Alumina Catalysts

Article pubs.acs.org/IECR

Natural Gas Steam Reforming over Rhodium/Alumina Catalysts: Experimental and Numerical Study of the Carbon Deposition from Ethylene and Carbon Monoxide Claudia Eßmann,† Lubow Maier,‡ Aijun Li,§ Steffen Tischer,‡ and Olaf Deutschmann*,†,‡ †

Institute for Chemical Technology and Polymer Chemistry and ‡Institute of Catalysis Research and Technology at Karlsruhe Institute of Technology (KIT), 76128 Karlsruhe, Germany § Research Center for Composite Materials (RCCM), Shanghai University (SHU), P.O. Box 321, Yanchang Rd. 149, 200072 Shanghai, People’s Republic of China ABSTRACT: Natural gas steam reforming (SR) over technically used rhodium/alumina (Rh/Al2O3) honeycomb catalysts is studied experimentally at temperatures between 923 and 1073 K and steam-to-carbon ratios (S/C) of unity, with regard to coke deposition caused by the decomposition of the product species ethylene (C2H4) and carbon monoxide (CO). Furthermore, the process is modeled using detailed reaction mechanisms, and numerical simulations are carried out to describe the coke formation on Rh/Al2O3 catalysts quantitatively. The amount of deposited carbon was detected and analyzed for varying feed mixtures of the products CO and C2H4 diluted in N2. During the decomposition of CO, the saturation of the amount of coke is monitored by feeding CO in high concentrations. No saturation occurs for the same amounts of coke resulting from the decomposition of C2H4. The coking rate caused by the decomposition of C2H4 is found to be ∼25 times higher than the coking rate caused by the decomposition of CO. The differences in coking behavior caused by C2H4 and CO, respectively, are described by coking models.

1. INTRODUCTION Conversion of hydrocarbon feedstock by steam reforming (SR) and catalytic partial oxidation (CPOX) of natural gas in order to produce hydrogen and synthesis gas (H2 and CO) is in the focus of industrial and academic interest. Syngas is utilized for the production of chemicals such as methanol, ammonia, and synthetic logistic fuels by Fischer−Tropsch synthesis.1,2 Rhodium is a potential catalyst for the production of syngas in high quality, because it ensures both a high hydrogen yield as well as a long catalyst lifetime, compared to other materials such as the commonly used nickel.3−5 During SR of hydrocarbon fuels, catalyst deactivation occurs depending on the process conditions, i.e., temperature, residence time, fuel and product composition. In general, catalyst deactivation is caused mainly by sintering, poisoning, or coking of the catalyst particles: • Sintering of catalyst particles results in a decrease of the catalytically active surface and, therefore, sintering decreases the conversion of the hydrocarbon fuels. • Poisoning of the catalyst surface is due to irreversible chemisorption of chemical species at the active sites of the catalyst. Poisoning of the catalyst can occur by addition of sulfur, which is contained in crude oils, natural gases, and other fuels. • Coking of the catalyst is initiated by the chemisorption or physisorption of hydrocarbon intermediates or CO on the catalyst surface, followed by the decomposition of these species. The decomposition of the carbon-containing intermediate species can lead to carbonaceous overlayers blocking the active sites of the catalyst.3,6,7 CO or the hydrocarbon intermediates that lead to coking of the catalyst are produced by main or side reactions of SR of © 2014 American Chemical Society

hydrocarbons. These carbon precursors are formed by surface reactions or gas-phase reactions, or simultaneously by both. Schädel et al.8−10 described coke formation during SR of propane−one of the minor constituents of natural gas−over Rh/Al2O3 catalysts in dependence on the operating temperature and the S/C ratio. Three different carbon deposition rates were identified, depending on the operating temperature. At temperatures above 923 K, gas-phase reactions occur and cracking as well as pyrolysis reactions result in the formation of coke gaseous precursors, which accelerate the deactivation of the catalyst by coke depositions. Using models originally developed for the formation of pyrolytic carbon composites,11,12 coke formation was successfully described in the high-temperature range above 973 K. These models were based on a detailed chemical scheme for pyrolysis of hydrocarbon species such as CH4, C2H2, C2H4, and C3H6.12,13 At temperatures lower than 973 K, the coking rate decreases strongly with decreasing temperatures. Here, rather slow surface reactions are presumably responsible for the coke formation on the catalyst surface. Coking results from a step-bystep formation of carbonaceous layers from CO or hydrocarbon intermediates. Depending on the precursor, the temperature, and the pressure, different carbon morphologies are detected. Coke is deposited as graphene-like or amorphous carbon, and the formation of carbon filaments and whisker structures is detected as well.3,6,7,12,14,15 Received: Revised: Accepted: Published: 12270

April 15, 2014 July 4, 2014 July 8, 2014 July 8, 2014 dx.doi.org/10.1021/ie5015525 | Ind. Eng. Chem. Res. 2014, 53, 12270−12278

