Ind. Eng. Chem. Res. 1990, 29, 1349-1363
1349
Multicomponent Three-phase Azeotropic Distillation. 1. Extensive Experimental Data and Simulation Results Brett P. Cairns and Ian A. Furzer* Department of Chemical Engineering, University of Sydney, Sydney, New South Wales, Australia 2006
The first comprehensive data set of compositions of a 13-component mixture on each tray of a distillation column operated in the 3-phase region is presented in this paper. The design of the experiments has been aimed a t covering all the significant regions of the composition space. Distillation columns producing a near heteroazeotrope as overheads can produce widely different bottom products depending on the feed conditions. The composition profiles have been simulated with a modified Naphtali-Sandholm method that includes a phase splitting algorithm. A wide range of group contribution methods has been critically examined for the evaluation of VLE and VLLE data. A comparison of the experimental and predicted composition profiles along the column indicates an overall column efficiency of 70%. Attempts to calculate Murphree vapor efficiencies require a thermodynamic model and except for the major components lead to widely varying numerical values. Overall column efficiencies are preferred for column simulation studies. The collected data set should be valuable to test all three phase distillation models. Three-phase azeotropic distillation is an important separation technique which finds wide use in industry. The process is so-named because the key azeotrope taken overhead condenses to form two liquid phases. This azeotrope is often called a heteroazeotrope, and the process is sometimes referred to as heteroazeotropic distillation. The overhead liquid-phase split raises the possibility of a similar occurrence on the column trays, and therefore, effective simulation of the process can only be accomplished by a three-phase distillation model. Despite the importance of azeotropic distillation and the numerous other examples of three-phase distillation found in industry, there appears to be virtually no experimental tray-to-tray composition data available for columns operating with two liquid phases. Kovach and Seider (1987a,b) have reported experimental results of the feed, distillate, and bottoms compositions but not the tray-totray compositions in a three-phase distillation column. This, combined with scarcity of vapor-liquid-liquid equilibrium (VLLE) data, makes the results obtained from simulation models difficult to accept. This paper presents an extensive set of tray-to-tray multicomponent three-phase azeotropic experimental data at total reflux. The first comprehensive data set for the system ethanol-water-2,2,4-trimethylpentanewas given by Furzer (1985). The new data in this paper complement the previous work of Cairns and Furzer (1987a,b). The system investigated was the dehydration of ethanol using an impure isooctane fraction, containing 11 components, as an entrainer. The experimental results are compared with a newly developed three-phase distillation model. Further, the predictions of the currently available group contribution activity coefficient models-UNIFAC-VLE, UNIFAC-LLE, modified UNIFAC, ASOG, modified ASOG-are critically assessed. Furzer (1988) has made a critical examination of the extremely limited VLLE data published and the modified UNIFAC model. Simulation A column operating at total reflux is depicted in Figure 1. At total reflux, it can be shown that “passing streams are equal”; that is, i = 1, ..., n, n = 1, ..., N - 1 (1) yn+i,l= x,,~ Further, because no streams are withdrawn, the following restriction on the initial charge to the still must be satisfied:
N
FxF,~=
n=l
mnxn,i
i = 1, ..., n,
(2)
where m, is the molar liquid holdup at each stage n. This assumes that the vapor holdup is negligible. Finally, the mole fraction summation for each stage must be held: nc
Cy,,i = 1
i=l
n = 1, ...,N
(3)
Equations 1-3 represent N ( n , + 1) equations in N(n, + 1) unknowns, Nn, liquid-phase mole fractions (x,,$, and N stage temperatures ( T J . By use of the Naphthali-Sandholm (1971) approach of Furzer (1986),the following deviation functions are written: (4) Fn,i = Y n + l , i - xn,i n,
(5)
(7)
These deviation functions equal zero at the solution. Grouping the equations as given in eqs 4-7 gives the following block structure, which has also been presented by Hirose et al. (1980):
or JAx = -F
The Newton-Raphson scheme then updates the variables Xk+l
= xk
+ Ax
k = iteration counter
(10)
and the process is terminated once the deviation function residuals have reached a suitable tolerance. Maximum projection limits are applied to the corrections at each application of eq 10. The limits are user supplied as AX,
osss-~~a~~~o~~~0 ~ 1990 ~ - ~American ~ 4 9 $Chemical 0 ~ . ~ Society 0/0
1350 Ind. Eng. Chem. Res., Vol. 29, No. 7, 1990
This in effect reduces eq 8 to
1
-
Condenser
Stage
1
1 Stage
X
1,i
vN-l
N '
4
1l!
