Modeling of an Autothermal Reactor for the Catalytic Oxidative

Article Cite This: Ind. Eng. Chem. Res. 2019, 58, 10264−10270

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Modeling of an Autothermal Reactor for the Catalytic Oxidative Decomposition of H2S to H2 and Sulfur Daniela Barba,† Vincenzo Vaiano,*,† Vincenzo Palma,† Michele Colozzi,‡ Emma Palo,‡ Lucia Barbato,‡ Simona Cortese,‡ and Marino Miccio‡ †

Department of Industrial Engineering of the University of Salerno, Via Giovanni Paolo II, 132, Fisciano 84084, SA, Italy KT- kinetics Technology S.p.A, Viale Castello Della Magliana, 27, Rome 00148, Italy

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‡

ABSTRACT: An adiabatic fixed-bed reactor for the simultaneous production of sulfur and hydrogen by catalytic oxidative decomposition of H2S was simulated using a steady-state one-dimensional heterogeneous reactor model. A nonlinear system of differential equations of the first order, characterized by material balances coupled with the energy ones, was solved by numerical integration with the Euler method. The influences of the reactor inlet temperature (TIN), molar feeding ratio (O2/ H2S), H2S inlet molar fraction (ZH2SIN), and pressure on the adiabatic reactor performances were investigated in terms of H2S conversion, H2 yield, and SO2 selectivity for the homogeneous and the catalytic systems. The reaction system in the presence of Al2O3-based catalyst was compared to the homogeneous one at the reactor inlet temperature of 873 K, highlighting how the catalyzed system is able to reach the final temperature (1395 K) more quickly than the homogeneous one. The variation of TIN, O2/H2S, and ZH2SIN improved both the H2S conversion and the H2 yield with the reaching of very high temperatures.

1. INTRODUCTION

Hydrogen is currently needed in large quantities as a chemical feedstock in the synthesis of ammonia and methanol, in the desulfurization and hydrocracking processes, as well as for the upgrading of various hydrocarbon resources such as heavy oil and coal.4 Nowadays, hydrogen is obtained starting from the light hydrocarbons via steam and dry reforming, partial oxidation, water gas shift reactions, etc.5 In the last years, several research papers have dealt with investigating the possibility of obtaining simultaneously sulfur and hydrogen by H2S catalytic thermal cracking (H2S = H2 + 1/2S2).6−9 Generally, sulfides and oxides of transition metals are used as catalysts for the heterogeneous high-temperature decomposition of hydrogen sulfide.10,11 Bimetallic-based Ru catalysts supported on γ-Al2O3 (MoS2−RuS2, MnS−RuS2, CuS−RuS2, NiS−RuS2, FeS2−RuS2, and ZnS−RuS2) have been studied for the H2S thermal decomposition reaction.12

Hydrogen sulfide (H2S) is one of the main byproducts in natural gas plants, refineries, heavy oil upgraders, and metallurgical processes. It is a toxic gas and classified as hazardous industrial waste.1 The standard commercial technology for treating H2S is the Claus process, in which H2S is converted to elemental sulfur and water. In the Claus process, the reactions are performed at very high temperatures and require costly equipment.2 In addition, the Claus process recovers only elemental sulfur, while hydrogen is lost as low-grade steam. Sulfur represents an important element in today’s world. Approximately six-sevenths of all of the sulfur produced is converted into sulfuric acid, which is used in the manufacture of fertilizers (phosphates and ammonium sulfate). Other important uses include the production of pigments, detergents, fibers, petroleum products, sheet metal, explosives, and storage batteries. Sulfur not converted to sulfuric acid is used in making paper, insecticides, fungicides, dyestuffs, as well as numerous other products,3 such as the manufacturing of fertilizers. © 2019 American Chemical Society

Received: Revised: Accepted: Published: 10264

March 22, 2019 May 21, 2019 May 30, 2019 May 30, 2019 DOI: 10.1021/acs.iecr.9b01595 Ind. Eng. Chem. Res. 2019, 58, 10264−10270

