Journal of Food Engineering 90 (2009) 463–470

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Journal of Food Engineering

journal homepage: www.elsevier .com/locate / j foodeng

Deacidification of olive oil by countercurrent supercritical carbon dioxideextraction: Experimental and thermodynamic modeling

Luis Vázquez a, Andrés M. Hurtado-Benavides b, Guillermo Reglero a, Tiziana Fornari a, Elena Ibáñez c,*,Francisco J. Señoráns a

a Sección Departamental Ciencias de la Alimentación, Facultad de Ciencias, Universidad Autónoma de Madrid, Ciudad Universitaria de Cantoblanco, 28049 Madrid, Spainb Facultad de Ingeniería Agroindustrial, Universidad de Nariño, Pasto, Colombiac Instituto de Fermentaciones Industriales, CSIC, Juan de la Cierva, 3, 28006 Madrid, Spain

a r t i c l e i n f o a b s t r a c t

Article history:Received 17 March 2008Received in revised form 8 July 2008Accepted 13 July 2008Available online 19 July 2008

Keywords:DeacidificationFatty acidsSupercritical fluid extractionGroup contribution equation of stateSimulation

0260-8774/$ - see front matter � 2008 Elsevier Ltd. Adoi:10.1016/j.jfoodeng.2008.07.012

* Corresponding author.E-mail address: elena@ifi.csic.es (E. Ibáñez).

Supercritical carbon dioxide was used as an extractive solvent to remove free fatty acids from cold-pressed olive oil. Crude oil of different acidity content (from 0.5 to 4.0 wt%) was extracted in a packedcolumn at 313 K and pressures of 180, 234 and 250 bar. The group contribution equation of state wasemployed to simulate the separation process, representing the oil as a simple pseudo-binary oleicacid + triolein mixture. Despite the simple representation of oil composition to simulate the deacidifica-tion process, a satisfactory agreement between the experimental and calculated yields and acidity of raff-inates was obtained. The thermodynamic model was employed to study a continuous countercurrentmultistage extraction process which yielded a raffinate having acidity lower than 0.7 wt%, when crudeolive oil with different FFA content was processed.

� 2008 Elsevier Ltd. All rights reserved.

1. Introduction

Olive oil is commercially obtained by cold-pressing processes.Depending on biological, meteorological, agricultural factors orprocessing conditions, the crude oil obtained contains differentamounts of free fatty acids (FFA). These substances are also suscep-tible to oxidation, leading to rancidity and conferring an undesiredflavor to the oil. Therefore, high amounts of FFA in the oil must beabsolutely avoided. Furthermore, the lower the FFA content in thevirgin olive oil, the higher is its commercial value. According to theEuropean Union regulations (the major world producer of oliveoil), cold-pressed olive oil with FFA content greater than 2.0%(lampante olive oil) is not acceptable for human consumption,and refining or deacidification is required prior to blending withvirgin olive oil. Additionally, high valued olive oil (extra virgin oliveoil) must contain less than 0.8% of FFA (European Council Regula-tion No. 1513/2001).

Therefore, the deacidification of crude olive oil is important notonly for consumer acceptance but also because it has the maxi-mum economic impact on production. The removal of FFA fromcrude oil is, therefore, a crucial step in olive oil production sinceit predominantly determines the quality of the final product. Thewell-known adverse effects of chemical or physical refining pro-

ll rights reserved.

cesses on the oil quality reduce its market value (Bondioli et al.,1992). New alternative deacidification processes proposed are re-esterification, solvent extraction, biological deacidification, mem-brane technology and supercritical fluid extraction (SFE). Each ofthese alternatives has its own advantages and drawbacks (Bosleand Subramanian, 2005).

SFE using carbon dioxide is a low temperature and a relativelypollution free operation. Its high selectivity permits the removalof FFA from the oil with minimum loss of neutral oil: triglyceridesand unsaponifiable matter (tocopherols, sterols and vitamins).Thus, when this technique is applied, the deacidification processcan be carried out without significant loss in yield or the nutri-tional properties (Brunetti et al., 1989).

