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International Journal of Greenhouse Gas Control 7 (2012) 230–239

Contents lists available at SciVerse ScienceDirect

International Journal of Greenhouse Gas Control

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

nvestigation of the dynamic behavior of different stripper configurations forost-combustion CO2 capture

ehdi Karimi ∗, Magne Hillestad, Hallvard F. Svendsenepartment of Chemical Engineering, NTNU, 7491 Trondheim, Norway

r t i c l e i n f o

rticle history:eceived 20 July 2011eceived in revised form 12 October 2011ccepted 28 October 2011vailable online 30 November 2011

eywords:ynamic simulationtripper different configurationsEA

a b s t r a c t

The dynamic behavior of three process configurations proposed for CO2 capture from flue gas is inves-tigated. In the previous paper (Karimi et al., 2011) the different configurations were investigated withrespect to energy requirement and capital costs and the best configurations were defined based on CO2

avoided cost and total capture cost. In addition to the economy, the dynamic behavior is importantwhen operation of the plant is considered. In this study the transient behavior of the vapor recompres-sion and the split-stream configurations are investigated when different disturbances happen, and theresults are compared with the conventional configuration as a benchmark. All process configurations areoperationally feasible with the proposed control configuration. The results show that the conventionalconfiguration is the most inherently resilient to disturbances, and that the vapor recompression config-

uration can handle disturbances better than split-stream configuration. The same relative reduction inreboiler duty has more negative effect (more reduction in capture ratio) for split-stream configurationthan two other configurations.

The control structure generated is able to control the plant for all configurations except the situationwhere the reboiler duty is an active constraint. In this situation the energy loss will increase and it isbetter to find another control variable to pair with reboiler duty. This will be done in future studies.

. Introduction

Fossil fuels like oil, coal and natural gas will have a major sharen energy sources for the future. Producing energy from fossil fuels

ill create large amounts of CO2 that is a main contributor tolobal warming. Improvement of energy efficiencies and a tran-ition to renewable energy sources will reduce CO2 emissions, butuch measures can reduce emissions significantly only in the longerm. In a shorter time frame, CO2 capture and storage (CCS) is nec-ssary to reduce global warming. Among different technologies,ost-combustion aqueous amine absorption/stripping is one of theost promising technologies and 30 wt% MEA is the most applied

olvent for post-combustion CO2 capture. However, high energyequirement is the main challenge for this technology. Alternativerocess configurations have been proposed to reduce operatingosts and in some cases also compared with capital cost of the

O2 capture process, as shown by Karimi et al. (2011), Schach et al.2010), Oyenekan and Rochelle (2006, 2007), Jassim and Rochelle2006), Leites et al. (2003), and Goff et al. (1996). These studies are

∗ Corresponding author. Tel.: +47 73 59 40 28; fax: +47 73 59 40 80.E-mail addresses: mehdi.karimi@chemeng.ntnu.no

M. Karimi), magne.hillestad@chemeng.ntnu.no (M. Hillestad),allvard.svendsen@chemeng.ntnu.no (H.F. Svendsen).

750-5836/$ – see front matter © 2011 Elsevier Ltd. All rights reserved.oi:10.1016/j.ijggc.2011.10.008

© 2011 Elsevier Ltd. All rights reserved.

all steady-state investigations and dynamic behavior of the plantsis not considered. In addition to affecting the economy, dynamicbehavior is important for any plant. Because of varying electricitydemand, a power plant may be operated in full-load in the peakhours and part-load in the off-peak hours. In addition to varyingelectricity demand, there are other factors like ambient conditionsand fuel composition that will change the flue gas flow rate andcomposition. Consequently dynamic simulations of the CO2 cap-ture plant may be a useful tool for investigating plant behavior intransient periods.