Article

Industrial & Engineering Chemistry Research

advantages, such as their relative compactness, lower capital cost, greater potential for economies of scale, and their flexibility with respect to the products’ composition.27 Although the H2S oxidative decomposition carried out in an autothermal reactor could have an interesting potential in the field of industrial chemistry, no paper discusses the design and simulation of this type reactor in current literature. For this reason, in this work the kinetics parameters evaluated in isothermal conditions were used for the mathematical modeling of the reactor working in autothermal conditions. To calculate the final adiabatic temperature and the gaseous composition at the outlet of the reactor, different cases studied were considered, such as the reactant inlet temperature, molar feeding ratio, and H2S inlet volumetric fraction.

Unfortunately, this reaction is highly endothermic and is strictly limited thermodynamically, because it is favored only by high temperatures.13 Therefore, technical hurdles such as great energy demand, recycling of H2S unconverted, fast quenching of the reaction products, and successive separation stages should result in high fixed and operating costs.14 Other approaches based on H2S decomposition have also been studied (such as electrochemical, photochemical, plasma methods, solar dissociation, and microwave), but are not applicable on an industrial scale due to the high energy demand.15 In a bid to solve some of the highlighted problems to make commercialization feasible and attractive, many H2S thermal decomposition schemes have been proposed. Reed (1986) proposed and obtained a U.S. patent for a modified Claus reaction furnace in which the considerable heat generated is used for the dissociation of H2S taking place in ceramic tubes loaded with a cobalt−molybdenum catalyst.16 Cox et al. (1998) carried out a simulation study together with an economic evaluation of two process schemes for the thermal decomposition of H2S using design data. The first scheme involves the recycling of H2S through a steam methane reforming type furnace so that all possible hydrogen could be extracted. In the second scheme, a once-through configuration was employed where H2S was passed through a decomposition/exchanger-type reactor, and then on to a conventional Claus sulfur recovery plant. From the economic evaluation, the production cost of hydrogen was significantly higher for the second scheme due to the lower hydrogen recovery obtained.17 The simultaneous recovery of hydrogen and sulfur can be realized through a novel process recently based on the catalytic oxidative decomposition of H2S at 1273−1373 K; an appropriate amount of oxygen could be added to use the heat produced by the H2S combustion (H2S + 1/2O2 → 1/2S2 + H2O) to obviate to the endothermicity of the H2S decomposition reaction (H2S ↔ H2 + 1/2S2), obtaining simultaneously sulfur, H2O, and H2 (2H2S + 1/2O2 → S2 + H2O + H2).18 A problem related to the presence of the oxygen is the possible formation of SO2 that is kinetically favored in a wide temperature range (373−1800 K).19,20 To overcome this last aspect, the catalyst plays a key role in depressing the possible SO2 formation, in compliance with the strictest emission regulations.21 Alumina-based catalysts are generally employed because they enhance H2 production and reduce SO2 formation by the Claus reaction (2H2S + SO2 = 3/ 2S2 + 2H2O).22 First, this reaction was studied in a homogeneous phase, by changing the main operating conditions.23 The results showed an approach of H2S conversion and H2 yield to equilibrium values only at high temperatures, but with SO2 selectivity (∼5%) higher than that expected from the thermodynamic equilibrium calculations (0.4%) at 1273 K.24 Very different results were observed in the presence of alumina-based catalysts, which allowed one to obtain a very low SO2 selectivity (0.05%) at 1273 K.25 Moreover, starting from the experimental data shown in the paper by Palma et al.,26 a mathematical model able to describe the reaction system in isothermal conditions with a good accuracy was developed. However, it is well-known that the realization of an isothermal reactor is very difficult at the industrial scale level. On the other hand, autothermal reactors have received considerable attention due to the significant number of

2. EXPERIMENTAL METHODS The experimental tests were carried out in the laboratory plant presented in Figure 1 and described in detail in a previous work.26

Figure 1. Scheme of the laboratory apparatus.