IyN,i
I
i
L*
LN-3
N-2,i
4
X
Reboiler
L
C1
N-1,i
c 2
B" C n BN-l
I
where
and ATmm with suggested values of 0.2 mole fraction and 10 K, respectively, and if
or
I&l> h m a x I A q > ATma
(23)
(11)
(12)
for any iteration the projections are reset to the maximum values. Furthermore, if the projections of the NewtonRaphson scheme produce negative mole fractions, these are reset as small positive numbers, typically lo*. Cairns and Furzer (198813) have shown that this technique provides a robust and flexible way of constraining the Newton-Raphson iterations. Equation 9 can be solved for Ax by using simple Gaussian elimination. The row of D matrices are eliminated by expanding eq 9 to give
Bl&l+ Clk2 = -F1 D1&1+ D2&2 + ... + DN&N = -FN
(13) (14)
Multiplying eq 13 by the product DIB1-l,we have
Dl&i+ DIB1-lC1&P = -DlBl-lFl Subtracting eq 15 from eq 14 gives
(15)
= -FN*(')
and
Wn+l = B,-lF,, or This allows standard Gaussian elimination to be used to solve eqs 26 and 27 and retrieve Zn+land Wn+l. The reason that the sum of the y values is used in eq 3 and not the sum of the x values is that, if the latter is used, the B matrices have a column of zeroes due to the temperature partial derivative. This arises because at total reflux the component deviation functions do not involve the vapor composition leaving the stage but rather the vapor entering the stage, that is, leaving the stage below, which contributes to the C matrices. Using eq 3 introduces the temperature derivative via the equilibrium condition Y n i = Kn,ixnj
(D2 - D1B1-1C1)~2 + D3h3 + ... + DN&N = -(FN -.D1B1-'FJ (16)
+ D3&3 + ... + DN&N
(22)
Furzer (1986) points out that the inverse of the B matrix does not have to be calculated since eqs 22 and 23 can be written as n = 1, ...,N - 1 (24) Dn+l*= Dn+l- DnZn+l FN*(n) = FN*(n-l) - DnWn+1 (25)
LN-l
Figure 1. Total reflux column.
(17)
where
D2* = D2 - DIBrlC1
(18)
- DlBl-'Fl
(19)
FN*(') = F N
B2
Dn+l*= Dn+l*- DnBn-lCn n = 1, ...,N - 1 FN*(n) = FN*(n-l)- DnB;lFn
N-2
1
D2*&2
cN-l
Dn DN-~ DN-
which allows the correction vectors Axn to be calculated by back-substitution using Gaussian elimination. In general, eqs 18 and 19 can be written as
N-2
yN-1,i
D3
This process is carried out N - 1times to give the following
2
N-3,i
Stage
--
Cn BN-l
D2*
x2, i
v[:yN-2,i
Bn
(28)
since Kni is a function of temperature. For an ideal gas, pa
.A,
.
n,i 1 n,i
Kn,i = P where P",i is evaluated by the Antoine equation and Y~~ is calculated by using an appropriate model. The models investigated in this work are UNIFAC (Fredenslund et al., 1975) with parameters correlated against VLE, UNIFACVLE (Gmehling et al., 1982), and LLE; UNIFAC-LLE
Ind. Eng. Chem. Res., Vol. 29, NO. 7, 1990 1351 Table I. Total Reflux Feed Compositions for Methanol-AcetoneChloroform (Mole Fractions) run methanol acetone chloroform 0.90 0.83 0.05 0.05
1 F.
fi
2 3 4
0.05 0.12
0.05 0.05 0.55 0.57
0.40 0.38
D = DISTILLATE B = BOlTOMS
.-8 0.8 0
02 01
-
00 00
"
"
"
01
02
03
04
05
.
06
'
*
07
"
'
08
'
09
IO
CHLOROFORM (mole fraction)
Figure 3. Total reflux simulation, UNIFAC thermodynamics: methanol (1)-acetone (2)-chloroform (3).
Figure 2. Bubble point temperature surface, methanol (1)-acetone (2)-chloroform (3). Temperature-composition axes. A = pure chloroform (61.2 "C), B = chloroform-acetone azeotrope (64.5 "C), C = pure acetone (56.4 "C), D = acetone-methanol azeotrope (54.6 "C),E = pure methanol (64.7 "C), F = methanol-chloroform azeotrope (53.5 "C).
(Magnussen et al., 1981);modified UNIFAC (Larsen et al., 1987); ASOG (Kojima and Tochigi, 1979); and modified ASOG (Tochigi et al., 1981).