There are some studies of CO2 capture plants that includedynamic behavior of the plant in presence of disturbances. Linet al. (2011) introduce plantwide control for the capture plant. Theyfound that lean solvent flow rate and lean loading are the mostimportant parameters that influence the CO2 removal rate. Theyalso mention that the water balance is important factor to achievestable long-term operation. They tested their control structure byinvestigating the dynamic response of different variables in pres-ence of disturbances, such as flue gas flow and composition andCO2 removal rate.

Panahi et al. (2010) used the Unisim Design software to find the

best control structure based on minimizing the energy loss withconstant setpoint for control variables when different disturbanceshappen. Lawal et al. (2010, 2009), used a rate-based dynamic modelin gPROMs. They found that the performance of the absorber is

M. Karimi et al. / International Journal of Gree

Nomenclature

CAM amine concentration (mol amine/m3)Cp solution heat capacity (kJ/kg)CV control variableDOF degree of freedomk steady-state gaink′ slope after responseKC proportional gain, tuning parameterMV manipulated variableN numberP proportionalP* equilibrium vapor pressure (kPa)Psat saturation pressure (kPa)PI proportional–integralQdes desorption heat (kJ/mol CO2)QR reboiler duty (kW)Qsens sensible heat (kJ/mol CO2)Qstrips stripping heat (kJ/mol CO2)t time (minute)TC cold temperature, condenser temperature of steam

turbine (K)TH hot temperature, extracted steam temperature for

solvent regeneration (K)TR stripper reboiler temperature (K)u manipulated valueWeq equivalent work (kW)WElect electricity consumption (kW)x mole fractiony control value

Greek letters˛Lean CO2 lean loading (mol CO2/mol amine)˛Rich CO2 lean loading (mol CO2/mol amine)� turbine efficiency for produce electricity from steam� delay, time where output does not change (minute)�1 process time constant, additional time to reach 63%

of final change (minute)�C desired closed-loop response time (minute)�I integral time, tuning parameter (minute)� density (kg/m3)� difference�Habs CO2

CO2 absorption reaction heat (kJ/mol CO2)�HVap water latent heat (kJ/mol H2O)

ml

idacosh

3. Process description

TS

ore sensitive to the L/G ratio than to the actual flow rates of theiquid solvent and flue gas.

Kvamsdal et al. (2009) have modeled a stand-alone absorbern gPROMs. They investigated startup and load reduction con-itions. The results of the constant lean solvent flow rate casend the simultaneous reduction solvent flow rate case wereompared. Ziaii et al. (2009) used Aspen Custom Modeler for devel-ping a rate-based dynamic model. They proposed two control

trategies and compared the results when different disturbancesappen.

able 1imulation results compare to pilot plant data.

Item Pilot plant data

CO2 capture ratio (%) 72.0Rich loading 0.491Reboiler duty (kW) 210

nhouse Gas Control 7 (2012) 230–239 231

These studies usually include a conventional capture plant orparts of it. As mentioned above, different configurations are pro-posed to reduce capture cost in the steady state situation butdynamic behavior, controllability and stability of these configura-tions are not investigated. These configurations usually are morecomplex and more difficult to control and also have less stabilitywhen disturbances happen. In previous work, Karimi et al. (2011), atechno-economical investigation was done for five different strip-per configurations and the best configurations were defined basedon CO2 avoided cost and total capture cost. Among these con-figurations it was found that the vapor recompression and thesplit-stream configurations had the lowest CO2 avoided cost. Inthis study these two configurations are investigated in a dynamicsituation to investigate the dynamic behavior of the plant whendifferent disturbances happen. The results are compared to theconventional configuration as a benchmark. The steady state is sim-ulated for 90% CO2 capture with aqueous 30 wt% MEA from the fluegas of a 150 MW supercritical coal power plant. Simulations areperformed in Unisim Design with the Amine Fluid package. Thenon-equilibrium stage steady state and dynamic model uses Kent-Eisenberg thermodynamic equations and modified Murphree masstransfer efficiencies are used for simulation. The modified Mur-phree efficiency is a function of the kinetic rate constants for thereactions between acid gas and the amine, the physico-chemicalproperties of the amine solution, the pressure, temperature andthe geometry of the column (Honeywell, 2010).

