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J. of Supercritical Fluids 41 (2007) 391–403

Sunflower oil hydrogenation on Pd in supercritical solvents:Kinetics and selectivities

A. Santana, M.A. Larrayoz, E. Ramırez, J. Nistal, F. Recasens ∗Chemical Engineering Department, ETSEIB, Universitat Politecnica de Catalunya, 08028 Barcelona, Spain

Received 13 June 2006; received in revised form 9 December 2006; accepted 15 December 2006

bstract

The partial hydrogenation of sunflower oil on a few supported Pd catalysts in supercritical (SC) dimethyl ether (DME) as reaction solvent wastudied to obtain hydrogenates with low trans C 18:1 and stearic contents.

The kinetics was determined on eggshell 0.5% Pd/Al2O3 and uniform 2% Pd/C catalysts using a sequential experimental design in a contin-ous, radial-flow, internal recycle reactor. The operating variables were temperature (456–513 K), pressure (18–23 MPa) and the space-velocityWHSV = 41–975 h−1). The rotation frequency and the molar feed concentration (oil:H2:DME) were held constant at 157 rad/s and 1:4:95 mol%,espectively. Kinetic scheme was based on that published before. Some reactor runs were simulated using mixed-flow assumption and the kineticsata for both systems with good results. A comparison was established between the eggshell 0.5% Pd/Al2O3 in DME and the data for 2% Pd/C

n propane with respect to trans production and stearic formation. trans seems to be lower using 2% Pd/C in propane, while the undesired stearicormation is less on the eggshell 0.5% Pd/Al2O3 catalyst in DME. An overview is presented on the merits of the catalysts available for the SCFrocess in terms of linoleic selectivity and trans yield on a few vegetable fats.

2007 Elsevier B.V. All rights reserved.

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eywords: Supercritical solvent; Vegetable oil hydrogenation; Phase equilibriu

. Introduction

Supercritical fluid technology is becoming important in theipid and food industry in a variety of fields. So far the empha-is, however, has been on extraction processes [1]. Reaction andatalytic applications lag behind other processes as can be seenn Brunner [2,3]. In general, the benefits of SCF in heteroge-eous catalysis have been emphasized in certain isomerizationeactions involving coke deposition and for enhancing intraparti-le diffusion [4–6]. In catalytic reactions, hydrogenation standsmong the important reactions in petroleum processes and in theood and fine chemical industries [7].

The purpose of partial hydrogenation in vegetable oil con-itioning is to obtain a more stable product (no oxidation ontorage), together with a suitable texture and melting-tempera-

ure range at human mouth conditions for use as margarine anddible shortenings. Catalytic slurry batch process is customaryn the vegetable oil hydrogenation industry with nickel (Raney

∗ Corresponding author. Tel.: +34 93 401 0939; fax: +34 93 401 7150.E-mail address: f.recasens@upc.edu (F. Recasens).

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896-8446/$ – see front matter © 2007 Elsevier B.V. All rights reserved.oi:10.1016/j.supflu.2006.12.009

netics; trans fatty acid content; Supported Pd catalysts

r supported), or supported Pd as catalysts, on an otherwise wellstablished catalysts and technology [7–9].

Fatty acids of natural triglycerides with unsaturation occurn nature in the cis isomeric form and multiple double bondsre isolated on either of side of a methylene chain. However,ydrogenation can cause isomerization to an unwanted transonfiguration and also to conjugated double bonds. In the con-entional low pressure hydrogenations, the trans content can bes large as 40 wt% [10]. Although the effects of trans isomersn human health are not clear, they are suspect to have similarffects as saturated fats, both affecting coronary health.

In contrast, the continuous processes using Pd as catalyst andupercritical fluids (propane or CO2) as solvent were introducedn the 1990s by Tacke et al. [11–14] and Harrod and Møller ando-workers [15–19,57] showed that the reaction is extremely fastnd the formation of trans C18 is reduced. Later, King et al. [20]sing SC carbon dioxide on supported Ni [20] in batch foundhat the reaction conditions had a strong influence on the char-

cteristics of the final product, Ramırez et al. [21] summarizedhe data of several authors regarding hydrogenation rates.

To present, Denmark has been the only country in the worldhere the Ministry of Health has limited by law the trans content

392 A. Santana et al. / J. of Supercritical Fluids 41 (2007) 391–403

Nomenclature

Ai pre-exponential factor (mol−0.5 m4.5 kg−1 s−1 ormol−1 m6 kg−1 s−1)

cis C18:1 oleic fatty esterCi mol concentration of fatty ester i (mol m−3)C

expq,i,d experimental observed concentration of the

species i (mol/m3)Cmdl

q,i,d calculated model concentration of species i

(mol/m3)E activation energy (J mol−1)Fi molar flow rate of species i (mol s−1)IV iodine value (g I2/100 g oil): 1 IV = 36 mol H2/

m3 oilkij kinetic rate constant (mol−0.5 (m4.5) kg−1 s−1 or

mol−1 (m6) kg−1 s−1)Ki VLE constant; Ki = yi/xi, i = 1 for SCF and i = 2

for sunflower oilNq number of hydrogenation runsNq,d number of experimental pointsNq,i number of componentsP pressure (MPa)Pc critical pressure (MPa)ri global reaction rate of species i (mol s−1 kg−1)Si specific isomerization, or trans yield, see Eq. (9)SLo linoleic selectivity, see Eq. (8)trans C18:1 elaidic fatty esterT temperature (K)W mass of catalyst (kg)WHSV mixture-weight-hourly-space-velocity (h−1)

Greek letterχ2 Chi-squared, see Eq. (7)

AcronymsCSTR continuous stirred-tank reactorE elaidic fatty acid, trans C18:1H2 hydrogeni component, i = L, O, E, S, H2L linoleic fatty acid, C18:2O oleic fatty acid, cis C18:1S stearic fatty acid, C18:0

Subscripts and superscriptsc criticald data pointexp experimentali componentmdl modelp partialq experimental0 inlet

Fo

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Ktas(tcthH(

dscotTta

ig. 1. Accepted triglyceride interconversion during hydrogenation of vegetableils (H2 not shown; Wisniak and Albright, [8]).

o less than 2% on fat components for human ingestion since May003. The committee of experts of Codex Alimentarius of theAO is on a debate regarding the inclusion of trans fatty acidsontent on food labels. On the other hand, the EU is favourableo include this on the label, but legal action is yet to be taken.n the US, the government (through the FDA) has put forwardcampaign (announced by the US Surgeon General in 2003)

o label by law the % trans content and the % saturated fat,efore 2006. The efforts of King et al. [20] to develop a lowrans process at the FDA, are in this direction.