In detail, a system of three-way valves allows one to feed the feed stream (H2S, O2, and N2) to the reaction section and the products to the analysis section (sampling line 2), while, in the bypass position, the reactants go directly to the analysis section to verify the composition of the feeding gas (sampling line 1). Experiments were performed in a fixed bed quartz tubular reactor specifically designed to operate at high temperatures. The exhaust stream was analyzed by a quadrupole filter mass spectrometer (Hiden HPR 20). The operating conditions used for the modeling of the adiabatic reactor are as follows: pressure, 1 atm; inlet reactor temperature (TIN), 298−873 K; H2S inlet concentration (ZH2SIN), 0.1−0.5 vol %; molar feeding ratio (O2/H2S), 0.15−0.25; total flow rate, 600 N cm3 min−1; GHSV = 12 000 h−1; and contact time, 300 ms. The H2S conversion (xH2S), SO2 selectivity (sSO2), and H2 yield (yH2) were evaluated according to the following relationships (eqs 1−3): x H 2S (%) =

sSO2 (%) = y H 2 (%) = 10265

(z H 2SIN − z H 2SOUT) ·100 z H 2SIN

zSO2OUT (z H 2SIN − z H 2SOUT) z H 2OUT z H 2SIN

·100

·100

(1)

(2)

(3) DOI: 10.1021/acs.iecr.9b01595 Ind. Eng. Chem. Res. 2019, 58, 10264−10270


yz jij d ijjj zz = jj∑ ν r ( −ΔH 0 (T )) jj∑ cpyQT z j i j r ,j dV j i i i zz jj i , j k { k yz z + ∑ νirjcat·ρcat ( −ΔHr0, j(T ))zzz zz i,j { The boundary conditions are as follows:

where zH2SIN is the inlet H2S volumetric fraction; zH2SOUT is the outlet H2S volumetric fraction; zSO2OUT is the outlet SO2 volumetric fraction; and zH2OUT is the outlet H2 volumetric fraction. The modeling of the adiabatic reactor was carried out for the homogeneous case and in the presence of Al2O3-based catalyst. Alumina in powder form was synthesized by the thermal treatment of pseudoboehmite samples at 900 °C for 12 h in static air to obtain the stabilization of the alumina phase.26 The nonlinear system of differential equations of first-order with constant coefficients was solved by a numerical procedure using the Euler method. It is the most basic explicit method for the numerical integration of ordinary differential equations (ODEs) with a given initial value. The mass balances were written considering a plug flow reactor (PFR) having a reaction volume of 3 cm3. The values of the apparent kinetic constants for each catalytic reaction were obtained by using the least-squares approach, based on the minimization of the sum of squared residuals between the experimental data obtained at different temperatures and the values given by the mathematical model with a molar feed ratio (O2/H2S) equal to 0.2. The evaluation of the equilibrium constant (Keq) and reaction heat (ΔH0r) as a function of temperature was performed with the Gaseq program, a software based on the minimization of Gibbs free energy, able to calculate the equilibrium product composition of an ideal gaseous mixture, when there are a lot of simultaneous reactions for which it is difficult to use the equilibrium constants only. The thermodynamic analysis was carried out considering that the following chemical species are present when the system reaches the equilibrium conditions: H2S, O2, SO2, S2, S6, S8, H2O, H2, and nitrogen.

(4)

(2) H 2S + 3/2O2 → SO2 + H 2O

(5)

V = 0 Tx = TxIN

The kinetic constants for the homogeneous (eqs 11−15) and catalytic reactions26 (eqs 16−19) as a function of temperature are reported in the following equations: k1 = e(22.655) e(−20033/ T ) [min−1 atm−1.5]

(11)

k 2 = e(21.845) e(−14477/ T ) [min−1 atm−2]

(12)

k 3 = e(12.274) e(−5079.5/ T ) [min−1 atm−1.5]

(13)

k4 = e(17.038) e(−11712/ T ) [min−1 atm−2]

