CO2 removal from natural gas by employing amine absorption and membrane technologyA technical and economical analysis

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Chemical Engineering Journal 172 (2011) 952– 960
Contents lists available at ScienceDirect
Chemical Engineering Journal
jo u r n al hom epage: www.elsev ier .com/ locate /ce j
O2 removal from natural gas by employing amine absorption and membrane
echnology—A technical and economical analysis
ars Petersa, A. Hussainb, M. Follmanna, T. Melina, M.-B. Häggb,∗
AVT-Chemical Process Engineering, RWTH Aachen University, Turmstraße. 46, 52064 Aachen, Germany
Department of Chemical Engineering, Norwegian University of Science and Technology, N-7491 Trondheim, Norway
 r t i c l e i n f o
rticle history:
eceived 22 March 2011
eceived in revised form 1 July 2011
a b s t r a c t
A technical and economic analysis of gas sweetening processes for natural gas with amine absorption
and membrane technology has been conducted. Amine absorption is still considered a state of the art
technology for gas sweetening but membranes have shown a great potential in this area, if the flux and
ccepted 3 July 2011
eywords:
as sweetening
spen Hysys
arbon dioxide capture
mine absorption
selectivity for CO2 is high enough. The PVAm/PVA (polyvinyl amine and polyvinyl alcohol) membrane,
developed at NTNU shows such qualities. For two different feed gas cases a simulation analysis with
Aspen Hysys for amine absorption and a membrane model interfaced within Aspen Hysys was performed.
Further, an economical analysis was also conducted to evaluate the total capital investment and gas
processing cost for both technologies.
© 2011 Elsevier B.V. All rights reserved.
ixed site carrier membrane
. Introduction
Natural gas sweetening is a very important issue for many rea-
ons; the acid gas concentration in natural gas can cause pipeline
orrosion problems during transport [1]. The removal of acid gases
educes the gas volume to be transported and increases the calorific
alue of sold gas stream [2].
The captured acid gases can be pumped back in the reservoir;
hich reduces the atmospheric pollution by impurities like H2S,
nd the emission of greenhouse gases as CO2 [1]. The pressure
nd composition of crude natural gas varies from well to well. The
ressure range mostly lies between 20 and 70 bar [4] but can be
ven higher for some wells. Table 1.1 demonstrates the wide range
f compositions in crude natural gas. Processing of natural gas is
o far the largest gas separation application worldwide. Almost all
rude natural gas streams require treatment to remove impurities
nd reduce higher hydrocarbons to match with tightly controlled
ipeline specifications (typical values are given in Table 1.2). CO2
sually has to be lower than 2 vol.% in sale (natural) gas [5], a
equirement that generally makes CO2 removal necessary.
The scope of this study is to investigate the removal of CO2 from
atural gas with amine absorption and membrane plants. Removal
f CO2 from crude natural gas by amine absorption is a well known
nd implemented industrial process [3,5,6] and is still considered
∗ Corresponding author. Tel.: +47 735 94033.
E-mail address: may-britt.hagg@chemeng.ntnu.no (M.-B. Hägg).
385-8947/$ – see front matter © 2011 Elsevier B.V. All rights reserved.
oi:10.1016/j.cej.2011.07.007
a state of the art technology. Likewise, there are quite a few mem-
brane plants installed around the world, but these membranes do
not have optimum performance with respect to flux and selec-
tivity, and therefore require fairly large membrane areas. Current
research in this area is focused on optimizing the membrane sep-
aration performance thus reducing both membrane area and CH4
loss. Membrane processes are considered as a promising alterna-
tive for offshore production processes. Since there will always be a
little loss of amine solution to the atmosphere during the process,
membrane processes also offer an environment friendly alterna-
tive.
Two different feed gas streams are considered in this study.
For each stream a technical and economical analysis will be
done to assess the optimal process conditions, required capital
investment and the resulting gas processing cost for amine absorp-
tion and different membrane configurations, respectively. Aspen
Hysys has been used to simulate the amine absorption and mem-
brane separation process. Hysys is a powerful process simulation
tool, which provides the opportunity to estimate physical proper-
ties, liquid vapor phase equilibrium, material and heat balances.
