430
Ind. Eng. Chem. Res. 1997, 36, 430-435
Mathematical Modeling of Sunflower Seed Extraction by Supercritical CO2 M. Perrut,† J. Y. Clavier,† M. Poletto,‡ and E. Reverchon*,‡ Separex, B.P. 9, F-54250 Champigneulles, France, and Dipartimento di Ingegneria Chimica e Alimentare, Universita` di Salerno, Via Ponte Don Melillo, I-84084 Fisciano (SA), Italy
Extraction of oil from crushed sunflower seeds with supercritical CO2 was performed at 280 bar and 40 °C on a laboratory apparatus of 0.15 × 10-3 m3 volume and on a pilot plant of 1.5 × 10-3 m3 volume. CO2 flow rate ranged between 5 and 45 kg h-1. To study the influence of the vegetable matrix on the extraction rate, a re-extraction experiment on the extraction product was also performed on the pilot scale apparatus. A mathematical model of the supercritical extraction process was developed on the basis of the experimental evidence. The equilibrium between the solids and the fluid phase appeared to be the controlling step during the extraction process. A simplified form of a sigmoidal-shaped equilibrium curve was adopted to fit the experimental data in the whole range of CO2 flow rates explored. The meaning of this nonlinear equilibrium relationship was also discussed. Introduction From the processing point of view, supercritical fluid extraction (SFE) of seed oil has been studied by several authors, and a wide range of seed species has been explored: wheat germ (Taniguki et al., 1985), oats (Fors and Ericksson, 1990), corn germ (Christianson et al., 1984), cottonseed (Stahl et al., 1980; List et al., 1984; Snyder et al., 1994), soybean (Stahl et al., 1980; Friedrich and List, 1982; Eggers et al., 1985; Eisenbach, 1988; Snyder et al., 1994), evening primrose (Favati et al., 1991), jojoba (Stahl et al., 1984), rice bran (Ramsay et al., 1991), rapeseed (Eggers et al., 1985), peanut (Snyder et al., 1994), and grape seed (Molero Gomez et al., 1994). In particular, sunflower seed oil extraction was tested by Stahl et al. (1980) at temperatures of 20 and 40 °C and at pressures up to 700 bar. They concluded that concentration of oil in CO2 is independent of the solvent flow rate. Despite the large number of seed species processed, only some models of SFE of seeds have been published. They all agree with the fact that at least an early part of the SFE process of seed oils is governed by the solubility equilibrium between the oil and the fluid phase. The equilibrium relationship has been generally supposed to be linear since other information on the equilibrium is not available in such complex systems. However, different descriptions have been proposed to model the SFE process. From the mathematical point of view, all models proposed show the same approach based on differential mass balance integration. Bulley et al. (1984), Lee et al. (1986), and Fattori et al. (1988) assumed that mass transfer resistance occurred only in the solvent phase. Sovova´ (1994), Sovova´ et al. (1994), and Goto et al. (1994) made a distinction in the solid phase between broken and intact cells which were considered with two separate mass balances and with different mass transfer resistances: in the solvent phase for the former kind of cells and in the solid phase for the latter. King and Catchpole (1993) used a shrinking core model to describe * Author to whom correspondence should be addressed. Phone: +[39](89)964116. Fax: +[39](89)964057. E-mail:
[email protected]. † Separex. ‡ Universita ` di Salerno. S0888-5885(96)00354-5 CCC: $14.00