Phase Splitting The inclusion of liquid-phase splitting has been outlined in part 2 (Cairns and Furzer, 1990). The liquid phase is checked for stability every time the activity coefficients are evaluated. Derivatives with respect to the average liquid-phase mole fractions (f,,Jfor unstable phases are calculated analytically as described in part 2. Examples. Two different azeotropic mixtures have been simulated at total reflux to illustrate the scope of the method. The first, methanol-acetone-chloroform, represents a homogeneous mixture which exhibits a saddle azeotrope. The second, ethanol-water-2,2,4-trimethylpentane, represents a three-phase azeotropic application which has relevance for the proposed experimental work. Methanol-Acetonexhloroform. Possibly the most investigated multicomponent homogeneous azeotropic mixture is methanol-acetone-chloroform, which was discovered by Ewe11 and Welch (1945). This system is interesting because it exhibits a ternary saddle point azeotrope as well as binary pair maximum and minimum azeotropes. To illustrate how these azeotropes combine to create distillation zones, Figure 2 shows a representation of the bubble point surface based on the predictions of UNIFAC-VLE. The figure clearly illustrates the saddle point and the relationships of the binary azeotropes. Four total reflux runs have been simulated in the different distillation zones, with the feed compositions given
in Table I. A 25 ideal stage column, including reboiler, operating at atmospheric pressure was used in the simulation. The reboiler holdup was assumed to be 85% of the overall mixture, and the remaining 15% was distributed evenly over the other stages. The results are shown in Figure 3 for UNIFAC-VLE thermodynamics with Antoine coefficients taken from Sinnot et al. (1983). The total reflux distillation paths shown in Figure 3 clearly illustrate the effect of the saddle azeotrope. Run 1, for example, had a mixture that was in a zone that produced a distillation path approaching pure methanol at the bottom and the methanol-chloroform azeotrope at the top. On the other hand, run 2 was such that the top of the column approached the methanol-acetone azeotrope while again producing methanol a t the bottoms. Runs 3 and 4 illustrate the same effect for a bottom that approaches the maximum boiling binary acetone-chloroform azeotrope. Cairns and Furzer (198810) have recently used this example to show that different thermodynamic models could place the same overall feed mixtures in different distillation regions and, as illustrated in Figure 3, predict very different distillation paths. Several workers, such as Hoffman (1964) and more recently Van Dongen and Doherty (1985) have noted that the simple distillation residue curve which passes through the bottom composition of a column coincides, in part, with some portion of the total reflux composition profile. Indeed, Maga et al. (1986) have reported experimental residue curves for the methanol-acetone-chloroform system that largely match those shown in Figure 3. Ethanol-Water-2,2,4-Trimethylpentane. The bubble point surface for the heteroazeotropic mixture ethanolwater-2,2,4-trimethylpentane,based on UNIFAC-VLE, is represented in Figure 4. The dew and bubble point surfaces would touch at the pure component corners of the ternary diagram, as well as a t the azeotropes. The heterogeneous azeotropes exhibit minimum lines of temperature rather than the minimum and maximum points, as shown by the homogeneous example in Figure 3. This is illustrated by the binary heteroazeotrope between 2,2,4trimethylpentane and water. The small solubility of the two components is reflected in the constant temperature
1352 Ind. Eng. Chem. Res., Vol. 29, No. 7 , 1990
Figure 5. Expanded view of bubble point surface along heteroazeotropic tie line, Temperature-composition axes.
+ - - - + LIQUD-LlQUlDTE.LwE t . OVERALL LIQUD-VAPOUR TIE-LNf
*-
Figure 4. Bubble point temperature, surface, ethanol (1)-water (2)-2,2,4-trimethylpentane(3). Temperaturecomposition axes. A = pure 2,2,4-trimethylpente (99.22 "C), B = 2,2,4-trimethylpentane-water azeotrope (78.8 "C), C = pure water (100.0 "C),D = water-ethanol azeotrope (78.2 "C), E = pure ethanol (78.3 "0, F= ethanol-2,2,4-trimethylpentaneazeotrope (71.8 "C), G = ternary heteroazeotrope (68.95 "C).
of the miscribility gap that sharply rises to the pure component temperatures at the diagram corners. The diagram also shows the "imprint" of the two-liquid-phase region and that only very small changes in the bubble point temperature occur within this region. Large increases in temperature occur at the binodal curve as the mixture becomes homogeneous. The ethanol corner of the diagram corresponds to the region in which an ethanol dehydration distillation column would operate. The surface shows that, if the distillation path of the column approached the pure ethanol corner along the zero water axis, a steep temperature front would be experienced in the bottom section of the column. On the other hand, if the path approached along a constant water composition line, the steep front would be at the top of the column, leaving the bottom to slowly increase in temperature. The ternary heteroazeotrope is less obvious in Figure 4 than the saddle azeotrope in Figure 3. This is because of the very small temperature gradients in the three-phase region. Figure 5 shows an expanded view of the heteroazeotropic tie line in which the temperature difference between the minimum heteroazeotropic plane and the maximum points shown is only 0.2 "C. The figure illustrates that the bubble point temperature is a minimum along this line. Only at the point where the dew point surface touches this line is the ternary mixture at the heteroazeotrope and the overall liquid composition equal to the vapor composition. A feature of three-phase mixtures is that the vapor composition in equilibrium with two liquid phases remains the same irrespective of the amount of each liquid phase. Furthermore, the bubble point of the mixture is inde-
00
0I
0.2
0.3
0.4
0.5
0.6
VAPOURLINE
07
OS
0.9
10
2.2.4.TRIMETHYLPEhTAhT (mole fraction)
Figure 6. Vapor line: ethanol (1)-water (2)-2,2,4-trimethylpnhne (3).
