bleaching kinetics

© WILEY-VCH Verlag GmbH, 69451 Weinheim, 2002 0931-5985/2002/0202-0098 $17.50+.50/0

98 Eur. J. Lipid Sci. Technol. 104 (2002) 98–109

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1 Introduction

1.1 Aim of this study

The bleaching stage of the refining process of edible oils,which consists of bringing the oil into contact with an ad-sorbent mineral clay, is regarded as the most importantstep to achieve and to guarantee an excellent quality ofthe final product. It is performed at temperatures of about100 °C and removes different undesired substances byadsorption and catalytic reaction. Since these substancesare primarily pigments like carotenoids and chlorophylls,the oil is lightened up – giving this process step its name.Beside pigments the clay also removes other impuritiessuch as soap, trace metals and phosphatides and re-duces the oxidation levels by catalysing the breakdown ofhydroperoxide primary oxidation-products. As a conse-quence, the quality of the oil is stabilised by bleaching,since precursors and catalysts of autoxidation are re-moved. The position of the bleaching step in the refiningprocess depends on the kind of deacidification used. Incase of chemical refining bleaching is placed after thechemical neutralisation and in case of physical refiningbefore the distillative deacidification.

The intention of this study is to optimise the bleachingprocess, focussing on the enhancement of mass transferand to reduce the running costs of the oil refineries.

During bleaching different impurities are removed simul-taneously by different, partly interfering mechanisms. Ac-cordingly the bleaching step is quite complex and difficultto optimise, since many parameters and componentshave to be taken into consideration and several thresholdvalues have to be reached [1, 2]. Furthermore the waver-ing quality of the crude oil complicates optimisation, whichis usually performed empirically in the oil mills. Generally,the goal of optimisation is to achieve a defined stabilityand transparency of the oil in a minimum of reaction timeusing as few clay as possible. The latter is desired sincethe clay effects the processing costs in manifold ways, in-cluding the costs of the clay itself, the costs for disposal ofspent clay, the running costs of filtration, and the clay in-duced oil loss from the process. This is based on the factthat spent bleaching clay contains about 30 wt-% of oil,which can only be recovered as oil of deteriorated quality.

To optimise the bleaching of vegetable oils several strate-gies are followed. One strategy is the optimisation of theclay itself, (capacity, pore structure, acidity, particle size)which is followed by the manufacturers of the clay. Thestrategy of the oil mills is to improve the use of the capac-ity of the clay by countercurrent bleaching or the empiricalenhancement of mass transfer between particle and oil.For economic reasons the process is not run until steadystate is reached in the reactor, since the closer the systemgets to the equilibrium, the smaller becomes the drivingforce, and consequently the slower the kinetics. The resi-dence time in the reactor is shortened by using more claythan necessary from the equilibrium point of view and bymaking use of the so-called ’post bleach effect’ takingplace during the sequencing filtration [3]. This strategy

Thies Langmaack, Rudolf Eggers

Technische UniversitätHamburg-Harburg, Verfahrenstechnik II,Wärme- und Stofftransport,Hamburg, Germany

On the bleaching kinetics of vegetable oils –experimental study and mass transfer-basedinterpretationThis study focuses on the systematic investigation of the parameters involved in ad-sorptive vegetable oil bleaching. These parameters range from the quality of oil andbleaching conditions chosen (e.g. amount and activity of clay, initial water content,temperature, pressure) to the effects of the different reactor systems used (agitatedvessel, bubble column) on the kinetics. The study of the latter aims to optimise the typeand intensity of power input. The most efficient process was shown to take place in thebubble column reactor. The findings result in a kinetic model for the transport phenom-ena of the pigments from the oil phase into the clay. The model splits the effects in phe-nomena due to mass transfer resistances inside and outside of the clay particle. By do-ing so it is possible to describe the bleaching kinetics in dependence of the amount ofclay added and on the size of the external mass transfer resistance, which is influ-enced by the power input into the reactor. Thus a on model-based analysis of the ef-fects as well as the optimisation of the bleaching reactor are made possible.

Keywords: Bleaching, rapeseed oil, adsorption, refining, reactor design.

Correspondence: Rudolf Eggers, Technische Universität Ham-burg-Harburg, Verfahrenstechnik II, Wärme- und Stofftransport,Eisendorfer Strasse 38, 21073 Hamburg, Germany. Phone: +49-40-42-878-3191, Fax: +49-40-42-878-2859; e-mail: [email protected]

could be optimised further by enhancing the externalmass transfer particle-pigment resulting in a reducedneed of clay or time. By enhancing the external masstransfer the clay is loaded more effectively. Thus in thesame time the same reduction in pigment in the oil can beobtained by using less clay. To enhance the rate of exter-nal mass transfer the key tool is the operating mode of thereactor to maintain a high slip velocity between particleand liquid. These flow conditions around the particle de-termine the thickness of the boundary layer around theparticle representing the external mass transfer resis-tance to the particle. They depend on the agitation systemand the power input used. Therefore it should be possibleto reduce the amount of clay and the residence time tosave running costs by optimising the amount and mannerof power input.

