Biodiesel Two Step Production

2nd Mercosur Congress on Chemical Engineering

4th Mercosur Congress on Process Systems Engineering

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Production of biodiesel by a two-step supercritical reaction process with

adsorption refining

C.R. Vera*, S.A. DIppolito, C.L. Pieck, J.M. Parera Instituto de Investigaciones en Catlisis y Petroqumica, INCAPE

(Facultad de Ingeniera Qumica, Universidad Nacional del Litoral, CONICET)

Abstract. A new catalyst-free, effluent-free process for the production of biodiesel is disclosed. The reaction of transesterification of triglycerides is carried out under supercritical conditions, i.e. at temperatures higher than the critical temperature of methanol. Raw materials for the reaction are methanol and triglycerides with any amount of free fatty acids thus enabling the production of biodiesel from cheap feedstocks such as beef tallow or high acidity yellow grease. The reaction proceeds without the aid of alkaline or acid catalysts, thus eliminating the need for neutralization steps downstream the reactor. In order to minimize the heat consumption and pumping power which are usually very high in the one-reactor configuration of all reported supercritical processes, two medium-pressure successive reactors with intermediate glycerol removal are used and a heat recovery scheme composed of heat exchangers and adiabatic flash drums is proposed. Process conditions and equipment design parameters are determined on the basis of the minimization of the heat duty and energy requirements. For the final removal of free glycerol, an alternative approach is presented which differs from the classical washing-distillation and washing-ultracentrifugation operations. Glycerol is cyclically retained in adsorption beds, desorbed in a swing step and recycled to the first reactor. Design parameters are obtained both from experimental data and estimation. No process water effluents are produced with this approach. The whole system is essentially dry and only small water amounts are produced by esterification of the free fatty acids in the feed. Glycerol purification is simplified by the absence of catalyst and low water content.

. Keywords: biodiesel, supercritical, swing-adsorption

1. Introduction

In the last years the need for processing cheaper feedstocks in order to decrease the high price of biodiesel

has been pointed out. Some plant oils like soy and sunflower oils have a highly fluctuating price and in some

years their use can be shifted from the making of biodiesel to the food market due to high international prices.

Beef tallow and yellow grease are inexpensive feedstocks discarded by some industries which could be

advantageously used in the making of biodiesel. They contain high amounts of free fatty acids (FFA, 5-30%)

and they cannot be directly processed by facilities working with the alkali catalyzed process. Many solutions

have been proposed for the handling of acidic feedstocks (Mittelbach and Koncar, 1998): (i) pre-neutralization

is easy but produces a net yield loss if the soaps are not recycled; (ii) pre-esterification of the FFA with methanol

or glycerol, catalyzed by strong acids (slow), followed by esterification in alkaline medium (fast); (iii) Fully

acid-catalyzed esterification. Acid catalysts are not as effective as the alkaline ones and take much longer

reaction times, thus leading to big reactor volumes. The process catalyzed by alkalis (NaOH, KOH) or by acids

(H2SO4) need of washing steps to eliminate the dissolved catalyst. Washing is also used to eliminate glycerol

and produces great amounts of effluents (3-10 wastewater litres/litre biodiesel) (Karaosmanoglu et al., 1996;

Demirbas, 2003). If washing is performed without flashing the unreacted methanol the (water-methanol-

* To whom all correspondence should be addressed. Address: INCAPE, Santiago del Estero 2654, 3000 Santa Fe Argentina. E-mail: [email protected]



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glycerol) mixture must be distilled in order to recycle the methanol.

An alternative process to the catalytic ones has been recently developed by japanese researchers (Kusdiana

and Saka, 2001; Sasaki et al., 2001). The process comprises the one-stage transesterification with supercritical

methanol, in the absence of any catalyst. At high temperaturas metanol is in a supercritical state (Tc=235 C)

and forms a homogeneous phase with the oil phase, something that does not occur in the other conventional

processes, unless a minimum threshold of 25% conversion is attained (Srivastava and Prasad, 2000). Reaction

by supercritical methanol has some advantages: (i) Glycerides and free fatty acids are reacted with equivalent

rates. (ii) The homogeneous phase eliminates diffusive problems. (iii) The process tolerates great percentages

of water in the feedstock (Kusdiana and Saka, 2004); catalytic process require the periodical renoval of water in

the feedstock or in intermediate stage to prevent catalyst deactivation. (iv) The catalyst renoval step is

eliminated. (v) If high methanol:oil ratios are used, total conversion of the oil can be achieved in a few minutes.