pendent of the overall liquid composition as long as the equilibrium liquid-phase compositions remain unchanged. This implies that a mixture that lies on any given liquid-liquid tie line has the same vapor composition and bubble point temperature irrespective of its position on the line. If this is extended for all liquid-liquid tie lines, then the series of vapor compositions can be linked to form what is called the "vapor line". This is illustrated in Figure 6 for ethanol-water-2,2,4-trimethylpentaneusing UNIFAC-VLE. Figure 6 shows different vapor compositions in equilibrium with various overall liquid compositions which lie within the three-phase region. The line connecting the vapor points is the vapor line. Any mixture that lies within the three-phase region would boil to produce a vapor composition that lies on the vapor line. Therefore, large differences in composition between the overall liquid and vapor can be experienced for mixtures that exist just inside the binodal curve. The vapor line intersects the water-2,2,4-trimethylp e n m e axis at the binary heteroazeotropic point. Moving away from this axis, the bubble point temperature of each ternary liquid-liquid tie line decreases, as shown in Figure 4, until the minimum heteroazeotropic tie line is reached. The bubble point temperature then increases as illustrated in the magnified view shown in Figure 5. The intersection of the vapor line and the heteroazeotropic tie line defines the heteroazeotropic point. For mixtures that lie on the heteroazeotropic tie line, the vapor composition would lie on the vapor line as well as the liquid-liquid tie line. Only at the heteroazeotropic point, however, will the overall liquid and vapor compositions be equal. Although perhaps less apparent from the bubble point surface shown in Figure 4, the heteroazeotropic system
Ind. Eng. Chem. Res., Vol. 29, No. 7 , 1990 1353 Table 11. Total Reflux Feed Compositions for Et hanol-Water-2,2,4-Trimethylpentane(Mole Fractions) run ethanol water 2,2,4-trimethylpente 1 0.940 0.005 0.055 0.025 2 0.950 0.025 0.025 3 0.850 0.125 0.025 4 0.500 0.475 0.025 5 0.050 0.925 0.025 6 0.010 0.965 0.965 7 0.010 0.025 0.950 8 0.040 0.010
t
______
02
ethanol-water-2,2,4-trimethylpentanealso has distillation zones that control the end products from a column. The total reflux method was again used to simulate the distillation paths for a variety of different feed mixtures, which are listed in Table 11. The, distillation paths are shown in Figure 7 and were calculated for a five-ideal-stage atmospheric column, including reboiler, assuming 85% reboiler holdup. UNIFAC-VLE thermodynamics and Antoine coefficients from Sinnott et al. (1983) were again used. Figure 7 shows that, as expected, all distillate compositions tend toward the heteroazeotrope; however, depending on the feed composition, the bottoms approach different corners of the ternary diagram. Run 1, for example, had a high ethanol feeed content with only traces of water, and the simulation predicts that the total reflux column would approach pure ethanol at the bottom, leaving the top to tend toward the heteroazotrope. Similarly run 2 also approached the ethanol of the ternary diagram; however, because more water and less 2,2,4-trimethylpentane was present in the system, the distillation path shows an approach along a constant water line which intersects the binary ethanol-water axis. As the ethanol content in the feed mixture is lowered, the total reflux paths show the influence of the binary ethanol-water azeotrope, and the bottoms approach the pure water corner of the ternary diagram (runs 3 and 4). For mixtures very low in ethanol and rich in water (runs 5 and 6), the total reflux path approaches the heteroazeotrope from below, as shown in Figure 7 . The distinguishing feature of heteroazeotropic systems is the liquid-phase split. Figure 7 illustrates the effect of the vapor line and demonstrates how it essentially divides the composition space in half. The requirement that passing streams be equal for total reflux means that the distillation paths in the three-phase region must approach the heteroazeotrope along the vapor line. This is because any ideal stage that has two liquid phases must be in equilibrium with a vapor that lies on the vapor line, and since the liquid leaving the stage above is equal to the vapor entering that stage, the liquid must also lie on the vapor line. This is well illustrated in Figure 7 , which shows that all runs tend toward the heteroazeotrope along the vapor line. Furthermore, large changes in composition
HETEROAZEOTROPE BINODAL CURVE AT BUBBLE POINT DISTILLATION PATH
-
2.2.4TMC5 (mole fraction)
Figure 7. Total reflux simulations, UNIFAC-VLE thermodynamics: ethanol (1)-water (2)-2,2,4-trimethylpentane (3).
from stage to stage can result because of this requirement, which is shown by the paths for runs 7 and 8. These runs have an overall feed composition rich in 2,2,4-trimethylpentane, and the total reflux paths show that the heteroazeotrope can be approached along the vapor line from either above or below depending on the amount of ethanol in the system (Table 11).
Experimental Section The objective of this section of the work was to collect at total reflux an extensive set of data on three-phase azeotropic distillation. This was to include measurements of stage temperatures, liquid-phase compositions, and split ratios, which would allow distillation paths to be constructed. The mixture investigated was ethanol, water, and an industrial isooctane (2,2,4-trimethylpentane) fraction referred to as “alkylate”. This mixture was chosen to examine the potential use of isooctane as an entrainer in the production of dehydrated ethanol. Two batches of the alkylate mixture were obtained and used in two distillation columns. Alkylate Analysis. An analysis utilizing high-performance gas chromatography and mass spectrometry was performed to identify and quantify each component in the alkylate batches. Details of the method can be found in Cairns et al. (1987). The first batch (batch 1)was found to contain 10 components ranging from C, to C9 with a boiling point range of 89.7-124.1 “C. The second batch (batch 2) was found to have one extra C, component with a boiling point of 80.5 OC. The results of the analyses are summarized in Table 111. Distillation Columns. Two distillation columns were used in these experiments. The first was a 104-mm internal diameter glass column containing nine stainless steel dual-flow trays at 320-mm tray spacing of 20.19% free area with 8.0-mm-diameter holes. The second was a 610-mm
Table 111. Alkylate Composition comDonent 1 2 3 4 5 6 7 8 9 10 11
name 2,4-dimethylpentane 2,3-dimethylpentane 2,2,4-trimethylpentane 2,5-dimethylhexane 2,4-dimethylhexane 2,2,3-trimethylpentane 2,3,4-trimethylpentane 2,3,3-trimethylpentane 2,3-dimethylhexane 3,4-dimethylhexane 2,2,5-trimethylhexane
symbol 2,4-DMCS 2,3-DMC6 2,2,4-TMCS 2,5-DMC6 2,4-DMC6 2,2,3-TMCS 2,3,4-TMC, 2,3,3-TMCS 2,3-DMC6 3,4-DMCB 2,2,5-TMCe
bp
batch 1, mole fraction
80.5 89.7 99.2 109.1 109.4 109.8 113.4 114.7 115.6 117.7 124.1
0.0105 0.6203 0.0453 0.0589 0.0127 0.1215 0.0346 0.0830 0.0062 0.0070
batch 2, mole fraction 0.0087 0.0172 0.6800 0.0254 0.0422 0.0078 0.1237 0.0318 0.0579 0.0032 0.0021
1354 Ind. Eng. Chem. Res., Vol. 29, No. 7 , 1990
09 0.R L
2n7 L Y nh
2 0.5
0
04
n3 n2 ' I ,
t
0I 8
0.0
0
1
3
2
4
ALKYLATE (mole fraction)
Figure 8. Summary of typical experimental distillation paths.
