Gas-Lift Reactor for Hydrogen Sulfide Removal - ACS Publications

Apr 14, 2005 - The internal loop gas lift or the autocirculation reactor is an appropriate reactor for the Lo-Cat process because it combines the reac...
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Ind. Eng. Chem. Res. 2005, 44, 6115-6122

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Gas-Lift Reactor for Hydrogen Sulfide Removal Sunun Limtrakul,*,† Sudtida Rojanamatin,† Terdthai Vatanatham,† and Palghat A. Ramachandran‡ Department of Chemical Engineering, Faculty of Engineering, Kasetsart University, Bangkok 10900, Thailand, and Department of Chemical Engineering, Washington University, St. Louis, Missouri 63130

Hydrogen sulfide (H2S) removal by the Lo-Cat process [reaction with iron(3+) chelates] consists of two steps: the removal reaction and catalyst regeneration. The internal loop gas lift or the autocirculation reactor is an appropriate reactor for the Lo-Cat process because it combines the reaction and regeneration steps in a single device. The removal reaction takes place in the riser, and the used catalyst solution is regenerated in the downcomer. This research involved the study of the performance of the Lo-Cat air-lift reactor on H2S removal under various operating conditions. H2S could be removed completely (>99.9%) when the superficial H2S velocity in the riser was 0.03 m/s and the superficial air velocity in the downcomer was 0.01 m/s. Iron(2+) ethylenediaminetetraacetic acid [Fe2+EDTA] was regenerated to Fe3+EDTA in the downcomer with a conversion of up to 79%. The conversions of H2S, Fe3+EDTA, and Fe2+EDTA decreased with increasing superficial gas velocity in the riser for a fixed gas velocity in the downcomer. This was due to the decrease in the liquid-phase circulation time in the riser and downcomer with an increase in the superficial gas velocity in the riser at a constant downcomer velocity. A mixing cell model for gas and liquid phases was also developed to describe the H2S degradation behavior of the reactor, and the model was found to be consistent with the measured data. 1. Introduction Lo-Cat is a gas-sweetening process that removes hydrogen sulfide (H2S) from gas streams and converts it to solid sulfur.1 In this process, a chelated iron(3+) ethylenediaminetetraacetic acid [Fe3+EDTA] solution is used as a catalyst for converting H2S to elemental sulfur. The regeneration of the reduced chelated Fe2+EDTA is carried out by using air. The Lo-Cat process promises a high sulfur recovery efficiency, often up to 99.9%.2 A suitable reactor for the Lo-Cat process is a modified gas-lift reactor. Gas-lift reactors are commonly used in bioprocessing and other fields. A gas-lift reactor consists of a riser, a gas-liquid separator, and a downcomer.3 Liquid circulation is induced by the density difference between the riser and downcomer usually by injecting gas at the bottom of the riser. The gas-lift reactor is economical and has no moving parts. The Lo-Cat process involves two gaseous streams, H2S and air; therefore, the gas-lift reactor cannot be used in its conventional form, where the gas is injected only at the base of the riser zone. Hence, for the Lo-Cat application, the gaslift reactor is modified by injecting H2S at the bottom of the riser, where the reaction occurs, and injecting air in the downcomer for catalyst regeneration or vice versa. Thus, both reactions (gas removal and regeneration) can be carried out in the same vessel, thereby reducing the capital costs. The gas holdup in the riser is more than that in the downcomer, and the resulting density difference creates a natural liquid circulation from the downcomer section of the vessel to the riser section. Thus, liquid circulation is established without the need * To whom correspondence should be addressed. Tel.: +662942-8555 ext. 1210. Fax: +66-2942-8555 ext. 1232. E-mail: [email protected]. † Kasetsart University. ‡ Washington University.