The first kinetic scheme of the hydrogenation reaction wasresented by Bailey [22], who proposed consecutive reactionsf the unsaturated triglycerides. The reaction rates for theechanism were found to be pseudo-first-order with respect to

ydrogen [23–26] due to excess triglycerides. Gut et al. [27]roposed a kinetic model accounting for the formation of the cisnd trans isomers during the hydrogenation of sunflower seedil. Their model assumed that both cis and TFAs are adsorbeddentically on the catalyst sites and that the adsorbed doubleond could isomerize while adsorbed. Several attempts haveeen made to treat the catalytic hydrogenation of triglyceride oilsinetically. The kinetic model in Fig. 1 is found to be applicableor the hydrogenation of cottonseed oils by Hashimoto et al. [28],sed by other authors for the design of hydrogenation plants.

From the concepts of Horiuti and Polanyi [29] and Allen andiess [30], it is accepted that hydrogenation occurs by reac-

ion of an adsorbed, dissociated hydrogen next to a fatty aciddsorbed on an adjacent site. The controlling mechanism is theurface reaction between adsorbed species (linoleic (L), oleicO) or elaidic (E) fatty esters) and adsorbed hydrogen atoms,herefore reactions are half order in the dissociating species (H2).is–trans isomerization takes place through the hydrogenationo a saturated intermediate therefore that is also expected to bealf order in H2. This is what is accepted in G/L kinetics with2 pressure of less than 5 bar. Under gas-phase, fluid conditions

i.e. 20 MPa) this can be different.In the Hashimoto model, it is assumed that adsorption and

esorption steps at equilibrium (Langmuir–Hinshelwood one-tep controlling) as well as the catalyst surface is sparselyovered by adsorbed components and thus the concentrationf unoccupied sites is essentially independent of the amount of

he catalyst and type of triglyceride adsorbing onto the surface.hence, the denominators of the Langmuir–Hinshelwood equa-

ions would be close to unity, hence assuming that the glyceridesdsorb with similar coverage.

A. Santana et al. / J. of Supercritical Fluids 41 (2007) 391–403 393

droge

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2

2

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Fig. 2. Continuous hy

We have used the methodology of design of experiments [31]nd response surface models [32] to achieve optimum hydro-enation conditions for the 2% Pd/C catalyst in SC propane inwell mixed-flow reactor. By modelling the system response

mpirically, we showed that it was possible to obtain a hydro-enate fat with less than 5% trans content in one pass throughhe reactor in a continuous process, with a iodine value (IV) ofbout 90 (starting with a value of 135). The kinetic formalismsed earlier for vegetable oil hydrogenation was used to deter-ine the kinetics on 2% Pd/C using SC propane as reaction

olvent.Selective hydrogenation of functional groups in substrates

nd partially hydrogenated fatty acids glycerides and fatty aciderivatives in SCF was reported by Harrod et al. [52], andacher and Holmqvist [58]. The first authors increase diffusion

ates by using a solvent with high diffusive properties and sol-ation capacity towards the substrate, as well as high hydrogenolubility conditions. Selectivity (linoleic to oleic) is enhancedy lowering the catalyst activity and temperature. As we willhow here, higher temperatures impair the linoleic selectivity.e showed in a patent that selectivities are improved if internal

s well as external diffusional resistances are kept low (Campst al. [55]). Using standard catalysts, like the metal uniformnes used here, better selectivities are obtained with shorteriffusion paths, sometimes without a need for lower activities54].

In this paper, the study of the hydrogenation of sunfloweril on a few standard supported Pd catalysts using SC dimethylther (DME) as reaction solvent was carried out. The aim ofhis work was to measure the kinetic constants also on eggshell.5% Pd/Al2O3 catalyst (dp = 2 mm). A uniform metal 2% Pd/Catalyst (dp = 0.47 mm) was also examined. The other purpose

as to establish a comparison between the SC solvents and thed-supported catalyst in order to see what is the best combination

o obtain final products with low trans C18:1 and saturates fattycid contents.

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nation reactor setup.

. Experimental methods

.1. Raw materials

A sunflower seed oil from Helianthus annuus from Sigma–ldrich (Barcelona, Spain) with an initial iodine value (IV) of30 and a fatty acid composition of 13.2% saturated ((C16:0,18:0); 18.4 wt% cis C18:1; 67.6% C18:2 and 0.8% C18:3)as used in the experiments. The (DME) used in this work was

upplied by Grit S.L (Barcelona, Spain) with a minimum purityf 99.9%.

The catalysts used were 0.5% Pd on alumina (spheres ofmm, surface area = 320 m2/g, pore volume = 0.45 cm3/g andulk density = 750 kg/m3. The metal is deposited on an eggshelln the pellets) from Johnson Matthey (Barcelona, Spain). Theecond catalyst used was 2% Pd on activated C (uniformetal distribution, pellets of 2 mm, surface area = 1530 m2/g,

ore volume = 1.3 cm3/g mostly in micropores and bulk den-ity = 360 kg/m3), from Degussa AG (Frankfurt, Germany).efore hydrogenation, this was crushed and sieved to 0.47 mm.he pellets were activated in situ by flushing with N2 (99.999%inimum purity grade) into the reactor in order to remove oxy-

en and then displacing nitrogen with pure hydrogen (105 STPm3/min) at high pressure and temperature for 2 h.

.2. Equipment

The experiments were performed on the same equipmentnd analyzed as reported in earlier work. A flow sheet ofhe laboratory-scale equipment is shown in Fig. 2. A detailedescription of the experimental apparatus along with the analyt-cal method can be found elsewhere [31].

