Process optimisation in sunflower oil extraction by supercritical CO2

Chemical Engineering Science 57 (2002) 2753 – 2764

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Process optimisation in sun"ower oil extraction by supercritical CO2 M. Bravi ∗ , R. Bubbico, F. Manna, N. Verdone Dipartimento di Ingegneria Chimica, Universita di Roma “La Sapienza”, via Eudossiana, 18, 00184 Roma, Italy Received 20 July 2001; received in revised form 9 January 2002; accepted 3 April 2002

Abstract A continuous process for the extraction of sun"ower oil using supercritical CO2 , featuring multiple extractors, one oil separator and three cascaded CO2 recovery vessels operating at di5erent pressures, was devised and studied. For every single equipment of the plant making up the process a mathematical model was built. Experimental tests—consisting in measurements of oil solubility in supercritical CO2 —were carried out in a laboratory-scale apparatus to characterise the behaviour of sun"ower oil in the separation from the supercritical "uid. The mathematical model of the whole process was coded in the commercial gPROMS process modelling environment where both its simulation and optimisation—this latter assuming the overall oil production cost as the objective function—were carried out. The process- and economics-related results are discussed and compared with those obtained with traditional and cold-pressing extraction. ? 2002 Elsevier Science Ltd. All rights reserved. Keywords: Supercritical "uid; Extraction; Optimisation; Food processing; Sun"ower seed oil; High-quality food products

1. Introduction In the last years, several studies investigating the supercritical "uid extraction (SFE) of lipidic compounds from a wide range of di5erent vegetal matrices (seeds, "owers, leaves, etc.) have been published. In most of these studies carbon dioxide is the solvent used ◦ because of its relatively low critical temperature (31:1 C), non-toxicity, non-"ammability, good solvent power, ease of removal from the product and low cost. In the following of the present article “solvent” indicates carbon dioxide. The high quality of the products obtained by SFE was pointed out by List, Friederich, and Pominski (1984a, b), who observed that the extracted products do not need any particular reBning operation as the vegetal material does not undergo any stressing treatments. As far as the extraction of sun"ower oil is concerned, the use of supercritical CO2 was tested by Stahl, Schultz, ◦ ◦ and Mangold (1980) at temperatures of 20 C and 40 C and at pressures up to 700 bar. Later, Perrut, Clavier, Poletto, and Reverchon (1997) provided the mathematical model of the extraction and estimated the relevant parameters (mass ∗ Corresponding author. Tel.: +39-6-44-585600; fax: +39-6-44-585451. E-mail address: [email protected] (M. Bravi).

transfer coeGcient, equilibrium relationship, etc.) for an ◦ SFE performed at 280 bar and 40 C, based on pilot-plant experimental data. Calvo, Cocero, and Diez (1994) and Calvo and Cocero (1996), on the other hand, investigated the quality of sun"ower oil extracted by SC-CO2 , the method to guarantee oxidative stability of the extracted product and the use of ethanol to improve the oil yield. The “basic” apparatus carrying out the SFE of edible or pharmaceuticals compounds uses pure carbon dioxide as the solvent, and features a single extractor operating batchwise, a "ash valve and a separator. ModiBed (i) plant layouts and operating (ii) modes and (iii) conditions have been proposed to increase the capacity and eGciency of the extraction operation. Eggers (1996) described a three-stage separator system for a single-extractor corn-germ extraction plant for increased energy recovery and a (patented) continuous extraction unit to avoid Blling and emptying the extraction vessel at atmospheric pressure. Eggers and Sievers (1989) optimised the solvent "ow rate=seed bed weight ratio in the extraction from milled corn germ, showing that the amount of extracted oil per unit residence time passes through a maximum. Goodrum, Kilgo, and Santerre (1996) carried out a two-level, half-factorial design experimental campaign to assess the e5ects of pressure, temperature, moisture, particle size and solvent "ow rate and a pressure–temperature investigation (at constant values of the

0009-2509/02/$ - see front matter ? 2002 Elsevier Science Ltd. All rights reserved. PII: S 0 0 0 9 - 2 5 0 9 ( 0 2 ) 0 0 1 4 5 - 8

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M. Bravi et al. / Chemical Engineering Science 57 (2002) 2753–2764

remaining three quantities), pointing out the complex relationships between pressure, temperature and SFE yield. Brunner and Peter (1982) and Calvo and Cocero (1996) investigated the enhancing action of ethyl alcohol added to the solvent as an entrainer. As far as process optimisation is concerned, Smith, Inomata, Kanno, and Arai (1999) carried out an energy analysis of supercritical extraction cycles and indicated the extractions operating conditions favouring the use of compressor or pump mode, pointing out that this latter is generally preferable when the extraction temperature is low and extraction pressure is high. Gani, Hytoft, and Jaksland (1997) presented a computer aided system for analysing and optimising the operating conditions of a SFE process and demonstrated its use in some selected "uid–liquid extraction cases where solvents other than carbon dioxide are used. Gani et al. (1997) showed that lowering the extraction temperature and pressure and increasing the operating pressure of the product recovery stage leads to a minimisation of the total production cost. Brunner (1998) discussed in detail a roadmap for the development of countercurrent multistage SFE processes, while Espinosa, Diaz, and Brignole (2000) analysed their optimal design taking into consideration the e5ect of temperature gradients and external re"ux in the extraction. Despite the process improvements proposed to reduce the live costs and increase the extraction yield, the current trend is to use SFE only to extract small quantities of high value substances in small-size plants. SFE as an alternative technique to traditional combined pre-pressing and hexane extraction of common oilseeds (sun"ower, corn, soybean, canola, etc.) still awaits a clear demonstration of its proBtability.