Membrane operations are not standard units in Hysys, but
it is possible to include user defined unit operation. Therefore
a membrane model (ChemBrane) developed at NTNU has been
implemented in this analysis. This model enables three different
membrane configurations (co-current, perfectly mixed and counter
current flow) to estimate realistic process conditions. ChemBrane
can handle vacuum and sweep operations [7]. Table 1.3 summarizes
the different cases for feed gas conditions. For amine absorption
L. Peters et al. / Chemical Engineering Journal 172 (2011) 952– 960 953
Table 1.1
Different natural gas composition [3].
Rio Arriba
County, New
Mexico
Western
Colorado
Canada
(Alberta)
Helium 0 0 0
Nitrogen 0.68 26.1 3.2
CO2 0.82 42.66 1.7
Hydrogen sulfide 0 0 3.3
CH4 96.91 29.98 77.1
p
a
C
c
F
p
a
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2
w
2
2
h
(
H
h
a
a
c
C
C
C
C
H
H
H
i
c
a
2
t
f
T
Table 1.3
Summary of feed gas conditions.
Case 1 Case 2a Case 2b
Flow rate [Nm3/d] 1.0 × 107 1.0 × 107 7.0 × 105
Flow rate [kmol/h] 18560 18560 1301
Temperature [◦C] 8 60 60
Pressure [bar] 115 90 90
Mole fraction
CO2 0.029 0.095 0.095
H2S – 20 ppm 20 ppm
H2O 10 ppm 10 ppm 10 ppm
CH4 0.971 0.724 0.724
Ethane – 0.089 0.089
Propane – 0.052 0.052
T
P
Ethane 1.33 0.55 6.6
Propane 0.19 0.28 3.1
Butanes 0.05 0.21 2
Pentanes and heavier 0.02 0.25 3
rocess and membrane application the property packages of Amine
nd Peng-Robinson are used, respectively. For amine absorption, a
O2 concentration of less than 0.5 vol.% in the sweet gas and a con-
entration above 90% CO2 in the acid gas stream were assumed.
or single stage membrane processes the aim was to match the
ipeline conditions, and to minimize the loss of CH4 in the perme-
te stream. In multi-stage membrane processes a similar goal for
O2 purity and recovery was set as in amine absorption; less than
% CO2 in the product stream and a CO2 concentration around 90%
ithin the permeate stream.
. Amine absorption process
.1. General chemistry
Treatment of natural gas with aqueous diethanolamine solution
as a long history and is state of the art technology. Diethanolamine
DEA) is a secondary amine and therefore less reactive with CO2 and
2S than primary amines like monoethanolamine (MEA). Hence it
as lower energy requirements for the regeneration [1], which is
 key factor for estimating gas processing cost. The absorption of
cid gases into aqueous amine solution is described by the following
hemical reactions [6,8]:
O2 + 2R1R2NH ↔ R1R2NCOO− + R1R2NH2
+ (2.1)
O2 + OH− ↔ HCO3 (2.2)
O2 + H2O ↔ HCO3
− + H+ (2.3)
O2 + R1R3COHCNH + OH− ↔ R1R3COCO2CNH− + H2O (2.4)
2O ↔ H+ + OH− (2.5)
2S ↔ H + HS− (2.6)
S− + R1R2NH ↔ S2−R1R2NH2+ (2.7)
The maximum concentration of diethanolamine in water is lim-
ted due to its corrosive behavior. With additives, which decrease
orrosion, a concentration up to 35 wt.% is feasible [3]. For this work
 30 wt.% solution is assumed.
.2. General process descriptions
Fig. 2.1 shows a simplified flow sheet of a typical amine absorp-
ion process. The crude natural gas enters a separator to remove
ree liquids before entering the bottom of the absorber column.
he lean amine solution enters the absorber column at the top and
able 1.2
ipeline specifications [5].
Component CO2 Total water H2S 
Specificationflows down in counter current mode with the up streaming crude
natural gas. The acid gas components react with amine and dissolve
into the liquid phase.