To calculate an economic optimum for this proceeding thedominating factors influencing the kinetics have to beknown, ranging from the effect of the external mass trans-fer rate to the effect of parameters like amount of clayused. To separate these interfering effects modelling workwas performed on one hand to describe the wholeprocess of adsorptive bleaching in a proper way and onthe other hand to compare the influence of the externalmass transfer on the kinetics with the internal mass trans-fer in the clay particle.

1.2 Theoretical background

Although the pigments of the oil are removed by severalinterfering reactions and mechanisms beside adsorption,the removal of pigments during bleaching can be de-scribed as an adsorptive equilibrium process following theFreundlich isotherm [2]. The final colour of the oil, respec-tively content of pigments, depends mainly on the amountof clay added.

The kinetics of bleaching have been investigated by sev-eral authors. The fundamental work was published byBrimberg [4, 5]. She described the kinetics of pigment re-duction in a linearised form when plotting the concentra-tion of remaining pigments against the square root ofbleaching time:

(1)

Accordingly the reduction of pigments takes place in threephases, characterised by the kinetic constants k1, k2 andk3, whereof the first step lasts the longest and reducesmost of the pigments. In further investigations Brimbergvaried the parameters activity of clay (natural or activatedclay), amount of clay, size of clay particles and amount of

ln* * *

.CC

k t t tk t t t tk t t t

with k k k0

1 1

2 1 2

3 2

1 2 3

0

=− < <− < <− <

> >

Eur. J. Lipid Sci. Technol. 104 (2002) 98–109 Kinetics of vegetable oil bleaching 99

initial water and showed how these parameters influencethe constants ki. ki was found to increase nearly propor-tionally with the amount of bleaching clay added; itincreased with decreasing particle diameter, increasingactivity of clay and in a certain area even with increasinginitial amounts of water when bleaching with minor acti-vated clays. Summarising the kinetics can be describedas a function of bleaching time with this model. The draw-backs of this approach are the impossibility to predict theduration of the bleaching phases t1 and t2 and to explainthe size of the kinetic constants with a physical model.

To circumvent this drawback, a different approach wasdeveloped by Brat et al. [6]. They determined the kineticsof bleaching in dependence of the amount of bleachingclay “a” and successfully correlated the kinetics using thefollowing approach

(2)

where the constants p1, p2, p3 and p4 are found by fittingthe measured data. The advantage of this approach isthat an algebraic solution for the whole reaction time isgiven and that the effect of the amount of clay used canbe calculated directly using this equation. Although the ki-netics can be described as function of time and amount ofclay with this model, it has some drawbacks. These arethat this approach is based on empirical fitting of the pa-rameters pi, a direct physical background of the parame-ters used is missing and an estimation or extrapolation ofthese parameters is not possible.

In this work a further generalisation is achieved yieldingthe expression of kinetics as function of time, amount ofclay and mass transfer between particle and pigmentbased on a physical model describing bleaching as ad-sorptive removal of pigments from oil by clay. This modelis based on the basic considerations on pure adsorptionon porous particles in liquid batch systems. In these sys-tems the mass transfer takes place in several steps:

– diffusive transfer of pigment from the bulk of the liquidthrough the boundary layer to the surface of the parti-cle,

– diffusive transfer of the pigments into the particle bysurface diffusion,

– adsorption of the pigments on the active sites at thesurface of the particle.

Hereof the last step has been shown to happen veryrapidly, thus the transfer to and in the particle being therate limiting steps. These different consecutive transportmechanisms can be combined to differential equationsand boundary conditions and solved numerically. Theseequations are similar to those equations describing thebatch treatment of water with powdered activated carbon

− =dcdt

p a cp p ap

12 3 4( ),

[7, 8]. To apply these models on bleaching systems someof the parameters as diffusion and mass transfer coeffi-cients for the system carotene-oil-small clay particles hadto be determined, since they are not known from litera-ture. This was done in this study using matching calcula-tion, where calculated kinetic curves are fitted to mea-sured kinetics curves by matching the internal diffusioncoefficients and external mass transfer coefficients.

2 Materials and methods

2.1 Oil

Rapeseed oil was used for all experiments. The main partof the experiments was performed using chemical neu-tralised rapeseed oil (Thywissen, Neuss, Germany;0.065 wt-% free fatty acids, 4.9 ppm phosphor, 0.082 ppmiron, 55 ppm carotene, 3.5 ppm chlorophyll) in this studyindicated as oil “A”. Since many oil mills are changingtheir refinery towards physical refining, some experimentswere done using a degummed rapeseed oil “B” preparedfor physical deacidification (Broekelmann, Hamm Ger-many; 0.64 wt-% free fatty acids, 50 ppm carotene,3.6 ppm chlorophyll). For some experiments a complete-ly refined rapeseed oil “C” (Broekelmann) enriched withsynthetic carotene (BASF, Ludwigshafen, Germany) wasused. In some cases this oil was further enriched withlecithin (Lucas Meyer, Hamburg, Germany).

