Download - Acetic anhydrid production
I
Abstract: The design of process to produce Acetic Anhydride from Acetone has
been discussed in this project. The purpose of this project is simply to
collect and present a set of simple calculation for this case study. For
calculation and design we use App A&C&E of Conceptual design of
chemical process by Douglas.
II
Table of Contents
Input Information ....................................................................................................................................................... 1
Decision Batch Versus continuous ............................................................................................................................. 5
Input-Out put structure ............................................................................................................................................... 7
Recycle structure ..................................................................................................................................................... 13
Separation system .................................................................................................................................................... 22
1
Level 0 - Input Information
The definition of reaction and importance information are as follow:
2
Input Information –level 0:
atmHCCoKetene
CHKeteneAcetoneC 1,700
2
10
42
4
+→
+→
atmdrideAceticAnhyAceticAcidKetene C 1,800→+
Product: Acetic Anhydride
Production rate: P = 16.5 mol/hr with 99% Purity
Cost Data:
Acetone:15.66 $/mol
Acetic Acide: 15 $/mol
Acetic Anhydride: 44.41 $/mol
Fuel: 4 $/106Btu
Reaction Data: ∆HR,acetone=34700 Btu/mol
∆HR,ketene=-27000 Btu/mol
∆HR,anhydride=20700 Btu/mol
Heat Values: Co: 0.122*106 Btu/mol CH4: 0.383*106 Btu/mol C2H4: 0.608*106 Btu/mol
S = moles of ketene leaving the pyrolysis reactor
moles of Acetone converted
3
selectivity data for this process is given in the 1958 AICHE Student contest problem, is as follows:
0.8 0.7 0.6 0.5 0.4 0.3 0.2 0.1 X
0.13 0.19 0.28 0.38 0.49 0.62 0.75 0.88 S
S=1-4/3 X
0
0.1
0.2
0.3
0.4
0.5
0.6
0.7
0.8
0.9
1
0 0.2 0.4 0.6 0.8 1
S
X
S=1- 4/3 X
4
Heat capacity data:
Component heat capacity
T=700oc Btu/mol T=80oc Btu/mol Co 7.89 6.97
16.97 9.15 CH4
11.55 C2H4 22.56 Acetic Anhydride 53.48 49.02 Acetone 38.25 19.98 Acetic Acid 31.66 23.43 Ketene 21.03 12.04
Vapor pressure Equation constants:
),()(
ln 0 KTKPaPCT
BAP sat ==++=
COMPONENT A B C
CO 41.66 -1109.88 5.46 CH4 31.35 -1307.5 -3.26 C2H4 48.11 -2473.7 -5.74 Acetic Anhydride 98.15 -8897.7 -12.16 Acetone 71.30 -5952 -8.53 Acetic Acid 61.34 -6768.8 -6.73 Ketene 14.00 -1849.2 -35.15
5
Level 1
Decision Batch Versus continuous
6
Choose a continuous process.
7
level 2
Input-Out put structure
We must answer these questions:
1. Should we purify the feed streams before they inter the
process?
2. Should we remove or recycle a reversible byproduct?
3. Should we use a gas recycle and purge stream?
4. Should we not bother to recover and recycle some reactants?
5. How many product streams will there be?
6. What are the design variables for the input-output structure,
and what economic trade-offs are associated whit these
variables?
8
Input-Out put structure-level 2:
1. Purify feed streams: the acetone and acetic acid feed streams are pure.
2. Reversible byproduct: we do not have such by product in our process.
3. Recycle and purge: hence we do not have any light component in the feed streams,
so no recycle and purge are needed.
4. Excess reactants: in the first reactor, as we have only one feed stream entering the
reactor, so there would be no case of using excess, but in the second reactor, we
consider using of excess acetic acid and try to monitor its effect on economic potential
of the process (trade-off).
5. Number of product streams:
The boiling point of the components and their destination are as follows:
9
COMPONENT NBP,(oF) DESTINATION
CO -312.6 Fuel byproduct CH4 -258.6 Fuel byproduct C2H4 -158.4 Fuel byproduct
Ketene -42.1 unstable reactant-completely converted
Acetone 133.2 Reactant-recycle to R1 – liquid
Acetic Acid 244.3 Reactant-recycle to R2 – liquid
Acetic Anhydride 281.9 Primary product So we have:
Two product streams: (CH4+C2H4+CO) and acetic anhydride.
6. Material balance and stream costs: by considering the production rate of
anhydride (16.58 mol/hr), we have:
S
PF =1 Acetone:
PF =2 Acetic acid:
We assume that the ketene is totally converted in the reactor.