(14)

k5 = e(14.981) e(−7171/ T ) [min−1 atm−2] (15) ÄÅ É Ñ ÅÅ cm 3 ÑÑÑ ÑÑ k1cat = e(20.327) e(−10678/ T ) ÅÅÅÅmin−1 atm−1.5 ÅÅÇ g ÑÑÑÖ (16) ÅÄÅ ÑÉ Å cm 3 ÑÑÑ ÑÑ k 3cat = e(16.911) e(−7229.7/ T ) ÅÅÅÅmin−1 atm−1.5 ÅÅÇ g ÑÑÑÖ (17) ÄÅ É Ñ ÅÅ cm 3 ÑÑÑ ÑÑ k1cat reverse = e(26.382) e(−22756/ T ) ÅÅÅÅmin−1 atm−1 ÅÅÇ g ÑÑÑÖ (18) ÄÅ É Ñ ÅÅ cm 3 ÑÑÑ ÑÑ k5cat = e(22.014) e(−8781.4/ T ) ÅÅÅÅmin−1 atm−2 ÅÅÇ g ÑÑÑÖ (19) It was assumed that reactions 2 and 4 and the reverse reaction 3 do not occur in the presence of the catalyst,26 for which their respective kinetic constants can be considered negligible (eq 20).

k 2cat = k 3cat reverse = k4cat = 0

(3) H 2S + 1/2SO2 ↔ 3/4S2 + H 2O

(6)

(4) H 2S + O2 → H 2 + SO2

(7)

(5) H 2S + 1/2O2 → 1/2S2 + H 2O

(8)

∑ νirj + ∑ νirjcat·ρcat i

i

(20)

For each reaction, the variation of the equilibrium constant (Keq) with the temperature was expressed with the following relationship (eq 21):

i −ΔGro yz zz Kjeq = expjjjj z (21) k RT { The equilibrium constants for the reactions 1 and 3 (eqs 22 and 23) and the reaction enthalpy for all of the reactions were calculated using the GasEq software (eqs 24−28):

It was considered that these reactions occur both in the homogeneous phase as well as in the presence of the aluminabased catalyst. The kinetic expressions employed in the model and the mass balance for each component (H2S, O2, S2, SO2, H2O, H2) are reported in a previous work.26 The material and energy balances are expressed in general form in accordance with eqs 9 and 10. d (yQ ) = dV i

(10)

V = 0 yi = yiIN

3. RESULTS AND DISCUSSION Starting from the experimental results previously obtained,26 a mathematical model was developed through the identification of the five main reactions able to describe the reaction system in isothermal conditions (eqs 4−8). (1) H 2 + 1/2S2 ↔ H 2S

Article

i 0.0485T − 89.455 yz zz [atm−0.5] K1eq = expjjj 0.008314T k {

(9) 10266

i −0.0303T + 23.354 yz zz [atm+0.25] K3eq = expjjj 0.008314T k {

(22)

ΔHr10 = −3 × 10−9T 3 + 10−5T 2 − 0.02T − 79.63

(24)

(23)

DOI: 10.1021/acs.iecr.9b01595 Ind. Eng. Chem. Res. 2019, 58, 10264−10270

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homogeneous system (∼2.5 cm3). Moreover, the achievement of the stationary conditions at different reaction volumes significantly influences the value of the outlet molar fraction of the main components (H2S, H2, and SO2) for the catalytic system with respect to the homogeneous one. In particular, the presence of the catalyst (Figure 2b) allows one to obtain a H2S conversion, H2 yield, and SO2 selectivity equal to 52, 12, and 0.04%, respectively. The rapid increase in temperature up to a maximum value is due to the presence of exothermic reactions, as confirmed by the peak of SO2 formation, which, in the absence of the catalyst (0.045%), is 1 magnitude order higher than that observed for the catalytic system (0.005%). The SO2 peak of both systems is better highlighted in Figure 3, where it is also possible to observe how, in the absence of

ΔHr20 = 8 × 10−10T 3 − 3 × 10−6T 2 − 0.0005T − 518.4 (25)

ΔHr30

−9 3

−5 2

= 4 × 10 T − 10 T − 0.01T + 21.43

(26)

ΔHr40 = 9 × 10−10T 3 − 6 × 10−6T 2 + 0.013T − 280.2 (27)