of pumps,4 resulting in savings in the operating costs. Gas Technology Product LLC has developed a commercial Lo-Cat process in which the regeneration air is sparged into the riser with considerably more air velocity than H2S velocity in the downcomer.5 Additional details on the modes of operations of autocirculation reactors are available in the work of Hardison2 and in the patent of Hardison.6 This research work involves the study of the performance of the Lo-Cat gas-lift reactor and its operating conditions on the H2S removal efficiency in an experimental small-scale reactor. In addition, a mixing cell model for gas and liquid phases is developed to describe the behavior of the reactor for design purposes. The model uses the available hydrodynamic parameters for mass transfer (see, for instance, the work by Chisti3). However, it may be noted that the published hydrodynamic models are for the conventional air-lift type of reactors (the operation with no gas flow in the downcomer). These are not strictly applicable for the gas-lift reactor with gas flow in both the riser and downcomer, and more studies on the hydrodynamics are clearly needed. Despite these limitations, the modeling exercise provides a basic understanding of the effect of operating parameters on the gas removal efficiency. 2. Experimental Methods 2.1. Apparatus. The gas-lift reactor consisted of two concentric pipes of diameters of 0.108 m (steel) and 0.15 m [poly(vinyl chloride)]. The inside pipe serves as the riser, while the annular space serves as the downcomer (see Figure 1). The heights of the riser and downcomer were 0.8 and 1.1 m, respectively. The H2S sparger in the riser was a single nozzle with a diameter of 25 mm. The downcomer air sparger was a ring sparger made from a perforated pipe with six equally spaced holes of 1 mm diameter (Figure 2).

10.1021/ie049166p CCC: $30.25 © 2005 American Chemical Society Published on Web 04/14/2005

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Figure 3. Axial concentration profiles of Fe3+EDTA in the riser and downcomer at U0gr ) 0.03 m/s and U0gd ) 0.01 m/s.

Figure 1. Gas-lift reactor.

Figure 2. Gas spargers.

The reactor was initially filled with a batch of a chelated Fe3+EDTA solution. The riser was injected with a feed gas containing about 5% H2S. H2S absorption and reactions took place in the riser as follows:

H2S(g) + 2Fe3+ f 2H+ + S0 + 2Fe2+ The downcomer was injected with air. Oxidation of Fe2+ in the downcomer regenerated the catalyst as follows:

O2(g) + 4Fe2+ + 2H2O f 4OH - + 4Fe3+ To prevent contamination of H2S with headspace gas, there was a large tube on the top of the riser. This tube

could be adjusted to the level of a downward-flowing liquid. The superficial H2S velocity in the riser, U0gr, was varied from 0.03 to 0.11 m/s, while the superficial air velocity in the downcomer, U0gd, ranged from 0.01 to 0.03 m/s. The concentration of Fe3+EDTA used is in the range of 7-14 mol/m3, depending on the superficial H2S velocity in the riser. The pH of the reactor was maintained in the range of 6.0-8.0 using acetate and tris(hydroxymethyl)aminomethane buffers.7 Experiments were carried out at room temperature and atmospheric pressure. 2.2. Measurement Methods. The concentration profile of the chelated iron catalyst solution was measured in the riser and downcomer at five axial positions. In the downcomer, liquid samples were drawn from the five sampling ports spaced along the column wall, as shown in Figure 1. In the riser, the samples were withdrawn by a level-adjustable sampling pipe inserted from the top. The liquid-phase concentration of chelated iron was measured using atomic absorption (AA) and titration with KMnO4. The AA method measures the concentration of total iron (Fe2+ and Fe3+), and the method of titration with KMnO4 measures only the Fe2+ concentration.8 The gas-phase concentration of H2S was measured at the inlet and outlet, using gas chromatography. 2.3. Results and Discussion. Figure 3 shows the axial concentration profiles of Fe3+EDTA in the riser and downcomer at superficial gas velocities in the riser and downcomer of 0.03 and 0.01 m/s, respectively. The concentration of Fe3+EDTA in the riser and downcomer decreases as the bed height increases. At the bottom of the riser, the liquid solution coming from the (regenerating) downcomer contains a high concentration of Fe3+EDTA. In the riser, this Fe3+EDTA solution flows concurrently upward and reacts with H2S gas. As expected, Fe3+EDTA keeps getting consumed with increasing height in the riser. However, near the surface of the liquid, the concentration of Fe3+EDTA increases slightly because of oxidation by the oxygen in the headspace. The spent chelated iron solution with a low concentration of Fe3+ then circulates downward into the downcomer and meets the injected air countercurrently. The Fe3+ concentration increases as this solution travels down because Fe2+ is regenerated. Figure 4 shows the concentration profiles of Fe3+EDTA as a function of time from the start-up of the reactor, for a fixed height of 0.55 m, in the riser and downcomer. The concentration in the riser decreases with time and finally reaches a steady-state value once the concentration profiles in both the riser and down-