The purpose of the present work is twofold: Brst, to devise an industrially feasible process layout and operation scheme that would make a continuous production of sun"ower oil and an eGcient recovery of CO2 possible; second, to identify the optimal operating conditions, from the quantitative point of view, for the devised process and to assess whether SFE is a potential candidate process for the extraction of sun"ower oil for the food market. This latter aim required the set-up of a comprehensive mathematical model of the whole process (extraction section and recovery section). Some missing thermodynamic data necessary to accomplish this latter task were determined by carrying out a suitable experimentation on an SFE pilot plant. 2. Layout of the proposed process The basic batch apparatus carrying out an SFE consists of an extractor, a "ash valve and a separator. The devised process for the continuous production of sun"ower oil from sun"ower seeds (Fig. 1) features, in its general form, multiple batch extractors in parallel, each containing a Bxed bed of seeds, and multiple stages in the CO2 recovery section (D1, D2, D3) operating at such pressures that the ratio of the prevailing pressure in separator S to that in D1 equals the ratio of the prevailing pressure in Di to that in Di+1 , while the prevailing pressure in D3 is kept constant and equal to 10 bar. Incidentally, the prevailing pressure in D2 is also close to that at which carbon dioxide is normally stored (20 bar), so that the solvent make-up (pure CO2 ) was assumed to be made there. The P& I scheme of the plant conBguration reported in Fig. 1 features three extractors and

Fig. 1. Plant scheme (the included control instrumentation and manipulation devices are in Brst-approximation detail).

M. Bravi et al. / Chemical Engineering Science 57 (2002) 2753–2764

corresponds to what will be referred to in the following as the “reference” conBguration. The operation sequence undergone by every batch extractor is: solids loading, pressurisation, extraction with a continuous solvent "ow, depressurisation, solids unloading. Every extractor begins this sequence at a di5erent time, so that there are always two extractors in the extraction phase and a continuous "ow of oil-rich SC-CO2 reaches the "ash valve. Due to the Bxed time which must be devoted to the loading and unloading of the seeds and to the pressurisation and depressurisation of the extractors (which, together, make up the “inactive” phase of the working cycle of an extractor), the number of extraction units featured by the plant is dependent upon the relative duration of the active (extraction) and inactive phases as well as the desired production regularity. The pressure reduction causes the extracted oil to separate in the form of liquid droplets inside the separation vessel (S) and to accumulate at its bottom. This way, if a suGcient number of extractors is installed, a continuous production of oil is accomplished. Solvent is then re-compressed and re-circulated to the extractors pool. During the depressurisation phase the extractor discharges its content into the Brst recovery vessel (D1), which is at the highest pressure. Once the available pressure gap has been exhausted, the extractor is connected to the second recovery vessel (D2) and, eventually, to the third- and last-one (D3); after the last recovery stage, the CO2 still remaining in the extractor is released to atmosphere. For every stage the solvent vapour is re-compressed and recirculated to the stage operating at the nearest higher pressure level or, for the Brst stage, directly to the extractive section, while the liquid is expanded in the stage operating at the nearest lower pressure level. The second expansion stage is Btted with a thermal circuit controlling the liquid level of the third vessel (D3). An oil blowdown line is connected to the expansion vessel operating at the lowest pressure. The adopted three-stage solvent recovery and compression system dramatically reduces (from 6678 down to 1820 kW) the energy requirement of solvent recompression compared to a singlestage one. A pneumatic conveyor system is used to load and unload the seeds into and out of the extractors.

3. Process modelling In order to set up a comprehensive model for the entire process, separate sub-models were written for each piece of equipment and piping item and then connected together. The plant items that were taken into account by modeling were: • • • • •

extractors; expansion valve; separator; compressors; solvent recovery vessels;

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• heat exchangers; • pipe union joints and branches. The last item group has only been modelled as nodes, i.e., equipment that fully obey the mass and energy conservation balances. Pressure drop were not considered in any piece of equipment but the expansion valve. A general hypothesis, that was applied to every piece of equipment and streams, is that the thermodynamic behaviour is established by carbon dioxide only. Oil content being low, the thermodynamic in"uence of this latter was deemed negligible and disregarded. The eight-parameter BWR equation of state (Benedict, Webb, & Rubin, 1942) was used to describe the pressure– temperature–density relationship, to calculate residual functions used in the calculation of thermodynamic properties and to calculate vapour pressure
 C0 P = RT + B0 T − A0 − 2 + (bRT − a)3 T
3 c (1 + 2 ) exp(−2 ); + a 6 + (1) T2  T ∗ Cp ∗ dT − Oh∗ ; h = h0 + (2) T0 T 
4C0 Oh∗ = −B0 RT + 2A0 + 2  T + 12 (−2bRT + 3a)2 −

c 6a 5 − [(22 4 − 2 − 6) 5 2T 2

×exp(−2 ) + 6];  Psat (vliq − vvap ) = −

vvap

vliq

(3) P dv;

