Separation of fish oil ethyl esters with supercritical carbon dioxide

Journal of Supercritical Fluids 17 (2000) 55–64 www.elsevier.com/locate/supflu

Separation of fish oil ethyl esters with supercritical carbon dioxide V. Riha a, *, G. Brunner b a F. Hoffmann-LaRoche AG, CH-4070 Basel, Switzerland b Technische Universita¨t Hamburg-Harburg, Arbeitsbereich Verfahrenstechnik 2, Eißendorfer Straße 38, D-21071 Hamburg, Germany Received 28 December 1998; received in revised form 8 September 1999; accepted 8 September 1999

Abstract Fatty acid ethyl esters from fish oil (FAEE) were continuously fractionated with supercritical carbon dioxide (SCCO ). A pilot plant countercurrent extraction column with an inner diameter of 68 mm and an effective height of 2 12 m packed with Sulzer CY wire mesh packing was used for the experiments. The operating conditions varied in the temperature range of T=40–80°C and the pressure range of p=6.5–19.5 MPa. All experiments focused on a separation between the low-molecular-weight components (LMC ), with carbon numbers from C14 to C18, and the high-molecular-weight components (HMC ), C20 and C22. This separation represents a possible first step in an enriching process for the commercial interesting components EPA and DHA. HMC concentrations greater than 95 wt% at a yield greater than 95% were achieved with this set-up. The pilot plant was equipped with a sampling system to measure the concentration profile along the column. These experiments allow the height of a theoretical plate to be calculated. Previously published phase equilibrium data [ V. Riha, G. Brunner, Phase equilibrium of fish oil ethyl esters with supercritical carbon dioxide, J. Supercrit. Fluids 15 (1999) 33–50] are employed as base information for these calculations. The methods of McCabe-Thiele and PonchonSavarit (R.E. Treybal, Mass-transfer Operations. Classic Textbook Reissue, 3rd edition, McGraw-Hill, New York, 1987) and a multi-component simulation with the flowsheeting program Aspen+ are used to compute the theoretical number of plates for each experimental run. All methods yield similar results. A height of approximately 0.3 m can be assumed for a theoretical plate with this set-up. © 2000 Elsevier Science B.V. All rights reserved. Keywords: Countercurrent extraction; Fatty acid ethyl esters; Multi-component simulation; Number of theoretical plates; Supercritical CO 2

1. Introduction The positive effects of highly unsaturated fatty acids on the organism have been the subject of * Corresponding author. Tel.: +49-40-42878-3040; fax: +49-40-42878-4072. E-mail address: [email protected] ( V. Riha)

numerous publications. Namely the EPA, C20:5 n-3, and the DHA, C20:6 n-3, are of major interest [3]. Along with this knowledge, an increasing amount of interest for suitable commercial processes to recover polyunsaturated fatty acids (PUFAs) in concentrated forms can be monitored. The conventional methods for fractionation and isolation of these components include vacuum

0896-8446/00/$ - see front matter © 2000 Elsevier Science B.V. All rights reserved. PII: S0 8 9 6 -8 4 4 6 ( 9 9 ) 0 0 03 8 - 8

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Nomenclature a separation coefficienta=K /K i j FAEE fatty acid ethyl esters HETP height equivalent for one theoretical plate HMC high-molecular-weight components, chain length from C20 to C22K , partition coefficient of i component i, K =y /x i i i LMC low-molecular-weight components, chain length from C14 to C18 LMV solvent-to-feed ratio NTU number of theoretical plates SCCO supercritical carbon dioxide 2 v reflux ratio x concentration of component i in liquid phase i y concentration of component i in vapor phase i Y yield distillation, urea crystallization, hexane extraction, or conventional crystallization having the disadvantages of requiring high-temperature processing resulting in degradation or decomposition of the thermally labile compounds or employing flammable or toxic solvents having adverse health effects, respectively [4]. Here, countercurrent extraction with supercritical fluids as solvents, especially carbon dioxide, offers new opportunities for the solution of separation problems [5]. It is well known that only modest enrichments in the EPA or DHA concentration from a triglyceride feedstock can be achieved. The complex triglyceride structure with different chain length ligands makes separation as a function of one specific component impossible. Therefore, the triglycerides are saponified and esterified to fatty acid methyl or ethyl esters. Eisenbach [6 ] demonstrated an EPA enrichment from 14.5 to 48.2 wt% out of FAEE with SCCO . He extracted his feedmaterial 2 in a batch-continuous operation. A batch of the feedmaterial is initially charged to a vessel, and SCCO passes continuously through the vessel. 2 The solvent starts to extract the low-molecularweight components with a short chain length as can be expected from phase equilibrium measurements. An additional column following the batch vessel enhances the fractionation. A temperature or pressure gradient along the column generates an internal reflux that is needed for an improved separation. Basically, a change in density, and therefore in the solubility, primarily causes the

