Distillation Tower Design
• As computer technology advances, the
fundamental aspects of plant design are
becoming a lost art. … N.P. Lieberman,
Refinery Manager, GHR Energy Inc., La
• The following steps are taken to design
and optimize a distillation tower:
R.A. Hawrelak, 22 Jan 02, CBE 497
(a) Select a Process Sequence
• Consider a five component feed as shown below.
Arrange in order of descending vapor pressure.
•
•
•
•
•
•
C2
C3
C4
C5
C6
Total
3
20
37
35
5
100
lb moles/hr
Process Sequence Cont’d
•
•
•
•
•
Make a split between C3 and C4
Show this as C2, C3 / C4, C5, C6
This called a depropanizer.
C3 is identified as the light key.
C4 is identified as the heavy key.
Establish Key Component
Specs
• C3, light key composition in bottoms
shall be 1.0 mole %. (2.0% sales spec)
• C4, heavy key composition in the
overheads shall be 1.5 mole %.
(3.0% sales spec).
Set Up Mass Balance for
Tower
Feed
C2
C3 lk
C4 hk
C5
C6
Feed
3
20
37
35
5
100
Ohds
3
20 - Y
X
23 + X - Y
Btms
Y
37 - X
35
5
77 + Y - X
Mass Balance Equations
1.00%
1.50%
Light Key In Btms =
Heavy Key In Ohds =
Y
77 + Y - X
0.01
Eqn 1
X
23 + X - Y
0.015
Eqn 2
Mass Balance Solution
C2
C3 lk
C4 hk
C5
C6
Feed
mf in Feed
Ohds
mf in Ohds
3
20
37
35
5
100
0.0300
0.2000
0.3700
0.3500
0.0500
1.0000
3.00
19.23
0.34
0.1330
0.8520
0.0150
22.56
1.0000
Btms
mf In Btms
0.77
36.66
35.00
5.00
77.44
0.0100
0.4734
0.4520
0.0646
1.0000
Obtain Antoine Constants
• Need Antoine constants for Vapor
Pressure
• Vap Press, VP = 10^(A + B / (t°C + C))
psia
Component
C2
C3 lk
C4hk
C5
C6
A
5.0120015
4.3742477
3.8201853
4.0537542
4.0165587
B
-823.03103
-587.76681
-367.50819
-539.73661
-545.39181
C
328.18
248.90
153.30
169.60
141.15
Feed Conditions
• Temperature of feed = 225 deg F =
107.22 deg C.
• Pressure of feed = 264.7 psia
Determine Bubble Point of
Feed
Assume T, deg F =
Component
C2
C3 lk
C4 hk
C5
C6
212.5714318
Moles/h, M
3
20
37
35
5
100.00
T, Deg C = 100.317462
At Press =
264.7
psia
VP, psia
1233.86
491.08
235.03
113.27
57.27
K = VP/P
4.66133
1.85523
0.88790
0.42792
0.21635
M*K
13.98400
37.10452
32.85248
14.97726
1.08173
99.999999
Determine Dew Point of Feed
Asssume T =
At Press =
Component
C2
C3 lk
C4 hk
C5
C6
269.7699145
T, Deg C = 132.094397
264.7
psia
Moles/hr, M
3
20
37
35
5
100.00
VP, psia
1674.442
678.469
340.773
183.966
104.851
K = VP/P
6.3258
2.5632
1.2874
0.6950
0.3961
M/K
0.4742
7.8029
28.7403
50.3600
12.6227
100.000001
Assess Feed Condition
•
•
•
•
Feed Bubble Point = 100.32 deg C
Feed Temp = 107.22 deg C
Feed Dew Point = 132.09 deg C
Feed temp between Bubble Pt. and
Dew Pt.
• Feed must be in a two-phase V / L
state.
• Special care will have to be taken for
feed distributor design on feed tray.
Determine V / L for Feed
Assume V/L =
Component
C2
C3 lk
C4 hk
C5
C6
0.17114
1
Moles/hr, M
2
K = VP/264.7
3
K(V/L)
4
1+K(V/L)
5
L=1/4
6
V=1-5
3
20
37
35
5
100.00
5.0000
2.0000
0.9700
0.4800
0.2500
0.856
0.342
0.166
0.082
0.043
1.856
1.342
1.166
1.082
1.043
1.617
14.900
31.732
32.343
4.795
85.387
1.383
5.100
5.268
2.657
0.205
14.613
Flash Fraction Vapor in Feed =
1-q = V / M = 0.14613121
V/L calc'd =
V/L Ass'd =
Diff =
0.17114
0.17114
0.00000
Solve For Ø, The Underwood
Parameter
• Example In article by J.M. Ledanois,
Hydrocarbon Processing, April, 1981, P231
• Trial and error solution with as many
solutions as there are components.