Some disadvantages of the one-stage supercritical method are clear: (vi) It operates at very high pressures

(25-40 MPa). (vii) The high temperatures bring along proportionally high heating and cooling costs. (viii) High

methanol:oil ratios (usually set at 42) involve high costs for the evaporation of the unreacted metanol. (ix) The

process as posed to date does not explain how to reduce free glycerol to less than 0.02% as established in the

ASTM D6584 or other equivalent international standards. In this sense washing can not be eliminated because

usual deglycerolization steps comprise intensive washing with water or washing/ultracentrifugation.

It has been pointed out that in a one-step continuous supercritical process a large quantity of the solvent is

required for the treatment of the feedstock as compared with the batch-type system (Minami et al, 2004). This is

not really accurate and a rebuttal to this and a solution to (vi)-(ix) is proposed in this work. The one-reactor

configuration of the classical supercritical process is replaced by two supercritical reactors operating in series.

An intermediate step of glycerol removal between the two reaction steps is implemented in order to allow the

reaction to proceed to completion with reasonably low methanol:oil ratios (6-10), lower temperatures (250-300

C) and lower pressures (2.5-5.0 MPa). The low methanol:oil ratio is expected to decrease the heat duty

required for the evaporation of the unreacted metanol. The lower pressure is expected to decrease costs of

pumping and of robustness of the equipment. For the pumping power, this issue is critical because in

continuous running applications, it costs more to operate a pump for one year than it does to buy the pump. In

some cases, a 20 percent reduction in operating costs can pay for the cost of a pump in a little over a year's time.

Additional heat recovery schemes are introduced to decrease the total heat duty of the process. Double tube

heat exchangers before the supercritical reactors are used to preheat the reacting mixture and adiabatic flash

drums downstream the reactors are used to evaporate the unreacted methanol.

A final modification is tried by studying the glycerol removal of the biodiesel by means of adsorption in a

packed bed. In this way the whole process would operate under anhydrous conditions and the issue of effluents

and the water content of the glycerol by-product can be reduced. Solid adsorbents have been used in the past for

removing glycerol from process streams (Griffin and Dranoff, 1963) and more recently to shift esterification

equilibria (Stevenson et al., 1994).



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Fig. 1. Flowsheets of 2 supercritical biodiesel production processes. (a) 1-step, no heat recovery, no methanol removal

(Minami et al., 2004). (b) 2-step, with heat recovery (this work). MeOH=methanol, Gly=glycerol, TG=triglycerides (oil,

grease), SCR= supercritical reactor, HE=heat exchanger, CON=condenser, FAME=fatty acid methyl esters (biodiesel),

DEC=decanter, AFD=adiabatic flash drum, BPR=backpressure regulator.

2. Model development

2.1. Fluid properties

Components are named with abbreviations throughout the text (see Figure 1 caption). Pure component

properties (methanol, glycerol) were mostly taken from the NIST on-line Webbook database. Triolein and

methyl oleate were taken as model compounds for triglycerides and biodiesel. Vapor pressure was calculated

with Antoines formula only in the case of methanol and glycerol. In the case of biodiesel the volatility data

(Goodrun, 2002) was fitted with a Clausius-Clapeyron formula and in the case of triglycerides a group-

contribution correlation was used (Ceriani and Meirelles, 2004). Activity coefficients were calculated with the

UNIFAC algorithm (Fredenslund et al., 1977) with group contributions by Gmehling et al (1993).