2. Model validation

In the techno-economical study, Karimi et al. (2011), we men-tioned that Unisim result have a good agreement base on pilot plantdata from Luo et al. (2009). Here again we validate the simulationresult against the pilot plant data from University of Texas at Austin(Dugas, 2006).

The pilot plant absorber and stripper diameter is 0.427 m and thetotal height of packing is 6.1 m for both columns. 48 different setsof data are available in this reference. Here the simulation resultsare validated with the case 43 pilot plant data. The simulation andexperimental data comparisons are shown in Fig. 1.

The input of the simulation for absorber is the column geometry,lean amine and flue gas streams properties (flow rate, temperature,pressure and composition). Then we compare the temperature pro-file, CO2 capture rate and rich loading. For the stripper the richamine flow properties, the geometry of the column and the load-ing of lean amine are the inputs. The output of the simulation isthe temperature profile and reboiler duty. In Unisim Design wecan use two different thermodynamic models; Kent-Eisenberg andLi-Mather. The simulation results for both models are compared tothe experimental data. The simulation result compare to pilot plantdata are shown in Table 1.

From Fig. 1 and Table 1 show that the simulation results fromUnisim with the Kent-Eisenberg thermodynamic model have agood agreement to pilot data.

This study is done for CO2 capture of a 150 MW supercritical coalpower plant. The flue gas specifications are given in Table 2.

Simulation Kent-Eisenberg Simulation Li-Mather

73.2 71.70.492 0.487202 201

232 M. Karimi et al. / International Journal of Greenhouse Gas Control 7 (2012) 230–239

ratur

ceh

3

oit

ia

3

st

oF

3

dr

4

osa

TF

where TH = TR + 10 [K] and TC = 313 K and � = 75%.The equality constraints are:

1. The capture rate of CO2 is 90%.

Fig. 1. Simulation validation for tempe

Dynamic simulations just include the capture section, while theompression section is not simulated. In the previous work (Karimit al., 2011), different configurations were explained in detail andere we introduce them shortly.

.1. Conventional process configuration

The conventional configuration is the simplest and the easiestne to control. In the other configurations there are some changesn order to reduce the energy requirement. However, at the sameime the process complexity will increase.

In Fig. 2, a flow diagram for a conventional process structures given, including absorber, stripper, cross heat exchanger, coolernd pumps.

.2. Split-stream configuration

In the split-stream configuration the rich amine is split into twotreams, going to two sections of the stripper after preheating withwo separate lean amine streams as shown in Fig. 3.

In the steady state study (Karimi et al., 2011) we found that it isptimal to send 42% of the rich amine to the top section of stripper.or dynamic simulation, we fix the split ratio at this value.

.3. Vapor recompression configuration

In the vapor recompression process configuration, a pressurerop is created in the lean amine stream and the resulting vapor isecompressed and sent to the stripper as shown in Fig. 4.

. Self-optimizing control

In the work of Panahi et al. (2010), we applied the self-ptimizing method for a CO2 capture plant. The aim ofelf-optimizing control is to identify the best controlled vari-bles (CVs) that can be kept at constant set points in the presence

able 2lue gas specification.

Temperature (◦C) 48Flow rate (kmol/h) 24 123Pressure (bar) 1.1Composition (mol fraction)

CO2 0.1186Nitrogen 0.7291Oxygen 0.0505H2O 0.1018

e profile, (a) absorber and (b) stripper.

of different disturbances. With self-optimizing control variableswe do not need to re-optimize the plant after a disturbance hashappened. For self-optimizing control, the stepwise procedureof Skogestad (2004) is used. This method is explained for theconventional configuration in the following and for the otherconfigurations it is similar.