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Figure 4. Concentration of Fe3+EDTA in the riser and downcomer at a height of 0.55 m as a function of time at U0gr ) 0.03 m/s and U0gd ) 0.01 m/s. Table 1. Conversion of H2S, Fe3+EDTA, and Fe3+EDTA in the Riser and Downcomer conversion (%) U0gr

(m/s)

U0gd

(m/s)

H2S riser

Fe3+ riser

Fe2+ downcomer

0.03 0.06 0.09 0.11 0.09 0.11

0.01 0.01 0.01 0.01 0.03 0.03

99.98 78.11 64.96 52.64 69.62 58.90

83.85 79.81 76.42 67.40 80.52 71.71

79.07 69.05 61.9 52.5 78.05 71.05

Figure 5. Schematic representation of the mixing cell model.

comer are established to the recirculation mode. The reactor takes about 15 min to reach steady-state conditions. Table 1 shows the conversion of H2S and Fe3+EDTA in the riser and the conversion of Fe2+EDTA in the downcomer. The conversion of H2S decreases with an increase in the superficial gas velocity in the riser because of the decreased contact time between H2S and the liquid. H2S is removed completely (>99.9% removal) at U0gr of 0.03 m/s and U0gd of 0.01 m/s. At any fixed value of the superficial aeration velocity in the downcomer, an increase in the superficial velocity in the riser increases the liquid circulation velocity in the reactor. This reduces the residence time of the liquid phase in both the riser and downcomer during any one circulation. Consequently, the conversions of Fe3+EDTA and Fe2+EDTA in the riser and downcomer, respectively, were lower at a higher riser superficial gas velocity. Under the best conditions tested, Fe2+EDTA was regenerated to Fe3+EDTA in the downcomer with a conversion of 79%. 3. Mixing Cell Model The mixing cell model is potentially useful to evaluate the performance of the reactions taking place in the riser and downcomer because the mixing characteristics of these zones are intermediate to the ideal situations of plug flow and backmixed flow.9 The mixing cell model divides the reaction zone into a series of connected wellmixed cells, as shown schematically in Figure 5. The values of n (i.e., the cell numbers) would be 1 and ∞ for well-mixed and plug-flow cases, respectively.10 Within each cell in the riser and downcomer, we have a bulk liquid phase, a bulk gas phase, and a liquid film, as shown in Figure 6. Gas reactants enter and dissolve in the liquid phase before they can react, and reaction occurs in the liquid-phase only. Gas and liquid flows in the riser are cocurrent. The flow pattern is countercur-

Figure 6. Three zones of mass balance in the mixing cell model, i.e., gas bulk, liquid film, and liquid bulk.

rent in the downcomer. The model further assumes the following: (a) a steady-state system; (b) isothermal operation; (c) negligible resistance to mass transfer in the gas phase; (d) well-mixed gas and liquid bulk; (e) reaction taking place only in the liquid phase (either liquid film or bulk); (f) nonvolatile liquid reactants and products; (g) constant volumetric flow rates. The material balances for the gas bulk, liquid bulk, differential equations describing the diffusion and reaction of the chemical species through the liquid film can be developed as follows.

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Figure 7. Control volume for the gas phase.

Figure 9. Control volume for the liquid film.

In the riser, the liquid-phase flows up, and the mass balance for i ) Fe3+ is

Vr ulrArCi(l)(j-1) - ulrArCi(l)(j) + agl Nfi(j)|x)δ ) 0 n

(3)

where ulr is the liquid circulation velocity in the riser and δ is the film thickness. In the downcomer, the liquid-phase flows down, and the mass balance for i ) Fe3+ is Figure 8. Control volume for the liquid bulk.

3.1. Mass Balances for the Gas Phase. The control volume for mass balances for the gas bulk is shown in Figure 7. The mass balances for H2S and O2 in the gas bulk of the mixing cell, j, are written as follows:

In the riser Vr f (j)|x)0 ) 0 QGrCH2S(g)(j-1) - QGrCH2S(g)(j) - agl NH 2S n (1) where the flux of H2S at the gas-liquid interface is f f NH ) -DH2S [dCH (j)/dx]x)0. This is obtained from 2S 2S(l) the film model described in section 3.3. The superscript f refers to film concentrations and x is the distance coordinate in the film measured from the gas-liquid interface. Also in eq 1, n is the number of mixing cells and Vr is the volume of the riser. Other variables are standard and are defined in the Notation section.