Sweet natural gas leaves the column at the top. The enriched
DEA solution leaves the absorber column at the bottom. The pres-
sure of the rich amine solution is reduced to nearly atmospheric
pressure and a separator split the gas stream from the liquid solu-
tion. The vent stream can contain higher hydrocarbons and CH4
which is often used as fuel for the regeneration column. The liquid
stream flows through the rich-lean-heat exchanger and enters the
regeneration column. Inside the regeneration column, the solution
is heated up to temperatures around 125 ◦C, to strip off the acid
gases from the solvent which leave the stripper column from the
top. A hot and regenerated DEA solution leaves the regeneration
column through the bottom and flows through the rich-lean-heat
exchanger to exchange heat with CO2 enriched solution. After the
rich-lean-heat exchanger, make up solvent and water are added
and the stream is re-pressurized and cooled until it matches the
absorber’s inlet conditions. The acid gas stream of the stripper col-
umn is cooled and flashed to recapture water. The water is piped
back into the regeneration column to reduce the water losses dur-
ing the process [3,6].
2.3. Optimal process conditions
Optimal process conditions lead to lower capital and invest-
ment costs which ultimately result in lower gas processing cost.
The key factors for amine absorption plants are thermal re-boiler
duty, working capacity and the solvent circulation flow rate [9].
The working capacity is limited through the corrosive behavior of
amine solution. Thermal re-boiler duty and solvent flow rate are a
function of the lean solution loading [mole acid gas/mole amine].
Fig. 2.2 shows the impact of lean amine loading on the solvent flow
rate and specific thermal re-boiler duty.
The lean solution loading increases linearly by increasing the
lean solvent flow rate. Since the rich amine loading is limited by the
stoichiometry, the working capacity of aqueous solution is reduced.
A larger solvent flow is required to match the simulation goals of
0.5% CO2 in the product stream. The specific thermal re-boiler duty
decreases with an increase in the lean solution loading. For low
loading values (thickness, pf and pp are pressure on feed and permeate side, xi and
yi are the fractions of component i in feed and permeate stream.
The above expression can be re-written as,
Ji = Di
l
(ci,f − ci,p) (3.2)
where Di is the Fick’s diffusion coefficient and ci,f and ci,p the con-
centration of component i on the feed and permeate side of the
membranes. Standard polymeric membranes loose part of their
separation efficiency when swollen. However, the properties of
developed fixed site carrier membrane will enhance when swollen
[20,21]. The PVAm/PVA membrane combines the advantages of
supported liquid membranes and solid polymeric membranes [18].
The membrane material polyvinylamine contains an amino group
which contributes to the transport of CO2 within the membrane
when the membrane is wet with water. Following reversible reac-
tions will occur in the membrane:
CO2 + H2O ↔ HCO−
3 + H+ (3.3)
[CH2 − CH2 − NH+
2 ] + HCO−
3 ↔ complex (3.4)
Hence the equation for the flux through PVAm/PVA membrane
is modified to
Ji = Di
l
(ci,f − ci,p) + Dic
l
(cic,f − cic,p) (3.5)
where Di is Fick’s diffusion coefficient and ci,f and ci,p are the con-
centration of component i on the feed and permeate side of the
956 L. Peters et al. / Chemical Engineering Journal 172 (2011) 952– 960
Table 2.2
Values economic analysis.
Value Unit Source Value Unit Source
Wellhead price for natural gas 4.2 Dollar/MMBTU [28] Direct labor 15 Dollar/h [25 MMSCFD feed] [23]
Solvent degradation 0.25 kg/t CO2 captured [6] Working time 8 h/d [23]
Cost for diethanolamine 2.4 Dollar/kg [27] Overall labor cost 2.15 Direct labor [23]
Cost for reboiler steam 14.5 Dollar/t [27] Maintenance cost 0.05 TCI [1]
Cost for electricity 0.07 Dollar/kWh Capital recovery cost 0.277 TCI [1]
Cost for cooling water 0.01 Dollar/t [27] Plant on stream factor 0.97
Cost for process water 0.5 Dollar/t [27]
m
s
f
a
m
h
0
p
m
c
u
0
p
A
u
p
u
b
c
p
3
m
[
T
P
Fig. 3.1. schematic transport through FSC membrane.
embranes and permeate stream, Dic the carrier facilitated diffu-
ion coefficient, and cic,f and cic,p are the complex concentration on
eed and permeate side, respectively (Fig. 3.1).