Brunetti et al. (1989) have investigated the extraction of fattyacids from fatty acid + triglyceride mixtures using supercriticalcarbon dioxide (SC-CO2). Experiments were carried out on sampleswith different FFA content (from 2.6 up to 20 wt%) at pressures of20 and 30 MPa and temperatures of 313 and 333 K. Besides thelimitations concerned with the use of batch equipment, they con-cluded that the SFE was particularly suitable for deacidification ofolive oils with FFA content lower than 10%, since the selectivityfactor for fatty acid extraction increases as the concentration ofFFA in the crude oil decreases. On the other hand, Simo�es andBrunner (1996) evaluated the possibility of using SC-CO2 to deacid-ify olive oil using a commercial oil containing squalene (around0.7 wt%), FFA (from 3 to 15 wt%) and triglycerides. Experimental

464 L. Vázquez et al. / Journal of Food Engineering 90 (2009) 463–470

phase equilibria measurements covering a wide range of extractionpressures and temperatures (from 313 to 353 K and up to 30 MPa)were performed and these data were used to simulate a counter-current packed column for the deacidification of olive oil by SFE.Simulated results showed an increase of extraction yield withCO2 density. Additionally, higher solvent-to-feed flow ratios im-proved the extraction yield, resulting in lower raffinate acidities.

The continuous CO2-SFE of olive oil from different geographicorigins (Italy, Spain and Tunisia) in a countercurrent packed col-umn of 3 m high was studied by Bondioli et al. (1992). The influ-ence of extraction temperature and pressure, CO2/oil flux ratio,oil injection point and use of temperature gradient was discussed.Of particular interest were the results obtained by varying theinjection point, which demonstrated the efficacy of a rectifyingsection (obtained by temperature gradient) to reduce the loss ofneutral oil.

In the present work, the deacidification of olive oil using CO2-SFE was experimentally studied and thermodynamically simu-lated, investigating pressures higher than those employed byBondioli et al. (1992). Extractions were performed in a packed col-umn 3 m high, without external or internal reflux and utilizing1.8 m of the column as the stripping section; the upper part ofthe column was used to avoid carryover of the crude oil. All exper-imental assays were performed at 313 K, with a CO2/oil flux ratio of20 and pressures of 180, 234 and 250 bar. The acidity of the oliveoil samples varied from 0.5 to 4.0 wt%, which is the normal rangeobserved in the olive oil obtained by conventional cold-pressing.The experimental results were simulated using the group contribu-tion equation of state (GC-EoS) (Skjold-Jørgensen, 1984) in a com-pletely predictive manner. The model was previously tested withthe experimental phase equilibrium data reported by Simo�es andBrunner (1996) and the countercurrent extraction data reportedby Bondioli et al. (1992). A reasonable agreement between modelpredictions and experimental phase equilibria compositions,experimental yield and refined oil acidity was achieved. The mainaim of the present work was to develop a simple predictive toolwhich systematically evaluated the extraction process conditionsthat guaranteed a target low acidity in the raffinate, when crude ol-ive oils having different acidity were deacidified.

2. Experimental

2.1. Raw materials and reagents

The olive oil samples containing different concentrations of FFA,employed in the SFE experiments, were donated by a local com-pany (Migasa, Sevilla, Spain). Carbon dioxide N38 (99.98%), waspurchased from AL Air Liquide España S.A. (Madrid, Spain). All sol-vents used were of HPLC grade and were obtained from Lab-Scan(Dublin, Ireland).

2.2. Equipment and extraction process

A schematic diagram of the SFE system employed in this studyis shown in Fig. 1. The heart of the pilot plant is a 316 stainlesssteel extraction column (18 mm i.d.) packed with Fenske rings(3 � 0.5 mm; Afora, S.A. Spain). The total height of the extractioncolumn is 300 cm with provision for sample introduction at threelevels: top, middle and bottom, as shown in Fig. 1. In this work,the olive oil samples are introduced in the middle, giving an effec-tive packed height of 180 cm (measured from the point of intro-duction to the CO2 feed point).

Ten thermocouples with temperature controllers are installedalong the length of the extraction column, located at heights of60, 150, 210, 270 and 300 cm measured from bottom.

The SFE system also comprises two separator cells where a cas-cade decompression takes place and a cryogenic trap at atmo-spheric pressure. Both CO2 and liquid feed sample are preheatedat the discharge of their respective pumps (Dosapro Milton Roy)before introduction into the SFE device. The temperature of theextraction column and separator units is maintained by using a sil-icone heater bath. The plant has computerized PLC-based instru-mentation and a control system, with several safety devicesincluding valves and alarms.