In the downcomer QGdCO2(g)(j-1) - QGdCO2(g)(j) - agl

Vd f )0 N (j)| n O2 x)0 (2)

f where the oxygen flux at the interface is NO ) -DO2 2 f [dCO (j)/dx] and V is the volume of the downx)0 d (l) 2 comer. 3.2. Mass Balance for the Liquid Phase. The mass balances in the liquid bulk for the dissolved gaseous and liquid reactants and the liquid products are based on the control volume depicted in Figure 8. From the values of the Hatta numbers (defined later), both the H2S removal reaction and catalyst regeneration confirm to the fast regime and the reaction is complete in the film itself. There is no accumulation of gaseous species in the bulk liquid, and no reaction occurs in the liquid bulk. Hence, the mass balance for the dissolved gas in the bulk liquid is not needed. Thus, the mass balances in the bulk liquid are needed only for the Fe species.

uldAdCi(l)(j) - uldAdCi(l)(j-1) + agl

Vd f N (j-1)|x)δ ) 0 n i (4)

where uld is the liquid circulation velocity in the downcomer. Also in the above equations, Nfi ) -Di f [dCi(l) (j)/dx]x)δ, which is the flux of the species into the bulk liquid, and this appears as a source term in the material balance for the bulk liquid. The values of Nfi into the bulk are related directly by stoichiometry to the interfacial fluxes of H2S/O2 because there is no bulk reaction. 3.3. Mass Balance in the Liquid Film. For the gaseous species diffusing through the liquid film, the steady-state mass balance based on the control volume shown in Figure 9 is given by

Di

f d2Ci(l) (j)

dx2

) -Ri(j)

i ) O2, H2S

(5)

where Ri is the rate of production of the ith species in cell j. The rate equation for the reaction of H2S catalyzed with chelated iron was formulated by DeBerry7 as follows:

-RH2S ) kH2SCH2SCFe3+ The rate equation for the reaction of oxygen with chelated iron has been given by Demmink and Beenackers11 as follows:

-RO2 ) kO2CO2CFe2+2 Two boundary conditions are required to solve eq 5. One of these boundary conditions is obtained by specifying the condition at the interface: f x ) 0, Ci(l) (j) )

Ci(g)(j) Hi

(6)

where Hi is the Henry’s law constant for species i expressed as the ratio of concentration in the gas and that in the liquid. The second boundary condition is

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obtained by specifying the concentration of the gaseous species in the liquid bulk (taken as zero here); thus f x ) δ, Ci(l) (j) ) 0

(7)

Equivalent equations for describing the diffusion and reaction of the liquid reactants and products in the liquid film are 2

Di

d

f Ci(l) (j)

dx2

) -Ri(l)(j)

(8)

where i ) Fe2+, Fe3+. The concentration condition of the liquid species in liquid bulk is defined as a boundary condition for eq 8 at x ) δ. The other condition is based on the fact that there are no fluxes of liquid reactants and products transferred across the gas-liquid interface. These conditions are stated as follows:

x ) 0,

f (j) dCi(l)

dx

)0

(9) (10)

The equations of mass balances for all components coupled with the corresponding boundary conditions can be solved for each cell. These boundary value problems can be numerically solved by the boundary element method.9 This approach is general and can also be used in conjunction with more detailed kinetic models for the two reactions. An alternate approach is to use simple approximate analytical solutions for the enhancement factor. The relevant equations are as follows: For the reaction A(g) + zB(l) f product with the rate ) kmnCAmCBn, the dimensionless interfacial concentration of B is given by

CfB(l)|x)0 CB(l)

)1+

1 φ q q

(11)

f DBCB(l)/zDACA(l) |x)0 and φ is the [defined as (Ni)x)0/kLC f(0)], which

where q ) enhancement factor is related to the “interfacial” Hatta number by the following equation:

φ)

Haxbin

(12)

tanh(Haxbi ) n

where Ha is the Hatta number defined as

Ha )

x

( )

CA(g) 2 DAakmn m+1 HA kL

riser

ref

(m3/mol‚s)

rate constant Henry’s law constant (kPa‚m3/mol) diffusivity of H2S (m2/s) diffusivity of Fe3+EDTA

9 1.95 1.44× 10-9 0.54× 10-9

downcomer

7 7 12 12 ref

rate constant (m6/mol2‚s) Henry’s law constant (kPa‚m3/mol) diffusivity of O2 (m2/s) diffusivity of Fe2+EDTA (m2/s)

