Mathematical modeling and optimization of carbon dioxide stripping tower in an industrial ammonia plant

International Journal of Greenhouse Gas Control 58 (2017) 42–51

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Mathematical modeling and optimization of carbon dioxide stripping tower in an industrial ammonia plant L. Mahmoodi, P. Darvishi ∗ Department of Chemical Engineering, School of Engineering, Yasouj University, Yasouj, Iran

a r t i c l e

i n f o

Article history: Received 7 May 2016 Received in revised form 11 September 2016 Accepted 4 January 2017 Keywords: Stripping tower Carbon dioxide Benfield solution Modeling Optimization

a b s t r a c t In the present study, performance of an industrial regeneration tower has been examined and validated for stripping of carbon dioxide from diethanolamine (DEA)-promoted hot potassium carbonate solution. The applied model employs a comprehensive non-linear rate-based method to take into account the coupling between material and energy balances, thermodynamic vapor-liquid equilibrium (VLE) relations and chemical kinetics. VLE relations are used to model the flash section located on the top of tower. The penetration theory provides an appropriate desorption rate and enhancement factor for the chemical desorption which incorporates an extensive set of important reactions. The model predictions were compared with the operating data obtained from ammonia plant of a petrochemical complex. The impact of parameters such as steam temperature, potassium carbonate concentration, inlet temperature of rich solution, promoter concentration and column pressure on the tower performance have been examined and optimized. Acceptable agreements have been attained between the simulation and measured industrial data. The optimum conditions were found to be 2.5 wt%, 1.6 barg, 401 and 378 K for promoter concentration, column pressure, steam temperature, and inlet temperature of rich solution which demonstrate a good compatibility in comparison to optimum operational conditions. © 2017 Elsevier Ltd. All rights reserved.

1. Introduction In the production of ammonia, the removal of CO2 from synthesis gas is an essential step to inhibit corrosion and increase the production rate. The carbon dioxide captured from ammonia plant is often recovered as a useful by-product in the production of urea and other fertilizers. Several approaches such as absorption, adsorption and membrane technology have been developed to remove CO2 from synthesis gas (Rahimpour and Kashkooli, 2004). Absorption/desorption process has been industrially adopted in many processes, such as synthesis gas production, natural gas purification, oil refining and hydrogen manufacture. The rich carbon dioxide solution discharged from absorber should be exposed to the low pressure stripping tower, in which CO2 is released. Among all solutions, promoted hot potassium carbonate, known as Benfield has gained wide industrial acceptance for CO2 removal which was initially developed by the U.S. Bureau of Mines, at Bruceton, Pennsylvania and the improvements to the process were made by Benson and Field in 1970s.

∗ Corresponding author. E-mail address: [email protected] (P. Darvishi). http://dx.doi.org/10.1016/j.ijggc.2017.01.005 1750-5836/© 2017 Elsevier Ltd. All rights reserved.

Aqueous potassium carbonate solution has several advantages, such as higher absorption capacity, lower solvent cost, lower energy requirement for regeneration, and high resistance of oxidation, in comparison with aqueous ethanolamine solutions. Amines have a relatively high rate of reaction with dissolved carbon dioxide; However, their performance as solvents is limited by a high heat of absorption, along with issues associated with amine loss and degradation and corrosion (Rahimpour and Kashkooli, 2004). In spite of the many advantages of aqueous potassium carbonate solution, it has a relatively slow CO2 absorption reaction rate and corrosiveness. Therefore, many investigators have used promoters to increase the CO2 absorption rate and inhibit the corrosion problems. One way to improve the overall system performance for CO2 capture is to blend a fast reacting amine with potassium carbonate to take advantage of the benefits of both solvents. Amine promoter could significantly enhance absorption rate, while the carbonatebicarbonate buffer provides a feasible regeneration. This offers other advantages such as high stripping capacity and economy. Addition of promoters to the solution which accelerate the rate of absorption of carbon dioxide, results in appreciable improvement of the process Utilizing the DEA activator for the potassium carbonate solution, has been resulted in producing high-purity treated gas with capital and operating costs substantially below those of the process using unactivated solutions. The DEA activator is inexpen-

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Nomenclature a CP,L Cp,k Cp,G E Eff. H hg kL kg.k Kg,CO2 G L Nk Pk PCO2,e PH2O,i TG TL xk yk Z HCO2 HH2 O