10
=
=
=
−=−=−+−+=
−+−+=⇒
−=
−=
−=−=
=
hr
molF
hr
mol
SF
hr
molP
hr
mols
SS
S
PSS
S
PP
SS
PS
S
P
S
PP
SS
PP
S
PYPHC
SS
PP
S
PYPCO
S
PYPCH
HC
CO
CH
58.16
58.16
58.16
)2
3
2
5(
58.16)
2
3
2
5())1(
2
1)1(1(
)1(2
)1(
)1(22
1*:
)1(*:
*:
2
1
1
1
142
1
14
42
4
As we know the relation between s and x, and we also have the cost data of each
component, we can write the economic potential of this level and plot it versus the
conversion (or selectivity).
EP2 = Product value + byproduct value- raw material costs ($/yr) which for anhydride
process would be:
(We assume working hours per year= 8150 hr/yr)
−−−+−++= )*()*()**)1(2
()**)1(*()**()*(815021424244 ,,,2 FFHCHCVCOCOVCHCHVP CPC
S
PCHS
S
PCHS
S
PCH
S
PCPEP
−++= −− )10*4*10*122.0*)1(58.16
()10*4*10*383.0*58.16
(14.44*58.16*8150 6666
2 SSS
EP
)15*58.16()66.15*58.16
()10*4*10*608.0*)1(2
58.16( 66 −−−+ −
SS
S
)()()( 22 xhEPxgSsfEP =→=→=
11
X s EP2($/yr)*10^-6
0.1 0.88 1.836081107
0.2 0.75 1.505404865
0.3 0.62 1.036057939
0.4 0.49 0.317669789
0.5 0.38 -0.674113043
0.6 0.28 -2.251949367
0.7 0.19 -5.092054749
0.8 0.13 -9.170154784
We see that EP decrease with X increase Because more Acetone convert to by product.
Economic Potential vs. Conversion
-10
-8
-6
-4
-2
0
2
4
0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9
conversion (X)
EP2($/yr)*10^-6
12
LEVEL 2 – Alternatives:
There are no recycle and purge streams and we do not have any reversible
by product; so, we have no alternatives at this level of design.
13
LEVEL 3
Recycle structure
Questions that should be answered at this level
1. How many reactor systems are required? Is there any
Separation between the reactor systems?
2. How many recycle streams are required?
3. Do we want to use an excess of one reactant at reactor inlet?
4. Is a gas compressor required? What are the costs?
5. Should the reactor be operated adiabatically, with direct
heating or cooling or a diluents or heat carrier is required?
6. Do we want to shift the equilibrium conversion? How?
7. How do the reactor costs affect the economic potential?
14
LEVEL 3 – Recycle structure:
1. Because we have two set of reactions, with two different temperature
conditions, we must use two reactors.
2. There are two recycle streams:
a) Acetone recycle stream
b) Acetic acid recycle stream
3. We use excess amount of acetic acid at the inlet of the second reactor.
4. Because the recycle stream is totally liquid, so we do not need to use
any compressor. Thus, we need pumps in the recycle loops which their
cost is neglectable in comparison with other equipment costs.
15
Recycle material balances:
Acetone:
Acetone leave the first reactor=R1
Total acetone enters the first reactor=F1+R1
X
XFRRXRF
)1()1)(( 1
1111
−=→=−+⇒
Total flow entering the first reactor=FR1
SX
P
X
F
X
XFFRFFR ==
−+=+= 11
1111
)1(
(1-X)Acetone enters the second reactor= P SX
16
Other materials which enter the second reactor are:
−=
−=−=
=
=
)1(2
)1(
42
4
SS
PHC
SS
PP
S
PCO
S
PCH
PKeten
)(*
, 2222
MRPacidacetic
P
FR
ketene
acidaceticMRFRacidacetic
=→
+==+=
Total flow entering the second reactor=FR2
+−+−
=
+−+−+++−=
2
5)12(
2
)1(
)()1(2
)1()1(
2
2
MRS
X
X
S
PF
PMRSS
PS
S
P
S
PPX
SX
PF
R
R
Reactor heat effects:
a. First reactor: two reactors are taken place in the first reactor (we have
the cp of materials from the information data part).