Liquified DME was pumped from the vessel with a dip tubesing a high pressure diaphragm pump (Milroyal D, Dosaproilton Roy, France) to the reactor in order to provide and main-

ain a system downstream pressure of 18–25 MPa, which was

3 critical Fluids 41 (2007) 391–403

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94 A. Santana et al. / J. of Super

et manually with a back pressure regulator (Model PR57, GO,uroval, Barcelona). The sunflower oil was pumped at con-tant flow rate using a HPLC pump (Gilson 305, France) and2 was compressed by a gas booster system (AG-62, Haskel,arcelona) equipped with a high pressure gas reservoir con-

rolled at 30 MPa. H2 flow was metered from the constantressure reservoir through a mass-flow indicating controller-alve (Model 5850S, Brooks Instruments, Euroval, Barcelona)ith computer output.The reaction mixture was made up by mixing known amounts

f solvent, hydrogen and sunflower oil together. The reactantixture was preheated to the desired operating temperature

efore entering the reactor. The reactor was heated with an elec-rical heating jacket. Internal reactor temperature was monitoredith a thermowell located in the bottom. The temperature incre-ent in the reactor relative to the feed was seldom more than

73 K above the temperature of the inlet heater. This is cer-ainly due to the relatively small adiabatic temperature rise foreactions in SCF solvents.

After leaving the reactor, the effluent was continuouslyxpanded to atmospheric pressure on an externally heated nee-le valve (10VRMM2812, Autoclave Engineers, Erie, PA) inrder to control the total flow of mixture through the reactor.he effluent was then sent to a condenser made of a seriesf glass U-tubes, immersed in a EG glycol–water (40%) batheld at 249 K in order to condense the oil from the solvent andnreacted H2 mixture. The flow rate of exhaust gas was mea-ured with a rotameter (2300, Tecfluid, Spain) and sent to anxplosion-proof hood system. Samples of the condensed reac-ion product were analyzed using silver ion high-performanceiquid chromatography (HPLC) after converting to FAME.

.3. Operation in one fluid phase

The phase separation depends on the mixture composition,emperature and pressure. Thus, a substantially homogeneousupercritical phase can be formed by choosing the right operat-ng conditions. From the phase behaviour study developed byans Sole and Camps [33] at 20 MPa and 473 K, the inlet reac-

ion mixture composition can range between 1 and 2 mol% ofunflower oil, 2–20 mol% of H2 and 78–97 mol% DME.

With regard to process economy, the addition of solvent haso be kept to a minimum to avoid high recycle volumes of thisolvent in the production plant. Also, excess hydrogen has toe recycled. Thus, the composition of the reaction mixture isery important for industrial practice. High substrate loadingsnd low concentration of hydrogen are desirable but the con-entration of substrate in the reaction mixture is limited becausehe viscosity of the reacting mixture increases rapidly as moreubstrate is dissolved in the solvent causing single-phase con-itions impossible in the reactor [34]. This is most important inhe laboratory as well as in the plant.

In addition, hydrogen is an anti-solvent in the reaction mix-

ure and reduces the solubility of both substrate and product,s reported Van den Hark and Harrod [35] who have suggestedhat hydrogen/substrate molar ratios must be kept around 10 orower. A similar conclusion was reached by Hitzler et al. [36]

fpap

ig. 3. VLE in the critical region for the binary system DME (1) and sunfloweril (2) (vaporisation constants calculated with PR-EOS).

ho found that at any temperature, the miscibility of the reactionystem (cyclohexane + hydrogen + CO2 in some cases) alwaysmproved when the concentration of H2 is less.

Hitzler et al. [36] have found that the most dramatic changesn the hydrogenated product composition happened near and justelow the critical pressure of the solvent (in the case of DME forpressure of Pc = 5.37 MPa). On the other hand, the total systemressure in the process must be above the critical pressure (Pc)f the mixture in order to insure single-phase conditions.

At the same time, the temperature operating range must beigher than the critical temperature in order to operate in a single-hase condition but not too high to avoid side-products (e.g.ver-hydrogenation of the reaction products, or product decom-osition, or hydrogenation of other functional groups within theubstrate molecule, or polymerization reactions or coke forma-ion).

Based on the considerations explained above, the composi-ion of the inlet reactants mixture for performing the experimentsas chosen as follows: 1 mol% sunflower oil, 4–9 mol% H2

nd 90–95 mol% DME as well as the operating pressure andemperatures, which ensure single vapor phase conditions.

The binary VLE diagram for sunflower oil in DME is shownn terms of the VLE equilibrium constants as a function of pres-ure at constant temperature in Fig. 3. It was calculated withhe Ki = yi/xi values from the ratio of the fugacity coefficientsbtained with the Peng-Robinson equation-of-state [4,37], withero interaction parameters. The calculations were performedith PE 2000 [38] in terms of the convergence pressure [39].lthough this is a rough approximation, the DME–oil binaryixture exhibits a convergence pressure of around 10 MPa,hich is quite below that reported in previous workers [21,31]

or the propane-oil system (14 MPa), estimated with the same

rocedure. Therefore, from the viewpoint of the reactor oper-ting pressure, it would be less expensive to use DME thanropane, and much less than carbon dioxide (convergence pres-

A. Santana et al. / J. of Supercritical Fluids 41 (2007) 391–403 395

Table 1Effect of pressure on conversion and product distribution in the sunflower oilhydrogenation on 0.5% Pd/Al2O3 using DME

Pressure (MPa) IV trans (wt%) C18:0 (wt%)

18 104.12 3.42 10.420 104.59 4.67 9.7822.5 106.34 4.95 9.68

TWf

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e

TEo

T

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Fig. 4. Ternary systems (mol%): (a) dew and bubble point curves for DME/H2a

it

twa

2

emperature = 483 K, feed composition (oil:H2:DME) = 1:4:95 mol%,HSV = 203 h−1. Feed composition trans C18:1 = 0.2 wt%, C18:0 = 7 wt%,

eed MW = 875 kg kmol−1.

ure about 30 MPa). Harrod et al. [52], used butane, which isnteresting in countries where oil refinery butane is pure enough,nd accepted in most countries on food legislation grounds.

Recently, Brake et al. [40] and Weidner et al. [41] haveeported that the principal thermodynamic behaviour of DME inriglyceride hydrogenation is similar to that in the propane sys-em but DME is a slightly better solvent due to its larger regionf complete system miscibility. A set of preliminary runs underhe same experimental conditions except for the reaction sol-ent used (propane or DME) suggests that the rate of reactions fairly higher (around 3%) in the case of DME, which is ingreement with the results of Brake et al. [40] and Weidner etl. [41] mentioned above.