Psat = P(T; liq ); Psat = P(T; vap ):

(4)

The adopted BWR parameters for carbon dioxide were derived from the Chemical Engineers’ Handbook (Perry & Green, 1984). The BWR equation of state was chosen over more recent and accurate ones such as the 42-parameter Span and Wagner one (Span & Wagner, 1996) or the 20-parameter Bender one (Bender, 1975) as a good trade-o5 between accuracy and computational weight. The BWR parameters for carbon dioxide were adopted from Perry and Green (1984). 3.1. Extractors (E) The mathematical model adopted for the extraction phase is that proposed by Perrut et al. (1997) and is based on a

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di5erential mass balance. Following Bulley, Fattori, Meisen, and Moyls (1984), Lee, Bulley, Fattori, and Meisen (1986) and Fattori, Bulley, and Meisen (1984) mass transfer resistance was assumed to occurr only in the solvent phase. Another important hypothesis is that, during SFE, oil can be well described by only one component (called the solute). Moreover, the enthalpy variations of the system during the extraction were deemed negligible and, therefore, were not considered. Summing up, mass balance equation of the solute for the "uid phase is 
@y @y f  + vi = J; (5) @t @z while, for the solid phase, it is s (1 − )

@x = −J; @t

(6)

where y is the solute concentration in the "uid phase, x is the solute concentration in the solid phase, vi is the interstitial velocity uniformly distributed in every section of the extractor (plug-"ow of SC-CO2 ), f is the "uid density, which is supposed not to be a5ected by the presence of the solute, s is the bulk density of the insoluble solid,  is the void fraction of the seed bed and J is the solute exchange rate between the phases. x and y only depend on time (t) and on the axial co-ordinate (z). According to Perrut et al. (1997), the solute mass transfer rate can be expressed as J = ap hf (y∗ − y);

(7)

where ap is the speciBc surface for mass exchange between the phases, h is the mass transfer coeGcient, and y∗ is the solute concentration at the solid–"uid interface. As far as the equilibrium relationship—which has a roughly sigmoidal shape—is concerned, the simpliBed model relationship proposed by Perrut et al. (1997) states that x ¡ x; S

y = Kx;

x ¿ x; S

y = y0 :

(8)

This expression clearly shows that an oil concentration in the bed below xS determines unfavourable mass transfer conditions. The extraction operating conditions chosen by Perrut et al. ◦ (1997) (40 C and 280 bar) were adopted for the extractors pool of the devised process. Consequently, all the relevant model parameters (K = 0:0036; y0 = 0:011; xS = 0:33; ap = 2000) were inherited. Once the operating parameters and the initial and boundary conditions have been assigned, the mathematical system constituted by Eqs. (5) – (8) can be solved, and the solutions yield the values of x; y and y∗ . At the end of the extraction phase, each extractor is isolated and the depressurisation phase begins. The mass and energy balances for the extractor during the discharge phase

can therefore be written as V

@ = −Fout ; @t

V

@(u) = −Fout hout : @t

(9) (10)

V being the volume of one extraction vessel,  and u the density and the internal energy of CO2 , fout and hout the "ow rate and the enthalpy of the outlet stream. Eqs. (1) – (3) were used to make the pressure–temperature– density and enthalpy–temperature–density relationships explicit, so that the mathematical system constituted by Eqs. (9) – (10) can be solved. The number of the real solutions obtained for the density (one root meaning that the system is in single-phase supercritical, liquid, vapour or gas state; two roots meaning that the system is inside the two-phase liquid–vapour bell-shaped zone) together with the current pressure–density state of the system relative to the critical point was used to determine the state of the CO2 -oil system. The state of the extractor outlet stream depends on the state of the content of the extractor: gas (or vapour) if only gas (or vapour) is present, liquid otherwise, with the heavier phase leaving the vessel Brst. Again, the system can be solved once the initial conditions and the outlet pressure have been set. 3.2. Expansion valve (V ) The outlet stream coming from the extractors goes through the expansion valve. The mass and energy balances for the valve are expressed by the following equations: Fin = in vin Sin = out vout Sout = Fout ;

(11)

2 2 = hout + 12 vout : hin + 12 vin

(12)

The system is completely deBned including Equation set (1) – (4) and can be solved for the values of the outlet variables once the outlet pressure has been assigned. 3.3. Compressor (C) The mass and energy balances for the compressor can be written as Fin = in vin Sin = out vout Sout = Fout ;

(13)

2 2 + L = hout + 12 vout : hin + 12 vin

(14)

From Eqs. (13) and (14), assuming Sin =Sout , the work done can be expressed by the following equation:

2  2  1 Fout out L= 1− + (hout − hin ): (15) 2 out Sout in

M. Bravi et al. / Chemical Engineering Science 57 (2002) 2753–2764

It is also possible to calculate the required electrical power: W=

Fin L ; &

(16)

where & is the overall eGciency of the compressor-electric motor group. The system can be solved once the outlet pressure is set; the outlet pressure is variable along the optimisation calculation and dependent on the operating pressure assigned in the separator for all the compression stages but the Bnal CO2 compression stage, where it is constant and equal to 280 bar. 3.4. Heat exchangers (H ) The mass and energy balances for the heat exchangers are

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3.7. Separator (S) A mass balance under steady-state conditions is suGcient to completely describe the settling of oil from solvent occurring in the separator: Fin = Fout ;

(26)

(Fin MCO2 )yin = (Fout MCO2 )yout + woil :

(27)

The residual solvent content in the recovered oil is neglected. The weight fraction of solute in the solvent leaving the separator must be known to estimate the amount of oil that can be recovered. Experimental tests were carried out to evaluate the solubility of oil in SC-CO2 at the prevailing temperature and pressure in the separator.