−=∆=
∆−+∆=
)1292(**
*)1(*
111
1
OUTacetonePRRR
keteneAcetoneR
TCPSX
PTCFQ
HSS
PH
S
PQ
−=
−−+=
)1292(*25.38*58.16
)27000(*)1(58.16
34700*58.16
1
1
OUTR
R
TSX
Q
SSS
Q
17
−=
−−
+=
)1292(*185.634
)1(447660575326
1
1
OUTR
R
TSX
Q
SSS
Q
[ ])882.705307.201(1292
3.11
SXT
XS
OUT +−=
⇒−=
b. The second reactor:
∆=
∆=
TCFQ
HpQ
PRR
AnhydridR
22
2
)176(*)*)(*()*2
)1(*()*)1(*()*()*()*
)1((
4242 OUTacidHCCOCHketeneacetonR TCPMRPCPS
SPCPS
S
PCP
s
pcppCP
SX
XPQ −
+−
+−+++−
=
34320620700*58.162 ==RQ
)176(*)43.23*)(*58.16()55.11*2
)1(58.16()97.6*)1(*
58.16()15.9*
58.16()4.12*58.16()98.19*
)1(58.16(2 OUTR TMR
S
SS
SsSX
XQ −
+−
+−+++−
=
NOTE: the heat capacity values for each component in each reactor, is used at
the reactor condition.
Now we can plot the reactors effluent stream temperatures versus conversion
for each reactor:
18
X s Tout ( F)
0.1 0.88 1209.75
0.2 0.75 1145.86
0.3 0.62 1100.31
0.4 0.49 1073.12
0.5 0.38 1057.23
0.6 0.28 1052.63
0.7 0.19 1057.2
0.8 0.13 1057.54
Because reactions are endothermic so T decrease with reaction progress.
Reactor 1 exit temperature vs. conversion
1040
1060
1080
1100
1120
1140
1160
1180
1200
1220
0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9
conversion (x)
T out (F)
19
MR=2 MR=3 MR=5
X s T out (F) T out (F) T out (F)
0.1 0.88 100.9284733 106.8078924 116.178103
0.2 0.75 62.43109104 75.36714614 94.03872714
0.3 0.62 44.17186029 61.28847212 84.93602812
0.4 0.49 40.13863227 58.24668457 83.02951251
0.5 0.38 43.90993531 61.09019948 84.81112102
0.6 0.28 55.84760157 70.23193138 90.66427054
0.7 0.19 75.91147661 86.09653086 101.2996208
0.8 0.13 94.29716962 101.2133133 112.041488
We see that with high MR T has changed little because asetic asid treat like heat carrier.
0
20
40
60
80
100
120
140
0 0.2 0.4 0.6 0.8 1
Tout
X
MR=2
MR=3
MR=4
20
REACTOR COST:
We assume that the pyrolysis reactor (first reactor of the process) cost is
calculated as a furnace cost, and if anhydride reactor cost is neglectable
in comparison with the first reactor cost, then we would have:
)27.1(*)10*52.5)(280
&($,cos 85.03
CFQSM
tInstalled +=
=
=
=
++=
pressurecatmospheriinrunsprocesstheasF
materialtuberadiantassteelcarbonuseweifF
typepyrolysisaforF
FFFF
P
m
d
pmdC
)7.14(0
0
1.1
Also we need M&S factor in our calculation, we assume M&S factor=1350
[ ]
+=
−−+=
→∆−+∆==
SS
Q
SSS
PQ
HSS
PH
S
PQQ keteneAcetoneR
27000770058.16
)27000(*)1(*58.16
)34700(*58.16
*)1(*1
)37.2(*10
)270007700(*58.16
)10*52.5)(280
1350($,cos
85.0
6
3
+=
SStInstalled
In order to write installed cost in annualized basis we should calculate capital charge factor (ccf): We assume: interest rate: i=0.2 Number of operating years: 10 yr
[ ] [ ]1
10
10
24.01)2.01(
)2.01(2.0
1)1(
)1( −=−+
+=
−+
+= yr
i
iiCCF
n
n
21
)*cos$,cos. CCFtinstalledtInstalledAnn =
tactorEPEP cosRe23 −=
X s EP2($/yr)*10^-6 EP3($/yr)*10^-6
0.1 0.88 1.836081107 1.826378412
0.2 0.75 1.505404865 1.49535338
0.3 0.62 1.036057939 1.025515021
0.4 0.49 0.317669789 0.306382362
0.5 0.38 -0.674113043 -0.686414373
0.6 0.28 -2.251949367 -2.265834331
0.7 0.19 -5.092054749 -5.108714167
0.8 0.13 -9.170154784 -9.190662659
By subtracting EP2 from reactor cost we can calculate EP3 as function
of design variable. Of course the optimum from this figure is not true
because we don’t consider distillation column cost.
