Fluidos Super Críticos_Aplicaciones en Alimentos

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Journal of Food Engineering 67 (2005) 21–33

Supercritical fluids: technology and application to food processing

Gerd Brunner *

Thermal Separation Processes, Technische Universitat Hamburg-Harburg, Eissendorfer Str. 38, D 21073 Hamburg, Germany

Received 10 October 2003; accepted 1 May 2004

Abstract

Supercritical fluids (SCFs) are substances at pressures and temperatures above their critical values. It is characteristic that prop-

erties of SCFs can be changed in a wide range. Their solvent power is the highest for non-polar or slightly polar components and

decreases with increasing molecular weight. They can easily be removed from the solutes by mere expansion to ambient pressure.

Carbon dioxide (CO2) is particularly advantageous for processing food materials. SCFs are used for batch extractions of solids, for

multi-stage counter-current separation (fractionation) of liquids, and for adsorptive and chromatographic separations. State of the

art design for commercial plants is available, and a number of installed plants are working. Special applications to food processing

include decaffeination of green coffee beans, production of hops extracts, recovery of aromas and flavours from herbs and spices,

extraction and fractionation of edible oils, and removal of contaminants, among others. The application of SCFs is now extended to

new areas like formulation or specific chemical reactions. Costs of SCF extraction (SCFE) processes are competitive. In certain cases

SCFE processing is the only way to meet product specifications.

� 2004 Published by Elsevier Ltd.

Keywords: Supercritical fluid; Carbon dioxide; Food processing; Extraction; Separation

1. Introduction: supercritical fluids and their solvent power

A pure component is considered to be in a supercriti-

cal state if its temperature and its pressure are higher

than the critical values (Tc and pc, respectively). At crit-

ical conditions for pressure and temperature, there is no

sudden change of component properties. The variation

of properties with conditions of state is monotonous,when crossing critical conditions, as indicated in Fig. 1

by the hatched lines. Yet the magnitude of the variation

can be tremendous, thereby causing different effects on

solutes and reactants within neighbouring conditions

of state. Similar effects to that of the supercritical state

can in some cases be achieved at near critical tempera-

tures in the liquid state of a substance for p > pc and

T < Tc.

0260-8774/$ - see front matter � 2004 Published by Elsevier Ltd.

doi:10.1016/j.jfoodeng.2004.05.060

* Tel.: +49 40 42878 3040; fax: +49 40 42878 4072.

E-mail address: [email protected]

Characteristic values for the gaseous, liquid, and

supercritical state are listed in Table 1. In the supercriti-

cal state, liquid-like densities are approached, while vis-

cosity is near that of normal gases, and diffusivity is

about two orders of magnitude higher than in typical

liquids.

In processes with supercritical fluids (SCFs), the driv-

ing potential for mass and heat transfer is determined bythe difference from the equilibrium state. The equilib-

rium state provides information about: (i) the capacity

of a supercritical (gaseous) solvent, which is the amount

of a substance dissolved by the gaseous solvent at ther-

modynamic equilibrium; (ii) the amount of solvent,

which dissolves in the liquid or solid phase, and the

equilibrium composition of these phases; (iii) the selec-

tivity of a solvent, which is the ability of a solvent toselectively dissolve one or more compounds, expressed

by the separation factor a; (iv) the dependence of

these solvent properties on conditions of state (p,T)

mailto:[email protected]

Fig. 2. Variations in the solubility of a low-volatility substance (liquid

or solid) in a subcritical (temperature < critical temperature, Tc) or

supercritical fluid (temperature> Tc) as a function of process temper-

ature, process pressure (segmented lines correspond to equal pressure

or isobaric conditions), and solvent density (doted lines correspond to

equal density conditions). (Adapted from Brunner (1994).)

Fig. 1. Definition of supercritical state for a pure component

(Brunner, 1994).

22 G. Brunner / Journal of Food Engineering 67 (2005) 21–33

and, (v) the extent of the two-phase area, as limiting

condition for a two-phase process like gas extraction.

The separation factor a is defined by

a ¼yi=yjxi=xj

; ð1Þ

where xi and xj are the equilibrium concentrations of

component i and j, respectively, in the condensed phase,in mole or mass fractions; and yi and yj, the correspond-

ing equilibrium concentrations of the same components

using equivalent units.

If capacity and selectivity are known, a good guess

can be made about whether a separation problem can

be solved with SCFs. Fig. 2 shows the solubility of a

substance of low volatility, like a triacylglyceride, caf-

feine, or naphthalene in a sub- and supercritical fluid,for instance carbon dioxide (CO2). The solubility in

the subcritical (liquid) fluid (solvent) increases at con-

stant pressure up to temperatures slightly below the Tc

of the solvent. A further increase in temperature leads

at ‘‘low’’ pressures to a decrease of the dissolved amount

of the low-volatility substance in the subcritical liquid

solvent and at ‘‘high’’ pressures still to an increase.

‘‘High’’ and ‘‘low’’ pressures refer to a ‘‘medium’’ pres-sure level which for most systems (including the above

Table 1

Characteristic values of gas, liquid and supercritical state (Brunner, 1987)

State of the fluid Density (g/cm3)

Gas

p = 1atm, T = 15–30�C (0.6–2.0) · 10�3

Liquid

p = 1atm, T = 15–30�C 0.6–1.6

Supercritical fluid

p = pc; T � Tc 0.2–0.5

p = 4pc; T�Tc 0.4–0.9

mentioned) is about 10MPa. The same dependence ofsolubility of the low volatile component in the solvent

remains at temperatures higher than Tc. At ‘‘low’’ pres-

sures, solubility of the low-volatility substance in the

supercritical and near critical solvent decreases with

temperature since density of the supercritical solvent

decreases rapidly with increasing temperature at near-

critical pressures. At ‘‘high’’ pressures, density changes

with temperature are far more moderate, so that the in-crease of vapor pressure is the dominating factor, while

at ‘‘low’’ pressures loss in solvent power induced by

lower density prevails. Analogous solubility behaviour

can be found in systems of a SCF with a solid substance

or a low-volatility liquid.