Specific surface area (m2 m−3 ) Molar specific heat of liquid (kcal kmol−1 K−1 ) Specific heat of kth component (kcal kmol−1 K−1 ) Molar specific heat of gas (kcal kmol−1 K−1 ) Enhancement factor Efficiency Solubility of carbon dioxide in solution (kmol m−3 atm−1 ) Heat transfer coefficient in gas phase (kcal m−2 h−1 K−1 ) Liquid side mass transfer coefficient (m h−1 ) Mass transfer coefficient of kth component in gas phase (kmol m−2 h−1 atm−1 ) Overall mass transfer coefficient for CO2 based on liquid phase (m h−1 ) Molar velocity of gas (kmol h−1 m−1 ) Molar velocity of liquid (kmol h−1 m−1 ) Mass transfer of kth component (kmol m−2 h−1 ) Partial pressure of kth component in gas phase (atm) Equilibrium vapor pressure of carbon dioxide in gas phase (atm) Partial pressure of water in gas- liquid interface (atm) Gas temperature (K) Liquid temperature (K) Mole fraction of kth component in liquid phase Mole fraction of kth component in gas phase Spatial variable along the height of column (m) Heat of reaction and stripping of CO2 (kcal kmol−1 ) Heat of vaporization of water (kcal kmol−1 )

sive, stable and introduces no operating problems. A large number of plants using this process are presently in operation. Due to the use of steam as the stripping phase, desorption manipulates the economy of the process. Optimal stripper operation is essential because desorption energy requirement accounts for 80% of the operating cost of absorption/stripping system. Besides, corrosion problems have been receiving a great deal of attention because they have substantial impacts on the plant’s economy, especially in terms of unplanned downtime, production losses, reduced equipment life, safety of plant personnel, and extra-expenditure for restoring the corroded equipment. These problems further prevent the absorption/stripping process from achieving energy efficient operations. A number of research literature reported the severity of corrosion in the Benfield process. Dissolved CO2 is the main contributor to such corrosion. Pure carbonate solutions without dissolved CO2 are not aggressive towards carbon steel. It was reported that corrosion problems in carbon steel equipment with potassium carbonate were encountered, especially where the solution loadings and conversion to bicarbonate are high or where carbon dioxide and steam are released by pressure reduction (Bienstock and Field, 1961). It is revealed that a 40 wt% potassium carbonate solution saturated with CO2 corroded carbon steel at the rate of 340 mils per year and the corrosion rate was considerably reduced in the presence of H2 S (Bienstock and Field, 1961). Numerical modeling provides a detailed understanding of the stripper operation, presents insight into the stripping phenomenon, and results in optimal designs. However, what is published on the performance of stripping towers is novel to the readers, since prior studies have scarcely examined the performance of stripper and the flashing section leading to it. A

43

number of studies include absorption/desorption (Alatiqi et al., 1994; Alié et al., 2006; Desideri and Paolucci, 1999; Aroonwilas et al., 2003), others include only the stripping (Tobiesen et al., 2008; Oyenekan and Rochelle, 2009). A few authors have attempted to establish the model of gas-liquid mass transfer in RPB as well. Many investigations have been already published include modeling, simulation and even optimization of the absorbing towers (Cullinance and Rochelle, 2004; Lin et al., 2009; Matsunaga et al., 2009; Schubert et al., 2001). Though, experimental results of a CO2 absorber/stripper by MEA, DEA, MDEA, AMP and their mixtures in an experimental hemispherical contactor was obtained (Jamal et al., 2006), its industrial performance has not reported. KLG method was used to simulate carbon dioxide absorber/stripper by hot potassium carbonate promoted by piperazine in a micro scale liquid film (Park, 2014). A rate-based absorber/stripper modeling tool was applied in a pilot plant by aqueous MEA solution to predict the amount of total desorbed CO2 and location of the operation (Zhang et al., 2009). A linearization scheme was applied in order to solve CO2 absorber/stripper analytical model with MEA solution (Meldon and Morales-Cabrera, 2011). A proposed model by Sanyal et al. simplified the calculations and gave reasonable predictions for Benfield process in packed bed (Sanyal et al., 1988). It would be of interest to study the corrosion behavior of carbon steel in Benfied solution for better understanding of the corrosion mechanisms and to have an idea about the main controlling parameters, which are valuable data in the design and operations of such units. Corrosion in carbonate system is affected by concentration of carbonate and bicarbonate ion, ratio of carbonate and bicarbonate, temperature and solution velocity. This information is obtained from profiles of temperature, flow rate and components mole fraction along the tower. This study presents a complex rigorous non-linear model for process of reactive carbon dioxide stripping using promoted hot potassium carbonate, considering its flashing section due to high pressure drop at the entrance of the stripper, as a new case study. The model is adjusted in order to properly consider insight into the phenomenon of gas and liquid mass transfer with chemical reaction at stripper conditions. Obtained results are allowed to determine the reboiler energy required to optimize the total costs. 2. Process description Fig. 1 shows a schematic diagram of CO2 stripping tower in the ammonia plant of Shiraz Petrochemical Complex. The feed gas containing acid gases enters the bottom of absorber and flows countercurrently to Benfield solution entering the top of tower. Rich solution containing absorbed acid gases is preheated and subsequently supplied to the top of the regenerator. Before entering rich Benfield solution to the stripper, it passes through a throttling valve installed in the closest proximity of the tower. This valve which causes high pressure drop at constant temperature, flashes the rich solution and thus two phases of vapor and liquid are formed. The produced vapor and stripped acid gases discharge from the top of tower. Liquid phase moves downward and countercurrently comes into contact with steam, produced by the reboiler, to strip acid gases. Finally regenerated liquid is cooled and recirculated to the absorber for further gas absorption. 3. Model development The presented model formulates the couples of fundamental mass and energy balance equations, VLE and kinetic relations around a differential height of the column. The system of differential equations is highly non-linear and should be solved over all the

