Article pubs.acs.org/IECR
Industrial Applied and Modeling Research on Selective H2S Removal Using a Rotating Packed Bed Zhi Qian,† Zhen-Hu Li,‡ and Kai Guo*,§ †
College of Resources and Environment, Graduate University of Chinese Academy of Sciences, Beijing 100049, People's Republic of China ‡ Beijing Research Institute of Chemical Industry, SINOPEC, Beijing 100013, People's Republic of China § Research Center of the Ministry of Education for High Gravity Engineering and Technology, Beijing University of Chemical Technology, Beijing 100029, People's Republic of China ABSTRACT: For the simultaneous absorption of H2S and CO2 into methyldiethanolamine (MDEA) solution, MDEA in fact is selective toward both H2S and CO2. It is kinetically selective toward H2S and thermodynamically selective toward CO2. Selective H2S removal using a rotating packed bed (RPB) solves the problem that a large amount of CO2 accompanying H2S is simultaneously removed in the conventional desulfurization process. An RPB highlights that MDEA is kinetically selective toward H2S and restricts its thermodynamically selectivity toward CO2. In this work, an industrial test for MDEA selective absorption of H2S was performed. Compared with a conventional tower the RPB has a substantial advantage in selective absorption of H2S. A reaction−equilibrium−mass transfer model based on penetration theory is developed to describe the selective absorption process and the inhibition effect on H2S absorption from the CO2−MDEA reaction in the liquid film.
1. INTRODUCTION Both hydrogen sulfide (H2S) and carbon dioxide (CO2) are acidic. The considerable quantity of CO2 accompanying H2S is simultaneously removed in a conventional desulfurization process, which seriously influences the standard desulfurization process, increasing the circulation volume of the absorption liquid, and also sharply aggravates the energy load of the solution regeneration system. The desorbed CO2 dilutes the concentration of H2S in the desorbed acid gas and subsequently inhibits the efficiency of sulfur recovery of the Claus unit, eventually leading to massive CO2 emission into the atmosphere, polluting the environment. Therefore, interest in H2S selective absorption is increasing for the gas desulfurization process in refining and petrochemical enterprises.1,2 The development of H2S selective removal technology may bring significant social and economic benefits. Commercially used alkanolamines for the approach are monoethanolamine (MEA), diethanolamine (DEA), diisopropanolamine (DIPA), methyldiethanolamine (MDEA), and 2amino-2-methyl-1-propanol (AMP). Among these alkanolamines, MDEA as an absorption solvent of acid gases is widely used today because it possesses characteristics such as higher H2S selectivity, bigger absorption capacity, lower regeneration energy, smaller hot degradation, and less corrosion.3 A packed tower and a valve tower are typically employed as the desulfurization equipment of choice for most refinery enterprises. An MDEA solution directly contacts a sulfur-bearing gas countercurrent in a tower, following which a rich solution enters a desorption tower and is desorbed at a temperature of 120 °C for solvent recycling. The desorbed acid gas is converted into highpurity sulfur by the Claus process. The conventional selective desulfurization process is based on the selectivity of the solvent toward H2S due to the difference in reaction rates between H2S− MDEA and CO2−MDEA. The longer gas−liquid contact time within conventional equipment inhibits MDEA selectivity © 2012 American Chemical Society
toward H2S and promotes the absorption of CO2, resulting in a CO2 coabsorption rate as high as 79.9%.2 In addition, existing desulfurization units have the problem of large volumes, high investment, and low H2S removal efficiency in later stages. The selective treating process requires multiple stages and minimum contact time to avoid any unwanted time-dependent reaction. Compared with a conventional contactor, a rotating packed bed (RPB) is ideal for process operations that need both a “short” contactor and multiple stages.4 The reaction of H2S with MDEA is essentially instantaneous while that of CO2 with MDEA is relatively slow; therefore, a high selectivity for H2S is achievable with an RPB. As shown in Figure 1, the RPB
Figure 1. A rotating packed bed (RPB).