The CO2 transport benefits from the reversible reaction between
mino group and the CO2. All other components can just per-
eate through the membrane by solution-diffusion. A membrane
aving a CO2/CH4 selectivity up to ∼40 and a permeance of
.3 m3(STP)/m2 h bar will be competitive with amine absorption
rocess for natural gas sweetening. Commercial cellulose acetate
embranes have a selectivity of 12–15 under normal process
onditions [18]. For this work a selectivity of 35 (although val-
es up to 45 are measured in the lab) and a permeance of
.3 m3(STP)/m2 h bar is assumed. The permeances for other com-
onents are derived from Baker’s Membrane Technology and
pplications [22]. Table 3.1 summarizes the permeance values,
sed in this analysis.
The procedure to evaluate total capital investment cost and gas
rocessing costs for membrane plants is similar to the procedure
sed for amine absorption. First the feasibility of different mem-
rane configurations is investigated to identify the optimal process
onditions and afterwards a technical and economical analysis is
erformed.
.2. Membrane configurations
In the past, many researchers have investigated different
embrane configurations, to obtain optimal process conditions
1,2,4,23,24]. Summing up on all those works, two different mem-
able 3.1
ermeance values [22].
CO2 H2S H2O CH4
Permeance
[mol/(k Pa m2)] 0.134 1E−09 1E−09 0.00381 
a The support membrane, PSf, is hydrophobic, hence very low H2O permeance has been
Fig. 3.2. Simplified single stage process.
brane configurations promised good results for our given feed
conditions and goals, a single stage configuration without any recy-
cle streams and a 2 stage cascade configuration. A single stage
process has always lower investment cost than a multistage process
but has higher gas processing cost [24].
3.2.1. Single stage configuration
Fig. 3.2 illustrates the simplified flow sheet for a single stage
membrane process, without any recycle streams. The crude natu-
ral gas flows over the feed side of the membrane. Along the way,
CO2 permeates through the membrane to the permeate side. The
retentate leaves the membrane with nearly the same pressure as
the feed. On the permeate side, a permeate stream enriched with
CO2 leaves the membrane.
3.2.2. Optimization
Since the feed conditions are given, the only influencing param-
eter on the simulation results is the permeate pressure. Figs. 3.3–3.5
show the influence of permeate pressure on permeate temperature,
required membrane area and the permeate composition for case 1
and 2a/b respectively. A decreasing permeate pressure leads to a
decreasing permeate temperature due to Joule Thompson effect
[25]. For case 2a/b this effect is stronger than case 1 due to pres-
ence of higher hydrocarbons in the feed. Low temperatures must
be avoided because of condensation on the membrane surface–this
can lead to undesired membrane fouling (Fig. 3.3) [5]. An increase
in the permeate pressure will give an increase in required mem-
brane area. The driving force for gas permeation is primarily based
on partial pressure difference between the feed and the permeate
side. An increase in permeate pressure reduces the driving force and
therefore an enlarged membrane area is necessary to reach the sep-
aration goals (Fig. 3.4). A variation in the permeate pressure has also
a major impact on the composition of the permeate stream. With
Ethane Propane Butane Pentane Hexane
0.00572a 0.00858a 0.0129a 0.0192a 0.029a
 measured.
L. Peters et al. / Chemical Engineering Journal 172 (2011) 952– 960 957
0
5
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20151050
permeate pressure [bar]
-15
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-5
0
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pe
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ea
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 [°
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pe
rm
ea
te
 te
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pe
ra
tu
re
 [°
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]
Case 2 a
Case 2b
F
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2520151050
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as
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2 
a:
 m
ol
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fr
ac
tio
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pe
rm
ea
te
permeate pressure [bar]
CO2
CH4
Ethane
Propane
Butane
Pentane
Hexane
0
0,1
0,2
0,3
0,4
0,5
0,6
2520151050
C
as
e 
2 
b:
 m
ol
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fr
ac
tio
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pe
rm
ea
te
permeate pressure [bar]
CO2
CH4
Ethane
Propane
Butane
Pentane
Hexane
permeate pressure [bar]
ig. 3.3. Influence permeate pressure on the permeate temperature, top case1, bot-
om case 2a/b.
ncreasing permeate pressure the mole fraction of CO2 decreases
Fig. 3.5).
Since the aim for an optimal single stage process is to reduce the
osses of CH4 and higher hydrocarbons to a minimum, the permeate
ressure should be kept as low as possible, but considering the
roblems related to low temperature and membrane fouling.