a variable external resistance where the solute balance on the solid phase determines the thickness of the mass transfer layer in the external part of the particles. Almost all the authors simplified the system of partial differential equations deriving from the differential mass balances to obtain an analytical solution. The only exception is represented by Bulley et al. (1984), who solved numerically their model with the method of characteristics. Despite the proliferation of small-scale studies, the scale-up of laboratory results is very difficult. Indeed, scale-up calculation are possible when the extraction results are available at different scales of operation and the mathematical model of the process is available. Therefore, the scope of this study was to perform the supercritical fluid extraction of sunflower seed oil at two operation scales spaced by 1 order of magnitude of extraction volume (0.15 × 10-3 and 1.5 × 10-3 m3). The extraction was studied at different CO2 flow rates, and the experimental results were used to develop a model of the seed oil extraction process. Mathematical Modeling General Hypotheses. Some hypotheses were made to model the extraction system. These are listed in the following: (a) Several components are generally involved in the extraction. However, we supposed that their behaviors with respect to the mass transfer phenomena are similar enough to be described by a single pseudocomponent, which will be called the “solute”. (b) The commonly accepted continuous description of the extraction bed has been assumed, with the implicit hypothesis that the relevant concentration gradients in the fluid phase develop at larger scales than the particle size. The solute concentration in the fluid phase depends only on time, t, and on the axial coordinate, z. Its value, y, is given in terms of solid mass per unit of solvent mass. (c) The solvent flow rate, with interstitial velocity u, is uniformly distributed in every section of the extractor. The pressure drop can be neglected, as well as temperature gradients within the column. Axial dispersion is negligible. (d) The solute concentration in the solid, x, is expressed in terms of solute mass per unit mass of © 1997 American Chemical Society
Ind. Eng. Chem. Res., Vol. 36, No. 2, 1997 431
nonsoluble solid. The value of x is an average value within the particle and depends on t and z. Concentration gradients within the particles are not resolved in this model. (e) The volume fraction of the fluid, , is not affected by the reduction of the solid mass during the extraction. Mass Balance Equations. According to the above hypotheses, the balance of the solute in the fluid phase is
F f
(∂y∂t + u∂y∂z) ) J ∂x ) -J ∂t
(2)
where Ff is the fluid density which is supposed not to be affected by the presence of the solute, Fs is the bulk density of the nonsoluble solid, and J is the solute exchange rate between the phases. The expression of J is strongly dependent on the structure and on the preprocessing operations performed on the natural extraction matrix. Schematically, we can ascribe the mass transfer resistance to the diffusion phenomena inside the particles and to the fluid dynamic mass transfer mechanisms outside the particles. The diffusion phenomena inside the particles can be either in the solid phase, if the diffusion phenomena concern the solute combined with the solid, or in the fluid phase, if the diffusion phenomena concern the solute dissolved in the fluid phase permeating the solid matrix. In the former hypothesis, the mass transfer coefficient should be independent of the solvent flow rate through the extractor. In the latter hypothesis, the mass transfer coefficient can be affected by the solvent flow rate inside the extractor or not, depending on whether the solvent permeating the solid matrix is affected or not by the solvent motion inside the extractor. Due to the large quantity of oil available in the seeds, we supposed that the presence of an external porous layer of solids was very likely, and therefore, we supposed that the diffusion took place in the fluid phase. Consequently, we put
J ) aphFf(y* - y)
(4)
where f is a function depending on temperature and pressure. Substitution of (3) in (1) and (2) leads to
Ff
(∂y∂t + u∂y∂z) ) a hF (y* - y)
(5)
∂x ) -aphFf(y* - y) ∂t
(6)
Fs(1 - )
p
f
The system of eqs 5 and 6 has a unique solution when
b.c., y|z)0 ) 0 (7)
