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Deacidification of olive oil by countercurrent supercritical carbon dioxide extraction: Experimental and thermodynamic modeling
Journal of Food Engineering 90 (2009) 463–470
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Journal of Food Engineering journal homepage: www.elsevier.com/locate/jfoodeng
Deacidification of olive oil by countercurrent supercritical carbon dioxide extraction: 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, Spain Facultad de Ingeniería Agroindustrial, Universidad de Nariño, Pasto, Colombia c Instituto de Fermentaciones Industriales, CSIC, Juan de la Cierva, 3, 28006 Madrid, Spain b
a r t i c l e
i n f o
Article history: Received 17 March 2008 Received in revised form 8 July 2008 Accepted 13 July 2008 Available online 19 July 2008 Keywords: Deacidification Fatty acids Supercritical fluid extraction Group contribution equation of state Simulation
a b s t r a c t Supercritical carbon dioxide was used as an extractive solvent to remove free fatty acids from coldpressed olive oil. Crude oil of different acidity content (from 0.5 to 4.0 wt%) was extracted in a packed column at 313 K and pressures of 180, 234 and 250 bar. The group contribution equation of state was employed to simulate the separation process, representing the oil as a simple pseudo-binary oleic acid + triolein mixture. Despite the simple representation of oil composition to simulate the deacidification process, a satisfactory agreement between the experimental and calculated yields and acidity of raffinates was obtained. The thermodynamic model was employed to study a continuous countercurrent multistage extraction process which yielded a raffinate having acidity lower than 0.7 wt%, when crude olive oil with different FFA content was processed. Ó 2008 Elsevier Ltd. All rights reserved.
1. Introduction Olive oil is commercially obtained by cold-pressing processes. Depending on biological, meteorological, agricultural factors or processing conditions, the crude oil obtained contains different amounts of free fatty acids (FFA). These substances are also susceptible to oxidation, leading to rancidity and conferring an undesired flavor to the oil. Therefore, high amounts of FFA in the oil must be absolutely avoided. Furthermore, the lower the FFA content in the virgin olive oil, the higher is its commercial value. According to the European Union regulations (the major world producer of olive oil), 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 with virgin olive oil. Additionally, high valued olive oil (extra virgin olive oil) must contain less than 0.8% of FFA (European Council Regulation No. 1513/2001). Therefore, the deacidification of crude olive oil is important not only for consumer acceptance but also because it has the maximum economic impact on production. The removal of FFA from crude oil is, therefore, a crucial step in olive oil production since it predominantly determines the quality of the final product. The well-known adverse effects of chemical or physical refining pro* Corresponding author. E-mail address: elena@ifi.csic.es (E. Ibáñez). 0260-8774/$ - see front matter Ó 2008 Elsevier Ltd. All rights reserved. doi:10.1016/j.jfoodeng.2008.07.012
cesses on the oil quality reduce its market value (Bondioli et al., 1992). New alternative deacidification processes proposed are reesterification, solvent extraction, biological deacidification, membrane technology and supercritical fluid extraction (SFE). Each of these alternatives has its own advantages and drawbacks (Bosle and Subramanian, 2005). SFE using carbon dioxide is a low temperature and a relatively pollution free operation. Its high selectivity permits the removal of FFA from the oil with minimum loss of neutral oil: triglycerides and unsaponifiable matter (tocopherols, sterols and vitamins). Thus, when this technique is applied, the deacidification process can be carried out without significant loss in yield or the nutritional properties (Brunetti et al., 1989). Brunetti et al. (1989) have investigated the extraction of fatty acids from fatty acid + triglyceride mixtures using supercritical carbon dioxide (SC-CO2). Experiments were carried out on samples with different FFA content (from 2.6 up to 20 wt%) at pressures of 20 and 30 MPa and temperatures of 313 and 333 K. Besides the limitations concerned with the use of batch equipment, they concluded that the SFE was particularly suitable for deacidification of olive oils with FFA content lower than 10%, since the selectivity factor for fatty acid extraction increases as the concentration of FFA in the crude oil decreases. On the other hand, Simoes and Brunner (1996) evaluated the possibility of using SC-CO2 to deacidify olive oil using a commercial oil containing squalene (around 0.7 wt%), FFA (from 3 to 15 wt%) and triglycerides. Experimental
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phase equilibria measurements covering a wide range of extraction pressures and temperatures (from 313 to 353 K and up to 30 MPa) were performed and these data were used to simulate a countercurrent packed column for the deacidification of olive oil by SFE. Simulated results showed an increase of extraction yield with CO2 density. Additionally, higher solvent-to-feed flow ratios improved the extraction yield, resulting in lower raffinate acidities. The continuous CO2-SFE of olive oil from different geographic origins (Italy, Spain and Tunisia) in a countercurrent packed column of 3 m high was studied by Bondioli et al. (1992). The influence 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 the injection point, which demonstrated the efficacy of a rectifying section (obtained by temperature gradient) to reduce the loss of neutral oil. In the present work, the deacidification of olive oil using CO2SFE was experimentally studied and thermodynamically simulated, investigating pressures higher than those employed by Bondioli et al. (1992). Extractions were performed in a packed column 3 m high, without external or internal reflux and utilizing 1.8 m of the column as the stripping section; the upper part of the column was used to avoid carryover of the crude oil. All experimental assays were performed at 313 K, with a CO2/oil flux ratio of 20 and pressures of 180, 234 and 250 bar. The acidity of the olive oil samples varied from 0.5 to 4.0 wt%, which is the normal range observed in the olive oil obtained by conventional cold-pressing. The experimental results were simulated using the group contribution equation of state (GC-EoS) (Skjold-Jørgensen, 1984) in a completely predictive manner. The model was previously tested with the experimental phase equilibrium data reported by Simoes and Brunner (1996) and the countercurrent extraction data reported by Bondioli et al. (1992). A reasonable agreement between model predictions and experimental phase equilibria compositions, experimental yield and refined oil acidity was achieved. The main aim of the present work was to develop a simple predictive tool which systematically evaluated the extraction process conditions that guaranteed a target low acidity in the raffinate, when crude olive 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 company (Migasa, Sevilla, Spain). Carbon dioxide N38 (99.98%), was purchased from AL Air Liquide España S.A. (Madrid, Spain). All solvents 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 study is shown in Fig. 1. The heart of the pilot plant is a 316 stainless steel extraction column (18 mm i.d.) packed with Fenske rings (3 0.5 mm; Afora, S.A. Spain). The total height of the extraction column is 300 cm with provision for sample introduction at three levels: top, middle and bottom, as shown in Fig. 1. In this work, the olive oil samples are introduced in the middle, giving an effective packed height of 180 cm (measured from the point of introduction to the CO2 feed point). Ten thermocouples with temperature controllers are installed along the length of the extraction column, located at heights of 60, 150, 210, 270 and 300 cm measured from bottom.
The SFE system also comprises two separator cells where a cascade decompression takes place and a cryogenic trap at atmospheric pressure. Both CO2 and liquid feed sample are preheated at the discharge of their respective pumps (Dosapro Milton Roy) before introduction into the SFE device. The temperature of the extraction column and separator units is maintained by using a silicone heater bath. The plant has computerized PLC-based instrumentation and a control system, with several safety devices including valves and alarms. A continuous flow of CO2 (2.7 kg/h) was introduced into the column at the bottom. When the set operating pressure and temperature were reached, 135 g/h of oil was pumped continuously over the duration of extraction (90 min). A CO2 to oil flux ratio of 20 was employed in all the experiments. When the extraction was completed, the oil flow was stopped and CO2 was pumped for another 30 min. Olive oil having different degree of acidity (0.5, 1.0, 2.5 and 4.0 wt%) was employed as the feed. The extraction pressures investigated for each sample were: 180, 234 and 250 bar. The column and separator units were maintained at 313 K in all experimental extractions. The pressure of the first separator unit was about 1.8 times lower than the column pressure, and the second separator was maintained at a low pressure (20–30 bar) during all the experiments. The bottom product (raffinate) and liquid fractions in the separators (extract) were collected, weighed and analyzed after the extraction was completed. The closure of material balance was ascertained in all experiments within an accuracy of 10%. 2.3. Analytical methods 2.3.1. HPLC analysis The composition of the neutral lipids was evaluated on a Kromasil 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 gradient module FCV-10AL VP, a DGU-14A degasser, and a evaporative light scattering detector ELSD-LT from Shimadzu (IZASA, Spain). The ELSD conditions were 2.2 bars, 35 °C, and gain three. The flow rate was 2 mL/min. A splitter valve was used after the column and only 50% of the mobile phase was directed through the detector. The column temperature was maintained at 35 °C. The mobile phase utilized has previously been reported by Torres et al. (2005). 2.3.2. Oleic acid analysis by titration with KOH The acid value of raffinates (deacidified oil) was determined according to the volumetric method described in the Official Analytical Methods for Edible Fats and Oils, published by the Ministry of Agriculture of Spain in 1974. The samples were dissolved in a mixture of ethanol:ethyl ether 1:1. The acid value was calculated by titration of the raffinates with an ethanolic solution of potassium hydroxide (0.1 M). This solution was calibrated with dipotassium phthalate and phenolphthalein as indicator, in order to determine its normality accurately. Taking into consideration that oleic acid is the main fatty acid in olive oil, the acid value was expressed as percentage of oleic acid:
Acid value ð% oleic acidÞ ¼ VMN=10P where V and N are, respectively, the volume (mL) and normality of the ethanolic solution of KOH, M is the molecular weight of oleic acid (282.45) and P is the weight of the sample (g). 3. Thermodynamic modeling Phase equilibrium behavior of olive oil + SC-CO2 has been studied using GC-EoS model (Chen et al., 2000). This model has two
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Fig. 1. Pilot plant SFE device.