38000 120.55 0.78× 10-9 0.54× 10-9

11 11 11 12

agree very closely. Further results shown in the paper are based on the analytical solutions. The general method (numerical) is useful if more detailed kinetics (e.g., multistep reactions) are to be included and is therefore formulated in this paper for completeness. This model requires several kinetic and physicochemical parameters as well as several hydrodynamic parameters. The parameters needed and the values used in the present simulation are discussed in the following sections. 4. Model Parameters

f (j) ) Ci(l)(j) x ) δ, Ci(l)

bi )

Table 2. Kinetic, Diffusion, and Solubility Parameters

m-1

CB(l)n

f f The flux terms NH and NO needed for bulk gas and 2S 2 liquid models are then directly calculated using the enhancement factor calculated from the above set of equations (11) and (12). Both approaches, exact numerical solution of the film model and approximate solution of the analytical model,

4.1. Kinetic, Diffusion, and Solubility Parameters. The kinetic rate constant, diffusion coefficient, and solubility parameters in the riser and downcomer are shown in Table 2, together with the reference from which they were obtained. The kinetic parameters vary over a range and depend on the type of chelate, ionic strength, pH, etc.; hence, reasonable estimates obtained from the sources indicated in Table 2 were used for model comparison. 4.2. Hydrodynamic Parameters. The main hydrodynamic parameters are gd and gr. In addition, the liquid recirculation velocity is an implicit variable that needs to be calculated (or measured) for given operating gas velocities in the riser and downcomer. These parameters as well as the number of mixing cells are expected to be functions of the relative slip velocity between the gas and liquid phases, defined as

ur,riser )

ugr ulr gr 1 - gr

ur,downcomer )

ugd uld + gd 1 - gd

(13)

(14)

Also, by mass balance, ulr and uld are related as follows:

ulrAr ) uldAd

(15)

A complete hydrodynamic model that predicts all of these simultaneously is not available. Hence, gas holdup in the downcomer, gd, was obtained from pressure-drop measurement, while gas holdup in the riser was calculated by the equation3

g )

Argr + Adgd Ar + Ad

(16)

where g is the measured total gas holdup obtained from bed-expansion measurement.

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Table 3. Gas Holdups and Circulation Velocities Used for the Simulation gas holdup U0gr (m/s)

U0gd (m/s)

total g

riser gr

downcomer gd

liquid velocity in the riser ulr (m/s)

0.03 0.06 0.09 0.11 0.09 0.11

0.01 0.01 0.01 0.01 0.03 0.03

0.059 0.087 0.106 0.128 0.128 0.141

0.0625 0.094 0.108 0.132 0.133 0.152

0.056 0.081 0.104 0.125 0.124 0.132

0.031 0.060 0.090 0.110 0.080 0.090

The liquid circulation velocity in the riser can be related to the gas holdup difference using the model for the gas-lift reactor having a single sparger:13

ulr )

x

2gHD(gr - gd) Ar 2 1 fB Ad (1 -  )2

()

klagl ) 0.5736Ug0.65g0.35(1 - g)0.65

Equation 18 was obtained for bubble columns at low gas velocities with no net flow of liquid and has been widely used in the literature for bubbly systems. Strictly speaking, it is not for the same conditions as the present work, but it is useful to get an estimate of the overall mass-transfer coefficient. Another approach in the absence of suitable correlations or actual data is to use estimates based on the bubble diameter for agl and the penetration model for kL. The relevant equations are as follows:

g agl ) 6 dB

(17)

