ARTICULOS

1. cw ~imring .S&nce. Vol. 40, No. 12. pp. 2281-2292. 1985. ooo9-2so9/85 s3.00 + 0.00
Fviltted in Gmlt Britain. 0 198s. ~‘
~rgamon FWSS Ltd.
HYPERAZEOTROPIC ETHANOL SALTED-OUT BY
EXTRACTIVE DISTILLATION. THEORETICAL EVALUATION
AND EXPERIMENTAL CHECK
DIEGO BARBA, VINCENZO BRANDANI and GABRIELE DI GIACOMO
Dipartimento di Chimica, Ingegneria Chimica e Materiali, Universita’ degli Studi, 67100 L’
Aquila. Italy
(Received 25 November 1984)
Abstract-A distillation process for the production of hy-perazeotropic ethanol from a dilute wine obtained
from the fermentation of biomass has been studied. This process utilizes the coupling of a soft
preconcentration stage and of a dehydration stage based on the salting-out effect produced by calcium
chloride on the ethanol in an aqueous solution, with the disappearance of the azeotrope. The salt is employed
in a closed cycle, due to the presence of a regeneration stage, therefore no consumption of calcium chloride is
noticed.
The distillation process utilizes one column consisting of two wtions operating at different pressures in
order to reach an efficient heat recovery.
In this paper, a simplified flow-sheet of the processand the principal operating conditions of the distillation
column are illustrated. When compared with other processes, conventional or under development, this one is
characterized by the promising reduction of the s-c energy requirement.
The operating conditions chosen for the distillation with salt have been experimentally checked using a
laboratory column running continuously with calcium chloride as salting-out agent. Moreover, the
experiments confirmed the reliability of the mathematical model of the process. Further experiments are in
progress with the aim of utilizing a mixture of salts which can be fed from the bottom of the dehydration
section back to the fermentor, so that the salt regeneration stage can be reduced.
INTRODUCIXON
In the field of research on the use of renewable
resources, the production of ethanol by fermentation is
especially interesting. There are already large quan-
tities of fermentable biomasses available from wastes
or agricultural surplus and more could be obtained
with careful agricultural planning.
In the Brazilian national economic plan, an annual
production of hy-perazeotropic ethanol of SO million
hectolitres is expected by 1985, using sugar cane as the
raw material Cl]_
At the end of the 197Os, the use of ethanol as fuel was
the object of several studies which aimed to show the
technical+conomic advantages of this choice [2,3].
However, the energy balance that emerged proved
rather unattractive when traditional industrial pro-
cesses are employed. In fact, the net energy benefit is
not great enough to justify investment costs for the
construction of the plants.
From an energy point of view, the bottle-neck of the
whole productive cycle is the distillation section where
a significant part of the total demand lies. In fact,
hyperazeotropic ethanol must be obtained from very
dilute fermentation broths, the concentration of which
is not usually greater than 7.5 o/0by weight [4]. Because
of the presence of an azeotrope two concentration
stages are usually necessary: in the first stage, hypo-
azeotropic ethanol is produoed, and in the second
stage, hyperazeotropic ethanol is produced by means
of a third component (benzene, pentane or diethyl
ether).
In a study developed by the Italian Research
Committee (C.N.R.) as a part of the “Progetto
Finalizzato Energitica” [4], the energy consumption
for this process was estimated to be 2000 kcal/kg of
anhydrous ethanol (about 70 “/, of the energy demand
for the whole productive cycle). It was also found that a
significant reduction of the specific energy requirement
can be achieved by adopting different processes rather
than the traditional one cited above.
The theoretical results of the C.N.R. study [4] show
that an energy reduction of about 25% could be
reached; this gives an energy consumption of about
1500-1600 kcal/kg of anhydrous ethanol.
An accurate study [S], comparing six conventional
distillation processes, shows that the energy con-
sumption lies between 31 and 64 0/0 of the heating
value of ethanol. This means a range of about
22-500 kcal/kg of anhydrous ethanol.
Moreover, the U.S. Department of Energy has
recently reported an evaluation of non-distillation
ethanol separation processes under development [6],
qualifying as promising the processes that assume
as a target an energy requirement in the range
1100-1500 k&/kg of ethanol.
With the aim of reaching this level of consumption
using a distillation process, which appears more re-
liable than non-conventional non-distillation pro-
cesses, we found it useful to reconsider critically the use
of extractive distillation based on the salting-out effect
when coupled with a low preconcentration stage.