contributions to the residual Helmholtz energy of the system: a repulsive free volume term and a contribution which accounts for attractive group interactions. The GC-EoS equations can be found in the Appendix. For a detailed description of the model the reader is referred to Skjold-Jørgensen (1984, 1988). The attractive term of the GC-EoS model is a group contribution version of the NRTL model, and has five pure group parameters (T*, 0 q, g*, g0 and g00 ) and four binary interaction parameters (kij , kij , aij and aji). The repulsive term is modeled assuming hard sphere behavior for the molecules; each species is characterized by its critical hard sphere diameter dci (see Appendix). Phase equilibrium calculation was performed by representing the multicomponent olive oil feed material as a mixture comprising oleic acid, representing the FFA fraction, and triolein representing the (more abundant) neutral oil fraction. The required pure component parameters (critical temperature, critical pressure and critical hard sphere molecular diameter) for oleic acid and triolein are given in Table 1. All pure group and binary interaction parameters used in this work to represent the vapor–liquid equilibrium behavior of the triolein + CO2 and oleic acid + CO2 binary systems are summarized in Table 2. The capability of the GC-EoS model to represent high-pressure phase equilibria of triglyceride + CO2 binary mixture has been discussed elsewhere (Espinosa et al., 2002; Florusse et al., 2004). Fig. 2 shows a comparison between experimental and calculated phase equilibrium compositions of the triolein + CO2 mixture at two different temperatures, 313 and 333 K, and pressures up to 50 MPa.
Table 1 Pure component parameters
Oleic acida Trioleinb a b
Tc (K)
pc (bar)
dc (cm/mol)
796.3 1043.3
12.4 4.57
7.457 11.839
Joback and Reid (1983). Espinosa et al. (2002).
Despite the discrepancies between the different sources of experimental data (Bharath et al., 1992; Chen et al., 2000; Weber et al., 1999; Nilsson et al., 1991), the GC-EoS model provides a reasonable prediction of the triolein + CO2 phase equilibrium behavior. In order to describe the oleic acid + CO2 binary mixture the COOH–CO2 group interaction parameters were optimized in this work using vapor–liquid equilibria data for oleic acid + CO2 mixture (Bharath et al., 1992; Zou et al., 1990; Yu et al., 1992). The parameters obtained in the regression procedure are given in Table 2. The absolute average relative deviation between experimental and .calcu P cal exp Z exp lated CO2 mole fractions AARD ¼ 1=N exp Z CO2 Z CO2 CO2 in the vapor and liquid phases were, respectively, 0.07% and 3.01%. The phase equilibria representation of the oleic acid + CO2 mixture obtained with the GC-EoS model and parameters reported in Table 2 is given in Fig. 3. Phase equilibria prediction of the oleic acid + triolein + CO2 ternary mixture was compared with the experimental data reported by Simoes and Brunner (1996) for mixtures of commercial olive oil + CO2 in the temperature range from 313 to 353 K and pressures up to 300 bar. The experimental liquid and vapor mass fractions reported by Simoes and Brunner (1996) correspond to the quaternary system FFA + triglycerides + squalene + CO2. Since the squalene content in commercial olive oil was very low (0.7 wt%), the oil was simulated as the binary oleic acid + triolein. The CO2/oil mass ratio was equal to 1. Fig. 4 shows a comparison between the experimental and predicted oleic acid mass fraction (CO2-free) in the liquid (deacidified) oil phases. AARD between experimental and calculated 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 and 15.3 wt% FFA. This means that model prediction worsens as the acidy of the olive oil increases, as can be observed in Fig. 4.
4. Results and discussion The extraction temperature (313 K) was selected in agreement with previous studies reported in the literature (Brunetti et al.,
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Table 2 GC-EoS pure group and interaction parameters used in this work Pure group parameters
Reference temperature T*
Group surface area q
Pure group energy parameters 0
g
g
g
316,910 356,080 403,590 1211745.4 346,350 531,890
0.9274 0.8755 0.7631 0.1105 1.3469 0.5780
CH3 CH2 CH@CH COOH (CH3COO)2CH2COO triglyceride group (TG) CO2
600 600 600 600 600 304.2
0.848 0.540 0.867 1.224 3.948 1.261
Binary group interaction parameters i
j
Attractive energy parameters
Reference 00
0.0 0.0 0.0 0.0 0.0 0.0
Skjold-Jorgensen (1988) Skjold-Jorgensen (1988) Pusch and Schmelzer (1993) Ferreira et al. (2003) Espinosa et al. (2002) Skjold-Jorgensen (1988)
Non-randomness parameters
0
aij
aij
CO2
CH3 CH2 CH@CH COOH TG
0.0898 0.874 0.948 0.789 1.094
0.0 0.0 0.0 0.0 0.112
4.683 4.683 0.0 1.9351 1.651
4.683 4.683 0.0 0.2403 1.651
Espinosa et al. (2002) Espinosa et al. (2002) Ferreira et al. (2003) This work Espinosa et al. (2002)
TG
CH3/CH2 CH@CH COOH
0.860 0.883 1.062
0.0 0.0 0.0
0.0 0.0 0.0
0.0 0.0 0.0
Espinosa et al. (2002) Espinosa et al. (2002) Ferreira et al. (2003)
COOH
CH3/CH2 CH@CHa
0.932 0.932
0.0 0.0
2.946 2.946
2.424 2.424
Ferreira et al. (2003) Ferreira et al. (2003)
a
kij
kij
Assumed to be equal to the CH3/CH2–COOH interaction parameters.
b
a
50
30
40
20
P (MPa)
P (MPa)
25
15
30
20 10
10
5
0
0 0.0
0.1
0.2
0.3
0.98
0.99
1.00
CO2 mass fraction
0.0
0.1
0.2
0.3
0.98
0.99
1.00
CO2 mass fraction
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.
1989; Simoes and Brunner, 1996; Bondioli et al., 1992). In the temperature 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 high fatty acid solubility in the extractive solvent. Higher temperatures have also been tested (353 K) in the literature; in this case, a good selectivity of FFA extraction towards squalene was found (Bondioli et al., 1992), while similar conclusions were drawn when comparing selectivity of FFA vs triglycerides. Bondioli et al. (1992) studied the SFE deacidification of olive oil in the operational range of 90–150 bar, 313–333 K and CO2/oil flux ratios from 20 to 170. They could reduce oil acidity from 6.3 wt% to values less than 1 wt% at 313 K, 130 bar and CO2/oil flux ratios greater 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 system. Thus, pressures greater than those employed by Bondioli et al. were explored (180, 234 and 250 bar). The higher pressures but lower CO2/oil flux ratio gave high raffinate yields, as those obtained at low extraction pressures but high CO2/oil flux ratios (Bondioli et al., 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 the lower 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) obtained in this work. Also given in the table is the acidity reduction factor, defined as the ratio of the crude oil acidity to raffinate acidity. As evident in Table 3, the acidity reduction factor is higher than 2 at pressures greater than 234 bar for all oil samples having FFA
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30
pressure (MPa)
25
20
15
10
5
0 0.00
0.10
0.20
0.30
0. 40
0.50 0.97 0.98 0.99 1.00
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-EoS calculations.