1.0
internal diameter mild steel column containing four stainless steel sieve trays with downcomers on a 610-mm tray spacing. The trays contained 10% free with 13.0mm-diameter holes with 50-mm weirs. Liquid samples were withdrawn from each tray, separated into two liquid phases where necessary, and analyzed for water by an automated Karl Fischer titration method and for ethanol and hydrocarbons with a GC equipped with a 50-m fused silica open tubular column. Full details of the distillation columns and analysis are given by Cairns (1988).
".y 0.8
5 6 P L A T E NUMBER
10
,
I 2.2.4TMC.5
- c -
2.3.4TMC5
0.I
Experimental Results Glass Column, 104-mm Diameter. A total of 29 total reflux runs, as described by Cairns (1988), were performed in the various distillation regions of the compositional diagram. Figure 8 shows 12 typical runs on a quasi-ternary diagram with all the hydrocarbon components summed to give an overall alkylate composition. The figure shows that all runs tended toward a common heterogeneous azeotrope overhead and that the bottoms approached different compositions along varying paths which were dependent on the makeup of the overall feed mixture. An example of the full results is given for run 1 in Table 1V. The complete set of results for all 29 runs are available as supplementary material. Run 1 was wholly within the three-phase region with the bottoms composition approaching pure alkylate. Table IV shows the composition and fraction of each phase a t each stage. Also given are the stage temperatures and estimated froth heights, along with the calculated F numbers ( F = up,"2). The raw measurements of the volumetric liquid flow rate leaving the bottom plate, plate 9, and the reboiler steam pressure are given, and an average mass balance error is also shown. This figure represents the average of the absolute deviations from 100% obtained when the independent water and gas chromatograph results are summed. The distillation paths have been shown graphically on quasi-ternary diagrams, and where a liquid-phase split was recorded, the tielines linking the liquid phases are also shown. Bottoms Approaching Pure Alkylate. Runs 5 and 6 shown in Figure 8 resulted in the entire column operating in the three-phase region, with an approach to pure alkylate at the bottoms and an approach to the heteroazeotrope at the distillate. Run 28, shown in part on Figure 8, avoided the three-phase region, with the compositions near the zero water axis and the bottoms approaching pure alkylate. Composition profiles along the column for runs 6 and 28 are shown for the major components in Figure 9 and the minor components in Figure 10. The regular profiles obtained for all components highlight the accuracy of the analytical techniques.
9
8
7
0.0 0
1
2
3
4
5
6
1
8
1
9
2,3DMC6
no40 0.045
L2ADM.6
t
z
c 0.035 e2 on30 3 0025 Li
1
I
//
0
0020 0.015 0.010
0.005 0.W
-
0.07
f2,3DMC6
2,4DMC6 2.5DMC6
0.02
o ni 0.00
0
I
2
3
4
6 7 8 PLATE NUMBER 5
9
IO
I1
12
Figure 10. Comparison of minor component profiles for runs 6 (top) and 28 (bottom).
Bottoms Approaching Pure Water. Runs 7 and 8 shown on Figure 8 resulted in the column also operating
Ind. Eng. Chem. Res., Vol. 29, No. 7, 1990 1355 1.0
A
-
-
-
-
-
hl
0.045
PHASE REGION
0.8
5
1
0.050 I
l
0.040
2,3DMC6
z
8
E
I
THREE PHASE REGION
0.035 0030
3 0.025
0
0.020
0.015
0.010
0.2 0.1
0.0
0
1
2
3
4
5
6
0.005
-
-
-
7
8
9
.
0
1
2
3
4
5
6
7
8
-
9
.
1
0
PLATE NUMBER
1.o
0.9
-
c
O.OO0
IO
PLATE NUMBER 0.025 THREE PHAPF
THREE PHASE REGION
I !