Carbon dioxide (CO2) is the most commonly SCF

used as a solvent in food applications. It is not only

cheap and readily available at high purity, but also safeto handle and physiologically sound to the very low

levels at which it is present in foods (because it is easily

removed by simple expansion to common environmen-

tal pressure values). Consequently, it is approved for

food processing without declaration. Furthermore,

when being recycled, it does not contribute to the

Diffusivity (cm2/s) Viscosity (g/cm/s)

0.1–0.4 (0.6–2.0) · 10�4

(0.2–2.0) · 10�5 (0.2–3.0) · 10�2

0.7 · 10�3 (1–3) · 10�4

0.2 · 10�3 (3–9) · 10�4

G. Brunner / Journal of Food Engineering 67 (2005) 21–33 23

environmental CO2-problem. The solvent power of

supercritical CO2 (SC-CO2) can be summarized by a

few rules (Brunner, 1987; del Valle & Aguilera, 1999):

(i) it dissolves non-polar or slightly polar compounds;

(ii) the solvent power for low molecular weight com-

pounds is high and decreases with increasing molecularweight; (iii) SC-CO2 has high affinity with oxygenated

organic compounds of medium molecular weight; (iv)

free fatty acids and their glycerides exhibit low solubili-

ties; (v) pigments are even less soluble; (vi) water has a

low solubility (<0.5% w/w) at temperatures below

100 �C; (vii) proteins, polysaccharides, sugars and mine-

ral salts are insoluble; and, (viii) SC-CO2 is capable

of separating compounds that are less volatile, have ahigher molecular weight and/or are more polar, as pres-

sure increases.

2. Overview on techniques

SCF extraction (SCFE) processes can be carried out

in different modes of operation. In most cases extractionfrom solids is concerned, which usually is carried out in

batch and single stage mode, since solids are difficult to

handle continuously in pressurized vessels and separa-

tion factors are high. Fluid mixtures often have separa-

tion factors which make necessary the application of

multi-stage contacting, which is carried out most effec-

tive in a counter-current mode. If separation factors

are approaching 1, many theoretical stages are necessaryfor separating the components. Chromatography is a

magnificent tool for separating similar compounds.

Although chromatography with SCFs as solvents and

mobile phases is presently mostly applied in analytical

separations, the advantageous possibilities of gaseous

solvents should also be applied in chromatographic

processes on a preparative and process scale.

2.1. Single stage supercritical fluid extraction (SCFE)

The extraction of valuable materials from solid sub-

strates by means of SCFs has been carried out on a com-

mercial scale for more than two decades. Large-scale

Fig. 3. Flow scheme of single stage processing with supercritica

processes are related to the food industry like the deca-

ffeination of coffee beans and black tea leaves and the

extraction of bitter flavours (a-acids) from hops. Smaller

scale processes comprise the extraction and concentra-

tion of essential oils, oleoresins and other high-value fla-

vouring compounds from herbs and spices, and theremoval of pesticides from plant material. The extrac-

tion of edible oils would be a large-scale process, but

as for all commodity products, the value-added is not

high, so the economy of the process is the main problem

and must be considered separately for each case. Here,

the term ‘‘large-scale’’ comprises vessels of P20m3 for

decaffeination and of �5m3 for extraction of hops; most

other plants consist of pressure vessels in the range of 1to several cubic meters internal volume. Early in the 80s,

pressure vessels of about 40m3 volume were built for the

extraction from solids. Later, vessels became smaller,

and today most are 61m3 in capacity. The maximum

throughput of a single plant for extraction from solids

is well above 10 thousand tons per year. SCFE from sol-

ids is carried out by continuously contacting the solid

substrate with the supercritical solvent. The solid sub-strate in most cases forms a fixed bed. The SCF flows

through the fixed bed and extracts the product compo-

nents until the substrate is depleted. This extraction

from solids consists of two process steps, namely, the

extraction, and the separation of the extract from the

solvent (Fig. 3). In the extraction, the SCF flows

through a fixed bed of solid particles and dissolves the

extractable components of the solid. The solvent is fedto the extractor and evenly distributed to the inlet of

the fixed bed. The loaded solvent is removed from the

extractor and fed to the precipitator. The direction of

flow of the SCF through the fixed bed can be upwards

or downwards. At high solvent ratios (ratio of flow of

SCF to the amount of solid) the influence of gravity is

negligible. The solid material will be depleted from the

extractable material in the direction of flow. Concentra-tion of extract components increases in the direction of

flow in the SCF and in the solid material. The shape of

the concentration curve depends on the kinetic extrac-

tion properties of the solid material and the solvent

l fluids (for example for the extraction of solid materials).

Fig. 4. Extraction of oil from rape seeds (25g substrate) using 200NL/

h of SC-CO2 as the solvent: influence of process temperature and

pressure on typical curves of integral extraction yield versus specific

solvent consumption. (Adapted from Brunner (1994).)

Fig. 5. Total number of pressure vessels >0.1m3 for supercritical fluid

extraction processes. (Data from Gehrig (1998) and Fukuzato (2003).)


power of the SCF which, in turn, depend on operating

conditions. For the solid as well as for the solvent, the

extraction is an unsteady process (Brunner, 1994).