2.2 Clay

In nearly all experiments the highly activated clay TonsilOptimum 210 FF (Suedchemie, Munich, Germany) was

used. Its particle size ranged from 1 to 180 µm. The latterwas determined for the coarse fraction by sieve analysisand for the fine fraction by laser diffraction analysis. TheSauter mean diameter calculated from this analysis wasfound to be 16 µm.

For the examination of the effect of the initial content ofwater in the oil on the kinetics in dependency on the ac-tivity of the clay used, the natural clay Terrana D (Sued-chemie) and the medium activated clay Tonsil StandardFF (Suedchemie) were used.

2.3 Reactor systems

As reactor systems a stirred vessel and a bubble columnreactor were used and run at various conditions. A sketchof these reactor systems is given in Fig. 1.

The bubble column reactor was made from glass, with aninner diameter of 0.06 m, a height of 1.6 m and a porousor perforated plate as bottom. It was heated electricallyand was run using nitrogen, carbon dioxide or superheat-ed steam. The gasflux was controlled, so that the superfi-cial velocity could be varied between 0 and 0.19 m/s andthus the reactor could be operated either with homoge-neous bubble flow (< 0.02 m/s), heterogeneous flow(< 0.1 m/s) or slug flow. The specific power input P(W/kgoil) in this reactor was calculated using equation 3[9] as directly proportional to the superficial gas velocityu0 (m/s):

P = g u0 (3)

100 Langmaack et al. Eur. J. Lipid Sci. Technol. 104 (2002) 98–109

Fig. 1. Sketch of the reactorsused. The table imbeddedgives some characterising da-ta of the impeller (Newtonnumber at 144 and 730 rpm in100 °C hot oil). The pitched-blade-stirrer is abbreviated asPBS.

The power input into the bubble column was varied be-tween 0 and 1.9 W/kgoil. The gas holdup was determinedby measuring the increase of the dispersed bed and wellpredicted using the relation by Bach and Pilhofer [9] foroily systems. The liquid phase dispersion was followedusing the steady-state method proposed by Deckwer [9]using carotene as tracer.

The cylindrical stirred vessel reactor was made fromstainless steel, having an inner diameter of 0.13 m, aheight of 0.38 m and was heated electrically. The impellerspeed could be varied between 144 to 730 rpm. Thefollowing impeller systems have been used: a Rushtonturbine (0.038 m), an anchor stirrer (0.0965 m) and ashaft equipped with one to three 2-45° pitched bladestirrers (0.08 m). The power input of these systems in thereactor in dependency of the impeller speed was deter-mined using the test station of TU Braunschweig [10] de-tecting the torsion of the impeller shaft. The power inputachieved increases with the speed of rotation, with the di-ameter of the stirrer and the number of stirrers used onthe shaft. From these measurements the Newton-num-bers characterizing the power input of the impeller in100 °C hot oil were calculated using equation 4:

(4)

with the specific power input P (W/kgoil) into the mass ofliquid ml (kg) with the density ρl (kg/m3) and the impellerwith diameter d (m) rotating at a speed n (1/s). The New-ton numbers obtained are given in Fig. 1. By varying thetype of impeller and the rotational speed a broad range ofpower input/turbulence was investigated, ranging from0.001 to 6,5 W/kgoil. In some cases the vessel was baffledto avoid the formation of a vortex at the shaft.

2.4 Methods

Batches of 2.7 kg oil were used per bleaching experi-ment. To perform a bleaching experiment the oil washeated up and simultaneously degassed and dried undervacuum. When the bleaching temperature of 98 °C wasreached, water (0.25 wt-%) and citric acid (0.06 wt-%)were added. The latter is necessary in order to complexand precipitate trace metals, water has been shown toenhance the kinetics additionally. Subsequently thebleaching was started by adding powderous clay (0.2 to1.0 wt-%). To avoid any contact of the clay-oil-system withoxygen and to maintain the water in the system the reac-tor was kept under an inert gas blanket at a constant pres-sure of 1 bar for the first five minutes after the addition ofthe clay. In the stirred tank reactor the pressure was thenreduced to a vacuum of 52 mbar, in contrast to the bubblecolumn reactor, where it was kept constant at 1 bar. From

NeP m

n dl= 3 5 ρ l

the start of the bleaching period, samples were withdrawnfrom the reactor in intervals of 0.5 min for 25 min. To avoidany contact of the hot oil/clay system with air and to stopthe bleaching process, the clay was separated immedi-ately from the oil by filtration when taking the sample via avacuum-filter system.

2.5 Analysis

The concentrations of chlorophyll and carotene were de-termined photometrically according to the method ofBrimberg [4] using 1-cm cuvettes and a spectrophoto-meter (Shimadzu UV-120-02, Shimadzu, Kyoto, Japan).Chlorophyll was detected at a wavelength of 660 nm,carotene at 445 nm. At high carotene concentrations thesample was diluted in ratio 1:10 (w/w) with isooctanol toremain in the reliable detection range of the spectropho-tometer. The colour according to the Lovibond chart wasadditionally determined for several samples in 1” and in5.25”-cells using a Lovibond Tintometer-PFX 990 (Tin-tometer, Dortmund, Germany). Additionally some sam-ples were deodorised in lab scale equipment (150 ml)and thereafter the colour according to the Lovibond chartwas determined. The Rancimat stability was determinedat 100 °C and an air flow of 20 l/h.