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 to 2.5 wt% can be deacidified by CO2-SFE to obtain a product with an acidity lower than 1 wt%. If the crude olive oil contains higher amounts of FFA, it is expected that higher pressures or higher CO2/oil flux ratios are required. For example, as mentioned before, Bondioli et al. (1992) were able to reduce the acidity from 6.3 to 1 wt% at lower pressures but employing CO2/oil flux ratios greater than 100. Although the experimental assays reported in Table 3 demonstrate that deacidification of olive oil using low CO2/oil flux ratios is possible, the extraction conditions employed imply a CO2 density in the range of 820–880 kg/m3, while olive oil density is around 900 kg/m3. This means that low differences between the liquid and supercritical phase would be expected and thus, the extractions reported in Table 3 should be considered as the results of a semi-continuous process. As previously mentioned, the upper part of the column (120 cm above the feed point) was employed to minimize the loss of neutral oil by carryover.
b
18
in liquid phase
%wt FFA (CO 2 free)
a
Besides the presumed semi-batch character of the experimental runs, a continuous countercurrent column with N theoretical stages was considered to simulate the process. The extraction process was simulated as a stripping operation, i.e., with the feed at the top of the column and no enriching section. The GC-EoS model was used to perform phase equilibria calculations. The number of theoretical stages that provided the best agreement between the experimental and the simulated data was N = 2. Crude olive oil was considered to be a binary oleic acid + triolein mixture. The SFE simulation provided a satisfactory prediction of the raffinate yield and acidity variation as a function of pressure and crude oil FFA content. The AARD between experimental and calculated yield was 1.6%, the AARD was 11.2% with respect to raffinate acidity for olive oil samples having 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 between experimental and calculated yield and acidity were, in general, higher at lower pressures. Such deviations may be attributed to the simplified olive oil composition assumed (i.e., the binary triolein + oleic acid mixture). The model predictions were also tested against the experimental data reported by Bondioli et al. (1992) for the SFE deacidification of olive oil with FFA content of 6.3 wt%. Again, a multistage continuous countercurrent column was assumed for the purpose of process simulation. In this case, the best agreement between the experimental data reported by Bondioli et al. and the simulated countercurrent column was achieved with N = 3. The variation of raffinate acidity in the pressure range examined by Bondioli et al. (80–150 bar) and explored in this work (180–250 bar) are represented in Fig. 5 and compared with the GC-EoS model predictions. Additionally, Fig. 6 compares model predictions with the variation in yield and acidity (Bondioli et al., 1992) as a function of CO2/oil flux ratio. As evident in Figs. 5 and 6, SFE deacidification is well-described by the GC-EoS model. However, Bondioli et al. have experimentally determined a minimum raffinate acidity of 0.7 wt% at 313 K, 130 bar and CO2/oil flux ratio = 100, which is not predicted by the model. Despite the simple binary composition assumed (oleic acid + triolein) the GC-EoS model can be used as a screening tool to analyze olive oil SFE deacidification over a wide range of temperatures (313–343 K) and pressures (130–300 bar). Taking into account the discrepancies between experimental and calculated phase equilibria compositions, the model is accurate for crude olive oils with FFA content up to ca. 10 wt%.
18
16
16
14
14
12
12
10
10
8
8
6
6
4
4
2
2 0
0 12
16
20
24
pressure / MPa
28
32
16
20
24
28
32
pressure / MPa
Fig. 4. Oil acidity + CO2 liquid–vapor equilibria. Experimental data from Simoes and Brunner (1996). Solid lines: GC-EoS predictions. (a) 313 K and (j) 2.9 wt%, (N) 7.6 wt% and (d) 15.3 wt% FFA in crude oil. (b) 323 K and (h) 2.9 wt% and (4) 7.6 wt% FFA in crude oil; (s) 353 K and 15.3 6 wt% FFA in crude oil.
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Table 3 Experimental 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 flux ratio = 20 Crude oil acidity
Extraction pressure (bar)
Raffinate yield (%)
Raffinate acidity
Experimental acidity reduction factor
Experimental
Calculated
Experimental
Calculated
0.5 0.5 0.5
250 234 180
87.3 87.3 95.2
87.4 87.2 90.1
0.27 0.31 0.45
0.17 0.19 0.29
1.82 1.59 1.11
1.0 1.0 1.0
250 234 180
86.7 87.9 93.0
87.1 87.2 90.4
0.40 0.42 0.50
0.35 0.39 0.58
2.50 2.38 2.00
2.5 2.5 2.5
250 234 180
85.8 87.3 92.6
86.2 86.4 89.5
0.93 0.88 1.21
0.89 0.99 1.50
2.68 2.84 2.07
3.0 3.0 3.0
250 234 180
86.1 86.4 92.3
86.0 86.2 89.8
0.99 1.29 1.49
1.07 1.21 1.81
3.03 2.33 2.01
4.0 4.0 4.0
250 234 180
84.1 86.1 92.9
85.5 85.8 89.5
1.43 1.85 2.31
1.46 1.64 2.45
2.80 2.16 1.73
7
10
100
8
90
6
80
4
70
2
60
5 4 3 2 1
raffinate yield (%)
raffinate acidity (%wt FFA)
raffinate acidity (%wt FFA)
6
0 70
100
130
160
190
220
250
pressure / MPa Fig. 5. Raffinate acidity (wt% FFA) obtained in SFE deacidification. Experimental data 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.
The GC-EoS was used to obtain the process conditions required to deacidify crude olive oil, with different FFA content, down to 0.7 wt% FFA. Considering the significant improvement in functional and organoleptic properties of crude olive oil SFE deacidification, this oil can potentially be blended with virgin olive oil to produce commercial olive oil of high quality. The SFE process was mathematically solved within a sequential process simulator that includes rigorous models for a high-pressure multistage extractor (Brignole et al., 1987). In these routines (FORTRAN language) thermodynamic phase equilibria calculations were performed using the GC-EOS model. The number of stages employed to carry out the calculations was 10; additional stages did not produce any significant changes in raffinate and extract compositions. Furthermore, considering that the height equivalent theoretical stage (HETP) is around 0.5–1.0 m (Hurtado-Benavides et al., 2004; Simoes and Brunner, 1996), a reasonably value can be deduced for the total height of the packed column. Fig. 7 shows the required variation of CO2/oil ratio (Fig. 7a) and process yield (Fig. 7b) as a function of the crude olive oil acidity, for different extraction conditions (temperature and pressure). Extrac-
0 0
50
100
150
50 200
CO2/oil flux ratio Fig. 6. Raffinate yield (d) and acidity (s) obtained in olive oil SFE deacidification at 313 K, 130 bar and different CO2/oil flux ratios (1992). Solid lines: GC-EoS calculations.
tion temperature and pressure were selected to maintain a CO2 density of 740 Kg/m3, so as to guarantee a reasonable difference between the down-flowing liquid and the up-flowing supercritical solvent. As observed in Fig. 7a, the CO2/oil ratio can be reduced twofold when extraction pressure is increased from 130 bar (extraction temperature 313 K) to 250 bar (extraction temperature 343 K). Additionally, raffinate yields are mainly determined by the quality of the crude oil processed (see Fig. 7b), with values decreasing from ca. 90% to 77% when crude oil acidity increases from 2 to 10 wt%. Besides the higher capital costs that a more intense set of extraction conditions would demand, it has to be noted that the power required for CO2 recirculation at 343 K and 250 bar is significantly 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 oil with 0.5–4 wt% FFA content was experimentally carried out and compared with SFE data reported in the literature. Also, a simple model to simulate the deacidification process was developed. The model was tested using experimental phase equilibria data for
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a
b
160
95
140
90
raffinate yield (%)
CO2/oil flux ratio
120 100 80 60
85
80
40 75 20 70
0 0
2
4
6
8
0
12
10
2
4
6
8
10
12
crude oil acidity (%wt FFA)
crude oil acidity (%wt FFA)
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.