The course of the extraction process can be followed

by determining the amount of extract against the time of

extraction or solvent consumption. The amount of ex-

tract accumulating during the course of the extraction

will be typically shaped as any one of the curves inFig. 4. The first part of the curve may be a straight line,

corresponding to a constant extraction rate, whereas in

the second part the extraction rate decreases as a limit-

ing yield value is approached, which is given by the total

amount of extractible substances in the substrate.

Although the slope of the first part of the graph can

be given by the equilibrium solubility, a straight line

can be caused also by constant resistance to mass trans-fer, and is no proof that equilibrium conditions are at-

tained during the extraction.

The example in Fig. 4 corresponds to the extraction

of pretreated rape seeds with supercritical carbon diox-

ide. Oils from oil seeds mainly consist of triglycerides

of C16-to C20-fatty acids, which are fairly soluble in

SC-CO2. The extraction rate clearly depends on condi-

tions of state for the extraction, which determines thesolvent power of the SCF. The total yield of oil is gen-

erally the same for the conditions of reasonable solubil-

ity, but is lower at relatively low pressures than at higher

pressures. The pretreatment of the oil bearing material

has also a major influence in total yield. If solvents are

used which are better solvents for triglycerides (for in-

stance short chain paraffines, like propane), the rate of

extraction can be enhanced at moderate conditions.The extraction of oil from oil seeds should be one of

the primary targets of gas extraction, if the value-added

by extraction for most vegetable oils were not too low.

However, there may remain specialty oils or valuable

components co-extracted with common oils that could

be targeted for commercial application of extraction

with SCFs.

The curve for the total amount of extract (Fig. 4) is a

response curve to the flow of SCF entering the extractor.

The response curve depends on process parameters and

all the phenomena occurring during the extraction in thefixed bed. Some of these phenomena include the radial

distribution of the solvent at the inlet, the back-mixing

of the SCF as it flows through the fixed bed due to the

uneven size, surface, and distribution of the solids, and

the self-diffusion of the solvent. Due to the kinetic of

mass transfer, the concentration of extracted substances

in the SCF has an axial concentration profile that corre-

sponds to an axial concentration profile in the solid.Moreover, a radial concentration profile in the solid

and in the gaseous phase overlays the axial concentra-

tion profile.

The course of extraction can also be followed by the

remaining amount of extractible components in the

solid. The extract is depleted monotonously in the solid

substrate with increasing time of extraction or amount

of solvent. When the resistance to mass transfer resist-ance is fully localized in the fluid phase only a straight

line results (where the slope of the line represents the

extraction rate), but when there exist resistances in both

the fluid and solid phase, the course of extraction fol-

lows an exponential function. There may be a total

depletion of the substrate of extractible components, if

the concentration of these compounds in the solvent is

zero and there are no irreversible reactions of theextractible compounds with the substrate. Otherwise

the extraction curve approaches an asymptote.

Quite a large number of industrials plants (maybe

around 100) of different size have been built during the

last 20 years for the extraction of solid material with

SCFs in a batch mode. Since the early 80s, a total num-

ber of about 100 vessels bigger than 100 l in volume have

been ordered for about 50 plants (Fig. 5). They aremostly distributed in Europe, the USA, Japan, and in

Fig. 6. Schematical drawing of a system used for counter-current

extraction of ethanol with SC-CO2. Solid lines represent fluxes of

liquid feed and/or supercritical fluid, and dashed lines represent fluxes

of gaseous solvent.


the South East Asian Countries. Standard designs are

available which can be acquired from various suppliers,

and may be customer tailored to the individual task

(Lack & Seidlitz, 2001). These standard designs typically

consists of one or several extraction vessel(s), a pressure

vessel (separator) for precipitating the extract, and acycle gas compressor or cycle pump, depending whether

the solvent (CO2) is recycled in gaseous or liquid state.

Heat exchangers have to be added to adjust the temper-

ature of process materials. Additional equipment com-

prises pipes, valves, measuring and controlling devices.

In many cases, two or more vessels are used in a plant,

for recovery of the solvent (usually CO2), and for step-

wise extraction simulating a few stages of counter-cur-rent extraction.

The simplest mode of operation for a fixed bed

extraction consists of contacting it until a certain

amount of extract has accumulated or a certain mean

residual concentration in the solid raffinate is achieved.

However, even for an extract consisting of a pure com-

ponent this is not the best way to carrying out the proc-

ess. During the extraction process, extraction kineticschange due to the depletion of the solid substrate and

therefore optimum process conditions change. In addi-

tion, loading of the solvent may be enhanced by carrying

out the extraction in several stages. For an extract mix-

ture of different compounds, an extraction in several

stages can yield different extract products. Each extrac-

tion stage may be designed for different process condi-

tions, even the solvent can be different.Multi-stage counter-current contacting is the most

effective mode. It reduces the amount of solvent and

makes possible continuous production of extract. Real

counter-current contact is not easily established for sol-

ids, since special effort is necessary for moving the solid,

with increased difficulties at elevated pressure. There-

fore, it is easier keeping the solid material stationary

and achieving counter-current contact by other means.Either several fixed beds can be applied and contacted

in such a way that the bed with the highest extract con-

centration is contacted with the gas with the highest

loading of extract and vice versa, or one fixed bed in a

column is used and individual sections are formed by in-

lets or outlets for the solvent streams.

