Ind. Eng. Chem. Res. 2007, 46, 9223-9231
9223
Reducing Diameters of Distillation Columns with Largest Calculated Diameter at the Bottom Phillip C. Wankat* School of Chemical Engineering, Purdue UniVersity, 480 Stadium Mall DriVe, West Lafayette, Indiana 47907-2100
Methods are developed to reduce column diameter for distillation when the calculated diameter is largest at the bottom of the column. This diameter can often be reduced by generating vapor above the reboiler with an intermediate reboiler or with two-enthalpy feed. For a saturated liquid feed that is 50% n-butane and 50% n-pentane the two-enthalpy feed system reduced column volume by 8.2% with no change in energy use. For distillation of dilute acetic acid (5%)-water (95%) mixtures the best system (an intermediate reboiler) resulted in decreases in column volume as large as 53.9% with no increase in energy use. For a 50% acetic acid mixture the two-enthalpy feed system resulted in up to 32.5% reduction in column volume, while vaporizing the entire feed resulted in 31.5% volume reduction. Vapor bypass was effective in concentrated absorbers. In an earlier paper1 we showed how to reduce the diameter of distillation columns with vapor feeds when the largest calculated diameter is in the enriching section. The driving philosophy behind this research was to consider a large difference in the calculated diameters as potential to either reduce the column volume or process more material in a retrofitted column. Essentially, by condensing some of the vapor below the column condenser, the vapor flow rate at the location with the largest calculated diameter was reduced. The methods explored were to use an intermediate condenser,1,2 condition the feed by cooling the entire feed stream,1,3,4 and use a twoenthalpy feed that cools only part of the feed stream.1,5 In this paper we explore reducing the diameter of distillation columns when the largest calculated diameter is at the bottom of the column. Most of the cases involve liquid feeds. The strategy employed is to reduce the vapor flow rate at the bottom of the column by generating some of the vapor required for the separation at a higher location in the column. Four approaches for doing this will be examined: (1) Use an intermediate reboiler,2 which consists of withdrawing a liquid stream from the column, vaporizing it, and then returning it to the column as a vapor. (2) Condition the feed by heating or vaporizing the entire feed stream.3,4 (3) Heating or vaporizing a part of the feed while leaving the other part vapor with a two-enthalpy feed system.5 In two-enthalpy feed operation the original liquid feed is split into a liquid part FL and a part FV that is heated or vaporized in a heat exchanger. (4) Use a vapor bypass to input vapor from the reboiler or from vaporized bottoms product at a location above the bottom of the column. With these systems there will either be an energy penalty resulting in an increase in overall heating and cooling loads or more stages will be needed to produce the desired purity. However, since the column diameter will be decreased, there will often be capital cost savings that can be significant. These methods for balancing the calculated diameters are applicable to standard and high capacity trays, and to standard and high capacity packings. Base Cases for Standard Distillation Systems To illustrate the differences in calculated diameters that are possible, we will first consider three base cases. The columns * Tel.: (765) 494-0814. Fax: (765) 494-0805. E-mail: wankat@ ecn.purdue.edu.
use sieve trays, and diameters were calculated with the procedure developed by Fair6 and incorporated in the commercial AspenPlus 2004 process simulator. The optimum feed trays that gave the maximum product purities are reported. Liquid Feed of 50-50 n-Butane and n-Pentane Base Case. A distillation column is separating 1000 kmol/h of a saturated liquid feed that is 50 mol % n-butane and 50 mol % n-pentane. The column has 28 stages plus a partial reboiler and a total condenser (N ) 30 in AspenPlus notation). Calculations were done with the AspenPlus simulator using the Peng-Robinson vapor-liquid equilibrium (VLE) correlation. For an external reflux ratio ) 0.879, the distillate was 99.90 mol % n-butane and the bottoms was 99.90 mol % n-pentane. Table 1 shows the results and conditions for this base case, and the calculated diameters are listed in Table 2. Although the vapor flow rate is highest at the top, the calculated diameter is highest at the bottom. Acetic Acid (5 mol %), Water (95 mol %) Liquid Feed Base Case. The acetic acid-water system is economically quite important and represents a rather difficult separation with highenergy loads. In fact, the separation is difficult enough that an entrainer is often used.7 Table 1 lists the details of this base case, and Table 2 shows the diameters calculated for this distillation column. The column has 40 stages plus a kettle reboiler, a total condenser (N ) 42), and L/D ) 2.3. Although the NRTL-HOC VLE correlation is recommended,7 we found a better fit to the VLE data in Perry’s Handbook8 with the Wilson HOC VLE correlation regressed to this data. The percentage error in measured and predicted values of (y - x) of the more volatile component (MVC) at equilibrium was calculated.
% error in (yMVC - xMVC) ) 100{[(y - x)predicted - (y - x)data]/(y - x)data} (1) This equation measures the error in the distance between the equilibrium curve and the total reflux operating line. The percent error in (yW - xW) ) 0.3% at xW ) 0.3063, 2.2% at xW ) 0.6750, and 4.6% at xW ) 0.9578. Thus, the fit is good at low and moderate mole fractions of water, but slightly overstates the difference between vapor and liquid mole fractions at high mole fractions of water. The results for the diameter calculations for this separation should be accurate, but the number of
10.1021/ie0709887 CCC: $37.00 © 2007 American Chemical Society Published on Web 11/23/2007
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Ind. Eng. Chem. Res., Vol. 46, No. 26, 2007
Table 1. Simulation Conditions and Results for Base Casesa NFeed
N
dia
19
30
3.10
25
42
6.14
9
20
1.20
A
vol
tray
QR
Liquid Feed of 50-50 n-Butane and n-Pentane Base Case 7.54 133.3 29 6 150
Qc
L/D
D
-5 860
0.879
500
2.3
950
3.6
50
Acetic Acid (5 mol %), Water (95 mol %) Liquid Feed Base Case 29.56 554.1 41 35 560 -35 550 Methanol (5 mol %), Water (95 mol %) Liquid Feed Base Case 1.13 9.8 2 2 410
-2 270
F ) 1000 kmol/h, all feeds are saturated liquids, D is distillate flow rate in kmol/h, and N ) number of trays + condenser + reboiler. Tray spacing is 24 in. ) 0.6096 m (butane-pentane) or tray spacing ) 18 in. ) 0.4572 m (acetic acid-water and methanol-water). Operation is at 80% of flooding. dia ) maximum calculated diameter in m, tray ) tray at which maximum diameter occurred, A ) maximum calculated column area in m2, vol ) column volume in m3 ) A(N - 2 + 1)(tray spacing), where N - 2 is the number of trays and the +1 is for disengagement space for liquid and vapor, QR and Qc are the reboiler and condenser duties in kW, p ) column pressure in atm ) 1.0, pressure drop in psi/tray ) 0, Nfeed is the optimum feed stage, and condenser is labeled as stage 1 in all cases. a
Table 2. Diameters Calculated for Standard Distillation Base Casesa n-butane, n-pentane, liquid
a
acetic acidwater, liquid
methanolwater, liquid
tray
V
dia
A
tray
dia
A
tray
dia
A
2 18 19 29
940 887 885 855
2.64 2.60 2.94 3.10
5.47 5.29 6.77 7.54
2 25 41
3.64 3.86 6.14
10.42 11.68 29.56
2 8 9 19
1.20 0.98 1.18 1.15
1.13 0.76 1.09 1.04
Conditions and results are listed in Table 1. Vapor flow rate V is in kmol/h.