where Table 4 Comparison of the countercurrent SFE deacidification of crude olive oil from 6 to 0.7 wt% FFA at different extraction conditions (CO2 density = 740 kg/m3) as predicted by the GC-EoS model Temperature (K)
Pressure (bar)
Raffinate yield (%)
CO2 fluxa (kg CO2/kg oil)
CO2 recirculation powerb (kJ/kg oil)
313 323 333 343
130 170 210 250
83.2 82.2 81.4 80.8
123.2 94.6 76.6 63.8
21,691 15,188 11,122 8292
a b
pk3 1 Y ¼ 1 6V
and kk ¼
NC X
k
ni di
ðA:3Þ
i
NC is the number of components, ni is the number of moles of component i and V is the total volume. A temperature dependent generalized expression is assumed for di:
di ¼ 1:065655dci f1 0:12 exp½2T ci =ð3TÞg
ðA:4Þ
dci is the pure component critical hard sphere diameter, given by:
dci ¼ ð0:08943RT ci =Pci Þ1=3
Required to achieve a raffinate oil with 0.7 wt% of FFA. Considering a single separator tank at 313 K and 50 bar.
olive oil (with different acidity) + CO2 mixtures, and was employed to study the required variation of process conditions to ensure a final acidity value lower than 0.7% in the raffinate. Acknowledgements This work has been financed under a R&D program of the Spanish Ministry of Education and Science, Project AGL-200407227-C02-01 and Project S-0505/AGR/000153 from the Comunidad Autónoma de Madrid. T.F. would like to acknowledge the financial support of the Ramon y Cajal Program from the Ministry of Education and Science. The authors acknowledge the financial support of Migasa (Spain). Appendix
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 component vapor pressure curve, usually the normal boiling point. Since vapor pressure data for low volatile or thermolabile substances is often not available or not reliable, infinite dilution activity coefficients can be used to estimate the dci parameter of high molecular weight compounds, such as alkanes and triglycerides, as demonstrated by Espinosa et al. (2002). For the evaluation of the attractive contribution to the Helmholtz energy, a group contribution version of a densitydependent NRTL-type expression is derived:
, NC NG NG NG X X X zX i A =RT ¼ ni mj qj ðhk g kj q~qÞ hl slj 2 i j i k att
NC X hj ¼ qj =q ni mij
q¼
The residual Helmholtz energy in the GC-EoS model is described by two terms:
sij ¼ exp aij Dg ij q~=ðRTVÞ
A ¼ Afv þ Aatt
Dg ij ¼ g ij g jj
The free volume contribution (Afv) is modeled assuming hard sphere behavior for the molecules, characterizing each substance i by a hard sphere diameter di. A Carnahan–Starling type of hard sphere expression for mixtures is adopted:
Afv =RT ¼ 3ðk1 k2 =k3 ÞðY 1Þ þ k32 =k23 Y þ Y 2 ln Y þ n ln Y ðA:2Þ
ðA:6Þ
where
i
ðA:1Þ
ðA:5Þ
NC X i
ni
NG X
mij qj
j
NG is the number of groups, z is the number of nearest neighbors to any segment (set to 10), mij is the number of groups type j in molecule i, qj is the number of surface segments assigned to ~ is the total number group j, hk is the surface fraction of group k, q of surface segments, gij is the attraction energy parameter for interactions between groups i and j, and aij is the NRTL non-randomness parameter (aij – aji). The interactions between unlike groups are calculated from
470
g ij ¼ kij ðg ii g jj Þ1=2
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ðkij ¼ kji Þ
with the following temperature dependencies for the interaction parameters:
g jj ¼ g jj 1 þ g 0jj T=T j 1 þ g 00jj ln T=T j n h . io kij ¼ kij 1 þ kij ln 2T T i þ T j being g jj the interaction parameter at the reference temperature T i . References Bharath, R., Inomata, H., Adschiri, T., Arai, K., 1992. Phase equilibria study for the separation and fractionation of fatty oil components using supercritical carbon dioxide. Fluid Phase Equilibria 81, 307–320. Bondioli, P., Mariani, C., Mossa, E., Fedelli, A., Muller, A., 1992. Lampante olive oil refining with supercritical carbon dioxide. Journal of the American Oil Chemists’ Society 69, 477–480. Bosle, B.M., Subramanian, R., 2005. New approaches in deacidification of edible oils – a review. Journal of Food Engineering 69, 481–494. Brignole, E.A., Andersen, P.M., Fredenslund, A.A., 1987. Supercritical fluid extraction of alcohols from water. Industrial and Engineering Chemistry Research 26, 254– 261. Brunetti, L., Daghetta, A., Fedeli, E., Kikic, I., Zanderighi, L., 1989. Deacidification of olive oils by supercritical carbon dioxide. Journal of the American Oil Chemists’ Society 66, 209–217. Chen, C., Chang, C.J., Yang, P., 2000. Vapor-liquid equilibria of carbon dioxide with linoleic acid, a-tocopherol and triolein at elevated pressures. Fluid Phase Equilibria 175, 107–115. Espinosa, S., Fornari, T., Bottini, S., Brignole, E.A., 2002. Phase equilibria in mixtures of fatty oils and derivatives with near critical fluids using the GC-EoS. Journal of Supercritical Fluids 23, 91–102. European Council Regulation No. 1513/2001. Descriptions and Definitions of Olive Oils and Olive-Pomace Oils. Official Journal of the European Communities 26.7.2001.
Ferreira, O., Fornari, T., Brignole, E.A., Bottini, S.B., 2003. Modelling of association effects in mixtures of carboxylic acids with associating and non-associating components. Latin American Applied Research 33, 217–222. Florusse, L.J., Van’t Hof, A., Fornari, T., Bottini, S., Peters, C.J., 2004. Phase behavior of binary systems of carbon dioxide and certain low-molecular weight triglycerides: measurements and thermodynamic modeling. Journal of Supercritical Fluids 31, 123–132. Hurtado-Benavides, A.M., Señoráns, F.J., Ibáñez, E., Reglero, G., 2004. Mass transfer evaluation of countercurrent supercritical fluid extraction of olive oil. Journal of Supercritical Fluids 28, 29–35. Joback, K.G., Reid, R.C., 1983. Estimation of pure-component properties from group contributions. Chemical Engineering Communications 57, 233–238. Nilsson, W.B., Gauglitz, E.J., Hudson, J.K., 1991. Solubilities of methyl oleate, oleic acid, oleyl glycerols, and oleyl glycerols mixtures in supercritical carbon dioxide. Journal of the American Oil Chemists’ Society 68, 87–91. Pusch, J., Schmelzer, J., 1993. Extension of the group contribution equation of state parameter matrix for the prediction of phase equilibria containing argon, ammonia, propene and other alkenes. Berichte der Bunsen-Gesellschaft fuer Physikalische Chemie 97, 597–603. Simoes, P.C., Brunner, R., 1996. Multicomponent phase equilibria of an extra-virgin olive oil in supercritical carbon dioxide. Journal of Supercritical Fluids 9, 75–81. Skjold-Jørgensen, S., 1984. Gas solubility calculations II. Application of a new groupcontribution equation of state. Fluid Phase Equilibria 16, 317–351. Skjold-Jorgensen, S., 1988. Group contribution equation of state (GC-EOS): a predictive method for phase equilibrium computations over wide ranges of temperatures and pressures up to 30 MPa. Industrial and Engineering Chemistry Research 27, 110–118. Torres, C.F., Vázquez, L., Señoráns, F.J., Reglero, G., 2005. Study of the analysis of alkoxyglycerols and other non-polar lipids by liquid chromatography coupled with evaporative light scattering detector. Journal of Chromatography A 1078, 28–34. Weber, W., Petkov, S., Brunner, G., 1999. Vapour-liquid-equilibria and calculations using the Redlich–Kwong–Aspen-equation of state for tristerarin, tripalmitin and triolein in CO2 and propane. Fluid Phase Equilibria 158–160, 695–706. Yu, Z.R., Rizvi, S.S.H., Zollweg, J.A., 1992. Phase equilibrium of oleic acid methyl oleate and anhydrous milk fat in supercritical CO2. Journal of Supercritical Fluids 5, 114–122. Zou, M., Yu, Z.R., Kashulines, P., Rizvi, S.S.H., Zollweg, J.A., 1990. Fluid-liquid phase equilibria of fatty acids and fatty acid methyl esters in supercritical carbon dioxide. Journal of Supercritical Fluids 3, 23–28.
