oexmann (2008)

i n t e r n a t i o n a l j o u r n a l o f g r e e n h o u s e g a s c o n t r o l 2 ( 2 0 0 8 ) 5 3 9 – 5 5 2

Post-combustion CO2-capture from coal-fired power plants:Preliminary evaluation of an integrated chemical absorptionprocess with piperazine-promoted potassium carbonate

Jochen Oexmann a,*, Christian Hensel b, Alfons Kather a

a Institute of Energy Systems, Hamburg University of Technology, Denickestr. 15, D-21073 Hamburg, GermanybEvonik Energy Services GmbH, Rellinghauser Str. 1-11, D-45128 Essen, Germany

a r t i c l e i n f o

Article history:

Received 11 December 2007

Received in revised form

21 March 2008

Accepted 1 April 2008

Published on line 19 May 2008

Keywords:

CO2 capture

Chemical absorption

Potassium carbonate

Piperazine

ASPEN Plus1

EbsilonProfessional1

a b s t r a c t

The simulation tool ASPEN Plus1 is used to model the full CO2-capture process for chemical

absorption of CO2 by piperazine-promoted potassium carbonate (K2CO3/PZ) and the sub-

sequent CO2-compression train. Sensitivity analysis of lean loading, desorber pressure and

CO2-capture rate are performed for various solvent compositions to evaluate the optimal

process parameters. EbsilonProfessional1 is used to model a 600 MWel (gross) hard coal-

fired power plant. Numerical equations for power losses due to steam extraction for solvent

regeneration are derived from simulation runs. The results of the simulation campaigns are

used to find the process parameters that show the lowest specific power loss. Subsequently,

absorber and desorber columns are dimensioned to evaluate investment costs for these

main components of the CO2-capture process. Regeneration heat duty, net efficiency losses

and column investment costs are then compared to the reference case of CO2-capture by

monoethanolamine (MEA).

CO2-capture by piperazine-promoted potassium carbonate with subsequent CO2-com-

pression to 110 bar shows energetic advantages over the reference process which uses MEA.

Additionally, investment costs for the main components in the CO2-capture process

(absorber and desorber columns) are lower due to the enhanced reaction kinetics of the

investigated K2CO3/PZ solvent which leads to smaller component sizes.

# 2008 Elsevier Ltd. All rights reserved.

avai lab le at www.sc iencedi rec t .com

journal homepage: www.e lsev ier .com/ locate / i jggc

1. Introduction

Growing public awareness of the ongoing climate change has

led to increasing research activity in the field of greenhouse

gas (GHG) mitigating technologies. Since it is envisaged that

renewable and nuclear energy will only provide part of the

world’s energy needs in the next decades, fossil fuels will

remain a key energy source. Of all fossil fuels, coal resources

are the largest and show a wide global distribution. The

* Corresponding author. Tel.: +49 40 42878 2771; fax: +49 40 42878 284E-mail address: [email protected] (J. Oexmann).

Abbreviations: CCS, carbon capture and storage; FGD, flue gas desuMEA, monoethanolamine; PZ, piperazine; RPP-NRW, Reference Power1750-5836/$ – see front matter # 2008 Elsevier Ltd. All rights reserveddoi:10.1016/j.ijggc.2008.04.002

continuing use of coal ensures a diversification of the energy

supply and thus safeguards security of supply, especially in

countries lacking their own natural gas and oil resources.

However, coal-fired power plants show the highest specific

CO2-emissions, currently about twice as large as those of

natural gas-fired combined cycle power plants.

CO2-emissions from coal-fired power plants can be reduced

by increasing the energy conversion efficiency or by capturing

and storing the emanating CO2. The latter is commonly

1.

lphurisation; GHG, greenhouse gas; K2CO3, potassium carbonate;Plant North-Rhine-Westphalia..

mailto:[email protected]

http://dx.doi.org/10.1016/j.ijggc.2008.04.002

List of symbols

A column cross-sectional area (m2)

AP,i required packing surface for stage i (m2)

C column capacity factor

Ctot total column costs (s)

CCE column external devices costs (s)

CCJ column jacket costs (s)

CCP column packing costs (s)

deq packing specific equivalent diameter (m)

E enhancement factor

FLV flow parameter

FP packing factor (1/m)

G0 correction factor (m/s)

Ha Hatta number

Ji mass transfer flux (kg/s3)

kj reaction rate of reaction j

Kj equilibrium constant of reaction j

KG mass transfer coefficient of the gas phase (m/s)

KL physical mass transfer coefficient of the liquid

phase (m/s)

Ktot,i mass transfer coefficient (m/s)

L solubility of CO2 in PZ solution

mCO2mass flow of CO2 in the flue gas (kg/s)

mCO2 ;cap amount of CO2 captured (kg/s)

mL;max maximum mass flow of liquid phase (kg/s)

mG;max maximum mass flow of gaseous phase (kg/s)

D pCO2 ;ipartial pressure difference between gas and

liquid phase on stage i (N/m2)

pext pressure of extracted steam (bar)

ploss overall specific power loss (kWhel/kg CO2)

Pcap power duty for CO2-capture (MW)

Pcomp power duty for CO2-compression (MW)

PCW power duty for additional cooling water supply

(MW)

Pel,net net power output of original power plant with-

out CCS (MW)

Pel,CCS net power output with CO2-capture and CO2-

compression (MW)

Ploss overall power loss (MW)

Preg power loss due to steam extraction for solvent

regeneration (MW)

q specific reboiler duty (GJ/t CO2)

R universal gas constant (8.3143 J/(mol K))

S packing specific length (m)

Tcond main condenser temperature in the water

steam cycle (K)

Ti temperature on stage i (K)

Tsatext saturation temperature of extracted steam (K)

usf superficial velocity (m/s)

VG;max maximum volumetric gas flow (m3/s)

xi mole fraction of component i

Greek symbols

gi activity coefficient of component i

e CO2-capture rate

heff effective efficiency

hCarnot Carnot efficiency

hL dynamic viscosity (mPa s)

vi; j stoichiometric factor of component i in reaction

j

rG gas-phase density (kg/m3)

rL liquid-phase density (kg/m3)

FCO2 ;i CO2 mole flow on stage i (mol/s)


referred to as carbon capture and storage (CCS). CCS would

permit the continuing use of coal and other fossil fuels in

power generation while significantly reducing GHG emissions.