gd

where fB is the friction loss coefficient for the connecting section. Equation 17 was obtained from a simple hydrodynamic model based on a macroscopic mechanical energy balance. The rate of energy input into the reactor is balanced by the rate of energy dissipation. The energy dissipation is due to wakes behind bubbles in the riser, stagnant gas in the downcomer, fluid turnaround at the bottom and top of the reactor, and friction in the riser and downcomer. The details of the model are shown by Chisti et al.13 and also in an earlier hydrodynamic model for air-lift reactors by Hsu and Dudukovic.14 The liquid circulation velocity can also be estimated from the measured data of the H2S concentration in the gas phase at the inlet and outlet of the riser and the measured data of change in the Fe3+EDTA concentration in the liquid phase by using mass balance, which is used in this study and indicated in Table 3. Liquid circulation velocities can also be directly measured using a colored pulse injection method. A detailed study and a model for liquid circulation velocities based on detailed hydrodynamic considerations will be published in a later study. Table 3 shows typical values of gas holdups and the values of ulr for the conditions used. These values were used in the simulation. Because the holdup differences in Table 3 are small, there may be some flow instability as Joshi et al.15 pointed out. They have modeled and have indicated the possibility of hydrodynamic instability for a certain range of operating conditions for the conditions of gas flow only in the riser or only in the downcomer. The study of Joshi et al.15 is very detailed and should be extended for systems with gas flow in both the riser and downcomer to identify the stable regimes of operating velocities. However, it may be noted that instability in operation was not observed in our studies for the range of flow velocities investigated. Also, the flow instability issues have not been reported in commercial operations either, and this was also confirmed in this work. Additional mass transport parameters needed are kL and agl in both the riser and downcomer. Note that separate values are needed and the overall volumetric mass-transfer coefficient is not sufficient. This is because Ha is dependent only on true kL and not kLagl. Empirical correlations are available for bubbly systems. For example, Kastanek16 measured the overall coefficient, kLagl, in bubble columns and suggested the following correlation:

(18)

kl ) 2

x

Dub πdB

(19)

(20)

where ub is the terminal rise velocity of the gas in the reactor and dB is the average bubble diameter. No empirical correlation is needed when using this approach, and the transport parameters are directly tied to the hydrodynamic parameters. The holdup and bubble diameter become the important parameters here. The overall volumetric mass-transfer coefficient in the liquid phase, klaal, obtained from the correlations of eqs 19 and 20 was 9.16 × 10-2 s-1 (for an assumed bubble diameter of 2 mm) and that from the correlation of eq 19 was 2 × 10-2 s-1 for the case at U0gr of 3 m/s and U0gd of 1 m/s. The two correlations are generally within reasonable ranges, although the value obtained from the correlation for the conditions with no gas injected in the downcomer is somewhat lower. For the simulation results shown in section 5, the penetration theory estimates are used. The overall volumetric mass-transfer coefficients klagl and agl can also be experimentally obtained using different concentrations of EDTA in the liquid feed and/ or concentrations of H2S in the gas feed at constant gas velocity (see details in work by Levenspiel10 and Wubs and Beenackers12). Different concentrations lead to different controlling regimes. At the conditions of the instantaneous reaction regime, klagl can be directly obtained from the rate information. At the conditions for a pseudo-first-order reaction regime, when the enhancement factor becomes equal to the Hatta number, the interfacial area, agl, can be directly obtained. Future work in this direction using data obtained over a wide range of gas and liquid concentrations would be useful to provide additional mass transport information on these types of reactors. Mass-transfer resistance in the gas phase was assumed to be negligible. The estimated maximum resistances (1/kga) in the riser and in the downcomer at U0gr of 3 m/s and U0gd of 1 m/s are only 0.6% and 0.07% of the total resistance, respectively. 5. Comparison of the Experiments with the Model Figure 10 compares the experimental and model predicted values for H2S and Fe3+EDTA conversion at different superficial gas velocities in the riser. Increasing the gas velocity in the riser decreases the H2S conversion. This is because of a decreasing residence

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Figure 12. Axial concentration profiles of Fe2+EDTA in the downcomer at U0gr ) 0.06 m/s and U0gd )0.01 m/s, number of cells ) 10, and bubble diameter ) 1.55 mm.

Figure 10. Comparison of experimental data and model predictions for H2S and Fe3+EDTA conversions.

Figure 11. Axial concentration profiles of Fe3+EDTA in the riser at U0gr ) 0.06 m/s and U0gd ) 0.01 m/s, number of cells ) 7, and bubble diameter ) 2 mm.

time of the gas phase in the riser. Fe3+EDTA conversion decreases with increasing riser gas velocity because a higher circulating velocity of the liquid reduces the residence time in the reaction zone. The results from the model agreed well with the experimental data. The bubble sizes used in the calculation were 2, 2, 2.5, and 3 mm for the riser superficial gas velocities of 0.03, 0.06, 0.09, and 0.11 m/s, respectively. The bubble diameter increases because of increased coalescence at higher velocity. Hence, the large bubble diameter of 3 mm used for higher velocities was physically realistic in view of the higher rate of bubble coalescence. The smaller value of the bubble diameter of 2 mm at low velocities is typical of ionized systems because the ions inhibit bubble coalescence. The total number of mixing cells, n, is lower when the flow behavior approaches mixed flow. The numbers of cells used in the model predictions were 10, 7, 5, and 4 for the superficial gas velocities of 0.03, 0.06, 0.09, and 0.11 m/s, respectively. Note that at higher gas velocities the backmixing increases and hence the value of n decreases as the gas velocity increases. Figure 11 shows the experimental- and modelpredicted axial profiles of the Fe3+EDTA concentration in the riser at a superficial gas velocity in the riser of 0.06 m/s and that in the downcomer of 0.01 m/s. The number of cells used in the calculation was 7, and the bubble diameter was 2 mm. Overall, there is a good