3 Results and discussion

The dependency of the bleaching kinetics on bleachingtemperature, pressure, quality of feed material, amount ofwater, amount of clay and conditions in the reactor sys-tems (type and speed of stirrer and type and velocity ofgas in the bubble column reactor) were determined. Inthis study the pigments of the oil were chosen as keycomponents, since it was shown, that when these pig-ments were removed from the oil, the other impurities hadbeen removed properly as well. In this case quality andstability of the oil after deodorisation were excellent.

3.1 Adsorption isotherms

The adsorption isotherms of carotene and chlorophyll cal-culated from values measured at steady state were foundto follow the Freundlich form as frequently mentioned inliterature [2].

3.2 Particle size

To evaluate the effect of the process conditions on theparticle size of the clay, the size of the particles was de-termined using a photosedimeter. With this apparatus thesedimentation of the particles in oil at ambient conditionswas observed. This method was chosen, as it allows the particles to remain in the original medium and thusadulteration of samples is avoided. Even intensive power


input by intensive stirring showed no significant effect on the size distribution. So neither conglomeration of theparticles nor destruction by abrasion seemed to have occurred during bleaching.

3.3 Wettability

The wettability of the clay by oil was determined in orderto estimate flotation effects in the bubble column reactorand to characterise the immersion behaviour of the clay inoil. The Washburn method for powders [11] was used de-tecting the velocity of liquid penetration induced by capil-lary forces into a sample of powder. By comparing the ve-locity detected for oil with the velocity detected for com-pletely wetting hexane, the wetting angle can be calculat-ed. The oil was found to wet the clay well at 98 °C, thewetting angle was calculated to be 34°. Because of thisgood wettability flotation effects are unlikely to occur inthe bubble column reactor, and in consequence the im-mersion takes place very fast.

3.4 Kinetics

In the following the findings concerning the kinetics ofbleaching are discussed. It was shown that the caroteneremoval follows nearly the same kinetics as the removalof chlorophyll. This effect also was stated by Brimberg [4].Therefore only the data found for carotene are presentedhere. To minimise the amount of data and graphs in thispublication these effects are summarised by discussingthe effect on the kinetic constant k1 when plotting the ki-netic curve of carotene removal according the approachof Brimberg vs. the square root of contact time. Thisanalysis is demonstrated in Fig. 2 showing the influenceof the amount of clay on the kinetics. The kinetic datagained by the kind of analysis discussed here are listed inTab. 1.

When comparing the data across the sections of Tab. 1has to consider the kind of oil that has been bleached. Especially when bleaching the model oil C (refined oil enriched with synthetic carotene), the kinetic constantsobtained are higher than those obtained with the ’normal’oils A and B at the same conditions.

3.4.1 Effect of oil quality

3.4.1.1 Physical refinement

Deacidified oil is easier to bleach than not deacidified oil.The Lovibond colours after deodorisation were lowerwhen using the same amount of clay.

The removal rate of chlorophyll is not affected. The re-moval of carotene seems to be faster in deacidified oil.

3.4.1.2 FFA addition

The variation of the initial free fatty acid (FFA) content ofthe oil by adding oleic acid showed neither a significanteffect on the Lovibond colour after deodorisation nor thedecolourisation rate. The FFA content in the oil remainednearly constant during bleaching, thus the FFA were notadsorbed by the clay and are not competing for activesites with the pigments.

3.4.1.3 Lecithin addition

The amount of lecithin present in the oil affects thebleaching rate. The more lecithin is present the slower therate. Since the lecithin content decreases during bleach-ing, it is likely that the lecithin is adsorbed by the clay andmight compete for active sites with the pigments. The de-gree of lecithin removal is influenced by the amount ofwater present in the system. If the amount of water isrisen from 0.02 to 0.25 wt-%, the removal of lecithin is en-hanced whereas the rate of decolourisation is not influ-enced significantly. The water might support the formationof micelles of lecithin which are removed from the oil dur-ing filtration.

3.4.2 The effect of the initial amount of water

Brimberg found that an enhanced initial amount of waterpresent in the system accelerates significantly the kinet-ics when bleaching rapeseed oil with Tonsil Standard FF,a medium activated clay [4]. When bleaching palm oil withTonsil Optimum FF, a highly activated clay, she could notfind an equivalent effect [5]. The same was found in thisstudy. When bleaching rapeseed oil in the stirred tank re-actor with Tonsil Optimum FF and varying the initial watercontent from 0 to 0.5 wt-% no significant effect of water onthe kinetics was found, the kinetics remained nearly con-stant.


Fig. 2. Analysis of kinetic data for batch bleaching as sug-gested by Brimberg; results from bleaching deacidifed oilwith Tonsil Optimum FF in a bubble column reactor.