equilibrium stages required in the enriching section may be slightly underpredicted. The distillate is 99.79 mol % water, and the bottoms is 4.00 mol % water. Table 2 shows there is little variation in the calculated diameter in the enriching section, and the diameter is largest at the bottom of the column. A single column with a diameter of ∼6.1 m would be built. Methanol (5 mol %), Water (95 mol %) Liquid Feed Base Case. Table 1 also shows the results and conditions for this base case, and the calculated diameters are listed in Table 2. The column has 18 stages plus a kettle reboiler and a total condenser (N ) 20). The distillate was 95.45 mol % methanol, and the bottoms was 99.76 mol % water. Calculations were done with the AspenPlus simulator using the NRTL VLE correlation. The largest calculated error in (yM - xM) was 3.6%; thus, the correlation fits the VLE data in Perry’s Handbook8 quite well. In this case the variation in the calculated column diameter is modest, and one would probably not try to improve performance with the methods employed in this paper. The first example in Table 2 shows a modest variation in column diameter, while the second example represents a fairly extreme, but not unusual case of diameter variation. Concentrated absorbers also show a larger diameter at the bottom of the column. There are also many systems similar to the separation of a liquid feed of methanol and water that show very modest variations in calculated column diameter. If there is no change in calculated diameter, then diameter control methods are not needed. Methods To Adjust Column Diameters The use of an intermediate reboiler2 to control the diameter of a column is well-known and is obviously related to our previous study with intermediate condensers for vapor feeds.1 This method is shown in Figure 1. Use of two-enthalpy feed to control column diameter for liquid feeds has not been reported in the literature. The two-enthalpy feed system is shown in Figure 2 when the original feed is a liquid.5 In normal operation the entire feed is often heated (set FL ) 0 in Figure 2) and will be partially or totally vaporized. In two-enthalpy feed operation the feed is split into a liquid part FL and a part FV that is vaporized in the heat exchanger. It is not necessary to totally
vaporize stream FV. What is of critical importance is to separate the feed into two parts before one part is heated or vaporized. The method is also useful if the original feed will flash and become a two-phase feed in the column. This approach is similar to the use of two-enthalpy feed to condense part of a vapor feed.1 Conditioning a liquid feed by heating or vaporizing the entire feed stream3,4 is well-known and is similar to condensing a vapor feed.1 We also discovered that vapor bypass can be useful for diameter control in certain distillation applications and for concentrated absorbers may be the method of choice. One version of vapor bypass for distillation is shown in Figure 3a, where part of the vapor from the reboiler bypasses the bottom stages in the column and is input higher up the column. This bypass vapor is in equilibrium with the bottoms product. An alternative if the reboiler does not have sufficient capacity is shown in Figure 3b, where part of the bottoms stream is vaporized and input into the column. This bypass vapor is the same mole fraction as the bottoms product. In general, the bypass can withdraw vapor from anywhere in the column and feed it at a different location. In absorption columns part of the feed gas bypasses the bottom one or more stages in the column (Figure 4). Application to Distillation of Liquid Feeds of Light Hydrocarbons In the distillation of homologous series the calculated diameter is usually largest at the bottom of the column when the feed is a liquid. The differences in calculated diameter depend upon the physical properties of the compounds. Generally speaking, as the ratio of molecular weights becomes larger, the difference in calculated diameters will also become larger. We will study applications of vapor bypass, two-enthalpy feed, and intermediate reboilers for the separation of n-butane from n-pentane and for the ternary separation of propane, n-butane, and n-pentane. Vapor Bypass for Separation of 50-50 n-Butane from n-Pentane. The vapor bypass systems shown in Figure 3 were studied, and the results are given in Table 3. In general, as the bypass rate increased at constant QR,total the calculated diameter at the bottom of the column (tray 29) decreased, the diameter
Ind. Eng. Chem. Res., Vol. 46, No. 26, 2007 9225 Table 3. Results for Vapor Bypass, Two-Enthalpy Feed, and Intermediate Reboiler for a Liquid Feed That Is 50% n-Butane and 50% n-Pentanea FV
NF,L
0 (base)
19
NF,V
N
dia
A
vol
tray
QR,total
Qc
) 30
3.10
7.54
133.3
29
6150
-5860
Vapor Bypass (N ) 30) FBy
NF,V
NBy,V
78.9 78.9 95.0
18 18 18
23 23 22
Figure
dia
A
vol
tray
QR,total
Qc
decr vol (%)
incr QR,total (%)
3b 3a 3a
2.99 2.99 2.97
7.03 7.03 6.94
124.3 124.3 122.7
23/29 23/29 22/29
6190 6190 6210
-5900 -5900 -5920
6.6 6.6 7.9
0.6 0.6 0.9
Two-Enthalpy Feeds FV
NF,L
NF,V
N
dia
A
vol
tray
QR,total
Qvaporize
Qc
decr vol (%)
incr QR,total (%)
107 115
19 19
22 23
30 31
2.93 2.92
6.74 6.69
119.2 122.3
22/29 23/19
6200 6150
757.8 814.5
-5920 -5860
10.5 8.2
0.9 0
NF
N
dia
A
vol
tray
QR,total
Qheat_feed
Qc
incr vol (%)
incr QR,total (%)
xD,B
19
31
3.26
8.34
152.5
19
6150
814.5
-5860
+14.4
0
0.9936
Two-Phase Feed (Qheat_feed ) 814.5 kW)
Intermediate Reboilers (N ) 30): Fwithdr
NF,liq
NL,with
NV,ret
dia
A
vol
tray
QR,total
Qc
decr vol (%)
incr QR,total (%)
89.8 107
19 19
22 21
23 22
2.95 2.94
6.82 6.80
120.6 120.2
22/23/29 21/19
6170 6190
-5880 -5890
9.4 9.8
0.3 0.5