developed in the 1980s is a novel highly efficient gas−liquid reactor5−15 which utilizes centrifugal acceleration to intensify mass transfer16,17 that has been applied to absorption, distillation, polymer devolatilization, and reactive crystallization.4,18−20 According to previous studies, an RPB has a higher gas−liquid Received: Revised: Accepted: Published: 8108
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Figure 2. Desulfurization process flow sheet.
was varied from 200 to 600 rpm and at 400 rpm for industrial production. Figure 3 shows an on-site photograph of the RPB desulfurization process. The packed tower located beside the
mass transfer efficiency which is demonstrated by the fact that the volumetric mass transfer coefficients within an RPB are of an order of magnitude higher than those in a conventional packed bed. The selective absorption of H2S by MDEA in an RPB both strengthens the MDEA selectivity toward H2S at reaction kinetics and highlights the equipment’s inherent selectivity toward H2S in time. It is thus an innovative and cutting-edge technology. There are reports concerning the selective removal of H2S in aqueous solutions of MDEA in RPB.4,21 It has been shown that an RPB with a shorter gas−liquid contact time is an ideal reactor for selective removal of H2S. However, an effective industrial application and fundamental research based on the model remain to be introduced. In this study, an industrial test and an application were performed, and a comparison between the RPB and a conventional tower was made. Simultaneously, the corresponding reaction−equilibrium−mass transfer model was established for quantitative analysis, which mainly concerned the inhibitory effect on H2 S absorption caused by MDEA thermodynamic selectivity toward CO2.
2. FLOW SHEET OF THE DESULFURIZATION PROCESS USING RPB Figure 2 presents the flow sheet of the RPB desulfurization process for FCC dry gas purification. An RPB was located in an upper desulfuration tower. Through an oil−gas separator, dry gas entered an RPB and contacted a poor solution countercurrent. The sweet gas left the RPB and went into a clear dry gas pipeline. An MDEA solution was fed into the RPB by a poor solution pump and evenly distributed on the surface of the inner edge of packing by the liquid distributor. The liquid sharply flowed and was dispersed toward the outer edge of the packing under centrifugal force, thereby contacting a gas countercurrent. From the RPB the MDEA-rich solution was conveyed into a liquid level control tank, entering the desorption process for recycling. The MDEA solution went through the desulfurization tower and left from the bottom, before being conveyed into the RPB by a pump. The poor liquid went directly into the RPB without any passing tower, which was just used as a dry gas channel for this work. The flow rates of refinery dry gas ranged between 8000 and 13 000 Nm3/h, the liquid flow rate was 21 t/h, H2S content in feed gas ranged from 10 000 to 15 000 mg/Nm3, and CO2 content in feed gas was approximately 4%. The volume of packing for an RPB was about 2.2% that for a packed tower. The rotating speed
Figure 3. A desulfurization site.
RPB stopped operating for 2 years. The great difference between the RPB and the conventional packed tower in volume and height is evident in the photograph. After desulfurization by RPB, the H2S content in sweet gas was less than 20 mg/Nm3 and the coabsorption rate of CO2 was approximately 8.9%. Industrial test research demonstrated that RPB desulfurization technology had the merits of high desulfurization efficiency, high selectivity toward H2S, small packing and equipment volumes, and low investment and energy consumption. Compared with towers, the volume mass transfer coefficient within RPBs increases 1−2 orders of magnitude, and the volume and weight of equipment are fractions of those in towers.