ΔHr50

−9 3

−6 2

= 3 × 10 T − 9 × 10 T + 0.007T − 158.4 (28)

The specific heats (Cpi) retrieved from Perry’s Handbook28 were expressed as in eq 29: ÄÅ É ÄÅ É ÅÅ c3i /T ÑÑÑ2 ÅÅ c5i /T ÑÑÑ2 Å Ñ Å ÑÑ Å Ñ Å Cpi = c1i + c 2iÅÅ Ñ + c4iÅÅ Ñ ÅÅÇ sinh(c3i /T ) ÑÑÑÖ ÅÅÇ cosh(c5i/T ) ÑÑÑÖ (29) where cni (n = 1, ...0.5) are characteristic constants for each component listed in Perry’s Handbook.28 Because the perfect adiabatic conditions are difficult to achieve in a reactor at laboratory scale, the data retrieved by the model will not be compared to the experimental data. Figure 2 shows the temperature profile and the volumetric

Figure 3. Comparison of the SO2 profile concentration for the homogeneous and catalytic systems (TIN = 873 K, ZH2SIN = 0.4, O2/ H2S = 0.2).

the catalyst, the maximum temperature value is about 60 K higher than that obtained in the presence of the catalyst (T = 1437 K), with it being observed for a higher reactor volume. At reaction volumes higher than 1 cm3 for the homogeneous system and higher than 0.2 cm3 for the catalytic system, the heat generated by the oxidation reaction of H2S to SO2 is used by endothermic reactions occurring in the system, allowing the reactor to reach the final temperature of about 1400 K and the steady-state conditions. The H2S, H2, and SO2 molar fractions obtained from the model for both systems were compared to those expected from the thermodynamic equilibrium. From the results listed in Table 1, it is important to note that the molar fractions calculated with the model are generally close to the thermodynamic equilibrium values. Table 1. H2S, H2, and SO2 Molar Fractions As Compared To the Equilibrium Valuesa

Figure 2. Molar fractions and reactor outlet temperature profiles as a function of the reaction volume for the homogeneous case (a) and in the presence of an alumina-based catalyst (b) (TIN = 873 K, ZH2SIN = 0.4, O2/H2S = 0.2).

fractions of the main components at the reactor outlet as a function of the reaction volume for the homogeneous system (a) as well as in the presence of the alumina-based catalyst (b) when the inlet reaction temperature (TIN) is equal to 873 K. It is worth noting how, by preheating the reactants at the temperature of 873 K, both of the reaction systems (a and b) are able to reach a final temperature of about 1400 K but with different reaction rates. In particular, the catalyzed system reaches the final temperature for a reaction volume of about 0.65 cm3, significantly lower than that required for the

a

species

equilibrium

homogeneous system

catalytic system

H2S H2 SO2

0.17 0.07 0.0013

0.18 0.06 0.0025

0.19 0.05 0

TIN = 873 K, ZH2SIN = 0.4, O2/H2S = 0.2, V = 3 cm3.

Figure 4 reports the outlet reactor temperature profile for the homogeneous and catalytic systems by increasing the inlet temperature from 298 to 873 K. The profiles look like a sigmoid curve. However, it is interesting to note that, starting from the inlet temperature of 573 K, the catalyzed reaction system reaches outlet temperatures higher than those in the 10267


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Figure 4. Reactor temperature profile as a function of the inlet reactor temperature (ZH2SIN = 0.4, O2/H2S = 0.2).