a Unless specified otherwise, mole fractions of n-butane in the distillate and n-pentane in the bottoms are both 0.9990. Pressure is 1.0 atm, and tray spacing is 24 in. ) 0.6096 m. Total condenser is labeled stage 1. FV is the kmol/h in the vapor portion of the feed for two-enthalpy feed, FBy is the kmol/h vapor in the vapor bypass, and Fwithdr is the kmol/h that is withdrawn and then vaporized in the intermediate reboiler. NF,L is the optimum feed stage for the liquid feed, NF,V is the optimum feed stage for the vapor portion of the feed in two-enthalpy feed systems, NBy,V is the feed location for the vapor bypass, NL,with is the optimum withdrawal plate for the liquid that is sent to the intermediate reboiler, NV,ret is the return location for the vapor for an intermediate reboiler. QR,total is the sum of the energy required in the reboiler QR plus the energy Qvaporize required to vaporize the vapor bypass, the portion of the feed, or in the intermediate reboiler, or QR,total ) QR + Qheat_feed, where Qheat_feed is the energy used to heat the entire feed to produce a two-phase feed.
Figure 1. Intermediate reboiler system.
Figure 2. Two-enthalpy feed for a liquid feed.
at the return stage increased, and the purity decreased. As the vapor was returned to a higher stage (lower number) in the column, the diameter at the return stage decreased, but the purity also decreased. Purity could be increased by increasing QR,total, which also increased both diameters. The rate of the vapor bypass, the stage to return the bypass, and the value of QR,total were varied until the calculated column diameters were balanced at stage 29 and the return stage. For example, in the Figure 3b result in Table 3, the calculated diameters were the same for return tray 23 and for tray 29. The notation 23/29 for “tray” indicates the largest calculated diameter occurred simultaneously at both trays 23 and 29. Note that there are also conditions where the diameters will match at return tray 24 and tray 29, at return tray 22 and tray 29, and so forth. By returning the bypass vapor higher up the column, increasing the bypass rate, and increasing QR,total, a slightly lower column diameter resulted. The calculation was first done for Figure 3b, and then the same conditions
were used for Figure 3a. Since the results are essentially identical, vapor bypass results for other systems will be done only for Figure 3b, which did not have the convergence problems that occurred for Figure 3a with high withdrawal rates. The improvements (6.6-7.9% decrease in column volume) shown in Table 3 for vapor bypass are modest. Two-Enthalpy Feed or Intermediate Reboiler for Separation of 50% n-Butane, 50% n-Pentane Liquid Feed. We can also use the two-enthalpy feed system for this separation by vaporizing a portion of the feed. The results in Table 3 show larger but still modest decreases in column volume. The result for FV ) 107 was obtained with the identical column as the base case by increasing QR until the purities matched the base case. The result for FV ) 115 used the same QR,total as the base case, but added an additional stage to match the base case purity. Because of the added height of the additional stage, the column volume decreases less than for the previous system. Since the
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Ind. Eng. Chem. Res., Vol. 46, No. 26, 2007
Figure 3. Vapor bypass systems for distillation. (a) Bypass part of vapor from reboiler. (b) Vaporize and bypass part of bottoms stream.
Figure 4. Vapor bypass for absorption.
column diameters were identical at two trays, increasing FV further had no effect on the column diameter and eventually will cause the purity to decrease below that of the base case. The results obtained for the intermediate reboiler (Table 3) are very similar to those obtained with the two-enthalpy feed system. Table 3 also shows the results for a two-phase feed using the same amount of energy to heat the entire feed as the twoenthalpy feed system used to vaporize FV (FV ) 115). Unfortunately, this system increases the maximum calculated diameter of the column and does not meet the product specification. The distillate purity drops because a two-phase feed increases the minimum reflux ratio;1 thus, operation at the same L/D as the two-enthalpy feed system is at a lower multiplier of the minimum reflux ratio. Separation of 25% n-Butane, 75% n-Pentane and 75% n-Butane, 25% n-Pentane Liquid Feeds. To explore the effect of feed concentration, the results for the separation of 25-75 and 75-25 mixtures of n-butane and n-pentane are given in Table 4. For all feed concentrations the two-enthalpy feed has the largest decrease in column volume, but at the price of a larger increase in QR. The intermediate reboiler system has slightly less decrease in column volume, but the smallest increase in QR. The vapor bypass system appears to work better at low feed mole fractions of the more volatile component than it does at 50 or 75% n-butane in the feed. The intermediate reboiler and two-enthalpy feed systems appear to have a higher volume reduction at 50 mol % n-butane and lower volume reduction at 75 mol % n-butane in the feed. The increases in energy use are very modest. Ternary Distillation of Propane, n-Butane, and n-Pentane. Although all the previous examples in this paper and in ref 1