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Journal of Food Engineering journal homepage: www.elsevier.com/locate/jfoodeng
Deacidification of olive oil by countercurrent supercritical carbon dioxide extraction: 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, Spain Facultad de Ingeniería Agroindustrial, Universidad de Nariño, Pasto, Colombia c Instituto de Fermentaciones Industriales, CSIC, Juan de la Cierva, 3, 28006 Madrid, Spain b
a r t i c l e
i n f o
Article history: Received 17 March 2008 Received in revised form 8 July 2008 Accepted 13 July 2008 Available online 19 July 2008 Keywords: Deacidification Fatty acids Supercritical fluid extraction Group contribution equation of state Simulation
a b s t r a c t Supercritical carbon dioxide was used as an extractive solvent to remove free fatty acids from coldpressed olive oil. Crude oil of different acidity content (from 0.5 to 4.0 wt%) was extracted in a packed column at 313 K and pressures of 180, 234 and 250 bar. The group contribution equation of state was employed to simulate the separation process, representing the oil as a simple pseudo-binary oleic acid + triolein mixture. Despite the simple representation of oil composition to simulate the deacidification process, a satisfactory agreement between the experimental and calculated yields and acidity of raffinates was obtained. The thermodynamic model was employed to study a continuous countercurrent multistage extraction process which yielded a raffinate having acidity lower than 0.7 wt%, when crude olive oil with different FFA content was processed. Ó 2008 Elsevier Ltd. All rights reserved.
1. Introduction Olive oil is commercially obtained by cold-pressing processes. Depending on biological, meteorological, agricultural factors or processing conditions, the crude oil obtained contains different amounts of free fatty acids (FFA). These substances are also susceptible to oxidation, leading to rancidity and conferring an undesired flavor to the oil. Therefore, high amounts of FFA in the oil must be absolutely avoided. Furthermore, the lower the FFA content in the virgin olive oil, the higher is its commercial value. According to the European Union regulations (the major world producer of olive oil), 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 with virgin olive oil. Additionally, high valued olive oil (extra virgin olive oil) must contain less than 0.8% of FFA (European Council Regulation No. 1513/2001). Therefore, the deacidification of crude olive oil is important not only for consumer acceptance but also because it has the maximum economic impact on production. The removal of FFA from crude oil is, therefore, a crucial step in olive oil production since it predominantly determines the quality of the final product. The well-known adverse effects of chemical or physical refining pro* Corresponding author. E-mail address: elena@ifi.csic.es (E. Ibáñez). 0260-8774/$ - see front matter Ó 2008 Elsevier Ltd. All rights reserved. doi:10.1016/j.jfoodeng.2008.07.012
cesses on the oil quality reduce its market value (Bondioli et al., 1992). New alternative deacidification processes proposed are reesterification, solvent extraction, biological deacidification, membrane technology and supercritical fluid extraction (SFE). Each of these alternatives has its own advantages and drawbacks (Bosle and Subramanian, 2005). SFE using carbon dioxide is a low temperature and a relatively pollution free operation. Its high selectivity permits the removal of FFA from the oil with minimum loss of neutral oil: triglycerides and unsaponifiable matter (tocopherols, sterols and vitamins). Thus, when this technique is applied, the deacidification process can be carried out without significant loss in yield or the nutritional properties (Brunetti et al., 1989). Brunetti et al. (1989) have investigated the extraction of fatty acids from fatty acid + triglyceride mixtures using supercritical carbon dioxide (SC-CO2). Experiments were carried out on samples with different FFA content (from 2.6 up to 20 wt%) at pressures of 20 and 30 MPa and temperatures of 313 and 333 K. Besides the limitations concerned with the use of batch equipment, they concluded that the SFE was particularly suitable for deacidification of olive oils with FFA content lower than 10%, since the selectivity factor for fatty acid extraction increases as the concentration of FFA in the crude oil decreases. On the other hand, Simoes and Brunner (1996) evaluated the possibility of using SC-CO2 to deacidify olive oil using a commercial oil containing squalene (around 0.7 wt%), FFA (from 3 to 15 wt%) and triglycerides. Experimental
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L. Vázquez et al. / Journal of Food Engineering 90 (2009) 463–470
phase equilibria measurements covering a wide range of extraction pressures and temperatures (from 313 to 353 K and up to 30 MPa) were performed and these data were used to simulate a countercurrent packed column for the deacidification of olive oil by SFE. Simulated results showed an increase of extraction yield with CO2 density. Additionally, higher solvent-to-feed flow ratios improved the extraction yield, resulting in lower raffinate acidities. The continuous CO2-SFE of olive oil from different geographic origins (Italy, Spain and Tunisia) in a countercurrent packed column of 3 m high was studied by Bondioli et al. (1992). The influence 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 the injection point, which demonstrated the efficacy of a rectifying section (obtained by temperature gradient) to reduce the loss of neutral oil. In the present work, the deacidification of olive oil using CO2SFE was experimentally studied and thermodynamically simulated, investigating pressures higher than those employed by Bondioli et al. (1992). Extractions were performed in a packed column 3 m high, without external or internal reflux and utilizing 1.8 m of the column as the stripping section; the upper part of the column was used to avoid carryover of the crude oil. All experimental assays were performed at 313 K, with a CO2/oil flux ratio of 20 and pressures of 180, 234 and 250 bar. The acidity of the olive oil samples varied from 0.5 to 4.0 wt%, which is the normal range observed in the olive oil obtained by conventional cold-pressing. The experimental results were simulated using the group contribution equation of state (GC-EoS) (Skjold-Jørgensen, 1984) in a completely predictive manner. The model was previously tested with the experimental phase equilibrium data reported by Simoes and Brunner (1996) and the countercurrent extraction data reported by Bondioli et al. (1992). A reasonable agreement between model predictions and experimental phase equilibria compositions, experimental yield and refined oil acidity was achieved. The main aim of the present work was to develop a simple predictive tool which systematically evaluated the extraction process conditions that guaranteed a target low acidity in the raffinate, when crude olive 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 company (Migasa, Sevilla, Spain). Carbon dioxide N38 (99.98%), was purchased from AL Air Liquide España S.A. (Madrid, Spain). All solvents 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 study is shown in Fig. 1. The heart of the pilot plant is a 316 stainless steel extraction column (18 mm i.d.) packed with Fenske rings (3 0.5 mm; Afora, S.A. Spain). The total height of the extraction column is 300 cm with provision for sample introduction at three levels: top, middle and bottom, as shown in Fig. 1. In this work, the olive oil samples are introduced in the middle, giving an effective packed height of 180 cm (measured from the point of introduction to the CO2 feed point). Ten thermocouples with temperature controllers are installed along the length of the extraction column, located at heights of 60, 150, 210, 270 and 300 cm measured from bottom.
The SFE system also comprises two separator cells where a cascade decompression takes place and a cryogenic trap at atmospheric pressure. Both CO2 and liquid feed sample are preheated at the discharge of their respective pumps (Dosapro Milton Roy) before introduction into the SFE device. The temperature of the extraction column and separator units is maintained by using a silicone heater bath. The plant has computerized PLC-based instrumentation and a control system, with several safety devices including valves and alarms. A continuous flow of CO2 (2.7 kg/h) was introduced into the column at the bottom. When the set operating pressure and temperature were reached, 135 g/h of oil was pumped continuously over the duration of extraction (90 min). A CO2 to oil flux ratio of 20 was employed in all the experiments. When the extraction was completed, the oil flow was stopped and CO2 was pumped for another 30 min. Olive oil having different degree of acidity (0.5, 1.0, 2.5 and 4.0 wt%) was employed as the feed. The extraction pressures investigated for each sample were: 180, 234 and 250 bar. The column and separator units were maintained at 313 K in all experimental extractions. The pressure of the first separator unit was about 1.8 times lower than the column pressure, and the second separator was maintained at a low pressure (20–30 bar) during all the experiments. The bottom product (raffinate) and liquid fractions in the separators (extract) were collected, weighed and analyzed after the extraction was completed. The closure of material balance was ascertained in all experiments within an accuracy of 10%. 2.3. Analytical methods 2.3.1. HPLC analysis The composition of the neutral lipids was evaluated on a Kromasil 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 gradient module FCV-10AL VP, a DGU-14A degasser, and a evaporative light scattering detector ELSD-LT from Shimadzu (IZASA, Spain). The ELSD conditions were 2.2 bars, 35 °C, and gain three. The flow rate was 2 mL/min. A splitter valve was used after the column and only 50% of the mobile phase was directed through the detector. The column temperature was maintained at 35 °C. The mobile phase utilized has previously been reported by Torres et al. (2005). 2.3.2. Oleic acid analysis by titration with KOH The acid value of raffinates (deacidified oil) was determined according to the volumetric method described in the Official Analytical Methods for Edible Fats and Oils, published by the Ministry of Agriculture of Spain in 1974. The samples were dissolved in a mixture of ethanol:ethyl ether 1:1. The acid value was calculated by titration of the raffinates with an ethanolic solution of potassium hydroxide (0.1 M). This solution was calibrated with dipotassium phthalate and phenolphthalein as indicator, in order to determine its normality accurately. Taking into consideration that oleic acid is the main fatty acid in olive oil, the acid value was expressed as percentage of oleic acid:
Acid value ð% oleic acidÞ ¼ VMN=10P where V and N are, respectively, the volume (mL) and normality of the ethanolic solution of KOH, M is the molecular weight of oleic acid (282.45) and P is the weight of the sample (g). 3. Thermodynamic modeling Phase equilibrium behavior of olive oil + SC-CO2 has been studied using GC-EoS model (Chen et al., 2000). This model has two
L. Vázquez et al. / Journal of Food Engineering 90 (2009) 463–470
465
Fig. 1. Pilot plant SFE device.