Fin = Fout ;

(17)

4. Experimental

Q = Fin hin − Fout hout = wc Cp; c OTc ;

(18)

If an SFE of sun"ower oil were carried out at 280 bar ◦ and 40 C as in the devised process, the expansion pressure for the recovery of oil (that is, the quantity to be optimised in this work) should be chosen in the range 100 –220 bar, given that the solubility of oil in CO2 is almost negligible below 100 bar, and that only a very limited amount of oil can be recovered by expanding the solvent above 220 bar. Solubility data for the CO2 -sun"ower oil system are available in the literature (Stahl et al., 1980). However, they do not cover the range of interest of this work. Therefore, some experimental runs were carried out in a pilot plant to derive solubility data for sun"ower oil in SC-CO2 in the particular range of interest, that is 100 –220 bar. Each experimental determination was carried out at the temperature calculated for the adiabatic "ash at the pressure value under concern. The procedure used consisted in dissolving sun"ower oil in SC-CO2 and recovering it by expanding the supercritical phase at atmospheric pressure. Weighing the recovered oil and dividing the obtained Bgure by the weight of the "owed carbon dioxide, the solubility at the prevailing pressure and temperature inside the extractor (expressed in kg of solute per kg of solvent) was calculated. This procedure is based on the following simplifying hypotheses:

where the outlet temperature of the hot stream is equal to the operating temperature in the extractors, i.e., a parameter. 3.5. Solvent recovery vessels (D) The recovery vessels have, in the most general case, four inlet (extractor, liquid from the upstream recovery vessel, compressed CO2 from the downstream recovery vessel and make-up CO2 ) and two outlet nozzles (vapour and liquid CO2 ), so that the balances can be written as V

@ = Fin1 + Fin2 + Fin3 + Fin4 − Fout1 − Fout2 ; @t

(19)

V

@(u) = Fin1 hin1 + Fin2 hin2 + Fin3 hin3 + Fin4 hin4 @t

(20)

− Fout1 hout1 − Fout2 hout2 :

(21)

Adding, again, Equation set 1– 4 and specifying the initial conditions still leaves two variables to be Bxed, e.g., pressure and liquid volume fraction. 3.6. Pipe union joints and branches • Three-inlet stream union joints: Fin1 + Fin2 + Fin3 = Fout ;

(22)

Fin1 hin1 + Fin2 hin2 + Fin3 hin3 = Fout hout :

(23)

• Branches having three outlet streams: Fin = Fout1 + Fout2 + Fout3 ;

(24)

hin = hout; i :

(25)

(1) oil is considered to be constituted by only one pseudo-component. This assumption co-implies that the same equilibrium relationship can be adopted for all the components contained therein and that they all share the same transport kinetics, as already stated in the section concerning the mathematical model of the extractor. Therefore, composition is not a function of either time or the operating conditions; (2) the carbon dioxide out"owing from the extractor is saturated and the amount of dissolved oil only depends on the prevailing pressure and temperature in the extractor itself; (3) solubility of oil in carbon dioxide and of carbon dioxide in oil is negligible at atmospheric pressure, so that

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M. Bravi et al. / Chemical Engineering Science 57 (2002) 2753–2764

Fig. 2. Scheme of the laboratory plant used for the sun"ower oil solubility experimental determinations. Cy: CO2 cylinder; C: air-driven CO2 compressor; D: heating coil; E: extractor; F: CO2 totaliser; H: oil drain; S: separator; T: thermostatic coil. Table 1 Values of the calculated parameters for the empirical correlation of oil solubility in SC-CO2 vs. temperature and density reported in Eq. (28) Parameter

Value

C0 C1 C2 C3

92.9144 −8:1226 × 104 1:1725 × 107 14:0422

r2

foundations set by Chrastil (1982), was used: C1 C2 (28) + 2 + C3 ln : ln y = C0 + T T The parameters C0 : : : C3 appearing in Eq. (28) were calculated by multiple regression by Btting the results of the experimental determinations performed. Their values, together with the correlation coeGcient of the performed regression, are reported in Table 1.