have been for binary systems, these methods also apply to multicomponent separations. For the ternary base case a distillation column is separating 1000 kmol/h of a saturated liquid feed that is 30 mol % propane, 25 mol % n-butane, and 45 mol % n-pentane. The column has 19 stages plus a kettle reboiler and a total condenser (N ) 21). Calculations were done with the AspenPlus simulator using the Peng-Robinson VLE correlation. Propane was selected as the light key and the distillate flow rate was 300 kmol/h. For an external reflux ratio ) 0.45, the distillate was 95.82 mol % propane and 4.18 mol % n-butane, and the bottoms was 1.79 mol % propane, 33.92 mol % n-butane, and the remainder n-pentane. The calculated diameters and vapor flow rates are stage 2 ) 1.55 m, V ) 435.0 kmol/h; stage 10 ) 1.51 m, V ) 407.3 kmol/h; stage 11 ) 1.92 m, V ) 398.4 kmol/h; and stage 20 ) 2.12 m, V ) 400.2 kmol/h. Although the vapor flow rate is highest at the top, the calculated diameter is highest at the bottom. Other information for the base case is in Table 5. The results for vapor bypass, two-enthalpy feed, and intermediate reboiler systems are shown in Table 5. The optimum values of the flow rate of material vaporized FV and the optimum input and (for intermediate reboiler) withdrawal stages were found by matching the column diameters at stage 20 and one of the input stages and then increasing the absolute value of Qc until the propane mole fraction in the distillate matched the base case (95.82 mol %). For this ternary hydrocarbon system the two-enthalpy feed system had the largest decrease in column area. Comparing the results in Table 5 to those in Tables 3 and 4 for the separation of n-butane from n-pentane, we see that the reductions in column area are larger for this ternary system for all three techniques. This occurred because the ratio of the smallest calculated diameter divided by the largest calculated diameter is smaller for the ternary base case (0.71) than for the 50-50 binary separation (0.84). Thus, there is more room for improvement with this ternary system. This ratio is smaller for the ternary base case to a large extent because the molecular weights of the distillate and bottoms differ by a larger amount. We also simulated a ternary distillation with the same system but with n-butane as the light key (D ) 550 kmol/h) (not shown in Table 5). The column had a total condenser, a kettle reboiler, 19 equilibrium trays (N ) 21), and liquid feed on stage 12. Both liquid and vapor flow rates decreased from stage 15 to stage 20, but the largest diameter was on the bottom tray (no. 20). Although there is less variation in calculated diameter in this case, the qualitative conclusions are essentially the same as for the base case in Table 5. Variations in physical properties can cause significant increases in the calculated diameter even when vapor or liquid flow rates decrease. Application of Processes for Distillation of Acetic Acid and Water The results in Table 1 show that, with a liquid feed that is dilute in acetic acid, the calculated diameter is largest at the bottom of the distillation column. We will see later in this section that this statement is also true for more concentrated (50-50) feeds. For a liquid feed of water-acetic acid the calculated diameters will be balanced by vaporization of part of the feed, with vapor bypass, or with an intermediate reboiler. Distillation of a 95 mol % Water, 5 mol % Acetic Acid Liquid Feed. The base case was presented in Table 1 and is repeated in Table 6, and the calculated diameters for the base case were presented in Table 2. The column had 40 equilibrium trays, a total condenser, and a kettle reboiler (N ) 42). The two-enthalpy feed system was employed by vaporizing a portion
Ind. Eng. Chem. Res., Vol. 46, No. 26, 2007 9227 Table 4. Results for Base Case, Vapor Bypass, Two-Enthalpy Feed, and Intermediate Reboiler for Liquid Feeds That Are 25 mol % n-Butane, 75% n-Pentane and 75% n-Butane, 25% n-Pentanea 25% n-Butane and 75% n-Pentane FV
NF,L
0 (base)
13
87
13
FBy
19
NF,V
NBy,V
13
21
58 Fwithdr
NF,V
dia
A
vol
2.70
5.72
101.1
2.57
5.17
Figure
dia
A
3b
2.60
5.32
NF,L
NL,with
NV,ret
dia
A
13
23
22
2.59
5.28
55
QR,total
Qc
29
4330
-4045
decr vol (%)
incr QR,total (%)
Two-Enthalpy Feed 91.4 19
4360
Vapor Bypass vol tray
-4070
9.6
0.7
QR,total
Qc
decr vol (%)
incr QR,total (%)
21/29
4340
-4050
7.0
0.2
Intermediate Reboiler vol tray
QR,total
Qc
decr vol (%)
incr QR,total (%)
4330
-4050
7.8
0.07
decr vol (%)
incr QR,total (%)
94.0
93.3
tray
22/29
75% n-Butane and 25% n-Pentane FV
NF,L
0 (base)
17
160
21
26
NF,V
NBy,V
21
24
NF,L 21
FBy 60 Fwithdr 49
NF,V
dia
A
vol
3.27
8.40
148.5
3.15
7.81
Figure
dia
A
3b
3.20
8.03
NL,with
NV,ret
dia
A
27
26
3.20
8.02
tray
QR,total
Qc
29
7410
-7260
Two-Enthalpy Feed 138.1 21
7410
-7260
7.1
0.07
Vapor Bypass vol tray
QR,total
Qc
decr vol (%)
incr QR,total (%)
142.0
29
7420
-7270
4.5
0.15
Intermediate Reboiler vol tray
QR,total
Qc
decr vol (%)
incr QR,total (%)
7410
-7260
4.6
0.02
141.8
21/29
The distillation system had 28 stages, a total condenser, and a kettle-type reboiler (N ) 30). F ) 1000 kmol/h, p ) 1.0 atm, and tray spacing is 24 in. ) 0.6096 m. For the 25% n-butane feed with L/D ) 1.582, the base purity was xD,C4 ) 0.9802 and xB,C5 ) 0.9934. For the 75% n-butane feed with L/D ) 0.550, the base purity was xD,C4 ) 0.9941 and xB,C5 ) 0.9822. Notation is the same as in Table 3. a
Table 5. Results for Ternary Distillationsa FV
NF,L
0 (base)
11
80
10
NF,V
16
dia
A
vol
2.12
3.52
42.9
tray
QR,total
Qc
2955
-2307
1.93
2.93
35.7
Two-Enthalpy Feed 16/20 2957
-2308
dia
A
vol
tray
QR,total
Qc
decr vol (%)
incr QR,total (%)
1.97
3.04
37.1
16/20
2957
-2308
13.8
0.05
20
decr vol (%)
incr QR,total (%)
16.9
0.06
Vapor Bypass FBy
NF,V
NBy,V
75
9
16
Figure 3b
Intermediate Reboiler Fwithdr
NF,L
NL,with
NV,ret
dia
A
vol
tray
QR,total
Qc
decr vol (%)
incr QR,total (%)
71
9
17
16
1.95
2.99
36.5
16/20
2957
-2308
15.0
0.05
a
Pressure is 1.0 atm, N ) 21, and tray spacing is 24 in. ) 0.6096 m. Distillate mole fraction xD,C3 ) 0.9582. Notation is the same as in Table 3.