0.8
I
0.020
0.7
2 OA 3
3 0.015
E 0.4
0.010
0.5
U.
w
d
0.3 0.2
0.005
0.1 0.0
Figure 11. Comparison of major component profiles for runs 7 (top) and 16 (bottom).
in the three-phase region, but this time the bottoms approached pure water. Runs 14,16, and 18 shown in part in Figure 8 contained compositions near the zero alkylate axis, exhibited maxima in ethanol composition, and approached the heteroazeotrope from above. Composition profiles showing the change from a twophase region to a three-phase region are shown for runs 7 and 16 for the major and minor components in Figures 11 and 12. Bottoms Approaching Pure Ethanol. Runs 26 and 27 shown in Figure 8 resulted in a bottom product approaching pure ethanol. In run 26, a three-phase region existed on the top trays of the column. In both runs 26 and 27, the compositions were close to the zero water axis on many trays. Composition profiles for the major components are shown in Figure 13, and the pronounced maxima for the minor components are shown in Figure 14. Temperature Profiles. The experimental temperature profiles are shown in Figure 15 for some of the typical runs in the different distillation regions. Figure 15a shows that the region around the pseudoternary heteroazeotrope(runs 3-6) had only a small temperature gradient. This is consistent with the ternary bubble point surfaces of Figure 4 for ethanol-water and 2,2,4-TMCS. Runs 7-9 in Figure 15a illustrate a sharp drop in temperature at the top of the column due to the presence of the alkylate components. This is also consistent with Figure 4 where a steep drop in temperature is shown for only small additions of 2,2,4-TMC5to dilute ethanol-water mixtures. The deep penetration of runs 5 and 6 toward the alkylate corner of the quasi-ternary diagram is reflected in Figure 15a by the sharp rise in temperature on the lower stages of the column. This compares with runs 3 and 4, which show only small temperature variations due to the comparatively small composition changes over the length of
O.Oo0
3
Figure 12. Comparison of minor component profiles for NIU7 (top) and 16 (bottom).
0.8
E
0.7
o,2&
~
0.1
0.0
j
~
WATER 2,3,4T C 0
1
2
3
5
4
6
7
8
9
IO
8
9
0
PLATE NUMBER
3 0.5 2,2,4TMCS
2.3.4TMC5 0.0
0
1
2
3
4
5
6
7
PLATE NUMBER
Figure 13. Comparison of major component profiles for runs 26 (top) and 27 (bottom).
the column. Furthermore, the differences between runs 5 and 6 and runs 3 and 4 highlight the steep temperature
1356 Ind. Eng. Chem. Res., Vol. 29, No. 7 , 1990 0027.5
0.02s0
1
,
I
THREE
&PHASE ,
REGION-
TRAY SPACIFiC
~
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0.0075
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2 n.0100 0.0075 0.0050 ao(125
O . m I.
1
2
3
4
5
6
R
7
9
10
PLATE NUMBER
Figure 14. Comparison of minor component profiles for runs 26 (top) and 27 (bottom). 100.0
1
RUN 9
95.0
t / //
I
w.O
2
80.0
n
i
2
3
4
5 6 7 PLATE NUMBER
s
9
in
11
RUN14
I
85.0
1 80.0
I-
p
RLW 16 RUN 26
Y
$ k! 5 . 75.0
70.0
65.0
2
3
4
,
.
~,
5 6 7 PLATE NUMBER
8
9
Figure 16. Froth height profile, run 8.
0.0025
0 . m
1
I
R L 3 27
-
0
I
2
3
4
5 6 7 PLATE NUMBER
8
9
10
11
Figure 15. Experimental temperature profiles (a, top; b, bottom).
front as the pure alkylate corner is approached (see Figure 4).
Figure 15b shows that comparatively gentle temperature gradients were recorded for the typical runs that approached the pseudoternary heteroazeotrope from above. Runs 14 and 16 tended toward pure water in the reboiler and, therefore, had higher temperatures than runs 26 and 27, which both approached a pure ethanol bottom composition. The differences between the distillation paths of runs 26 and 27 are also highlighted in Figure 15b. Run 27 had only a small temperature gradient because it approached the ethanol corner of the diagram along the zero water axis, whereas run 26 approached the same composition along a constant water composition, leading to the ethanol-water axis. These differences in temperature profiles are well illustrated in the bubble point surfaces of ethanol-water and 2,2,4-TMC6 shown in Figure 4. Operating Observations. The observations of Davies et al. (1987) and the work of Ross and Nishioka (1975, 1981) indicate that the capacity of a three-phase distillation column is limited by foaming on the trays that have a composition close to the binodal curve. The froth height observations during the experimental runs tend to agree with this, although in some cases large foams were recorded on trays away from the binodal curve. This is illustrated in Figure 16 for run 8 where the upper most tray had a froth height twice that of a normal tray spacing. It was noted that the froth was very