Economic Potential vs. conversion
-10
-8
-6
-4
-2
0
2
4
0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9
conversion (x)
EP3 ($/yr)*10^-6
85.02700077000
*308.1/$,cos.
+=
S
SyrtInstalledAnn
22
level 4
Separation system
We don’t have gas recovery system so we must design liquid sepration
system.
To synthesize the liquid separation system the following decisions should
be made:
1. How should the light ends be removed if they might
contaminate the product?
2. What should be the destination of the light ends?
3. Do we recycle components that form zoetrope’s with the
reactants, or do we split the azeotropes?
4. What sequence of columns do we use?
5. How should we accomplish separations if distillation is not
feasible?
23
Separation system-level 4:
Light ends: We assume most of light ends product (CH4, C2H4 , CO)
departed in flash, and with neglecting amount of light ends product at
liquid part.
Azoetrope’s: we don’t have any zoetrope’s.
Sequence of columns: we consider two atmospheric distillation column
in separation system. The heuristic of lightest first, most plentiful first,
and favor equimolar splits all suggest direct sequence.
Note: we are neglecting Acetone, Acetic Acid, Anhydride losses in flash.
With assuming X=.3 , s=.6 , MR=2 we design two distillation columns.
(P=16.58)
P/S(1-X) = 19.3:Fresh acetone Feed
P(MR)-P=16.58:Fresh Acetic Acid Feed
P=16.58:Produced Anhydride
24
653.0)58.16*2(4.62
4.62=
+=FX
Temperature at top is N.P of Acetone 329.2K , and Temperature at
bottom is mean N.P of Anhydride and Acid Asetic:
97.4
34.36.61
206:401@
4.758.5
56.43:2.329@
=→
=→
=
=
=→
=
=
avg
bottomsat
acidacitic
sat
acetone
topsat
Acetone
sat
Acetone
P
PK
P
PK
α
α
α
With assuming 99.5% recovery Acetone and 99.5% Anhydride and Acid
Asetic we have:
62.4*0.995= 62.088 :Acetone at top
16.58*(1-0.995) = 0.0829 Acetic acid at top:
9986.00829.0088.62
088.62=
+=DX
00934.058.16497.16312.0
312.0=
++=WX
25
We consider feed is saturated liquid:
386.0653.0*)197.4(
1
*1
1=
−=
−≈
F
mX
Rα
463.0386.0*2.12.1 === mRR
006.7
)97.4ln(
)00934.0
00934.01*
10*4.1
9986.0ln(
)ln(
)1
*1
ln(3
=→
−
=
−
−=
−
m
avg
W
W
D
D
m
N
X
X
X
X
EqNFenskeα
528.19
)463.1
386.0463.0(175.0
1
006.7
)1
(175.01
:
5688.0
5688.0
=→
−−=
+
−
+
−−=
+
−
N
N
N
R
RR
N
NNEqGilland mm
Tower height:
With considering: 2 ft towers space and 15ft at top and bottom end we
have:
.H=2*40+15=95 ft
401)5.0
528.19(
5.0
528.19
0
=
+=→== actact NE
NN
26
Tower Cross – sectional Area and Diameter:
VTbMA *)460(*10*124.2 4 += −
For bottom of distillation column:
075.812
1.10205.60
2=
+=
+= anhydrideacidacetic
avg
MMM
FTb o
avg 1.2632
9.2813.244=
+=
ftA
D
ftA
DRV
44.24
677.4
95.90)0829.0088.62(*463.1)1(
2
==
=
→=+=+=
π
yrtAnn
$3741824.0*)118.2(*)95(*)44.2(*)9.101(*)
280
1350($,cos. 802.0066.1 =+=
Acetone column condenser:
)90120(***** −=∆=∆= PCmCCVc CWTAUVHQ
We consider water at 900F enter and exit at 1200F:
.
95.90
12530)..(
10002
=
=∆
=
V
H
ffthr
BTUU
V
C
[ ]FTm
082.8)1202.133()902.133(ln
90120=
−−−
−=∆
21292
82.8*100
95.90*12530ftAC ==
27
Install cost for column:
)29.2(*)3.101)(280
&($, 65.0
CFASM
+=
yrtAnn
$4059924.0*)29.3(*)1292)(3.101)(
280
1350($,cos 65.0 ==
Cooling Water cost & cooling tower :
yryr
hr
hr
lb
lb
gal
galtAnn
$2227)8150(*
30
95.90*12530)
341.8
1)(
1000
$06.0($,cos =
=
Acetone column reboiler:
mRRSSVR TAUHWVHQ ∆=∆=∆= **
The boiling point of mixture at bottom is 4100k (assuming ∆� = 30�� ) so
we can use stream with 67 PSI pressure and 4120k temperature.