• Solution is a Newton convergence
method.
• Not all cases converge.
Solve For Ø, The Underwood
Parameter, Cont’d
Next Ø
1.5976753
from Neqton Convergence
Assumed Ø =
1.5976753
Solution 2
1-q=
0.1461312
enter no. manually value from cell P55 - Fraction Vapor In Feed
Temp =
Press, PT =
Moles/Hr
Feed
C2
C3 lk
C4 hk
C5
C6
225
264.7
Fi
3
20
37
35
5
xF1
0.0300
0.2000
0.3700
0.3500
0.0500
100
1.0000
°F
psia
Ki = VPi / PT
5.0000
2.0000
0.9700
0.4800
0.2500
107.22
rel volatility
Alpha i
5.1546
2.0619
1.0000
0.4948
0.2577
°C
Alpha*xFi
0.155
0.412
0.370
0.173
0.013
∑=
Eqn 13-43
Alpha*xFi
Aplha - Ø
0.043
0.888
-0.619
-0.157
-0.010
0.1461312
Ist Estimate
Of Ø
Avg Alpha
Adj Compts
3.6082
1.5309
0.7474
0.3763
= 1-q
Solve For Ø, The Underwood
Parameter, Cont’d
1-q=
Feed
C2
C3 lk
C4 hk
C5
C6
0.1461312
Alpha*xFi
0.155
0.412
0.370
0.173
0.013
∑=
Eqn 13-43
Alpha*xFi
Aplha - Ø
0.043
0.888
-0.619
-0.157
-0.010
0.1461312
Ist Estimate
Of Ø
Avg Alpha
Adj Compts Solution No.
3.6082
1
1.5309
2
0.7474
3
0.3763
4
= 1-q
Final Est
Of Ø
Avg Alpha
Ø Solutions
4.80084859
1.59767531
0.64185818
0.26706697
Calc Minimum Reflux Ratio by
Underwood
• See Perry VI, Chem Eng HB, Page 1336
• Solution For Minimum Reflux Ratio By
Solving For ∑ [Alpha*xDi / (Alpha - Ø)]
= L/D min. + 1
• Ø, The Underwood Parameter, was
determined above.
Calc Minimum Reflux Ratio by
Underwood, Cont’d
Underwood parameter =
Temp =
225
Pressure =
264.7
Ohds
C2
C3 lk
C4hk
Moles/hr
3.00
19.30
0.30
22.6
1.5977
°F
psia
xD1
0.1327
0.8540
0.0133
1.0000
107.22
Ki = VPi / PT
5.0000
2.0000
0.9700
°C
Alphai
Ki/K hk
5.1546
2.0619
1.0000
Alpha*xDi
0.684
1.761
0.013
L/Dmin. + 1=
L/D min. =
Eqn 13-42
Alpha*xDi
Aplha - Ø
0.192
3.793
-0.022
3.9635
2.9635
Determine Minimum No. Trays
by Fenske - Underwood
• Assume top and bottom pressure equal
feed pressure of 264.7 psia for now.
• Assume overhead distillate is removed
as a vapor from the condenser.
Determine Minimum No. Trays
by Fenske – Underwood, Cont’d
Dew Point of Overhead Vapor Stream and Alpha of Keys
Temp =
Pressure =
Ohds, D
C2
C3 lk
C4hk
118.3
264.7
Moles/hr, M
3.00
19.30
0.30
°F
47.94
psia
xD1
0.1327
0.8540
0.0133
1.0000
22.6
Alpha Top = KC3 / KC4 =
°C
Ki = VPi / PT
2.5180
0.9364
0.3726
M/K
1.19
20.61
0.81
22.61
2.51
For Distillate at 118 °F
Determine Minimum No. Trays
by Fenske – Underwood, Cont’d
Bubble Point Of Bottoms Stream and Alpha of the Keys
Temp =
Press =
C3 lk
C4hk
C5
C6
MT =
274.0
264.7
Moles/hr, M
0.7
36.7
35
5
°F
psia
134.44
xF1
0.009044
0.474160
0.452196
0.064599
1.000000
77.4
Alpha Btm = KC3 / KC4 =
Ki = VPi / PT
2.6196
1.3189
0.7175
0.4119
°C
M(K)
1.834
48.405
25.112
2.060
77.411
1.99
For Btms at 274°F
Determine Minimum No. Trays
by Fenske – Underwood, Cont’d
• Determine geometric Average Alpha
between top and bottom of the tower.