2.2. Counter-current multi-stage processing

Counter-current operation of a separation device re-

duces the amount of solvent necessary, increases

throughput, and enables higher extract concentrations

in the solvent and lower residual concentrations in the

raffinate than does single-stage or multi-stage cross-cur-

rent operation. Counter-current operation is therefore

useful for separations with high separation factors, as,for example, the extractions from solid substrates de-

scribed above. But counter-current operation is abso-

lutely necessary for achieving a reasonable separation

between two substances with a relatively low separation

factor.

In Fig. 6 the process scheme of a counter-current gas

extraction for the separation of two components into

practically pure substances is shown (Brunner, 1994).Process equipment consists of the separation column

where gaseous and liquid phases are contacted coun-

ter-currently, a separator at the top for separating sol-

vent and extract, devices for feeding reflux to the

column, for recovering top product, for delivering feed

to the column, for recovering product at the lower end

of the column, and for recycling the solvent. When the

feed is introduced at an intermediate position, the sepa-ration column consists of two separation cascades: in

the enriching section (the upper part) the bottom prod-

uct compounds are separated from the top product com-

pounds and rejected to the lower or stripping section; in

the stripping section the top product compounds are

separated from the bottom product compounds and

transported to the enriching section.

At the top of the column the separator removes theextract from the solvent. From the extract a specified

part is separated and introduced at the top of the col-

umn as reflux. The remaining part of the extract is the

top product. The solvent is reconditioned (filtered,

sometimes liquefied and again evaporated for removing


trace substances, pressure and temperature are adjusted)

and recycled by the cycle pump as the supercritical sol-

vent at the bottom of the column. The feed is introduced

at an intermediate location in the column by the feed

pump; for a mixture of two components of about the

same concentration, this location is at about the middleof the column. Although multi-component mixtures pre-

vail, it is common practice to try to reduce multi-compo-

nent separations to two-component separations, because

these still retain the main features of a multi-component

separation but are simpler to analyze as described

above.

Fig. 7 presents an example of a multi-component sep-

aration, corresponding to the separation of fatty acidethyl esters (C16- and C18-compounds from C20- and

C22-compounds) (Riha, Meyer, Birtigh, & Brunner,

1996; Tiegs, Steiner, Riha, & Brunner, 1996). Concen-

tration profiles are typical for a multi-component and

multi-stage separation of key compounds (C18 and

C20) and accompanying compounds (C14, C16 on the

high-solubility or extract side and C22 on the low-solu-

bility or raffinate side). Key compounds exhibit typical

Fig. 7. Counter-current separation in a column of fatty acid ethyl

esters according to the length of the carbon backbone, using SC-CO2

as the solvent. Concentration profiles along a column are presented for

(h,j) C14, ( , ) C16, (n,m) C18, (,,.) C20, and ( , ) C22 fatty

acid ethyl esters (closed symbols and solid lines represent compositions

of the liquid phase, whereas open symbols and dotted lines those of the

gaseous phase). The bottom product contained >97% C20/C22-

fraction, whereas the top product contained >90% of C16/C18-

fraction. The dashed horizontal line for theoretical plate 23 represents

the feed plate.

maximum values in concentration near the ends of the

column, while accompanying compounds are trans-

ported to the column ends and depleted rapidly in the

other direction. In an analogue manner, C20 can be sep-

arated from C22 in different runs. The process in Fig. 7

is developed for industrial application (Fleck, Tiegs, &Brunner, 1998). Due to the relative high solubility of

the fatty acid ethyl- or methyl-esters in SC-CO2, the

process is competitive to distillation. In addition, distil-

lation cannot produce the highly enriched fractions––up

to 90%––which are possible with SCF counter-current

separation.

Capacity and diameter of a column for counter-cur-

rent multi-stage separation must be determined fromthe planned throughput for the design of industrial scale

plants. Costs of counter-current separation are also

determined by maximum possible throughput. In a

gravity driven counter-current column, the parameter

to know is the flooding point. The experimental deter-

mined flooding diagram for SCF–liquid counter-current

columns is essentially the same as for gas–liquid

columns operating at normal pressures. At high pres-sures, however, the effect of not only the type of mass

transfer equipment should be taken into account, but

also the influence of density on the gaseous phase.

The capacity of SCF–liquid counter-current columns

often is much higher than anticipated. Low viscous sys-

tems will allow a relatively large linear gas velocity

of about 40–50mm/s. Thus, relatively high quantities

of SC-CO2 are needed for optimum mass transfer: thethroughput of SC-CO2 in a counter-current column

increases from 49kg/h for a 25mm-diameter column to

785kg/h for a 100mm-column, for a measured flooding

point of 100ton CO2/(m2h) and a linear gas velocity of

46mm/s (Fleck, 2000). Scale-up of column diameter is

possible maintaining similar separation performance.

Yet capacity is essentially determined by the type of

mass transfer equipment. With Sulzer CY regular pack-ing, capacity is in the range of 80–100ton CO2/(m

2s).

The diameter of a column will then increase from about

200mm for a capacity of 200ton per year, to 450mm for

1000ton per year (Fig. 8).

Industrial scale applications have been developed for

the fractionation of the n-3 fatty acid esters, and for the

enrichment of tocopherols and tocotrienols. Other

possible applications include: (i) separation of alcoholand water; (ii) separation of aroma compounds (essen-

tial oils); (iii) fractionation of citrus oils; (iv) recovery

of squalene from edible oils; (v) de-acidification of edible

oils; and, (vii) enrichment of carotenoids. Columns of

100 and 200mm diameter have been successfully oper-

ated for counter-current separation, which have allowed

a throughput of about 10ton per year. However,

columns of up to 500mm diameter with a throughputof several hundred tones per year can safely be designed,

if demand shows up.