CCS has been discussed since the 1980s, but the lack of

economic incentives as well as political and legal uncertainties

have only allowed for a few realisations worldwide. Besides

the technical and economic aspects, safety and regulatory

issues with respect to transport and storage of CO2 remain

unclear. Until today the concept of CCS for coal-fired power

plants has not been realised on a large scale.

In post-combustion capture the CO2 is separated from the

flue gas of a conventional power plant. The CO2 content in the

flue gas of typical coal-fired power plants lies in the range of

12–15 vol% (wet) and the flue gas is present at atmospheric

pressure. There are a number of possibilities to implement

post-combustion capture. Since the partial pressure of CO2 in

the flue gases of coal-fired power plants is comparatively low,

technologies driven by high CO2 partial pressure differentials,

such as physical solvents or membranes, are not efficiently

applicable for post-combustion capture. In the relevant range

of low CO2 concentrations in power plant flue gas, only

chemical solvents show an absorption capacity large enough

to be applicable for CO2-capture. It is commonly agreed that of

all possible post-combustion capture concepts, the imple-

mentation of a chemical absorption scrubber is the most

developed and best suited for deployment in the near-term

(Aroonwilas and Veawab, 2006; Simmonds and Hurst, 2004).

CO2-capture from coal-derived flue gas by wet chemical

absorption using monoethanolamine (MEA) is considered the

most developed technology as two commercialised processes

exist: Fluor’s Econamine FGSM (Chapel and Mariz, 1999) and

Kerr McGee/ABB Lummus Global’s absorption/stripping pro-

cess (Alstom, 2001). These processes, however, show some

major disadvantages: besides corrosion issues and degrada-

tion of the solvent with respect to SOx, NOx and O2, this is

mainly the large amount of heat required for solvent

regeneration of around 4 GJ/t CO2 (Alie et al., 2005).

Potassium carbonate (K2CO3) is used for CO2 removal in

high-pressure applications such as natural gas sweetening or

the production of pure hydrogen for ammonia synthesis (Kohl,

1997). The commercial processes Benfield and Catacarb use 20

to 30 wt% aqueous K2CO3 solutions for CO2 removal. However,

these processes are limited by selectivity and slow rates of

absorption under the conditions present in coal-fired power

plants.

One approach to improve the performance of CO2-capture

with potassium carbonate is the promotion by an amine. In

this work a CO2-capture process using piperazine-promoted

potassium carbonate is analysed. Piperazine (PZ) is a diamine,

thus in contrast to common amines such as MEA, it can in

theory react with 2 mol of CO2 per mole of amine. Coupled

i n t e r n a t i o n a l j o u r n a l o f g r e e n h o u s e g a s c o n t r o l 2 ( 2 0 0 8 ) 5 3 9 – 5 5 2 541

with the potassium carbonate in solution, the blended solvent

has the potential for a higher CO2 capacity. Two amine

functional groups and the high acid dissociation constant (pKa)

also favour a higher rate of absorption (Cullinane, 2005).

Furthermore, piperazine is less sensitive to oxidative degrada-

tion than MEA (Alawode, 2005).

Fundamental research concerning the thermodynamics of

the system H2O–CO2–K2CO3–PZ has been performed in the

recent past, in particular by the group of Professor Gary

Rochelle of the University of Texas (Cullinane and Rochelle,

2004, 2005, 2006). Oyenekan (2007) developed a non-rigorous

model of the complete CO2-capture process based on

empirical expressions for CO2-solvent vapour–liquid equilibria

(VLE) including piperazine-promoted potassium carbonate.

There is a lack of process modelling and simulation in

particular to evaluate the potential of CO2-capture by K2CO3/

PZ in comparison to other solvents such as MEA and when

considering the integration of the capture and compression

sub-processes into the overall process of a coal-fired power

plant. In this work focus is therefore put onto the overall

process, integrating the three sub-processes capture, com-

pression and power plant into one comprehensive model.

Sensitivity analysis on the key process parameters such as

solvent composition, lean loading, desorber pressure and CO2-

capture rate are performed in order to determine the optimum

values for these parameters with respect to the overall specific

energy requirement. The size of absorber and desorber

columns is estimated to evaluate investment costs of these

components. Finally, a preliminary comparison of CO2-

capture by K2CO3/PZ, based on the results of this work, and

by MEA, based on information from literature, is carried out.

2. Methodology

2.1. Thermodynamics of CO2-capture by piperazine-promoted potassium carbonate

The chemical absorption of CO2 in aqueous solutions of

potassium carbonate and piperazine can be described by

considering the molecular species water (H2O), K2CO3,

potassium bicarbonate (KHCO3), PZ and carbon dioxide

(CO2), the ionic species potassium cation (K+), hydronium

cation (H3O+), hydroxide anion (OH�), bicarbonate anion

(HCO3�), carbonate anion (CO3

2�), protonated PZ (PZH+), PZ

carbamate (PZCOO�), PZ dicarbamate (PZ(COO�)2) and proto-

nated PZ carbamate (HPZCOO) in the following reaction

scheme (Hilliard, 2005):

2H2O $ H3Oþ þOH� (1)

CO2þ 2H2O $ H3Oþ þHCO3� (2)

HCO3� þH2O $ H3Oþ þCO3

2� (3)

PZHþ þH2O $ PZ þ H3Oþ (4)

PZ þ HCO3� $ PZCOO� þH2O (5)

PZCOO� þHCO3� $ PZðCOO�Þ2 þH2O (6)

HPZCOO þ PZ $ PZCOO� þPZHþ (7)

K2CO3 ! 2Kþ þCO32� (8)

KHCO3 ! Kþ þHCO3� (9)

Eqs. (1)–(7) describe the equilibrium of the system and Eqs. (8)

and (9) express the dissociation of K2CO3 and KHCO3 in water.

The equilibrium reactions can be described by the equilibrium

constant:

Kj ¼Y

i

ðxig iÞvi; j ; (10)

where xi is the mole fraction, gi the activity coefficient of

component i in the solution and vi; j the stoichiometric factor

of component i in reaction j. The component activity is

described by the electrolyte non-random two liquid (ElecNRTL)

model which was originally proposed by Chen et al. (1982) and

later extended to mixed solvent electrolyte systems by Mock

et al. (1984, 1986) and Chen and Evans (1986).