Moreover, we supposed that the loading time of the solvent into the extractor was long enough to enable the fluid to reach the equilibrium concentration before the extraction started. Therefore, we have
y0 ) f(x0)
(8)
This means that, before extraction starts, a part of the solute in the untreated solid saturates the fluid. The solute mass balance relates x0 and y0 to xu, the solute concentration in the untreated solid:
xuFs(1 - ) ) x0Fs(1 - ) + y0Ff
(9)
Numerical Solution. The system of eqs 1 and 2 is hyperbolic and linear, has constant coefficients, and is not homogeneous. The characteristic curves of the system of eqs 1 and 2 belong to two bundles of parallel lines, zI ) zI(t) and zII ) zII(t), on which the total differential of the variables y and x lie. Therefore, in this simple case, the Riemann invariants are y and x themselves; that is, y ) ψ′(t) ) ψ(zI(t),t) and x ) ξ′(t) ) ξ(zII(t),t). We have
dψ′ ∂ψ ∂ψ dzI ) + ∂t ∂t ∂z dt
(10)
dzI )u dt
(11)
We put
and use (10) and (11) in (1) to obtain
dy J(x,y) ) dt Ff
on
zI ) R + ut
(12)
where the second equation is the integral of (11) and R is an integration constant. Similarly,
(3)
where ap is the specific mass exchange surface between the phases, h is the mass transfer coefficient, and y* is the local solute concentration at the interface with the solid solute. The thickness of the oil seed is such that the hypothesis of constant mass transfer coefficient should apply. Therefore, we assumed that there was no resistance in the solid phase and that there was an equilibrium relationship between x and y*:
y* ) f(x)
i.c., x|t)0 ) x0, y|t)0 ) y0;
(1)
and that in the solid phase is
Fs(1 - )
the initial conditions (i.c.) on x and y and the boundary condition (b.c.) on y are given:
dξ′ ∂ξ ∂ξ dzII ) + dt ∂t ∂z dt
(13)
dzII )0 dt
(14)
In turn, we put
and use (13) and (14) in (2) to obtain
J(x,y) dx )dt Fs(1 - )
on
zII ) β
(15)
where the second equation is the integral of (14) and β is another integration constant. The system of eqs 12 and 15 can be solved numerically on a grid of points defined by the intersection of two sets of equally spaced parallel lines, each set belonging to one of the two different bundles of characteristics. The time increment of the variable y along the first bundle of characteristics given by the second of eqs 14 and that of the variable x along the second bundle of characteristics given by the second of eqs 15 can be obtained with a single step of an ODE integration algorithm. In the present work, a
432 Ind. Eng. Chem. Res., Vol. 36, No. 2, 1997
a
Figure 1. Extraction yield as a function of extraction time. Extractor of 0.15 × 10-3 m3 volume: f, 5 kg h-1. Extractor of 1.5 × 10-3 m3 volume: 9, 5; [, 10; 2, 20; 1, 25; and b, 45 kg h-1. s, Model predictions.
b
fourth-order Runge-Kutta algorithm was used. The initial condition used in the integration was taken from eqs 7 and 8:
i.c., t ) 0, x ) x0, y ) f(x0) ) y0
(16)
Experimental Section Two extraction units were used throughout this study, both located at Separex (Nancy, France). The first unit had an extractor of 0.15 × 10-3 m3 internal volume (length (L) ) 155 mm, internal diameter (i.d.) ) 35 mm) and has been operated at a CO2 flow rate of 5 kg h-1. In each extraction test, 50 g of crushed sunflower seeds was charged into the extractor. The second unit was equipped with an extractor of 1.5 × 10-3 m3 internal volume (L ) 290 mm and i.d. ) 82 mm) and was used at CO2 flow rates of 5, 10, 15, 20, and 25 kg/h. A charge of 550 g of crushed seeds was loaded into this extractor before each run. Both units were operated at 280 bar and 40 °C. The mean particle size of the crushed seeds was about 3 mm, and a wide distribution of particle diameters was observed. Indeed, a fraction of about 50% of the particles was smaller than 1 mm, and a fraction of 15% was even smaller than 0.5 mm.
Figure 2. Extraction yield as a function of the specific mass of solvent. Extractor of 0.15 × 10-3 m3 volume: f, 5 kg h-1. Extractor of 1.5 × 10-3 m3 volume: 9, 5; [, 10; 2, 20; 1, 25; and b, 45 kg h-1. s, Model predictions.