Fig. 9. Chromatograph for preparative-scale separation of phytol

isomers. The process conditions were as follows: Column = Li chro-

spher Si 60 (10lm); temperature = 40�C; pressure (before col-

umn) = 240bar; flow rate = 3NL/min of CO2; modifier = 3% w/w

isopropanol; injection volume = 20ll of phytol pure (dotted line

�17mg/ml of crude phytol sample––) or diluted in n-hexane (dashed

line �0.85mg/ml of crude phytol sample––); UV-detection wave-

length = 221nm. Peaks for retention time <2min correspond to

samples impurities, whereas peaks in the 4–6.5min range correspond

to the isomers of phytol (Figure adapted from Depta et al. (1999).)

Fig. 8. Column diameter for a separation column separation C18 to

C22 fatty acid ethyl esters. In all cases, the loading of the gas phase was

4% (w/w) and the separation factor was 1.3, so that 52 theoretical

stages, each 0.3m in height, were required to reach a product purity of

99% (w/w) when employing a reflux ratio of 10. Curves indicate

maximum throughput for two regular packings, namely Sulzer EX for

the lower throuphput (Q = 27.5ton/m2/h), and Sulzer CY for the

higher throughputs (Q = 86.2–100ton/m2/h). (Figure adapted from

Riha et al. (1996).)


2.3. Preparative chromatographic separation

Another separation technique where SCFs may beapplied is chromatography. In chromatographic separa-

tion, the supercritical solvent is used as mobile phase.

Analytical packed columns have diameters up to

4.6mm. Columns can be directly scaled-up in diameter

if measures are taken that the fixed bed of the packed

column with particle diameters in the range of

20–40lm remains unchanged during operation and

start-up and shut-off operations. One way to achieve thisis by axial compression, where a piston compresses the

fixed bed in the cylindrical column to keep it under

higher pressure conditions than operational levels all

the time the device is in use. This has allowed long oper-

ational time use of months in our laboratories (Depta,

2000).

The separation of phytol isomers is shown in Fig. 9 as

an example of preparative chromatographic separation.Analytical and preparative separations were carried out

in the elution mode which is a batch separation process.

However, a continuous chromatographic separation can

be designed, for example as a simulated moving bed

(SMB). (The flow scheme of such an apparatus was re-

ported by Depta, Giese, Johannsen, & Brunner (1999),

who arranged eight columns in a circle, to allow a

switching in their function by a multi-port rotatingvalve.) The advantages of SMB-chromatography are

lower solvent ratio, higher throughput, and higher prod-

uct concentration. With 30mm-diameter columns about

10ml/h of phytol isomers can be separated to more than

95% purity.

Applications of chromatographic separations com-

prise analytical and preparative separations. There aremany publications available on analytical separations.

Advantages of this type of analysis is that compounds

spanning a wide range in terms of volatility and polarity

can be separated in a single run. The analysis of edible

oil compounds is particularly advantageous.

Preparative separations can be used to separate

enantiomers (ibuprofen, phytol), produce standards

(tocotrienols), and purify pharmaceuticals (vitamin D).Chromatographic separations have been carried out so

far using columns of up to about 100mm-diameter,

allowing throughputs which are far below 1ton per year.

The maximum diameter employed in our laboratory was

70mm, but it is technically possible to scale-up these

processes.

3. Special applications of supercritical fluids to food

processing

Some products possibly produced by SCF technology

may be found on our everyday�s table. Some examples

of products are listed in Fig. 10, which can be obtained

by the methods discussed above. Processes to obtain

vitamin additives, de-alcoholize beverages, de-fat potato

Fig. 10. Supercritical fluid technology applied to everyday�s food.

Fig. 11. Percent oil removed from salted and unsalted potato chips as

a function of solvent usage when employing SC-CO2 at 55 ± 1 �C and

40.8 ± 0.7MPa as the solvent. The separation conditions were

30 ± 3 �C and 0.1MPa in both the cases. (Figure adapted from Vijayan

et al. (1994).)


ships, and encapsulate liquids will be discussed in more

detail below. For more information on the other exam-

ples the reader is referred to the literature, in particular

King and Bott (1993), Brunner (1994, 2000), McHugh

and Krukonis (1994), Bertucco and Vetter (2001), and

Arai, Sako, and Takebayashij (2002).

3.1. Removal of fat from foods

Edible oils and their components have been the target

of SCF processing since the early 70s. Although triacyl-

glycerides are only fairly soluble in SC-CO2, the advan-

tages of organic solvent-free processing have stimulated

research and development in various areas. One of theseis the removal of fat from food. As an example in Fig. 11

the extraction curves for fat from potato chips is shown.

The process has been fully designed for commercial

application, using the aforementioned standard design.

The process has the advantage of producing fat-free or

fat-reduced potato chips. According to the expected

taste the amount of remaining fat in the potato chips

can easily be controlled.

3.2. Enrichment of vitamin E from natural sources

SCFE offers several advantages for the enrichment of

tocochromanols over conventional techniques such as

vacuum distillation, in particular a lower operating tem-

perature (Brunner, 1994). As starting material one can

use various edible oils or their distillates. Most promis-

ing as feed materials are crude palm oil (CPO) and soy-

bean oil deodorizer distillate (SODD). CPO containsseveral tocotrienols and tocopherols at a total concen-

tration of approximately 500ppm. SODD may contain

(after several conventional concentration steps) about

50% tocopherols. Both materials can be used for the

production of enriched fractions of tocochromanols.