4.1. Step 1: define objective function and constraints

The total equivalent work is considered as the objective func-tion that should be minimized. We assume that the temperatureof steam in the reboiler (TH) is 10 ◦C higher than the reboiler tem-perature and that steam condenses at 40 ◦C in the turbine (TC). Thetotal equivalent work for the plant is then:

Weq = QR

(1 − TC

TH

)× � + WElect (1)

Fig. 2. Conventional process configuration.

M. Karimi et al. / International Journal of Greenhouse Gas Control 7 (2012) 230–239 233

Rich-LeanHeatexchanger

Cooler

Flue Gas

Rich Amine

CO2

LeanAmine

Clean Gas

StripperAbsorber

RichAmine

LeanAmine

Cooler

ream

23

Fig. 3. Split-st

◦

. The temperature of lean amine to the absorber is 45 C.. Because of degradation problems at high temperature, the strip-

per top pressure is kept at 1.8 bar to avoid too high temperaturein the reboiler.

Stripper

Rich-LeanHeatexchanger

Rich Amine

CO2

LeanAmine

V-11 atm

2 atm

Fig. 4. Vapor recompression configuration.

configuration.

4. The stripper condenser temperature is assumed to be 30 ◦C.5. The amine concentration of lean amine is 30 wt%.

4.2. Step 2: determine the degrees of freedom for optimization

There are 10 degrees of freedom (DOF) in the plant. In the previ-ous work, Panahi et al. (2010), there were 9 valves because amineconcentration was not controlled. Fig. 5 shows the conventionalconfiguration with 10 dynamic DOFs.

However, 4 degrees of freedom are used to control the levels (2in the stripper, 1 in the absorber, 1 in the make-up tank) and thenumber of degrees of freedom (DOFs) for steady-state optimizationis 6.

4.3. Step 3: identification of important disturbances

The flue gas flow rate and composition are considered to be themain disturbances. From the active constraint the stripper pressurehas a significant effect on the objective function and we considerit as a third disturbance. We consider ±5% in flue gas flow rate,composition and stripper pressure as the disturbances.

4.4. Step 4: optimization with and without disturbances

There are 5 equality constraints and 5 DOFs of 6 steady stateDOFs needed for controlling them in the constraint value. Thereforeone DOF is left for optimization.

Nopt.free = 6 − 5 = 1 (2)

The objective function, Weq, should be minimized subject to the

5 equality constraints in Section 3.1 and the following inequalityconstraint:

0.005 < CO2 mole fraction at the bottom of stripper < 0.05.

234 M. Karimi et al. / International Journal of Greenhouse Gas Control 7 (2012) 230–239

Absorber

Desrober

Pump 1

Reboiler

V-6

Rich/Lean Exchanger

Surge TankPump 2

V-9

Condenser

Cooler

V-5

V-2

V-4

V-10 V-3

Steam

CoolingWater in

AmineMakeup

WaterMake up

Flue Gas fromPower Plant

V-8

To Stack

CO2

Condensate

CoolingWater out

CoolingWater out

CoolingWater in V-7

V-1

22950 kmol/hr110 kPa48 oCCO2:12.4%N2 :76.1%O2 : 5.3%

21600 kmol/hr100 kPa50 oCCO2: 1.3%N2 :80.8%O2 : 5.6%H2O : 12.3%

2614 kmol/hr180 kPa30 oCCO2:97.7%H2O: 2.3%

119 oC190 kPa

79723 kmol/hr101 kPa45 oCCO2: 2.30%H2O: 86.78%MEA: 10.92%

4019 kmol/hr150 kPa25 oC

25 kmol/hr150 kPa25 oC

163 ton/hr300 kPa133.6 oC

280 kPao

2E5 kmol/hr20 oC

30 oC

1.3E4 kmol/hr20 oC

30 oC 326 kW

212 kW

dynam

sl

4

ta

12

4

vP

T

tad

5

ef

In this study we have tried to limit the complexity and weuse simple proportional (P) controllers for level and propor-tional–integral (PI) controllers for other parameters. The tuning

Table 3Scaled gain for different candidate CVs.