In contrast when bleaching rapeseed oil in the bubble col-umn reactor using the less activated clay Tonsil Standardor the natural clay Terrana D the presence of water in-creased the rate in the magnitude of factor 1.2 to 1.5.Thus the effect of initial water seems to be related recip-rocal to the degree of activation of the clay.

Generally an accelerating effect of water on the bleachingkinetics can be explained by the fact that the presence ofwater supports the removal of soaps and phospholipidswhich would block pores of the clay otherwise. But sincethe soap and lecithin content of the oils used in these ex-periments was negligible a different effect has to be con-


Tab. 1. Results of parameter analysis during bleaching, given as the kinetic constant k1 when analysing the kinetic data asproposed by Brimberg.

Parameter varied Experimental conditions Kinetic constant k1 for carotene

Lecithin Bubble column reactor, oil C, Lecithin content [ppm]70 ppm carotene 0.6 wt-% clay initial finalwater content varied: < 25 <25 / <25 1.02 / 1.040,02 wt-% / 0,25 wt % 910 155 / <25 0.81 / 0,85

6000 682 / 270 0.64 / 0.59

FFA Bubble column reactor, oil B FFA content [wt-%]0.6 wt-% clay initial final

0.6 0.6 0.442.0 2.0 0.405.1 4.9 0.46

Water at different Stirred tank reactor, oil A Initial water [wt-%]clay activities 0.6 wt-% T. Optimum FF 0.02 0.55

0.10 0.460.25 0.520.40 0.500.50 0.51

Bubble column reactor, oil C, 0.02 0.7053 ppm carotene 0.25 0.830.6 wt-% T. Standard 0.50 0.69

Bubble column reactor, oil C, 0.02 0.05453 ppm carotene 0.25 0.0811.0 wt-% Terrana D 0.50 0.040

Pressure Stirred tank reactor, 0.6 wt-% clay, [rpm] [rpm] [rpm]various impeller speeds 144 244 500

54 mbar 0.54 0.52 0.562 bar 0.48 0.55 0.53

Bubble column reactor, [wt-%] [wt-%] [wt-%]various amount of clay 0.2 0.4 0.6

500 mbar 0.16 0.32 0.551 bar 0.16 0.37 0.62

Temperature Bubble column reactor, oil B, [°C]0.8 wt-% clay 100 0.65

83 0.6060 0.42

Amount of clay Bubble column reactor, oil A [wt-%]0.2 0.160.4 0.370.6 0.620.8 0.801.0 0.96

State of refining Bubble column reactor, [wt-%]different amounts of clay 0.2 0.4 0.6 0.8 1.0

Oil A 0.16 0.37 0.62 0.8 0.96Oil B 0.14 0.29 0.44 0.65 0.81

sidered. The different effect of initial water might be due toa different wetting of the less and highly activated clay bywater, which has to be further investigated.

From the experiments done in the bubble column reactorat ambient pressure even an upper limit for the supportingamount of water can be shown. If this level is exceeded,the removal of pigments is disturbed. When adding 0.5wt-% of water, the hot oil is no longer a single phase butdroplets of free water are formed. In these visible dropletsa part of the clay is trapped, separated from the oil andthus not removing pigments. Accordingly the rate of re-moval decreases. In the end the colour of the filter cake isgrey instead of black due to the white clay particles nothaving adsorbed pigments. The fact that this limit of0.5 wt-% of water was not observed in case of using Ton-sil Optimum FF should be due to the fact that this experi-ment was performed at vacuum conditions.

3.4.3 Effect of pressure

There was no effect of the pressure (54 mbar or 2 bar) tobe seen on the kinetics when using Tonsil Optimum FFclay in the stirred tank reactor. This might be connectedwith the fact that no effect of initial water amount could beseen when using this highly activated clay. However, it themain prerequesite seems to be that contact of the oil withoxygen is avoided, irrespective whether this is achievedby the application of vacuum or a protective layer of inertgas.

3.4.4 Effect of temperature

Bleaching at lower temperatures (tested for 83 °C and60 °C) decreases the rate of decolourisation. This reduc-tion correlates well with the enhancement of oil viscositywith decreasing temperature (being 8.5 cP at 100 °C,11.6 cP at 83 °C, 21.5 cP at 60 °C [12]) and thus the slow-down of transport mechanisms due to decreasing diffu-sion rates. The final value obtained at steady state itselfwas not found to be influenced significantly by the varia-tion of temperature in this range.

3.4.5 Amount of clay

Besides the final content of pigment in the oil the amountof clay influences the kinetics of bleaching. A duplicationin amount of clay yields a duplication in bleaching rateand thus the effect on the kinetics is higher than on the fi-nal content.

3.4.6 Effect of time

While the removal of pigments continues in an exponen-tial way till the equilibrium value is reached, the sum pa-rameters as Lovibond colour and stability of the oil after

deodorisation remain constant much earlier as can beseen from Tab. 2. Therefore it is not necessary to contin-ue bleaching until equilibrium and to remove as much pig-ment as possible to gain a certain Lovibond colour of thecompletely refined oil.