Although it is possible to recover tocochromanols di-

rectly from CPO, it is better to produce esters of the tri-

glycerides in order to be able to more easily separatethese compounds from the tocochromanols. In this

method, the triglycerides are subject to an esterification

with methanol to form fatty acid methyl esters, which

are easily extractable with CO2. That means that the toc-

ochromanols, together with other unsaponifiable matter

(squalene, sterols, etc.) are enriched in the bottom phase

of an extraction column. This attempt is described in

more detail by Jungfer (2000).For a discussion of enriching tocochromanols, phase

equilibrium data have to be considered first. In Fig. 12

the distribution coefficients Ki (distribution between liq-

uid and gaseous phase) of the components of interest are

shown for the example of CPO with SC-CO2 at 67 �C.Free fatty acids (FFA) and tocochromanols exhibit a

much higher solubility in CO2 than the triglycerides.

Hence, these components are enriched in the gaseousphase, expressed by a distribution coefficient being high-

er than one. The distribution coefficient of the triglycer-

ides is smaller than one, whereas that for the carotenes is

much smaller than one, meaning that these components

stay in the liquid oil phase. Thus, tocochromanols can

be extracted as the top phase product in a separation

column, whereas carotenes remain in the bottom phase

product together with triglycerides. For recovering thecarotenes together with the tococromanols the above

mentioned esterification to volatile (CO2 soluble) methyl

Fig. 13. Counter-current separation of squalene, tocopherols and

sterols in a column using SC-CO2 as the solvent. Concentration

profiles along a column are presented for (h, j) squalene, ( , )

tocopherols, and (n, m) sterols (closed symbols and solid lines

represent compositions of the liquid phase, whereas open symbols and

dotted lines those of the gaseous phase). The dashed horizontal line for

theoretical plate 3 represents the feed plate. The experiment was

carried out at 353K and 23MPa by employing 86g/h of feed, 3.6kg/h

of SC-CO2, and 55g/h of reflux. (Figure adapted from Saure (1996).)

Fig. 12. Distribution coefficients for (h) tocochromanols, ( ) free

fatty acids, (n) triglycerides, and (,) carotenes of crude palm oil

(700ppm tocochromanols, 1–5% (w/w) free fatty acids, 95–99% (w/w)

triglycerides, and 800ppm carotenes) in SC-CO2 as a function of

pressure. The segmented line separates the less volatile components

that are enriched in the raffinate (below the line) from the more volatile

compounds that are enriched in the extract. (Figure adapted from

Jungfer (2000).)


esters makes possible to recover tocochromanols and

carotenes (together with squalene and sterols) as bottom

product from this natural source.

When the glycerides (in case of the esterification) or

the FFAs from deodorizer distillates have been re-

moved, then there is a feed material available for obtain-

ing enriched fractions of tocochromanols and carotenesof much higher concentration. In this feed material, toc-

ochromanols and carotenes (in case of palm oil) are the

main components and have to be separated from other

unsaponifiable substances present, such as squalene

and sterols. Of these compounds, squalene has the high-

est solubility in SC-CO2, all phytosterols have a rather

low solubility in CO2 (and remain in the oil phase),

and tocochromanols exhibit an intermediate solubilitybetween the two.

In a second separation step tocochromanols are sep-

arated from phytosterols. For the first step concentra-

tion profiles of an experiment are shown (Fig. 13) that

were conducted in a pilot scale plant described in more

detail by Gast and Brunner (2001). The core of this

apparatus is an extraction column having an effective

separation height of 6m equipped with Sulzer EX pack-ing. The inner diameter of the column is 17.5mm and

can be operated either in a true stripping mode, with

the solvent introduced at the bottom and the feed at

the top of the column, or in reflux mode. In the latter

mode of operation the feed is introduced at the middle

of the column and part of the extract is reintroduced

at the top of the column to ensure counter-current flow.

The gaseous solvent is recycled.Investigations employing SODD as feed material for

the extraction were performed at temperatures of 353

and 363K and at pressures of 23 and 26MPa. The sol-

vent-to-feed (S/F) ratio was varied in the range between

33 and 171, and the reflux ratio was set between 1 and

38. Experiments showed that, as expected, squalenewas enriched in the top phase product. In fact, it could

not be found in the bottom phase. Consequently, it was

enriched from 3.1% (w/w) in the feed to 18.8% (w/w) in

the top phase product. The sterols were completely en-

riched in the bottom phase, occasionally to more than

50% (w/w). If the concentration approaches 50%

(w/w), the products become too viscous and clog the

packing mash and the column (Saure, 1996). For furtherpurification, the sterols and the tocochromanols have to

be separated first. The tocochromanols, as intermediate-

soluble components, were neither distinctly enriched in

the top nor in the bottom phase product.

In order to simulate a taller extraction column, the top

phase product of some experiments was collected and

introduced as feed a second time. With this sterol-free

feed material, the separation of squalene and tocochro-manols was now possible. Starting from a tocochroma-

nol concentration of 48.3% (w/w) an enrichment to

94.4% (w/w) in the bottom product was obtained (Gast,

Jungfer, Saure, & Brunner, 2003). This separation was

performed at 23MPa, 353K, having a S/F ratio of 110

and a reflux ratio of 4.6. The contained squalene could

be completely enriched in the top phase product.

Products with a tocotrienol and/or a tocopherol con-centration in the range of 70–90% are obtainable with


SCF separation techniques. A further purification of

these compounds is possible, e.g. with adsorptive or

chromatographic techniques, again using SCFs.