Based on our previous experimental experience of oil hydro-enation on 2% Pd/C using propane as reaction solvent [31]long with several preliminary experiments, operating pressuresnd temperatures were determined. Tables 1 and 2 show theffect of pressure and temperature on conversion (as related tohe final IV) and the formation of trans-mono-unsaturated andaturated species as these concentrations are extremely impor-ant for getting the appropriated product plasticity in margarinesroduction with a low trans content. From the table, it can beealized that the variation of pressure at constant temperatureas no remarkable effect on either reaction conversion or prod-ct distribution. With respect to the temperature, the results ofable 2 show that an increase in this operating variable at con-tant pressure involves an increase in conversion as well as inhe trans-mono-unsaturated and saturated content as expected.

The estimated critical pressure and temperature for the cho-en molar compositions at 20 MPa and 483 K were around.5 MPa and 420 K, respectively. These values were estimated

sing the method of Chueh-Prausnitz [39].

The phase diagram for the ternary system, calculated for sev-ral temperatures at 20 MPa in triangular form [33], is shown

able 2ffect of temperature on conversion and product distribution in the sunfloweril hydrogenation on 0.5% Pd/Al2O3 using DME

emperature (K) IV trans (wt%) C18:0 (wt%)

56 112.9 2.4 9.483 104.6 4.7 9.813 101.5 6.1 15.2

ressure = 20 MPa, feed composition (oil:H2:DME) = 1:4:95 mol%, WHSV =03 h−1. trans C18:1 feed composition = 0.2 wt%, C18:0 feed composi-ion = 7 wt%. Feed MW = 875 kg kmol−1.

0o

rc

TRS

V

THW

2/sunflower oil system estimated with the PR-EOS at 423, 453 and 473 K and0 MPa (Sans, [33]). (b) DME/H2/soybean oil system modelled by Weidner etl. [41] at different pressures and temperatures.

n Fig. 4 along with that reported by Weidner et al. [41] for theernary system soybean oil–H2–DME.

In our experiments, non-condensing conditions were moni-ored by watching the reactor pressure oscillations in the sameay as Van den Hark and Harrod [35]. Condensation manifests

s a wetting of the solid catalyst evident when reactor is opened.

.4. Sequential experimental design

The hydrogenation reactions were carried out on an industrial.5% Pd on alumina-supported catalyst, eggshell pellets of 2 mm

f nominal size.

Based on the above phase equilibrium considerations, theanges of operating variables (see Table 3) were selected to avoidondensation. A Robinson–Mahoney reactor was used. This

able 3anges of variables for the vegetable oil hydrogenation on 0.5% Pd/Al2O3 inC DME

ariable Range

emperature (K) 456–513

2 (mol%) 4–9HSV (h−1) 200–600

396 A. Santana et al. / J. of Supercritical Fluids 41 (2007) 391–403

Table 4Effect of reaction conditions on sunflower oil hydrogenation on eggshell 0.5% Pd/Al2O3 with DME as reaction solvent (oil feed: 1 mol%; composition: cis C18:1,18.0%; C18:0, 7.0%)

Run Temperature (K) H2 (mol%) WHSV (h−1) IV C18:2 (wt%) cis C18:1 (wt%) trans C18:1 (wt%) C18:0 (wt%) Si SLo

1 483 4 567 118.0 55.9 18.3 0.4 7.5 2 –2 456 4 567 127.8 57.5 26.1 0.3 7.8 4 19.23 483 4 283 108.2 49.4 18.4 2.4 7.9 9 –4 456 4 283 117.0 53.8 20.7 1.0 8.6 6 1.925 483 4 203 104.7 45.4 20.6 4.5 9.5 15 1.056 483 8 203 102.8 45.0 19.8 3.8 8.7 12 1.087 483 9 203 100.2 46.4 14.2 3.8 10.9 11 –8 456 4 204 112.9 54.2 14.2 2.3 8.8 10 –9 456 7 204 111.5 53.5 14.0 2.2 9.0 9 –

10 456 9 204 110.1 53.1 13.6 2.0 8.9 8 –11 513 4 203 101.3 45.7 15.0 5.7 14.2 17 –12 513 7 203 99.7 45.3 14.2 5.6 14.9 16 –13 513 9 203 93.8 43.4 12.2 4.8 18.6 12 –11

P

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osupt

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3

3w

Tp(cfR

r

r

r

r

r

iefifclit

Given the above kinetic scheme, it is possible to optimize thekinetic constants by fitting the concentrations in the outlet fluidfor the multiple reaction system as follows. The steady stateconservation equations for multiple reactions in a mixed-flow

4 513 4 283 109.4 50.35 513 7 284 104.4 48.5

ressure = 20 MPa. SLo from Eq. (8b), Si from Eq. (9b).

as a recirculated radial fixed bed with catalyst held betweencreens. The total system pressure, the molar oil concentrationnd the stirrer speed were kept constant at 20 MPa, 1 mol% and31 rad/s, respectively. This stirrer speed value is enough foraking kinetic measurements independent of agitation [31].The amount of catalyst was varied between 0.4 and 1 g in

rder to obtain a conversion similar to that reported earlier forunflower oil hydrogenation using propane as solvent [31]. Wese the space-velocity (WHSV, grams of reactant mixture feder hour per gram of catalyst) as experimental variable. This ishe reciprocal of space-time, W/F.

The operating variables were varied according to a sequen-ial design, where each of the experiment chosen depending onhe results obtained in the previous one. This design is espe-ially appropriate for building kinetic studies because it can beeduced the number of experiments to performance. A summaryf experimental operating conditions carried out is presented inable 4 [56].

. Results and discussion

.1. Kinetic analysis of CSTR data for oil hydrogenationith 0.5% Pd/Al2O3 and DME as SC solvent

The results from the 15 experimental runs are summarized inable 4. The kinetic expressions are assumed for the reaction netroposed by Elbid and Albright [42] using the same assumptionssee Fig. 1). These features have been put in terms of fluid phaseoncentrations as follows, where the rates are the net rates oformation for each species, based on the hypothesis made byamırez et al. [31]:

L = −(k21 + k22)CL√

CH2 (1)

O = k21CL√

CH2 − kOCO√

CH2 + kECE√

CH2 − k11COCH2

(2) F

14.7 5.8 16.1 23 –13.3 5.3 18.0 17 –

E = k22CL√

CH2 + kOCO√

CH2 + kECE√

CH2 − k12CECH2

(3)

S = k11COCH2 + k12CECH2 (4)