3.4.7 Effect of reactor system

To show the effect of the running conditions of the reactoron the kinetics of bleaching, Fig. 3 summarizes some ki-netic curves obtained in the stirred tank reactor systemwith different stirrers and impeller speeds and a kineticcurve gained in the bubble column reactor, using 0.6 wt-%of Tonsil Optimum clay in all cases. The same carotenecontent is reached at the end, but the way how fast it isachieved differs significantly. The bubble column reactorshows the fastest kinetics. When running the latter no sig-nificant difference in kinetics was found neither whenvarying the kind of bottom (perforated or porous plate norwhen varying the gaseous phase (carbon dioxide, nitro-gen, superheated steam).

If the desired pigment concentration to be obtained by ad-sorptive bleaching is set to be 20% of the initial value(which is the case for the oil used in this study, since suchoils lead after deodorisation to a good Lovibond colourand stability) the goal is reached in the best case within 7 min and in the worst case within more than 20 min usingthe same amount of bleaching clay. So the operationmode of the reactor has a significant influence. This effectis mainly due to the degree of suspension of the particlesand to the intensity of mass transfer between suspendedparticle and surrounding liquid. In the case of poor kinet-ics at low power input the particles seem to be not sus-pended properly or not distributed homogeneously. If thepower input exceeds a certain limit the kinetics decline.


Fig. 3. Bleaching oil A in the stirred vessel reactor using0.6 wt-% of Tonsil Optimum clay: Kinetics obtained whenvarying kind, size and rotation speed of impeller and thusthe power input. For comparison one curve obtained inthe bubble column reactor is added.

This is due to the formation of a vortex in the reactor andthe resulting inhomogeneous distribution of particles isslowing down the kinetics. When successfully avoidingthe formation of a vortex by baffling the reactor the kinet-ics did not decrease with increasing power input.

To investigate the influence of the operating mode on thekinetics the following model-based analysis was per-formed, interpreting the process as adsorptive processand splitting the mass transfer of pigments from the bulkof the oil into the clay in internal and external mass trans-fer phenomena.

3.5 External mass transfer

The external mass flux of pigments from the bulk of theliquid to the surface of the clay particle n· f (kg/m2 s) is given by

(5)

with βf [m/s] representing the external liquid-particle masstransfer coefficient, C [kg/m] the concentration of pigmentin the bulk of the liquid, Cs [kg/m] the concentration of thepigment in oil at the outer surface of the clay particle (seeFig. 4). The external mass transfer coefficient βf can be interpreted as ratio of film diffusion coefficient Df [m2/s]and thickness δ [m] of the boundary layer around the particle. From this setting it can be seen directly that thereduction of the boundary layer by proper circumflow orreduction of the viscosity should enhance the externalmass transfer rate. For the case of bleaching vegetableoil the latter can be reached by enhancing the temper-

˙ ( ) ( )n C CD

C Cf f Sf

S= − = −βδ


Tab. 2. Effect of time of bleaching on pigment removal, Lovibond colour and Rancimat stability after deodorisation.

Bleaching oil A in a After time t After time t of bleachingstirred tank reactor of bleaching and subsequent deodorisation

Amount of clay Time of bleaching t Chlorophyll Lovibond [5,25” cell] Rancimat stability[wt-%] [min] [ppm] yellow/red [h]

0.2 5 2.4 17/2.1 13.810 1.7 14/2.015 1.5 15/2.0 14.220 1.3 15/2.025 1.1 15/2.0 14.9

0.4 5 1.5 14/2.2 14.110 0.7 13/2.015 0.5 11/2.7 14.920 0.4 12/1.725 0.4 13/2.0 15.2

0.6 5 0.8 12/1.7 15.710 0.3 12/1.715 0.2 11/1.6 15.720 0.1 11/1.625 0.1 11/1.6 14.5

Fig. 4 A. Surface Diffusion Model to describe adsorptionin liquids. B. Matching measured data using the SurfaceDiffusion Model with setting the inner diffusion coefficientto be constant 3.5 E-15 m2/s and the outer resistance tobe 10 E-6 m/s.

A

B

ature or by bleaching in the presence of supercritical car-bon dioxide dissolved in the oil [13].

The size of the external mass transfer coefficient βf canbe derived from experimental data by analysing the initialkinetics if it is supposed that in this period only externalmass transfer takes place. This means that in the begin-ning pigments are only covering the outer surface of thecompletely homogeneously suspended spherical parti-cles.

Thus in the initial period of decolourisation the effect ofthe internal mass transfer resistance on the rate is ne-glected. Therefore the external mass transfer coefficientβf0 can be determined from the slope at the beginning ofthe process by graphical derivation [7, 8]. This initial valuefor βf0 was observed to be constant for the whole period ofbleaching, meaning that the thickness of the boundarylayer and the diffusion coefficient in the film remain con-stant, thus it is supposed to be equal to βf. This value βf

was determined for all batch bleaching experiments donein this study and it is found to vary between 0.3 E-6 m/sand 18 E-6 m/s and to be independent of the amount ofclay added. To correlate the values, they were plotted(Fig. 5) in dependence of the power input. This kind ofplot is frequently used in literature [9] for the correlation ofsolid-liquid mass transfer coefficients for suspended par-ticles.