• Geometric Avg = (Alpha Top*Alpha
Btm)^0.5
• Avg Alpha = ((2.51)(1.99))^0.5 = 2.23
Determine Minimum No. Trays
by Fenske – Underwood, Cont’d
• Min. Trays = LN((C3 lkD / C4 hkD)*
(C4 hkB / C3 lkB)) / LN(Alpha Avg)
• Minimum No. Trays, Sm = 10.11
Determine Trays versus Reflux
Ratio by Gilliland Method
• Use Chang equation to represent Gilliland.
• Huan Yang Chang, HC Proc, Oct 1981, P146
• A partial condenser and a reboiler represent
two theoretical trays.
• No. trays = S – 2.
• Assume the economic reflux ratio is 1.2 times
the minimum reflux ratio,
• Plot the results.
Determine Trays versus Reflux
Ratio by Gilliland Method, con’d
L/D Min. =
2.9635
Sm =
10.11
L/D =
3.56
Chang Factor = (S - Sm) / (S + 1) = 1 - EXP(1.49+0.315*C-1.805/C^0.1)
where C = (L/D - L/Dmin) / (L/D + 1)
Chang Factor
A
B
C
(S - Sm)
L/D
L/D - L/Dmin
B/(A +1)
(S + 1)
S
3.1
0.14
0.0333
0.65
30.30
3.2
0.24
0.0563
0.59
26.29
3.3
0.34
0.0783
0.56
24.07
3.56
0.59
0.1301
0.49
20.97
4
1.04
0.2073
0.43
18.39
4.5
1.54
0.2794
0.38
16.81
6
3.04
0.4338
0.29
14.53
By Chang
N=S-2
28.30
24.29
22.07
18.97
16.39
14.81
12.53
Plot of Trays Versus Reflux Ratio
Reflux Versus Number of Trays
40.00
35.00
30.00
Number of Trays
25.00
20.00
Example
15.00
10.00
5.00
0.00
2.00
2.50
3.00
3.50
4.00
4.50
Reflux Ratio, L/D
5.00
5.50
6.00
6.50
Determine Feed Tray Location
Feed Tray Location By Kirbride Equation
Oil And Gas Journal, Oct. 20, 1980, P-138, by Henry Y. Mak
Feed Tray Location = EXP[0.206*LN(B/D*fhk/flk*((blk/B)/(dhk/D))^2)]
B = 77.44
D = 22.56
Kirbride Feed Tray Location =
No theoretical trays =
Trays Above =
Trays Below =
Ratio A/B =
10.55
8.45
1.25
fhk = 37
flk = 20
1.25
dhk = 0.34
blk = 0.77
= # trays above fd / # trays below fd
19
11
8
1.38
final selection
10
9
1.11
Determine Reflux Flow & Comp’n
Temp =
Pressure =
L/D =
Ohds, D
C2
C3 lk
C4hk
D=
Moles/hr, M
3.00
19.23
0.34
22.5641
118.65
264.7
2.963
Mole Frac
y, D
0.1330
0.8520
0.0150
1.0000
°F
psia
48.14
L=
Ki = VPi / PT
2.5246
0.9392
0.3741
°C
66.868
Dew Pt.
M/K
1.19
20.47
0.90
Comp of
Liq Reflux
x=y/K
0.05266
0.90724
0.04009
Reflux
Stream,
mole/hr, L
3.522
60.666
2.681
22.5641
1.00000
66.868
Calculate Overhead Vapor
Flow from Top Tray = 19
Minimum Reflux Ratio =
D=
L = (2.963)(D) =
V=L+D=
2.963
22.56
66.868
89.432
moles/hr
moles/hr.
moles/hr
Calc Vapor Composition from
Top Tray 19
Temp =
Pressure =
Ohds, D
C2
C3 lk
C4hk
Moles/hr, M
3.00
19.23
0.34
22.5641
118.65
264.7
Mole Frac
y, D
0.1330
0.8520
0.0150
1.0000
°F
psia
Ki = VPi / PT
2.5246
0.9392
0.3741
48.14
°C
Dew Pt.