Industrial & Engineering Chemistry Research

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In summary, coke formation may result from either surface reactions or gas-phase reactions followed by adsorption of rather large precursors, or from a combination of both, depending on the operating conditions such as temperature, pressure, S/C ratio, and residence time of the gases in the catalyst region. In addition, the interactions occurring between catalyst particles and support material increase the complexity of understanding the coke formation on technical catalysts. Coke formation was found to be preferentially caused by surface reactions at medium temperatures and low steam-to-carbon ratios.8,9,16 With increasing process temperatures, gas-phase reactions gain significance. The effect of main and side product species resulting from SR on the coking of catalysts was investigated in several studies. Ferrizz et al.17 performed a temperature-programmed desorption (TPD) study of C2H4 and CO on Rh-based catalysts on different support materials. C2H4 and CO were separately fed to the catalyst at room temperature and the product distribution, as well as the amount of deposited carbon, was detected and analyzed. According to their study C2H4 undergoes complete dehydrogenation on supported Rh particles at room temperature, Snoeck et al.15 investigated the influence of the Boudouard reaction on the coking behavior of nickel-based catalysts, and created a kinetic model for the formation of coke from CO and for gasification caused by H2O, H2, and CO2. In this study, the SR of natural gas over Rh/Al2O3 catalysts as well as the resulting coke formation on the catalysts is investigated by performing experiments and numerical simulations. In particular, the simulation of coke formation over Rh/Al2O3 catalysts is, for the first time, directly related to a detailed chemistry model, describing main and side reactions during SR process, leading to the formation of coke precursors. Detailed surface reaction mechanism and a gas phase reaction mechanism are combined with a transient coke formation model describing coke formation reasonably. Thereby, we focus on coke formation from the product species C2H4 and CO at different feed conditions. Therefore, catalytic decomposition of C2H4 and CO diluted in N2 is experimentally studied. Technical Rh/Al2O3 honeycomb catalysts are coked systematically using a conventional flow reactor. Then, temperature-programmed oxidation (TPO) is carried out over these coked catalysts in order to identify the kinetics of carbon deposition and the saturation concentrations of the respective species in a concentration range comparable to the SR process of natural gas (i.e., 923 K and S/C = 1). The catalytic decomposition of C2H4 is described by a modified carbon deposition model originally developed for hightemperature pyrolysis of light hydrocarbon species.12 Furthermore, a reaction mechanism for carbon deposition by means of the Boudouard reaction is developed.

Figure 1. Reactor design and position of the catalyst inside the quartz tube reactor with corresponding temperature profile of the reactor, measured in argon flow (1 L/min).

chamber filled with quartz glass spheres at a temperature of 453 K before entering the quartz glass reactor. The process temperature in the quartz reactor is adjusted by a tube furnace with a rather small temperature variation over the catalyst (Figure 1); the profile is measured for a gas flow of 1 L/min of argon. Natural gas (“Nordsee H”, Air Liquide), nitrogen (5.0, Air Liquide), and oxygen (4.8, Air Liquide) are dosed by mass-flow controllers (MFC/El-Flow, Bronkhorst). The supply of deionized water is realized by a MFC for liquids (Liquid-Flow, Bronkhorst). The water is vaporized at a temperature of 453 K using an evaporator operated at a low pulsation level. The product gas composition is analyzed by Fourier trnaform infrared (FT-IR) spectroscopy and mass specotroscopy (MS). The same flow reactor is used, as well for the catalytic decomposition of C2H4 and CO. The gases (3.0 C2H4, 4.7 CO, Air Liquid) are first fed to a mixing chamber by MFCs (El-Flow, Bronkhorst) and then directed to the quartz tube reactor. While the reactants are entering the reactor, no heating of the steel tubes was applied in order to avoid gas-phase reactions of the species in front of the catalyst region. The reactants are fed to the reactor diluted in N2 and with varying concentrations. The formation of coke on the catalysts caused by steam reforming and catalytic decomposition is quantified by temperature-programmed oxidation (TPO) (4.8 N2 and 5.0 O2, Air Liquide). More details on the reactor can be found in previous studies.8,9,18−20 2.2. Catalyst. The technical honeycomb catalysts used consists of cordierite coated with rhodium supported by γ-Al2O3. The rhodium particles, ∼1−2 nm in size, are welldispersed on the γ-Al2O3 washcoat. The catalytic honeycomb monolith is 10 mm in length and 19 mm in diameter, with a cell density of 900 channels per square inch (cpsi). CO chemisorption measurements of the catalytically active surface yielded a value of 0.64 m2/g. This catalytically active surface corresponds to a total amount of 2.2 × 10−5 mol Rh of the technical catalyst used. 2.3. Measurement Procedure. 2.3.1. Steam Reforming of Natural Gas. Steam reforming was studied in a temperature range of 923−1073 K in steps of 50 K. The S/C ratio was kept at approximately unity. The mixture of natural gas and steam was fed to the reactor at a gas hourly space velocity (GHSV) of