Candidate CV Conventional Split-stream Vapor recompression

Point 1 (Top) 45 695 8Point 2 27 724 40Point 3 18 847 171Point 4 91 1021 528Point 5 204 1134 1716Point 6 366 1024 1746Point 7 592 679 1952Point 8 880 230 2532Point 9 1211 256 3633

Point 10(Middle) 1531 541 5230Point 11 1760 1169 6763Point 12 1789 1815 3968Point 13 1620 1993 2160Point 14 1295 2119 1275Point 15 953 1569 802Point 16 676 978 529Point 17 473 618 358Point 18 329 401 244

H2O: 6.3%

Fig. 5. Process with 10

The optimum value for the CO2 mole fraction at the bottom oftripper is found to be 0.024 which is equivalent to about 0.20 inean loading.

.5. Step 5: identification of candidate control variables

We select reboiler duty for the remaining DOF and want to findhe best control variable to pair it with. The candidate control vari-bles are

. Temperature in different points in the stripper.

. CO2 composition at the bottom of the stripper.

.6. Step 6: evaluation of loss

For selecting the best control variable from the candidateariable we use the maximum scaled gain rule (Skogestad andostlethwaite, 2005).

The scaled gains for the different control variables are shown inable 3.

The results from the self-optimizing study show that tempera-ure around the middle of the stripper is a good control variable forll cases. The best place for temperature control changes a little forifferent configurations.

. Control structure and tuning of controllers

Before switching to dynamic mode in Unisim simulator, all thequipment and valves are sized and the flow specifications definedor input and output streams. The make-up tank is sized based on

130 C

ic degrees of freedom.

10 min liquid residence time and the absorber and stripper sumpheights are assumed to be 2 m. The control strategy and necessarycontrol loops are implemented as shown in Fig. 6.

Point 19 224 261 164Point 20 145 166 106

Reboiler temp. 90 102 84Bottom CO2 mole frac. 636 998 814

M. Karimi et al. / International Journal of Greenhouse Gas Control 7 (2012) 230–239 235

AbsorberDesrober

Pump 1

Reboiler

Rich/Lean Exchanger

Surge TankPump 2

V-9

Condenser

Cooler

V-5

V-2

V-4

V-10

V-3

Steam

CoolingWater in

AmineMakeup

WaterMake up

Flue Gas fromPower Plant

ToStack

CO2

Condensate

CoolingWater out

CoolingWater out

CoolingWater in V-8

LC

LC

TC

XC

XC LC

PC

TCLC

TC

V-7

ention

p2ad

F

Fig. 6. Process flowsheet of the conv

arameters are calculated using the SIMC method (Skogestad,

003). Here a step change is made in the manipulated variablend the control variable change is recorded. For the first-order pluselay process these changes are shown in Fig. 7.

ig. 7. Step response of first-order plus time delay process (Skogestad, 2003).

V-6

al configuration with control loops.

k = �y(∞)�u

(3)

k′ = k

�1(4)

The tuning parameters for the PI controller are calculated from Eqs.(5) and (6).

KC = 1k′ · 1

(� + �C )(5)

�I = min(�1, 4(� + �C )) (6)

�C is the tuning parameter in this method. For the initial guess weconsider �C = �. The controller performance with calculated KC and�I must be tested. We can test the controller performance by intro-ducing a step change in setpoint or a disturbance to the system.If the control variable is oscillating we increase the �C and calcu-late new tuning parameters. This is continued until the controlledvariable has an acceptable response.

Some parts of the process may behave like a pure integratorsuch as the absorber level. The response of the control variable foran integrating process is shown in Fig. 8.

In these systems k′ and �I are calculated as below:

k′ = �y

�t · �u(7)

�I = 4(� + �C ) (8)

236 M. Karimi et al. / International Journal of Greenhouse Gas Control 7 (2012) 230–239

Table 4CVs and MVs in the conventional configuration.