A continuous flow of CO2 (2.7 kg/h) was introduced into the col-umn at the bottom. When the set operating pressure and temper-ature were reached, 135 g/h of oil was pumped continuously overthe duration of extraction (90 min). A CO2 to oil flux ratio of 20 wasemployed in all the experiments. When the extraction was com-pleted, the oil flow was stopped and CO2 was pumped for another30 min.

Olive oil having different degree of acidity (0.5, 1.0, 2.5 and4.0 wt%) was employed as the feed. The extraction pressures inves-tigated for each sample were: 180, 234 and 250 bar. The columnand separator units were maintained at 313 K in all experimentalextractions. The pressure of the first separator unit was about 1.8times lower than the column pressure, and the second separatorwas maintained at a low pressure (20–30 bar) during all theexperiments.

The bottom product (raffinate) and liquid fractions in the sepa-rators (extract) were collected, weighed and analyzed after theextraction was completed. The closure of material balance wasascertained in all experiments within an accuracy of 10%.

2.3. Analytical methods

2.3.1. HPLC analysisThe composition of the neutral lipids was evaluated on a Kroma-

sil silica 60 column (250 mm by 4.6 mm, Análisis Vinicos, Tomelloso,Spain) coupled to a CTO 10A VP 2 oven, a LC-10AD VP pump, a gradi-ent module FCV-10AL VP, a DGU-14A degasser, and a evaporativelight scattering detector ELSD-LT from Shimadzu (IZASA, Spain).The ELSD conditions were 2.2 bars, 35 �C, and gain three. The flowrate was 2 mL/min. A splitter valve was used after the column andonly 50% of the mobile phase was directed through the detector.The column temperature was maintained at 35 �C. The mobile phaseutilized has previously been reported by Torres et al. (2005).

2.3.2. Oleic acid analysis by titration with KOHThe acid value of raffinates (deacidified oil) was determined

according to the volumetric method described in the Official Ana-lytical Methods for Edible Fats and Oils, published by the Ministryof Agriculture of Spain in 1974. The samples were dissolved in amixture of ethanol:ethyl ether 1:1. The acid value was calculatedby titration of the raffinates with an ethanolic solution of potas-sium hydroxide (0.1 M). This solution was calibrated with dipotas-sium phthalate and phenolphthalein as indicator, in order todetermine its normality accurately. Taking into consideration thatoleic acid is the main fatty acid in olive oil, the acid value was ex-pressed as percentage of oleic acid:

Acid value ð% oleic acidÞ ¼ VMN=10P

where V and N are, respectively, the volume (mL) and normality ofthe ethanolic solution of KOH, M is the molecular weight of oleicacid (282.45) and P is the weight of the sample (g).

3. Thermodynamic modeling

Phase equilibrium behavior of olive oil + SC-CO2 has been stud-ied using GC-EoS model (Chen et al., 2000). This model has two

Fig. 1. Pilot plant SFE device.

L. Vázquez et al. / Journal of Food Engineering 90 (2009) 463–470 465

contributions to the residual Helmholtz energy of the system: arepulsive free volume term and a contribution which accountsfor attractive group interactions. The GC-EoS equations can befound in the Appendix. For a detailed description of the modelthe reader is referred to Skjold-Jørgensen (1984, 1988).

The attractive term of the GC-EoS model is a group contributionversion of the NRTL model, and has five pure group parameters (T*,q, g*, g0 and g00) and four binary interaction parameters (k�ij, k0ij, aij

and aji). The repulsive term is modeled assuming hard spherebehavior for the molecules; each species is characterized by itscritical hard sphere diameter dci (see Appendix).

Phase equilibrium calculation was performed by representingthe multicomponent olive oil feed material as a mixture compris-ing oleic acid, representing the FFA fraction, and triolein represent-ing the (more abundant) neutral oil fraction. The required purecomponent parameters (critical temperature, critical pressureand critical hard sphere molecular diameter) for oleic acid and tri-olein are given in Table 1. All pure group and binary interactionparameters used in this work to represent the vapor–liquid equi-librium behavior of the triolein + CO2 and oleic acid + CO2 binarysystems are summarized in Table 2.

The capability of the GC-EoS model to represent high-pressurephase equilibria of triglyceride + CO2 binary mixture has been dis-cussed elsewhere (Espinosa et al., 2002; Florusse et al., 2004). Fig. 2shows a comparison between experimental and calculated phaseequilibrium compositions of the triolein + CO2 mixture at two dif-ferent temperatures, 313 and 333 K, and pressures up to 50 MPa.