2. EXPERIMENTAL SETUP AND MEASUREMENT PROCEDURE 2.1. Experimental Setup. The reactor consists of a quartz tube with an inner diameter of 19 mm. The honeycomb catalyst is placed in the quartz glass tube 15 cm downstream from the gas inlet in the flow direction (Figure 1). Before the reactants reach the catalyst, they pass through an alumina foam monolith (Al2O3, 85 pores per linear inch, ppi) that is packed in front of the catalyst and serves as a flow homogenizer, heat shield, and fixation for the front thermocouple. The reactor inlet configuration was design based on CFD simulations of the mixing process, resulting in a homogeneous mixing and pulse-free feed supply.18 To ensure uniform mixing, natural gas and steam are mixed in a cylindrical 12271

dx.doi.org/10.1021/ie5015525 | Ind. Eng. Chem. Res. 2014, 53, 12270−12278


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16000 h−1 with respect to standard conditions (298.15 K, 1013.25 mbar). Before starting the SR experiments, the catalysts were burned off at 923 K in air (one standard liter per minute (SLPM) to ensure the removal of organic contaminants from the catalyst surface. SR was started by feeding the water first (0.45 SLPM) and adding natural gas after a few seconds (0.45 SLPM). The oven temperature was increased to 1073 K, in steps of 50 K. For each temperature, the product distribution was analyzed by FT-IR and MS after reaching steady state. 2.3.2. Carbon Deposition from C2H4 and CO. In the SR experiments of natural gas at 923 K, the catalyst outlet temperature was 823 K. In accordance with these SR experiments, the catalytic decomposition experiments were carried out at an oven temperature of 823 K. Varying concentrations of CO and C2H4, diluted with N2, were separately fed to the reactor. Depending on the concentration, the GHSV was altered between 23000 h−1 and 47000 h−1. Thereby, the concentration of C2H4 and CO varied over ranges of 0.07− 1.13 vol % and 1.20−10.92 vol %, respectively. Before starting the experiments, the catalysts were burned off at 823 K with 1 SLPM air. Afterward, the catalysts were reduced by temperatureprogrammed reduction (TPR) in H2/N2 mixtures (4.8 vol % H2 in N2) at 823 K. The potential coke precursors, C2H4 and CO, were separately fed in different dilutions at temperatures of 823 K for 20 min. In the experiments, the temperature. as well as the concentration of reactants and products. were monitored by FT-IR and H-Sense (QMS). After 20 min, the feed of the coke precursors was stopped and the catalysts were cooled in N2. After they reached room temperature, TPO of these catalysts was carried out with 20 vol % O2 in N2 at a temperature ramp of 20 K/min up to a temperature of 923 K while recording the concentration of the formed CO2 by FT-IR. The amount of deposited coke was quantified by analyzing the amount of the CO2 formed.

Figure 2. Sketch of the algorithm of the computational tool DETCHEMRESERVOIR for the simulation of the deposition of coke on the catalyst surface coupled with the chemically reactive flow through a single monolith channel during the SR of natural gas.

transient model of the entire reactor. Based on the gas-phase concentrations, the rate of deposition is determined. This rate is applied to solve the transient equation for the deposited species: dCidep = si̇(T , ci , θi , Cidep) dt

(1)

Here, denotes the concentration of deposited species; Ṡi is the rate of generation or consumption of species i due to adsorption, desorption, or surface reaction, depending on the local temperature (T), the gas-phase concentration (ci), surface coverages (θi), and depositions. The amount of carbon deposited may change the accessible catalytic surface area. This is quantified by the ratio of the active catalytic surface area to the geometric surface area (Fcat/geo). The reservoir model keeps track of this number. To close the cycle, the steady-state plug flow simulation is repeated with updated inlet conditions and the new state of the surface. 3.1. Steady-State Flow-Field Model. The steady-state plug flow simulation uses the code DETHCEMPLUG, which is designed for a nondispersive, one-dimensional flow of a chemically reacting, ideal gas mixture under steady-state conditions. The system of differential algebraic equations describing the plug-flow reactor consists of the continuity equation, the species conservation equation, the energy equation, and the equation of state: Cdep i

3. MODELING APPROACH The reactor model of carbon deposition is a combination of a steady-state plug flow field simulation of a single channel of the catalytic monolith and the transient description of the coke formation on the wall. The numerical simulation is conducted by using the modules DETCHEMPLUG and DETCHEMRESERVOIR of the DETCHEM software package.21 It is assumed that the time scales of gaseous flow and carbon deposition are decoupled. The residence time of the gas inside a single channel is between 50 and 80 ms. However, the deposition time is much larger, on the order of minutes. Thus, it can be assumed that, during the flow of an individual “gas volume” through the catalyst, the properties of the wall remain unchanged. In this case, the flow field and the deposition model do not have to be solved simultaneously−they can call each other sequentially. This procedure is illustrated by the cycle of red arrows in Figure 2. In a single steady-state fluid flow simulation, we assume a given state of deposition on the surface, i.e., a certain coverage that, however, varies with the local position in the reatcor. Two types of surface species are distinguished: fast-reacting species (nondepositing) and slow-reacting surface species (depositing). The time-varied inlet conditions and the locally resolved coverage of deposited species are provided by the RESERVOIR model. The concentrations of the gas-phase species and the fast reacting nondepositing surface species are calculated for each time step of the transient RESERVOIR simulation by a quasisteady-state plug-flow simulation. The gas-phase concentrations at various positions along the channel are passed back to the