0
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em
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ar
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²]
permeate pressure [bar]
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9000
0
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60000
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as
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2 
b:
 m
em
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an
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ar
ea
 [m
²]
C
as
e 
2 
a:
 m
em
br
an
e 
ar
ea
 [m
²]
permeate pressure [bar]
Case 2 a
Case 2 b
ig. 3.4. Influence permeate pressure on the membrane area, top case1, bottom case
 a/b.
Fig. 3.5. Influence permeate pressure on the membrane area, top case1, middle case
2a, bottom case 2b.
3.2.3. Single stage simulation results
Table 3.2 summarizes the simulation results from Aspen Hysys
for a single stage process. A content of maximal 2% CO2 in the reten-
tate stream is set as target for the simulation. For all 3 cases the goal
of 2 vol.% in the retantate could be achieved. The permeate pressure
Table 3.2
Simulation results summary.
Case 1 Case 2a Case 2b
Membrane area [m2] 6636 49278 3482
Permeate temperature [◦C] 3.6 4.84 4.84
Permeate pressure [bar] 2.5 7 7
Product (retentate) flow [kmol/h] 18120.86 15191 1061
Mole fraction
CO2 0.02 0.02 0.02
H2S – 0.000245 0.000245
H2O – 2.3E−16 2.3E−16
CH4 0.98 0.801 0.801
Ethane – 0.0940.091
Propane – 0.051 0.051
Butane – 0.026 0.026
Pentane – 0.005 0.005
Hexane – 0.003 0.003
Permeate flow [kmol/h] 468.70 3398 240
Mole fraction
CO2 0.377 0.431 0.431
H2S – 1.28E−15 1.28E−15
H2O – 2.0E−15 2.0E−15
CH4 0.623 0.381 0.381
Ethane – 0.067 0.067
Propane – 0.055 0.055
Butane – 0.042 0.042
Pentane – 0.0115 0.0115
Hexane – 0.0124 0.0124
CO2 recovery [%] 32.8 83.2 83.2
CH4 loss [%] 1.6 9.6 9.6
958 L. Peters et al. / Chemical Engineering Journal 172 (2011) 952– 960
ge membrane cascade with retentate recycle.
f
a
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a
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w
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3
o
f
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p
l
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m
c
i
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3
a
o
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o
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o
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a
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7000
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se
co
nd
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ag
e:
 m
em
br
an
e 
ar
ea
 [m
²]
m
em
br
an
e 
ar
ea
 [m
²]
feed pressure second membrane stage [bar]
membrane area 1
total membrane area
membrane area 2
0
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30000
35000
40000
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20000
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se
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ag
e:
 m
em
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an
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ar
ea
 [m
²]
m
em
br
an
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ar
ea
 [m
²]
feed pressure second membrane stage [bar]
membrane area 1
total membrane area
membrane area 2
500
1000
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7000
se
co
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e:
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ar
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 [m
²]
m
em
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ar
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 [m
²]
membrane area 1
total membrane area
membrane area 2
Fig. 3.6. simplified flowhseet of a 2-sta
or case 1 is set to 2.5 bar and for case 2a/b to 7 bar. CO2 recovery
nd CH4 loss are calculated as
 = CO2p
CO2f
× 100% (3.6)
nd
H4loss =
(
1 − CH4s
CH4f
)
× 100% (3.7)
here CO2p, CO2f, CH4s and CH4f as the concentration of CO2 and
H4 in permeate (p), sweet (s) and feed (f) gas stream. A recovery of
2 and 83% of CO2 is calculated for case 1 and case 2a/b. The losses
f CH4 in the process are calculated to be 1.6% for case 1 and 9.7%
or case 2a/b. The difference in the permeate pressure is caused by
he different influence of Joule Thompson effect. It is quite clear
hat for case 1 with the low CO2-content in the feed, a single stage
rocess is not acceptable as the CO2 recovery is too low and CH4
oss is too high. A multistage process was therefore investigated.
.2.4. Multi stage configuration
Fig. 3.6 illustrates a simplified flow sheet of a two stage cascade
embrane system. A multistage configuration reduces the hydro-
arbon losses to a minimum, however, those plants have higher
nvestment costs than single stage configurations [2,4,23,24]. The
ermeate stream of the first membrane serves as feed for the second
embrane. Therefore, the permeate stream needs to be recom-
ressed and cooled. The retentate stream of the second membrane
tage is recompressed, cooled and recycled as feed to the first stage.
he retentate stream from the first stage is collected as product gas.