H2 = −3k21CL√

CH2 − 3k22CL√

CH2 − 3k11COCH2

−3k12CECH2 (5)

n which an order one-half with respect to hydrogen was consid-red for the reactions involving di- and mono-unsaturates andrst-order for the formation of saturates from oleic or elaidicatty acids. The reaction order with respect to hydrogen washecked by means of plotting log (rate) data as a function ofog (hydrogen partial pressure). The slope of the regression lines 0.42 as shown in Fig. 5. The obtained value is close to theheoretical value 0.5.

ig. 5. Linearized plot of hydrogen reaction rate: ln rH2 vs. ln pH2 for 483 K.

critical Fluids 41 (2007) 391–403 397

r

F

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rtcaamtec

χ

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itgta

fc

Table 5Effect of reaction conditions on sunflower oil hydrogenation

Run IV trans (wt%) C18:0 (wt%)

1 116.6 0.4 7.42 128.7 0.3 7.43 108.7 2.4 7.94 117.9 0.9 7.55 104.6 4.7 9.86 102.8 4.0 9.17 100.2 3.8 10.98 112.9 2.4 9.49 111.7 2.3 9.3

10 110.4 2.0 9.111 101.5 6.1 15.212 99.0 5.8 15.513 94.0 5.0 19.314 109.7 5.7 15.715 104.7 5.2 17.8

Oil concentration: 1 mol%; P = 20 MPa; catalyst: 0.5% Pd/Al2O3; solvent:DME.

Fk

TF

P

kkkkkkχ

P

A. Santana et al. / J. of Super

eactor are:

i0 − Fi + riW = 0, i = 1–5 (6)

here W is the mass of catalyst in the reactor and ri is the globalate of formation for species i per unit mass of catalyst (L, O,, S and H2). Substitution of the rate expressions given before

n the balance equations provides a system of equations in theoncentrations and in terms of parameters kij. So in principlehere would be six parameters to fit at every temperature. Andix more activation energies, if at least two temperatures werevailable.

The final expression results in a non-linear problem to fit theeactor outlet concentrations by guessing the six (kij) parame-ers. The system of equations for the CSTR was solved for theoncentrations using the non-linear Newton–Raphson methodlong with simultaneous optimization of the parameters usingnon-linear least squared method with restrictions in order toinimize a χ2-objective function [43], defined as the devia-

ion between experimental and model concentrations for Nq,dxperimental points of the Nq hydrogenation runs with Nq,iomponents:

2 = 1

Nq

Nq∑

q=1

1

Nq,iNq,d

Nq,i∑

i=1

Nq,d∑

d=1

(Cexpq,i,d − Cmdl

q,i,d)2

(Cexpq,i,d)

2 (7)

here Cexpq,i,d and Cmdl

q,i,d are the experimentally observed andhe calculated model concentration values for the qth experi-

ent and the ith component and dth data point, respectively.ote that the Chi-square used weights the errors differently,

o it weights more the smaller concentrations. In Table 4 arehe results on the isomer formation, to be discussed later.nd in Tables 5 and 6 are the results obtained for the kineticarameters.

Despite the selection of the experiments for the kinetic stud-es should cover conversion ranges between 30 and 80% [44], inhe case of selective hydrogenation of vegetables oils for mar-arines production, the expected conversion is around 50%. Inhe present case study, the maximum conversion reached was

round 31%.

Fig. 6 depicts the predicted and experimental data for theatty acid compositions using the optimal values of the kineticonstants. As can be observed in this figure, the accuracy of

doTp

able 6itted kinetic parameter value

arameter T (K)

456

21a 3.41 × 10−6 ± 5.29 × 10−7

22a 2.78 × 10−6 ± 8.38 × 10−7

11b 1.12 × 10−5 ± 1.93 × 10−6

12b 2.83 × 10−6 ± 5.12 × 10−7

Oa 3.06 × 10−6 ± 9.40 × 10−7

Ea 2.92 × 10−5 ± 1.83 × 10−6

2 8.23 × 10−5

d/Al2O3–DME reaction system.a mol−1/2 (m3)3/2 kg−1 s−1.b mol−1 (m3)2 kg−1 s−1.

ig. 6. Parity plot of component concentrations in CSTR vs. those predicted byinetic model for vegetable oil hydrogenation on 0.5% Pd/Al2O3 in SC DME.

ata for the trans content is not completely good whereas thether components are well predicted with the kinetic model.he explanation to this deviation could be the lack of analyticalrecision.

483 513

4.64 × 10−6 ± 1.77 × 10−6 6.00 × 10−6 ± 8.78 × 10−7

1.01 × 10−5 ± 3.56 × 10−6 1.81 × 10−5 ± 3.89 × 10−6

1.52 × 10−5 ± 2.16 × 10−6 1.97 × 10−5 ± 2.08 × 10−6

3.11 × 10−5 ± 2.94 × 10−6 6.26 × 10−5 ± 4.21 × 10−5

3.63 × 10−5 ± 2.40 × 10−5 2.42 × 10−4 ± 4.78 × 10−5

3.93 × 10−4 ± 1.83 × 10−6 8.26 × 10−4 ± 2.37 × 10−5

1.91 × 10−4 2.60 × 10−4

398 A. Santana et al. / J. of Supercritical Fluids 41 (2007) 391–403

Table 7Kinetic parameters for hydrogenation of sunflower oil using 0.5% Pd/Al2O3 ascatalyst and DME as solvent

Reaction Ei (J/mol) Ai (mol−1/2 (m3)3/2 kg−1 s−1)

21 19,283 5.56 × 10−4

22 64,069 67.9611 19,025 1.71 × 10−3a

12 106,382 5.64 × 10+6a

O 152,444 8.97 × 10+11

E

ec

3P

tanhtls

sa1HrioaatT

ndttw

TE

R

1234567

O

Table 9Fitted kinetic parameters values for 2% Pd/C–DME system at 483 K and 20 MPa

k21a 4.92 × 10−4 ± 1.08 × 10−5

k22a 1.75 × 10−4 ± 1.08 × 10−5

k11b 7.87 × 10−4 ± 2.03 × 10−5

k12b 7.90 × 10−5 ± 7.68 × 10−5

kOa 5.66 × 10−4 ± 6.56 × 10−6

kEa 2.94 × 10−3 ± 5.63 × 10−5

χ2 4.85 × 10−4

rirt

o0htca

auTTpFtwb

3selectivities

Table 10 establishes a comparison between the estimatedrate constants for three different sunflower oil hydrogena-

194,615 1.09 × 10+17

a mol−1 (m3)2 kg−1 s−1.