Extractive distillation with salts was proposed as far
back as 1922 [7] and it is also known [8,9] that some
industrial plants have been used for the production of
absolute ethanol from aqueous solutions. However,
this process was never really successful because of the
2287
Ll!.>
4”: IL-l

2. 2288 D. BARBA et al.
technical problems encountered in the dissolution and
the subsequent recrystallization of the salt [lo] and
because of the need for special construction materials
to avoid corrosion problems. On the other hand, at
that time, energy saving was not such a priority so as to
encourage researchers to work on these problems.
Our goal was to check experimentally a new process
[11] which, by combining the salt etfect with soft
preconcentration of the initial dilute tilution, reduces
the energy requirement to 1200 kcal/kg of anhydrous
ethanol. After presentation of the process flow-sheet
and of the values of the operating conditions obtained
by the mathematical model, we report the experimen-
tal results obtained using calcium chloride as the
“salting-out” agent in a programme of experiments
conceived to confirm the calculated values. In the
meantime, research is in progress with other types of
salts, useful in the fermentation, which can be fed from
the bottom of the dehydration column back to the
fermentor.
PROCESS FLOWSHEET
When a salt is dissolved into a liquid solution, the
activities of the other components change as a con-
sequence of the structural modifications experienced
by the liquid phase For example, Fig. 1 shows the
effect of the dissolution of calcium chloride into
ethanol-water mixtures at atmospheric pressure.
The equilibrium curve for the binary ethanol-water
system is compared with the experimental data for the
ethanol-water system with calcium chloride added at a
constant weight ratio [12] and also with the cor-
responding predicted equilibrium curve obtained by a
model presented elsewhere [ 131.In this case, ethanol,
the concentration of which in the vapour phase
increases, is salted-out, whereas water is salted-in. This
effect can be utilized in extractive distillation where a
salt, rather than a liquid, is used to separate an
azeotropic mixture or mixtures of components with
close boiling points.
Experlmentals [NO]
Calculated [q]
Ethanol-Water System
ov ’ I 1 1 I
0 0.2 04 0.6
X,,,”
08 1.0
Fig. 1. Vapour-liquid relationships for ethanol-water and
ethanol-water-calcium chloride at 1 bar.
In some systems, such as ethanol-water, the use of a
salt allows easier separation of the components and, as
shown later, a reduction in the energy consumption
compared to conventional azeotropic distillation.
In the case of the ethanol-water system, several salts
can be utilized in extractive distillation because they
determine the disappearance of the azeotrope at
atmospheric pressure. Calcium chloride [12],
cobalt(II)chlo [14],cupricchloride [15], nickel(II)
chloride [15], strontium bromide [16], sodium and
potassium acetates [17, 18], calcium nitrate [19],
sodium and potassium iodides [20,21] are among the
most efficient salts.
Calcium chloride, which was used in our laboratory
column, is very attractive because it greatly increases
the relative volatility, as shown in Fig. 1.
Figure 2 shows a schematic flow-sheet of the studied
process which produces hyperaxeotropic ethanol start-
ing from dilute wine. The two distillation sections
operating at different pressures are arranged in a single
column, because of the limited number of total plates.
After a pre-heating treatment, the dilute solution,
coming from the fermentor, enters the lower section of
the column, which is actually a stripping tower operat-
ing without salt. Since in this section only a limited
degree of preconcentration is required (about 50 % by
weight of ethanol), the binary system operates in a
region of high relative volatility. As a consequence, a
lower load is required compared with a reflux column
designed to obtain ethanol with a composition near to
the azeotropic composition.
The overhead vapour, after recovery of its heat
content, is fed to the upper section of the distillation
column which operated at lower pressure than the
stripping section. The upper section of the tower,
where the salt enters dissolved in the reflux stream,
gives hyperazeotropic ethanol as overhead product.
The bottom aqueous saline solution, except for the
Fig. 2. Simplifkd flow-sheet of the process for the production
of hyperazeotropic ethanol.

3. Extractive distillation of hyperszeotropic ethanol 2289
portion sent to the fermentor, is fed to an evaporative
crystallizer and, after final drying, the salt is recycled
back.
It is worth noting that the bottom stream for the
extractive distillation section of the column comes
from the overhead vapour stream which leaves the
evaporator.
The principle of the studied flow-sheet is based on
the choice of a coupling criterion which determines the
operating conditions for both the column sections
characterized by a high average relative volatility.