0.9896

5. Optimisation results and discussion the amount of oil recovered after expansion equals that dissolved in the solvent before the expansion. This justiBes the expression of weight solubility used. The laboratory apparatus used to carry out the experimental measurements is represented in Fig. 2. The extractor has the volume of 0:2 × 10−3 m3 . The "ow rate of carbon dioxide was adjusted in the range 0.15 –0:24 m3 h−1 . Glass spheres (with diameter in the range 3–5 mm) were introduced in the extractor, in order to avoid that CO2 bubbled through the liquid without allowing a suGcient contact to reach saturation. The packing height was adjusted to half of the total internal length of the extractor in order to avoid entrainment of the oil phase. For the same purpose, a disk of porous steel was placed at the exit of the extractor, located on the top face. The adopted devices and the high residence times, compared to those used by Bulley et al. (1984), Lee et al. (1986) and Perrut et al. (1997) completely justify the hypothesis of saturation in the outlet stream. In the literature di5erent types of relationship are available to correlate the solubility data of oils and fats in SC-CO2 with pressure and temperature or density and temperature. Here, the functional form of the relationship proposed by del Valle and Aquilera (1988), derived from the theoretical

The oil content of the sun"ower seeds was assumed to be 0:492 kg kg−1 , as in Perrut et al. (1997). The steps followed in the optimisation of the plant conBguration and operating conditions were: (1) Design of a process set-up considering the “classical” approach to SFE; (2) Quantitative optimisation of the plant layout and operating conditions. The performance criterion (objective function) adopted for the optimisation was the unit oil production cost, expressed in kg−1 . This latter Bgure was estimated by rigorously accounting the operating costs directly referenced by the process design (e.g., compression costs and duty requirements) and applying short-cut techniques for all the remaining operating costs (e.g., manpower) and for the equipment cost estimates (by the cost index method). As far as this latter cost item is concerned, the criteria adopted in the choice of the extractor diameter and of its height=diameter ratio include the adopted limit for the shell thickness (70 mm) and take into consideration the ease of the loading=unloading operations; these criteria are further discussed by Caputo (1997). The extractor was designed according to the required capacity of the reference plant and the resulting size was maintained in all

5.1. Reference case The base, reference plant conBguration was set up by simultaneously addressing the following two problems: (i) extracting, and then recovering, as much oil as possible from the seed bed and (ii) ensuring a continuous operation of the plant itself. For such a high-quality (and, as it can be anticipated, fairly costly) product a production capacity of the same order of that found in the production plants using hexane as the solvent (e.g., thousands of kg h−1 ) is exceedingly high: a capacity an order of magnitude lower is probably appropriate; the target production capacity of the reference plant was therefore set at the (arbitrarily assigned) level of 350 kgoil h−1 . After the necessary design round-o5s, the actual plant capacity was found (by simulation) to be 345 kg h−1 . The Brst 20 min of the time proBle of the yield of a single extraction unit, considered independently of any other piece of equipment, were simulated to calculate the expected instantaneous oil yield of an operating extractor as a function of the time past since the beginning of the SC-CO2 "ow (that is, after recompression, which is carried out in the absence of outlet "ow from the extractor, has been completed). As shown in Fig. 3, 20 min (1200 s) allow the extraction of most of the oil that SC-CO2 can dissolve under the dic-

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0.45 0.40 0.35 0.30 0.25

0

200

400

600

800

1000

1200

Time (s) Fig. 3. Average oil content inside the extractor as a function of time.

SC phase composition (kg/kg)

the subsequent optimisation calculations. The adopted solvent "ow rate (30 000 kg h−1 ), on the other hand, ensures the accomplishment of the desired oil recovery rate and provides a residence time (1:3 min) of the same order of that used by Perrut et al. (1997). A detailed description of the applied costing procedure is reported in Appendix A. Due to the intrinsic approximation of the adopted costing methods, the resulting Bgure must be considered approximate as well. The aim of the optimisation can, therefore, be synthetically restated as the identiBcation of the minimum of the function C = -(te ; Ps ), where C is the unit oil production cost, te is the time allowed for the extraction phase on each seed batch and Ps is the prevailing pressure in the oil separator. The following quantities were kept constant during the optimisation: the size of each extraction vessel, the number of simultaneously "owed extraction vessels and the circulating solvent "ow rate (and, therefore, its residence time inside the extraction vessels). A further simplifying assumption adopted was that the extractor loading and unloading facility only served one extractor at a time; this assumption is important in that it permits to exclude this facility from the list of the “main plant items” to which the direct costing procedure is applied. For the numerical integration and the visualisation of the results the commercial, general-purpose process simulation package gPROMS (Process Systems Enterprise, UK) was used.

Average residual oil (kg oil / kg raw seed)

M. Bravi et al. / Chemical Engineering Science 57 (2002) 2753–2764

0.012 0.011 0.010 0.009 0.008 0.007 0.006 0.005 0.004 0.003 0.002 0.001 0

200

400

600

800

1000

1200

Time (s)

Fig. 4. Oil content in the supercritical phase leaving from a single extractor during a single extraction phase.