continuity equation: Ng

AC

∂(ρu) = AS ∑ si̇Mi ∂z i=1

(2)

species conservation equation: AC

∂(ρuYi ) = Mi(ASsi̇ + A Cωi̇ ) ∂z

(3)

energy equation: ρuA C

∂(cpT ) ∂z

Ng

+

Ng

∑ ωi̇ hiMiA C + ∑ si̇hiMiAS i=1

= UAS(Tw − T )

i=1

(4)

equation of state:

ρM̅ = ρRT

(5)

In the above equations, ρ is the density, u the velocity, Ac the cross-sectional area of the channel, As the catalytic active surface 12272



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coefficients on the surface coverage of species i due to possible changes in the energy potential of the surface. Sticking coefficients are commonly used for adsorption reactions and are converted to conventional rate coefficients according to

area per unit reactor length, Ng the number of gas-phase species, Ṡi the molar rate of the production of species i by surface reactions, ω̇ i the molar rate of the production of species i by the gas-phase reaction, Mi the molecular mass of species i, Yi the mass fraction of species i, cp the specific heat capacity of species i, hi the specific enthalpy of species i, U the overall heat-transfer coefficient, Tw the wall temperature, T the gas temperature, p the pressure, and M̅ the average molecular weight. For the solution of the energy equation, the measured temperature profile is taken as a boundary condition (shown in Figure 1). The system of equations is solved using the differential algebraic equation solver LIMEX. 3.2. Gas-Phase Chemistry Model. Chemical reactions may not only occur on the catalytic surface but also in the gaseous flow. With vik′ (reactants), vik″(products) being the stoichiometric coefficients of species i in the reaction k, the chemical source term of homogeneous reactions can be expressed by ajk Ng ⎛ Yjρ ⎞ ⎜ ⎟ ″ ′ ωi̇ = ∑ (νik − νik)k fk ∏ ⎜ Mj ⎟⎠ j=1 ⎝ k=1

= k fads k

(6)

Here, Kg is the number of gas-phase reactions, kfk the Arrhenius rate coefficient, and ajk the order of reaction k related to the concentration of species j. In case of elementary reactions, the reaction orders ajk in eq 6 equal the stoichiometric coefficients v′ik. A variety of elementary reaction mechanisms are available for modeling homogeneous gas-phase reactions of hydrocarbons, it is referenced by Warnatz et al.22 and references therein. 3.3. Surface Chemistry Model. The dynamics of the locally varying surface coverage of adsorbed species (θi) is determined by (7)

Here, σi indicates the number of surface sites that are occupied by species i; Si̇ the rate of generation or consumption of species i due to adsorption, desorption, or surface reaction; and Γ the surface site density (i.e., the molar number of adsorption sites per catalytic surface area). The surface site density can be calculated from the catalyst material. For rhodium, a surface site density of Γ= 2.77 × 10−9 mol/cm2 is used in this model. At steady state, eq 7 becomes a set of coupled, nonlinear algebraic equations. The total molar production rate of species i by surface reactions is given by Ks

si̇ =

Ng + Ns

∑ (νik″ − νik′ )k fk

∏

k=1

j=1

ν″

c j jk (8)

in which Ks is the number of surface reactions (including adsorption and desorption), vik′ (reactants) and vik″ (products) are the stoichiometric coefficients, and Ng and Ns are the number of gas-phase and surface species, respectively. The concentrations cj of adsorbed species are given in mol/m2. The temperature dependence of the rate coefficients is described by a modified Arrhenius expression: ⎡ε θ ⎤ ⎛ −E ⎞ Ns μ k fk = Ak T βk exp⎜ ak ⎟ ∏ θi ik exp⎢ ik i ⎥ ⎝ RT ⎠ ⎣ RT ⎦ i=1

RT 2πMi

(10)