The parameters of this model for the systems H2O–K2CO3–

CO2, H2O–PZ and H2O–K2CO3–PZ–CO2 are regressed using the

Data Regression System (DRS) within the simulation tool ASPEN

Plus1, version 2006 (Aspen Plus, 2006). The data regression

was performed in accordance to the work done by Hilliard

(2005). Hilliard used vast experimental data in the form of

water vapour depression, mean ionic activity coefficient, heat

capacity, CO2-solubility, acid dissociation constant and proton

nuclear magnetic resonance (NMR) speciation to put up a

model which satisfactorily correlates the experimental data of

the mixed solvent electrolyte system over a wide range of

temperature, mixed solvent concentration and CO2-loading.

2.2. Sub-process I: CO2-capture

The simulation tool ASPEN Plus1 with the regressed para-

meter set is used to build a complete flow sheet for the CO2-

capture process as shown in Fig. 1. The boundary conditions of

the CO2-capture process are given in Table 1. The water-

saturated flue gas comes from the flue gas desulphurisation

(FGD) unit of the power plant at 47 8C. The flue gas passes a

blower, which becomes necessary to overcome the introduced

additional pressure losses of the flue gas stream in the

downstream absorber column.

The flue gas enters the absorber column at the bottom with a

temperature of 62 8C. The absorber is modelled as a multiple

stage equilibrium unit. It shows that five stages are a good

compromise between calculation precision, computational

effort and convergence behaviour of the model. As no flue gas

cooler is applied, the temperature profile in the absorber is

dominated mostly by the hot inlet temperature of the flue gas

rather than by the exothermic absorption reaction. Therefore,

the absorber temperature decreases continuously from the

bottom to the top of the column and does not show a

pronounced bulge over the absorber height. A further increase

in theoretical equilibrium stages leads to a change in the results

intermsofrich loading and specificreboiler dutyof less than5%.

The CO2-rich solvent leaves the absorber column at the

bottom and is pumped via a heat exchanger, recovering heat

Fig. 1 – Simplified flow sheet of CO2-capture by wet chemical absorption.

Table 1 – Boundary conditions for the CO2-captureprocess

Flue gas mass flow 577 kg/s

Flue gas temperature

from FGD

47 8C

Flue gas temperature

at absorber inlet

62 8C

Flue gas pressure 1.01325 bar

Flue gas CO2 concentration 14.2 vol% (wet), 15.9 vol% (dry)

Absorber solvent inlet

temperature

40 8C

Lean–rich heat exchanger

logarithmic mean

temperature difference

5 K

Reboiler temperature difference 10 K

CO2 pressure at compressor outlet 110 bar


from the lean solvent, to the top of the desorber column. The

solvent pump is necessary to overcome the geodetic height of

the desorber column as well as the pressure drop in the heat

exchanger. In the simulation, a constant logarithmic mean

temperature difference of 5 K is kept in the lean–rich heat

exchanger.

In the desorber the CO2 is released by introducing latent

heat from low-pressure (LP) steam in the reboiler. The

desorber is also modelled as a five-stage equilibrium unit.

The product stream at the top of the desorber column is cooled

to 40 8C and condensing water is used as washing water in the

washing section at the top of the absorber column. In the

washer, vaporised piperazine is recovered leading to a

piperazine slip of below 1 ppmv in the vent gas; the recovered

amine is reintroduced at the top of the absorber.

Table 2 – Solvent compositions

Solvent short name K2CO3 molality (m) PZ molality (m

S2.5,2.5 2.5 2.5

S3.2,1.6 3.2 1.6

S4.8,0.6 4.8 0.6

S3.0 3.0 0.0

The regenerated lean solvent is pumped via the lean–rich

heat exchanger to the absorber. Before it enters the absorber

column on the first stage the solvent is cooled to 40 8C.

Three different solvent compositions for piperazine-pro-

moted potassium carbonate and additionally a pure potash

solvent are analysed (Table 2).

When increasing the amount of K2CO3 one must consider

the precipitation of a solid KHCO3 as a product of CO2

absorption. The highest concentration of HCO3� at the lowest

temperatures is found at the absorber bottom. Simulation

results show a maximum HCO3� concentration of 5.36 mol/kg

H2O at 51 8C for the S4.8,0.6 solvent which is very close to the

solubility product at this temperature of approximately

5.4 mol KHCO3/kg H2O (Linke and Seidell, 1965). A further

increase of the K2CO3 concentration in the solvent is therefore

not advisable.

The S3.0 solvent is included for comparison and represents

the commercialised Benfield and Catacarb processes, which

use 20–30 wt% K2CO3 solutions (Kohl, 1997). One should keep

in mind that the Benfield as well as the Catacarb process

operate at high absorber pressure to achieve reasonable

absorption (reaction) rates and absorber columns of econom-

ical size. CO2-absorption by pure potassium carbonate cannot

be realised at atmospheric pressure due to the kinetically

inhibited absorption, which would lead to unreasonable large

components.

2.3. Sub-process II: CO2-compression

The CO2-compression train is also modelled within ASPEN

Plus1. The liquefaction is performed by five intercooled

) K2CO3 mass fraction (wt%) PZ mass fraction (wt%)

22.1 13.8

28.0 8.7

38.7 3.0

29.3 0.0


compressor stages with a water extraction after each stage,

and a subsequent booster unit in which the final pressure of

110 bar is realised. After each compression stage the CO2-rich

gas stream is cooled to 40 8C.

2.4. Sub-process III: power plant

In order to determine the total efficiency decrease when

integrating the CO2-capture and CO2-compression processes

into a power station, the model of a 600 MWel (gross) hard coal-

fired power plant within the simulation tool EbsilonProfes-

sional1, version 6.00 (Steag Ketek, 2006), was used. The model

follows a concept study which aimed at an economically and

ecologically optimised hard coal-fired steam power plant (VGB

PowerTech, 2004). The so-called Reference Power Plant North-

Rhine-Westphalia (RPP-NRW) features live steam parameters of

285 bar and 600 8C, a reheat temperature of 620 8C, 8 stages of

feedwater pre-heatingand 45 mbar condenser pressure leading

toanetpoweroutputof555.5 MWelwithanoverallnetefficiency

of45.9%(LHV). It representsastate-of-the-artpowerplantasit is

being realised today (e.g. Datteln (E.ON, 2007)). The flow rate and

CO2 concentration of the RPP-NRW flue gas are given in Table 1.