3. MODEL DEVELOPMENT The model established in this paper mainly concerns the influence on the reaction thermodynamic stability of H2S− MDEA from the CO2−MDEA reaction, and the reaction− equilibrium−mass transfer model, which is based on three assumptions presented by Qian,22 is developed in this section. The rotor in an RPB consists of 240 layers of packing for purposes of this study. 8109
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3.1. Reactions of H2S and CO2 in Aqueous MDEA Solutions. When H2S and CO2 are absorbed into an aqueous solution of MDEA, several equilibrium reactions occur in the solution, which are as follows: k 2,MDEA , K1
CO2 + R3N + H 2O ←⎯⎯⎯⎯⎯⎯⎯→ R3NH+ + HCO3− + K2
R3N + H ↔ R3NH
+
K4 =
CO2 + H 2O ↔ HCO3− + H+ K4
k OH− , K5
(5)
K6
H 2O ↔ H+ + OH−
(6)
The reaction between H2S and aqueous amines involves proton transfer and is instantaneously fast with respect to mass transfer.23 Everywhere in the liquid phase, including the interfacial liquid film, H2S−MDEA equilibrium exists always. R3N + H 2S ↔ R3NH+ + HS−
3
x8 = c H0 3S ,
K 2b1x 7 + x7 − 1 + K 2x 7 −
With the H2S loading of aqueous MDEA solution, αH2S, in kmol of H2S/kmol of MDEA, the overall H2S balance is (10)
The charge balance is 3
(11)
In the reaction scheme of reactions 2−6, four equilibrium constants, K2, K3, K4, and K6, are independent. The other can be obtained by an appropriate combination of these independent equilibrium constants. The equilibrium constants of reactions 2−4, 6, and 7 are as follows: K2 =
K3 =
c R0 2NH+ c R03Nc H0+
(12)
0 0 −c + c HCO 3 H 0 cCO 2
x1 = c R03NH+ ,
b2 = c R3N,totalα , 0 , x3 = cCO 2
0 −, x4 = c HCO 3
0 −, x6 = cOH
x 7 = c H0+ ,
0 − x 9 = c HS
K6 =0 x7
b2 x7 K3
+1+
K4 x7
−2
K4 x7
b2 x7 K3
+1+
K4 x7
(17)
Equation 17 was solved by the dichotomy programmed with FORTRAN for x7, and the error was within the last step length, where x7′ was the solution. Other data were obtained by inversion. 3.3. Simultaneous Absorption of H2S and CO2 into MDEA Solutions in an RPB. When CO2 dissolves in a solvent, it binds chemically to the amine at finite rates of reaction, forming reaction products which are stable, requiring heat and stripping vapor to decompose them and reverse the reaction. When H2S dissolves in an amine, it converts immediately to HS− via basic instantaneous proton transfer, a protonation reaction which is immediately reversible. Although CO2 reacts relatively slowly and H2S reacts rapidly, the stable products formed from the CO2−MDEA reaction will influence the reaction equilibrium of H2S and MDEA. If the gas−liquid contact time is sufficient, the selectivity of MDEA toward CO2 appears. The stable reaction products of CO2 and MDEA weaken the alkalinity of the solution and shift the equilibrium of H2S and MDEA toward the reactant direction. The concentration of free H2S in liquid film increases accordingly, which leads to a decline of the H2S mass transfer rate, and the absorption of H2S is inhibited. The reaction between CO2 and MDEA generates stable reaction products which are scarcely decomposed unless there is stripping at high temperature, a reaction which is essentially an irreversible process.26 The reaction between H2S and MDEA is a reversible reaction. Higbie’s penetration theory is used to set up
(9)
0 0 0 − + c c R03NH+ + c H0+ = cOH HCO3− + 2cCO 2 −
(16)
An implicit function for x7 was obtained by solving eqs 8−16:
With the CO2 loading of an aqueous MDEA solution, αCO2, in kmol of CO2/kmol of MDEA, the overall CO2 balance is
0 − c R3N,totalαH2S = c H0 2S + c HS
c H0 2Sc R03N
0 x5 = cCO 2− ,
(8)
3
0 0 −c c HS R3NH+
x 2 = c R03N ,
3.2. Bulk Liquid Equilibrium Model. The initial liquid bulk concentrations of all chemical species are estimated from the total concentration of MDEA, the CO2 loading of the MDEA solution, and the assumption that all reactions are at equilibrium at the liquid outlet. The following equations are utilized for the liquid bulk concentrations c0R3NH+, c0R3N, c0CO2, c0HCO3−, c0CO32−, c0OH−, c0H+, cH0 2S, and c0HS−. The overall balance for MDEA is
0 0 − + c0 c R3N,totalαCO2 = cCO + c HCO CO 2 − 2 3
(15)
b1 = c R3N,total ,
(7)
c R3N,total = c R03NH+ + c R03N
0 − K 6 = c H0+cOH
The nine algebraic equations, eqs 8−16, are solved for the concentrations in the bulk liquid at the liquid outlet: c0R3NH+, cR0 3N, c0CO2, c0HCO3−, c0CO32−, c0OH−, c0H+, cH0 2S, and c0HS−. The Newton method is typically used, but it is complicated due to error propagation. In this study, the dichotomy is applied to solve the equations. We introduce an implicit function with the unknown c0H+, which ranges from 0 to 1 mol/L determined by the pH indicator. The equilibrium constants required to solve the mathematical model are well-known, as discussed by Zhang24 and Austgen.25 For convenience, the concentrations of chemical species are renamed as follows:
(4)
CO2 + OH− ←⎯⎯⎯⎯→ HCO3−
(14)
K7 =
(3)
HCO3− ↔ H+ + CO32 −
3
0 − c HCO 3
(1) (2)
K3
0 0 cCO 2 −c H+
(13) 8110
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distribution of nine chemical species varying with time and penetration depth in liquid film. Initial and Boundary Conditions for the Model. At t = 0 (x ≥ 0) and x → ∞ (t ≥ 0), the concentrations of all chemical species are equal to their liquid bulk concentrations:
the reaction−equilibrium−mass transfer partial differential equations. Among reactions 1−6, reactions 1, 3, and 5 have influence on the absorption of CO2. However, it is well-known that the rate of reaction 3 is very slow and may usually be neglected.26 In this study, it is assumed that there is a fast reaction of CO2 and OH− in parallel, i.e., reaction 5, with another rapid pseudo-first-order reaction between CO2 and MDEA, i.e., reaction 1. They are following the base catalysis mechanism27 and zwitterion mechanism,28 respectively. Therefore, the CO2 balance can be written as ∂cCO2 ∂t
ci = ci0 , i = R3NH−, R3N, CO2 , HCO3− , CO32 − , OH−, H+ , H 2S, HS−
At x = 0 (gas−liquid interface), the flux of the nonvolatile chemical species is equal to zero, which leads to the following equation:
2
= DCO2
∂ cCO2 ∂x 2
− (k CO2 − MDEAc MDEA
+ k CO2 − OH−cOH−)cCO2
∂ci =0 ∂x
(18)
The total carbon dioxide balance is ∂cCO2 ∂t
+
∂c
= DCO2
HCO3−
∂t ∂ 2cCO2 ∂x
2
+
∂cCO3 ∂t
∂ 2c HCO3− ∂x
2
+ DCO32−
∂ 2cCO32− ∂x
−Di
2
∂t
∂c R3N ∂t
= DR3NH+
∂ 2c R3NH+ ∂x
2
+ DR3N
∂ci = kg, i[Pi − Hici(0, t )] ∂t
∂ 2c R3N ∂x
Pi −
2
ci(0, t ) =
∂c HCO3− ∂cCO32− ∂c − ∂c H+ − OH − −2 ∂t ∂t ∂t ∂t ∂t 2 2 + ∂ c ∂c − ∂ c H+ R3NH − HS = DR3NH+ + D H+ 2 ∂t ∂x ∂x 2 2 ∂ 2c HCO3− ∂ cOH− − − DOH− − D HCO3 ∂x 2 ∂x 2 ∂ 2cCO32− ∂ 2c HS− − − 2DCO32− − D HS ∂x 2 ∂t 2 +
ci(0, t ) = ci* =
(21)
(22)
For instantaneous reactions assumed to be at equilibrium c R NH+ K2 = 3 c R3Nc H+ (23) cCO32−c H+ c HCO3−
(24)
K 6 = c H+cOH−
(25)
K7 =
c HS−c R3NH+ c H2Sc R3N
t>0
x = 0,
t>0
(30)
Pi Hi
x = 0,
t>0
(31)