3.3. Removal of alcohol from wine and beer, and related

applications

De-alcoholized wine or beer is achieved by removing

ethanol from water. Distillation is well known for this

purpose with the disadvantage that aroma compounds

will also be removed. New techniques like membrane

separation (pervaporation) emerge, and in between these

is SCFE with CO2.

Starting from an aqueous solution with about 10%(w/w) ethanol, ethanol can be removed by SC-CO2 in

a stripping column as shown in Fig. 14. The rate of eth-

anol removal depends strongly on temperature. Reduc-

ing the alcohol content to values well below 0.5% (w/

w) requires about 2.5h at 45 �C under non-optimized

conditions. Much shorter times for the ethanol removal

can be obtained if flow rates and mass transfer equip-

ment are carefully selected. With the information avail-able in the literature, for instance from Budich (1999), a

column for dealcoholizing aqueous solutions can be de-

signed. Recovery of aroma compounds is achieved by a

side column in which a separation from ethanol is car-

ried out.

A related process that can be mentioned is the recov-

ery of absolute alcohol. Many studies were carried out

at conditions of complete miscibility of ethanol andCO2 in order to get a high solubility of ethanol in the

vapor phase (Gilbert & Paulaitis, 1986; Kreim, 1983).

At these conditions, anhydrous ethanol cannot be pro-

duced. However, ethanol can be concentrated above

Fig. 14. Effect of process temperature and extraction time on the

removal of ethanol from an aqueous solution using SC-CO2 at

15.5MPa (superficial solvent velocity = 0.12cm/s). (Figure adapted

from Brunner and Kreim (1985).)

azeotropic composition whenever the pressure in the ter-

nary mixture CO2 + ethanol + water is below the critical

pressure of the binary mixture CO2 + ethanol (Budich &

Brunner, 2003; Furuta, Ikawa, Fukuzato, & Imanshi,

1989; Nagahama, Suzuki, & Suzuki, 1988). Fig. 15

illustrates the separation factor of ethanol–water as afunction of the concentration of ethanol in the solvent-

free liquid phase. The separation factor decreased from

around 30 at infinite dilution of ethanol in water to

approximately 1.25 at infinite dilution of water in etha-

nol. No azeotrope was formed at the conditions investi-

gated. Separation factors are larger compared to data at

atmospheric conditions.

Counter-current multi-stage extraction was carriedout in an extraction column of 6m total height

(25mm ID, equipped with of Sulzer EX packing).

The experimental set-up is shown in Fig. 6. Extraction

conditions were set to 333K and 10MPa. Solvent and

extract were separated by pressure reduction down to

5MPa and washing the extract phase with liquid CO2

in counter-current flow. With a feed of 94% (w/w) eth-

anol, an extract with 99.5% (w/w) ethanol was pro-duced at a reflux ratio of 4 and a S/F ratio of 60

(using 9kg CO2/h and 150g feed/h). Twelve equilibrium

stages were achieved, with a height equivalent of a the-

oretical stage (HETS) of 0.33m. Liquid solvent reflux

was not required during this experiment because the

raffinate was still very rich in ethanol (87% w/w). S/F

ratios are relatively small (30kg/kg) compared to other

counter-current gas extraction processes (Budich, 1999).This is due to large separation factors and a solubility

of pure ethanol in CO2 of 5% w/w at the conditions

investigated.

Fig. 15. Equilibrium diagram for the separation of ethanol–water with

SC-CO2 at 333K and 10MPa. The plot includes experimental values of

the separation factor determined using both ( ) a conventional and

( ) a modified sampling method, as well as ( ) reference values

experimentally assessed at normal pressure (0.1MPa). The line

represents the trend for experimental data at high pressure. (Figure

adapted from Budich and Brunner (2003).)


3.4. Encapsulation of liquids for engineering solid

products

A liquid product can be entrapped by adsorption

onto solid particles (liquid at the outside of solid parti-

cles), by agglomeration (liquid in the free volumes be-tween the solid particles), or by impregnation (liquid

within the pore system of the solid particles). Micro-

spheres or larger capsules can be formed, totally encap-

sulating the liquid. The solid material provides a coating

for the liquid inside.

Such particulate products can be achieved by means

of SCF processing also. An example is the so-called con-

centrated powder form-process, wherein CO2 is mixed(dissolved) in the liquid feed by static mixing. The

CO2–liquid feed mixture is then sprayed into a spray

chamber at ambient conditions together with the sub-

strate material. The CO2 is suddenly released from the

liquid, and the liquid forms small droplets. During the

spraying process, solid substrate and liquid droplets

are intensively mixed and combined to a solid particu-

late product of the type described above. The productis finally removed from the chamber as a free flowing

powder and separated from the outgoing gas stream

by a cyclone. With this type of process, a wide variety

of solid substrates can be applied to uptake liquids of

different kind and up to about 90%. As advantages

can be claimed the easier handling and storage, preven-

tion of oxidation processes, and easier dosage.

Solid products can also be formed under high pres-sure conditions. As an example for such a type of proc-

ess, the encapsulation or adsorption of tocopherol

acetate on silica gel. Here, about 50% of tocopherol ace-

tate can be incorporated onto the silica gel without

apparent change of morphology and flow properties of

the powder. The powder with 50% loading is still free

Fig. 16. Adsorption isotherms for tocopherol acetate (TA) in SC-CO2

on silica gel (Fleck, 2000). The autoclave was operated at 333K and

20MPa, and was fed with ( ) 73% (w/w) TA or (h) 97% (w/w) TA.