It is assumed that the CO2-capture process is integrated

into a greenfield power plant; thus power plant components

and process design are optimised to operate under the

specified boundary conditions. Therefore, all components

are operated in their individual design points.

The power loss of the power plant due to the steam

extraction for solvent regeneration can be calculated by

assuming that the extracted steam would be used in a Carnot

cycle. The power loss Preg can be described by the following

equation:

Preg ¼ heffhCarnotmCO2 eq; (11)

where the effective efficiency heff represents any additional

losses in the power cycle, mCO2is the mass flow of CO2 in the

flue gas, e is the CO2-capture rate and q is the specific reboiler

duty(e.g. inGJ/tCO2).FortheCarnotefficiencyhCarnot itholdsthat

hCarnot ¼ 1� Tcond

Tsatextð pextÞ

; (12)

Fig. 2 – Column de

where Tsatext is the saturation temperature of the extracted

steam at the extraction pressure pext and Tcond the tempera-

ture of the main condenser in the water steam cycle. Due to

the condensation of a large fraction of the LP steam in the

reboiler of the CO2-capture process, the cooling water mass

flow in the steam turbine condenser can be reduced in com-

parison to the operation of the power plant without CO2-

capture. Therefore, the auxiliary power needed for the cooling

water pumps decreases. This effect is incorporated in heff by

performing a sensitivity analysis of the RPP-NRW model in

EbsilonProfessional1, varying CO2-capture rate, specific reboi-

ler heat duty and the steam extraction parameters. The power

demand needed for additional cooling water in the capture

and compression sub-processes are considered separately

(Section 2.5).

It shows that the power loss due to reboiler steam

extraction can be adequately represented by Eq. (11) and

(12) with an effective efficiency of

heff ¼ 0:7855þ 0:01485 pext (13)

or

heff ¼ 0:6102þ 0:00165ðTsatext � 273:15Þ (14)

where pext and Tsatext have to be inserted in bar and K. The

application of Eqs. (11)–(14) yields accurateresults forPreg within

0.5% compared to the values given by the simulation model for

extraction pressures between 2.8 and 5.5 bar (131 and 155 8C).

2.5. Overall process

The integration of the CO2-capture and CO2-compression sub-

processes intoa powerplant leads toa decrease inthe net power

output. This power loss consists of four major contributors:

� T

sig

he decrease of power output due to the extraction of large

amounts of low-pressure steam from the steam cycle for

solvent regeneration (Preg).

� T
he auxiliary power demand for the lean and rich solvent
pumps and the flue gas blower in the CO2-capture sub-

process (Pcap).

n approach.

Table 3 – Optimal lean loading, rich loading and lean–rich


� T
loading difference in mol COtot
2 =ðmol K2CO3 þmol PZÞ for
he power demand for the CO2 compressors and booster
unit in the CO2-compression sub-process (Pcomp).
varying solvent compositions and CO2-capture rates � T CO2-capture rate (%) S2.5,2.5 S3.2,1.6 S4.8,0.6 S3.0
CO2-lean loading

90 1.013 1.119 1.261 1.331

70 1.054 1.183 1.370 1.435

50 1.063 1.200 1.399 1.446

CO2-rich loading

90 1.101 1.250 1.447 1.612

70 1.125 1.275 1.493 1.644

50 1.129 1.280 1.504 1.646

Lean–rich loading difference

90 0.088 0.131 0.186 0.281

70 0.071 0.092 0.123 0.209

50 0.066 0.080 0.105 0.200

he power demand for additional cooling water in both sub-

processes CO2-capture and CO2-compression (PCW).

The sum of the four contributions leads to an overall power

output Pel,CCS which is smaller than that of the original power

plant (Pel,net):

Pel;CCS ¼ Pel;net � Ploss

¼ Pel;net � ðPreg þ Pcap þ Pcomp þ PCWÞ< Pel;net (15)

Commonly, the power loss Ploss is given as a specific value with

respect to the amount of captured CO2 mCO2 ;cap (e.g. in kWhel/

kg CO2):

ploss ¼Ploss

mCO2 ;cap: (16)

2.6. Column design

Within ASPEN Plus1 both absorber and desorber are modelled

as multiple equilibrium stages. To be able to compare the

results of the column design to the reference process with

MEA, the methodology of Abu-Zahra et al. (2007a) is closely

followed and adapted to kinetic data of the K2CO3/PZ system

taken from Cullinane and Rochelle (2006). The results from the

equilibrium calculations with the simulation tool are used to

design the column diameter and the height of each theoretical

stage, taking into consideration kinetic effects (Fig. 2). For a

more detailed explanation of the column design calculations

refer to Appendix A and to Abu-Zahra et al. (2007a).

3. Results

Convergence of the simulation in ASPEN Plus1 is difficult to be

achieved with a model that features a closed solvent loop,

unless the starting values for the tear streams are already very

close to the solution. Therefore, the solvent loop in the model is

left open, defining the composition and with it the CO2-loading

of the lean solution stream at the absorber inlet before each

simulation run as an input to the model. The rich loading and

thus the loading difference between the lean and rich stream

(pick-up range) is then a result of the defined lean loading and

the temperature and pressure profile in the absorber.

The lean solvent flow is varied until the desired CO2-capture

rate is achieved. Subsequently, the reboiler duty is varied until

the material balance over the solvent loop is closed (absorber

inlet = desorber outlet). The water and piperazine balance is

closed by providing make-up streams to the lean solution.

3.1. Optimisation of the CO2-loading in lean solution

As was shown before (e.g. Freguia and Rochelle, 2003) there is a

discrete minimum in the specific reboiler duty for a certain

lean loading due to two opposing effects. If a low lean loading

is to be reached (large pick-up range) the amount of stripping

steam required to desorb the CO2 from the solution dominates

the specific heat demand (per t CO2) in the reboiler. If the

process is operated at a higher lean loading (smaller pick-up

range) a larger solvent flow is required and the sensible heat to

bring the solution from absorber to desorber temperature is

dominant in the thermal energy requirement.

During the analysis of varying CO2-capture rates, desorber

pressures and solvent compositions in this work, the optimal

lean loading is always determined in advance for each

operational point. The dependency of the optimal loading

on varying desorber pressures for one solvent shows to be

negligible. Table 3 shows the determined optimal lean loading

of the four analysed solvents for CO2-capture rates between 50

and 90%. The results compare well to the findings of Oyenekan

(2007), who set up a non-rigorous equilibrium based model for

the evaluation of a K2CO3/PZ process.