with S0i being the sticking coefficient and τ the number of sites occupied by adsorbed species. 3.4. Reaction Mechanisms. The steady-state plug flow simulations (DETCHEMPLUG model) for steam reforming of natural gas were performed using the detailed gas-phase reaction mechanism (see section 3.4.1 below) in the coupling with surface reaction mechanism of natural gas steam reforming, which is the combination of detailed C1 surface kinetic model, developed before in our research group27 with lumped reactions for higher hydrocarbons (see section 3.4.2 below). Transient simulations of surface carbon deposition from C2H4 and CO (DETCHEMRESERVOIR model) were carried out under isothermal conditions, using a detailed gas-phase reaction mechanism (section 3.4.1) and detailed surface mechanisms: surface kinetic model for natural gas steam reforming (section 3.4.2) at the case of coking from C2H4; detailed surface mechanism of Boudouard reaction (see section 3.4.3 below), at the case of coking from CO. The both kinetic models C2H4 and CO were coupled with pyrolytic carbon deposition model of Li and Deutschmann11 (see section 3.4.4 below), which were implemented in DETCHEM RESERVOIR code to estimate quantitatively the coke deposition on the surface. 3.4.1. Gas-Phase Reaction Mechanism for Steam Reforming of Natural Gas. Since the available gas-phase reaction mechanisms still have quantitative uncertainties in the prediction of gas-phase pyrolysis,23 we use two different mechanisms given in the literature: (1) the mechanism by Golovitchev et al.,24 which was developed to describe the oxidation of n-heptane and iso-octane, consisting of 690 gas-phase reactions between 130 species, and (2) the mechanism of Curran et al.,25 which was developed for the oxidation of natural gas, consisting of 3128 gas-phase reactions between 289 species. Both gas-phase mechanisms were coupled with the surface mechanisms discussed below. 3.4.2. Surface Reaction Mechanism for Steam Reforming of Natural Gas. A detailed surface reaction mechanism based on C1 chemistry was applied, which was originally developed for the CPOX of methane26 and further adopted for diesel reforming.27 Thereby, the pre-exponential factor for the desorption of CO2(s) was adjusted to be 1.0 × 108 s−1. In addition, the reactions of the C2−C4 alkane species (C2H6, C3H8, C4H10) contained in the natural gas were modeled by global reactions.8 These global reactions were extended in this work by taking into account the adsorption reactions of the olefins: ethylene and propylene. In these reactions, hydrocarbons are cracked down to the surface carbon and hydrogen species. The rates of the global reactions are expressed by Arrhenius function,

Kg

∂θi σs ̇ = ii Γ ∂t

Si0 Γτ

(9)

Here, Ak represents the pre-exponential factor, βk the temperature exponent, and Eak the activation energy of reaction k. The coefficients μik and εik describe the dependency of the rate

⎛ −E ⎞ a b s ̇ = AT β exp⎜ a ⎟c HC csites ⎝ RT ⎠ 12273

(11)



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in which A is the pre-exponential factor, Ea the activation energy, cHC the concentration of the regarded hydrocarbon species, and csites the concentration of the free adsorption sites on the surface. The reaction orders a and b are derived from comparison of experimentally determined and modeled reaction rates. The reactions, as well as the Arrhenius parameters of the global reactions, are shown in Table 1.

supposed to occur in the so-called particle-filler model to describe growth of coke layers.29 The activation energy of acetylene addition reactions for heavy aromatics is low, on the order of 0.1−1 eV in the HACA model.30 Therefore, adsorption reactions of the dominant gas-phase intermediate species on surface active sites can be supposed to be an associative chemisorption with a very low binding energy and do not involve an activation energy. In the present work, no activation is assumed for adsorption reactions of hydrocarbons. All other surface reaction kinetic data were derived from the CVD experimental results of the formation of pyrolytic carbon from hydrocarbons.31,32 Reactions and reaction rate constants of the pyrolytic carbon deposition model are given in Table 3. In the

Table 1. Reactions and Kinetic Parameters for the Adsorption of Ethylene and Propylene on the Rhodium Surface (See eq 11) parameter

value/remark

reaction A (mol, m, s, K) Ea (kJ/mol) a b

CnH2n + 3nRh(s) → 2nH(s) + nC(s) 1138 48 0.495 0.94

Table 3. Reactions and Reaction Rate Constants for Pyrolytic Carbon Growth (1) (2) (3) (4) (5) (6) (7) (8) (9) (10) (11)

3.4.3. Surface Reaction Mechanism for CO Disproportionation. The kinetics for the Boudouard reaction (Table 2) is Table 2. Surface Reaction Mechanism for Coke Formation Caused by the Catalytic Decomposition of CO reaction (1) (2) (3) (4) (5) (6) (7) (8) (9) a

CO2 + Rh(s) → CO2(s) CO2(s) → CO2 + Rh(s) CO + Rh(s) → CO(s) CO(s) → CO + Rh(s) CO(s) + Rh(s) → C(s) + O(s) C(s) + O(s) → CO(s) + Rh(s) CO(s) + O(s) → CO2(s) + Rh(s) CO2(s) + Rh(s) → CO(s) + O(s) C(s) → Cdep + Rh(s)

Ea A (mol, cm, s, K) (kJ/mol) 1.0 × 10−5a 1.0 × 108 5.0 × 10−1a 1.0 × 1013 3.7 × 1018 5.0 × 1021 5.5 × 1018 3.7 × 1021 1.0 × 108

0 21.7 0 133.4 169.0 97.9 121.6 115.3 125.5

ref 27 this work 27 27 this work 27 this work 27 this work

reaction

reaction rate constant, k (s−1)

H2 + 2* → H2(*) H2(*) → H2 + 2* C2H4 + 2* → C2H4(*) C2H4(*) → C2H4 + 2* C2H4(*) → 2Cdep + 2H2(*) C2H2 + 2* → C2H2(*) C2H2(*) → C2H2 + 2* C2H2(*) → 2Cdep + H2(*) C6H6 + 6* → C6H6(*) C6H6(*) → C6H6 + 6* C6H6(*) → 6Cdep + 3H2(*)

1.0 × 108 7.5 × 109 2.5 × 107 1.0 × 106 0.9 × 102 2.2 × 108 1.5 × 106 5.4 × 103 1.8 × 106 3.0 × 106 1.0 × 109

present work, this carbon deposition model was implemented in the DETCHEMRESERVOIR code, and corresponding kinetic data were adjusted to satisfy both the process conditions of the coke formation at low temperatures and the surface properties of the rhodium catalyst.