M/K
1.19
20.47
0.90
Comp of
Liq Reflux
x=y/K
0.05266
0.90724
0.04009
Reflux
Stream,
mole/hr, L
3.522
60.666
2.681
Vapour
Stream
moles/hr, V
6.52
79.89
3.02
Vapor
Comp'n
mole frac
0.0729216
0.893316
0.0337624
22.5641
1.00000
66.868
89.432
1.00000
Show Molar Balance Around
Top Tray 19
Top Tray n, Assume equal molar flow
Molar balance = V18 = V19 + L19 - L
V19 = 89.4324
Tray 19
V19 = L + D
Reflux L
66.868
L19
66.868 Liq fr Tray 19
V18 = 89.4324
Calc Dew Pt of Vapor Fr T19
and Liquid Comp’n Fr T19
Temp =
Pressure =
Ohds, V
C2
C3 lk
C4hk
Moles/hr, V
6.52
79.89
3.02
126.5378
264.7
Mole Frac
y, V
0.0729
0.8933
0.0338
1.0000
°F
psia
Ki = VPi / PT
2.6753
1.0035
0.4091
89.4324
Moles/hr, V
Diff =
52.52
Dew Pt.
M/K
2.44
79.61
7.38
89.4325
89.4324
0.0001
°C
Comp of
Liq fr Tray n
x=y/K
0.02726
0.89022
0.08252
1.00000
Vapor Comp’n From Tray 18
C2
C3 lk
C4hk
Vapor, V19
moles/hr
6.52
79.89
3.02
Reflux, L
moles/hr
3.52
60.67
2.68
L19 Comp
mole fr
0.0273
0.8902
0.0825
L19 Flow
moles/hr
1.82
59.53
5.52
Vap, V18
moles/hr
4.82
78.75
5.86
Vap, V18
mole frac
0.0539
0.8806
0.0655
89.43
66.87
1.0000
66.87
89.43
1.0000
Calc Dew Point of Vapor V18
Temp =
Pressure =
133.0915
264.7
Vap V18
C2
C3 lk
C4hk
V18
Moles/hr, M
4.82
78.75
5.86
°F
psia
Mole Frac
y, V18
0.0539
0.8806
0.0655
1.0000
56.16
°C
Ki = VPi / PT
2.8044
1.0587
0.4394
Dew Pt.
M/K
1.72
74.39
13.33
Moles/hr V18
Diff
89.4326
89.4326
0.0001
89.4326
Design Data For Top of Tower
C2
C3 lk
C4 hk
MW
30.07
44.10
58.12
MW =
Vap, V18
lb/hr
145
3,473
340
3,958
44.26
L19 Flow Liq Density
lb/hr
lb/cf
55
22.2
2,625
31.6
321
35.1
3,001
31.80
44.87
Dew Point of Vapor from Tray 18 = Vapour to Top Tray 19
Temp =
Pressure =
133.0915
264.7
Deg F
psia
Vapor Density = (MW)(Psia) / (10.73) / (Deg Rankine) =
1.841
lb/cf
Input Shortcut Tower Dia.
(FWG)
Input Data
Vapor To Tray, V, lb/hr =
Vapor Density, Dv, lb/cf =
Liquid From Tray, L, lb/hr =
Liquid Density, DL, lb/cf =
System Factor, SF =
Tray Spacing, TS, Inches =
Spray Ht. / Tray Spacing, =
Assume Minimum Valves / AA =
Downcomer Flood, % DCF =
3,958
1.841
3,001
31.80
1.00
18
50.00%
8
50.00%
CFS Vapor =
0.60
CFS Vapor = V / DV / 3600
USGPM =
11.76
USGPM = L / DL / 60 * 7.481
Non Foaming System
FWG 4900 / 5, Table 1b
Assume 18 inches or 24 inches as a First Try. 24 inch TS is preferred.
Assume 70% as a default value.
No. Valves / sf AA A Higher value, eg 13, may lead to an unsuitable, smaller to
Assume 60% as a default value.
Preliminary Sizing
Downcomer Area (One Side), sf =
Active Area, AA, sf =
Tower Area sf =
Tower Dia, ft. =
0.14
1.39
1.66
1.45
DCA = (L*7.481 / DL /60)/(7.5*TS^0.5*(DL-Dv)^0.5*SF) / %DC Flood
AA = V/3600/Dv/(((((TS*SH/TS)-4.5213)/4.3662)*DL/Dv)^0.5)*78.5/(No Valves /
AT = AA + 2*DCA
D = (AT*3/p)^0.5
Select Tower Diam =
2.00
ft.