Table 1Pure component parameters

Tc (K) pc (bar) dc (cm/mol)

Oleic acida 796.3 12.4 7.457Trioleinb 1043.3 4.57 11.839

a Joback and Reid (1983).b Espinosa et al. (2002).

Despite the discrepancies between the different sources of experi-mental data (Bharath et al., 1992; Chen et al., 2000; Weber et al.,1999; Nilsson et al., 1991), the GC-EoS model provides a reasonableprediction of the triolein + CO2 phase equilibrium behavior.

In order to describe the oleic acid + CO2 binary mixture theCOOH–CO2 group interaction parameters were optimized in thiswork using vapor–liquid equilibria data for oleic acid + CO2 mixture(Bharath et al., 1992; Zou et al., 1990; Yu et al., 1992). The parametersobtained in the regression procedure are given in Table 2. The abso-lute average relative deviation between experimental and calcu-lated CO2 mole fractions AARD¼ 1=Nexp

� �PZcal

CO2�Zexp

CO2

� �Zexp

CO2

. ������� �in the vapor and liquid phases were, respectively, 0.07% and 3.01%.The phase equilibria representation of the oleic acid + CO2 mixtureobtained with the GC-EoS model and parameters reported in Table2 is given in Fig. 3.

Phase equilibria prediction of the oleic acid + triolein + CO2 ter-nary mixture was compared with the experimental data reportedby Simo�es and Brunner (1996) for mixtures of commercial oliveoil + CO2 in the temperature range from 313 to 353 K and pressuresup to 300 bar. The experimental liquid and vapor mass fractions re-ported by Simo�es and Brunner (1996) correspond to the quaternarysystem FFA + triglycerides + squalene + CO2. Since the squalenecontent in commercial olive oil was very low (0.7 wt%), the oilwas simulated as the binary oleic acid + triolein. The CO2/oil massratio was equal to 1. Fig. 4 shows a comparison between the exper-imental and predicted oleic acid mass fraction (CO2-free) in the li-quid (deacidified) oil phases. AARD between experimental andcalculated FFA content in the liquid phase were 8.5%, 10.2%,15.4% and 30.1% for olive oils with, respectively, 2.9, 5.2, 7.6 and15.3 wt% FFA. This means that model prediction worsens as theacidy of the olive oil increases, as can be observed in Fig. 4.

4. Results and discussion

The extraction temperature (313 K) was selected in agreementwith previous studies reported in the literature (Brunetti et al.,

Table 2GC-EoS pure group and interaction parameters used in this work

Pure group parameters Reference temperature T* Group surface area q Pure group energy parameters Reference

g g0 g00

CH3 600 0.848 316,910 �0.9274 0.0 Skjold-Jorgensen (1988)CH2 600 0.540 356,080 �0.8755 0.0 Skjold-Jorgensen (1988)CH@CH 600 0.867 403,590 �0.7631 0.0 Pusch and Schmelzer (1993)COOH 600 1.224 1211745.4 �0.1105 0.0 Ferreira et al. (2003)(CH3COO)2CH2COO triglyceride group (TG) 600 3.948 346,350 �1.3469 0.0 Espinosa et al. (2002)CO2 304.2 1.261 531,890 �0.5780 0.0 Skjold-Jorgensen (1988)

Binary group interaction parametersi j Attractive energy parameters Non-randomness

parameters

kij k0ij aij aij

CO2 CH3 0.0898 0.0 4.683 4.683 Espinosa et al. (2002)CH2 0.874 0.0 4.683 4.683 Espinosa et al. (2002)CH@CH 0.948 0.0 0.0 0.0 Ferreira et al. (2003)COOH 0.789 0.0 1.9351 0.2403 This workTG 1.094 0.112 �1.651 �1.651 Espinosa et al. (2002)

TG CH3/CH2 0.860 0.0 0.0 0.0 Espinosa et al. (2002)CH@CH 0.883 0.0 0.0 0.0 Espinosa et al. (2002)COOH 1.062 0.0 0.0 0.0 Ferreira et al. (2003)

COOH CH3/CH2 0.932 0.0 �2.946 �2.424 Ferreira et al. (2003)CH@CHa 0.932 0.0 �2.946 �2.424 Ferreira et al. (2003)

a Assumed to be equal to the CH3/CH2–COOH interaction parameters.