high-molecular-weight components to precipitate. Individual chain-length fractions are collected in a separator as a function of time. Various papers deal with the extraction of EPA and DHA from fatty acid methyl or ethyl esters with SCCO . Most of the papers follow the idea 2 proposed by Eisenbach, e.g. Ref. [7]. In some cases, especially high EPA or DHA concentrations up to 95% were achieved with a previous ureacrystallization step. Since most researchers assume a high demand for EPA and DHA, in the future there will be a trend towards a continuous operating process. In the batch-continuous operation, the solubility of the material in the autoclave changes permanently with the composition. This effect causes a constant change in the internally generated reflux stream. Therefore, the semi-batch data cannot be used for a scale-up of a continuous process. Brunner [5] compared different extraction methods, focusing on scale-up parameters. For the scale-up of a countercurrent extraction, a stationary operation point with an invariant mass flow inside the column is needed. Krukonis [8] designed a continuous countercurrent extraction column with a well-defined reflux ratio on the basis of multi-component phase equilibrium data by Nilsson [9]. He assumed a height of 0.6 m for one theoretical plate. Van Gaver [10] separated various fatty acid methyl esters (FAME ) from seed oils in the pilot plant described in this paper. He achieved in his experiments a separation

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V. Riha, G. Brunner / Journal of Supercritical Fluids 17 (2000) 55–64

between the C16- and C18-FAME components. Van Gaver calculated a height close to 0.3 m for one theoretical plate on the basis of his experimental results.

Packard 5890 Series II gas chromatograph with a 30 m×0.32 mm ID Stabilwax column from Resteck Corp., USA. 2.2. Experimental set-up

2. Experimental 2.1. Materials Two similar fatty acid ethyl ester mixtures derived from sardine oil are used as feed materials. Experiments P1–P6 ( Table 1) refer to one FAEE feed mixture. The tabulated results for experiments P7–P10 belong to the second feed charge. Basically, both feed charges differ in the fraction of low- and high-molecular-weight components. The purity of both materials related to the identified FAEE components is at least 95 wt%. The CO from Carba Gas, Switzerland, had a 2 quality of 4.0 and a purity higher than 99.99 vol.%. Solvents for the chromatographic analysis are of analytical grade and are obtained from Merck, Germany. All samples are analyzed with a Hewlett

The separation experiments were carried out in a pilot-plant column. Fig. 1 shows a simplified process flow sheet of the plant set-up. For clarity, detailed information, e.g. controls and buffers, has been omitted from the figure. The main part of the process is the extraction column (1) with an inner diameter of 68 mm and an effective height of 12 m. The whole column is packed with Sulzer CY wire mesh packing. The feed (2) enters the column in the middle, 5.6 m above the SCCO feed. A metering pump (9) adds 2 the reflux from the separator (5) at the top of the column. Regenerated SCCO enters the column at 2 the bottom and flows upwards. Because of differences in density, the liquid feed and reflux streams flow downward. The liquid downstream comes into contact with the SCCO . SCCO extracts the 2 2 low-molecular-weight components from the liquid

Table 1 Results of separation experiments Run Temperature Pressure Extract C14 C16 C18 C20 C22 Feed C14 C16 C18 C20 C22 Raffinate C14 C16 C18 C20 C22 CO 2 Reflux