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Fig. 2. Differential section of packed bed.

differential equations for gas flow, mole fraction of CO2 and H2 O and temperature of gas phase: dG = (NCO − NH2 O )a 2 dz

Fig. 1. A schematic diagram of stripping tower.

dyCO2 nodes and grids to predict the performance of the tower. The model is proposed based on the following assumptions:

=

dz dyH2 O

– – – –

Steady-state conditions prevail. Amine vaporization rate is negligible. The reaction takes place in the liquid phase. CO2 and H2 O are the only components transported across the interface. – Pressure drop across the column is negligible. The basic reaction for desorption of CO2 from hot potassium carbonate solutions is represented by the following reversible reaction: 2KHCO3 ⇔ K2 CO3 + H2 O + CO2

(1)

Modeling of this complex operation requires an adequate modeling framework. There are two main approaches to model strippers: the equilibrium (Oyenekan and Rochelle, 2009) and ratebased approaches (Oyenekan and Rochelle, 2009). The application of the rate-based approach to model multicomponent reactive separation processes has proved to be superior to other concepts like the equilibrium stage model. The key element of the rate-based model is an axial segment of a packed column in which simultaneous mass transfer and chemical reaction are described according to the extended film model (Fig. 2). By the rate-based approach which is used in this work, reaction kinetics with heat and mass transfer rates are considered directly, vapor and liquid phases are balanced separately and their simultaneous and unequal fluxes across the interface are accounted (Matsunaga et al., 2009; Schubert et al., 2001; Jamal et al., 2006). It is assumed mass transfer resistances in both phases and thermodynamic equilibrium at the interface (Park, 2014; Saimpert et al., 2014; Meldon and Morales-Cabrera, 2011). 3.1. Gas phase Fig. 2 shows an elemental volume in a differential packed height (Z) of the stripper, consisting of the gas and liquid phases. Differential mole and energy balances for the gas phase around a differential height of packed bed (Z) give rise to the following

[NH2 O yCO2 + NCO2 (1 − yCO2 )]a

(3)

G

=−

dz

(2)

[NCO2 yH2 O + NH2 O (1 − yH2 O )]a G

(4)

a(NH2 O − NCO2 )TG dTG = G dz +

(NCO2 Cp,CO2 TL − NH2 O CP,H2 O TG )a G.CPG

−

hg a(TG − TL ) GCp,G

(5)

3.2. Liquid phase Similarly, differential mole and energy balances for the liquid phase around the Z give rise to the following differential equations for liquid flow, mole fraction of KHCO3 , K2 CO3 and H2 O and temperature of the liquid phase: dL = −NH2 O .a dz dxK2 CO3 dz dxKHCO3 dz dxH2 O dz

= =

=−

(6)

[NH2 O xK2 CO3 − NCO2 ]a L [NH2 O xKHCO3 + 2NCO2 ]a L [NCO2 (1 + xH2 O ) + NH2 O (1 − xH2 O )]a L

(7) (8) (9)

a(NCO2 Cp,CO2 TL − NH2 O CP,H2 O TG ) aNH2 O TL dTL + = − L LCP,L dz a(NCO2 HCO2 − NH2 O HH2 O ) hg a(TG − TL ) + LCp,L LCp,L

(10)

Applying the transport equations, fluxes of CO2 and H2 O are calculated by the following expressions: NH2 O = kg,H2 O (PH2 O − PH2 O,i ) NCO2 =

kg,CO2 kL EH kg,CO2 + kL EH

(PCO2 ,e − PCO2 ) = Kg,CO2 (PCO2 ,e − PCO2 )

(11) (12)

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45

Table 1 Characteristics of the stripping tower and packing system. Parameter

value

Unit

Height of packing Diameter of packed bed Packing size Packing shape Specific surface of packing Pressure of the tower