0

5

10

15

20

25

30

0.0 0.1 0.2 0.3 0.98 0.99 1.00 0.0 0.1 0.2 0.3 0.98 0.99 1.00

CO2 mass fraction CO2 mass fraction

P (M

Pa)

P (M

Pa)

0

10

20

30

40

50

a b

Fig. 2. Phase equilibria for the binary triolein + CO2 mixture at (a) 313 K and (b) 333 K. Comparison between different sources of experimental data: (d) Bharath et al. (1992);(j) Chen et al. (2000); (N) Weber et al. (1999) and (H) Nilsson et al. (1991). Solid lines: GC-EoS calculations.

466 L. Vázquez et al. / Journal of Food Engineering 90 (2009) 463–470

1989; Simo�es and Brunner, 1996; Bondioli et al., 1992). In the tem-perature range normally used in SFE oil processing (308–343 K),the solubility of unsaturated fatty acids in supercritical CO2 is 3–6 times greater than the solubility of triolein (Brunetti et al.,1989). Low extraction temperature (high CO2 density) ensures highfatty acid solubility in the extractive solvent. Higher temperatureshave also been tested (353 K) in the literature; in this case, a goodselectivity of FFA extraction towards squalene was found (Bondioliet al., 1992), while similar conclusions were drawn when compar-ing selectivity of FFA vs triglycerides.

Bondioli et al. (1992) studied the SFE deacidification of olive oilin the operational range of 90–150 bar, 313–333 K and CO2/oil fluxratios from 20 to 170. They could reduce oil acidity from 6.3 wt% tovalues less than 1 wt% at 313 K, 130 bar and CO2/oil flux ratiosgreater than 100. In the present work, the CO2/oil flux ratio was

set at 20, a values limited by the capacity of the CO2 pumping sys-tem. Thus, pressures greater than those employed by Bondioli et al.were explored (180, 234 and 250 bar). The higher pressures butlower CO2/oil flux ratio gave high raffinate yields, as those obtainedat low extraction pressures but high CO2/oil flux ratios (Bondioli etal., 1992). The crude olive oil samples contain, respectively, 0.5, 1.0,2.5, 3.0 and 4.0 wt% of FFA. Samples of acidity lower than 1 wt%were studied to check the reliability of the GC-EoS model at thelower end of acidity range.

Table 3 reports the process yield (defined as: mass of raffinate/mass of crude olive oil) and the raffinate acidity (wt% of FFA) ob-tained in this work. Also given in the table is the acidity reductionfactor, defined as the ratio of the crude oil acidity to raffinate acid-ity. As evident in Table 3, the acidity reduction factor is higher than2 at pressures greater than 234 bar for all oil samples having FFA

0.30 0. 0. 0.97 0.98 0.99 1.000

5

10

15

20

25

30

35

0.00 0.10 0.20 40 50

pres

sure

(MPa

)

CO2 mass fraction

Fig. 3. Phase equilibria for the binary oleic acid + CO2 mixture. Experimental data:(Bharath et al., 1992) at (�) 313 K, (N) 333 K and (j) 353 K; (Zou et al., 1990) at (�)313 K and (4) 333 K; (Yu et al., 1992) at (�) 313 K and (s) 333 K. Solid lines: GC-EoScalculations.

L. Vázquez et al. / Journal of Food Engineering 90 (2009) 463–470 467

content greater than 1 wt%. Therefore, at 313 K, 234 bar and CO2/oil flux ratio = 20, crude olive oil with a FFA content up to2.5 wt% can be deacidified by CO2-SFE to obtain a product withan acidity lower than 1 wt%. If the crude olive oil contains higheramounts of FFA, it is expected that higher pressures or higherCO2/oil flux ratios are required. For example, as mentioned before,Bondioli et al. (1992) were able to reduce the acidity from 6.3 to1 wt% at lower pressures but employing CO2/oil flux ratios greaterthan 100.

Although the experimental assays reported in Table 3 demon-strate that deacidification of olive oil using low CO2/oil flux ratiosis possible, the extraction conditions employed imply a CO2 densityin the range of 820–880 kg/m3, while olive oil density is around900 kg/m3. This means that low differences between the liquidand supercritical phase would be expected and thus, the extrac-tions reported in Table 3 should be considered as the results of asemi-continuous process. As previously mentioned, the upper partof the column (120 cm above the feed point) was employed to min-imize the loss of neutral oil by carryover.