.2.5. Optimization of the 2-stage membrane process
Optimization of 2-stage system is similar to the optimization of
 single stage process. Favre et al. [26] concluded that compression
f a stream at the feed side is always more expensive than on the
ermeate side. Hence the recompression of first permeate stream
ust be considered carefully. Figs. 3.7 and 3.8 present the impact
f recompression on the required membrane area and the specific
O2 capture duty [kJelec./kg CO2 captured]. The permeate pressure
f the first membrane stage was set to 2.5 and 7 bar, similar to
he single stage process. Fig. 3.7 shows the impact of increasing
eed pressure on the second membrane stage. As a reminder at this
oint, a content of 2% CO2 in the retentate stream of the first stage
nd a content of 90% CO2 in the permeate stream of the second
00
353025201510
feed pressure second membrane stage [bar]
Fig. 3.7. Impact of increasing feed pressure on membrane area, top case 1, middle
case 2a, bottom case 2b.
L. Peters et al. / Chemical Engineering Journal 172 (2011) 952– 960 959
980
1000
1020
1040
1060
1080
1100
1120
1140
1160
1180
50403020100sp
ec
 e
le
c 
C
O
2 
ca
pt
ur
e 
du
ty
 [
kJ
/k
g 
C
O
2 
ca
pt
ur
ed
]
feed pressure second membrane stage [bar]
440
445
450
455
460
465
470
475
480
353025201510
sp
ec
 e
le
c 
C
O
2 
ca
pt
ur
e 
du
ty
 [k
J/
kg
 C
O
2 
ca
pt
ur
ed
]
feed pressure second membrane stage [bar]
440
445
450
455
460
465
470
475
480
353025201510
sp
ec
 e
le
c 
C
O
2 
ca
pt
ur
e 
du
ty
 [k
J/
kg
 C
O
2 
ca
pt
ur
ed
]
F
m
m
s
d
n
c
s
s
o
g
O
[
o
r
3
c
a
Table 3.3
Summary of simulation results for 2-stage membrane system.
Case 1 Case 2a Case 2b
Membrane area 1 [m2] 8141 57556 4091
Membrane area 2 [m2] 2863 21154 1491
Total membrane area [m2] 11004 78710 5582
Permeate pressure membrane 1 [bar] 2.5 7 7
Feed pressure membrane 2 [bar] 24 18 18
Permeate pressure membrane 2 [bar] 1 1 1
Permeate temperature 1 [◦C] 3.6 5.1 5.1
Permeate temperature 2 [◦C] 43.33 48.2 48.2
Compression energy required [hp] 20700 10100 690
Product (retentate) flow [kmol/h] 18389 17010 1190
Mole fraction
CO2 0.02 0.02 0.02
H2S – 2.1E−16 2.1E−16
H2O – 2.1E−16 2.1E−16
CH4 0.98 0.788 0.788
Ethane – 0.096 0.096
Propane – 0.056 0.056
Butane – 0.030 0.030
Pentane – 0.006 0.006
Hexane – 0.004 0.004
Permeate flow [kmol/h] 200.5 1576 111
Mole fraction
CO2 0.90 0.90 0.90
H2S – 2.0E−5 2.0E−5
H2O − 8.0E−31 8.0E−31
CH4 0.10 0.04 0.04
Ethane – 0.011 0.011
Propane – 0.014 0.014
Butane – 0.017 0.017
Pentane – 0.006 0.006
feed pressure second membrane stage [bar]
ig. 3.8. Impact of increasing feed pressure on spec. CO2 capture duty, top case 1,
iddle case 2a, bottom case 2b.
embrane stage is desired. By increasing the feed pressure of the
econd membrane stage the total membrane area of the process
ecreases. For both cases the area of the first membrane stage stays
early constant. The decrease in total required membrane area is
aused by the increase in driving force in the second membrane
tage, whose area also decreases. Fig. 3.8 illustrates the change of
pecific CO2 capture duty with increasing feed pressure in the sec-
nd membrane stage. For both cases the specific CO2 capture duty
oes through a minimum. This effect is caused by different reasons.
n the one hand more energy is needed to compress a feed stream
26] than recompressing a retentate or permeate stream. On the
ther hand the increasing driving force reduces the total size of the
ecycling stream to the first stage.