26 5.5 40 Metal mini rings 197 1.6

m m mm – m−1 barg

3.3. Flashing region In the regenerator, CO2 stripping occurs in three regions: (a) at the inlet of tower where flashing can occur if the sum of equilibrium partial pressures of CO2 and water is greater than the operating pressure of the stripper; (b) within the section of packing due to normal mass transfer; (c) in the reboiler under boiling conditions. As a result, the presented model divides the stripper into a flash region at the top, packing bed and an equilibrium reboiler. The flash region quantifies the effect of flash phenomena using the ratebased approach and electrolyte NRTL method in removing of acid gases. The main equations in this region are as follows: F=L+V

(13)

Fzi = Vyi + Lxi

(14)

ki =



yi xi

(15)

(ki − 1)zi =0 1 + (ki − 1)(V/F)

(16)

Fig. 3. Variation of (a) liquid (b) gas mass transfer coefficient versus height of the tower in the presence and absence of DEA.

3.4. Input data Table 1 presents physical characteristics of the installed tower in Shiraz Petrochemical Complex. The physical properties of materials, packing specific fluid dynamics, relevant mass transfer correlations and the component specific reaction rates are presented by adequate relations in Table 2. Flow rate, composition and temperature of the gas leaving the top of the stripper and the flow rate, loadings and temperature of the rich solution entering the top of the stripper were known as input data. Initial guesses for the segment temperatures, partial pressures and loadings were also provided. 3.5. Numerical solution In order to simulate the CO2 removal in stripper, the proposed differential Eqs. (2)–(10) which are substituted by Eqs. (11) and (12) need to be solved numerically in MATLAB software by taking into account the required parameters and boundary conditions. Estimating the composition of the gas leaving the stripper is an essential step in simulation which helps the equations be integrated by Runge-Kutta method up to the outer edge of stripper. By using the shooting method, a new value of CO2 composition in the outlet gas was estimated for further iteration until satisfactory convergence was obtained. Finally, temperature and composition profiles, regeneration efficiency and phase flow rates were calculated. 4. Results and discussion According to Eqs. (11) and (12), addition of DEA as a promoter is included in the parameter called Enhancement factor. The promoter directly induces higher CO2 removal, as a result of higher mass transfer coefficient in the presence of amine.

In accordance to Fig. 3, values of mass transfer coefficients in the liquid and gas phases are compared. Due to enhancement factor, mass transfer coefficient in the presence of amine is higher than the case without amine. The increase in mass transfer coefficient from the rich to the lean end is due to the presence of more free K2 CO3 in the liquid available for reaction. The mass transfer process in the stripper can be separated into its component mechanisms. According to Fig. 3, the resistance associated with diffusion of reactants and products in the liquid phase, dominates the stripper operation. In addition, the increased reaction rates at high temperature causes the stripping operation to be liquid phase controlled. This can also be confirmed by evaluating the overall mass transfer coefficients in the liquid and gas phases. There are still appreciable contributions by the gas phase resistance. Raising the system pressure will increase the contribution of gas phase resistance. The gas resistance is greater at the lean end of the stripper. At nearatmospheric pressure, most studies on stripping operations have assumed the stripping operation to be controlled by diffusion of reactants and products.

4.1. Model verification Table 3 presents a verification of the predicted model results by comparison with the plant data, under the design specification for the column summarized in Table 1. Overall, for all cases the obtained agreement is satisfactory. In order to examine the impact of stripper operation parameters on corrosion problems, simulation results are plotted as profiles of dependent variables versus height of the tower bed. Besides, the effects of various parameters on the performance and economy of stripper were investigated and verified with industrial plant data.

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Table 2 Equilibrium, kinetic and transport parameters of the proposed model. Parameter

Expression

PCO2,e

˛ PCO2,e = 1.95 × 109 × m0.4 1−˛ exp( −8160 ); m = molarity of solution T

Source 2

L



HCO−





HCO−

3



3

=

Astria (1983)

˛

˛=

L

L = 12.9275 − (0.114TL ) + (4.118 × 10−4 TL2 ) − (5.797 × 10−7 × TL3 )

g

g =

H

H = 10(−4.3856+

kgk

kgk = 5.32

kDEA

kDEA = 6.4 × 108 × exp(14.97(1 −

kOH

log(kOH ) = 14.131-(2895/TL )

PH2 O,i

L LogPH2 O,i = −( 2.303R )( T1 − 2.45 × 10−3 ) − 1.672 + C

Tosh et al. (1959)

C

C = 1.2014 + (0.2857/XC ) − (0.0537/XC2 )

Tosh et al. (1959)

XC

XC =

[K + ]

Astria (1983)

2m

P.MWgas RTG

–

aDgk RT

867.4932 ) TL

G ( a ) g

0.7

× (0.025m) 

( g Dg )

(1/3)

gk

Astria (1983)

(adP )

−2

353 )) TL

L

0.691(%KHCO3 ) %K2 CO3 +0.691(%KHCO3 )