A certain minimum amount of power input is necessary tosuspend the particles sufficiently. If the power input fallsbelow this value the mass transfer coefficient is quite low.The height of this minimum power input depends on thetype of impeller used and the flow structure close to thebottom of the vessel. In the bubble column reactor a high-

er external mass transfer coefficient is obtained than inthe stirred tank reactor. In the state of heterogeneousbubble flow (> 0.2 W/kgoil) the mass transfer coefficientincreases with increasing power input (respectivly super-ficial velocity), in the area of slug flow (> 1.0 W/kgoil) it de-creases again. The latter might be due to the pulsatingflow in the reactor causing an heterogeneous distributionof the clay and thus resulting in a slower kinetic. In thestirred tank reactor a certain plateau is reached when thepower input reaches 0.2 W/kgoil. A further increase ofpower input does not enhance the external mass transferfurther since it does not result in a higher slip velocity be-tween particle and fluid. Consequently there is no need torun a stirred reactor with a higher power input.

To generalise these findings on the external mass transferthe Sherwood numbers obtained in the system were cal-culated from the βf measured using equation 6:

(6)

with dp [m] as Sauter mean diameter of the particles andDf (m2/s) as film diffusion coefficient of carotene in oil, calculated to be about 6 E-11 m2/s for carotene in oil at100 °C using the Stokes-Einstein-Equation [14]. The Shnumbers ranged between 2.0 and 5.0. These experimen-tally obtained numbers were compared with calculatednumbers using power input-based correlations recom-mended in literature. The best fit was obtained using thecorrelation proposed by Sano et al. [10]:

(7)

in which ϕ c (–) is the Karman surface factor (being 1.0 fora spherical particle), P (W/kg) the specific power input, ν(m2/s) the kinematic viscosity of the oil and Sc (–) theSchmidt number defined as the ratio of kinematic viscosi-ty and film diffusion coefficient.

The experimentally based numbers determined in thisstudy were found to deviate in a range of 30% from thecalculated values. This might be due to the fact that thecorrelations published so far are based on the transfer de-tected between liquid and particles of a 5 to 10 times big-ger size than the particles used in this study. In the caseof such small particles as used in this study the slip veloc-ity between particles and liquid is close to zero at least forthe smallest particles. This is supported by the fact thatthe Sherwood numbers achieved in this study are close to2.0 representing a slip velocity of zero. This coincideswith the finding that enhancement of power input higherthan 0.2 W/kgoil in the stirred tank reactor does not accel-erate the kinetics further. Due to their small size the clay

ShP d

Sccp= +

ϕ

ν2 0 4

4

3

1 4

1 3.

/

/

Shd

Df p

f=

β


Fig. 5. Experimentally gained external mass transfercoefficient in dependence on power input and reactorsystem when varying kind, size and rotation speed of im-peller resp. superficial velocity of gas using 0.6 wt-% ofclay.

particles might move with the motion of eddies or gettrapped in small eddies where viscous effects are prevail-ing in spite of high turbulence in the bulk liquid [15].

3.6 Internal mass transfer

To rate the effect of mass transfer resistance inside theparticle the whole process was modelled and the intra-particle diffusion coefficients were determined by matchcalculations.

To do so a model approach has been developed in thisstudy based on the differential pigment balances aroundand in the particle considering external and internal masstransfer resistance resulting in differential equations.

A fairly good approach was obtained by describing the in-traparticle transport by pure surface diffusion with a con-stant diffusion coefficient. This proceeding resulted in the2 parametric Surface Diffusion Model (SD-model) basedon the external mass transfer coefficient βf and the intra-particle diffusion coefficient DS. The results of using thismodel are given here.

The principles used for this approach are depictured inFig. 4A, the basic equations can be found in [7].

Around the clay particle an external boundary layer of thethickness δ exists, the pigment is transported through thislayer by film diffusion powered by the difference in con-centration in oil at the surface of the particle CS and thebulk concentration C (see equation 5). The concentrationCS is assumed to be in equilibrium with the adsorbedamount of pigment at the surface qS via the adsorption-isotherm. The adsorbed pigment is transported via sur-face diffusion into the core of the particle powered by agradient of amount of pigments adsorbed.

In case of a constant surface diffusion coefficient the dif-ferential equation describing this transfer is [7]:

(8)

with DS (m2/s) representing the intra-particle diffusion coefficient, r (m) the radius of the particle and t (s) thetime.

The internal transfer phenomena are connected to the ex-ternal transfer phenomena (equation 5) via the boundarycondition.

(9)

with R (m) representing the radius of the clay particle. Thedifferential equation obtained was solved numerically.

t r R Dqr

C Cp S f S> = = −0, : ( )ρ∂∂

β

∂∂

∂∂

∂∂

qt

Dq

r rqrS= +

2

2

2

To determine the size of the intra particle diffusion coeffi-cient DS this model was used to match the kinetic curvesgiven in Fig. 3. The external mass transfer coefficient forevery single experiment was determined by graphical dif-ferentiation of the curves as described above, given inFig. 4B. A fairly good matching was obtained when settingthe intra particle diffusion coefficient DS to be 3.5E-15 m2/s. This number can be used to predict bleachingkinetics as can be seen from Fig. 6. Here calculated andmeasured data gained from oil bleaching in a bubble col-umn reactor, using different amounts of clay, are listed.For the calculation the external mass transfer coefficientwas set to be 10 E-6 m/s, a value which is reasonable forthese conditions in a bubble column reactor (as can beseen from Fig. 5).