For a calculation result getting closer to a practical situation, the boundary condition near the interface is improved in this study. Both gas and liquid film resistances are significant for H2S absorption,29,30 so the boundary condition is accordingly modified. The iteration method is employed to approach the real partial pressure of H2S at the interface. First, the value of eq 31 as an initial value is substituted into the model to obtain the concentration distribution of H2S in the liquid film (cH2S). Then cH2S is substituted into eq 30 to obtain cH2S(0,t). Finally, comparing cH2S(0,t) with the initial value, if in agreement, the initial value is regarded as the real partial pressure of H2S at the interface. If not, cH2S(0,t) as a new initial value is calculated again as per the above procedure until convergence. 3.4. Mass Transfer Coefficient and Mass Balance in an RPB. For brevity, for the mass transfer coefficient and mass balance for CO2, please refer to the previous study.31 For H2S absorption, both gas and liquid film resistances are critical.29,30 A prior study reported that the values of the gas-side mass transfer coefficient (kG) in an RPB lie in a range similar to those in a static absorption plant.32 Hence, the gas-phase resistance to mass transfer was determined by experiments
According to reaction 7, the H2S balance is
K4 =
x = 0,
For simplifying the calculation and reducing the complexity of the model, Mandal29 suggested that the diffusion process of gas i on the gas−liquid interface might be regarded as pure gas i diffusion. For the case of pure gas i in the gas phase, the interfacial partial pressure of the gas, P*i , is the same as the bulk partial pressure of the gas, Pi, and there is no mass transfer resistance in the gas phase. Hence, the boundary condition for gas i at the gas− liquid interface is simplified as
The electroneutrality balance is given by
∂ 2c H2S ∂c H2S ∂ 2c HS− ∂c HS− = DHS− + + D H 2S ∂t ∂t ∂x 2 ∂x 2
−Di ∂ci kg, i ∂t
Hi
(20)
∂c R3NH+
(28)
Therefore, the real partial pressure (or concentration) of component i at interface is obtained:
The total MDEA balance is +
t>0
(29)
(19)
∂c R3NH+
x = 0,
for all i except i = CO2 and H2S. For the volatile components CO2 and H2S, the mass transfer rate in the gas near the interface is equal to the mass transfer rate in the liquid near the interface:
2−
+ DHCO3−
(27)
(26)
There are nine partial differential−algebraic equations for the MDEA solvent, which are solved for the concentration 8111
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to solve the partial differential equations such that the concentration distributions of H2S and MDEA in liquid film are obtained. The program converges within 30 min. Based on the numerical solution to partial differential equations of the reaction−equilibrium−mass transfer coupled model, concentration distributions of H2S in liquid film at a penetration time of 5.0 × 10−11 s for the absorption of CO2 and H2S with 30% MDEA at rotating speeds of 200 rpm (case a) and 450 rpm (case b) are obtained. Comparing case a and case b, the penetration depth of case a is somewhat deeper than that of case b. This is mainly caused by the thermodynamic selectivity of MDEA toward CO2. The stable products, R3NH + and HCO3−, formed from the reaction of CO2 and MDEA shift the equilibrium of H2S and MDEA toward the reactant direction. The “R3NH+ + HS−” system is unstable, and HS− readily deprotonates as the H2S starts to desorb. Therefore, the H2S cannot be rapidly consumed at