Onada and Takeuchi (1968)

Leader (1971) Tseng and Savage (1988)



hg =

Bocard and Maryland (1962)

Tosh et al. (1959)

yk hgk

hg

Reid et al. (1977) hgk = kyk Cpk 2/3

DL,CO2

log(DL,CO2 ) = - 3.0188 - (586.9729/TL ) − 0.4437

kL

kL (

1/3

2 L

/DCO2 = 0.015( aL )( 

)

2 g

L

L

L L DCO2

)

1/3

Savage et al. (1980)

Van Krevelen and Hoftijzer (1948)

HCO2

HCO2 = 27228.2 + 81.37ω + 5.32ω2 − 0.1313ω3 + 1.654 ∗ 10−3 ω4

Bocard and Maryland (1962)

HH2 O

HH2 O = 9800(kcal/kmol)

–

ω

ω=

DAB

DAB =

%K2 CO3 +0.691(%KHCO3 ) %K2 CO3 +0.781(%KHCO3 )+%H2 O 10−3 ×T 1.75 ( G 

105 ×P[

 E

E=

1+

Bocard and Maryland (1962)

1 + 1 ) MWA MWB



V 0.3 +

V 0.3 ]

Kohl and Risenfield (1985)

2

B

A

DL (kOH [OH− ]+kAm [Am])

–

k2 L

C  P,L

C  P,L = 0.998 − (0.00914ω) − (0.1063 × 10−3 ω2 ) + (0.3058 × 10−5 ω3 )

Kohl and Risenfield (1985)

L

L = 1.0679 + 0.01ω − (9.4 × 10−4 TL )

Bocard and Maryland (1962)

PCO2,e

˛ PCO2,e = 1.95 × 109 × m0.4 1−˛ exp( −8160 ); m = molarity of solution T

˛

˛=

L

L = 12.9275 − (0.114TL ) + (4.118 × 10−4 TL2 ) − (5.797 × 10−7 × TL3 )

g

g =

H

H = 10(−4.3856+

kgk

kgk = 5.32

2

L



HCO−



3



HCO−

[K + ]



3

=

Astria (1983)

2m

P.MWgas RTG

aDgk RT

Astria (1983)

Bocard and Maryland (1962) –

867.4932 ) TL

G ( a ) g

0.7

× (0.025m) 

( g Dg ) gk

(1/3)

(adP )

Astria (1983) −2

Onada and Takeuchi (1968)

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47

Table 2 (Continued) Parameter

Expression

Source 8

353 )) TL

kDEA

kDEA = 6.4 × 10 × exp(14.97(1 −

kOH

log(kOH ) = 14.131-(2895/TL )

PH2 O,i

L LogPH2 O,i = −( 2.303R )( T1 − 2.45 × 10−3 ) − 1.672 + C

Tosh et al. (1959)

C

C = 1.2014 + (0.2857/XC ) − (0.0537/XC2 )

Tosh et al. (1959)

XC

XC =

Leader (1971) Tseng and Savage (1988)



L

0.691(%KHCO3 ) %K2 CO3 +0.691(%KHCO3 )

hg =



Tosh et al. (1959)

yk hgk

hg

Reid et al. (1977) hgk = kyk Cpk 2/3

DL, CO2

log(DL,CO2 ) = - 3.0188 - (586.9729/TL ) − 0.4437

kL

kL (

2 L

2 g

1/3

)

L L DCO2

/DCO2 = 0.015( aL )(  L

L

)

Savage et al. (1980)

1/3

Van Krevelen and Hoftijzer (1948)

HCO2

HCO2 = 27228.2 + 81.37ω + 5.32ω2 − 0.1313ω3 + 1.654 ∗ 10−3 ω4

Bocard and Maryland (1962)

HH2 O

HH2 O = 9800(kcal/kmol)

–

ω

ω=

DAB

DAB =

E=

E

%K2 CO3 +0.691(%KHCO3 ) %K2 CO3 +0.781(%KHCO3 )+%H2 O 10−3 ×T 1.75 ( G 

105 ×P[



1+

Bocard and Maryland (1962)

1 + 1 ) MWA MWB



V 0.3 +

V 0.3 ]

Kohl and Risenfield (1985)

2

B

A

DL (kOH [OH − ]+kAm [Am])

–

k2 L

C  P,L

C  P,L = 0.998 − (0.00914ω) − (0.1063 × 10−3 ω2 ) + (0.3058 × 10−5 ω3 )

Kohl and Risenfield (1985)

L

L = 1.0679 + 0.01ω − (9.4 × 10−4 TL )

Bocard and Maryland (1962)

Efficiency

Efficiency =

Volumetric flow rate of stripped carbon dioxide(m3 h−1 )