Activity coefficients for VLE were valid only in the subcritical range. In the supercritical range, an empirical

ad-hoc P-T-Xv relationship was built from the data of Sasaki et al. (2001) and Saka and Kusdiana (2001):

P (MPa) = 0.4092* T (C) + 65.19107 * Xv 134.4197 (1) Xv = volume fraction of MeOH; 250 < T (C) < 475

2.2. Units

Tubular supercritical reactor. The transesterification of oil is known to proceed via a three step

consecutive reaction network (Noureddini and Zhu, 1997) comprising the conversion of oil to a diglyceride, of

diglycerides to monoglycerides, and of monoglycerides to glycerol. There are no kinetic models of oil

methanolysis that make a difference between the different positional isomers of di- and monoglycerides and

constitutional isomers (different acyl groups) of the glycerides. The 3-step oil methanolysis model might be

enlarged to include the reaction of free fatty acids (FFA) and water but this is not necessary in the supercritical

reactor model. In contrast to the alkaline catalyzed system, where FFA methanolysis is difficult to occur due to

the formation of soaps, the catalyst-free supercritical methanolysis of FFA proceeds faster than the methanolysis

of glycerides and their kinetic constants are very similar (Warabi et al., 2004). The same authors have also



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demonstrated that the reacting system is extremely tolerant to the presence of water (Kusdiana and Saka, 2004).

For this reason we do not include FFA or water in the formulation of the model and the feedstock is considered

to be composed only of oil and methanol.

The 3-step reaction network was further simplified to a one-step lumped kinetic model, where FAME is a

structural lump containing all monoglycerides and TG is another lump comprising all triglycerides. The

merging of the intermediate steps is justified by the fact that the TG-to-diglyceride step is reported to be much

slower than the consecutive diglyceride-to-monoglyceride and monoglyceride-to-FAME steps (Noureddini and

Zhu, 1997). The final kinetic model is described by the global reaction:

TG + 3 MeOH 3 FAME + Gly (2) Kinetic data as a function of temperature were taken from the work of Kusdiana and Saka (2001) and

Demirbas (2003) and correlated to the second-order kinetics form written below. In this way k(T) was obtained.

(3)

The equilibrium constant (Keq) was taken from the reported data of Noureddini and Zhu (2001). Cj is the

concentration of the j species. The reactor was considered to be heated by means of jacket with a constant

temperature (Tr) fluid.

(4)

Double tube heat exchanger. The pressure is considered to be high enough in order to assume that no phase

change occurs. In the inner and outer tube (i=inner, hot side; o=outer, cold side) the properties of the

methanol:oil mixture were calculated from known mixing rules.

(5)

Adiabatic flash drum. A Rachford-Rice algorithm (Henley and Seader, 1981) was used for solving the flash

drum problem. Fugacity coefficients were taken as unity and fugacity was taken as the product of the vapour

pressure multiplied by the mole fraction of the component.

Fixed bed adsorber. Several assumptions were made: (i) The whole system is isothermal. (ii) The flow

pattern was supposed to be of the axially dispersed plug-flow type; radial concentration and temperature

gradients in the adsorption bed are considered negligible. (iii) The Langmuir isotherm can be reduced to a linear

one (high dilution regime); axial diffusion DL and film coefficient kf obey the Wakao and Funazkri (1978)

correlations. (iv) The transfer through the film can be described by a linear driving force equation (eq. (9)).

Applying these assumptions to a mass balance through a packed bed equations (6-11) were obtained. The

constants for the adsorption of glycerol were taken from reported data for several adsorbents. An analytical

solution to (6-11) was used, based on the quasi-lognormal distribution approximation (Li et al., 2004). The Q-

LND approximation is known to become weak when the Biot number (Bi) is low, but in this case Bi is

FAMEGlyeqMeOHTGGlyGlyMeOHFAME CCKkCkC

dzdC

dzdC

dzdC

dzdC

)/(33 ====

)(4 oipt

io TTCDh

dzdT

dzdT ==

densitytcoefficientransferheathdiametertubeDTTCDh

dzdT

trSCRpt

r ==== ;;)(4



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prevented to reach negligible values due to the small value of the intraparticle diffusivity in liquid medium.