the surface of the liquid film and gradually penetrates the liquid film. The absorption rate of H2S declines accordingly. This decline is not so apparent compared with the overall mass transfer rate. According to quantitative analysis, the difference of the penetration depth between the two operation conditions is approximately 5 × 10−10 m. However, this marginal difference is enough to influence the final gas treatment effect and determines whether the outlet gas concentration meets the standard. The effect of the CO2− MDEA reaction on the selective absorption of H2S is not apparent at the total amount of mass transfer, but it is essential for the high H2S removal efficiency required during the desulfurization process. For validating the conclusion based on the reaction− equilibrium−mass transfer model above, more industrial experiments with the simultaneous absorption of H2S and CO2 into MDEA solutions in an RPB were performed. MDEA aqueous solutions with mass concentrations from 10 to 30% and a liquid flow rate of 21 t/h entered the RPB and contacted the gas−mixed stream countercurrent at a gas flow rate of 11 000 N m3/h. Figure 4 depicts the dependence of Ka on the rotating speed. The KGα
concerning the absorption of H2S into 10% aqueous hydroxide solutions from a 1% H2S−99% N2 gas mixture as described by Savage et al.33 The gas-phase mass transfer coefficient (KG,H2S) of RPB for H2S absorption is determined by the following equation: 1 K G,H2S
=
1 k G,H2S
+
H H 2S kL,H2S
(32)
Considering an element of the RPB with radial length dr, the differential mass balance equation for H2S is given as 8.314 TPα(yH S − yH* S )(2π )hR dR 2 2 101.325 ⎞ ⎛ y H 2S ⎟ = G N2 d⎜⎜ ⎟ ⎝ 1 − yH2S ⎠
K G,H2S
(33)
where yi and y*i respectively represent the mole fraction of component i in the gas bulk and that of component i in equilibrium with the bulk liquid. P denotes total pressure. For the reaction between H2S and MDEA, the equilibrium mole fraction in the gas phase y*H2S is not negligible because of reversible reaction. The equilibrium concentration of H2S cH2S,eq is calculated from reaction 7 and written as c H2S,eq =
c HS−(c R3N,total − c R3N) K 7c R3N
(34)
The overall transfer balance for H2S is given by ⎡ ⎤ H c HS−(c R3N,total − c R3N) ⎥ K G,H2SR gTα⎢yH S − (2π )hR dR P K 7c R3N ⎢⎣ 2 ⎥⎦ ⎞ ⎛ y H 2S ⎟ = G N d⎜⎜ ⎟ ⎝ 1 − yH2S ⎠
(35)
⎡ ⎤ H c HS−(c R3N,total − c R3N) ⎥ K G,H2SαP ⎢yH S − (2π )hR dR P K 7c R3N ⎢⎣ 2 ⎥⎦ = Q d(c HS−)
(36)
⎡ ⎤ H c HS−(c R3N,total − c R3N) ⎥ (2π )hR dR −K G,H2SαP ⎢yH S − P K 7c R3N ⎢⎣ 2 ⎥⎦ = Q d(c R3N)
(37)
There are three differential equations for H2S mass transfer, which are solved for the three unknown functions. The variation of yH2S in the radial direction of packing and the final gas concentration at the outlet are obtained. The Runge−Kutta− Fehlberg method is employed to solve the model equations as discussed by Qian.22 3.5. Solution Methodology. The finite element method (FEM) based on the variation principle is employed to numerically solve the diffusion−reaction partial differential equations. The FEM is a powerful technique for finding the approximate solution of a partial differential equation where the domain boundaries of the given problem are so complex that standard methods such as the finite difference method (FDM) may have difficulties or fail. A procedure is compiled in Flex PDE
Figure 4. Effects of rotating speed on the overall volumetric mass transfer coefficient (KGα).