(next 6" increment)
Shortcut Method by Dr
Prakash
Kv =
cm/s =
vel =
A=
D metric =
D English =
0.04070
0.01691
0.16419
0.103
0.362
1.188
by Kv = -0.17*(TS)^2+0.27*(TS)-0.047 where TS is in meters
cm/s
by cfs/3.2808^3
m/s
by Vmax =Kv(DL-DV/DV)^0.5, the Souder Eqn
sm
cm/s / m/s
m
D = (sm/.785)^0.5
ft
D = m*3.2808
Check Tower Mole Balance
89.43
22.56
66.87
reflux
Tray 19
89.43
66.8684
14.61
100.00
Flashing feed
Feed Tray Molal Balance
Feed In =
241.69
Feed Out = 241.69
74.82 Feed Tray
85.39
74.82
152.26
Tray No. 1
Vapour to Tray No. 1 = 74.82
In equilibrium with bottoms
152.26 Liquid From Tray No. 1
Note: Tower simulation usually assume this config
for the reboiler. In actual practice it is not quite like this
See tower sketch for reason
Reboiler
77.44
Composition From Btms Bubble Pt.
Calc Bubble Point of Bottoms
Temp =
Pressure =
Btms, B
C3 lk
C4 hk
C5
C6
Moles/hr, M
0.77
36.66
35.00
5.00
77.4359
273.757833 °F
264.7
psia
Mole Frac
x, B
0.0100
0.4734
0.4520
0.0646
1.0000
Ki = VPi / PT
2.6163
1.3171
0.7162
0.4110
134.31
Bub Pt.
M*K
2.03
48.29
25.07
2.06
77.4358
77.4359
-0.0001
°C
y = K(x)
Vap to Tr 1
mole frac
0.02616
0.62359
0.32371
0.02654
1.00000
Vapor Rate To Tray 1
Btms, B
C3 lk
C4 hk
C5
C6
y = K(x)
Vap to Tr 1 Vap to Tr 1
mole frac Moles/hr, V1
0.02616
1.96
0.62359
46.66
0.32371
24.22
0.02654
1.99
1.00000
74.82
Molecular
Weight
44.10
58.12
72.15
86.18
Vap to Tr 1
lb/hr
86
2,712
1,747
171
63.04
4,717
Final Vapor & Liquid Data to Tr 1
Btms, B
C3 lk
C4 hk
C5
C6
Btms, B
moles/hr
0.77
36.66
35.00
5.00
77.44
Temp =
Press =
Vap to Tr 1 Liq Fr Tray 1
Moles/hr, V1
Moles/hr
1.96
2.73
46.66
83.32
24.22
59.22
1.99
6.99
74.82
273.76
264.70
152.26
Molecular
Weight
44.10
58.12
72.15
86.18
Liq Fr Tray 1 Liq Dens
lb/hr
lb/cf
120
31.6
4,842
35.1
4,273
37.2
602
40.7
9,838
36.31
Deg F
psia
Vapor Density = (MW)(Psia) / (10.73) / (Deg Rankine) =
2.119
lb/cf
Tower Diameter For Bottom Tray 1
Vapor To Tray, V, lb/hr =
Vapor Density, Dv, lb/cf =
Liquid From Tray, L, lb/hr =
Liquid Density, DL, lb/cf =
System Factor, SF =
Tray Spacing, TS, Inches =
Spray Ht. / Tray Spacing, =
Assume Minimum Valves / AA =
Downcomer Flood, % DCF =
Preliminary Sizing
Downcomer Area (One Side), sf =
Active Area, AA, sf =
Tower Area sf =
Tower Dia, ft. =
Select Tower Diameter
4,717
2.119
9,838
36.31
1.00
18
50.00%
8
50.00%
0.36
1.45
2.17
1.66
2
CFS Vapor =
0.62
USGPM =
33.78
FWG 4900 / 5, Table 1b
Assume 18 inches or 24 inches as a Firs
Assume 70% as a default value.
No. Valves / sf AA A Higher value, eg 1
Assume 60% as a default value.
DCA = (L*7.481 / DL /60)/(7.5*TS^0
AA = V/3600/Dv/(((((TS*SH/TS)-4.5
AT = AA + 2*DCA
D = (AT*3/p)^0.5
ft.