A mathematical model has been studied to simulate
the distillation process shown in Fig. 2. Due to high
non-ideality of the system, a matrix method of compu-
tation, studied for multicomponent distillation
[22,23], has been selected and adapted to the present
case, where one of the components is non-volatile.
Lt is known that the material and heat balances for a
multi-stage unit operation can be written as matrix
equations:
Xj=A,:*bj j=l,...,M (1)
V=H_‘
g. (2)
The summation over j of eq. (1) represents the
congruence equation expressed by:
C&+A&j)-‘
bj=CXj=a
i j
where A, has been decomposed into Aoj, functions of
the initial temperature profile t,, and A&j, functions
of the unknown correction temperature profile At,.
The algorithm studied in Refs 22 and 23 allows eq.
(3) to be made explicit with regard to correction
temperature profile:
At, = xA$’ A;,,
i )-‘
(4Xoj-u)- f4)
In Table 1 the type, size and expression of each
general term of these matrices are summarized.
The equilibrium ratios shown in Table 1 are ex-
pressed by:
K.. = YijxijpZij
U
pXi,j
(5)
where the vapour pressures P$,j corrected for the non-
ideality of the vapour phase, are calculated by using the
Hayden and O’
Connel correlation [24] for the predic-
tion of the second virial coefficient and where the
activity coefficients y, for the two solvents in the liquid
phase are calculated using an ad hoc thermodynamic
model described elsewhere [13] with the numerical
values of the parameters reported in Table 2.
The mathematical model has been used to obtain the
process parameters reported in Table 3, from which a
specific energy requirement of about 1200 kcal/kg of
anhydrous ethanol can be calculated.
In Fig. 3 the calculated concentration profiles of the
liquid along the extractive section of the column are
shown.
Table 2. Numerical values of the parameters for the calcu-
lations of the activity coefficients of the system ethanol
(1jwater (2)+zalcium chloride (3) [13]
r1 = 2.11 r2 = 0.92
41 = 1.97 92 = 1.40
4; = 0.92 92 = 1.00
logD, = 1.3979-0.00264 logD, = 1.903615
(t - 20) - 0.001969(t - 20)
(A, /R) = 8754- 18.1x T (L,/R) = 7363- 9.0 x T
+ 0.0202x T= + 0.00695x TZ
a12 = - 640.029 021 = 1996.392
+ 196,353/T - 543,307fl
ST* = 7.911- 1836/T ST* = 1.796+ 386.37/T
513 = 1.0377- 323.82/T rzJ = 0.0208-27.16/T
kI3 = 0.229-59.92/T k,, = 0.0424- 7.46/T
Table 1. Type, size and expression for the generic elements of the matrices in eqs (l), (2) and (4)
Tridiagonal matrix z%,,~(N x N)
f
qi-1 = L-l a;.i =
i
-(L,+D) i=l
-(V;,Ki,j+Li) i > 1
Bidiagonal matrix A;, j (N x N)
a&+ 1 = - V1+*
[
dKi+1 j
AX.
at,
a+l,j
1
Vector bj (IV x 1)
Bidiagonal matrix H [(N - 1) x (N - I)]
i # feed stage
i = feed stage
Ahi. i =
(Hi - hD) i=2
(Hi-hi-,) i>2
hi+, = (h--h,-,)
Vector g[(N-1)x l]
QC
gi = (hi_, -hi_2)(1N-F,_-FV)
(hi- 1 - hi)(&q- F, - F,) + QR
i=2
i12
i=N

4. 2290 D. BARBAer al.
Table 3. Processparametersof the distillationeohmm for the productionof hyperazeotropic
ethanolusingcalcium
chloride
Preconcentration
(8 stages, 1.5 bar) Dehydration(22 stages, 190 torr)
(&
Flow rate EtOH salt Flow rate EtOH
(kg/h) (wt%) (wt%) (& (kg/h) (wt%)
Feed 95.7 100,ooo 7.50 - 92.4 15,338 48.60 -
Distillate 103.3 15,338 48.60 - 46.8 7527 99.00 -
Bottom produet 111.7 97,946 0.04 - 65.6 14,737 8.40
Bottom column
negl.
vapour 111.7 132284 - - 76.9 5688 0.00 -
CaC12.wN%
;,i;
EtlJH.
w% salt freebasii
Fig. 3. Calculatedconcentrationprofilesin the liquid phase
along the column.