4. RESULTS AND DISCUSSION 4.1. Steam Reforming of Natural Gas. 4.1.1. Catalytic Steam Reforming over Technical Rhodium Catalysts. Steam reforming of natural gas over technical Rh/Al2O3 catalysts was carried out in the temperature range of 923−1073 K. The H and C yields YH/C of the product species P (H2, CO, CO2) were P calculated by

Sticking coefficient.

derived in this work. It consists of the adsorption/desorption steps and surface reactions of CO and CO2 on rhodium from our former mechanism27 with small corrections of some collision factors to experimental data. The mechanism is based on the key reaction intermediate, adsorbed atomic oxygen produced by CO dissociation, because this is known to play an important role in the mechanism of steam reforming and the water−gas shift reaction. The unity bond index-quadratic exponential potential (UBI-QEP) formalism28 was used to determine the heat of adsorption for molecular adsorbents, reaction enthalpy changes, and activation energies for elementary steps. The kinetic parameters for the carbon formation step (9) in the deposed carbon layer were derived from comparison of experimentally determined and modeled reaction rates. 3.4.4. Surface Reaction Mechanism for Pyrolytic Carbon Deposition. The coke formation mechanism used in the present work was originally developed for the deposition of pyrocarbon from light hydrocarbons in the high-temperature range over 1100 K.11,12 The surface reaction mechanism takes into account the inhibition effect of hydrogen on coke formation and includes the reversible adsorption/desorption reactions of the dominant gas-phase intermediate species (hydrogen, ethylene, acetylene, and benzene) and the irreversible dehydrogenation−carbon-deposition reactions of carbon precursor adsorbates. The chemisorption of light hydrocarbons, as well as the physisorption of aromatics, are

YPH/C =

vPH/CnṖ H/C in vfuel n fuel ̇

where ṅP is the molar flow rate of the product and ṅinfuel is the molar inlet flow rate of the fuel; vH/C fuel denotes the number of H or C atoms in the product species P or the fuel, respectively. The calculated H and C yields correspond to the mole fractions of H2, CO, and CO2 shown in Figure 3. With increasing temperature, the yields of H2 and CO increase, whereas the yield of CO2 decreases. The maximum yields of H2 and CO are 75% and 39%, respectively, for a temperature of 1073 K, whereas the yield for CO2 shows a maximum of 11% at 973 K. The experimentally measured distribution of main products during SR of natural gas can be satisfactorily modeled, especially in the low-temperature regime. Figure 4 reveals the experimentally determined and numerically predicted mole fractions of the olefin byproducts C2H4 and C3H6 during SR of natural gas. For low temperatures (923 K), the concentrations of C2H4 and C3H6 in the exhaust gas are at the ppm level. For high temperatures (above 973 K), the mole 12274



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as the cause for the production of C2H4 and C3H6 during SR of natural gas at 923 K. Simulations of SR of natural gas in the gas phase were performed using the Curran mechanism.25 The simulated mole fractions of C2H4 and C3H6 matches the experimental results accurately; even the concentrations of byproducts are very low (at the ppm level). (For C2H4, the experimentally measured mole fraction was 2542 ppm, while the simulated mole fraction was 2205 ppm. For C3H6, the experimentally measured mole fraction was 804 ppm, whereas the simulated mole fraction was 831 ppm.) This extremely good agreement supports the conclusion from above, i.e., accurate models for radial diffusion as well as adsorption and desorption of intermediates is crucial for understanding the interaction of gas-phase and surface kinetics in high-temperature catalysis. 4.2. Coke Formation from Gas-Phase Precursors C2H4 and CO. Both C2H4 and CO are well-known coke precursors; we refer the reader to refs 3, 6, and 12 and refs 7, 14, and 15, respectively. The impact of different feed concentrations of C2H4 and CO was studied to identify the rate of coke formation. Thereby, the process conditions, i.e., feed concentration and temperature, were adjusted to correspond to those of SR of natural gas at 923 K and an S/C of unity over Rh catalysts. In our previous SR experiments33 of natural gas under the given conditions, the CO concentration was measured to vary between 6 vol % and 8 vol %, whereas the C2H4 concentration varies between the ppm level and 0.25 vol % (both concentrations measured in the dried product gas flow). 4.2.1. Coking by C2H4. The amount of coke deposited was analyzed depending on the C2H4 feed concentration. The experiments were performed twice for each C2H4 concentration tested to ensure reproducibility. Figure 5 presents the experimentally

Figure 3. Mole fractions of main products as a function of inlet temperature in steam reforming of natural gas at S/C = 1, GHSV = 16000 h−1; symbols represent experimental data, lines represent computational data.