of the feed. With FV ) 600 kmol/h and the same N and same energy inputs (second entry in Table 6) the column volume decreased significantly, but the purity also dropped. The purity could be made to match the base case purity by increasing QR,total and Qc with the same N (third entry), or by increasing N with the same QR,total and Qc (fourth entry). In all cases the largest diameter is at the bottom of the column and diameters are not balanced. For the fourth entry the diameters were 5.53 m at stage 45, 3.84 m at stage 32 (vapor feed stage), and 3.64 m at stage 2. Increasing the amount vaporized to FV ) 800 further decreases the column diameter, but more stages need to be added to obtain the same purity as the base case (entry 5). The energy used to vaporize the feed with FV ) 800 is 9006.4 kW. Even with all of the feed vaporized (FV ) 1000 kmol/h, which is identical to the normal system of conditioning all of the feed) the diameter drops significantly, but so does the purity (not shown). The base case purity could again be obtained either by increasing QR or increasing the number of stages (entries 6 and 7, respectively), but the required increase in number of stages becomes quite large. If one wants to operate with the same
energy loads as the base case, the maximum volume reduction occurs close to FV ) 800 kmol/h. Table 6 also shows the results for a two-phase feed using the same amount of energy to heat the entire feed (9006.4 kW) as the two-enthalpy feed system with FV ) 800. To meet the product specification, the number of stages had to be increased. The two-phase feed system has less volume reduction than either the two-enthalpy feed with FV ) 800 or the fully vaporized feed. Vapor bypass results using the configuration in Figure 3b are also shown in Table 6. First, at a bypass rate of 500 kmol/h there is an 8% decrease in diameter, but the product purities are less than those of the base case. By adding one stage and adjusting the tray location for return of the bypass vapor, the base purities are met. In general, with a lower location in the column for the return of the bypass (larger NBy,V), the purities are slightly better but the diameter is slightly larger. The bypass runs were repeated with FBy values of 750 and 1000. Note in Table 6 that with FBy ) 1000 the largest diameter was calculated at the bypass return, not the bottom of the column. The optimum
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Ind. Eng. Chem. Res., Vol. 46, No. 26, 2007
Table 6. Results for Distillation of 95% Water, 5% Acetic Acid, Liquid Feeda Base FV
NF,L
NF,V
0
25
600
22
28
600 600 800 1000 1000
23 25 26
28 32 34 27 38
N
dia
A
vol
tray
QR,total
Qc
42
6.14
29.6
554.1
41
35 560
-35 550
decr vol (%)
incr QR,total (%)
Two-Enthalpy Feed 5.48 23.6 442.7 41 35 560 -35 550 20.1 Distillate (xD,W ) 0.9974) and bottoms (xB,W ) 0.0490) do not match base purities. 42 5.60 24.7 463.0 41 36 430 -36 400 16.4 46 5.53 24.0 493.8 45 35 560 -35 550 10.9 49 5.31 22.2 487.2 48 35 560 -35 550 12.1 42 5.30 22.0 412.4 41 37 730 -37 720 25.6 55 5.08 20.24 498.7 54 35 560 -35 550 10.0 42
2.4 6.1
Two-Phase Feed (Qheat_feed ) 9006.4 kW) NF
N
dia
A
vol
tray
QR,total
Qheat_feed
Qc
decr vol (%)
34
51
5.30
22.1
505.2
50
35 560
9006.4
-35 550
8.8
A
incr QR,total (%)
Vapor Bypass (Figure 3b) FBy
NF,L
NBy,V
N
dia
500 750 1000
24 25 25
38 38 40
43 43 43
25
31
42
47
5.89 27.3 524.2 42 35 560 -35 550 5.76 26.1 501.2 42 35 560 -35 550 5.94 27.7 531.9 40 35 560 -35 550 Two-enthalpy feed (FV ) 600) plus vapor bypass (FBy ) 750): 5.10 20.5 431.1 46 35 560 -35 500
Fwithdr
NF,L
NL,with
200 1800
25 25
33 33
34 42 5.94 27.8 521.1 41 35 560 -35 550 34 42 4.17 13.6 254.9 33/41 35 560 -35 550 Distillate (xD,W ) 0.9978) and bottoms (xB,W ) 0.0402) do not match base purities.
1800
25
33
34
vol
tray
QR,total
Qc
decr vol (%)
incr QR,total (%)
5.4 9.6 4.0 22.2
Intermediate Reboiler NV,ret
N
43
dia
4.11
A
13.3
vol
255.4
tray
42/33
QR,total
35 560
Qc
-35 550
decr vol (%)
incr QR,total (%)
6.0 54.0 53.9
F ) 1000 kmol/h, D ) 50 kmol/h, tray spacing is 18 in. ) 0.4572 meters, operation is at 80% of flooding, N ) 30, and p ) 1.0 atm. Same notation as in Table 3. Base results are in Tables 1 and 2. Base purities are xD,W ) 0.9979 and xB,W ) 0.0400. Unless specified otherwise, distillate and bottoms match or slightly exceed base purities. a
Table 7. Results for Distillation of 95% Water, 5% Acetic Acid, Saturated Vapor Feeda Fwithdr
NF
0
27
50 1100
27 27
NL,with
33 33
NV,ret
34 34
N
dia
42
Vapor Feed Base Case 5.30 22.0 412.4
A
vol
42 43
5.24 3.92
Intermediate Reboiler 21.6 404.9 12.1 232.3
tray
QR,total
Qc
41
26 470
-37 720
41 42/27
26 470 26 470
-37 720 -37 720
decr vol (%)
1.8 43.7
a F ) 1000 kmol/h, D ) 50 kmol/h, tray spacing is 18 in. ) 0.4572 m, operation is at 80% of flooding, p ) 1.0 atm, x D,W ) 0.9979, and xB,W ) 0.0400. Same notation as in Table 3. All simulations have the same value of QR,total. The decrease in column volume is compared to the vapor feed base case. All results match base case distillate and bottoms purities.