Results Experimental Data. Results of the extraction tests are reported in Figure 1, where the yield, Y, i.e., the quantity of oil extracted divided by the weight of the initial charge, is given as a function of the extraction time, t. The same results are shown in Figure 2, where the yield is given as a function of the specific mass of solvent used, i.e., the ratio between the mass of CO2 used and the mass of the seeds charged in the extractor. In this figures, we can observe that experimental results at different solvent flow rates coincide during the first part of the extraction process. Therefore, we can conclude that, during this part of the process, the extraction yield does not depend on the solvent flow rate. For extraction yields larger than 0.22, this is no more true, since the extraction curves in Figure 2 show a considerable spreading for these values of the yield. For very large extraction times, we can identify an asymptotic value of all extraction curves of about 0.33. Besides the ordinary extraction tests, a further experiment was performed with a solvent flow rate of 10 kg h-1 on the 1.5 × 10-3 m3 apparatus in which the oil extracted during the conventional extraction test had
Figure 3. Mass of solute extracted as a function of time. Extractor of 1.5 × 10-3 m3 volume, 10 kg h-1 flow rate: [, seed extraction; +, deposited solute re-extraction.
been spread over a filling of PALL rings. In Figure 3, the results of this test are compared with those of the original extraction from seeds, where the mass of solute extracted, me, is given as a function of extraction time, t. In the re-extraction case, no mass transfer resistance can be present in the solid phase during the whole extraction process, because the solid is extracted from an external layer deposited on the ring surface. In the initial part of the extraction process, the two tests show the same yield. In this region, equilibrium conditions are likely to occur; therefore, the results are independent of the different extraction matter. Only in the final part of the extraction process do the two extraction curves show a split. In particular, the re-extraction curve is straighter than the original extraction curve,
Ind. Eng. Chem. Res., Vol. 36, No. 2, 1997 433
solvent at the exit of the system reached the equilibrium value. In this case, the initial oil concentration in the fluid y0 can be evaluated from the experimental plot of the oil yield as a function of the mass of solvent flown, ms:
y0 )
Figure 4. Schematic view of the solute equilibrium curve: - - -, realistic shape; s, our model.
indicating a certain influence of the natural matrix in the final part of the extraction process. Model Evaluations. The first requirement of our model was to use an equilibrium relationship, which had to be able to describe some of the features shown by the extraction curves. The most important are (a) the initial part of all the extraction curves is linear and almost coincident independently from the solvent flow rate and (b) the extraction curves tend to be very close to one another in the yield vs the specific mass of solvent. Point a indicates that the fluid exiting the system is saturated with the solute for high solute concentration in the solid. A reasonable explanation for point b can be that equilibrium is still relevant in the final part of extraction when the solute concentration in the solid is low. The flat shape of the curves suggests that unfavorable equilibrium should apply at low concentration values. Our conclusion was that the shape of an appropriate equilibrium curve should be sigmoidal, such as the hyphenated curve shown in Figure 4. Indeed, this curve shows an upward convexity at low concentration values that is typical of unfavorable isotherms. Furthermore, it shows a saturation value, y0, which is asymptotically reached at high concentration values. In the middle, we have a transition region where an inflection of the curve occurs. Several functional forms can reproduce such a curve. However, in the lack of theoretical or experimental suggestions, any form would be arbitrary. To be able to estimate model parameters, we tried to simplify it as much as possible. We extended the limiting behaviors at the extremes of the concentration range to a single intermediate value of solute concentration, xj, where a sharp transition between these two different behaviors was hypothesized. The shape of the equilibrium function is shown by the continuous line in Figure 4. Its analytical description is