Comparison of the energy requirement with other
conventional and non-conventional processes is
shown in Table 4, where the distillation process con-
sumptions are derived from the work of Black CS] and
the energy consumptions of the non-distillation pro-
cesses are shown in the SERI report [6].
EXPERIMENTAL
In order to verify the reliability both of the thermo-
dynamic model (which does not require any ternary
adjustable parameter) and of the simulation model of
the distillation column together with the adopted
numerical procedure, some experimental runs in the
area of the chosen operating conditions were carried
out and the results are listed in Table 5. In the
experimental runs, a glass packed column was used.
Prevention of heat losses was achieved by the use of
vacuum jackets. Our laboratory column was made by
connecting two or three sections 60 cm long, having an
i.d. of 5 cm. Preliminary runs, carried out with
ethanol-water binary mixtures, indicated that each
section was equivalent to about five theoretical stages.
In each run, the feed was introduced between the
first and second sections starting from the bottom. A
schematic representation of the apparatus used for the
experimental runs is shown in Fig. 4. The reflux, at a
constant salt concentration, was realized, as shown in
Fig. 4, by introducing two streams into a mixer at a
constant flow rate. The first of these two streams was
made up of condensed over-head vapour and the
second of condensed over-head vapour presaturated
with calcium chloride. In each run, start-up of the unit
was carried out by filling the saturator still with the
axeotropic mixture. The feed flow rate and the reflux
were set up, in each run, by regulating the metering
pumps. The salt concentration in the reflux was
periodically controlled by taking small samples for
analysis. The weight percentage of calcium chloride
was determined by weighing the initial solution and
the solid obtained after complete drying. The distillate
Table 4. Comparison of processesfrom the energyrequirementpoint of view
Type of process
SpeaSc energy Fraction of EtOH
consumption heatingvalue Range of concentration
(kcal/kg EtOH) (%)= W %)
Proposed process
Distillationprocesses[S]
Low pressuredistillation
Azeotropic distillation:
Pentane
Benzene
Diethyl ether
Extractivedistillation:
Gasoline
Ethyleneglyeoi
Non-distillationprocesses[63
Solventextraction
Membrane pervaporation
1200
2800
Z%
3ooo
2200
4500
1500 21 10-98
1100 16 8-99.5
17
40
34
38
43
31
64
7.5-99
6.698
6A-dry EtOH
6.4-dry EtOH
6.4-dt-y EtOH
6.4-dry EtOH
6.&dry EtOH

5. Extractive distillationof hypera=o tropic ethanol 2291
Table 5. Operatingconditions,experimental
resultsandcomparison with theoretical simulation for the extractive distillation
with calcium chloride
R&lx Feed Distillate Bottom product
Ratio c&l, g/h EtOH g/h EtOH g/h C&l, EtOH
No. Enriching w %I (wt%I w %I (wt %I wt %I
Run of stages stages exp. cak. exp. cak. exp. CalC.
1 10 5 5.6 6.0 960 42.0 70 99.0 98.2 914 - 2.7 (36.6) 36.6
f 10
10 5
5 20
1.0 6.0
2.2 1330
1000 44.0
43.5 275
151 98.5
95.0 97.0
95.6 1071
855 -
- 0.8
1.5 (29.3)
(34.1) 29.7
34.0
4 10 5 3.7 8.1 z: 47.7 102 98.0 97.4 895 4.5 3.6 (40.3) 40.3
5 10 5 5.6 21.2 46.3 62 99.3 98.7 1008 7.9 9.0 (38.9) 39.0
6 15 10 3.4 16.5 990 44.7 101 99.3 99.6 975 6.6 7.0 (35.8) 35.8
7 ‘
15 10 5.5 8.8 1220 7.9 98.0 98.4 1192 2.7 2.6 (3.2) 3.2
8 15 10 4.0 8.1 370 46.3 99.3 98.9 315 7.9 9.3 (27.9) 28.0
Values in parentheses were obtained from the material balance.
Re‘
,ux Solu,,on
Fig. 4. Schematic representation of the apparatus used for the experimental runs of extractive distillation
with calcium chloride.
flow rate was checked periodically by collecting this
stream into measuring cylinders. Its composition was
determined by measuring at 20°C the density of the
liquid with a Parr precision densimeter. During the
column transient behaviour, the distillate concentra-
tion was measured with an Abbe-type refractometer
with the prism temperature maintained at 20°C. This
kind of measurement was preferred because of its
rapidity and because of the small samples required for
the analysis. The average flow rate and composition of
the bottom product were usually calculated from the
material balance. For some runs, the composition of
the bottom product was checked following the pro-
cedure used to determine the composition of the reflux.