homogeneous case. This trend reverses for inlet temperatures higher than 650 K. The outlet H2S conversion, H2 yield, and SO2 selectivity as a function of the inlet reactor temperature are shown in Figure 5. As expected, the presence of the catalyst allows one to reach steady-state values at temperatures lower than those required by the homogeneous system. This aspect is even more evident for the H2 yield and SO2 selectivity. The catalyst enhances the H 2 formation and the minimization of the SO2 selectivity for an inlet temperature higher than 600 K. The values of H2 yield and SO2 selectivity obtained at TOUT = 1366 K, corresponding to an inlet temperature of 823 K, are, respectively, 11% and 0.045%. Conversely, in the homogeneous case, starting from an inlet temperature of about 800 K (TOUT ≈ 1370 K), the formation of SO2 is not negligible with a selectivity value close to 2%. It is worth noting that for the catalyzed system, the formation of SO2 is limited to a narrow interval temperature (TIN = 595− 600 K) corresponding to an outlet reactor temperature in the range 884−1220 K. For higher temperatures, the formation of SO2 is almost completely depressed. This behavior can be explained considering that the catalyst promotes the Claus reaction (2H2S + 3/2SO2 = S2 + 2H2O), thus involving SO2 consumption.26 Similar to the SO2 formation, the H2 production starts to occur from 823 K, reaching a yield higher than 10% at 1370 K. In addition, for the homogeneous system, the H2S decomposition reaction that leads to the H2 formation begins at higher temperatures than in the catalyzed system (Figure 5). The effect of the feed molar ratio (O2/H2S) on the reactor outlet temperature was simulated in the presence of the alumina-based catalyst (Figure 6). By increasing the O2/H2S ratio from 0.15 to 0.25, two effects on the temperature profile can be observed. First, a higher O2 content involves the achievement of a final temperature always much higher because the exothermic reactions are ever more favored, determining an increase of about 230 K, by varying the O2/H2S ratio from 0.15 (T = 1281 K) to 0.25 (T = 1512 K). Furthermore, by increasing the O2/H2S ratio, the temperature peak is shifted toward lower reaction volumes due to the high rate of the oxidation reactions. The modulation of the oxygen concentration in the feed stream involves a general increase of the catalytic performance and unfortunately also of the SO2 selectivity, as it is possible to see in Figure 7. The catalytic performance calculated by the

Figure 5. H2S conversion (a), H2 yield (b), and SO2 selectivity (c) as a function of the reactor inlet temperature (ZH2SIN = 0.4, O2/H2S = 0.2, V = 3 cm3).

Figure 6. Reactor outlet temperature profiles as a function of the reaction volume by varying the O2/H2S in the range 0.15−0.25 (TIN = 873 K, ZH2SIN = 0.4, V = 1.5 cm3).

model agrees well with those expected by the thermodynamic equilibrium (continuous line). The obtained data are also summarized in Table 2, where it is evident that the final temperatures (TOUT) achieved for the catalytic and homogeneous systems are very similar. 10268


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although there is an appreciable increase of the H2S conversion and H2 yield by increasing the H2S inlet concentration, the SO2 formation is in all cases very low due to the presence of the alumina-based catalyst.26 Finally, the effect of the pressure was investigated for the reaction system in the presence of the catalyst. The obtained results are summarized in Table 4. Table 4. Catalytic Performance and Outlet Temperature As a Function of Pressurea

Figure 7. Catalytic performance as a function of O2/H2S (TIN = 873 K, ZH2SIN = 0.4, V = 1.5 cm3). a

Table 2. Comparison of the Catalytic Performance and Final Temperature with the Homogeneous System by Varying O2/H2Sa XH2S,%

0.15 0.2 0.25

38 52 66

YH2,%

SSO2,%

O2/H2S

XH2S,%

YH2,%

SSO2,%

TOUT, K

0.15 0.2 0.25

18 43 61

1 12 24

4 5 8

1287 1408 1453

8 0.02 12 0.04 16 0.09 Homogeneous System

a

TOUT, K 1280 1395 1511

3

From the comparison of the two systems, it is possible to observe that, at the same final temperature, relevant differences are observable especially for the SO2 selectivity. Even if the oxidation reactions are the more kinetically favored, the catalyst is however able to minimize the SO2 selectivity that does not exceed 0.1% with the highest O2/H2S molar ratio. On the other hand, in the absence of the catalyst, the SO2 selectivity is always higher than 4%. The hydrogen yield is similar for the two reaction systems because the H2S decomposition reaction is particularly promoted at such a high reaction temperature (T > 1300 K). Additionally, the increase of the H2S inlet concentration involves an important effect on the adiabatic temperatures and the catalytic performance (Table 3). The H2S conversion and H2 yield are, respectively, equal to 55% and 15% by feeding a stream containing a molar fraction of H2S equal to 0.5, which is probably due to the high final temperature reached (∼1460 K) that has promoted in a particular way the H2S decomposition reaction to hydrogen and sulfur. Furthermore, it is important to observe that,