x < xj
f(x) ) Kx
x g xj
f(x) ) y0
(17)
Correspondingly with our requirements, we used values of K < y0/xj. Therefore, the discontinuity of f in xj is given by
f(xj) ) y0 > lim f(x) ) Kxj
(18)
xfxj-
This is a simple way to obtain an unfavorable equilibrium relationship at low concentration values. Further requirements of our model are the values of some of the initial concentrations y0 and x0. In the cases studied, the first part of the yield curve is linear. As pointed out by some authors (for example, Lee et al., 1986), this means that the solute concentration in the
me me m0 Y ) ) ms m0 ms ms/m0
(19)
where me is the mass of the extracted oil and m0 the initial mass of seed. The last term of eq 19 is the slope of the linear section of the yield curve. In the hypothesis that exhausted seeds are free of solute, the initial mass of the oil available in the solid can be supposed to be equal to the asymptotic value of the mass of oil extracted, me∞. With Y∞ the asymptotic yield, then it is
x0 )
me∞ Y∞ ) m0 - me∞ 1 - Y∞
(20)
The overall mass balance also defines the relationship between the density of the insoluble solids in the seeds Fs and the overall seed density Fc:
Fsx0 ) FcY∞
(21)
From our experimental data, we evaluated Y∞ ) 0.34. The physical properties of the treated solids are Fc ) 922 kg m-3, ap ) 2000 m-1, and ) 0.65, and the density of the fluid is Ff ) 900 kg m-3. Using these data with eqs 19-21, we evaluated Fs ) 637 kg m-3, x0 ) 0.492, and y0 ) 0.011. With these solid and fluid properties, in the 0.15 × 10-3 m3 volume reactor, the interstitial velocity corresponding to 5 kg h-1 is 2.45 × 10-3 m s-1; in the 1.5 × 10-3 m3 volume reactor, the interstitial velocities corresponding to 5, 10, 20, 25, and 45 kg h-1 are 0.499-, 0.898-, 1.80-, 2.25-, and 4.05 × 10-3 m s-1, respectively. The model developed has three further adjustable parameters. These are the mass exchange coefficient h and the two parameters of the equilibrium function K and xj. According to the hypotheses made, these parameters were independent of the experimental condition tested, and therefore, the same value of these parameters had to be used to fit all experimental results. The best-fit values were h ) 2.2 × 10-5 m s-1, K ) 0.0036, and xj ) 0.33. The model evaluations obtained for all the experimental conditions tested are the continuous curves in Figures 1 and 2. An example of the concentration profiles predicted by our model is given in Figure 5, relative to the case of extraction performed on the second unit at 25 kg h-1 solvent flow rate. In this figure, values of y and x are given as a function of the axial coordinate along the bed axis and are parametric in the extraction time. Discussion Concentration profiles in Figure 5 show a cusp corresponding to the switch in the equilibrium curve from the favorable part in the equilibrium isotherm at high solute concentration to the unfavorable part at lower solute concentration. Unfavorable equilibrium is found near the solvent inlet, where solute concentration in the fluid is smaller and particles have been extracted for a longer time. The favorable part of the isotherm can be found in the initial part of the extraction moving toward the extractor outlet, where the extraction takes place
434 Ind. Eng. Chem. Res., Vol. 36, No. 2, 1997
a
attributed to an external matrix made of a spongy structure with a very open porosity, in which the solvent velocity was affected by the external flow rate. However, the application of the appropriate modification of that model to this case was unable to give a reasonable fit of experimental data. We concluded, therefore, that only an increase of the solute bonding to the solid at lower solute concentrations can explain the observed extraction behavior. Literature data regarding sunflower seed oil SFE (Stahal, 1980), however, tend to exclude significant variations of extract composition during the process.
b Conclusions
Figure 5. Solute concentration profiles calculated for the 1.5 × 10-3 m3 volume extractor operated at 10 kg h-1 solvent flow rate: (a) fluid phase and (b) solid phase.