Control variable (CV) Manipulated variable (MV) Setpoint

Ratio of CO2 in clean gas to flue gas Lean amine flow rate 90%Lean amine temperature Cooler duty 45 ◦CAbsorber level percent Rich pump duty 50%Condenser pressure Valve actuator 190 kPaCondenser temperature Condenser duty 30 ◦CCondenser level percent Valve actuator of reflux flow 50%Temperature inside the column Reboiler duty a

Reboiler level percent Valve actuator 50%Amine concentration of lean amine Valve actuator of amine flow 30 wt%Make-up tank level percent Valve actuator of water make-up 50%

setp1 ession

Tt

umct

6

vt

6

Ilspiirgflsd

ai

Qstrip =Psat

H2O(TTop,Des)xH2O,freebasis

P∗CO2

(TTop,Des˛Rich)�Hvap

H2O (9)

a The temperature profile is not the same for the different configurations and the09.5, 107 and 100 ◦C for the conventional, the split-stream and the vapor recompr

Now we can calculate the tuning parameters for all controllers.he control structure includes 10 feedback loops for the conven-ional configuration as listed in Table 4.

For all manipulated variables the upper and lower bound val-es are defined. The bounds depend on process limitation such asaximum available utility. The upper bounds for the reboiler and

ondenser duties are 20% and for the pumps and cooler 40% aboveheir values at steady state.

. Dynamic simulation results

Here we show the dynamic behavior of the split-stream and theapor recompression configurations and compare with the conven-ional configuration for different operational situations.

.1. Step changes in reboiler duty

The heat required for stripping is the most important parameter.n the proposed control structure the reboiler duty is a manipu-ated variable and it used to control the stripper temperature. Theource of this energy is a steam stream from the power plant. It isossible to reduce the reboiler duty even if the power plant load

s not changing, for example if extracted steam from the turbines used somewhere else, or if the heat transfer coefficient in theeboiler decreases. So, for the first case study we want to investi-ate the plant behavior if the reboiler duty decreases but the flue gasow rate and composition are constant. The supposed controller iswitched to manual mode and a step change of −10% in reboiler

uty is introduced to each configuration at time = 20 min.

The stripper pressure controller is the first responding controllernd acts about two minutes after the disturbance. The lean load-ng changes immediately after the disturbance, but the lean amine

Fig. 8. Step response of integrating process (Skogestad, 2003).

oint of the temperature controllers in the stripper are not equal. The set points areconfigurations respectively.

valve (CO2 capture rate controller) acts after longer time (about10 min) because the residence time in the make-up tank and pipelines between stripper and absorber. The last controller to act is theabsorber level controller.

For all configurations the CO2 capture rate is near constant for awhile but starts to decrease fast and at the end reaches a constantvalue (Fig. 9a). However, the produced CO2 decreases rapidly at firstand after sometime the rate of decrease reduces and finally reachesa constant value as is shown in Fig. 9b.

The reason for this behavior is that by decreasing the reboilerduty, less CO2 is stripped and the produced CO2 decreases rapidly.At the same time the lean loading increases rapidly, but the CO2emission controller forces CO2 capture rate to be constant byincreasing the lean amine flow rate as shown in Fig. 9b.

Stripping heat, sensible heat and reaction heat are three terms ofheat that is required in the stripper (Erga et al., 1995). These termsare calculated as:

Fig. 9. Dynamic response of the plant for 10% reboiler duty reduction.

of Greenhouse Gas Control 7 (2012) 230–239 237

Q

Q

iap

rdpa

tsrcursr

fdtp

taAz

6

mitcoia(a

rstl

dacCwafi

tiapdvm

Fig. 10. Dynamic response of absorber for 10% change in flue gas flow rate.

M. Karimi et al. / International Journal

sens = �Cp�T

(˛Rich − ˛Lean)CAM(10)

des = �Habs CO2(11)

When the amine flow rate increases, the sensible heat term willncrease and the other two terms will decrease. Consequently, themount of CO2 released in the stripper decreases. Therefore theroduced CO2 decreases more, but with a smaller gradient (Fig. 9b).

