.; Banigan, T. F.; Weber, C. W.; Reid, B. L.; Trei, J. E.; Nelson, E. D. Detoxification and analyes of jojoba meal. In Proceedings of the Third International Conference on Jojoba, Yermanos, D. M., Ed.; University of California: Riverside, CA, 1979;p 185. Vilker, V. L.; Colton, C. K.; Smith, K. Concentration polarization in protein ultrafiltration Part 11: Theoretical and experimental study of albumin ultrafiltrated in an unstirred cell. AZChE J . 1981,27,637-645. Watanabe, A.; Komazawa, IC;Nabetani, H.; Nakajima, M.; Nakao, S. Concentration of polyphenols in alkaline extracts from larch bark with self-rejection type of dynamically formed membrane. Membrane 1986,11 (2),109-114. Watanabe, A.; Shoji, T.; Nakajima, M.; Nabetani, H.; Ohtani, T. Electronmicroscopic observation of self-rejection-type dynamic membrane formed with water soluble proteins in waste water from fish paste process. Nippon Nogeikagaku Kaishi 1988,62 (7),1061-1066. Weber, C. W.; Berry, J . W.; Cook, E. M. Influence of jojoba meal upon growth and reproduction in mice. In Jojoba and Its Uses Through 1982,Proceedings of the Fifth International Conference on Jojoba and Its Uses; Elias-Cesnik, A., Ed.; Ofiice of Arid Lands Studies, University of Arizona: Tucson, AZ,1983;p 93. Wijmans, J. G.; Nakao, S.; Smolders, C. A. Flux limitation in ultrafiltration: Osmotic pressure model and gel layer model. J. Membr. Sei. 1984,20,115-124. Wijmans, J . G.; Nakao, S.; van den Berg, J. W. A.; Troelstra, F. R.; Smolders, C. A. Hydrodynamic resistance of concentration polarization boundary layer in ultrafiltration. J. Membr. Sci. 1985,22,117-135. Wolf, W. J.; Schaer, M. L.; Abbott, T. P. Protein extractability of defatted jojoba meals: Effect of pH and salt concentration. In Proceedings of the Seventh International Conference on Jojoba and Its Uses; Baldwin, A. R., Ed.; American Oil Chemists’ Society: Champaign, IL, 1988;p 430. Wolf, W. J.; Schaer, M. L.; Abbott, T. P. Nonprotein nitrogen content of defatted jojoba meals. J. Sei. Food Agric. 1994,in press. Received for review August 9, 1994 Revised manuscript received February 8, 1995 Accepted February 23, 1995@
IE9404801
Abstract published in Advance ACS Abstracts, April 15, 1995. @