0

2

4

6

8

10

12

14

16

18

12 16 20 24 28 32

pressure / MPa

%w

t FFA

(CO

2 fre

e)

in li

quid

pha

se

a

Fig. 4. Oil acidity + CO2 liquid–vapor equilibria. Experimental data from Simo�es and Bruand (d) 15.3 wt% FFA in crude oil. (b) 323 K and (h) 2.9 wt% and (4) 7.6 wt% FFA in cru

Besides the presumed semi-batch character of the experimen-tal runs, a continuous countercurrent column with N theoreticalstages was considered to simulate the process. The extractionprocess was simulated as a stripping operation, i.e., with thefeed at the top of the column and no enriching section. TheGC-EoS model was used to perform phase equilibria calculations.The number of theoretical stages that provided the best agree-ment between the experimental and the simulated data wasN = 2. Crude olive oil was considered to be a binary oleicacid + triolein mixture. The SFE simulation provided a satisfac-tory prediction of the raffinate yield and acidity variation as afunction of pressure and crude oil FFA content. The AARD be-tween experimental and calculated yield was 1.6%, the AARDwas 11.2% with respect to raffinate acidity for olive oil sampleshaving 1–4 wt% of FFA. In the case of very dilute samples(0.5 wt% of FFA) deviations were significantly higher (ca. 35%).Additionally, for a given crude oil acidity, deviations betweenexperimental and calculated yield and acidity were, in general,higher at lower pressures. Such deviations may be attributedto the simplified olive oil composition assumed (i.e., the binarytriolein + oleic acid mixture).

The model predictions were also tested against the experimen-tal data reported by Bondioli et al. (1992) for the SFE deacidifica-tion of olive oil with FFA content of 6.3 wt%. Again, a multistagecontinuous countercurrent column was assumed for the purposeof process simulation. In this case, the best agreement betweenthe experimental data reported by Bondioli et al. and the simulatedcountercurrent column was achieved with N = 3. The variation ofraffinate acidity in the pressure range examined by Bondioli et al.(80–150 bar) and explored in this work (180–250 bar) are repre-sented in Fig. 5 and compared with the GC-EoS model predictions.Additionally, Fig. 6 compares model predictions with the variationin yield and acidity (Bondioli et al., 1992) as a function of CO2/oilflux ratio. As evident in Figs. 5 and 6, SFE deacidification is well-de-scribed by the GC-EoS model. However, Bondioli et al. have exper-imentally determined a minimum raffinate acidity of 0.7 wt% at313 K, 130 bar and CO2/oil flux ratio = 100, which is not predictedby the model.

Despite the simple binary composition assumed (oleicacid + triolein) the GC-EoS model can be used as a screening toolto analyze olive oil SFE deacidification over a wide range of tem-peratures (313–343 K) and pressures (130–300 bar). Taking intoaccount the discrepancies between experimental and calculatedphase equilibria compositions, the model is accurate for crude ol-ive oils with FFA content up to ca. 10 wt%.

0

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12

14

16

18

16 20 24 28 32

pressure / MPa

b

nner (1996). Solid lines: GC-EoS predictions. (a) 313 K and (j) 2.9 wt%, (N) 7.6 wt%de oil; (s) 353 K and 15.3 6 wt% FFA in crude oil.

Table 3Experimental and calculated yield (mass of raffinate/mass of crude olive oil) and acidity (wt% of FFA) of the raffinate obtained in the SFE of crude olive oil at 313 K and CO2/oil fluxratio = 20

Crude oil acidity Extraction pressure (bar) Raffinate yield (%) Raffinate acidity Experimental acidity reduction factor

Experimental Calculated Experimental Calculated

0.5 250 87.3 87.4 0.27 0.17 1.820.5 234 87.3 87.2 0.31 0.19 1.590.5 180 95.2 90.1 0.45 0.29 1.11

1.0 250 86.7 87.1 0.40 0.35 2.501.0 234 87.9 87.2 0.42 0.39 2.381.0 180 93.0 90.4 0.50 0.58 2.00

2.5 250 85.8 86.2 0.93 0.89 2.682.5 234 87.3 86.4 0.88 0.99 2.842.5 180 92.6 89.5 1.21 1.50 2.07