contributions to the residual Helmholtz energy of the system: a repulsive free volume term and a contribution which accounts for attractive group interactions. The GC-EoS equations can be found in the Appendix. For a detailed description of the model the reader is referred to Skjold-Jørgensen (1984, 1988). The attractive term of the GC-EoS model is a group contribution version of the NRTL model, and has five pure group parameters (T*, 0 q, g*, g0 and g00 ) and four binary interaction parameters (kij , kij , aij and aji). The repulsive term is modeled assuming hard sphere behavior for the molecules; each species is characterized by its critical hard sphere diameter dci (see Appendix). Phase equilibrium calculation was performed by representing the multicomponent olive oil feed material as a mixture comprising oleic acid, representing the FFA fraction, and triolein representing the (more abundant) neutral oil fraction. The required pure component parameters (critical temperature, critical pressure and critical hard sphere molecular diameter) for oleic acid and triolein are given in Table 1. All pure group and binary interaction parameters used in this work to represent the vapor–liquid equilibrium behavior of the triolein + CO2 and oleic acid + CO2 binary systems are summarized in Table 2. The capability of the GC-EoS model to represent high-pressure phase equilibria of triglyceride + CO2 binary mixture has been discussed elsewhere (Espinosa et al., 2002; Florusse et al., 2004). Fig. 2 shows a comparison between experimental and calculated phase equilibrium compositions of the triolein + CO2 mixture at two different temperatures, 313 and 333 K, and pressures up to 50 MPa.
Table 1 Pure component parameters
Oleic acida Trioleinb a b
Tc (K)
pc (bar)
dc (cm/mol)
796.3 1043.3
12.4 4.57
7.457 11.839
Joback and Reid (1983). Espinosa et al. (2002).
Despite the discrepancies between the different sources of experimental data (Bharath et al., 1992; Chen et al., 2000; Weber et al., 1999; Nilsson et al., 1991), the GC-EoS model provides a reasonable prediction of the triolein + CO2 phase equilibrium behavior. In order to describe the oleic acid + CO2 binary mixture the COOH–CO2 group interaction parameters were optimized in this work using vapor–liquid equilibria data for oleic acid + CO2 mixture (Bharath et al., 1992; Zou et al., 1990; Yu et al., 1992). The parameters obtained in the regression procedure are given in Table 2. The absolute average relative deviation between experimental and .calcu P cal exp Z exp lated CO2 mole fractions AARD ¼ 1=N exp Z CO2 Z CO2 CO2 in the vapor and liquid phases were, respectively, 0.07% and 3.01%. The phase equilibria representation of the oleic acid + CO2 mixture obtained with the GC-EoS model and parameters reported in Table 2 is given in Fig. 3. Phase equilibria prediction of the oleic acid + triolein + CO2 ternary mixture was compared with the experimental data reported by Simoes and Brunner (1996) for mixtures of commercial olive oil + CO2 in the temperature range from 313 to 353 K and pressures up to 300 bar. The experimental liquid and vapor mass fractions reported by Simoes and Brunner (1996) correspond to the quaternary system FFA + triglycerides + squalene + CO2. Since the squalene content in commercial olive oil was very low (0.7 wt%), the oil was simulated as the binary oleic acid + triolein. The CO2/oil mass ratio was equal to 1. Fig. 4 shows a comparison between the experimental and predicted oleic acid mass fraction (CO2-free) in the liquid (deacidified) oil phases. AARD between experimental and calculated 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 and 15.3 wt% FFA. This means that model prediction worsens as the acidy of the olive oil increases, as can be observed in Fig. 4.
4. Results and discussion The extraction temperature (313 K) was selected in agreement with previous studies reported in the literature (Brunetti et al.,
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L. Vázquez et al. / Journal of Food Engineering 90 (2009) 463–470
Table 2 GC-EoS pure group and interaction parameters used in this work Pure group parameters
Reference temperature T*
Group surface area q
Pure group energy parameters 0
g
g
g
316,910 356,080 403,590 1211745.4 346,350 531,890
0.9274 0.8755 0.7631 0.1105 1.3469 0.5780
CH3 CH2 CH@CH COOH (CH3COO)2CH2COO triglyceride group (TG) CO2
600 600 600 600 600 304.2
0.848 0.540 0.867 1.224 3.948 1.261
Binary group interaction parameters i
j
Attractive energy parameters
Reference 00
0.0 0.0 0.0 0.0 0.0 0.0
Skjold-Jorgensen (1988) Skjold-Jorgensen (1988) Pusch and Schmelzer (1993) Ferreira et al. (2003) Espinosa et al. (2002) Skjold-Jorgensen (1988)
Non-randomness parameters
0
aij
aij
CO2
CH3 CH2 CH@CH COOH TG
0.0898 0.874 0.948 0.789 1.094
0.0 0.0 0.0 0.0 0.112
4.683 4.683 0.0 1.9351 1.651
4.683 4.683 0.0 0.2403 1.651
Espinosa et al. (2002) Espinosa et al. (2002) Ferreira et al. (2003) This work Espinosa et al. (2002)
TG
CH3/CH2 CH@CH COOH
0.860 0.883 1.062
0.0 0.0 0.0
0.0 0.0 0.0
0.0 0.0 0.0
Espinosa et al. (2002) Espinosa et al. (2002) Ferreira et al. (2003)
COOH
CH3/CH2 CH@CHa
0.932 0.932
0.0 0.0
2.946 2.946
2.424 2.424
Ferreira et al. (2003) Ferreira et al. (2003)
a
kij
kij
Assumed to be equal to the CH3/CH2–COOH interaction parameters.
b
a
50
30
40
20
P (MPa)
P (MPa)
25
15
30
20 10
10
5
0
0 0.0
0.1
0.2
0.3
0.98
0.99
1.00
CO2 mass fraction
0.0
0.1
0.2
0.3
0.98
0.99
1.00
CO2 mass fraction
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.
1989; Simoes and Brunner, 1996; Bondioli et al., 1992). In the temperature 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 high fatty acid solubility in the extractive solvent. Higher temperatures have also been tested (353 K) in the literature; in this case, a good selectivity of FFA extraction towards squalene was found (Bondioli et al., 1992), while similar conclusions were drawn when comparing selectivity of FFA vs triglycerides. Bondioli et al. (1992) studied the SFE deacidification of olive oil in the operational range of 90–150 bar, 313–333 K and CO2/oil flux ratios from 20 to 170. They could reduce oil acidity from 6.3 wt% to values less than 1 wt% at 313 K, 130 bar and CO2/oil flux ratios greater 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 system. Thus, pressures greater than those employed by Bondioli et al. were explored (180, 234 and 250 bar). The higher pressures but lower CO2/oil flux ratio gave high raffinate yields, as those obtained at low extraction pressures but high CO2/oil flux ratios (Bondioli et al., 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 the lower 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) obtained in this work. Also given in the table is the acidity reduction factor, defined as the ratio of the crude oil acidity to raffinate acidity. As evident in Table 3, the acidity reduction factor is higher than 2 at pressures greater than 234 bar for all oil samples having FFA
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L. Vázquez et al. / Journal of Food Engineering 90 (2009) 463–470 35
30
pressure (MPa)
25
20
15
10
5
0 0.00
0.10
0.20
0.30
0. 40
0.50 0.97 0.98 0.99 1.00
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-EoS calculations.