The optimal lean–rich loading differences (pick-up ranges)

turn out to be very small. One should keep in mind that

equilibrium is assumed to be reached in the absorber column,

thus the calculated rich loadings represent a theoretical

maximum for the given temperature and pressure profile in

the absorber. During operation of a real plant, equilibrium

might not be reached in the column. Since a small deviation

from the equilibrium (rich) loading leads to an even smaller

pick-up range, the total solvent circulation rate in real plant

operation might have to be increased significantly, leading to a

higher specific reboiler duty. If the solvent circulation rate

stays unchanged, the CO2-capture rate will decrease with the

lower pick-up range.

3.2. Effect of desorber pressure

The operation of the desorber below ambient pressure

(vacuum stripping) can show advantages in terms of specific

reboiler duty for solvents with low heat of desorption. As

mentioned earlier (e.g. Tobiesen and Svendsen, 2006), this

effect is due to the relative amounts of the heat of desorption

and the heat which generates water vapour in the stripper. For

low heat of desorption solvents (such as K2CO3), vacuum

desorber operation generates a smaller amount of water

vapour relative to operation at normal pressure. For high heat

of desorption solvents (such as PZ and MEA) the effect is the

contrary, thus vacuum stripping will lead to larger amounts of

Fig. 3 – Specific reboiler duty for varying desorber pressures

and solvent compositions at 90% CO2-capture.

Table 4 – Specific reboiler duty, solvent mass flow andreboiler temperature for varying solvent compositionsand CO2-capture rates

CO2-capture rate (%) S2.5,2.5 S3.2,1.6 S4.8,0.6 S3.0

3.0a 0.3a 0.3a 0.3a

Specific reboiler duty (GJ/t CO2)

90 2.44 3.07 3.00 3.16

70 2.07 2.68 2.57 2.92

50 2.05 2.65 2.52 2.92

Solvent mass flow (kg/s)

90 9581 6667 4512 4350

70 9326 7491 5412 4596

50 7163 6169 4535 3442

Reboiler temperature (8C)

90 125.4 75.6 78.9 76.5

70 117.8 71.2 73.4 73.1

50 116.1 69.5 71.4 72.6

a Desorber pressure (bar).


water vapour in the desorber, while raising the pressure leads

to a favourable increase in the ratio of CO2 and H2O partial

pressure in the gas phase ðpCO2=pH2OÞ.

This effect is especially important for the application of

piperazine-promoted potassium carbonate as a solvent, since

the solvent composition represents an additional degree of

freedom which has a direct effect on the heat of desorption.

Fig. 3 shows that those solvents with no or low piperazine

fraction (i.e. low heat of desorption solvents) show an increase

in the specific reboiler duty when the desorber pressure is

increased. This trend is reversed when considering a higher

level of piperazine promotion (S2.5,2.5, high heat of desorption

solvent). For the latter, it is therefore advisable to increase the

desorber pressure as much as possible. Additionally, with an

increase in desorber pressure the power demand for CO2-

compression is reduced (Section 3.3).

It should be noted that the heat which is transferred in the

lean–rich heat exchanger increases dramatically because of an

increasing temperature difference between absorber and

desorber when increasing the desorber pressure. This will

ultimately result in a much larger heat exchanger and a

significant increase in investment costs. Therefore, the

desorber pressure for this composition is a critical process

Fig. 4 – Specific reboiler duty, specific power loss and steam

extraction temperature for varying desorber pressures for

S3.2,1.6 at 90% CO2-capture.

parameter and should be subject to future techno-economic

optimisation. Since the lean–rich heat exchanger in this work

is not subject of detailed design and investment cost

considerations, a maximum desorber pressure of 3 bar for

the S2.5,2.5 is set for further considerations. In the case of power

plant integration with the RPP-NRW this will limit the amount

of heat transferred to around 980 MW at 90% CO2-capture rate

(in comparison to 240 MW for a desorber pressure of 0.3 bar).

Table 4 shows the specific reboiler duty and the solvent

mass flow for different solvent compositions, CO2-capture

rates and for desorber pressures of 0.3 or 3 bar at the optimal

lean-loading for each case. Additionally, the reboiler tem-

perature, which determines the quality of the needed steam

for solvent regeneration, is given in Table 4 for each case.

The specific reboiler duty for the S2.5,2.5 solvent decreases

with decreasing CO2-capture rate from 2.44 GJ/t CO2 for a CO2-

capture rate of 90% to around 2.1 GJ/t CO2 at 70 and 50% CO2-

capture rate. The higher desorber pressure of 3 bar for this

solvent leads to reboiler temperatures between 116 and 125 8C.

The S3.2,1.6 and S4.8,0.6 solvents show larger specific reboiler

duties between 3.1 GJ/t CO2 (at 90% CO2-capture rate) and

2.5 GJ/t CO2 (at 50% CO2-capture rate). The lower desorber

pressure of 0.3 bar for these solvents corresponds to lower

reboiler temperatures of between 70 and 79 8C.

The solvent circulation rate is increasing with increasing

PZ concentration, due to the fact that the higher the PZ

concentration, the smaller the difference in CO2-loading

between lean and rich solution at optimal lean-loading

(Table 3). The larger solvent mass flow will also lead to larger

column diameters or more columns for the S2.5,2.5 solvent as it

is explained further below.

The results agree with the findings of Oyenekan (2007), who

determined an increase in the specific reboiler duty of 4.8% for

the S3.2,1.6 solvent when operating the desorber at a pressure of

160 kPa instead of 30 kPa (this work: +7.4%), while the S2.5,2.5

shows a 12% lower reboiler duty at higher pressure desorber

operation (this work: �16.3%). The absolute values for the

specific reboiler duty in this work lie in a constant range

between 10.8 and 15.7% below the values given by Oyenekan

for these two solvents.

Fig. 5 – Specific power loss for varying desorber pressures

and solvent compositions at 90% CO2-capture.


3.3. Power plant integration

When integrating a CO2-capture process into a power plant

process, heat for solvent regeneration can be taken from the LP

part of the power plant turbine. The desorber temperature is

directly connected to the operation pressure of the column.