for H2S absorption rises with an increase in the rotating speed in the range from 200 to 450 rpm, while the Ka for CO2 absorption has the opposite trend. The residence time increases when the rotating speed declines and more CO2 reacts with MDEA, which exerts a negative impact on H 2 S. MDEA gradually is thermodynamically selective toward CO2 with the decline in rotating speed. By contrast, the residence time declines when the rotating speed increases and there is less likelihood of CO2 8112
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Figures 5 and 6 above, it is evident that the RPB provides a substantial desulfurization effect and the operation is steady. The industrial test was performed under operation conditions of a gas flow rate varying from 8000 to 12 500 Nm3/h, H2S content in feed gas of 12 000 mg/Nm3, a rotating speed of 400 rpm, a liquid flow rate of 21 t/h, and an MDEA mass percentage of 30%, and the experimental points of the H2S content in desulfurized gas are shown in Figure 7. These experimental
reacting with MDEA. MDEA gradually is kinetically selective toward H2S with the increase in rotating speed. The two points marked by two circles in Figure 4 are the experimental points which are simulated by the reaction−equilibrium−mass transfer model. Point a indicates that more CO2 is participating in the reaction with MDEA because of the relatively longer gas−liquid contact time. The penetration depth of H2S in liquid film is deeper. Point b indicates that H2S is rapidly relatively exhausted in liquid film because of the relatively shorter gas−liquid contact time. The absorption of CO2 is restrained, and the penetration depth of H2S in liquid film is shallower.
4. INDUSTRIAL APPLICATION RESULTS AND DISCUSSION 4.1. Experimental Values and Model Predicted Values. The operation conditions for industrial application are a gas flow rate of 8000 Nm3/h, a rotating speed of 400 rpm, a liquid flow rate of 21 t/h, and an MDEA mass percentage of 30%. Figure 5
Figure 7. Experimental vs predicted values.
points are simulated with the model developed in this study, and the predicted result is shown as the line in Figure 7. The results show that the change trend of experimental points is consistent with that of the calculated points and the average error between the experimental and predicted values is 29%. The reason the positive deviations occurred in the predicted values is that the gas resistance for H2S absorption may have been slightly overestimated. 4.2. Laboratory Scale, Pilot Scale, and Industrial Scale Test Results. The RPB employed in the laboratory has an inner diameter of 42 mm, an outer diameter of 146 mm, and a height of 20 mm, as shown in Figure 8. The experiment was performed
Figure 5. Distribution of H2S content in FCC dry gas.
shows the distribution of H2S content in feed gas, which ranges from 10 000 to 15 000 mg/Nm3. Figure 6 shows that the content
Figure 6. Distribution of H2S content in desulfurized gas. Figure 8. Photograph of RPB in the laboratory.
of H2S in desulfurized dry gas is steadily under 20 mg/Nm3. At the beginning of the test, the adding of another stream of dry gas leads to an actual gas flow rate greater than the design value of 8000 Nm3/h, which causes the H2S concentration in desulfurized gas to fall below the standard at several experimental points. At the latter stage of the test, the H2S concentration is high because of the slipping of the belt which slows down the rotating speed of the RPB leading to the enhancement of the CO2 absorption rate and the restraint of H2S removal. From the comparison between
with an operation condition of a gas flow rate varying from 600 to 1100 L/h, a liquid flow rate from 4 to 9 L/h, a rotating speed from 600 to 1300 rpm, and an MDEA mass percentage of 30%. For the pilot test, the RPB had an inner diameter of 200 mm, an outer diameter of 600 mm, and a height of 220 mm. The test was performed with an operation condition of a gas flow rate varying from 800 to 2800 m3/h, a liquid flow rate of 10 m3/h, a rotating speed from 300 to 500 rpm, and an MDEA mass percentage of 8113
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30%. A photograph of the pilot test site is shown in Figure 9. Laboratory, pilot, and industrial data were simulated with the
Table 2. An RPB vs a Conventional Packed Tower RPB
Figure 9. Pilot test site.
model established in this study, and the experimental and predicted values were consistent. The comparison among laboratory, pilot, and industrial data is presented in Table 1. It
feed gas of CO2 feed gas of H2S gas−liquid ratio coabsorption rate of CO2 H2S content after absorption model calculation
pilot
industrial application
4−8% 10 000 mg/ Nm3 100−260 9−16%
6−9% 6000−10 000 mg/Nm3 80−280 10−12%
3−6% 10 000−15 000 mg/Nm3 380−600 8.5−10%
15−30 mg/ Nm3 error within 15%31
10−30 mg/Nm3