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 to 2.5 wt% can be deacidified by CO2-SFE to obtain a product with an acidity lower than 1 wt%. If the crude olive oil contains higher amounts of FFA, it is expected that higher pressures or higher CO2/oil flux ratios are required. For example, as mentioned before, Bondioli et al. (1992) were able to reduce the acidity from 6.3 to 1 wt% at lower pressures but employing CO2/oil flux ratios greater than 100. Although the experimental assays reported in Table 3 demonstrate that deacidification of olive oil using low CO2/oil flux ratios is possible, the extraction conditions employed imply a CO2 density in the range of 820–880 kg/m3, while olive oil density is around 900 kg/m3. This means that low differences between the liquid and supercritical phase would be expected and thus, the extractions reported in Table 3 should be considered as the results of a semi-continuous process. As previously mentioned, the upper part of the column (120 cm above the feed point) was employed to minimize the loss of neutral oil by carryover.
b
18
in liquid phase
%wt FFA (CO 2 free)
a
Besides the presumed semi-batch character of the experimental runs, a continuous countercurrent column with N theoretical stages was considered to simulate the process. The extraction process was simulated as a stripping operation, i.e., with the feed at the top of the column and no enriching section. The GC-EoS model was used to perform phase equilibria calculations. The number of theoretical stages that provided the best agreement between the experimental and the simulated data was N = 2. Crude olive oil was considered to be a binary oleic acid + triolein mixture. The SFE simulation provided a satisfactory prediction of the raffinate yield and acidity variation as a function of pressure and crude oil FFA content. The AARD between experimental and calculated yield was 1.6%, the AARD was 11.2% with respect to raffinate acidity for olive oil samples having 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 between experimental and calculated yield and acidity were, in general, higher at lower pressures. Such deviations may be attributed to the simplified olive oil composition assumed (i.e., the binary triolein + oleic acid mixture). The model predictions were also tested against the experimental data reported by Bondioli et al. (1992) for the SFE deacidification of olive oil with FFA content of 6.3 wt%. Again, a multistage continuous countercurrent column was assumed for the purpose of process simulation. In this case, the best agreement between the experimental data reported by Bondioli et al. and the simulated countercurrent column was achieved with N = 3. The variation of raffinate acidity in the pressure range examined by Bondioli et al. (80–150 bar) and explored in this work (180–250 bar) are represented in Fig. 5 and compared with the GC-EoS model predictions. Additionally, Fig. 6 compares model predictions with the variation in yield and acidity (Bondioli et al., 1992) as a function of CO2/oil flux ratio. As evident in Figs. 5 and 6, SFE deacidification is well-described by the GC-EoS model. However, Bondioli et al. have experimentally determined a minimum raffinate acidity of 0.7 wt% at 313 K, 130 bar and CO2/oil flux ratio = 100, which is not predicted by the model. Despite the simple binary composition assumed (oleic acid + triolein) the GC-EoS model can be used as a screening tool to analyze olive oil SFE deacidification over a wide range of temperatures (313–343 K) and pressures (130–300 bar). Taking into account the discrepancies between experimental and calculated phase equilibria compositions, the model is accurate for crude olive oils with FFA content up to ca. 10 wt%.
18
16
16
14
14
12
12
10
10
8
8
6
6
4
4
2
2 0
0 12
16
20
24
pressure / MPa
28
32
16
20
24
28
32
pressure / MPa
Fig. 4. Oil acidity + CO2 liquid–vapor equilibria. Experimental data from Simoes and Brunner (1996). Solid lines: GC-EoS predictions. (a) 313 K and (j) 2.9 wt%, (N) 7.6 wt% and (d) 15.3 wt% FFA in crude oil. (b) 323 K and (h) 2.9 wt% and (4) 7.6 wt% FFA in crude oil; (s) 353 K and 15.3 6 wt% FFA in crude oil.
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Table 3 Experimental 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 flux ratio = 20 Crude oil acidity
Extraction pressure (bar)
Raffinate yield (%)
Raffinate acidity
Experimental acidity reduction factor
Experimental
Calculated
Experimental
Calculated
0.5 0.5 0.5
250 234 180
87.3 87.3 95.2
87.4 87.2 90.1
0.27 0.31 0.45
0.17 0.19 0.29
1.82 1.59 1.11
1.0 1.0 1.0
250 234 180
86.7 87.9 93.0
87.1 87.2 90.4
0.40 0.42 0.50
0.35 0.39 0.58
2.50 2.38 2.00
2.5 2.5 2.5
250 234 180
85.8 87.3 92.6
86.2 86.4 89.5
0.93 0.88 1.21
0.89 0.99 1.50
2.68 2.84 2.07
3.0 3.0 3.0
250 234 180
86.1 86.4 92.3
86.0 86.2 89.8
0.99 1.29 1.49
1.07 1.21 1.81
3.03 2.33 2.01
4.0 4.0 4.0
250 234 180
84.1 86.1 92.9
85.5 85.8 89.5
1.43 1.85 2.31
1.46 1.64 2.45
2.80 2.16 1.73
7
10
100
8
90
6
80
4
70
2
60
5 4 3 2 1
raffinate yield (%)
raffinate acidity (%wt FFA)
raffinate acidity (%wt FFA)
6
0 70
100
130
160
190
220
250
pressure / MPa Fig. 5. Raffinate acidity (wt% FFA) obtained in SFE deacidification. Experimental data 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.
The GC-EoS was used to obtain the process conditions required to deacidify crude olive oil, with different FFA content, down to 0.7 wt% FFA. Considering the significant improvement in functional and organoleptic properties of crude olive oil SFE deacidification, this oil can potentially be blended with virgin olive oil to produce commercial olive oil of high quality. The SFE process was mathematically solved within a sequential process simulator that includes rigorous models for a high-pressure multistage extractor (Brignole et al., 1987). In these routines (FORTRAN language) thermodynamic phase equilibria calculations were performed using the GC-EOS model. The number of stages employed to carry out the calculations was 10; additional stages did not produce any significant changes in raffinate and extract compositions. Furthermore, considering that the height equivalent theoretical stage (HETP) is around 0.5–1.0 m (Hurtado-Benavides et al., 2004; Simoes and Brunner, 1996), a reasonably value can be deduced for the total height of the packed column. Fig. 7 shows the required variation of CO2/oil ratio (Fig. 7a) and process yield (Fig. 7b) as a function of the crude olive oil acidity, for different extraction conditions (temperature and pressure). Extrac-
0 0
50
100
150
50 200
CO2/oil flux ratio Fig. 6. Raffinate yield (d) and acidity (s) obtained in olive oil SFE deacidification at 313 K, 130 bar and different CO2/oil flux ratios (1992). Solid lines: GC-EoS calculations.
tion temperature and pressure were selected to maintain a CO2 density of 740 Kg/m3, so as to guarantee a reasonable difference between the down-flowing liquid and the up-flowing supercritical solvent. As observed in Fig. 7a, the CO2/oil ratio can be reduced twofold when extraction pressure is increased from 130 bar (extraction temperature 313 K) to 250 bar (extraction temperature 343 K). Additionally, raffinate yields are mainly determined by the quality of the crude oil processed (see Fig. 7b), with values decreasing from ca. 90% to 77% when crude oil acidity increases from 2 to 10 wt%. Besides the higher capital costs that a more intense set of extraction conditions would demand, it has to be noted that the power required for CO2 recirculation at 343 K and 250 bar is significantly 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 oil with 0.5–4 wt% FFA content was experimentally carried out and compared with SFE data reported in the literature. Also, a simple model to simulate the deacidification process was developed. The model was tested using experimental phase equilibria data for
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a
b
160
95
140
90
raffinate yield (%)
CO2/oil flux ratio
120 100 80 60
85
80
40 75 20 70
0 0
2
4
6
8
0
12
10
2
4
6
8
10
12
crude oil acidity (%wt FFA)
crude oil acidity (%wt FFA)
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.