foamy, while the lower stages had a spray/foam mixture with clearly visible liquid droplets. Figure 8 shows that the top section of the column for run 8 was well away from the binodal curve. In general, the froths on the trays with three-phase mixtures tended to appear foamy with no visible small spray droplets. The samples withdrawn for analysis always entered the syringe as a cloudy emulsion, which then separated into two phases in the syringe barrel. It was not possible to distinguish if one phase was dispersed in the other, and at no time was a tray observed to have a alkylate layer resting above an aqueous layer. Runs 1-6 where the entire column operated within the three-phase region were, in general, difficult runs to perform. This was because the liquid-phase split in the bottom section of the column created problems with the reboiler during start-up. Although a large thermosyphon circulation of liquid was usually established, this did not appear to be sufficient to mix the liquid phases, and often phase separation was observed in the pipework leading to the reboiler. The effect of a layered flow into the reboiler at the beginning of the run was to set up a cyclic vapor flow pattern which could only be stabilized by manipulating the steam load to the reboiler. The problem was that the hydrocarbon phase tended to vaporize first, sending a large pulse of vapor to the top of the column and flooding that section. As the water phase entered the reboiler, less vapor was generated, and the result was that the hydrocarbon at the top of the
Ind. Eng. Chem. Res., Vol. 29, No. 7, 1990 1357 Table IV. Ethanol-Water-Alkylate Distillation Experimental Results" DISTILLATION PATH RUN I
TlEUHB
t----a
0.8
0.2-
ai
-
0.0
0.1
component
cond
1
2
water ethanol 2,3-DMC, 2,2,4-TMCS 2.5-DMCa
0.342 79 0.639 36 0.001 15 0.016 31 0.000 10 0.000 13 0 0.000 16 0 0 0
0.328 44 0.644 59 0.001 56 0.024 36 0.000 22 0.000 30 0 0.000 33 0 0.000 21 0 0
0.330 99 0.643 37 0.001 25 0.023 02 0.000 26 0.000 37 0 0.00041 0 0.000 23 0 0.000 10
0
water ethanol 2,3-DMC5 2,2,4-TMC5 2,5-DMC6 2,4-DMCe 2,2,3-TMCS 2,3,4-TMC5 2,3,3-TMCS 2,3-DMCe 3,4-DMC6 2.2.5-TMCa
0.007 30 0.159 83 0.05308 0.749 77 0.006 56 0.008 06 0.001 72 0.007 99 0.001 17 0.004 51 0 0
staee cond 1
2 3 4 5 6 7 8 9 rboil
0.2
0.008 62 0.008 89 0.145 85 0.109 70 0.043 49 0.037 85 0.760 43 0.790 76 0.008 62 0.01044 0.011 12 0.013 42 0.002 87 0.002 97 0.011 11 0.014 86 0.00181 0.002 5 0.006 08 0.008 56 0 0 0 0 phase splits, mole aaueous 0.631 23 0.639 24 0.622 57 0.644 23 0.613 37 0.642 26 0.60000 0.624 60 0.536 30 0.373 00 0.211 57
0,4 0.5 0.6 0.7 ALKYLATE (MOLE FfUCnOh')
stage 3 4 5 6 Aqueous Phase, Mole Fraction 0.33856 0.34846 0.37837 0.398 12 0.636 65 0.630 89 0.605 84 0.589 20 0.000 95 0.000 62 0.000 36 0.000 26 0.021 94 0.017 77 0.013 56 0.01062 0.000 31 0.000 34 0.000 30 0.000 26 0.000 43 0.000 55 0.000 40 0.000 37 0.00015 0.00020 0 0 0.000 60 0.000 64 0.000 60 0.000 63 0 0.000 10 0.000 11 0.000 11 0.000 40 0.000 44 0.000 45 0.000 43 0 0 0 0 0
0
0
0
0.8
0.9
1.0
7
8
9
rboil
0.443 75 0.547 55 0.000 16 0.006 94 0.000 18 0.000 30 0 0.000 54 0.000 14 0.000 44 0 0
0.546 09 0.451 27 0 0.002 10 0.00008 0.000 11 0 0.000 22 0 0.000 12 0 0
0.728 76 0.271 24 0 0 0
0.960 88 0.039 12 0 0 0 0
0.002 96 0.041 40 0.011 57 0.711 23 0.033 74 0.043 83 0.009 53 0.074 41 0.018 46 0.047 66 0.002 99 0.002 25
0.001 50 0.021 32 0.008 71 0.668 02 0.040 15 0.05343 0.011 38 0.097 58 0.026 70 0.063 18 0.004 44 0.003 58
Alkylate Phase, Mole Fraction 0.008 37 0.007 94 0.007 94 0.006 01 0.004 85 0.100 33 0.094 49 0.082 96 0.076 01 0.062 43 0.030 39 0.024 89 0.019 67 0.017 40 0.014 57 0.785 41 0.775 75 0.761 81 0.751 84 0.739 91 0.013 84 0.017 18 0.021 15 0.023 44 0.027 29 0.017 58 0.02208 0.027 24 0.03144 0.035 78 0.003 89 0.004 77 0.005 84 0.006 86 0.007 72 0.022 19 0.028 87 0.039 28 0.046 39 0.055 84 0.004 53 0.005 88 0.008 56 0.010 82 0.012 93 0.012 91 0.017 39 0.023 99 0.02846 0.035 37 0.000 55 0.000 77 0.001 55 0.001 33 0.001 95 0.001 35 0 0 0 0 fraction alkvlate tema. "C froth heieht. 0.368 77 0.360 76 68.0 15.5 0.377 43 68.2 13.5 0.355 77 12.5 0.386 63 68.5 11.5 0.357 74 69.1 10.5 0.400 00 10.5 0.375 40 69.8 10.0 0.463 70 71.4 10.0 0.627 00 77.6 10.5 0.78843
cm
0 0 0 0 0 0 0
0 0 0 0 0 0
O.OO0 57 0.004 17 0.005 69 0.565 37 0.04890 0.066 56 0.014 27 0.13832 0.041 46 0.098 81 0.007 51 0.007 71
F. SI units 1.704 1.703 1.721 1.695 1.715 1.675 1.706 1.671 1.752
"Run 1;date performed, 20/5/1986; comments, none. Column operating conditions: reboiler steam pressure = 16.0 kPa g; liquid flow rate leaving plate 9 = 2.34 L/min; column average F = 1.705 SI units; average mass balance error = 1.317%.