223.11411250
95.90*14130
910,11250,14130
ftA
lb
BtuHUH
R
m
SR
==
=∆==∆
yrtAnn
$839024.0*)29.3(*)23.114)(3.101)(
280
1350($,cos 65.0 ==
The stream cost:
yryr
hr
hr
lb
lbtAnn
$25897)8150(*
910
95.90*14130*)
1000
$25.2($,cos =
=
28
Anhydride column (second column):
Acetone =0.005*62.4=0.312
Acetic Acid=0.995*16.58=16.49
Anhydride =16.58
494.0312.058.16497.16
497.16=
++=FX
With assuming 99.5% recovery Anhydride and 99% Anhydride we have:
Anhydride at bottom: 16.58*0.995=16.497
Acetic Acid at bottom:(16.497)/(0.99)*(1-0.99)=0.167
Acid Acetic at top : 16.497-0.167=16.33 Anhydride at top of column: 16.58-16.497=0.083 Acetone at top : 0.312
976.0312.0083.033.16
33.16=
++=FX
310*84.9497.16167.0
167.0 −=+
=WX
Temperature at Top is N.P of Acetic Acid (319K) and anhydride at
bottom(412K):
29
85.2
86.26.80
16.28:412@
84.268.16
42.47:391@
=→
=→
=
=
=→
=
=
avg
bottomsat
acidacitic
sat
ahydrid
topsat
Anhydrid
sat
acidAcetic
P
PK
P
PK
α
α
α
We consider feed is saturated liquid:
094.1494.0*)185.2(
1=
−=mR
313.1)094.1(*2.1 ==R
94.7)85.2ln(
)10*84.9
)10*84.9(1*
976.01
976.0ln(
:3
3
=
−
−=−
−
mNEqFenske
035.19)313.2
094.1313.1(175.0
1
94.7: 5688.0 =
−−=
+
−
N
NEqGilland
3915.0
035.19=
+=actN
Tower height:
H=2*39+15=93 ft
30
Tower Cross-sectional Area and Diameter:
24
4
26.2)685.38(*)4609.281(*1.102*10*124.2
685.38)083.033.16312.0(*)1313.1()1(
*)460(*10*124.2
ftA
DRV
VTbMA
=+=
=+++=+=
+=
−
−
ftA
D 7.14
==π
yrtAnn
$2502524.0*)18.3(*)93(*)7.1(*)9.101(*)
280
1350($,cos. 802.0066.1 ==
Anhydride column condenser:
)90120(***** −=∆=∆= PCmCCVc CWTAUVHQ
[ ]FTm
082.8)1203.244()903.244(ln
90120=
−−−
−=∆
244082.8*100
685.38*10030ftAC ==
yrtAnn
$2015824.0*)29.3(*)440)(3.101)(
280
1350($,cos 65.0 ==
Cooling Water cost & cooling tower :
yryr
hr
hr
lb
lb
gal
galtAnn
$758)8150(*
30
685.38*10030)
341.8
1)(
1000
$06.0($,cos =
=
Anhydride column reboiler:
The boiling point of mixture at bottom is 4120k (assuming ∆� = 30�� ) so
we can use stream with 90 PSI pressure and 4330k temperature.
31
235.6011250
685.38*17550
895,11250,17550
ftA
lb
BtuHUH
R
m
SR
==
=∆==∆
yrtAnn
$552524.0*)29.3(*)35.60)(3.101)(
280
1350($,cos 65.0 ==
stream cost:
yryr
hr
hr
lb
galtAnn
$13910)8150(*
895
685.38*17550)
1000
$25.2($,cos =
=
Acetone column($/yr) Anhydride Column($/yr) cost
37418 25.25 Ann. Install column cost
40599 20158 Ann. Condenser cost
2227 758 Ann. Cooling water cost
8390 5525 Ann. Reboiler cost
25897 13910 Ann. Stream cost
114531 65376 Total cost
Total cost of two column : 179907 $/yr
For x=0.3 we have:
EP4= EP3-Costs of two towers=1025515-179907
EP4=845608 $/yr for x=0.3
It seems that for getting good profit we have to consider x below 0.3
NOTE: We have numerous alternative for this part like condenser instead
flash .
32
THE END