general agreement between the model and measured data. The measured Fe3+EDTA concentration at the top of the riser is slightly higher than the predicted value. This is likely due to the air oxidation occurring at the surface. At the height of 0.1-0.3 m, the experimental concentrations are lower than the model-predicted values. This discrepancy may be due to the use of average values of the hydrodynamic parameters such as the bubble diameter, circulating velocity, and gas holdup for the entire riser. The model should be able to predict better if these parameters can be measured locally and local values are used in the model. The local measurements of the liquid circulation velocity and the gas holdup using noninvasive techniques such as computer-automated radioactive particle tracking and γ-raycomputed tomography17-19 are useful in this context. An accurate estimation of the bubble diameter is also important because the model is shown to be sensitive to the bubble diameter. Overall, we find that the model agreement is reasonable considering the uncertainties in many parameters. Figure 12 compares the measured and predicted axial concentration profiles of Fe2+EDTA in the downcomer. A number of mixing cells of 10 was used for the downcomer, which is different from the number for the riser. The assumption made here is that, in the downcomer, the flow is less backmixed than that in the riser because of lower gas velocities. The model-predicted conversion of Fe2+ is lower than the experimental values, and the match of the data for the downcomer is not as good as that for the riser. More accurate parameter values for the downcomer are required for obtaining a better prediction. 6. Conclusions Gas-lift reactors are suitable for the Lo-Cat process because they can remove H2S to greater than 99.9%. In addition, Fe2+EDTA can be regenerated to Fe3+EDTA in the downcomer with a conversion as high as 79%. Thus, the catalyst can be recycled in the same reactor, leading to lower capital and operational costs. The conversions of H2S, Fe3+EDTA, and Fe2+EDTA decrease with increasing superficial gas velocity in the riser because of decreasing residence times in the various reaction zones. The comparison of the predictions of the mixing cell model with the measured data shows essentially similar trends. The mixing cell model is therefore a good learning model for this complicated process and can be used to guide the design process. Further work on more accurate hydrodynamics parameter values is required to attain better predictions and to understand this novel type of air-lift reactor.

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Acknowledgment

Literature Cited

The financial support from the Kasetsart University Research and Development Institute (KURDI), CHEADB Graduate Research and Education Development Program in Chemical Engineering at Kasetsart University, and Kasetsart University Graduate School is acknowledged.

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Notation A ) cross-sectional area, m2 agl ) gas-liquid interfacial area per unit of reactor volume, m2/m3 Ci(l) ) concentration of the liquid reactant or product species in the liquid bulk, mol/m3 Ci(g) ) concentration of the reactant in the gas phase, mol/ m3 f Ci(l) ) concentration of the ith diffusing species in the liquid film, mol/m3 dB ) bubble diameter, m Di ) molecular diffusivity of solute i, m2/s fB ) friction loss coefficient Q ) flow rate, m3/s Ha ) Hatta number HD ) liquid height in the riser, m Hi ) Henry’s law constant of species i k ) rate constant kl ) mass-transfer coefficient, m/s n ) number of mixing cells in the reactor Nfi ) flux of diffusing species in the film, mol/m2‚s Ri ) rate of consumption of species i per unit of liquid volume in the liquid film, gmol/m3‚s ug ) superficial gas velocity, m/s ul ) liquid circulating velocity, m/s Vr ) volume of the riser, m3 Vd ) volume of the downcomer, m3 x ) distance in the liquid film measured from the interface, m g ) gas holdup δ ) thickness of the liquid film Subscripts A ) reactant in the gas phase B ) reactant in the liquid phase d ) downcomer f ) film g ) gas phase i ) species index j ) cell number index l ) liquid phase r ) riser

Received for review September 2, 2004 Revised manuscript received February 26, 2005 Accepted March 11, 2005 IE049166P