operation of this system is close to a bypass rate of 750 kmol/ h. We also tried combining the two-enthalpy feed system with vapor bypass (Table 6). The addition of two-enthalpy feed plus vapor bypass is an improvement compared to each individually. The best results for this separation were obtained with an intermediate reboiler (Table 6). With low liquid withdrawal rates (Fwithdr ) 200 in Table 6) the intermediate reboiler meets the purity specifications with no further changes. However, the decrease in volume is modest and the column diameters are not balanced since the diameter at stage 41 (5.94 m) remains much greater than the diameter at the vapor return stage (4.21 m). With much larger withdrawal rates (Fwithdr ) 1800), large reductions in diameter are possible and the column diameters essentially balance (4.13 m at stage 41 and 4.17 m at stage 33). By increasing the number of stages by one, we can match the base purities and balance the column diameters (4.11 m at stage 42 and 4.10 m at stage 33). For this water-acetic acid separation the use of an intermediate reboiler has the advantages that liquid can be withdrawn away from the pinch point that occurs at high water concentrations and the withdrawal rate can be greater than the feed rate. With Fwithdr ) 1800 the amount of energy needed
to vaporize the liquid withdrawal, 19 690 kW, is greater than the amount required in the reboiler, 15 870 kW, but QR,total is unchanged. Distillation of a 95 mol % Water, 5 mol % Acetic Acid Vapor Feed. If the feed is input as a vapor, we would normally expect the diameter to be largest at the top of the column.1 However, with this system the parameter effect on the diameter is very large and the diameter is still largest at the bottom of the column. The new base case with a vapor feed can be obtained from the two-enthalpy feed results with total vaporization of the feed in Table 6, but with QR,total reduced by the energy required to vaporize the liquid feed (Qvaporize ) 11 086 kW). This result is shown as the base case in Table 7. This system is unusual in that not only is less energy needed with a vapor feed, but also the column diameter is significantly less than with a liquid feed because much of the energy required is input with the vapor feed away from the stage that requires the largest diameter. Base case column diameters are tray 2, 3.76 m; tray 27, 3.91 m; tray 28, 3.38 m; and tray 41, 5.30 m. The size of a column with vapor feed can be reduced with vapor bypass or with an intermediate reboiler, but since the intermediate reboiler
Ind. Eng. Chem. Res., Vol. 46, No. 26, 2007 9229 Table 8. Results for Distillation of a Liquid Feed of 50% Water, 50% Acetic Acida FV
NF,L
0 (base)
30
NF,V
N
dia
A
vol
tray
QR,total
Qc
42
5.15
20.8
389.9
41
22 704
-22 679
decr vol (%)
Intermediate Reboiler Fwithdr
NF,L
NL,with
NV,ret
N
dia
A
vol
tray
QR,total
Qc
decr vol (%)
625 900 1000
31 31 31
34 34 34
35 35 35
43 43 43
4.57 4.56 4.58
16.4 16.3 16.5
314.9 313.0 316.8
35/42 34 34
22 704 22 704 22 704
-22 679 -22 679 -22 679
19.2 19.7 18.7
Two-Enthalpy Feed FV
NF,L
NF,V
N
dia
A
vol
tray
QR,total
Qc
decr vol (%)
430 500 980 1000
30 31 31
32 33 33 34
43 44 46 47
4.69 4.62 4.03 4.02
17.3 16.7 12.8 12.7
324.3 328.3 263.3 267.1
42 43 45 34
22 704 22 705 22 704 22 705
-22 679 -22 678 -22 679 -22 678
16.8 15.8 32.5 31.5
a Feed rate is 1000 kmol/h and operation is at 1.0 atm. L/D ) 3.0, D ) 500 kmol/h, tray spacing is 18 in. ) 0.4572 m, p ) 1.0 atm, x D,W ) 0.994 33, and xB,W ) 0.005 67. All simulations have essentially the same value of QR,total. Decrease in volume is compared to the base case in this table. Notation is the same as in Table 3. All systems meet purity requirements.
was much more effective only this result is shown (Table 7). The intermediate reboiler system with Fwithdr ) 50 kmol/h matches the base case product purities, but the reduction in column volume is only 1.8%. If the withdrawal rate is increased to Fwithdr ) 75 kmol/h (not shown in the table), the bottoms purity requirement is no longer met. If we increase N to 43 and keep increasing the withdrawal rate until the diameters at the feed stage and the bottom tray balance (Fwithdr ) 1100 kmol/ h), we obtain the same purities as the base case, but with a column volume that is 43.7% smaller (Table 7). Distillation of a 50 mol % Water, 50 mol % Acetic Acid Liquid Feed. The base case column had 40 equilibrium trays, a kettle reboiler, and a total condenser (N ) 42). External reflux ratio was 3.0. The optimum feed location was NF ) 30. Product purities are xD,W ) 0.994 33 and xB,W ) 0.005 67. Results are shown in Table 8. The calculated diameters are 2.94 (stage 2), 4.06 (feed stage ) 30), and 5.15 m (stage 41). This more concentrated feed is still a difficult separation, but the diameters and energy loads are reduced despite a considerably higher purity bottoms product. The intermediate reboiler results (Table 8) show a decrease in calculated column volume as Fwithdr increases up to Fwithdr ) 625 kmol/h, where the calculated column diameters are identical at the bottom of the column and stage 35 (vapor return stage). Above this withdrawal rate (e.g., to Fwithdr ) 900 kmol/h) stage 34 becomes the stage with the largest calculated diameter and there is very little advantage to increasing the withdrawal rate. Increasing the withdrawal rate to Fwithdr ) 1000 kmol/h causes the diameter to increase, and the purities just match the base case. Further increases of the withdrawal rate will cause the purities to not meet the product specifications. The two-enthalpy feed method was also simulated for this separation (Table 8). The run with FV ) 430 is better than the run with FV ) 500 because the increase in the amount of feed vaporized requires an extra tray to reach the desired purity. The best results were obtained with FV ) 980, although the fully vaporized feed system (FV ) 1000) is almost as good and is simpler. With this feed concentration the two-enthalpy feed system works well because the feed stage is no longer in the region of very low volatilities. Subtracting the energy required to vaporize the feed (9171.5 kW) from the FV ) 1000 results in Table 8 gives the results for a saturated vapor feed of the 50-50 mixture. Since the diameter at stage 34 (4.02 m) is almost identical to the diameter at stage 46 (4.01 m), no further reductions in column volume are possible for a saturated vapor feed.