.2.6. 2-stage simulation results summary
Table 3.3 summarizes the simulation results for 2 stage cascade
onfiguration. The CO2 recovery has the same order of magnitude
s achieved in the single stage application, but the CO2 purity in
Hexane – 0.004 0.004
CO2 capture duty [kJelec/kg CO2 captured] 986 451 451
CO2 recovery [%] 33.5 81 81
CH4 loss [%] 0.11 0.47 0.47
the permeate is now increased to 90%. The feed pressure for the
second membrane stage was set to 24 and 18 bar for case 1 and
2a/b. The losses of CH4 could be reduced to 0.11% for case 1 and
0.47% for case 2a/b. Therefore a specific CO2 capture duty of ∼1000
and ∼450 kJelec./kg CO2 captured is required.
3.3. Equipment sizing and economic analysis
Sizing and costing of membrane plants, follows a similar way as
sizing and costing of the amine absorption plant. The main invest-
ment cost consists of 3 parts; costs of membranes, heat exchanger
and compressors. The cost for heat exchangers is estimated in the
same way as for the amine absorption case. For the membrane
costs a price of 50 Dollar per m2 is assumed, including the module
costs. The investment costs for the needed compressors are calcu-
lated with an expression presented by Bhide et al. [1]. All values
are updated to the year 2008 with the Chemical Engineering Cost
Index.
Table 3.4 summarizes the values used for economic analysis. The
membrane life is set to 4 years and a specific replacement cost of 25
US Dollar per square meter is assumed. Table 3.5 shows the total
investment cost and gas processing cost for the single stageand
2-stage configuration. As expected the investment cost for single
stage operation are very low in comparison to multi stage oper-
ation. Based on the high losses of CH4 for case 2 (9.6%), the gas
processing costs are higher for single stage operation than for the
2 stage cascade.
For case 1, 2-stage process yields higher gas processing costs
than the single stage process. In this case the investment cost for
compressors and costs for recompression are higher than the costs
resulted by the loss of CH4 in single stage operation. The achieved
values for GPC are lower than the values found by other researchers
[1,6,23,24]. These differences can be explained by higher feed
960 L. Peters et al. / Chemical Engineering Journal 172 (2011) 952– 960
Table 3.4
Economic parameters.
Value Unit Source
Wellhead price for natural gas 4.2 Dollar/MMBTU [28]
Cost for electricity 0.07 Dollar/kWh
Cost for cooling water 0.01 Dollar/t [27]
Cost for process water 0.6 Dollar/t [27]
Membrane cost 50 Dollar/m2
Membranes replacement cost 25 Dollar/m2
Membrane life 4 Year
Direct labor 15 Dollar/h [25 MMSCFD feed] [23]
Working time 8 h/d [23]
Overall labor cost 2.15 Direct Labor [23]
Maintenance cost 0.05 TCI [1]
Capital recovery cost 0.277 TCI [1]
Plant on stream factor 0.97
Table 3.5
TCI in million Dollars and GPC in Dollars per mscf product for membrane configuration.
Single stage process 2 stage process
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Case 1 Case 2a 
Total capital investment 0.57 2.9 
Gas processing cost 0.08 0.40 
ressure (115 and 90 bar) in our simulation and higher flux and
electivity of the PVAm/PVA membrane in comparison to other
embranes reported.
. Conclusions
During this work a technical and economic analysis for gas
weetening with amine absorption and membrane technology was
onducted. For 2 different feed gas compositions a technical anal-
sis was done with a membrane model “ChemBrane” interfaced
ithin Aspen Hysys. For both technologies, an optimization was
one to minimize total capital investment and gas processing cost.
ith both technologies it was possible to reach the given simula-
ion goals. For amine absorption 90% CO2 within the acid gas stream can be reached.
arget for the single stage membrane configuration was a content
f 2% CO2 in the product gas stream, furthermore for the 2 stage
onfiguration a purity of 90% CO2 within the permeate stream of
he second membrane stage. Moreover an economical model was
stablished to calculate the total capital investment and gas pro-
essing cost on the basis of these simulation results. The simulation
esults show that purity of the achieved gas streams is for instance
ower when using a membrane process, while the higher purity
f sold and vent gas in amine process is paid by high total capital
nvestment and a potentially more harmful environmental process.
he membrane process yields lower purity than the amine pro-
ess with respect to CO2 in the sweet gas, but meets the sales gas
tandards according to spec. (