With this approach the kinetics of bleaching can be de-scribed similar to an adsorptive process.

Using this SD-Model the effect of the external mass trans-fer on the total kinetics of bleaching can be investigateddoing model calculations where the size of the externalmass transfer is varied. The results of this kind of calcula-tion are given in Fig. 6 where the size of the externalmass transfer coefficient has been varied in the range of1.0 to 20.0 E-6 m/s when bleaching with 0.6 wt-% of clayand using an inner diffusion coefficient of 3.5 E-15 m2/sfor calculation.

As can be seen from this graph the effect of the externalmass transfer coefficient influences the total kinetics inthis case only significantly if it is smaller than 8 E-6 m/s. Ahigher coefficient does not enhance the total kineticsmuch further compared to other experiments using thesame amount of clay. This interaction of external and in-ternal mass transfer can be summarised when calculatingthe Biot number Bi with the numerator representing the


Fig. 6. Effect of external mass transfer coefficient on total kinetics of bleaching when bleaching with 0.6 wt-%respectively 0.8 wt-% of clay (calculated data).

external transfer rate to the particle and the denominatorrepresenting the internal transfer rate in the particle. Forthe case of surface diffusion this number is defined as [7]:

(10)

with C0 [kg/m] as initial pigment concentration, rp [m] asthe Sauter mean radius of the particle, ρp [kg/m] as thedensity of the particle and q*

0 [kg/kg] as the loading of theparticle with pigment being in equilibrium with C0. Thehigher the Biot number is the more likely it is that the in-ternal transfer is dominating the rate. If it exceeds 50.0the mass transfer resistance of the film around the parti-cle can be neglected [7]. In the graphs shown in Fig. 6 theBiot number varies between 1.6 and 32.0.

From this figure it even can be seen that the effect of ex-ternal mass transfer coefficient is highest during the firstminutes of bleaching. The closer the systems gets to theequilibrium state of bleaching the effect of the internal re-sistance of mass transfer is dominating the total kinetics.

To demonstrate the approach of saving costs by enhanc-ing mass transfer, a curve calculated for the use of 0.8 wt-%of bleaching clay and gained with a poor external masstransfer of 3.0 E-6 m/s was added to this graph. Compar-ing this curve to the curve gained with 0.6 wt-% of clayand a proper external mass transfer coefficient of 7.5E-6 m/s shows that a proper mass transfer allows thesaving of clay by reaching a reduction of pigment to 20%in about the same time in the reactor for both conditions.

However, in order to get a proper external transfer coeffi-cient a higher power input is necessary, thus increasingthe running costs again. But calculations using the pre-sented model allow to optimise the total costs and e.g.lead to the conclusion, that significant savings of clay arepossible when investing in better external mass transfereither by building a bubble column reactor or by changingthe stirrer system. Thus the possible reduction in runningcosts by saving clay can be balanced with the increase inrunning and investment costs by investing in a better ex-ternal mass transfer to obtain this.

An attractive compromise could be the application of asemi-continuous bleaching reactor run as a cascade ofseveral reactors. Here the external mass transfer shouldbe highest in the first vessel obtained by an intensive mix-ing, followed by tanks where the power input is kept justas high as to keep the particles suspended.

4 Conclusions

The mode of running the bleaching reactor, thus deter-mining the external mass transfer rate of the pigment to

BiC r

D qf p

S p*

=β

ρ0

0

the particle was shown to influence the rate of vegetableoil bleaching significantly up to a certain amount of powerinput. To evaluate this operating point a model is devel-oped describing the bleaching kinetics as an adsorptiveprocess combining extra and intra particle diffusive masstransfer phenomena. This model can be used to predictthe effects of the amount of clay and of the external masstransfer resistance on the kinetics. From these calcula-tions can be deduced an upper optimum for the externaltransfer rate. Exceeding this optimum by enhancing thepower input to the reactor does not enhance the kineticsmuch further since the intra-particle mass transfer domi-nates the kinetics. The rate of the latter depends on claycharacteristics such as particle size and distribution ofpores.

Acknowledgements

The practical and consultative support by Dr. T. Krausefrom Pilot Pflanzenoeltechnologie Magdeburg e.V., theSuedchemie AG, the Institute of Process and NuclearTechnology at TU Braunschweig and the Krüss GmbH isgratefully acknowledged. This research project was sup-ported by the FEI (Forschungskreis der Ernaehrungsin-dustrie e.V., Bonn), the AiF and the Ministry of Economicsand Technology. Project No: 11567 B.

References

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[Received: August 27, 2001; accepted: October 30, 2001]

bleaching kinetics

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