TOUT, K

XH2S, %

YH2, %

SSO2, %

1050 1202 1395 1461

39 45 52 55

0.5 5 12 15

0.1 0.05 0.04 0.04

YH2, %

SSO2, %

52 50 49

12 10 9

0.04 0.05 0.06

TIN = 873 K, O2/H2S = 0.2, ZH2SIN = 0.4, V = 3 cm3.

S8 ↔ 4S2

(30)

3S8 ↔ 4S6

(31)

S6 ↔ 3S2

(32)

H 2S ↔ H 2 + 1/2S2

(33)

4. CONCLUSIONS An autothermal reactor for the simultaneous production of sulfur and hydrogen by catalytic oxidative decomposition of H2S was simulated using a steady-state one-dimensional heterogeneous reactor model. The nonlinear system of differential equations of the first-order was solved by numerical integration with the Euler method. The effect of the main operating conditions on the reactor adiabatic temperature and on the composition at the reactor outlet was studied by the model calculation. The variations of the preheating temperature of the reactants (298−873 K), the molar feeding ratio (O2/H2S = 0.15−0.25), and the H2S inlet molar fraction (0.1−0.5) have involved a decisive improvement of the catalytic performances in agreement with the high reactor outlet temperature. A high reaction temperature promotes H2 formation through the H2S thermal decomposition reaction, with a H2 yield always higher than 10%. For the homogeneous system, it was observed that a reactor volume higher (2.5 cm3) than the catalytic case (0.3 cm3) is necessary for the reaching of the steady-state conditions in terms of adiabatic temperature and outlet H2S, H2, and SO2 concentrations. On the basis of the obtained results, the mathematics enables the identification of optimal inlet reactor temperatures and other operating conditions to maximize the conversion of

Table 3. Catalytic Performance and Outlet Temperature As a Function of the Inlet H2S Concentrationa 0.1 0.2 0.4 0.5

XH2S, %

1396 1413 1424

In addition to the H2S decomposition reaction, there are the polyatomic sulfur reactions that are endothermic when the sulfur dissociation occurs. Therefore, the results reported in Table 4 can be explained considering the previous reactions that evolve with an increase of the mole number and are thus favored at lower pressures.

TIN = 873 K, ZH2SIN = 0.4, V = 1.5 cm .

ZH2SIN

TOUT, K

1 2 3

The change of the total pressure from 1 to 3 atm determines a slight decrease of the H2S conversion and H2 yield with a quasi-negligible variation of the SO2 selectivity. In the reaction system, the following endothermal reactions may occur (eqs 30−33):