(°C ) (MPa) (kg/h) (wt%) (wt%) (wt%) (wt%) (wt%) (kg/h) (wt%) (wt%) (wt%) (wt%) (wt%) (kg/h) (wt%) (wt%) (wt%) (wt%) (wt%) (kg/h) (kg/h)

P1

P2

P3

P4

P5

P6

P7

P8

P9

P10

60 14.0 2.6 10.84 44.33 28.57 14.76 1.50 4.2 6.47 26.91 20.90 24.02 21.71 1.6 0.00 1.16 9.57 37.70 51.57 297.0 8.5

60 14.5 1.5 10.60 43.78 33.96 11.16 0.49 2.3 6.47 26.91 20.90 24.02 21.71 0.8 0.00 0.50 0.47 44.12 54.90 299.0 11.0

60 14.5 1.9 9.58 39.63 30.79 19.61 0.39 2.7 6.47 26.91 20.90 24.02 21.71 0.8 0.00 0.45 0.35 33.18 66.02 299.0 9.7

60 14.5 1.3 11.28 46.92 36.45 5.00 0.35 2.3 6.47 26.91 20.90 24.02 21.71 1.0 0.00 0.00 0.00 49.58 50.42 298.0 11.3

60 14.5 1.3 11.60 48.26 36.57 3.12 0.45 2.4 6.47 26.91 20.90 24.02 21.71 1.1 0.00 0.00 1.15 50.35 48.50 300.0 10.6

60 15.0 2.3 8.50 35.25 27.43 26.56 2.25 2.9 6.47 26.91 20.90 24.02 21.71 0.6 0.00 0.42 0.18 15.94 83.46 298.3 11.4

40 9.6 1.2 15.07 60.69 23.63 0.60 0.00 2.2 8.38 33.73 22.26 21.23 14.40 1.0 0.00 0.00 20.54 47.05 32.42 280.3 10.4

60 14.5 1.6 12.55 50.54 32.89 3.99 0.04 2.4 8.38 33.73 22.26 21.23 14.40 0.8 0.01 0.02 0.94 55.82 43.22 303.0 11.6

70 17.0 1.8 12.65 50.93 32.88 2.10 1.44 2.7 8.38 33.73 22.26 21.23 14.40 0.9 0.00 0.00 1.44 58.75 39.81 269.6 12.7

80 19.5 2.1 13.20 52.01 32.28 2.44 0.07 3.3 8.38 33.73 22.26 21.23 14.40 1.2 0.00 0.00 2.57 56.90 40.53 209.3 14.9

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V. Riha, G. Brunner / Journal of Supercritical Fluids 17 (2000) 55–64

3. Results and discussion

Fig. 1. Principle flowscheme of pilot plant.

stream. The components with a high molecular weight are enriched in the downward flowing liquid stream. The raffinate is discharged at the bottom. A backpressure valve (3) regulates the pressure in the column with an accuracy of ±0.1 MPa. A pressure reduction below the critical pressure of CO and a heater (4) allow the extract stream in 2 the separator (5) to be split into the solvent and the extract. The extract basically consists of the low-molecular-weight components. A condenser (6) controls the pressure in the separator and liquefies the solvent. A pump (7) and a heater (8) close the CO circle and adjust 2 the pressure and temperature of the SCCO . 2 All experiments are carried out under isothermal and isobaric conditions inside the column. The temperature inside the column is controlled within ±1 K via three separate heating jackets. The experimental operation conditions are listed in Table 1. Coriolis-type mass flow meters measure each main stream indicated in Fig. 1. The column is equipped with special devices for gas- respective liquid phase sampling in order to monitor concentration profiles along the column. The first sample position is 1.4 m above the SCCO feed position. The next sampling devices 2 follow each 1.4 m up to a column height of 11.2 m.