(next 6" incremen
Shortcut Method by Dr
Prakash
For Bottom of Tower
Kv =
cm/s =
vel =
A=
D metric =
D English =
0.04070
0.01751
0.16348
0.107
0.369
1.212
by Kv = -0.17*(TS)^2+0.27*(TS)-0.047 where TS is in meters
cm/s
by cfs/3.2808^3
m/s
by Vmax =Kv(DL-DV/DV)^0.5, the Souder Eqn
sm
cm/s / m/s
m
D = (sm/.785)^0.5
ft
D = m*3.2808
Tray Efficiency
O'Connell and Drickamer / Bradford Tray Efficiencies
Basis, Perry VI, p 18-14, & Ludwig, Applied Process Design For Chemical Plant Design
Assume avergae column conditions at Feed temp =
Comp
C2
C3 lk
C4 hk
C5
C6
Xi, Feed
0.03000
0.20000
0.37000
0.35000
0.05000
1.00000
Liq Visc
cP
0.02
0.06
0.1
0.15
0.16
(Xi)(cP)
0.0006
0.0120
0.0370
0.0525
0.0080
0.1101
225
Deg F
Vap Press Vap Pre Ratio
psia
Alpha lk/hk (Alpha lk/hk)(Xi)
529.40
256.76
= X For Drickamer
O'Connell Y =
Drickamer Y =
0.1101
2.0619
0.2270
0.2270
Tray Efficiency cont’d
Box A
O'Connell Tray Efficiency =
70.39%
Perry VI, Fig 18-23a
Perry VI, Eqn 18-14 Tray Eff'y = IF Y > 4, (46.514*(Y)^-0.2052)/100
IF Y > 1 Tray Eff'y = (48*(Y)^-0.228)/100
IF Y > 0.45 Tray Eff'y = (48*(Y)^-0.2797)/100
Else, Tray Eff'y = (49.83*(Y)^-0.233)/100
Ans -->
63.06%
67.31%
72.67%
70.39%
Drickamer Tray Eff'y =
76.43%
Drickamer Tray Eff'y = -27.3*LN(Drickamer Y / 1.81) / 100
Ludwig, Applied Process Design For Chemical Plant Design
And Petroleum Plant, Vol Ii, Gulf Publishing, Circa 1960.
Recommend Use Average =
73.41%
(O'Connell + Drickamer) / 2
Actual No. of Trays & Feed
Tray Location
No. Theoretical Trays =
Traty Efficiency =
Actual Trays =
Kirbride Feed Tray Ratio =
Trays above Feed =
Feed Tray Location =
Selected Feed Tray Ratio =
19
73.41%
26
1.25
15
11
1.3636
vs
1.25
by Kirbride
Tower Dimensions
26
3 ft. top trat to top tan line
14 spcs at 1.5 ft./spc
21 ft.
52 ft. Tan to Tan
12
11
3 ft. feed tray space
10 spcs at 1.5ft/sp
15 ft.
1
10 ft. to first tray
10 ft. shirt
Vessel Specs
Vessel Specs
Operating pressure =
Design Pressure =
Max Operating temp =
Design temp =
Material =
Corrosion allowance =
264.7
300
273.758
650
SA-516 Gr 70
0.0625
psia
psia
deg F
deg F
inches.
Check Flange Ratings
Flg Rating, psig
Flg Press, psig
Des Temp, °F =
300
541
650
(150/300/400/600)
Flgs O.K.
deg F
Cost of Towers Database v1.1
Tag No., T Description
Flow Sheet No.
No. Eqt Items
90 Actual Cost
90 Est'd Cost
Tower Type
Tower Dia., ft
T-T Length, ft
Design Press, psig
Corr Allow, in.
Yield Eff'y
Tower Material No.
Tray Option
No Trays
Tray Mtl No.
Packing Option
Packing Ht, ft.
Tray Cost-88
100
Depropanizer
1000
1
(26)
(4)
Tray Tower
2
52
300
0.0625
0.85 10% X-RAY
4
SA-516 Gr 70
1
Valve
26
1
T-410 SS
Cost Estimate for Tower with Trays
Select Tower No., TdB Item No.
Description
Flowsheet No.
No. Eqt Items
90 Actual Cost
90 Est'd Cost
Tower Type
Tower Dia., ft
T-T Length, ft
Design Press, psig
Back Calc'd des press =
Corr Allow, in.
Yield Eff'y
Tower Material No.
Tray Option
No Trays
Tray Mtl No.
Packing Option
Packing Ht, ft.
Tray Cost-88
Btm Wall Thickness =
Top Wall Thickness =
Vessel Wt =
Skirt Ht =
Skirt Wt =
Tray Wt =
Total Tower Wt =
301
(1 to 16 in Tower dB) Time Period
9
Fab Eqt Index
Depropanizer
CND$/US$
0
Duty US>CAN
1
0
Shell Cost =
0
Tray Cost =
Tray Tower
Platforms =
2
Tot Tower Cost =
52
300
306.06996 psig
0.0625
0.85
10% X-RAY
4
SA-516 Gr 70
1
Valve
26
1
T-410 SS
0
#N/A
0
0
0.53125 Inches
0.3125 Inches
5,800
lbs
12
ft.