tate of the relevant equilibrium parameters; shorter extraction times still permit a good oil recovery but, as it will be explained later, entail a lower utilisation of the plant capacity. Therefore, a 20-min extraction duration was assumed for the base case. The time proBle of the oil content in the SC-CO2 stream leaving a single extractor during each extraction phase is shown in Fig. 4 which clearly shows that, after an initial period (the Brst 7 min—about 400 s—of the extraction) where the oil recovery yield is almost constant, this quantity rapidly decays and reduces to very low values after 13 min, i.e., about 800 s. A single-extractor plant cannot produce oil continuously. Under the hypothesis that the loading and unloading of the solids and the decompression and recompression of an extraction unit altogether require 10 min, as estimated by Caputo (1997), a minimum of two independent (i.e., operated separately) extraction units are required to attain continuous operation. In the following, we will denote these independent extraction units simply as “extractors”. However, due to the large variation of the oil recovery yield during an extraction phase (see Fig. 4), a two-extractor plant would feature a strongly irregular oil recovery rate. In order to reduce the "uctuations in the oil recovery the number of simultaneously "owed extractors must be at least two. This minimum number was adopted and maintained in all

M. Bravi et al. / Chemical Engineering Science 57 (2002) 2753–2764

Time (min)

10

Extractor 1

Extraction

20

30

40

50

60

Table 3 Operating conditions of the industrial plant in the Brst simulation (three extractors)

70

DULP

Extractor 2 Extractor 3

Fig. 5. Operating schedule of the reference case. Symbols meaning: DULP: Depressurisation, solid unloading, solid loading and pressurisation. Table 2 Main characteristics and operating conditions of the industrial plant

60 000 kg h−1 (a) 100 bar 20 min 2190 kg h−1 45, 20 and 10 bar

Total "ow rate of SC-CO2 Separation pressure Extraction time Seeds treated Pressure of the three recovery stages (a) 30000 kg h−1 per extractor

◦

Pressure and temperature in the extractor

280 bar, 40 C

Volume available to seeds per extractor Extractor height Extractor internal diameter Seeds bed height Seeds load per extractor Solids loading=unloading+ extractor pressurisation=depressurisation

1:13 m3 4:8 m 600 mm 4m 365 kg 10 min

0.20 0.18 0.16 0.14

Foil (kg/s)

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0.12 0.10 0.08 0.06 0.04

the subsequent optimisation calculations, thus bringing the total number of extractors of the plant, in the reference case, to three while, for the general case, the minimum number of extractors is given by   t e + tr ne = 2 ; (29) te where ne is the number of extractors, te is the extraction time and tr is the total vessel depressurizing=solids unloading, solids loading=vessel pressurizing (DULP in Fig. 5) time. The meaning of the given relationship is that ne is equal to the argument if this latter is an integer number or is equal to the smaller integer greater than the argument otherwise. The operation sequence of such a three-extractor plant where, at any time, two extractors are in the extraction phase while the third one is in the discharge=recharge phase, is shown in Fig. 5. Allowing more than two simultaneously in the extraction phase, which would further reduce the oscillations in the oil recovery rate, would require either that the loading=unloading facility be able to serve more than one extractor at a time or that more that one loading facility be provided; based on the assumptions made, this case was excluded from the present analysis. The main characteristics of the simulated plant conBgurations and the operating conditions that were not changed in the simulations are shown in Table 2, while Table 3 shows the other operating conditions used for the simulation of the reference case. The plant operation in its reference conBguration was simulated in order to observe the time proBle of the operating conditions of the three extractors from the plant startup to the establishment of its steady-state behaviour. This gave sensitivity concerning the role played by the adjustable parameters PS and te . Fig. 6 shows the oil "ow rate leaving the separator. The strong deviation from the periodic behaviour of the oil recovery rate which can be observed during the

0.02 0.00 0

1000

2000

3000

4000

5000

6000

Time (s)

Fig. 6. Instantaneous rate of the oil out"owing from the separator (Foil ).

Brst 16 min (1000 s) of operation, compared to the regular periodic behaviour reported in the Bnal part of the same curve, is due to the fact that the plant is assumed to be using pure (oil-free) solvent at startup. The collapse observed after the Brst 10 min of extraction in Fig. 4 is the primary cause of the observed periodicity of the oil recovery rate. The calculated oil production cost for the base case was found to be equal to 1:55 kg−1 . 5.2. Plant optimisation A reduction of the required power expenditure for solvent recompression (from the operating pressure of the Brst separator to 280 bar) can be obtained by increasing the operating pressure of the former; doing so also entails a smaller oil recovery. This is the object of the Brst one-dimensional minimum-cost search of the adopted optimisation procedure, the second being the duration of the extraction phases. Incidentally, the adopted base-case pressure in the separator of 100 bar was also found to be the optimal recovery pressure for the relevant plant conBguration. Based on the results plotted in Fig. 4 it can be seen that, in any extraction phase, productivity collapses after the Brst period in which it is almost constant. Fig. 7 shows that, after 10 min, most of the seed bed has a very low oil concentration, which creates unfavourable extraction conditions and causes a reduction of productivity: after 20 min, the oil recovery rate is essentially zero. From an economic point of view this means that product return decreases while the operating cost remains unchanged, thus leading to an increase of the average oil production cost.

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Table 4 Operating conditions in the simulation of a 4-extractor plant carried out with the optimal values found for the extraction time and pressure in the separator

0.50

0.40 0.35 0.30 0.25 0

1

2

3

4

z (m) t=0s t = 600 s t = 1200 s

Total "ow rate of SC-CO2 Separation pressure Extraction time Pressure of the three recovery stages Seeds treated (a) 30000 kg h−1 per extractor

60; 000 kg h−1 (a) 167 bar 10 min 65, 26 and 10 bar 4380 kg h−1

t = 300 s t = 900 s

Fig. 7. Residual oil content (x) along the seed bed (axial coordinate z) of an extractor during a single extraction phase (Reference plant conBguration).