In Table 5,the experimental results are reported and
compared with those obtained by the theoretical
simulation for each of the eight experimental runs. For
the same reflux, number of theoretical stages, salt
concentration in the reflux, flow rate and composition
of the feed stream, and flow rate of the distillate, the
agreement between the experimental and calculated
composition of the distillate and bottom streams is
quite satisfactory.
CONCLUSIONS
The results obtained indicate that it is possible to
simulate the extractive distillation of ethanol with
calcium chloride with confidence. It is therefore poss-
ible to fix the optimum ranges of the principal
operating variables that characterixe this process.
Using calcium chloride, the energy requirement of
the proposed distillation process has been Assessed to
be 1200kcal/kg. About 10% of this energy is required
to eliminate the crystallization water in order to obtain
anhydrous calcium chloride to be recycled.
When compared with other distillation processes
[S], the consumptions of which range between 2200
and 4500 kcal/kg, this figure appears appealing.
Moreover, the assessed energy requirement lies in the
range typical of the non-distillation processes under
development and recently reviewed by the SERI [6].
But the proposed process appears quite more reliable
than the latter ones.
We believe it is possible to achieve a lower energy
consumption for unit of product by using other, less
hygroscopic, salts or salt mixtures. Particular attention

6. 2292 D. BARBA et al.
in our research is also devoted to the salts used in the
fermentation, such as KCl, M&l, etc.
As a matter of fact, these salts can be directly fed to
the fermentor from the bottom of the dehydration
column with consequent reduction of the salt regener-
ation section.
Acknowledgement-The authors are grateful to Ministero
dehaFubblica Istruzionefox fmancialsupportof this work.
a
.‘
i
akj
ti
D
Oi
L
FL
Fv
9
hi
Hi
kj3
Ki. j
li
JG
M
N
P
P3.j
93
Qc
QR
‘
j
SjFi
T
t
t‘
3
t
U
vi
V
xi. i
NOTATION
generic element of the matrix A, defined in
Table 1.
generic element of the matrix A;j defined in
Table 1
binary parameters of the thermodynamic
model defined in Table 2
generic element of the vector bj defined in
Table 1
distillate Row rate, kmol/h
dielectric constant of component j defined in
Table 2
salt flow rate, kmol/h
liquid feed flow rate, kmol/h
vapour feed flow rate, kmol/h
generic element of the vector g defined in
Table 1
molar liquid enthalpy, stage i, kJ/kmol
molar vapour enthalpy, stage i, kJ/kmol
binary parameters for the thermodynamic
model, defined in Table 2
equilibrium ratio, stage i and component j,
defined in eq. (5)
total liquid flow rate, stage i, kmol/h
solvent flow rate, stage i, kmol/h
number of components
number of stages
total pressure, bar
corrected vapour pressure defined in Ref. 13,
stage i and component j, bar
structural area parameter for component j,
defined in Table 2
modified structural area parameter for com-
ponent j, defined in Table 2
condenser heat load, kJ/h
reboiler heat load, kJ/h
structural volume parameter for component j,
defined in Table 2
mean ionic solvation number at infinite di-
lution in solvent j, defined in Table 2
temperature, K
temperature, “C
vector temperature profile to be corrected by
eq. (41
generic vector temperature profile
unit vector
vapour flow rate, stage i, kmol/h
vector whose elements are V,
liquid mole fraction, stage i and component j
xi,i
&
Yi,j
ZLj
zVj
liquid mole fraction on salt-free base, stage i
and component j
vector whose elements are Xi,i
vapour mole fraction, stage i and component j
mole fraction of componentj in the liquid feed
mole fraction of component j in the vapour
feed
Greek letters
YY
Ah
At,
nj
63
rt1
c21
c31
c41
::;
EC:
c91
Cl01
Cl11
Cl21
Cl31
CL41
CL51
CL61
Cl77
El81
[tQl
PO1
:z:;
c231
1241
activity coefficient, stage i and component j
element of the matrix i
vector correction of the temperature profile to
latent heat of vaporization of component j
defined in Table 2
binary parameters for the thermodynamic
model defined in Table 2
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