2,345
lbs Assuming 3/4" Thk
980
lbs Assuming 10 ga. Wt.
9,125
lbs
2001
454
1.54
1
$79,204 Shop Fab
$9,268 Shop Installed
$8,218 Shop Installed
$96,690
Shortcut
Method for
Packed
Towers
Ekert Packing Factors
Some Ekert Wet Dumped Packing Factors, SF/Cf, for shortcut method
Diam. Inches
0.625
0.75
Material
Ceramic Super Intalox
Plastic Super Intalox
Ceramic Intalox Saddles
Metal Hy-Pak Rings
Plastic Pall Rings
97
Metal Pall Rings
70
Ceramic Berl Saddles
Ceramic Raschig Ring
380
Plastic Tellerettes
Plastic Mapak
User Choice - See Perry VI, P-18-23
Select Packing Factor =
155
145
170
255
sf/cf
1.00
60
33
98
42
52
48
110
155
40
FP
Note:
1.50
52
40
28
65
95
2.00
30
21
440
18
25
20
45
65
20
32
for shortcut method only
Approximate HETP
• From Tray Tower design, TS = 18 inches.
• For Approximated Packed Tower Design
assume one HETP = one Tray Spacing.
• HETP = 18 inches.
• Determine Tower Dimensions as for a Trayed
Tower.
• Allow 6 ft. for feed tray and top tray for liquid
distributer.
• No packing height should exceed 20 ft.
• If packing height exceeds 20 ft., must
redistribute liquid which adds another 6 ft.
FRI Packed Tower V1.2
FRI Packed Tower **FRIPT** Version 1.0, 14 Nov 93
Case Study =Bottom of Tower Example for CBE 497
By =
RAH
Dwg No. =
Tag No. =
Example from Perry III
CBE - 497
Input Data
Liquid Flow =
Liquid Density =
Liquid Viscosity =
Vapor Flow =
Vapor Density =
Vapor Viscosity =
No. Theoretical Trays =
Tower I.D. =
S = mG/L Factor =
9,933
36.31
0.08
4,812
2.117
0.01
19
2.00
1.01
lb/hr
lb/cf
Centipoise
lb/hr
lb/cf
Centipoise
Packing Factor =
Packing Type No. =
Packing Type =
Packing Size =
ft.
Minimum Value = 1.01
155
sf/cf
18
Ceramic Raschig Rings / Wet Packed
1
Inches
FRI Packed Tower Results
For 2 ft. Diameter Tower
From Generalized Eckert Pressure Drop Correlation
X = (WL/WG)(RHOG/RHOL)^0.5 =
Y = (G)^2(FP)/(gc*RHOG*RHOL) =
At X above, Y Flooding = YF =
% Flood At (X,Y) = (Y/YF)^0.5*100 =
Eckert Presure Drop =
0.4984
0.0113
0.0444
50.56%
0.20
Shortcut Method
Inches H2O/ft.
DP =
3.87
In. H2O
FRI Packed Tower Results
For 1.5 ft. Diameter Tower
From Generalized Eckert Pressure Drop Correlation
X = (WL/WG)(RHOG/RHOL)^0.5 =
Y = (G)^2(FP)/(gc*RHOG*RHOL) =
At X above, Y Flooding = YF =
% Flood At (X,Y) = (Y/YF)^0.5*100 =
Eckert Presure Drop =
0.4984
0.0358
0.0444
89.88%
1.77
Shortcut Method
Inches H2O/ft.
DP =
35.14
In. H2O
FRI Detailed Method for
Designing a Packed Tower
• Select a Packing Factor from 18
selected packing types.
• FRI have determined the design factors
which are too numerous to list here.
• FRI Packed Tower V1.2 will use this
packing data and the other data in the
shortcut method to design % Flood and
estimate the HETP.
FRI Detailed Method For PT
From New FRI Packing Correlations, Reports 92, 94, and 95, 1984
FS Factor = (V)(RHOG)^0.5 =
USGPM =
USGPM / SF =
a / Epsilon^3 factor =
Packing Height =
Packing Volume =
0.29
34.11
10.86
146.08
19.42
61.01
USGPM
USGPM/sf
sf/cf
Ft.
Cu. ft.
Dry Packing Pressure Drop =
Wet Packing Pressure Drop
0.46
0.67
In. H2O / ft.
In. H2O / ft.
Detailed Method
Top tower = 6 ft.
Feed tray 6 ft.
Btm packing =9 ft.