Figure 4. Mole fractions of byproducts as a function of inlet temperature in steam reforming of natural gas at S/C = 1; symbols represent experimental data, lines represent computational data.

fraction of C2H4 strongly increases, whereas the mole fraction of C3H6 increases slightly and linearly with temperature. The gas phase was simulated using the Curran mechanism,25 combined with global reactions for olefin cracking (Table 1) Even though, the model significantly underpredicts the formation of ethylene and overpredicts the formation of propylene, the model is able to match the temperature dependence and is on the same order of magnitude. Taking into account infinite radial mass transfer (plug-flow model) and the simplifications concerning the coupling radical reactions at the gas/surface interface, a better quantitative agreement was not expected. 4.1.2. Noncatalytic Steam Reforming in the Gas Phase. SR of natural gas at 923 K and S/C = 1 was performed in the gas phase using an empty quartz glass tube. Thereby, the main products of SR (H2, CO, and CO2) were formed only in insignificant amounts. For the chosen conditions, gas-phase reactions are not significant for the conversion of natural gas. However, in the gas phase, SR side products (C2H4 and C3H6) are produced. The maximum flow of C2H4 in the product stream corresponds to 5 mmol/min and that of C3H6 corresponds to 2 mmol/min. These amounts of C2H4 and C3H6 are on the same order as in catalytic SR study, which indicates gas-phase reactions

Figure 5. Time-integral amount of coke deposited on the honeycomb catalyst as a function of C2H4 feed concentration measured by temperature-programmed oxidation (TPO) and computed by transient simulations of the deposition process; time-on-stream 20 min at 823 K; symbols represent experimental data, lines represent computational data.

measured amount of coke resulting from feeding the precursor C2H4 at 823 K for 20 min through the catalytic reactor described above. The amount of carbon deposited was determined by TPO of the coked catalyst (20 K/min, up to 923 K). With increasing C2H4 concentration in the feed (rest nitrogen), the amount of coke increases monotonically. During the catalytic decomposition of C2H4, the formation of several species (e.g., H2, CH4, and C2H6) resulting from cracking processes was detected at the outlet of the reactor. 12275



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computed concentration profiles of C2H4 and H2 as a function of time for the case of an initial C2H4 feed of 1.13 vol %. With increasing time-on-stream, the mole fraction of C2H4 decreases without reaching steady state; the coke layer still grows. Also the concentration of C2H4 decreases along the catalyst length, for all time steps. Thus, the concentration profile of C2H4 along the catalyst channel corresponds to the deposited amount of coke for lowering C2H4 concentrations (Figure 6). Simultaneously, with a decreasing amount of C2H4, the amount of H2 rapidly increases within the first millimeter. Downstream one millimeter of the catalyst channel, however, the H2 spatial profile becomes constant for all reaction times. Figure 8 shows the dependency of the computed amount of coke from C2H4 as a function of both the catalyst length and the

The simulation of the catalytic decomposition of C2H4 was performed by utilizing the software DETCHEMRESERVOIR with concentrations of C2H4 varying between 0.07 vol % and 1.13 vol %. Thereby, first a surface site density of the storage species (coke) was approximated by the calculated site density of graphene (Γgraphene = 6.6 × 10−9 mol/cm2, based on an C−C atom distance of 0.142 nm) and was then fitted to experiments, resulting in a site density for coke of Γcoke = 9 × 10−9 mol/cm2. The concentration range (0.07−1.13 vol %) represents the amount of C2H4 built in the SR of natural gas at temperatures of 923−1073 K and at S/C = 1. Using an updated version of the mechanism by Li12 (Table 3), the coke formation during SR caused by C2H4 could be described quantitatively (see Figure 5). In Figure 6, the computed axial profiles of the amount of carbon deposited along the catalyst channel is shown at varying

Figure 6. Computed differential (length scale) amount of coke deposited along a single catalyst channel as a function of position in the catalyst for different C2H4 feed concentrations; time-on-stream = 20 min at 823 K.

Figure 8. Catalytic decomposition of C2H4: amount of coke deposited as a function of time-on-stream and position inside the catalyst (x = 0.15−0.16 m) for C2H4 concentration of 0.07 vol % at 823 K.