lately. The exit of the cusp of the concentration profile out of the extractor corresponds to the appearance of a less marked though visible cusp in all the extraction profiles in Figures 1 and 2. Such sharp transitions between different behaviors depend on the model simplicity. A smoother transition between the two equilibrium regimes is more realistic, and, in fact, smoother extraction curves are obtained in the experiments. However, apart from a small region around the cusp in the model extraction curve, the agreement between the experimental data and the model in Figures 1 and 2 is fairly good, especially for larger flow rates. Probably, these discrepancies would be smaller if a smoother equilibrium curve had been used in the model, but results show that the simple equilibrium curve adopted is able to reproduce the essential features of the extraction process. A physical explanation for the transition in the equilibrium curve between the different regimes can be found in the extraction of differently bonded fractions of the solute. Perhaps the fraction of the solute that is more readily extracted is easily available, while the fraction of solute that is extracted in the last portion of the process is more tightly bonded to the cell structure. To a certain extent, this explanation is similar to that given by Sovova´ (1994) to explain a double regime in extraction. In that case, however, the transition in the final part of extraction involved larger resistances to mass transfer internal to the solid phase. Our data, instead, suggest that mass transfer increases with flow rate, and this excludes a controlling internal resistance. There is another possibility in the modeling that is consistent with points a and b in the Model Evaluation section above. It is to consider a shrinking core model similar to that developed by King and Catchpole (1993), in which the effects of an increasing external resistance due to the presence of an external porous matrix are taken into account. A similar approach was used by Reverchon and Poletto (1996) to model extraction from flower concretes. In that case, the variation of the external resistance with the solvent flow rate was
Supercritical extraction of crushed sunflower seeds was carried out on a laboratory scale apparatus and on a supercritical extraction pilot plant at different flow rates. Extraction results appear to be mostly dependent on the mass of solvent flowed, indicating that equilibrium could be a controlling factor. A re-extraction test indicates that the solid matrix affects the extraction process, excluding the possibility of a process purely controlled by an external resistance. Our model with a constant resistance to mass transfer and an appropriate equilibrium curve is able to correctly account for the major features in the extraction process. It seems to indicate that the equilibrium curve is unfavorable for solute extraction at low solute concentrations, while solvent saturation is found for larger solute concentrations. Differences in the solute bonding to the solid at different extraction degrees can explain the observed extraction behavior. List of Symbols ap ) specific surface of the solids, m-1 h ) mass exchange coefficient, m s-1 J ) solute flux from the solid to the solvent per unit volume, kg m-3 s-1 K ) equilibrium constant for low solute concentrations m0 ) mass of the seed charge, kg me ) mass of the extracted oil, kg ms ) mass of the flowed solvent, kg t ) extraction time, s u ) interstitial velocity of the solvent, m s-1 x ) solute concentration in the solid in terms of mass of solute per unit mass of nonsoluble solid xj ) solute concentration in the solid controlling the transition in the equilibrium curve x0 ) solute concentration in the solid at the extraction start xu ) solute concentration in the untreated solid y ) solute concentration in the solvent in terms of mass of solute per unit mass of solvent Y ) extraction yield y0 ) solute concentration in the solvent at the extraction start y* ) equilibrium mass of solute per unit mass of solvent Y∞ ) maximum extraction yield z ) axial coordinate in the extractor, m zI ) coordinate along the solvent characteristic, m zII ) coordinate along the solid characteristic, m Greek Letters ) voidage of the extraction bed ξ ) solute concentration in the solid along the characteristic Fc ) density of the untreated particles, kg m-3 Ff ) solvent density, kg m-3