This work focuses on a separation between the C18 and C20 components. We assumed from phase equilibrium measurements [1,9] that separation due to the FAEE chain length is possible. Therefore, all components of the mixture were lumped into a set of five pseudocomponents as a function of their chain length. Experimental results for the pilot plant are listed in Table 1. The experiments represent steady-state conditions in the column. A steady state was defined as a constant mass flow and corresponding concentrations in two successive samples from the product streams at a time interval of 1 h. An experimental time of 8 h after the last adjustments proved to be sufficient to obtain both relevant samples. Experiments P1–P6 determine the best operating conditions at a constant temperature of T= 60°C. Phase equilibrium measurements suggest a good solubility of FAEE in SCCO at that temper2 ature. For each experiment, the parameters pressure, SCCO - and feed mass flow are fixed. By 2 varying the reflux mass flow it is possible to control the distribution between both product streams. From an operational point of view, a pressure of p=14.5 MPa at a fixed temperature of T=60°C produced the best results with the first feed charge. This experiments showed the strong impact of SCCO density on the operability. The SCCO 2 2 density varied between r=560 kg/m3 at r= 14 MPa and r=603 kg/m3 at p=15 MPa. In particular, the operating point at p=15 MPa lay close to the flooding point. Small changes in the mass flows at that operating point led to an unstable process with inconsistent samples from the product streams. In this series, the best results with a HMC concentration of x=96 wt% and a yield of Y= 95% were obtained with experiment P5. In the second set of experiments, P7–P10, the influence of temperature on the separation was studied. Therefore, the SCCO density was fixed 2 to r#583 kg/m3 via the combination of temperature and pressure. According to phase equilibrium measurements [1,9], the solubility of FAEE in SCCO increases with temperature. This effect can 2 be monitored with the solvent to feed ratio LMV, Eq. (1). Less solvent is required at higher temper-

V. Riha, G. Brunner / Journal of Supercritical Fluids 17 (2000) 55–64

atures to fractionate a given amount of feed: LMV= v=

m ˙ SCCO

2

m ˙ Feed

m ˙ Reflux m ˙ Extract

.

3.1. Modeling the separation (1) (2)

The reflux ratio, v [Eq. (2)], represents the excess amount of FAEE that must be dissolved in the SCCO . On one hand, the reflux is needed to 2 create a countercurrent flow regime inside the column, and on the other hand, the reflux should be kept at a minimum for economical reasons. Within this set of experiments, the solvent-to-feed ratio decreases from LMV=127 to LMV=63.3, while at the same time, the reflux ratio also decreases from v=8.7 to v=7.0. Experiments P7– P10 give comparable product qualities and yields as run P5. Comparing the results of P8 and P10, the throughput at T=80°C is 1.4 times higher, while at the same time, the SCCO flow is reduced 2 by a factor of 0.7. Fig. 2 shows the dependence of the feed throughput and the amount of FAEE in the extract phase on operating temperature. This figure directly indicates the strategy to optimize the throughput. The increase in the solubility of FAEE in SCCO allows the column to be operated 2 at a reduced reflux ratio as well as at a reduced solvent-to-feed ratio. These effects lead to an economical operation point. Therefore, further experiments should focus on higher temperatures in the extraction column.

Fig. 2. Influence of temperature on throughput.

59

In a first step, the number of theoretical plates are calculated on basis of the established McCabe– Thiele and Ponchon–Savarit methods, Treybal [2]. Both methods treat a pseudo-binary system with respect to the involved FAEE components. In addition to the binary, solvent-free McCabe– Thiele method, the ternary Ponchon–Savarit method also deals with the solvent. The fractionation of a real, multi-component fish oil system into two product streams is dominated by the selectivity for specific components. In our case, the separation efficiency between the C18 and C20 components controls the quality of the complete separation. Phase equilibrium measurements [1] as well as the results in Table 1 show that the degree of difficulty during a separation of FAEE increases with decreasing difference in the chain length. The separation factor [Eq. (3)] defines the separability of two components. In this case, the separation of the LMC from the C20 component, as well as the separation between the C18 and C22 components, should be examined more closely. Table 2 lists the partition coefficients [Eq. (4)] for each component of the first feed charge and the calculated separation factors. For detailed phase equilibrium data, see Ref. [1] (table 7, EE-4). As expected, the minimum separation factor appears at the C18– C20 ratio. The separation factors for the C16–C20 and C18–C22 components are nearly equal. This means that the more volatile C14 and C16 components are easily enriched in the extract fraction, while the high volatile C22 components remain in the liquid raffinate fraction: K y x a= 1 2 = 1 x y K 1 2 2