Table 9. Relative Heat Exchanger Areas, A ) Q/(U∆T), for Distillation of a Liquid Feed of 5% Acetic Acid and 95% Water Comparing the Intermediate Reboiler System (Fwithdr ) 1800 kmol/h) with the Base Case (see Table 6)a steam T (°C)
∆TR
∆Tint_reb
Atotal/Abase
122 132 142 152
5.6 15.6 25.6 35.6
20.9 30.9 40.9 50.9
0.60 0.73 0.79 0.83
a The column reboilers are at 389.55 K, and the liquid withdrawn from stage 33 in the intermediate reboiler system is at 374.25 K. QR,total ) 35 560 kW for both systems; for the intermediate reboiler system QR ) 15 870 kW, and Qint_reb ) 19 690 kW. Heat transfer coefficient U is assumed to be the same for all systems. Atotal/Abase is the ratio of the total heat exchanger area of the intermediate reboiler system divided by the heat exchanger area of the base case.
Heat Exchangers The effects of intermediate reboilers and two-enthalpy feed on the cost of heat exchangers and the cost of heating depends on the situation in each plant. Since the intermediate reboiler and the heater for the feed operate at lower temperatures than the column reboiler, they can use a lower temperature heating medium. If this lower temperature heating medium is less expensive, then operating costs can be reduced.2,5 Since the cost per area of heat exchangers decreases as the area increases,9 splitting the single reboiler of the base case into two heat exchangers for the intermediate reboiler and twoenthalpy feed systems would be expected to increase the costs of the heat exchangers. However, if the same heating medium is used, the driving force ∆T for heat transfer is larger for the intermediate reboiler and the feed heater than for the column reboiler. This will result in a smaller total area for systems with an intermediate reboiler or feed heater when QR,total is constant. This reduction in area can be large if the temperature of the heating medium is close to the boiling temperature in the reboiler. For example, consider the comparison of the intermediate reboiler system to the base case for the acetic acid-water distillation delineated in Table 9. For the lower steam temperatures the difference between the two driving forces is quite large, and the total heat exchanger area for the intermediate reboiler system is considerably less than for the base case. Despite the cost advantage per area of the larger heat exchanger, the cost for the heat exchangers for the intermediate reboiler system will be very competitive. Since the smaller distillation column of the intermediate reboiler system is less expensive than the base case distillation column,9 the overall system costs
9230
Ind. Eng. Chem. Res., Vol. 46, No. 26, 2007
Table 10. Results for Absorption of n-Pentane from Methanea FBy
NBy,V
dia
A
vol
tray
L0
LN
V1
decr vol (%)
incr L0 (%)
0 (base) 140
4
2.01 1.90
3.17 2.85
15.5 13.9
7 7/4
466 473
666 673
800 800
10.1
1.4
Methane purity in outlet gas matches the base case purity (99.03%). Both gas and liquid are at 20 °C and 3.0 atm. Column pressure is 3.0 atm, the column is adiabatic, and the plate spacing is 24 in. ) 0.6096 m. The feed is 99.41% vapor and 0.59% liquid. FBy is the kmol/h vapor in the vapor bypass, and NBy,V is the feed location for the vapor bypass. Flow rates of entering solvent L0, exiting liquid LN, and treated vapor V1 are in kmol/h. a
will be less. With higher steam temperatures the heat exchanger areas of the two systems start to approach each other and the heat exchangers for the intermediate reboiler system will become more expensive; however, the overall costs may still be less for the intermediate reboiler system. For accurate economic estimates, the costs need to be determined on a case-by-case basis. Vapor Bypass for Concentrated Absorbers Since gas enters at the bottom of the absorber and much of the solute is transferred to the liquid, the vapor flow rate decreases as one goes up the absorber. This effect is particularly marked for concentrated feed streams. Thus, the largest calculated diameter for concentrated absorbers is at the bottom of the column. Although in the particular chemical system studied here condensation of a portion of the gas stream is reasonably convenient, in many cases this is not true. When the gas stream is difficult to partially condense, neither two-feed nor intermediate reboiler systems are convenient. Thus, we will explore only the use of vapor bypass (Figure 4) for concentrated absorbers. Absorption of n-Pentane from Methane Gas with Heavy Solvent. Table 10 shows the results for the base case absorption column with seven equilibrium stages processing 1000 kmol/h of a feed stream that is 80 mol % methane and 20 mol % pentane. The inlet solvent is pure and is assumed to have the same properties as n-nonane. Calculations were done with the AspenPlus simulator using the Peng-Robinson VLE correlation. For the base case the inlet liquid flow rate was 466 kmol/h; the outlet gas stream is 99.03 mol % methane, 0.73% n-pentane, and 0.25% solvent; and the outlet liquid stream is 1.13% methane, 29.16% n-pentane, and 69.71% solvent. The calculated diameters and vapor flow rates for the base case are stage 1 ) 1.67 m, V1 ) 800 kmol/h; and stage 7 ) 2.01 m, V7 ) 952 kmol/h. For vapor bypass part of the concentrated feed gas bypasses the bottom of the column where the calculated diameter is largest. The best results with the same outlet methane mole fraction (99.03%) as the base case are shown in Table 10. Putting the bypass lower in the column improved the purity and required less increase in solvent flow rate to match the purity of the base case, but there was less reduction in vapor flow rate and hence a larger final diameter. Putting the bypass higher in the column reduced the purity more but reduced the calculated diameter more. Repeated simulations showed that stage 4 was the best compromise to input the bypass vapor feed. The calculated diameters and vapor flow rates for the vapor bypass run with FBy ) 140 are stage 1 ) 1.67 m, V1 ) 800 kmol/h; stage 4 ) 1.91 m, V4 ) 887 kmol/h; and stage 7 ) 1.91 m, V7 ) 816 kmol/h. Discussion Since the diameter balancing method is a processing technique for distillation and absorption, it is generally applicable to any type of equipment used for these unit operations. Thus, the methods developed in this and the previous paper1 can be