Ind. Eng. Chem. Res., Vol. 36, No. 2, 1997 435 Fs ) density of the nonsoluble solid, kg m-3 ψ ) solute concentration in the solvent along the characteristic
Literature Cited Bulley, N. R.; Fattori, M.; Meisen, A.; Moyls, L. Supercritical fluid extraction of vegetable seeds. J. Am. Oil Chem. Soc. 1984, 61, 1362-1365. Christianson, D. D.; Friedrich, J. P.; List, G. R.; Warner, K.; Bagley, E. B.; Stringfellow, A. C.; Inglett, G. E. SFE of dry milled corn germ with carbon dioxide. J. Food Sci. 1984, 49, 229233. Eggers, R.; Sievers, U.; Stein, W. High pressure extraction of oil seed. J. Am. Oil Chem. Soc. 1985, 62, 122-1230. Eisenbach, W. O. Extraction and fractionation of natural products. In Proc. 1st Int. Symp. Supercrit. Fluids Perrut, M., Ed. 1988, 719-725. Fattori, M.; Bulley, N. R.; Meisen, A. Carbon dioxide extraction of canola seed: oil solubility and effect of seed treatment. J. Am. Oil Chem. Soc. 1988, 65, 968-974. Favati, F.; King, J. W.; Mazzanti, M. Supercritical carbon dioxide extraction of evening primrose oil. J. Am. Oil Chem. Soc. 1991, 68, 422-427. Fors, S. M.; Ericksson, C. E. Characterization of oils extracted from Oats by supercritical carbon dioxide. Lebensmit. Wiss. U. Technol. 1990, 23, 390-395. Friedrich, J. P.; List, G. R. Characterization of soybean oil extracted by supercritical CO2 and hexane. J. Agric. Food Chem. 1982, 30, 192-193. Goto, M.; Sato, M.; Hirose, T. Extraction of peppermint oil by supercritical carbon dioxide. J. Chem. Eng. Jpn. 1993, 26, 401407. King, M. B.; Catchpole, J. R. Physico-chemical data required for the design of near critical fluid extraction process. In Extraction of natural products using near-critical solvents; King, M. B., Bott, T. R., Eds.; Chapman and Hall: New York, 1993; pp 184228. Lee, A. K. K.; Bulley, N. R.; Fattori, M.; Meisen, A. Modelling of supercritical carbon dioxide extraction of canola oilseed in fixed beds. J. Am. Oil Chem. Soc. 1986, 63, 921-925.
List, G. R.; Friedrich, J. P.; Pominski, J. Characterization and processing of cottonseed oil obtained by extraction with supercritical carbon dioxide. J. Am. Oil Chem. Soc. 1984, 61, 18471849. Molero Gomez, A.; Huber, W.; Pereyra Lopez, C.; Martinez de la Ossa, E. Extraction of grape seed oil with liquid and supercritical carbon dioxide. Proc. 3rd Int. Symp. Supercrit. Fluids Perrut, M., Brunner, G., Eds. 1994, 2, 413-418. Ramsay, M. E.; Hsu, J. T.; Novak, R. A.; Reigtler, W. J. Processing rice bran by SFE. Food Technol. 1991, 98-104. Snyder, J. M.; Friedrich, J. P.; Christianson, D. D. Effect of moisture and particle size on the extractability of oils from seeds with supercritical CO2. J. Am. Oil Chem. Soc. 1994, 61, 18511856. Sovova´, H. Rate of the vegetable oil extraction with supercritical CO2sI. Modelling of extraction curves. Chem. Eng. Sci. 1994, 3, 409-414. Sovova´, H.; Kucera, J.; Jez, J. Rate of the vegetable oil extraction with supercritical CO2sII. Extraction of grape oil. Chem. Eng. Sci. 1994a, 49, 415-420. Sovova´, H.; Komers, R.; Kucera, J.; Jez, J. Supercritical carbon dioxide extraction of caraway essential oil. Chem. Eng. Sci. 1994b, 49, 2499-2505. Stahl, E.; Schultz, E.; Mangold, H. K. Extraction of seed oils with liquid and supercritical CO2. J. Agric. Food Chem. 1980, 28, 1153-1157. Stahl, E.; Quirin, K.; Blagrove, R. Extraction of seed oils with supercritical carbon dioxide: effect on residual proteins. J. Agric. Food Chem. 1984, 32, 930-940. Taniguki, M.; Tsuji, T.; Shibata, M.; Kobayashi, T. Extraction of wheat germ with supercritical carbon dioxide. Agric. Biol. Chem. 1985, 49, 2367-2372.
Received for review June 24, 1996 Revised manuscript received October 8, 1996 Accepted October 22, 1996X IE960354S Abstract published in Advance ACS Abstracts, December 15, 1996. X