The fixed adsorber was operated at 353K using a feed of 20g/min of

SC-CO2.

flowing. The amount which can be adsorbed at high

pressures is comparable to that of normal pressure. Only

at very high densities, the equilibrium loading decreases

(Fig. 16). In the experiments that are reported in Fig. 16,

the autoclave was used to saturate the SC-CO2 current

with tocopherol acetate, and the density of the solventwas changed in the nozzle where the loaded SC-CO2

phase was fed to the adsorber. This adsorption at high

pressures makes possible the direct product formation

in the SCF, with the advantageous effect that the super-

critical solvent can easily be recycled without substantial

compression.

4. Costs of separation processes with supercritical fluids

Reported costs for production rates around 1000ton/

year of solid feed are in the range of 3US$/kg feed.

Economy of scale may bring costs down to less than

0.5U$/kg for batch operation. Continuous operation

would further reduce costs. The lower curve in Fig. 17

represents an anticipated fivefold increase in productiv-ity (mostly due to shorter use of high pressure volume

for extraction) reducing operating costs in the same pro-

portion. Yet continuous extraction has been so far car-

ried out only on laboratory scale. It has also to be

considered that cost estimates of such a type have an

agreed variability of ±30%. Therefore, some data of dif-

ferent authors and processes are shown in Table 2.

Another interesting feature is the cost structure forSCF processing. Lack and Seidlitz (2001) published data

for the production of extracts from hops. For extraction

pressure of 350bar, a recovery pressure for the extracts

of 45bar and 20 processing weeks they calculated

processing costs of 1€/kg of feed material. The cost

breakdown without considering the cost of the raw

material is as follows: interest and depreciation = 36.1%;

labour = 24.5%; utilities = 17.2%; taxes = 20.5%; admin-istration = 1.0%. Investment costs in this case are about

Fig. 17. Economy of scale for SC-CO2-extraction of solids.

Table 2

Cost estimates for processing of solids with supercritical carbon

dioxide (Lack and Seidlitz, 2001)

Type

of process

Capacity

(ton/year)

Cost estimate

(€/kg feed)

Decaffeination of

coffee beans

(batch, non-isobaric)

3500 1.1

4500 0.9

7000 0.75

Decaffeination of

coffee beans

(batch, isobaric)

3500 0.85

4500 0.7

7000 0.55

Solids extraction, general 10,000 0.7

Removal of pesticides

(from ginseng)

600 14

1000 9

2000 5


one third of total processing costs. This value is still

higher than for equipment operated at ambient pressure

but lower than the 50% mostly cited.

Fleck, Brunner, and Karge (2000) and Fleck (2000)

have carried out a cost analysis for counter-current

processing raw tocopherol acetate to industrial grade

tocopherol acetate (about 92% pure) using SC-CO2 with

the addition of some propane (10–20%). For a plantwith a 300mm ID column the production rate per year

was determined to 450ton (20% propane) with resulting

production costs (without costs for the raw material) of

1.4€/kg product. For a plant with 500mm ID column

the respective values for capacity and cost have been

estimated as 1090–1250 ton/year (for 10% or 20% pro-

pane) and 1.05€/kg product.

Table 3 summarizes the effects of using propane-CO2

mixtures as the solvent on operating costs. The base case

is high vacuum distillation (p < 1mbar). Costs are twice

as high with SC-CO2 and solvent recovery by pressure

reduction. If propane is added, costs can be reduced to

20% of the base case. If the favourable case occurs that

the solute can be adsorbed, and the adsorbate used as

product, cycle costs are only 50% of the base case, even

with pure CO2. If, in addition, propane is added, costsmay be less than 10% of the base case.

Table 3

Operating costs for the purification of tocopherol acetate for different

solvent mixtures and different methods for precipitating the extract

(Fleck, 2000)

Method

of precipitation

Supercritical

solvent

Operating

costa (%)

Pressure reduction CO2 200

CO2 + C3H8 20

Adsorption CO2 50

CO2 + C3H8 8

a Values are expressed as a percentage of the base case (operating

cost = 100% for high vacuum distillation).

Another possibility for reducing costs is a low pres-

sure drop along the solvent cycle. To achieve this, sepa-

ration of the solute from the supercritical solvent

by membranes would be an ideal means (Sartorelli &

Brunner, 2000). In cooperation with the Institute of

Chemistry of GKSS in Hamburg, Germany, our groupsucceeded in finding a polymer membrane which can

be applied (Sartorelli, 2001). An active multiple layer

of poly-tetra-fluoro-ethylene supported by a ceramic or

organic material proved to be applicable at high CO2-

pressures (tested up to 35MPa). The solute is retained

and SC-CO2 permeates through the membrane at high

rates. Still the membrane has to be tested in a module,

and the integration into a process cycle has to be proven.

5. Conclusions

Processing materials with SCFs is a proven and

industrially applied technology. It is readily available

for extraction from solids (also in simulated counter-

current mode) and multi-stage counter-current separa-tion. Chromatographic techniques with SCFs as a

mobile phase are primarily used for analytical purposes

by now. Some plants for preparative scale are available

as for SMB chromatography. Processing costs can be

very competitive to other processes but sometimes

supercritical processes are unique in their ability to pro-

duce solvent-free products and handle high viscous

material. New technologies are emerging, for instancefor particle design, sterilization, and separation of

enantiomers.

Acknowledgments

Useful discussions with Jose M. del Valle are grate-

fully acknowledged as is the funding by FONDECYT(International Cooperation project 703-0033) from

Chile, DFG (Deutsche Forschungsgemeinschaft) from

Germany, and F. Hoffmann-La Roche AG (Basel,

Switzerland).

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