The higher the reboiler temperature, the higher the quality of

the steam that is needed for the regeneration of the solvent

and the larger are the power losses in the power plant. Fig. 4

shows this interrelation for the S3.2,1.6 solvent, when a

temperature difference in the reboiler between the conden-

sing low-pressure steam from the power plant and the CO2-

loaded solvent of 10 K is assumed.

The observed minimum in the specific power loss for S3.2,1.6

at 0.5 bar is due to three opposing effects. First, with an

Fig. 6 – Absorber diameter and height for varying

increase in desorber pressure the power demand for the CO2-

compression train decreases as the compressor inlet pressure

is equal to the desorber outlet pressure. Second, the specific

reboiler duty increases with a rise in desorber pressure. Third,

the desorber temperature also increases with higher pressure,

and therefore a higher quality steam extraction is needed for

solvent regeneration, leading to larger losses in the power

plant sub-process.

For the S4.8,0.6 and the pure potassium carbonate solvent

S3.0, the increase in reboiler duty (Fig. 3) together with the need

for higher quality steam overcompensates the decreasing

power demand for the compression sub-process, as shown in

Fig. 5. The S2.5,2.5 solvent shows the lowest specific energy

requirement. Similar to the reboiler duty, the specific power

loss for this solvent decreases with increasing desorber

pressure. However, in comparison to the specific reboiler

duty the reduction is lower and shows an asymptotic

behaviour. A further increase of the desorber pressure above

3 bar does not result in significant energy savings.

When using the S2.5,2.5 solvent, the specific power loss

decreases by 8.9% when raising the desorber pressure from 30

to 160 kPa, where Oyenekan (2007) found a decrease of 10.6%.

The absolute values for the specific power loss of this work are

approximately 10% higher than those determined by Oyene-

kan. This can be attributed to the larger power demand for

CO2-compression and the larger power loss in the power plant

which was determined by integrating the heat demand of the

reboiler into the power plant model.

3.4. Column sizing and investment costs

Within the simulation tool, chemical and phase equilibrium is

assumed for all theoretical stages in the absorber as well as in

the desorber column. As was explained earlier, the results of

CO2-capture rates and solvent compositions.

Fig. 7 – Desorber diameter and height for varying CO2-capture rates and solvent compositions.


the equilibrium considerations are used to evaluate the size

and the cost of the columns, taking into account kinetic effects

which determine the required column design (Appendix A for

details). Correlations for the structured packing MELLAPAK

125Y are used for the column sizing. For the column

diameters, a maximum size of 12.8 m is set. This represents

an upper limit due to construction and transportation

limitations. For larger CO2-capture plants in commercial

Fig. 8 – Relative column investment costs for varyin

operation, the use of greater diameter columns may be

economic and columns with a rectangular cross-section

might become favourable (Chapel and Mariz, 1999).

With a decrease in piperazine concentration the solvent

mass flow can be reduced (Table 4). As can be seen from

Eqs. (A.1)–(A.4), the reduction of L/G result in slightly narrower

absorber columns (Fig. 6). It also shows that the diameter is

rather independent of the CO2-capture rate, since flue gas and

g CO2-capture rates and solvent compositions.


solvent flow stay almost unchanged. For an application to the

RPP-NRW with a flue gas flow of 577 kg/s and a CO2-

concentration of 15.9 vol% (dry), three absorber columns are

necessary for each case if the limitation of the absorber

diameter at 12.8 m is considered.

Absorber height is increasing with decreasing piperazine

concentration in the solvent due to reduced reaction kinetics.

The absorber height doubles when decreasing the PZ

concentration from 2.5 to 0.6 m (S2.5,2.5! S4.8,0.6), since higher

packing sections are needed to reach equilibrium in each

theoretical stage.

The absorber height can also be reduced when reducing the

CO2-capture rate. This can be attributed to a changed CO2-

concentration profile over the column height. The driving

force of mass transfer, i.e. the partial pressure difference, in

particular in the lower stages is enhanced, reducing the

required surface area of the packing and ultimately decreasing

the absorber height.

The desorber diameter and height is calculated following

the same methodology as for the absorber column. The

S3.2,1.6 and S4.8,0.6 solvents require only one or two desorber

columns to follow the limitation of the maximum column

diameter of 12.8 m (Fig. 7). The use of the S2.5,2.5 solvent

demands three or four desorber columns due to the larger

required solvent flow.

Fig. 8 shows relative investment costs for the columns,

where the total investment costs for absorber and desorber

columns at a CO2-capture rate of 90% using the S2.5,2.5 solvent

represent 100%. Even though three or four desorber columns

are needed when using the S2.5,2.5, total investment costs are

lower as compared to the other two solvents. This reflects the

lower required column heights due to the enhanced reaction

kinetics and therefore lower investment costs especially for

the required packing (Eq. (A.10)). Independent of the CO2-

capture rate, investment costs are about twice as high when

using S4.8,0.6 instead of S2.5,2.5.

3.5. Comparison to MEA process

Chemical absorption processes using MEA are considered the

reference in post-combustion CO2-capture from flue gas, as

Table 5 – Boundary conditions for MEA and K2CO3/PZ CO2-cap

Abu-Za

ME

CO2-capture rate (%)

CO2-outlet pressure (bar)

Desorber pressure (bar)

Absorber pressure (bar)

Flue gas mass flow (kg/s)

Flue gas temperature (8C)

Flue gas CO2 concentration (vol% (wet))

Lean solvent temperature (8C)

Specific solvent flow (m3/t CO2)

Specific cooling water flow (m3/t CO2)

Lean loading (mol CO2/mol solvent)

Rich loading (mol CO2/mol solvent)

a Solvent.

commercialised processes exist. However, the technical and

commercial feasibility of the MEA process is also yet to be

demonstrated for large-scale coal-fired power plants.

Abu-Zahra et al. (2007a) have analysed a CO2-capture

process using MEA with a similar methodology as the one that

has been applied in this work. They have also used ASPEN

Plus1 as a simulation tool and have followed the same

strategy for column sizing and design. In the following, the

results of this work are compared to the results of Abu-Zahra

et al. (2007a,b).

Table 5 shows the boundary conditions for the two

compared processes. The S2.5,2.5 solvent at desorber pressure

of 3 bar is used for the comparison as it showed the best

performance in the preceding analysis.