(6)

(7)

(8)

(9)

(10)

(11)

Notation: u=space velocity; =bed porosity; p=pellet density; Rp=pellet radius; Re=Reynolds number; Sc=Schmidt number; Dm=molecular diffusivity; q*=equilibrium adsorption capacity of the adsorbent;

qm=saturation adsorption capacity; q=actual adsorption load; b=Langmuir coefficient; K=Henry coefficient;

k*=global mass transfer coefficient

2.3. Process conditions and restrictions

The total heat duty is calculated as the sum of the enthalpy difference between the entrance and the exit of

each reactor. The pumping power is the sum of the pumping power in the first and second reaction steps.

Q = Heat duty = F*Cp*(TSCR_outlet TSCR_inlet) (12)

PW = Pumping power = Fv_1*(PSCR_1 Patm) + Fv_2*(PSCR_2 Patm) (13)

F is the molar flowrate and Cp is the heat capacity of the mixture. Fv is the volumetric flowrate. The whole

problem is subjected to the following set of constraints:

(14)

The restrictions have been worked out from the limits established in the quality standard ASTM D6584. In

order to study the design requirement on the glycerol fixed-bed restrainer an equilibrium glycerol concentration

of 0.5% at the outlet of the last decanter was calculated from solubility correlations reported by Kimmel (2001).

3. Results and discussion

Supercritical reactors. The fitting of the kinetic data produced the following results:

ksub = 16.025 * Exp(- 4628.3 / (T+273)); T (C), k (liter mol-1 min-1), T < 275 C (subcritical) (15)

ksupr = 698.78 * Exp(-5664.3 / (T+273)); T (C), k (liter mol-1 min-1), T > 275 C (supercrit.) (16)

Results for the correlation of the kinetic constant and the P-T-Xv relation (eq. (1)) are presented in Figure 2

(b). A comparison between several kinetic constants seems mandatory. At 50 C the value of non-catalytic ksub

is 0.00001 liter mol-1 min-1 (negligible) while the corresponding value for the alkali catalyzed reaction is 0.05

liter mol-1 min-1 (Noureddini and Zhu, 1997). ksupr is 0.036 liter mol-1 min-1 at 300 C and 0.08 liter mol-1 min-1

at 350 C. It can be seen that the magnitude of the reaction rate constant of the supercritical process is similar to

%01.0%;1.0%;6.99 __ productGlyproductMeOHTG CCConversion

pp

p

f

p

m

p

mf

p

L

pl

DR

kR

k

qqktq

CKbCbCq

q

RD

Sck

ScuRD

tq

zuC

zCD

tC

153*1

)*(*

1*

2)Re1.10.2(

5.0Re

202

01)(

2

6.03/1

2

2

+=

=

+=

+=

+=

=+

+



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that of the alkaline process.

Fig. 2. (a) Arrhenius plot of the kinetic constant.(second order model). (b) Parity plot of the experimental and predicted

pressure values in the supercritical and subcritical region.

Fig. 3 contains the values of conversion and (T/Tr) (ratio between the reactor and heating fluid temperatures).

Figure 3-(a) shows the influence of the temperature at the inlet of the reactor, for a fixed value of Tr. The effect

of the temperature is very marked when trespassing the threshold limit between the subcritical and critical

regimes which produces a discontinuity in the slope of the conversion curves. Fixing the value of Ti and setting

Tr as variable generates a plot with a similar pattern, though the dependence on Tr is stronger.

The value of Tr sets the maximum pressure of the process, because at the exit of the reactor the temperature of

the product stream approaches Tr. The pressure dependence on the temperature is very high and therefore a

small value above the threshold limit is the most convenient, i.e., Tr = 280-290 C.

Fig. 3. Temperature and conversion values. (a) In the 1st reactor, as a function of Tr and Tinlet. (b) Temperature and

conversion profiles along both reactors; conditions: Tr=320 C; inlet temperature Tinlet=250 C; oil flowrate at the entrance

of the first reactor w=1000 litres h-1.

Figure 3-(b) contains conversion and temperature values in the 1st and 2nd reactors, as a function of the

methanol-to-oil ratio (Rf) in the entrance of each reactor:.