–

Mass flow rate of steam used in reboiler(kg h−1 )

Table 3 Comparison of calculated results with the observed plant data (Shiraz Petrochemical Complex). Parameter −1

Gas flow rate (kmol h ) Mole fraction of CO2 in gas phase Mole fraction of H2 O in gas phase Liquid flow rate (kmol h−1 ) Mole fraction of H2 O in liq. phase Mole fraction of K2 CO3 in liq. Phase Mole fraction of KHCO3 in liq. phase Gas temperature (K) Liquid temperature (K)

Inlet (observed)

Outlet (observed)

Outlet (computed)

Error (%)

1656.83 0.02 0.98 30,863.51 0.7579 0.0116 0.1161 401 379

1447.632 0.98 0.02 30,299.77 0.864 0.064 0.062 379 401

1447.64 0.818 0.18 30,272.59 0.873 0.0637 0.0633 378.3 401

0.0006 17.38 88.9 0.0897 1.042 0.469 2.097 0.185 0

4.1.1. Gas temperature profile According to Fig. 4, due to the heat exchange between two phases and water vapor condensation into liquid phase taking place in the interface zone, temperature of gas drops gradually moving from the bottom to the top of the tower. Temperature has a significant impact on corrosion. Upon an increase in temperature from 20 to 80 ◦ C, the corrosion rate increased twice in pure carbonate-bicarbonate solution. Increasing solution temperature from 75 to 96 ◦ C led to increases in anodic current density of about an order of magnitude and a decrease in the extent of the passive region (Frolova et al., 1997).

4.1.2. Flow rate profiles Due to gas- liquid contact, water and carbon dioxide are transferred between the two phases, leading to variation in gas-liquid flow rates over the packed bed, which illustrated in Fig. 5. At the first half height of stripper, changes in liquid flow rate due to CO2 desorption are significant and then gradually decrease. On the other hand, higher water vapor condensation into liquid phase compared to carbon dioxide released into gas phase is the reason for decreased gas flow rate. Solution flow rate has a significant impact on corrosion. Higher erosion is obtained due to the increasing shear stress exerted by the increasing solution velocity and impingement of gas and solu-

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Fig. 4. Variation of gas phase temperature versus height of the stripping tower bed.

Fig. 5. Variation of (a) liquid (b) gas phase flow rate versus height of the tower bed.

tion on metal surfaces (Nielsen et al., 1995). At high velocity, both erosion and velocity-dependent corrosion play an important role. In the case of inhibited systems, a protective film is developed to cover the metal surface and suppress the excessive corrosion. However, this film can be removed or damaged by the shear force of a high velocity fluid stream. In the presence of solid contaminants such as iron carbonate, the solution velocity can cause even more severe erosion-corrosion (Meisen et al., 1996). In a system without a protective film, the corrosion rate is completely controlled by solution velocity (Videm and Dugstad, 1989a,b). Raising the solution velocity reduces thickness of mass transfer film, allowing corrosive chemicals to reach the metal surface at a higher rate. Thus, if corrosion rate is controlled by the rate of mass convection through the film, corrosion is undoubtedly higher. 4.1.3. Component mole fraction profiles Fig. 6 indicates the mole fraction of water and carbon dioxide in the gas phase and potassium carbonate in the liquid phase. Vapor phase moves upward along the stripper and supplies necessary

Fig. 6. Mole fraction of (a) water (b) carbon dioxide in gas phase and (c) potassium carbonate in liquid phase versus height of the stripper bed.

heat for the desorption reaction over the packed bed, causing the increased mole fraction of CO2 in gas phase along the tower. Due to water vapor condensation and increasing the CO2 fraction in gas phase, water mole fraction in gas phase decreases gradually. Water vapor condensation is due to temperature which leads to water vapor pressure reduction. Variation in water vapor pressure affects potassium carbonate solution, which is incredibly sensitive to temperature changes. Due to the abundance of temperature changes at the top of tower, condensation occurs mostly in this region. According to the main Reaction (1), the concentration of potassium bicarbonate in liquid phase decreases from top to the bottom of stripper. In contrast, mole fraction of K2 CO3 increases downward the tower, considering the fact that most of these changes occur in the first half of tower.