C2H4 concentration in the feed gas. For each C2H4 concentration, the maximum of the coke deposition rate is located within the first millimeter of the catalyst, decreasing downstream along the channel. For the highest concentration of 1.13 vol % C2H4, the amount of coke deposited is uniform downstream of ∼1 mm catalyst length; at low ethylene concentration, the amount of coke deposited decreases continuously along the axial coordinate. Figure 7 reveals the

reaction time for the case of 0.07 vol % C2H4 feed concentration, as discussed in a summary of the findings. The reaction time was chosen to be 20 min. The highest amount of coke deposition is identified along the first millimeter of the catalyst channel with an almost-linear rate. 4.3.2. Coking by CO. Analogous experiments and computations were performed with a feed containing CO at varying concentrations at 823 K and a time-on-stream of 20 min. The CO feed concentration was varied in the range of 1.2−10.92 vol %, corresponding to the amount of CO detected during the SR of natural gas at temperatures of 923−1073 K at S/C = 1. Aside from amount of coke experimentally determined by TPO, the integral amount of gaseous CO2 is reported as measured at the outlet (Figure 9). The computation is performed with DETCHEMRESERVOIR and uses the reaction mechanism shown in Table 2. In the simulations of the coke deposition by CO, the same surface site density of the storage species (coke) (Γcoke = 9 × 10−9 mol/cm2) was used as for the simulations of the catalytic decomposition of C2H4. The coke precursor CO leads to a different coking behavior on the rhodium catalyst compared to the coking behavior of C2H4. With increasing CO feed concentration, the coke formation increases until an apparent saturation is reached at higher CO feed concentration. CO2 is formed simultaneously with the deposition of coke during the CO stream over the catalyst.

Figure 7. Computed concentration profiles of C2H4 and H2 along the catalyst for varying time-on-stream of C2H4 feed concentration of 1.13 vol % at 823 K in coking by ethylene. 12276



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concentrations. For increasing CO feed concentration, the total coke deposits increase, having their local maxima at the catalyst inlet. However, for the highest CO concentration of 10.92 vol %, the shape of the axial deposition profile is flatter. This tendency for higher CO concentrations is explained by the saturation of the coke deposition, i.e., all adsorption sites of the catalyst are already occupied from the beginning of the catalyst channel onward. In addition, the computed gas-phase species profiles of CO and CO2 along the catalyst channel are shown in Figure 11 for the case of 1.2 vol % after a reaction time of 300 s. The mole fraction of the reactant species CO decreases by 0.024 × 10−2 from the feed value of 1.200 × 10−2 to 1.176 × 10−2 at the catalyst exit, while CO2 is continuously formed reaching 0.012 × 10−2 at the catalyst exit, according to the Boudouard reaction.

5. CONCLUSIONS In this study, steam reforming (SR) of natural gas in the temperature range of 923−1073 K (inlet) was investigated focusing on the deactivation of Rh/Al2O3 honeycomb catalysts by coking. Aside from SR, coking experiments were carried out by decomposition of the coke precursors ethylene and carbon monoxide formed in SR. The experimental data were interpreted by comparison with steady-state and transient numerical simulations using detailed surface and gas-phase reaction mechanisms. The comparison of catalytic SR and gas-phase SR revealed that the coking precursors C2H4 and C3H6 are already formed in the gas-phase before the natural gas reaches the catalytic zone. Therefore, the effect of the coke precursors C2H4 and CO on the coke formation was investigated more systematically. A newly developed simple surface reaction and deposition mechanism based on the mean-field approximation is able to model the coke formation rate as well as the production of CO2 via the Boudouard reaction. The amount of coke resulting from the catalytic decomposition of C2H4 is ∼25 times higher than the amount of coke resulting from CO. These differences in the rate of coke formation suggest that coke formation from CO could be negligible compared to coke formation from C2H4 in modeling coke formation during SR of natural gas. However, the analysis of the amount of coke formed by catalytic decomposition in the concentration range of C2H4 or CO occurring in the SR process of natural gas over Rh/Al2O3 catalysts at 923 K reveal similar significance of both species. Consequently, modeling coke formation in SR of natural gas calls not only for a detailed description of coking by the Boudouard mechanism but also for the detailed description of gaseous coke precursors such as ethylene and propylene. Therefore, elementary-step reaction mechanisms for the potential gas-phase reactions are recommended to be implemented into the numerical simulation of the SR reactor.

Figure 9. Time-integral amount of coke deposited on the honeycomb catalyst (blue rectangles rerpresent experimental data; green line represents the simulation) and time-integral amount of CO2 leaving the reactor (red triangles) as a function of CO feed concentration; time-on-stream = 20 min at 823 K.

Figure 10. Computed differential amount (length scale) of coke deposited along a single catalyst channel as a function of position in the catalyst for different CO feed concentrations; time-on-stream = 20 min at 823 K.

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AUTHOR INFORMATION

Corresponding Author

*Tel.: +49 721 608 4 3064. Fax: +49 721 608 4 4805. E-mail: [email protected]. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS The authors would like to thank the Umicore AG & Co KG for providing the commercial catalysts. The German Research Foundation (DFG) is gratefully acknowledged for financial support. The authors would also like to thank Canan Karakaya for performing the chemisorption measurements.

Figure 11. Computed concentration profiles of CO and CO2 in the gas phase along a single catalyst channel during the catalytic decomposition of CO (1.2 vol %) after a reaction time of 300 s.

In Figure 10, the computed amounts of coke deposited along the catalyst channel are compared for the different CO feed 12277



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