Rf = CMeOH / (CTG+CFAME) (17)

For the sake of simplicity identical values of Rf and Tr in both reactors were employed in the simulation runs.

As expected from the results of Sasaki et al (2001) if low ratios are employed (Rf=5-10) in order to keep the total



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pressure low, temperatures equal or higher than 270 C are needed to have a high conversion in the first reactor

and total conversion in the second. If the reactor and the process conditions provide a sufficiently high

residence time the level of conversion is dictated by the thermodynamics of the reaction and depends only on the

value of Rf.(as in the case of the results of Fig. 3). The small amount of glycerol entrained in the oil phase in the

first decanter has a negligible effect on the conversin at the outlet of the second reactor. The results indicate that

a minimum value Rf=10 is needed in order to fulfill the biodiesel purity specification of the quality norms.

Values of Rf > 10-15 seem unconvenient because they involve higher working pressures and higher evaporation

duties.

Flash drums. The MeOH content must be reduced to a sufficiently small value in the first decanter in order

to reduce MeOH losses and decrease the solubility of Gly in the oil phase. In the 2nd flash drum the methanol

content must be reduced to 0.1% in order to keep the flash point and cetane values inside the diesel range.

Fig. 4. MeOH content (weight %) (a) and flashing temperature (b) in the liquid downstream the first flash drum as a

function of the Rf ratio and the feed temperature before the flashing valve (Tv).

A plot of the content of MeOH in the liquid phase in the 1st flash drum as a function of the Rf ratio and Tfi, the

temperature of the flash drum, is shown in Fig. 4. The results indicate that a minimum value of Tfi=140-160 C

at Rf=6-10 is needed in order to keep a maximum methanol content of 3% in the liquid stream issuing from the

first flash drum. Higher Rf ratios, e.g. greater than 20, are not possible because in this case the methanol cannot

be completely adiabatically flashed even if the heat exchanger is eliminated. Another concern is that of the

degradation of glycerol. At high temperatures polymerization occurs, the quality is decreased and the

purification becomes difficult due to the increase in viscosity. Ueoka and Katayama (2001) have suggested that

flash towers for glycerol purification should operate at Tfl < 140 C. Upstream the flashing valve temperature is

not a concern because glycerol is highly solvated by methanol.

The second flash drum has a similar behaviour and therefore it can be inferred from Fig. 4-(a) that a 0.1%

MeOH cannot be obtained with a single flashing stage at the chosen conditions. The limit value of 0.1% can be

attained however by operating the 2nd flash drum at a reduced pressure (28.5 Hg) and Tv=190 C .

Heat and pumping duties. The heat duty in the reactor depends directly on the inlet temperature of the

flashing valve (Tfi). If we consider that the temperature at the outlet of the reactor is practically equal to Tr and

that there are no heat losses in the heat exchanger, the energy balance in the heat exchanger gives:

TSCR_inlet THE + Tr - Tv (18)



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THE is the temperature of the feedstock entering the preheating heat exchanger (20-60 C). Tr should be set at

280-290 C just above the sub/supercritical threshold, in order to have high conversion values in the reactor.

Higher temperatures are prohibitive because they generate exceedingly high pressures in the system (eq. (1)). Tv

should be set at 140-150 C. This is the minimum temperature that ensures low methanol levels in the liquid

streams entering the decanters. Greater temperatures are discouraging because they increase the heat duty in the

reactors, as explained below. A value of Rf=10 is adopted in order to allow a total conversion value of 99.6%.