(3)

y K= i. i x i

(4)

Based on the phase equilibrium observations, the FAEE mixtures are reduced to a set of binary data with the pseudo-components LMC and HMC. For the McCabe–Thiele calculations, two different cases were studied. The first case uses a

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V. Riha, G. Brunner / Journal of Supercritical Fluids 17 (2000) 55–64

Table 2 Partition coefficients and separation factors for the first feed charge on a solvent-free basis T (K)

313.2 313.2 313.2 313.2 313.2 313.2 313.2 313.2 333.2 333.2 333.2 333.2 333.2 333.2 333.2 333.2 343.2 343.2 343.2 343.2 343.2 343.2 343.2 343.2

p (MPa)

9.0 10.0 11.0 11.5 12.0 12.5 13.0 13.5 12.0 13.0 14.0 14.5 15.0 16.0 17.0 18.0 14.0 14.5 15.0 16.0 17.0 18.0 19.0 20.0

Partition coefficient, K (−) i

Separation factor, a (−)

C14

C16

C18

C20

C22

C14/C20

C16/C20

C18/C20

C18/C22

2.1 1.7 1.5 1.4 1.4 1.3 1.2 1.2 2.3 2.0 2.0 1.8 1.8 1.6 1.4 1.3 2.1 2.1 2.1 1.9 1.7 1.6 1.5 1.4

1.4 1.3 1.2 1.2 1.2 1.1 1.1 1.1 1.5 1.3 1.3 1.3 1.3 1.2 1.2 1.1 1.4 1.3 1.4 1.3 1.3 1.2 1.2 1.2

0.9 1.0 1.0 1.0 1.0 1.0 1.0 1.0 0.8 0.9 1.0 1.0 1.0 1.0 1.0 1.0 0.9 0.9 0.9 0.9 1.0 1.0 1.0 1.0

0.6 0.8 0.8 0.9 0.9 0.9 0.9 0.9 0.5 0.7 0.7 0.7 0.7 0.8 0.8 0.9 0.6 0.6 0.6 0.7 0.7 0.8 0.8 0.9

0.4 0.5 0.6 0.7 0.7 0.8 0.8 0.9 0.3 0.4 0.4 0.5 0.5 0.6 0.7 0.8 0.4 0.4 0.4 0.5 0.5 0.6 0.7 0.7

3.4 2.2 1.9 1.7 1.6 1.4 1.3 1.3 4.3 3.0 2.9 2.5 2.4 2.0 1.7 1.5 3.5 3.2 3.4 2.7 2.3 2.0 1.7 1.6

2.2 1.6 1.5 1.4 1.3 1.2 1.2 1.2 2.8 2.0 2.0 1.8 1.7 1.5 1.4 1.3 2.3 2.1 2.2 1.9 1.7 1.6 1.4 1.4

1.4 1.2 1.2 1.1 1.1 1.1 1.1 1.1 1.5 1.4 1.4 1.3 1.3 1.2 1.2 1.1 1.5 1.4 1.5 1.4 1.3 1.3 1.2 1.2

2.4 1.7 1.6 1.4 1.4 1.3 1.2 1.2 2.9 2.1 2.2 1.9 1.9 1.7 1.5 1.3 2.4 2.2 2.4 1.9 1.8 1.7 1.5 1.5

constant separation factor, a=1.33. The fixed value results out of the average of all individual separation factors for the C18–C20 components in Table 2. The second case assumes a dependency of the separation factor on the concentration of the liquid mixture. Fig. 3 shows a plot of the avalues as a function of composition and system pressure. The data points are taken from phase equilibrium measurements [1]. A linear equation [Eq. (5)] with the parameters composition and pressure is sufficient to correlate the data at a defined temperature, e.g. T=60°C: a=2.6221−0.0802p−0.2636x.

(5)

In addition to the solvent-free data used for the McCabe–Thiele method, the Ponchon–Savarit method deals with two FAEE pseudo-components and the solvent distribution between both phases. The phase boundary lines for the Ponchon–Savarit

Fig. 3. Separation factor as a function of composition.