Table 7 shows the activation energies Ei and the pre-xponential values A obtained from a plot of ln ki versus 1/T,orresponding to the Arrhenius type of dependence for the ks.

.2. Estimation of kinetics of oil hydrogenation on 2%d/C with DME as reaction solvent

In order to establish an appropriate discussion about how theype of supported Pd catalyst employed in the hydrogenationffects the reaction conversion and product distribution, it wasecessary to make a rough kinetic estimation of sunflower oilydrogenation on 2% Pd/C catalyst using DME. As the sameime, this kinetic data allow to make a comparison with a pub-ished work [31] in order to determine the effect of the reactionolvent on the oil hydrogenation.

In this case, a small set of experiments was carried out. Theystem pressure, the temperature, the molar oil concentrationnd the stirrer speed were set at 20 MPa, 483 K, 1 mol% and31 rad/s, respectively. The ranges for the experimental variables2 mol content and WHSV were 4–8 mol% and 40–975 h−1,

espectively. The catalysts was crushed and sieved to 0.47 mmn order to reduce internal mass transfer limitations. The amountf catalyst was varied between 0.25 and 1 g in order to achieveconversion degree similar to the previous case study as well

s to obtain the desired WHSV. The results from the seven runs,ogether with the respective reaction conditions are presented inable 8.

As can be noticed in Table 8, the effect of H2 is not sig-ificant either on the rate of hydrogenation or on the product

istribution. For the same degree of conversion (final IV = 102),he trans-mono-unsaturated and saturated contents are similar tohose found for sunflower oil hydrogenation on 0.5% Pd/Al2O3hich was previously studied. This feature indicates that the

able 8ffect of reaction conditions on sunflower oil hydrogenation

un H2 (mol%) WHSV (h−1) IV trans (wt%) C18:0 (wt%)

4 975 113.3 2.8 12.18 975 117.6 1.5 11.24 164 101.6 4.7 9.14 327 111.7 3.3 18.74 491 116.2 2.7 15.34 655 116.3 2.6 13.84 41 49.5 2.1 11.4

il concentration: 1 mol%; 483 K; 20 MPa; 2% Pd/C; solvent: DME.Fk

a mol−1/2 (m3)3/2 kg−1 s−1.b mol−1 (m3)2 kg−1 s−1.

eaction mechanism is the same for both catalysts. Therefore,t is possible to assume in this case that reaction order withespect to overall hydrogen consumption is 0.5 and 1 for theriglycerides.

Another fact to remark is the amount of catalyst needed tobtain a similar degree of conversion. In the case of the eggshell.5% Pd/Al2O3 catalyst, this amount is approximately five timesigher than that for uniform 2% Pd/C. A possible explanationo the less hydrogenation rate obtained in the case of eggshellatalyst is that a catalyst with a uniform metal location is morective at high pressures than an eggshell catalyst.

The kinetics for sunflower oil hydrogenation using 2% Pd/Cs catalyst and DME as reaction solvent, were investigatedsing the same approach than that in the former case study.he obtained results for the kinetic parameters are presented inable 9. Fig. 7 shows a comparison between experimental andredicted fatty acid concentrations. As can be observed fromig. 7, the accuracy data for the hydrogen content relatively less

han those of the other components, these being better predictedith the kinetic model. The explanation to this deviation coulde the lack of analytical precision.

.3. Comparison between different solvents, catalysts and

ig. 7. Parity plot of component concentrations in CSTR vs. those predicted byinetic model for vegetable oil hydrogenation on 2% Pd/C in SC DME.

A. Santana et al. / J. of Supercritical Fluids 41 (2007) 391–403 399

Table 10Estimated rate constants per kg Pd for sunflower oil hydrogenation on supported Pd catalysts using DME or propane as SC solvent (at 483 K and 20 MPa)

Parameters DME Propane

0.5% Pd/Al2O3 2% Pd/C 2% Pd/C

k21a 9.27 × 10−4 ± 3.53 × 10−4 2.46 × 10−2 ± 5.39 × 10−4 9.17 × 10−3 ± 6.09 × 10−4

k22a 2.01 × 10−3 ± 7.12 × 10−4 8.75 × 10−3 ± 5.40 × 10−4 1.53 × 10−4 ± 1.58 × 10−5

k11b 3.05 × 10−3 ± 4.31 × 10−4 3.93 × 10−2 ± 1.01 × 10−3 3.04 × 10−2 ± 1.27 × 10−4

k12b 6.22 × 10−3 ± 5.88 × 10−4 3.95 × 10−3 ± 3.84 × 10−3 9.94 × 10−4 ± 2.48 × 10−4

kOa 7.25 × 10−3 ± 4.79 × 10−3 2.83 × 10−2 ± 3.28 × 10−4 2.63 × 10−2 ± 2.614 × 10−3

kEa 7.85 × 10−2 ± 3.65 × 10−4 1.47 × 10−1 ± 2.82 × 10−1 1.09 × 10−1 ± 1.41 × 10−2

�2 1.91 × 10−4 4.85 × 10−4 2.28 × 10−5

−rH2c 0.88 1.1 0.99

tP

svisccfos

s2

hcsheb

c

o

S

ies(

tr

S

os

TL

C

2111SSTTT

S

a mol−1/2 (m3)3/2 kg Pd−1 s−1.b mol−1 (m3)2 kg Pd−1 s−1.c Total H2 uptake rate in mol s−1 kg Pd−1.

ion systems: 0.5% Pd/Al2O3–DME, 2% Pd/C–DME and 2%d/C–propane [31].

From the table presented above, it can be noticed that for theame type of catalyst (2% Pd/C), the estimated rate constantsalues are higher in the case of DME as reaction medium thann propane whereas the isomerization rate constants are about theame in both solvents. On the other hand, for the same operatingonditions using DME as reaction solvent, the estimated rateonstants for 2% Pd/C catalyst are somewhat higher than thoseor 0.5% Pd/Al2O3. In the last row of Table 10, a comparisonn the relative global hydrogenation rates is given. The rates areimilar with a slight advantage to the 2% Pd/C in DME.

The net rate of hydrogen uptake per active metal weight atimilar conversion (around 20%) is slightly higher in the case of% Pd/C catalyst irrespective of reaction solvent (see Table 10).