The specific reboiler duty is taken directly from the

reference. The specific power loss is recalculated considering

the given values for volumetric flue gas flow, solvent

circulation rate and cooling water demand. Energy demand

for CO2-compression is recalculated with the same model that

is used for the K2CO3/PZ process, taking into account the

changed inlet pressure in the MEA case. To determine an

overall efficiency decrease when these CO2-capture processes

are used for a large-scale power plant, the results are applied

to the RPP-NRW. Absorber and desorber column height and

diameter in the MEA case are estimates which lead to the

values for investment costs that are given in Abu-Zahra et al.

(2007b).

Table 6 shows the results of the comparison. It should be

noted, that the value for ‘‘power demand capture’’ includes

the power which is needed for the supply of additional cooling

water to the capture and compression sub-processes.

Power demand for the solvent pumps in the CO2-capture

sub-process is higher for the K2CO3/PZ process, since almost

three times as much solvent is being circulated. The largest

fraction of auxiliary power demand, however, can be

attributed to the additional flue gas blower. The blower shows

a similar power demand for both cases as the volumetric flue

gas flow is similar and as comparable pressure drops for the

absorber columns are assumed. As an equilibrium model for

the absorber is applied in the simulation, the corresponding

pressure drop of the absorber column is not a result of the

ture process comparison

hra et al. MEA This work K2CO3/PZ

A 30 wt%a K2CO3/PZ 22.1/13.8 wt%a

90 90

110 110

2.1 3

1.1 1.1

616 577

48 47

13.3 14.2

30 40

27.8 74.4

103 82.1

0.32 1.013

0.49 1.101

Table 6 – Results for MEA and K2CO3/PZ CO2-capture process

Abu-Zahra et al. MEA This work K2CO3/PZ

CO2 captured (kg/s) 112.5 110.4

Specific reboiler heat duty (GJ/t CO2) 3.29 2.44

Specific power loss (kWh/kg CO2) 0.342 0.288

Power loss for solvent regeneration 0.230 0.170

Power demand capture 0.033 0.047

Power demand compression 0.079 0.071

Power plant net efficiency (% LHV) 34.6 36.4

Efficiency decrease (%pts.) 11.3 9.5

Number of absorbers 2 3

Absorber height (m) 29a 12.0

Absorber diameter (m) 11a 12.7

Number of desorbers 1 2

Desorber height (m) 15a 6.9

Desorber diameter (m) 10a 11.3

Column investment costs (Ms 2007) 10.9 8.84

Specific column investment costs (s/(t CO2/h)) 352 288.3

a Estimated.


simulation and therefore has to be fixed in advance. However,

as it is shown below, the K2CO3/PZ process actually has lower

absorber columns than the MEA process, which is equivalent

to less required packing and thus a lower pressure drop. This

would lead to a smaller energy demand in the flue gas blower

and ultimately means an even more advantageous perfor-

mance of the K2CO3/PZ process when compared to the MEA

process.

Specific energy requirement for CO2-compression is

higher in the MEA case, since the desorber is operated at a

lower pressure (limited due to carbamate polymerisation of

MEA above �125 8C). Piperazine is less sensitive towards

thermal degradation and shows a lower vapour pressure than

MEA. The desorber can therefore be operated at higher

pressures and temperatures. In this case it was set to 3 bar to

limit the amount of transferred heat in the rich–lean heat

exchanger and with it the investment costs for this

component. A complete techno-economic evaluation of the

process has to take into account the energetic advantage of a

further increase in desorber pressure, the increase in specific

costs and decrease in size for elevated pressure vessels and

the increasing size and investment costs of the heat

exchanger.

More absorber and desorber columns are needed for the

K2CO3/PZ process than for the MEA case due to the much

higher solvent circulation rate (Table 4). However, both

absorbers and desorbers are lower in height due to the faster

reaction kinetics for piperazine-promoted K2CO3 in compar-

ison to CO2 absorption by MEA. Cullinane and Rochelle (2004)

state that a 1.9 m (20 wt%) K2CO3 solution promoted by 0.6 m

PZ has a comparable absorption rate as a 24 wt% MEA solution.

Fig. 6 shows that the S4.8,0.6 has indeed a similar absorber

height of approximately 25 m as was determined by Abu-

Zahra et al. for the MEA case (approximately 29 m). An

increase in PZ concentration leads to lower column heights,

since ‘‘the apparent rate constant of 1 M PZ is a factor of 20

greater than MEA’’, where ‘‘the rapid reaction of piperazine

with CO2 can be attributed to its unique, cyclic diamine

structure’’ (Cullinane and Rochelle, 2006). This ultimately

leads to 18% lower specific column investment costs for the

K2CO3/PZ process using the S2.5,2.5 solvent in comparison to the

MEA process.

4. Discussion and conclusion

A post-combustion CO2-capture process using a solution of

aqueous K2CO3 promoted by PZ was analysed with the

simulation tool ASPEN Plus1. The most critical parameter

which needs to be chosen carefully is the loading of the CO2-

lean solvent, as the specific reboiler duty for solvent

regeneration shows a discrete minimum when varying this

parameter. The impact of solvent composition, desorber

pressure and CO2-capture rate on the specific reboiler duty

as well as the specific power loss in the overall power plant

process were discussed.

A mixture of 2.5 m K2CO3 and 2.5 m PZ (S2.5,2.5) at a

desorber pressure of 3 bar shows energetic and economic

advantages over the other analysed solvent compositions.

Both specific reboiler duty and specific power duty of the

integrated overall process show a decrease with increasing

desorber pressure when using S2.5,2.5. Keeping in mind the

increase in transferred heat in the rich–lean heat exchanger

and the associated increase in investment costs, the

optimisation potential that lies in this process parameter

is limited.

The optimal solvent composition cannot be determined

unless a complete economic evaluation including operational

costs is performed. Piperazine has about five times the costs of

MEA, thus piperazine concentration, degradation and solvent

make-up rate will play a major role in the overall process

economics (Alawode, 2005).

It was shown that the use of S2.5,2.5 may significantly reduce

the required heat duty for solvent regeneration in a CO2-

capture process in comparison to MEA. It reaches values as

low as 2.4 GJ/t CO2 with a CO2-capture rate of 90% (compared to

i n t e r n a t i o n a l j o u r n a l o f g r e e n h o u550

3.3 GJ/t CO2 for MEA) and 2.1 GJ/t CO2 when reducing the CO2-

capture rate to 70% or 50%. It may also reduce the overall

energy requirement when considering the integration of a

CO2-capture and CO2-compression process into a coal-fired

power plant to 0.288 kWhel/kg CO2 compared to 0.342 kWhel/

kg CO2 for MEA.