Eventually the lean amine valve saturates and the amine flowate cannot increase more. Therefore the CO2 capture amount isecreasing fast and reaches a constant value as visible in the secondart of Fig. 9a. In this time the produced CO2 reaches a constant rates shown in Fig. 9b.

The different configurations have similar behavior in responseo a reboiler duty reduction, but they reach different new steadytates after the transient period. From Fig. 9 we can conclude thateboiler duty reduction has more negative effect (more reduction inapture ratio) for split-stream configuration than two other config-rations. The CO2 capture rate is 72.4% instead of 90% after the planteaches steady state for this configuration. The CO2 capture ratiotabilizes at 79.8% and 76.6% for the conventional and the vaporecompression configurations respectively.

Sometime after the reboiler duty reduction the CO2 capturerom the flue gas is nearly constant although the produced CO2 hasecreased. This means that there should be some accumulation inhe plant. Fig. 9c shows this accumulation rate during the transienteriod.

The curves show the accumulation rate and the total accumula-ion is equal to the area under each curve. Every curve has a peaknd indicating the time that lean amine valve goes into saturation.fter valve saturation the accumulation rate reduces and becomesero when the plant reaches a steady state.

.2. Flue gas flow rate change

Flue gas flow rate change is one of the main disturbances thatay happen in the CO2 capture plant because the power plant load

s changed with the electricity demand. Here we want to inves-igate the plant behavior when the flue gas rate is changing. Thehange may happen gradually or very fast. In addition to the sizef a disturbance, the speed of change will affect the plant stabil-ty. Therefore, we introduce ±10% changes in flue gas flow rate in

ramp type function over 30 and 10 min, as well as step changeFig. 10a). There is a period of time, about 50 min, needed to obtainlmost steady state in the plant after every change.

For this disturbance the first controller to act is the CO2 captureate controller that acts about one minute after the disturbancetarted. After that the absorber level controller acts and the last con-rollers to respond are the controllers in the stripper (temperature,evels and pressure).

All configurations perform well with regard to dampening theisturbance effect. By increasing the flue gas flow rate, the leanmine flow rate increases (Fig. 10b) to keep the CO2 capture rateonstant. Immediately after a change there is some deviation in theO2 capture rate from the setpoint (90%). The deviation increasesith the speed of the disturbance, i.e. a step change is worst, and

lso ramping is best. The deviations observed for split-stream con-guration is larger than for the other two configurations (Fig. 10c).

The CO2 capture rate changes in the absorber and consequentlyhe CO2 product changes in the similar shape (Fig. 11a). Increas-ng the amine flow rate causes a reduction in stripper temperaturend the reboiler duty increases to prevent this reduction. The tem-

erature in the stripper in the alternative configurations is affectedifferently by the amine flow rate. The largest changes are in theapor recompression configuration and the reboiler duty showsore serious changes immediately after the disturbances (Fig. 11b).

Fig. 11. Dynamic response of stripper for 10% change in flue gas flow rate.

The lean loading changes are shown in Fig. 11c. The lean load-ing changes are very smooth, even for the step change in flue gas

flow rate, for the split-stream configuration compared to other twoconfigurations.

2 of Gree

6

flwwrs

6

tahitfbrfCrtphlti

Fr

38 M. Karimi et al. / International Journal

.3. Flue gas CO2 composition change

The plants were tested with a change in CO2 composition of theue gas. The changes were ±10% in flue gas CO2 concentration andere introduced as a ramp type with 30 and 10 min time span, asell as a step change. The dynamic responses of different configu-

ations are very similar to those for the flue gas flow rate change,o the results are not shown here.