3.0 250 86.1 86.0 0.99 1.07 3.033.0 234 86.4 86.2 1.29 1.21 2.333.0 180 92.3 89.8 1.49 1.81 2.01

4.0 250 84.1 85.5 1.43 1.46 2.804.0 234 86.1 85.8 1.85 1.64 2.164.0 180 92.9 89.5 2.31 2.45 1.73

0

1

2

3

4

5

6

7

70 100 130 160 190 220 250

pressure / MPa

raffi

nate

aci

dity

(%w

t FFA

)

Fig. 5. Raffinate acidity (wt% FFA) obtained in SFE deacidification. Experimentaldata from this work at 313 K, CO2/oil flux = 20 and crude oil acidity of (N) 4 wt%, (d)2.5 wt% and (j) 0.5 wt%; (s) experimental data from Bondioli et al. (1992) at 313 K,CO2/oil flux = 100 and 6.3 wt% crude oil acidity. Solid lines: GC-EoS calculations.

0

2

4

6

8

10

0 50 100 150 200

CO2/oil flux ratio

raffi

nate

aci

dity

(%w

t FFA

)

50

60

70

80

90

100

raffi

nate

yie

ld (%

)

Fig. 6. Raffinate yield (d) and acidity (s) obtained in olive oil SFE deacidification at313 K, 130 bar and different CO2/oil flux ratios (1992). Solid lines: GC-EoScalculations.

468 L. Vázquez et al. / Journal of Food Engineering 90 (2009) 463–470

The GC-EoS was used to obtain the process conditions requiredto deacidify crude olive oil, with different FFA content, down to0.7 wt% FFA. Considering the significant improvement in functionaland organoleptic properties of crude olive oil SFE deacidification,this oil can potentially be blended with virgin olive oil to producecommercial olive oil of high quality.

The SFE process was mathematically solved within a sequentialprocess simulator that includes rigorous models for a high-pres-sure multistage extractor (Brignole et al., 1987). In these routines(FORTRAN language) thermodynamic phase equilibria calculationswere performed using the GC-EOS model. The number of stagesemployed to carry out the calculations was 10; additional stagesdid not produce any significant changes in raffinate and extractcompositions. Furthermore, considering that the height equivalenttheoretical stage (HETP) is around 0.5–1.0 m (Hurtado-Benavideset al., 2004; Simo�es and Brunner, 1996), a reasonably value canbe deduced for the total height of the packed column.

Fig. 7 shows the required variation of CO2/oil ratio (Fig. 7a) andprocess yield (Fig. 7b) as a function of the crude olive oil acidity, fordifferent extraction conditions (temperature and pressure). Extrac-

tion temperature and pressure were selected to maintain a CO2

density of 740 Kg/m3, so as to guarantee a reasonable differencebetween the down-flowing liquid and the up-flowing supercriticalsolvent. As observed in Fig. 7a, the CO2/oil ratio can be reducedtwofold when extraction pressure is increased from 130 bar(extraction temperature 313 K) to 250 bar (extraction temperature343 K). Additionally, raffinate yields are mainly determined by thequality of the crude oil processed (see Fig. 7b), with values decreas-ing from ca. 90% to 77% when crude oil acidity increases from 2 to10 wt%. Besides the higher capital costs that a more intense set ofextraction conditions would demand, it has to be noted that thepower required for CO2 recirculation at 343 K and 250 bar is signif-icantly lower (see Table 4) than that required at 313 K and 130 bar,because of the lower CO2/oil flux ratio required.

Therefore, in the present work deacidification of crude olive oilwith 0.5–4 wt% FFA content was experimentally carried out andcompared with SFE data reported in the literature. Also, a simplemodel to simulate the deacidification process was developed. Themodel was tested using experimental phase equilibria data for

0

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40

60

80

100

120

140

160

0 4 8 10 12crude oil acidity (%wt FFA)

CO

2/oil

flux

ratio

70

75

80

85

90

95

0 8 10 12crude oil acidity (%wt FFA)

raffi

nate

yie

ld (%

)

2 6 2 4 6

a b

Fig. 7. SFE deacidification of crude olive oil (CO2 density = 740 kg/m3) to attain a raffinate with 0.7 wt% of FFA, as predicted by the GC-EoS model. (j) 130 bar and 313 K; (N)170 bar and 323 K; (d) 210 bar and 333 K; (�) 250 bar and 343 K.