Top packing 12 ft.
btm tower = 12 ft.
Tan - tan = 42 ft.
Skirt =
10 ft.
Dry DP =
9.01
In. H2O
Wet DP =
13.05
In. H2O
Maximum Allowable Vapor Rate For Calculated HETP =
% Load or % Capacity =
4812 / 9150 =
9,150
52.59%
lb/hr
Maximum Stable Or Flood Vapor Rate At Unknown HETP =
% Flood =
4812 / 12279 =
12,279
39.19%
lb/hr
skirt ht. =
12 ft.
FRI HETP Values for 2 ft.
Diam.
FRI Vapor And Liquid Transfer Unit Values
Vapor Back Mixing Transfer Unit =
Liquid Phase Transfer Unit =
Vapor Phase Transfer Unit =
Overall Gas Phase Transfer Unit =
HDUG =
HTUL =
HTUG =
HTOG =
3.53
4.56
4.19
12.33
Inches
Inches
Inches
Inches
At Above Design Vapor / Liquid Rates,
HETP =
12.27
Inches
HETP Message =
Within 20% to 80% Capacity Limits
HETP Calculations (Good Only Between 20% to 80% Of Capacity
And For Level, FRI Tubed Drip Pan Distributor.)
Total Height of packing =
19.42
ft.
Packed Tower Cost Estimate
Select Tower No., TdB Item No.
Description
Flowsheet No.
No. Eqt Items
90 Actual Cost
90 Est'd Cost
Tower Type
Tower Dia., ft
T-T Length, ft
Design Press, psig
Back Calcd Design Pressure, psig =
Corr Allow, in.
Yield Eff'y
Tower Material No.
Tray Option
No Trays
Tray Mtl No.
Packing Option
Packing Ht, ft.
Tray Cost-88
Btm Wall Thickness =
Top Wall Thickness =
Vessel Wt =
Skirt Ht =
Skirt Wt =
Packing
Total Tower Wt =
302
(1 to 16 in Tower
TimedB)
Period
10
Fab Eqt Index
Depropanizer
CND$/US$
0
Duty US>CAN
1
$0
Shell Cost =
$0
Packing Cost =
Packed Tower
Platforms =
2
Tot Tower Cost
42
300
6.07
0.0625
0.85
10% X-RAY
4
SA-516 Gr 70
0
#N/A
0
0
#N/A
1
Ceramic Raschig Rings, 1 in
21
$0
0.4375
Inches
0.3125
Inches
4,482
lbs
12
ft.
2,345
lbs Assuming 3/4" Thk
0
Packing Installed In Field
6,827
lbs
2001
454.00
1.54
1.00
$61,120
$1,406
$7,319
$69,845
Shop Fab
Field Installed
Shop Installed
Summary
Diameter
Shell t-t
No. Theo Tr
Efficiency
No. trays
Tray Spg
HETP, inches
Pkg ht.
Shell Cost
Tr/Pkg Cost
Platforms
Total
Shortcut
Trays
2
52
19
73.41%
26
18
$79,204
$9,268
$8,218
$96,690
Shortcut
Pkd Twr
2
53
19
Detailed
Pkd Twr
2
45
19
18
29
12.27
21
$80,441
$1,942
$8,345
$90,728
$61,120
$1,406
$7,319
$69,845
Word of Caution – Trayed Towers
• Towers with trays are huge mixing
devices. Any slight restriction will cause
flooding.
• Three controlling factors:
•
(1) % Flood by Liquid and Vapor
Load
•
(2) % Spray Height by number of
holes.
•
(3) % Downcomer flood.
• Trays must be level and well supported.
Word of Caution – Trayed Towers
• Vendors will often quote towers with
many holes to reduce diameter and
obtain the bid.
• Later on detailed design, they find they
must reduce holes for specified
diameter.
• This increases spray height beyond
acceptable level and entrainment will be
too high.
• Buyers must be aware of all design
details.
Word of Caution – Packed
Towers
• Packed towers are low pressure drop
systems. Flows don’t always go where
they should.
• HETPs offered by vendors are
optimistic.
• Vendors claim a wide range of
operation.
• In actual practice there is a narrow
range.
Word of caution – Packed
Towers cont’d
• Uniform liquid distribution is difficult.
• If packing ht. Exceeds 20 ft.. Liquid
must be redistributed. This adds cost.
• Vapor is easily misdirected to walls.
• Vapor distributors are often required.
Good Luck On Your
Distillation Tower Design
• Presented to CBE 497
• 22 Jan 02
• R.A. Hawrelak