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V. Riha, G. Brunner / Journal of Supercritical Fluids 17 (2000) 55–64 Table 3 Results of ternary temperature-dependent correlation ka ij

kb ij

i, j

j=CO 2

j=LMC

j=HMC

j=CO 2

j=LMC

j=HMC

i=CO 2 i=LMC i=HMC

0.0 0.0 0.0

−0.002 0.0 0.0

0.0 0.0 0.0

0.270 −0.013 −0.103

0.092 0.0 −0.014

0.0 −0.589 0.0

Table 4 Calculated number of theoretical plates Run

P1

P2

P3

P4

P5

P6

P7

P8

P9

P10

McCabe–Thiele, a=1.33 McCabe–Thiele, a=f (x) Ponchon–Savarit

27 29 21

33 31 29

32 27 26

40 40 37

44 48 41

31 25 24

35 21 76

40 44 38

40 49 41

40 58 43

method are calculated using the Peng–Robinson equation of state [11] [Eq. (6)] combined with a temperature-dependent Melhem mixing rule [12] [Eq. (7)]: p=

RT v−b

−

a (T ) m v2+2bv−b2

n n a =∑ ∑ x x a ; m i j ij i=1 j=1

m T 0.45724R2T2 2 c 1+m 1− ; a= i p T c c m=0.37464+1.54226v−0.26992v2

C A S BD

0.0778RT n c. b =∑ x b ; b = m i i i p i=1 c

C

(6)

evaluation of the calculated NTUs yields an average value of NTU=38 with a standard deviation of 11. The height for one theoretical plate, HETP, results from the NTU value and the column height. Fig. 5 plots the HETP for each experiment and all three methods. An average value of HETP=0.3 m represents the separation with this system. No direct influence of temperature or pressure on the HETP value was detected in this set of experiments. In the next step, a multi-component separation was simulated with the flowsheeting program Aspen+. An external routine to calculate the partition coefficients, K , replaced the build in i

D

x i a =Ea a 1−k +(k −k ) ij ii jj ij ij ji x +x i j k =ka +kb T. (7) ij ij ij Table 3 presents the regressed values for the Melhem mixing rule. For detailed information, see Ref. [1]. Interpolation of the experimental separation factors on a solvent inclusive basis over the concentration range analogous to the method described above give the needed tie lines. Fig. 4 shows the construction of the theoretical plates for experiment P8 and the Ponchon–Savarit method. Table 4 lists the calculated number of theoretical plates, NTU, for both presented McCabe–Thiele methods and the Ponchon–Savarit method. The

Fig. 4. Ponchon–Savarit diagram for experiment P8.

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V. Riha, G. Brunner / Journal of Supercritical Fluids 17 (2000) 55–64

Fig. 5. Calculated height of theoretical plates.

Fig. 7. Modular column set-up for the Aspen+ simulation.

posed in a previous paper [1]. Equation (8) is used to model the K values of the FAEE pseudoi components, while a function, as obtained from Eq. (9), sufficiently describes the behavior of SCCO . The parameters for Eqs. (8) and (9) are 2 listed in Table 5:

Fig. 6. Correlation between the coefficients for the pseudo-component vs. molecular weight of the mixture.

K =a +a e(−a3M˜) (8) i 1 2 ˜a . K =a M (9) CO2 1 2 For the simulation, the column was split into a cascade of flash modules. Each module is connected to two other modules in accordance with the theory of a theoretical separation unit, as

equation of state models. This allows the K values i to be fitted to experimental phase equilibrium data. This step improved the prediction of the K i values for the system with five FAEE pseudo components and SCCO . Fig. 6 shows the depen2 dence of the FAEE partition coefficients on the average molecular weight of the mixture, as proTable 5 Fitting functions for the multi-component Aspen+ procedure p=14.0 MPa