Table 4Comparison of the countercurrent SFE deacidification of crude olive oil from 6 to0.7 wt% FFA at different extraction conditions (CO2 density = 740 kg/m3) as predictedby the GC-EoS model

Temperature(K)

Pressure(bar)

Raffinateyield (%)

CO2 fluxa (kgCO2/kg oil)

CO2 recirculationpowerb (kJ/kg oil)

313 130 83.2 123.2 21,691323 170 82.2 94.6 15,188333 210 81.4 76.6 11,122343 250 80.8 63.8 8292

a Required to achieve a raffinate oil with 0.7 wt% of FFA.b Considering a single separator tank at 313 K and 50 bar.

L. Vázquez et al. / Journal of Food Engineering 90 (2009) 463–470 469

olive oil (with different acidity) + CO2 mixtures, and was employedto study the required variation of process conditions to ensure a fi-nal acidity value lower than 0.7% in the raffinate.

Acknowledgements

This work has been financed under a R&D program of theSpanish Ministry of Education and Science, Project AGL-2004-07227-C02-01 and Project S-0505/AGR/000153 from the Comuni-dad Autónoma de Madrid. T.F. would like to acknowledge thefinancial support of the Ramon y Cajal Program from the Ministryof Education and Science. The authors acknowledge the financialsupport of Migasa (Spain).

Appendix

The residual Helmholtz energy in the GC-EoS model is de-scribed by two terms:

A ¼ Afv þ Aatt ðA:1Þ

The free volume contribution (Afv) is modeled assuming hardsphere behavior for the molecules, characterizing each substancei by a hard sphere diameter di. A Carnahan–Starling type of hardsphere expression for mixtures is adopted:

Afv=RT ¼ 3 k1k2=k3ð Þ Y � 1ð Þ þ k3

2=k23

� ��Y þ Y2 � ln Y� �

þ n ln Y

ðA:2Þ

where

Y ¼ 1� pk3

6V

� ��1

and kk ¼XNC

i

nidki ðA:3Þ

NC is the number of components, ni is the number of moles ofcomponent i and V is the total volume. A temperature dependentgeneralized expression is assumed for di:

di ¼ 1:065655dcif1� 0:12 exp½�2Tci=ð3TÞ�g ðA:4Þ

dci is the pure component critical hard sphere diameter, given by:

dci ¼ 0:08943RTci Pci=ð Þ1=3 ðA:5Þ

when the compound match with a group (e.g., H2O, CO2, H2, etc.).For the remaining cases dci is fitted to a point of the pure componentvapor pressure curve, usually the normal boiling point. Since vaporpressure data for low volatile or thermolabile substances is oftennot available or not reliable, infinite dilution activity coefficientscan be used to estimate the dci parameter of high molecular weightcompounds, such as alkanes and triglycerides, as demonstrated byEspinosa et al. (2002).

For the evaluation of the attractive contribution to theHelmholtz energy, a group contribution version of a density-dependent NRTL-type expression is derived:

Aatt=RT ¼ � z

2

XNC

i

ni

XNG

j

mijqj

XNG

k

ðhkgkj~qqÞXNG

i

hlslj

,ðA:6Þ

where

hj ¼ qj=q� �XNC

i

nimij q ¼

XNC

i

ni

XNG

j

mijqj

sij ¼ exp aijDgij~q=ðRTVÞ

Dgij ¼ gij � gjj

NG is the number of groups, z is the number of nearest neigh-bors to any segment (set to 10), mi

j is the number of groups type jin molecule i, qj is the number of surface segments assigned togroup j, hk is the surface fraction of group k, ~q is the total numberof surface segments, gij is the attraction energy parameter for inter-actions between groups i and j, and aij is the NRTL non-randomnessparameter (aij – aji). The interactions between unlike groups arecalculated from

470 L. Vázquez et al. / Journal of Food Engineering 90 (2009) 463–470

gij ¼ kijðgiigjjÞ1=2 ðkij ¼ kjiÞ

with the following temperature dependencies for the interactionparameters:

gjj ¼ g�jj 1þ g0jj T=T�j � 1� �

þ g00jj ln T=T�j� �� �

kij ¼ k�ij 1þ kij ln 2T T�i þ T�j� �. ih on

being g�jj the interaction parameter at the reference temperature T�i .

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