applied to different types of trays and different random and structured packings. Retrofitting, which is illustrated elsewhere,1 is probably easier with tray systems than with packed systems. Once the final design has been determined, a tray rating or packing rating analysis should be done to ensure proper hydraulic behavior.1 The ternary example showed that this method of diameter balancing can be applied to multicomponent systems. This also applies to azeotropic and extractive distillation if there is a major change in column diameter in the column. The processing method that maximized volume reduction with constant energy use depended upon the separation being performed and the feed concentration. Other factors such as ease of operation or low-cost retrofitting may also be important. For example, if a distillation system with excess reboiler and condenser capacities is limited by the vapor flow rate at the bottom of the column, modest debottlenecking can easily be achieved with vapor bypass (Figure 3a). Vapor bypass would probably not be the method of choice for a new column. When there are very large changes in calculated column diameter, distillation systems are often constructed with two columns of different diameters. Use of the methods in this paper will reduce the diameter of the larger column (in this paper the stripping section), and may allow use of a single, constant diameter column. Obviously, the economically optimum system for any given separation problem depends upon the detailed economic costs and needs to be calculated for each specific case. Keller10 found for distillation that the ratio of capital cost to utility costs ranged from about 0.55 to 3.5; thus, the capital cost is always significant. The large reductions in column volume for the acetic acid systems will clearly result in significant capital savings. The significance of the savings for the other systems needs to be determined from more detailed economic calculations. Nomenclature A ) column cross-sectional area, m2 D ) distillate flow rate, kmol/h dia ) calculated column diameter, m F ) feed rate, kmol/h FBy ) rate of vapor bypass, kmol/h FL ) feed rate of liquid in two-enthalpy feed column, kmol/h FV ) feed rate of vapor in two-enthalpy feed column, kmol/h Fwithdr ) withdrawal rate of liquid stream sent to intermediate reboiler, kmol/h Lj ) liquid flow rate leaving stage j, kmol/h L/D ) external reflux ratio N ) number of trays + condenser + reboiler NBy,V ) feed location for vapor bypass Nfeed ) optimum feed stage with condenser labeled as stage 1 NF,L, NF,V ) optimum feed locations for the liquid and vapor portions of feed NV,ret ) optimum location to return vapor from intermediate reboiler NV,with ) optimum withdrawal plate for the vapor Fwithdr p ) column pressure, atm Qc ) condenser duty, kW
Ind. Eng. Chem. Res., Vol. 46, No. 26, 2007 9231
Qheat_feed ) energy used to heat the entire feed to produce a two-phase feed, kW QR ) reboiler duty, kW QR,total ) total heating duties including intermediate reboiler and vaporizing portion of feed, kW Qvaporize ) energy required to vaporize the vapor bypass, the portion of the feed, or in the intermediate reboiler, kW Vj ) vapor flow rate leaving stage j, kmol/h vol ) column volume, m3 xB,i ) mole fraction component i in bottoms product xD,i ) mole fraction component i in distillate Subscripts AA ) acetic acid C3 ) propane C4 ) n-butane C5 ) n-pentane feed_st ) feed stage M ) methanol W ) water
(3) Liebert, T. Distillation Feed PreheatsIs it Energy Efficient? Hydrocarbon Process. 1993, 72 (10), 37. (4) Ognisty, T. P. Distillation/Energy Management. In Encyclopedia of Separation Science; Wilson, I. D., Ed. in Chief, Cooke, M., Poole, C. F., Eds.; Academic Press: New York, 1999; Vol. 3, pp 1005-1012. (5) Wankat, P. C.; Kessler, D. P. Two-Feed Distillation: Same Composition Feeds with Different Enthalpies. Ind. Eng. Chem. Res. 1993, 32 (12), 3061-3067. (6) Fair, J. R. Gas Absorption and Gas-Liquid System Design. In Perry’s Chemical Engineers’ Handbook, 7th ed.; Perry, R. H., Green, D. W., Eds.; McGraw-Hill: New York, 1997; pp 13-23 to 13-38. (7) Huang, H.-P.; Lee, H.-Y.; Gau, T. K.; Chien, I. L. Design and Control of Acetic Acid Dehydration Column with p-Xylene or m-Xylene Feed Impurity. I. Importance of Feed Tray Location on the Process Design. Ind. Eng. Chem. Res. 2007, 46, 505-517. (8) Seader, J. D. Distillation. In Perry’s Chemical Engineers’ Handbook, 7th ed.; Perry, R. H., Green, D. W., Eds.; McGraw-Hill: New York, 1997; Section 13. (9) Turton, R.; Baillie, R. C.; Whiting, W. B.; Shaeiwitz, J. A. Analysis, Synthesis and Design of Chemical Processes, 2nd ed.; Prentice Hall PTR: Upper Saddle River, NJ, 2003; Appendix A. (10) Keller, G. E., II. Separations: New Directions for an Old Field; AIChE Monograph Series; American Institute of Chemical Engineers: New York, 1987; No. 17.
Literature Cited (1) Wankat, P. C. Balancing Diameters of Distillation Column with Vapor Feeds. Ind. Eng. Chem. Res. 2007, 46, DOI: 10.1021/ie0705554. (2) King, C. J. Separation Processes, 2nd ed.; McGraw-Hill: New York, 1981; pp 704-707.
ReceiVed for reView July 20, 2007 ReVised manuscript receiVed September 19, 2007 Accepted October 2, 2007 IE0709887