Catalytic System O2/H2S

pressure, atm

a

TIN = 873 K, O2/H2S = 0.2, V = 3 cm3. 10269


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(12) Cao, D.; Adesina, A. A. Fluidised bed reactor studies of H2S decomposition over supported bimetallic Ru catalysts. Catal. Today 1999, 49 (1−3), 23. (13) Slimane, R. B.; Lau, F. S.; Dihu, R. J.; Khinkis, M. Production of hydrogen by superadiabatic decomposition of hydrogen sulfide. Proceedings of the U.S. DOE Hydrogen Program Review; NREL/CP610-32405, 2002. (14) Luinstra, E. Hydrogen from H2S-a review of the leading processes. Proceedings of 7th Sulphur Recovery Conference; Gas Research Institute: Chicago, IL, 1995; p 149. (15) Zaman, J.; Chakma, A. Production of hydrogen and sulfur from hydrogen sulphide. Fuel Process. Technol. 1995, 41, 159. (16) Reed, R. L. Modified Claus furnace. U.S. Patent 4575453, 1986. (17) Cox, B. G.; Clarke, P. F.; Pruden, B. B. Economics of thermal dissociation of H2S to produce hydrogen. Int. J. Hydrogen Energy 1995, 23, 531. (18) Colozzi, M.; Barbato, L.; Angelini, F.; Palo, E.; Palma, V.; Vaiano, V. Catalyst for a sulphur recovery process with concurrent hydrogen production, method of making thereof and the sulphur recovery process with concurrent hydrogen production using the catalyst. WO 2014/073966A1, 2014. (19) Palma, V.; Barba, D. Low temperature catalytic oxidation of H2S over V2O5/CeO2 catalysts. Int. J. Hydrogen Energy 2014, 39, 21524. (20) Salisu, I.; Raj, A. Kinetic Simulation of Acid Gas (H2S and CO2) Destruction for Simultaneous Syngas and Sulfur Recovery. Ind. Eng. Chem. Res. 2016, 55, 6743. (21) Colozzi, M.; Cortese, S.; Barbato, L. Innovative Way to Achieve “Zero Emissions” in Sulphur Recovery Facilities. Industrial Plants, Italian Engineering, Contracting and Plant Components Suppliers-animp 2016, 44−51. (22) Reshetenko, T. V.; Khairulin, S. R.; Ismagilov, Z. R.; Kuznetsov, V. V. Study of the Reaction of High-Temperature H2S Decomposition on Metal Oxides (γ-Al2O3; α-Fe2O3;V2O5). Int. J. Hydrogen Energy 2002, 27, 387. (23) Palma, V.; Vaiano, V.; Barba, D.; Colozzi, M.; Palo, E.; Barbato, L.; Cortese, S. H2 Production by Thermal Decomposition of H2S in the Presence of Oxygen. Int. J. Hydrogen Energy 2015, 40, 106. (24) Palma, V.; Vaiano, V.; Barba, D.; Colozzi, M.; Palo, E.; Barbato, L.; Cortese, S. H2S Oxidative decomposition for the simultaneous production of sulphur and hydrogen. Chem. Eng. Trans. 2016, 52, 1201. (25) Palma, V.; Vaiano, V.; Barba, D.; Colozzi, M.; Palo, E.; Barbato, L.; Cortese, S.; Miccio, M. Catalytic Oxidative Decomposition of H2S for Hydrogen Production. Chem. Eng. Trans. 2018, 70, 325. (26) Palma, V.; Vaiano, V.; Barba, D.; Colozzi, M.; Palo, E.; Barbato, L.; Cortese, S. Oxidative Decomposition of H2S over Alumina-Based Catalyst. Ind. Eng. Chem. Res. 2017, 56, 9072. (27) Piña, J.; Borio, D. O. Modeling a simulation of an autothermal reformer. Lat. Am. Appl. Res. 2006, 36, 289. (28) Perry, R. H.; Green, W. Perry’s Chemical Engineers’ Handbook; McGraw-Hill Co., Inc.: New York, 2008.

the reactants and the H2 yield with the simultaneous minimization of the SO2 selectivity.

■

AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. ORCID

Daniela Barba: 0000-0001-5163-2823 Vincenzo Palma: 0000-0002-9942-7017 Notes

The authors declare no competing financial interest.

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ABBREVIATIONS Q = total flow rate [cm3 min−1] yi = volumetric fraction of component i (H2S, O2, SO2, H2O, S2, H2) [−] V = reaction volume [cm3] νi = stoichiometric coefficient of component i (H2S, O2, SO2, H2O, S2, H2) [−] rj = reaction rate (for j = 1−5) Cpi = specific heat of component i [J mol−1 K−1] ΔHr0j = reaction enthalpy [J mol−1] T = temperature [K] Keqj = equilibrium constant kj = kinetic constant for the homogeneous reactions kjcat = kinetic constant for the catalytic reactions ρcat = catalyst bulk density = 0.1 g cm−3 ΔG0 = Gibbs free energy [J mol−1] R = 8.314 [J mol−1 K−1] cni = characteristics constant for each component (n = 1, ...0.5) sinh = hyperbolic sine [K−1] cosh = hyperbolic cosine [K−1]

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REFERENCES

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