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A 40% solution of potassium carbonate has been found to be too concentrated for commercial use in some cases because of its tendency to form a slurry of bicarbonate crystals if the solution is cooled at any point in the circuit. In the presence of bicarbonate slurry, carbon steel cases and impellers of pumps were found to last only a few hours. Reducing the solution concentration below 30% eliminates this severe erosion. High carbonate concentration can cause corrosion damage to the system. By increasing the concentration of K2 CO3 in Benfield solution, the corrosion rate of carbon steel increases. The increase in %K2 CO3 yields more HCO3 − , which enhances the direct reduction of bicarbonate to carbonate ion at cathodic sites, through which carbon steel corrosion takes place. The effect of bicarbonate ion concentration on corrosion rate of low and high strength steel in the 1 N sodium carbonate solutions investigated (Frolova et al., 1997). They reported that the corrosion rate increased with increasing bicarbonate concentration. The corrosion rate began to increase at 1 mg/L HCO3 and became twice as large at 100 mg/L. It is also revealed that high corrosion rate due to high bicarbonate concentration could not be easily reduced by introducing metavanadate corrosion inhibitor to the system (Banks, 1967). In addition to the total concentration of carbonates, a ratio of K2 CO3 /KHCO3 plays a key role in corrosion rate (Lunarska and Szyprowski, 1994). They reported that minimum corrosion rate would be expected to occur at a K2 CO3 /KHCO3 ratio of about 0.1–0.2. KHCO3 concentration should be maintained at 1.0–1.5 M and K2 CO3 /KHCO3 ratio at 0.2–1.0 to minimize the possibility of passive layer degradation and corrosion. A solution concentration beyond 30 wt% is highly vulnerable from the corrosion point of view. This will always maintain the bicarbonate in solution which would otherwise precipitate and start fouling the system (Kohl and Nielsen, 1997). Fig. 7. Effect of inlet temperature of rich solution on (a) mole fraction of potassium carbonate in lean solution (b) efficiency.

4.2. Parameters optimization The energy provided to the reboiler is employed to heat the rich solution entering the stripper, to drive the CO2 desorption reaction and to vaporize water to generate stripping vapor. Understanding the behavior of stripper is a necessary condition to improve its operation and this requires a critical study of many factors that affect the process performance. In order to provide insight into the stripping phenomenon, the effect of process parameters like the steam temperature, potassium carbonate concentration, inlet temperature of rich solution, promoter concentration and column pressure on reboiler energy consumption and regeneration of Benfield solution have been investigated. In the following section, the reboiler energy consumption has been quantified as the efficiency parameter in Table 2. 4.2.1. Inlet temperature of rich solution The influence of rich solution inlet temperature on the performance of regenerator is demonstrated in Fig. 7. A reduction in liquid temperature leads to a decrease in K2 CO3 mole fraction. Since desirable CO2 stripping process requires entering liquid stream with a temperature close to the tower’s temperature, when entering liquid’s temperature reduces, a greater area of the tower will be used to warm up the stream and CO2 desorption decreases. Furthermore, it leads to reduced reaction rate which is followed by a decrease in overall mass transfer coefficient in liquid phase. Due to mentioned reasons, a decrease in solution temperature causes a decrease in stripped carbon dioxide amount in top sections of the tower and the process efficiency decreases accordingly. On the other hand, by reducing solution temperature, water vapor pressure reduces and its condensation intensity increases at the top of the tower, since equilibrium water vapor pressure on Benfield solution depends strongly on temperature.

It is also demonstrated that by increasing the inlet temperature of rich solution, the thermal efficiency of column will be significantly increased. Preheating the feed before it is introduced into the column is important because it reduces the amount of heat added to the column to achieve the desired separation. However, the optimum temperature depends on other factors such as corrosion rate.

4.2.2. DEA concentration Fig. 8(a) reflects that raising amine concentration improves the enhancement factor, leads to a considerable increase of overall mass transfer coefficient in the liquid phase. As a result, mole fraction of K2 CO3 in the lean solution increases. It is noticeable that raising the amine content beyond a specific amount has no significant effect on the exit CO2 concentration. The possible explanation for this behavior is that increasing amine concentration reflects higher overall mass transfer coefficient in the case of liquid-phase controlled mass transfer. With more increasing the amine concentration, the gas phase mass transfer is considered the major factor controlling desorption process, therefore CO2 stripping is unaffected by increasing the promoter concentration. Fig. 8(b) represents regenerator efficiency changes in the range of 0.39–0.66 while concentration of DEA changes from 0 to 4. At low amine concentrations, a little change in DEA concentration leads to a significant increase in regeneration efficiency, while at high amine concentrations, due to increased mass transfer coefficient in liquid phase, resistance in gas phase has greater importance than low amine concentration, so the effect of amine increase on stripping tower total efficiency decreases.

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Fig. 8. Effect of DEA concentration on (a) mole fraction of potassium carbonate in lean solution (b) efficiency. Fig. 9. Effect of stripper pressure on (a) mole fraction of potassium carbonate in lean solution (b) efficiency.