column could not be sustained and would fall back toward the reboiler. Eventually the hydrocarbon would reenter the reboiler and the pulsing cycle would repeat. It is significant to note that to vaporize 1 mol of 2,2,4-trimethylpentane only requires approximately three-quarters
of the energy needed to vaporize the same quantity of water. By manipulating the steam flow to the reboiler during the upward surge of hydrocarbon vapor, it was possible, after some time, to arrive at a condition in which the two-liquid-phase flow into the reboiler produced a
1358 Ind. Eng. Chem. Res., Vol. 29, No. 7 , 1990 Table V. Composition of Simulation Feed Mixtures feed composition, mole fraction comDonent run 27 run 1 0.207 50 0.001 92 0.048 66 0.887 15 0.000 54 0.005 07 0.057 45 0.432 33 0.006 44 0.035 60 0.007 99 0.04841 0.001 73 0.01038 0.01771 0.099 98 0.00504 0.029 86 0.012 03 0.071 31 0.000 91 0.005 39 0.001 08 0.005 51
u
.
-
0.9
-
t
o.6
".,
- SMWLATION
* \
OM
0.05
0.10
0.1s
0.20
0.2s
0.30
0.35
ALKYLATE (mole fraction)
0.40
0.4.5
1
1.0
050
1
steady exit vapor flow.
Simulation of Experimental Results Runs 27 and 1were simulated by using the total reflux method described above. All 12 components were included in the simulations, and the predictions of the group contribution thermodynamic models (UNIFAC, modified UNIFAC, ASOG, and modified ASOG) were examined. The binary interaction models NRTL (Renon and Prausnitz, 1968) and UNIQUAC (Abrams and Prausnitz, 1975) could not be used, as no data exist to cover all possible interactions. Antoine coefficients for all components were taken from Sinnott et al. (1983), and a 1 kPa pressure drop per stage was assumed. The reboiler was assumed to retain 90% of the overall liquid mixture, with the remaining 10% distributed evenly amongst the other stages. Seven ideal stages were found to best represent the experimental data for both runs, and hence the simulations indicated that the column was operating at 70% overall column efficiency. The results of the simulations for both runs 27 and 1 are discussed in more detail below. The compositions of the feed mixtures are given in Table V. Run 27 Simulation Comparisons. Figure 17a compares the experimental distillation path with the simulation path obtained by using UNIFAC-VLE. This figure represents an expanded view of the distillation path, with all the hydrocarbon components being summed to give an alkylate composition a t each stage. The composition of the heteroazeotrope predicted by UNIFAC-VLE appears to be richer in hydrocarbon than suggested by the experimental path. The result is that the simulation path does not match the observed values a t the top of the column, as shown in Figure 17a. The results of scaling up the simulation UNIFAC-VLE predictions for a seven-ideal-stage column using an overall column efficiency of 70% are shown in the component profiles of Figure 17b,c. The predicted major component profiles (Figure 17b) show good agreement with the experimental values for the bottom section of the column, while some deviations are apparent toward the top. This is due to the errors of UNIFAC-VLE in predicting the composition of the heteroazeotrope. The predictions of the minor component profiles (Figure 17c) show excellent agreement considering the small quantities involved, and again the largest errors occur in the top section of the column. The predictions of the other thermodynamic models are shown in Figure 18 for 70% overall column efficiency. All models predict an accurate bottom composition; however, the overhead vapor compositions varied. The predictions obtained by using modified UNIFAC gave the best esti-
.
-
PLATE NUMBER
Figure 17. Simulation Vs,run 27. (a, top) Seven ideal stages, 90% reboiler holdup, UNIFAC-VLE thermodynamics. (b, middle) Major components, 70% overall efficiency, UNIFAC-VLE thermodynamics. (c, bottom) Minor components, 70% overall efficiency, UNIFAC-VLE thermodynamics. 1.0
0.9
0 0.8
2 0.6 3
I
F
ETHANOL
0.7
. W A C ---- MODIJNIFAC ------ ASOO MODASOO
OS
'
0.4
0.3
2.2.4TMC5 P---
==--=
Y -
0.2 0.1
nn
WATER 0
1
2
3
4
5 6 PLATE NUMBER
7
8
9
,
800 I
'2-7 7 . 5
IO
1 UNlFAC ___UNIFAC -_ _ -_ _ MOD ASOG MOD A S M
650
0
'
1
2
3
4
5
6
7
8
9
I
IO
PLATE NUMBER
Figure 18. Simulation V's, run 27. Comparison of thermodynamic models. (a, top) Composition profiles. (b, bottom) Temperature profiles.
Ind. Eng. Chem. Res., Vol. 29, NO. 7, 1990 1359 1.0
p
I
EXPERIMENr __-. SIMULATION
0.8
0.1 0.5
-
.-S
0.1
0.0
0.2 0.3 ALKYLATE (mole fraction)
0.0
0.5
0.4
0
0.0800
1.0
I
-e8
1
g
EXP" WAC-VLE WAC-LLE
4
5 6 PLATE NUMBER
7
8
1
10
9
12.3DMC6
z o.Ln00
2 0.9
3
2
-
I
EYPERIMENT
0.06oO
E 0.0500 Y 0 0.0000
v
0.8
z