On the other hand, at the end of the 20-min extraction cycle the seed bed still contains an oil residue (0:27 kg oil per kg of non-soluble solid on average). Consequently, if re-circulated solvent (still containing some solute, the actual amount depending on the adopted operating mode) rather than pure solvent is used as the extraction medium, the seeds cannot be completely exhausted. Making the total extraction time shorter would increase the oil production rate, even if this would occur at the expense of a lower oil recovery from the seed matrix. As Fig. 3 clearly shows, if the extraction were stopped after the Brst 10 min, we would still be inside the limits in which the extraction conditions are favourable; only a negligible amount of oil could be extracted in the following 10 min. If the extraction time were doubled (i.e., allowing the previously quoted reference time of 20 min) the average residual oil content of the seeds would change from 0:312 kgoil kg−1 insoluble matter to the previously quoted Bgure of 0.27; if the extraction time were decreased to 6:67 min (6 min 40 s), the average residual oil content would increase to 0.40. Intermediate extraction times (i.e., di5erent from 6.67, 10 or 20 min), although technically possible, would involve a lower utilisation of the plant (the extractors would not be kept busy all the time). Therefore, once the criterion of partially overlapping two subsequently starting extraction phases has been adopted, an extraction time of 20, 10 or 6:67 min makes the best possible use of a plant with 3, 4, or 5 extractors, given that the 10-min requirement for extractor loading, unloading, pressurisation and depressurisation is Bxed. If a more complete exhaustion of the seeds is desired, a further extraction step with hexane can be performed, at the same or at another site. The oil recovered from this further extraction step should be marketed separately from the oil extracted by SC-CO2 . For a 4-extractor plant (each carrying out the exhaustion of the seeds bed for 10 min) with an oil separator operating at 100 bar, an unit oil production cost of 0:78 kg−1 was found. After optimising the operating pressure with

Cost ( / kgoil)

x (kg/kg)

0.45

0.20

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1.2 1.1 1 0.9 0.8 0.7 0.6 0.5 0.4 0.3 0.2 80

100

120 140 160 180 200 220 Expansion pressure (bar) Capital costs Overall costs Operating costs

1.2 1.1 1 0.9 0.8 0.7 0.6 0.5 0.4 0.3 0.2 240

Fig. 8. Capital, operating and total costs as a function of the prevailing pressure in the separator for a four-extractor plant.

this plant conBguration, a unit oil production cost equal to 0:67 kg−1 was found. Increasing the number of extractors to 5 (and reducing, correspondingly, the extraction time to 6:67 min) the minimum unit oil production cost was found to be 0:79 kg−1 at 165 bar. This clearly shows that a four-extractor plant (operating with a 10-min solids exhaustion time) with a pressure at the separator of 167 bar features the lowest unit oil production cost (Table 4). The proBle of the capital cost contribution, of the operating cost and of the total production cost vs. pressure in the oil separator for the 4-extractor plant conBguration is given in Fig. 8, where it can be seen that the production cost shows a limited (12%) increase with pressure below the optimal pressure and above 100 bar. The main geometric (heat exchange surface) and operational (compression power, oil production rate and CO2 make-up) characteristic Bgures of the reference and optimised plants are reported in Table 5. Compared to the Bgures obtained before the optimisation was applied, the power required by the extraction section only is reduced by more than 50% (669 kW instead of 1400 kW, as reported in Table 5), while the required heat exchange surface is reduced by 35%. The production rate increases reaching 553 kg h−1 (against the pre-optimisation value of 345 kg h−1 ). In this condition the amount of oil leaving the separator together with (i.e., dissolved in) the supercritical CO2 and recirculated to the extractors is more than an order of magnitude greater than the value obtained for the reference case.

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Table 5 Main geometric and operational parameters of the reference (3-extractor, 100 bar in the separator) and optimised (4-extractor, 167 bar in the separator) plants

Item

Unit

Reference plant

Optimised plant

Compression power req. (total) Compression power req. (extraction section) Exchanger surface Required CO2 make-up Average residual oil content in the seed bed Oil production rate Oil content in the CO2 at the extractor inlet

kW kW m2 kg h−1 — kg h−1 —

1820 1400 120 324 0.270 345 3:9 × 10−4

1244 669 76 628 0.346 553 3:4 × 10−3

However, it would not be fair to compare the absolute values of the calculated production cost to that of the traditional process which is based on extraction by hexane, as this latter product has a much lower quality and does not share the same prospective market. Indeed SFE, combined with a cultivation of sun"ower under strict biological control could give a product of better quality. This should satisfy the demands of many consumers that exclusively desire “natural” foods and are willing to pay a premium price for a premium food. In such context, SFE-extracted sun"ower oil is a competitor of cold pressing-extracted oil, which is normally available at retail prices around 5 kg−1 . Therefore SFE-extracted sun"ower oil would have its own clear market-although, at least today, still a niche one. Although the presented results suggest that SFE-extracted sun"ower oil has a market future, further information should be gathered concerning the transport phenomena at varying operating conditions, the quality of the oil recovered at high pressure (150 bar or more) and the proBle of oil quality during the extraction. Further improvements could be obtained if a counter-current "ow between seeds and SC-CO2 were realised.