Higher Rf values are discouraging because they increase the total mass flowrate and the heat duty. With the use

of (12) and (18) and considering Tr TSCR_outlet: Qi = F Cp (TSCR_outlet TSCR_inlet) F Cp (Tv THE) (19) Q = Q1 + Q2 (20)

Q1 is the heat duty of the first stage and Q2 the heat duty of the second stage. Q2 is higher than Q1 because of

the higher value of Tv needed to almost completely flash the methanol. Now we can use (19) and (13) to

compare the pumping power and heat duties of the one-reactor and two-reactor biodiesel production schemes:

One reactor (no heat recovery, no MeOH removal,

Rf=42, Tr=300 C): PW=8.3 kWatt h litre-1, Q= 2166 kJ litre-1

Two reactors (with heat recovery, with MeOH

removal, Rf=10, Tr=290 C): PW=1.9 kWatt h litre-1, Q= 1382 kJ litre-1

Glycerol adsorber. A spatial velocity recommended for solid-liquid systems with intrapellet mass transfer

limitations (Teo and Ruthven, 1997) was adopted for the sake of simplicity (u=3 cm min-1) and maintained

throughout all calculations. The molecular diffusion coefficient was calculated as Dm=6.08x10-9 m s-1.

(Kimmel, 2004). The glycerol content of the biodiesel stream at the outlet of the decanter (C0) is approximately

0.2% at 25 C (Kimmel, 2004). Penetration curves for a fixed bed glycerol restrainer are plotted in Figure 5.

Fig. 5. Breakthrough curves for different values of the adsorption constant K. Adsorption conditions: u=3 cm min-1;

bed length=2 m; temperature=25 C, C0=0.2% glycerol in biodiesel.

The value of K depends strongly on the kind of support: K < 0.01 (titania-alumina, Chen et al., 2001), K=0.07

(silica, de Roodea et al., 2001), K=0.5 (Amberlyst 15, Griffin and Dranoff, 1963). Glycerol adsorption was

modelled on a monocomponent base and multicomponent adsorption was ruled out. In this sense, Nijhuis et al.



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(2002) have reported that in organic media the adsorption of C8 esters over Nafion/silica is negligible (K0), while the adsorption of short-chain alcohols and water is important.

The breakthrough curves of Fig. 5 indicate that K values between 0.005 and 0.02 are enough to keep the

glycerol content of the biodiesel product stream below 0.002% for a fresh bed of adsorber and for a time of

operation of 1-3 h. Too high K values may render impossible the desorption of glycerol and hence the

regeneration of the bed. In this sense a packed bed of silica seems to the most appropriate for a cyclic

adsorption-desorption operation of the glycerol restrainer.

Two regeneration schemes for the bed seem possible. In one case a stream of pure biodiesel can be passed at

a higher temperature and recycled to the first reactor. This scheme seems odd because it needs the supply of

extra heat and decreases the biodiesel yield. A better possibility seems to flush the packed bed with a stream of

methanol and to recycle this stream to the first reactor.

4. Conclusions

The operating pressure for the production of biodiesel by the reaction of oils in supercritical methanol can be

advantageously reduced if the reaction is allowed to proceed in two successive steps with intermediate removal

of glycerol. The one-reactor setup works with molar methanol:oil=42 and a pressure of 14-43 MPa (270-350

C). In the two-reactor setup the decrease of methanol:oil to 10 results in a reduction of the working pressure to

less than 4 MPa and of the pumping power to less than 25% the value for the one-reactor lay-out.

A low methanol:oil ratio also enables the use of adiabatic flash drums to vaporize the unreacted methanol at

the outlet of the reactor, with a substantial decrease of the heat duty of the process. An additional heat

exchanger contacting the streams entering and exiting the reactor enables an additional recovery of heat. The

final heat duty can be reduced to almost half the value corresponding to the one-reactor lay-out with no heat

recovery and no methanol removal.

The supercritical catalyst-free process can be run under fully dry conditions if glycerol removal is performed

by adsorption in a packed bed. Common acid adsorbents (silica, acid resins) can retain glycerol and keep the

biodiesel product inside the glicerol content specification of the quality norms. A convenient regeneration

procedure seems to be the flushing of the adsorbed glycerol with a methanol stream and the recycling of this

stream to the first supercritical reactor.

Acknowledgments

This work was financed with the funding of Universidad Nacional del Litoral (CAI+D 2002 grant, Project

146) and the National Agency for Promotion of Science and Technology (ANPCyT PICT 2002 grant, Project

14-08282).

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Biodiesel Two Step Production

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