In a preliminary comparison it was shown that the K2CO3/

PZ process could reduce the specific investment costs for

absorber and desorber columns by around 18% in comparison

to a CO2-capture process using MEA.

The K2CO3/PZ process has to be further analysed. In

particular the impact of key process parameters such as

CO2-capture rate, solvent loading, desorber pressure and

solvent composition on cost of electricity and ultimately on

the CO2-avoidance cost (i.e. s/t CO2 captured) has to be

investigated. In such a techno-economic analysis, capital

costs (CAPEX) including the investment costs for all major

components, installation, engineering and construction,

but also operating costs (OPEX), for example for solvent

make-up, maintenance, taxes and insurance, have to be

considered.

Acknowledgements

The authors would like to thank Marcus Hilliard of the

University of Texas in Austin, USA for providing data and

support on the ASPEN Plus1 Data Regression for the H2O–CO2–

K2CO3–PZ system.

Appendix A

A.1. Column diameter

The calculation of column diameter is based on the work by

Kister (1992). The flood-point criterion in combination with the

Sherwood–Eckert generalised pressure drop correlation is

applied. The column diameter is sized according to the

theoretical stage which shows the largest gas and liquid

phase mass flows.

Following the flood-point criterion the column is designed

for 75% of the velocity at the flood-point. Additionally an

application related security factor has to be considered. For

absorption and desorption processes this factor is usually 0.85

(Kister, 1992).

The flow parameter FLV which is calculated from the

maximum mass flows of the liquid and gaseous phase is used

to determine the capacity factor C with correlations for the

structured packing MELLAPAK 125Y by Sulzer. Together with

the density of the liquid and the gaseous phase and the

dynamic viscosity from simulation results, the capacity factor

C and a packing factor FP are used to calculate the superficial

velocity usf. With the maximum volumetric gas flow VG;max one

can then calculate the cross-sectional area A and the diameter

of the column.

FLV ¼mL;max

mG;max

ffiffiffiffiffiffirG

rL

r; (A.1)

G0 ¼ 0:75� 0:85CðFLVÞ

F0:5P

rL

hL

� �0:05

; (A.2)

s e g a s c o n t r o l 2 ( 2 0 0 8 ) 5 3 9 – 5 5 2

usf ¼ G0ffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi

rG

rL � rG

r; (A.3)

A ¼ VG;max

usf: (A.4)

A.2. Column height

The calculation of the column height is based on the film

theory, in which a bulk liquid and gas phase and an interface

between the two phases are distinguished. The column height

of each theoretical stage is mainly determined by mass

transfer between the gas and the liquid phase, where the main

driving force is the difference in the partial pressure of CO2 in

the gas and in the liquid phase.

The required packing surface AP,i and the total packing

volume can be calculated from

AP;i ¼FCO2 ;iRTi

Ji; (A.5)

where FCO2 ;i is the CO2 mole flow on stage i going from the gas

into the liquid phase (absorption) or vice versa (desorption),

and Ji is the mass transfer flux which is determined from

Ji ¼ D pCO2 ;iKtot;i; (A.6)

where D pCO2 ;iis the change in the logarithmic CO2 partial

pressure difference between gas and liquid phase from stage

i to stage i + 1. The values for D pCO2 ;iare taken from the

simulation results.

Ktot,i is the mass transfer coefficient which comprises the

mass transfer coefficient of the gas phase KG and of the liquid

phase taking into account the chemical reactions.

Ktot;i ¼1

ð1=KGÞ þ ð1=ðL E KLÞÞ: (A.7)

KG depends on the diffusion of CO2 in the gas phase, a packing

specific equivalent diameter deq (MELLAPAK 125Y:

deq = 0.018 m) and the Sherwood number Sh which represents

the ratio of the characteristic length of the system to the

diffusive boundary layer thickness.

L is the temperature-dependent solubility of CO2 in

aqueous PZ solutions. Following the Bravo Fair’s correlation

(Kister, 1992), the physical mass transfer coefficient KL is a

function of the effective velocity of the liquid, the diffusion of

CO2 in the liquid phase and a packing specific length S

(MELLAPAK 125Y: S = 0.017 m).

Following Westerterp et al. (1993) the enhancement factor E

can be determined iteratively taking into account the

dimensionless Hatta number Ha, which compares the rate

of absorption of a solute in a reactive system to the rate of

absorption of the same solute in the case of physical


absorption. As the carbamate reactions are dominant for the

determination of the overall absorption rate (Cullinane and

Rochelle, 2006), only the following reactions are considered:

PZþ CO2 þ

OH� @kPZ�OH�

PZCOO� þH2O

H2O @

kPZ�H2O

PZCOO� þH3Oþ

PZ @kPZ�PZ

PZCOO� þ PZHþ

CO32�

@

kPZ�CO3

2�

PZCOO� þHCO3�

PZCOO� @kPZ�PZCOO�

PZCOO� þHPZCOO

26666666664

37777777775

(A.8)

PZCOO� þ CO2

þ

H2O @

kPZCOO��H2O

PZðCOO�Þ2 þH3Oþ

PZ @kPZCOO��PZHþ

PZðCOO�Þ2 þ PZHþ

CO32�

@

kPZCOO��CO3

2�

PZðCOO�Þ2 þHCO3�

PZCOO� @kPZCOO��PZCOO�

PZðCOO�Þ2 þHPZCOO

266666664

377777775

(A.9)

With the rates of reaction kj taken from Cullinane and Rochelle

(2006), the Hatta number and with it the enhancement factor E

can be calculated.

Finally, with the column diameter the packing volume for

each equilibrium stage i is determined fromAP,i. A safety factor

of 25% and additional spacing for any additional equipment

such as distributors is added to determine the total packing

volume which is needed to reach equilibrium conditions on

each stage.

A.3. Investment costs

Total cost for absorber and desorber columns (Ctot) can be

divided into three groups:

Ctot ¼ CCJ þ CCP þ CCE; (A.10)

where CCJ is the cost for the column jacket, CCP the investment

cost for the packing and CCE is additional cost for any external

devices such as ladders and platforms. Costs are determined

following the strategy of Vatavuk and Neveril (1982) and scaled

to s in 2007 by taking into account the change in the M&S

index and s–US$ exchange rate.

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