extraction unit placed downstream of the SFE unit and the product extracted therefrom could be marketed separately. The results of the optimisation yielded the optimal values of the operating conditions and the optimal operating schedule for a given "ow-rate of supercritical carbon dioxide; for the proposed parametric process set-up, a four-extractor conBguration, with the extraction phase of every extractor lasting 10 min, is most proBtable. The single most important operating condition is the prevailing pressure in the Brst oil-solvent separation vessel which was calculated to be 167 bar. Correspondingly, the production cost was estimated to be 0:67 kg−1 , a Bgure that, although high, is still acceptable for a product destined to the ever-growing market of high-quality food products. Appendix A The approximate costing procedure adopted in this work (after Cena, Damiani, & Giorgi, 1987) is outlined in the Table 6. Notation

6. Conclusions In order to establish whether SC-CO2 o5ers a viable technique to extract sun"ower oil from sun"ower seeds a parametric process layout was devised and a comprehensive mathematical model of it was built. The simulation of the behaviour of the modelled oil extraction plant with three, four and Bve extractors was carried out and the process operation was optimised on all of them. The most important conclusion drawn from the simulation phase is that an energy-wise implementation of SFE (which requires operating the recovery stage at a high pressure) cannot be expected to exhaust the oilseeds due to the incomplete separation of oil from the solvent and to the early interruption of the extraction phase. However, this does not necessarily involve an oil waste; indeed, if the SFE plant were installed within an existing traditional oil seed extraction plant, the residual seed panel could be treated in a hexane

B0 ; A0 ; C0 ; a; b; c; ;  BWR parameters C0; 1; 2; 3 solubility equation parameter C heat capacity, J kg−1 K −1 ∗ C ideal gas heat capacity, J kmol−1 K −1 C unit oil production cost, kg−1 F mass "ow rate, kmol s−1 h enthalpy, J kmol−1 h∗0 ideal gas enthalpy at T0 , J kg−1 ∗ Oh residual enthalpy, J kmol−1 K −1 J solute "ux per unit volume, kg m−3 s−1 K equilibrium constant, dimensionless L work, J kg−1 M molecular weight, dimensionless P pressure, Pa Q heat exchanged, J R universal gas constant, Pa m3 kmol−1 K −1

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Table 6

Capital cost Cost of the main pieces of equipment Extractors weight[kg] × inox steel price [ kg−1 ] exchanger surface reference cost × (exchanger surface [m2 ])0:6 Reference cost per kw × (compressor power [kw])0:84 COST 1 [ ] 20% COST 1 COST 1b = 1:2 × COST1[ ] 20% COST 1b 15% COST 1b

Extractors Heat exchangers Compressors Solid handling equipment Pipes and valves Instrumentation Buildings and site development Insulation and paintings

10% COST 1b 5% COST 1b COST 2 = 1:5 × COST 1b [ ] 45% COST 2 10% COST 2

Plant construction Engineering Transportation and construction tools

20% COST 2 COST 3 = 1:75 × COST 2 [ ]

Capital recovery Capital recovery factor n i Working hours per year Capital-invested hourly cost

i(1+i)n (1+i)n −1

Fixed-capital-return period (years) Interest rate H(h) [ h−1 ] COST4 = (COST3∗R) H

Operating cost Hourly cost of energy requirement

Compressors power [kw] × price kWh [ kWh−1 ] CO2 make-up "owrate [kg h−1 ] × CO2 price [ kg−1 ] COST 5 [ h−1 ]

Hourly cost of CO2 make-up

Final production cost Intermediate production cost

COST 6 [ kg−1 ] = (COST 4 + COST 5) × Oil production [kg h−1 ] 20% COST 6 15% COST 6 CTot = 1:35 × COST 6 [ kg−1 ]

Maintenance cost Raw material cost Overall oil production cost

r2 S T t u v V Vˆ x xS x0

correlation coeGcient, dimensionless cross section, m2 temperature, K time, s internal energy, J kmol−1 K −1 velocity, m s−1 volume, m3 speciBc volume, m3 kmol−1 solute concentration in the solid phase, kgsolute =kgsolid solute concentration in the solid phase controlling the transition in the equilibrium curve, kgsolute =kgsolid solute concentration in the solid phase at the beginning of the extraction, kgsolute kgsolid

y y0 y∗ W w z

solute concentration in the solvent, kgsolute kg−1 solvent solute concentration in the solvent at the beginning of the extraction, kgsolute kg−1 solvent solute concentration in the solvent at equilibrium, kgsolute kg−1 solvent power, W "ow rate, kg s−1 axial coordinate in the extractor, m

Greek letters O  & 

di5erence void fraction of the seed bed, dimensionless eGciency, dimensionless density, kmol m−3

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f s

M. Bravi et al. / Chemical Engineering Science 57 (2002) 2753–2764

solvent density, kg m−3 nonsoluble solids density, kg m−3

Subscripts c i in, out p S sat oil

coolant interstitial inlet, Outlet constant pressure separator saturation oil

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