The overall linoleate selectivity (SLo) of the catalyst is definedere as the ratio of the rate of formation of cis-mono-unsaturatedompound from linoleate to the rate of formation of saturatedpecies (stearate). Following the reaction network of gas–liquidydrogenation, Elbid and Albright [42] proposed an equation to

valuate linoleic selectivity, that is rather simplified, that cannote applied to our case.

In the case that the mechanism is available, the selectivityalculated from the different kinetic equations, based on ratios

il

t

able 11inoleic selectivity and trans yield in the partial hydrogenation of vegetable oils in li

atalyst (dp) Process

5% Ni–silicaa Batch/slurry% Pd/deloxan AP II (0.18–1.8 mm)a Continuous trickle bed% Pd/deloxan AP II (0.18–1.8 mm)a Continuous fixed bed% Pd/deloxan HK I (0.1–0.4 mm)a Continuous fixed bedupported Pdc Harrod Res. AB Continuous fixed bed (rapesupported Ptd Harrod Res. AB Continuous fixed bed (soybhis group 2% Pd/Cb (0.55 mm) (Degussa AG) Continuous recycle reactorhis group 2% Pd/Cb (0.47 mm) (Degussa AG) Continuous recycle reactorhis work 0.5% Pd/Al2O3 (J. Matthey) Continuous recycle reactor

i = −100(�Ctrans/�(IV)), SLo = �CO/�CS.a Tacke et al. [14].b Ramirez et al. [31].c Harrod et al. [52].d idem (personal communication).

f rates is [54]:

Lo = L hydrogn. → Oleic

O hydrogn. → Stearic= k21CL

k11CO√

CH + k12CE√

CH

(8a)

This indicates that the selectivity decreases with increas-ng hydrogen concentration, but not proportionally to it. Forxample, if the H2 mole fraction at the reactor is doubled, theelectivity will not be reduced by 50%, but only a small fraction30% at most). As far as we can see, it is rather insensitive.

If the mechanism and the kinetic coefficients are not known,he linoleic selectivity can still be computed from [59] as theatio of oleic increase to stearic increase, in mol O/mol S, as:

Lo = �FO

�FS= �CO

�CS(8b)

Eq. (8b) is easier to apply than Eq. (8a), and can be used basedn the GC/HPLC analyses of inlet and outlet oil. Note that forome runs of Table 4 this selectivity is negative, because oleic

s hydrogenated to stearic more rapidly than it is formed frominoleic. For the runs in Table 4, SLo = 2–19.

Another parameter is the Specific Isomerization defined ashe ratio of the percentage increase in trans-isomer content and

quid and supercritical fluids (sunflower oil except where noted otherwise)

T (K) Pressure (MPa) SLo Si

393 0.3 10.8 Large333 2 2.1 18333 10 (CO2) 3.0 10333 10 (CO2) 13.1 29

eed oil) 323 20 (DME) 7.7 26ean oil) 298 20 (Butane) 1.7 3.1

483 20 (C3H8) 4 15483 20 (DME) 2 14456 20 (DME) 19.2 4

4 critica

t

S

wb

S

a

3

tonnotf

bichtbeaihsd

csApr[tD

r(w

cS

wftIicsTbelilFtc

dgIslStisp

3m

tavmli

a4

TV

TPFSC

00 A. Santana et al. / J. of Super

he decrease in IV by Coenen [45]:

i = trans C18 : 1 wt%

IV0 − IV(9a)

hich is equivalent (by a factor of 100) to the trans yield [59]ased on the iodine value reduction, and is written [52] as:

i = −100�CE

�(IV)(9b)

The values of the Si of Table 4 range from Si = 2 to 23, withminimum of 2.

.4. Summary and comparison

Table 11 presents the typical SLo and Si values obtained inhis study along with those published by several authors. Onef the problems is the basis of comparison that sometimes isot known. In batch hydrogenation reactions with supportedickel, the linoleate selectivity is high whereas a large amountf undesirable trans-mono-unsaturated specie is formed in theriglyceride molecule as a consequence of hydrogen mass trans-er control. This is carried out in gas–liquid reactors.

The continuous hydrogenation with Deloxan AP II in a trickleed reactor results in a significant decreased in cis/trans isomer-zation. However, the linoleate selectivity is very low. In thease of continuous hydrogenation in a fixed bed, the increase inydrogen partial pressure results in an improved linoleate selec-ivity keeping a low cis/trans isomerization value. This fact coulde explained from the presence of supercritical CO2, which low-rs the viscosity of reaction medium and increases mass transfernd diffusivity. With Deloxan HK I/Pd, the linoleate selectiv-ty is further increased. In comparison to the commercial batchydrogenation, this catalyst in combination with a supercriticalolvent gives a higher linoleate selectivity and a significantlyecreased cis/trans isomerization rate.

The continuous hydrogenations with uniform Pd/C in a recy-le reactor results in a higher linoleate selectivity with a similarpecific isomerization value (Si) in comparison to the DeloxanP II catalyst. A possible explanation is the use of small sizearticles in combination with supercritical propane or DME aseaction solvent, which enhances the solubility of fats and oils15–19,46]. Between both solvents, the selectivities values inhe case of propane as reaction solvent are better than those inME.

In continuous hydrogenation with eggshell Pd/Al2O3 in a

ecycle reactor, the linoleate selectivity value is very goodSLo = 19.2), better than that of Deloxan HK I/Pd (SLo = 13.1)ith a very good Isomerization Index (Si = 4). Table 4 indi-

Prsa

able 12ariables of reaction for simulation of vegetable oil hydrogenation on supported Pd c

emperature 483 Kressure 20 MPaeed composition (oil:H2:solvent) 1:4:95 molvent C3H8, Datalyst 2% Pd/

l Fluids 41 (2007) 391–403

ates that for 0.5% Pd/alumina in DME, SLo = 2–23 and maxi = 4.

The selectivity values for eggshell Pd/Al2O3 in combinationith DME as solvent are significantly better than those for uni-

orm Pd/C irrespective of the solvent. The possible reason tohis behaviour is the metal location along with the particle size.n eggshell catalyst, only the surface reaction is rate limitingndependently of the particle size whereas in the case of Pd/Catalyst, its uniform metal location as well as the large particleize could lead to the diffusion limits the overall rate of reaction.his fact are in concordance, at low pressure, with that reportedy Tsuto et al. [47], Coenen [45], Westerterp et al. [48], Colent al. [49] and Veldsink et al. [50] who have found that in theow pressure hydrogenation of edible oils on porous catalyst,ntraparticle mass transfer limitations not only reduce the cata-yst effectivities but also may change the product selectivities.or high pressure, only our group has given a complete study on

he intraparticle diffusion mechanisms in SCF in pore diffusionontrol regime [54].