4.2.3. Tower pressure Carbon dioxide removal performance is affected by the system pressure. Fig. 9(a) shows how the concentration of K2 CO3 in lean solution would vary if the operating pressure of the tower changes from its design value. As stripper pressure reduces below its optimum value, kinetics becomes the dominant mechanism for mass transfer in the stripper. On the other hand, as the stripper pressure increases, so does the operating temperature which increases the reaction rates. The combined effect of these factors would lead to improve the desorption performance and increase the K2 CO3 concentration in the lean solution. After the optimum pressure of system, the increase in stripper pressure results in a decrease in the ratio of the partial pressure of water to that of CO2 in equilibrium with the potassium carbonate solution. In this case, increasing the partial pressure of CO2 would result in a lower driving force. Besides, by increasing the stripper pressure, the resistance associated with diffusion of reactants and products dominates the stripper operation. Therefore, the gas phase resistance to mass transfer is increased and the desorption performance reduces. Fig. 9(b) shows how the decrease in partial pressure of water results in a decreased stripping vapor requirement. Since, in this case the stripping vapor is steam; an increased stripper pressure results in a decrease in the total reboiler duty. The operating pressure of a steam stripper can influence the efficiency and reliability of the column. Lower operating pressures will also result in lower operating temperatures and the better the volatility achieved. Stream strippers operating at vacuum pressure can be highly efficient. The drawback of operating at vacuum conditions is the added expense of the operation and maintenance of vacuum equipment. Typically most steam strippers are designed at or near atmospheric operating pressure. Operating at or near atmospheric pressure allows the designer to take advantage of higher

Fig. 10. Effect of inlet concentration of K2 CO3 on efficiency.

volatilities and lower operating temperatures without having the added expense of operating vacuum equipment.

4.2.4. Inlet concentration of potassium carbonate According to Fig. 10, an increase in the inlet potassium carbonate concentration induces a rapid increase in reboiler efficiency while raising the K2 CO3 content beyond a specific amount has no effect on the regeneration efficiency. The possible explanation for this behavior is that the solution capacity changes were effected by varying the degree of stripping of the lean solution. As can be seen, attempting to increase the solution-carrying capacity resulted in a steam-consumption increase.

L. Mahmoodi, P. Darvishi / International Journal of Greenhouse Gas Control 58 (2017) 42–51

Fig. 11. Effect of steam temperature on efficiency.

4.2.5. Steam temperature The energy supplied by steam to the reboiler is the most important component of the operating cost and is strongly influenced by the process design. Fig. 11 demonstrates the effect of steam temperature on the reboiler thermal efficiency. It can be seen that a little increase in steam temperature leads to a significant increase in regeneration efficiency, because steam temperature has a great effect on the amount of heat added to the column to attain the desired separation. 5. Conclusion A rigorous rate model based on penetration theory has been presented to investigating the performance of an industrial regeneration tower for stripping of carbon dioxide from hot promoted potassium carbonate solution. It integrates a series of highly nonlinear differential equations to couple the material and energy balances, VLE relations and chemical kinetics. The results indicate acceptable compatibility with industrial data which validates our assumption. Obtained results from numerical solution and investigated variables effects on the tower’s performance shows increasing stripping efficiency by changing rich solution temperature and amine concentration. Moreover, entering solution temperature affects rate constant of reaction in liquid phase and the amount of water vapor condensation. Furthermore, adding amine to potassium carbonate solution causes a decrease in mass transfer resistance in liquid phase. The model provides a procedure to investigate the effect of new promoters on stripping performance. References Alatiqi, Imad M., Dadkhah, Ali A., Akbar, Ali M., Hamouda, M.F., 1994. Comparison between dynamics and control performance of mesophilic and thermophilic anaerobic sludge digesters. Chem. Eng. J. Biochem. Eng. J. 55, 55–66. Alié, C., Ferauche, F., Léonard, A., Lambert, S., Tcherkassova, N., Heinrichs, B., Crine, M., Marchot, P., Loukine, E., Pirard, J.P., 2006. Pd–Ag/SiO2 xerogel catalyst forming by impregnation on alumina foams. Chem. Eng. J. 117, 13–22. Aroonwilas, Adisorn, Chakma, Amit, Tontiwachwuthikul, Paitoon, Veawab, Amornvadee, 2003. Mathematical modeling of mass-transfer and hydrodynamics in CO2 absorbers packed with structured packings. Chem. Eng. Sci. 58, 4037–4053. Astria, G., 1983. A Gas Treating with Chemical Solvent. Wiley, New York. Banks, W.P., 1967. Corrosion in hot potassium carbonate system. Mater. Prot. 6, 37–41. Bienstock, D., Field, J.H., 1961. Corrosion inhibitors for hot-carbonate systems. Corrosion 17, 571–574. Bocard, G.P., Maryland, B.J., 1962. Investigation of the equilibrium between vapor, carbon dioxide and hot potassium carbonate solutions. Hydro. Proc. Pet. Ref. 41, 128–135.

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