where Table 4 Comparison of the countercurrent SFE deacidification of crude olive oil from 6 to 0.7 wt% FFA at different extraction conditions (CO2 density = 740 kg/m3) as predicted by the GC-EoS model Temperature (K)
Pressure (bar)
Raffinate yield (%)
CO2 fluxa (kg CO2/kg oil)
CO2 recirculation powerb (kJ/kg oil)
313 323 333 343
130 170 210 250
83.2 82.2 81.4 80.8
123.2 94.6 76.6 63.8
21,691 15,188 11,122 8292
a b
pk3 1 Y ¼ 1 6V
and kk ¼
NC X
k
ni di
ðA:3Þ
i
NC is the number of components, ni is the number of moles of component i and V is the total volume. A temperature dependent generalized expression is assumed for di:
di ¼ 1:065655dci f1 0:12 exp½2T ci =ð3TÞg
ðA:4Þ
dci is the pure component critical hard sphere diameter, given by:
dci ¼ ð0:08943RT ci =Pci Þ1=3
Required to achieve a raffinate oil with 0.7 wt% of FFA. Considering a single separator tank at 313 K and 50 bar.
olive oil (with different acidity) + CO2 mixtures, and was employed to study the required variation of process conditions to ensure a final acidity value lower than 0.7% in the raffinate. Acknowledgements This work has been financed under a R&D program of the Spanish Ministry of Education and Science, Project AGL-200407227-C02-01 and Project S-0505/AGR/000153 from the Comunidad Autónoma de Madrid. T.F. would like to acknowledge the financial support of the Ramon y Cajal Program from the Ministry of Education and Science. The authors acknowledge the financial support of Migasa (Spain). Appendix
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 component vapor pressure curve, usually the normal boiling point. Since vapor pressure data for low volatile or thermolabile substances is often not available or not reliable, infinite dilution activity coefficients can be used to estimate the dci parameter of high molecular weight compounds, such as alkanes and triglycerides, as demonstrated by Espinosa et al. (2002). For the evaluation of the attractive contribution to the Helmholtz energy, a group contribution version of a densitydependent NRTL-type expression is derived:
, NC NG NG NG X X X zX i A =RT ¼ ni mj qj ðhk g kj q~qÞ hl slj 2 i j i k att
NC X hj ¼ qj =q ni mij
q¼
The residual Helmholtz energy in the GC-EoS model is described by two terms:
sij ¼ exp aij Dg ij q~=ðRTVÞ
A ¼ Afv þ Aatt
Dg ij ¼ g ij g jj
The free volume contribution (Afv) is modeled assuming hard sphere behavior for the molecules, characterizing each substance i by a hard sphere diameter di. A Carnahan–Starling type of hard sphere expression for mixtures is adopted:
Afv =RT ¼ 3ðk1 k2 =k3 ÞðY 1Þ þ k32 =k23 Y þ Y 2 ln Y þ n ln Y ðA:2Þ
ðA:6Þ
where
i
ðA:1Þ
ðA:5Þ
NC X i
ni
NG X
mij qj
j
NG is the number of groups, z is the number of nearest neighbors to any segment (set to 10), mij is the number of groups type j in molecule i, qj is the number of surface segments assigned to ~ is the total number group j, hk is the surface fraction of group k, q of surface segments, gij is the attraction energy parameter for interactions between groups i and j, and aij is the NRTL non-randomness parameter (aij – aji). The interactions between unlike groups are calculated from
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g ij ¼ kij ðg ii g jj Þ1=2
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ðkij ¼ kji Þ
with the following temperature dependencies for the interaction parameters:
g jj ¼ g jj 1 þ g 0jj T=T j 1 þ g 00jj ln T=T j n h . io kij ¼ kij 1 þ kij ln 2T T i þ T j being g jj the interaction parameter at the reference temperature T i . References Bharath, R., Inomata, H., Adschiri, T., Arai, K., 1992. Phase equilibria study for the separation and fractionation of fatty oil components using supercritical carbon dioxide. Fluid Phase Equilibria 81, 307–320. Bondioli, P., Mariani, C., Mossa, E., Fedelli, A., Muller, A., 1992. Lampante olive oil refining with supercritical carbon dioxide. Journal of the American Oil Chemists’ Society 69, 477–480. Bosle, B.M., Subramanian, R., 2005. New approaches in deacidification of edible oils – a review. Journal of Food Engineering 69, 481–494. Brignole, E.A., Andersen, P.M., Fredenslund, A.A., 1987. Supercritical fluid extraction of alcohols from water. Industrial and Engineering Chemistry Research 26, 254– 261. Brunetti, L., Daghetta, A., Fedeli, E., Kikic, I., Zanderighi, L., 1989. Deacidification of olive oils by supercritical carbon dioxide. Journal of the American Oil Chemists’ Society 66, 209–217. Chen, C., Chang, C.J., Yang, P., 2000. Vapor-liquid equilibria of carbon dioxide with linoleic acid, a-tocopherol and triolein at elevated pressures. Fluid Phase Equilibria 175, 107–115. Espinosa, S., Fornari, T., Bottini, S., Brignole, E.A., 2002. Phase equilibria in mixtures of fatty oils and derivatives with near critical fluids using the GC-EoS. Journal of Supercritical Fluids 23, 91–102. European Council Regulation No. 1513/2001. Descriptions and Definitions of Olive Oils and Olive-Pomace Oils. Official Journal of the European Communities 26.7.2001.
Ferreira, O., Fornari, T., Brignole, E.A., Bottini, S.B., 2003. Modelling of association effects in mixtures of carboxylic acids with associating and non-associating components. Latin American Applied Research 33, 217–222. Florusse, L.J., Van’t Hof, A., Fornari, T., Bottini, S., Peters, C.J., 2004. Phase behavior of binary systems of carbon dioxide and certain low-molecular weight triglycerides: measurements and thermodynamic modeling. Journal of Supercritical Fluids 31, 123–132. Hurtado-Benavides, A.M., Señoráns, F.J., Ibáñez, E., Reglero, G., 2004. Mass transfer evaluation of countercurrent supercritical fluid extraction of olive oil. Journal of Supercritical Fluids 28, 29–35. Joback, K.G., Reid, R.C., 1983. Estimation of pure-component properties from group contributions. Chemical Engineering Communications 57, 233–238. Nilsson, W.B., Gauglitz, E.J., Hudson, J.K., 1991. Solubilities of methyl oleate, oleic acid, oleyl glycerols, and oleyl glycerols mixtures in supercritical carbon dioxide. Journal of the American Oil Chemists’ Society 68, 87–91. Pusch, J., Schmelzer, J., 1993. Extension of the group contribution equation of state parameter matrix for the prediction of phase equilibria containing argon, ammonia, propene and other alkenes. Berichte der Bunsen-Gesellschaft fuer Physikalische Chemie 97, 597–603. Simoes, P.C., Brunner, R., 1996. Multicomponent phase equilibria of an extra-virgin olive oil in supercritical carbon dioxide. Journal of Supercritical Fluids 9, 75–81. Skjold-Jørgensen, S., 1984. Gas solubility calculations II. Application of a new groupcontribution equation of state. Fluid Phase Equilibria 16, 317–351. Skjold-Jorgensen, S., 1988. Group contribution equation of state (GC-EOS): a predictive method for phase equilibrium computations over wide ranges of temperatures and pressures up to 30 MPa. Industrial and Engineering Chemistry Research 27, 110–118. Torres, C.F., Vázquez, L., Señoráns, F.J., Reglero, G., 2005. Study of the analysis of alkoxyglycerols and other non-polar lipids by liquid chromatography coupled with evaporative light scattering detector. Journal of Chromatography A 1078, 28–34. Weber, W., Petkov, S., Brunner, G., 1999. Vapour-liquid-equilibria and calculations using the Redlich–Kwong–Aspen-equation of state for tristerarin, tripalmitin and triolein in CO2 and propane. Fluid Phase Equilibria 158–160, 695–706. Yu, Z.R., Rizvi, S.S.H., Zollweg, J.A., 1992. Phase equilibrium of oleic acid methyl oleate and anhydrous milk fat in supercritical CO2. Journal of Supercritical Fluids 5, 114–122. Zou, M., Yu, Z.R., Kashulines, P., Rizvi, S.S.H., Zollweg, J.A., 1990. Fluid-liquid phase equilibria of fatty acids and fatty acid methyl esters in supercritical carbon dioxide. Journal of Supercritical Fluids 3, 23–28.