.4. Flue gas flow rate change with constant reboiler duty

In this case study we want to test the different configurations ifhe reboiler duty is constant and the flue gas flow rate increases forperiod of time. This condition may happen in reality if a problemappens for temperature control loop in the stripper. The flue gas

ncreases 10%, as a step change, for 30 min and then returns the ini-ial value (Fig. 12a). The interesting result of this case study is thator all configurations the CO2 capture rate drops from the set point,ut the time when the drop starts is different for different configu-ations. If the duration of the disturbance increases, the start timeor the drop will decrease. When the flue gas flow rate increases theO2 capture rate can remain constant by increasing the amine flowate. But the extra captured CO2 cannot strip in the stripper becausehe reboiler duty is constant. Therefore the CO2 accumulates in thelant and when the flue gas flow rate returns to the initial value we

ave a lean amine with higher loading compared to the initial lean

oading. So when the flue gas flow rate returns to the initial value,he lean amine flow rate needed for 90% CO2 removal is more thants initial value. As we mentioned before when the amine flow rate

ig. 12. Dynamic response of the plant when flue gas flow increase with constanteboiler duty.

nhouse Gas Control 7 (2012) 230–239

increases, more energy is needed for separation or, on the otherhand, we will have less separation with the constant reboiler duty.This increases the lean loading more and at the end the lean aminevalve will saturate and the CO2 capture rate will decrease. Fig. 12shows the dynamic response of the plant in this case.

The results show that in the split-stream configuration, thecapture rate will reduce after a short time interval. The vaporrecompression configuration can operate for a longer time with-out CO2 capture reduction, while the conventional configurationhas the longest time. It means that if there is problem in the con-trol loop of the stripper temperature, we have more time to repairit without CO2 capture reduction, even some disturbances happen,if the conventional configuration is used.

7. Conclusions

The dynamic behavior of the split-stream and the vapor recom-pression configurations are investigated, and the plant responsesare compared with the conventional configuration as a benchmark.The controllers are tuned using the SIMC method (Skogestad, 2003).Different disturbances are introduced to the plant. Different con-figurations have different responses to the disturbances becausethe flow rates, the unit operations number and capacities are notthe same and consequently they have different complexities. First,a 10% step reduction is made in reboiler duty. The response of thischange is investigated for the different configurations. The CO2capture rate is constant for a period of time before it falls. Thevapor recompression configuration will hold the CO2 capture ratenear 90% for the longest time and the split-stream configurationfor the shortest time. After the transient period, the conventionalconfiguration has the highest capture rate and the split-streamconfiguration has the lowest value.

All configurations show good performance to dampen a 10% dis-turbance effect in flue gas flow rate and CO2 concentration. Thelean amine valve is not saturated for this amount of change, but ifthe disturbance is so large that the lean amine valve saturates, thebehavior of the different configurations will change.

In the last case study, the flue gas flow rate is increased for aperiod of time and then returned to the initial value with a constantreboiler duty. In this case, for all configurations, the CO2 capture ratedropped from the set point, but over a varying period of time. In thesplit-stream configuration the capture ratio falls over a short periodand for the conventional configuration this period is the longest.

From these case studies we can conclude that the conventionalconfiguration has the best dynamic behavior and is the most stableone. Of the other two configurations we can say that the vaporrecompression configuration can handle disturbances better thanthe split-stream configuration. Decreasing the CO2 capture rate toa lower value in a short time is the most important weak pointcompared to the vapor recompression configuration in this study.

Finally we should mention that the presented result is for astraight forward control configuration. There is a possibility toimprove the plant control if we change the control configurationor use other control methods like model predictive control (MPC).In addition, changing some parameters such as liquid holdup in theplant (make-up tank volume, absorber and stripper sump volumeand piping volume) will change the dynamic behavior. In futurestudies we will investigate the effect of different control structuresand liquid holdups on the dynamic behavior of the plant.

Acknowledgements

Financial support provided through the CCERT project (182607),by the Research Council of Norway, Shell Technology Norway AS,

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etso Automation, Det Norske Veritas AS, and Statoil AS is greatlyppreciated.

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