C14 C16 C18 C20 C22 CO 2

p=16.0 MPa

a 1

a 2

a 3

AAD

a 1

a 2

a 3

AAD

0.037 0.023 0.017 0.012 0.008 0.130

4.5 E12 5.6 E11 1.2 E13 4.0 E11 1.0 E12 0.382

0.110 0.104 0.115 0.104 0.108

2.3 1.6 1.3 1.0 1.2 9.4

0.071 0.048 0.035 0.028 0.022 0.102

9.9 E10 1.1 E10 1.0 E10 1.6 E10 1.6 E11 0.418

0.092 0.085 0.085 0.087 0.096

9.8 7.7 8.0 7.7 4.3 9.4

E−3 E−3 E−3 E−3 E−3 E−3

E−3 E−3 E−3 E−3 E−3 E−3

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V. Riha, G. Brunner / Journal of Supercritical Fluids 17 (2000) 55–64 Table 6 Results of Aspen+ simulation for experiment P8 and a variation of NTU NTU

P8

10

20

30

40

45

50

60

Reflux ratio Reflux (kg/h) Extract (kg/h) C14 (wt%) C16 (wt%) C18 (wt%) C20 (wt%) C22 (wt%) Raffinate (kg/h) C14 (wt%) C16 (wt%) C18 (wt%) C20 (wt%) C22 (wt%)

7.58 11.61 1.53 12.6 50.1 33.4 3.9 0.0 0.85 0.0 0.0 1.0 55.6 43.4

7.58 11.35 1.50 13.0 50.6 24.5 10.9 1.0 0.88 0.1 3.1 18.8 39.8 38.3

7.58 11.61 1.53 12.8 51.1 28.5 7.5 0.1 0.85 0.0 0.3 11.3 47.1 41.3

7.58 11.69 1.54 12.7 51.0 30.8 5.5 0.0 0.84 0.0 0.0 6.8 51.3 41.8

7.58 11.72 1.55 12.7 50.8 32.2 4.4 0.0 0.83 0.0 0.0 3.9 53.8 42.3

7.58 11.73 1.55 12.6 50.7 32.6 4.0 0.0 0.83 0.0 0.0 3.0 54.6 42.4

7.58 11.77 1.55 12.6 50.5 33.0 3.9 0.0 0.83 0.0 0.0 2.1 55.2 42.6

7.58 11.75 1.55 12.6 50.6 33.5 3.3 0.0 0.83 0.0 0.0 1.3 56.2 42.5

sketched in Fig. 7. The separator at the top of the column is represented by two split modules. A constant split ratio accounts for the reflux. In all calculations, the feed enters the column in the middle. Aspen+ used the specified phase equilibrium conditions to solve the mass balances. Energy balances are neglected in this case. Aspen+ requires the number of theoretical plates as an input value to solve the problem in this method of computation. Therefore, the appropriate number of plates must be determined iteratively. Table 6 and Fig. 8 show the dependency of the calculated extract and raffinate fractions on the number of theoretical plates for one example, P8. Looking at a value of NTU=45, the purities

Fig. 8. Product purity versus number of theoretical plates.

of the measured and calculated extract fractions are in good agreement. To obtain a corresponding result for the raffinate fraction, the number of plates must be increased to a value above NTU= 60. This in turn increases the deviation between the calculated and measured quality of the extract fraction.

Fig. 9. Measured vs. calculated column profile for five pseudocomponents.

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V. Riha, G. Brunner / Journal of Supercritical Fluids 17 (2000) 55–64

For a better comprehension of the separation, the column profile is examined more closely. Fig. 9 compares the measured column profile for experiment P8 and the calculated profile, assuming NTU=45 plates. The calculated profile for this multi-component system represents the measured points over the complete column very well. This confirms the assumption that a model with 45 theoretical plates sufficiently describes the separation in this column with a multi-component FAEE system. The same simulation with an increased number of plates to NTU=60 leads to greater deviations at each sample point. Table 6 displays the results for the experiment P8 and various NTU values in the range of NTU=10–60. A variation of the feed point could lead to a further improvement if required. Since all three methods yield similar numbers of theoretical plates, NTU#40, an average height of HETP=0.3 for one theoretical plate can be assumed. This value, in combination with the good results of the multi-component modeling, represents a basis for the further design of a production plant.

Acknowledgements The support of this work by F. HoffmannLa Roche AG, Basel/Switzerland is gratefully acknowledged. Special thanks are due to Drs. K. Steiner and C. Tiegs of the High Pressure Laboratory ( VFH ).

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