In Table 11, the best system as regards to minimal trans pro-uction, is the supported Pt catalyst, tried for soybean oil hydro-enation. The isomerization parameter is not very high, so theV should be kept low, as that selectivities go down with conver-ion. Not much experience is available for such systems. In fact,inoleate selectivity is best with Pd (rapeseed oil hydrogenation)Lo = 7.7, but the trans selectivity is Si = 26. It is worth notinghat at room temperature, either on Pd or Pt, the solvent is subcrit-cal so it is in compressed liquid form (20 MPa), and hydrogenhould be in the form of bubbles. The data of Harrod on suchrocess (Table 11) are very interesting and suggest new ideas.

.5. Experimental trans and stearic data and check ofodel

For the estimated rate constants for sunflower hydrogena-ion, a calculation was performed in order to study the stearicnd trans fatty acid concentration profiles as a function of con-ersion (expressed as final IV) in an isothermal continuous wellixed-flow reactor using an non-linear solver [51]. The simu-

ated reaction conditions are given in Table 12, and the resultsn Figs. 8 and 9.

The mol feed composition (oil:H2:solvent), system pressurend temperature were kept constant at 1:4:95%, 20 MPa and83 K, respectively. The particle size was set at 0.47 mm for 2%

d/C catalyst and 2 mm for 0.5% Pd/Al2O3. Fig. 8 presents theesults both as experimental data points and model results. Forhortening and margarines applications, the target iodine valuesre in the order of 70–100. With this in mind, it could be noticed

atalyst using SCFs as solvent

ol%ME

C (dp = 0.47 mm, 5 g), 0.5% Pd/Al2O3 (dp = 2 mm, 25 g)

A. Santana et al. / J. of Supercritica

Fig. 8. Measured and simulated trans and stearate vs. IV. Continuouslines are from mixed-flow reactor simulation [60] (20 MPa, 483 K, molarfWs

fsivitttPb

FoI

Fatt1c

faciwvtfaow

4

sfPptP

scs

eed oil:H2:SC solvent 1:4:95%, WHSV0.5% Pd/Al2O3 = 10–100 h−1 andHSV2% Pd/O = 50–500 h−1): (a) trans fatty acid composition profiles and (b)

tearic increase with conversion.

rom the figure presented above that the supercritical reactionolvent, which presents the least trans fatty acid content is DMErrespective of the Pd catalyst employed. Fig. 8a gives the modelalues and the experimental values for trans for the IV values. Its clear that both model and experiment the best trans values arehose of 0.5% Pd/alumina in DME, from 456 to 456 K. These

rans values are below 3.5%. For 2% Pd on C in DME a similarrend is seen but trans is a little higher (5% at IV = 100). For 2%d/C in propane (448–458 K) a trans value of about 5% woulde obtained.

ig. 9. Experimental and predicted fatty acid composition profiles as a functionf space-time (20 MPa, 483 K, 2% Pd/C, molar feed composition 1:4:95%, finalV = 50, well mixed-flow reactor, Eqs. (1)–(6)). W/F in kg cat kg−1 feed s−1.

svo

ttsstcm

safsff

A

D

l Fluids 41 (2007) 391–403 401

In Fig. 8b, the newly formed stearic, for the same runs ofig. 8a, is presented. The model lines indicated the same trends before: the best combination is the 0.5% Pd/Al2O3 in DME,hen 2% Pd/C in propane, and finally the worst case would behat of 2% Pd/C in DME. Note that the best system gives only.5 new stearate (S = 10.5%) over stearate in for IV = 100. Thisorresponds to 0.5% Pd in DME (456 and 483 K).

From this discussion, it is evident that the best combinationor sunflower oil hydrogenation in order to obtain a consider-ble conversion with less trans content seems to be the 2% Pd/Catalyst with DME as SC solvent. But if a linolenic selectivitys desired, so that less stearic is produced, the best combinationould be an eggshell 0.5% Pd/Al2O3 catalyst in SC DME sol-ent. Fig. 8a and b substantiate these findings. Fig. 9 depictshe fatty acid concentration profile along with experimental datarom a laboratory-scale reactor as a function of the space-time,nd as can be seen, the accuracy of data prediction for the mostf the oil component is fairly good. An operating space-timeould be about 20–30 s, for the Pd/C catalyst.

. Conclusions

In this work, we report in the fluid phase hydrogenation onunflower oil in dense DME with 1 mol% of vegetable oil in TGorm, 4–9% of H2 and 90–95% of DME on several supportedd catalysts. The reaction is run continuously in a single-fluidhase, using a laboratory setup with a Robinson–Mahoney reac-or. The catalysts are uniform 2% Pd/C and eggshell 0.5%d/Al2O3.

In comparison to commercial batch hydrogenation with aupported catalyst, the 0.5% Pd/Al2O3 catalyst with supercriti-al DME as a solvent gives a similar linoleate selectivity and aignificant decreased specific isomerization value.

However, these overall promising facts have been observedo far only at higher iodine values (IV > 90) because target iodinealues for shortening and margarine applications are in the orderf 60–90.

For the estimated rate constants for sunflower oil hydrogena-ion, a simulation was performed in order to study the stearic andrans fatty acid concentration profiles as a function of conver-ion in an isothermal continuous well mixed reactor. The resultsuggest that the best combination for sunflower oil hydrogena-ion in order to obtain a considerable conversion with less transontent seems to be the 2% Pd/C catalyst with DME as reactionedium.The results presented in Table 11 on the catalyst based on

upported Pt, for the hydrogenation of soybean oil, are encour-ging in the sense that the trans could be completely eliminatedorm the process. At the same time, however, the production oftearic esters is also increased so there is a need to limit also itsormation. Also, non-diffusion limited catalyst will improve theormation of stearate.

cknowledgements

We acknowledge the donation of free catalyst samples fromegussa AG (Germany) and Johnson Matthey (Barcelona). A

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02 A. Santana et al. / J. of Super

rant from the FI Program (Generalitat de Catalunya, Barcelona)o E. Ramirez and one to A. Santana from the FPI ProgramSpanish Ministry of Science and Technology, Madrid), areppreciated. Financial support to the project was received fromICYT-FEDER (Madrid-Brussels), project grant no. AGL2006-5156.

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