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Comparison of the performance of a direct-contact bubble reactor and an indirectly heated tubular reactor for solar-aided methane dry reforming employing molten salt
Chemical Engineering and Processing 83 (2014) 56–63
Contents lists available at ScienceDirect
Chemical Engineering and Processing: Process Intensification journal homepage: www.elsevier.com/locate/cep
Comparison of the performance of a direct-contact bubble reactor and an indirectly heated tubular reactor for solar-aided methane dry reforming employing molten salt Khalid Al-Ali ∗ , Satoshi Kodama, Hidetoshi Sekiguchi Department of Chemical Engineering, Tokyo Institute of Technology, 2-12-1 Ookayama, Meguro-ku, Tokyo 152-8552, Japan
a r t i c l e
i n f o
Article history: Received 13 June 2014 Received in revised form 21 July 2014 Accepted 24 July 2014 Available online 4 August 2014 Keywords: Dry reforming Direct contact bubble reactor Tubular reactor Heat transfer Simulation
a b s t r a c t A theoretical approach is presented for the comparison of two different atmospheric pressure reactors—a direct-contact bubble reactor (DCBR) and an indirectly heated tubular reactor (IHTR)—to evaluate the reactor performance in terms of heat transfer and available catalytic active surface area. The model considers the catalytic endothermic reactions of methane dry reforming that proceeds in both reactors by employing molten salts at elevated temperatures (700–900 ◦ C) in the absence of catalyst deactivation effects. The methane conversion process is simulated for a single reactor using both a reaction kinetics model and a heat transfer model. A well-tested reaction kinetics model, which showed an acceptable agreement with the empirical observations, was implemented to describe the methane dry reforming. In DCBR, the heat is internally transferred by direct contact with the three phases of the system: the reactant gas bubbles, the heat carrier molten salts and the solid catalyst (Ni-Al2 O3 ). In contrast, the supplied heat in the conventional shell-and-tube heat exchanger of the IHTR is transferred across an intervening wall. The results suggest a combination system of DCBR and IHTR would be a suitable configuration for process intensification associated with higher thermal efficiency and cost reduction. © 2014 Elsevier B.V. All rights reserved.
1. Introduction There is currently great interest in the development of solar fuels as concentrated energy carriers with long-term storage capacity. Solar fuels, such as hydrogen, are expected to become an important contributor to our future energy systems. At the present time, one of the most promising processes for producing hydrogen at a reasonable cost is hydrocarbon reforming, especially from methane. Methane dry reforming is particularly appealing because both methane and carbon dioxide are relatively inexpensive, naturally abundant, and readily available. The highly endothermic reaction of methane dry reforming, which requires substantial energy input, can be rendered by solar-aided, thereby offering, in principle, a real possibility to lower the cost for introducing renewable hydrogen production technologies to the market by a combination of fossil fuels and solar energy [1,2]. This research field is currently growing very rapidly. Many studies of catalytic dry reforming have been reported. Nickel-based and noble metal catalysts have been used for the
∗ Corresponding author. Tel.: +81 3 5734 2110; fax: +81 3 5734 2110. E-mail addresses: [email protected], [email protected] (K. Al-Ali). http://dx.doi.org/10.1016/j.cep.2014.07.005 0255-2701/© 2014 Elsevier B.V. All rights reserved.
reaction because they are relatively cheap and stable, respectively [3,4]. In conventional fired tubular reactors, the required heat for the reforming reaction is transferred indirectly by the flue gas that flows through the annulus of a concentric-type heat exchanger. The high endothermicity of the reforming reaction and limited wall heat transfer coefficient necessitate considerable energy input for such reactors [5]. Additionally, the reformer tubes are subjected to very large stresses because they operate at high temperatures [6]. In contrast to fired tubular and exchange reforming, autothermal reforming, where the heat for the reforming is supplied internally by combustion of part of the reactants, combines exothermic oxidation and endothermic reforming reactions in a single reactor [7]. Although the autothermal processes are efficient, they are not flexible in terms of the choice of fuel and the cost of the supplied heat is high in comparison with free, abundant solar energy. Moreover, the cost of the air separation unit can account for up to 40% of the total cost of a syngas plant [2]. Accordingly, solar energy has been exploited in many investigations addressing the practicality of solar-aided methane reformers. However, the intermittent nature of solar energy, its low spatial density and inconvenient distribution are disadvantages compared with conventional resources [8]. Catalytic methane reforming with CO2 requires stable operation despite the fluctuations caused by
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Nomenclature Acronyms DCBR direct contact bubble reactor IHTR indirectly heated tubular reactor Latin and Greek s˙ molar surface rate of production by surface kinetic reaction per unit surface area (mol/cm2 s) ˙ m mass flow rate (kg/s) y mole (or mass) fraction molar enthalpy (kJ/mol) H Cp specific heat capacity (kJ/kg K) T temperature (K or ◦ C) TCASA total catalytic active surface area available in a system (DCBR) (cm2 ) TCASA ϕ total catalytic active surface area per unit reactor length (IHTR) (cm) TCASA ∗ critical state of a catalyst at which performance of DCBR = IHTR (cm2 ) residence time (s) res gas density (kg/cm2 ) V gas volume in reactor (cm3 ) reactor diameter of inner tube (m) Dt convective heat transfer coefficient (W/m2 K) h molecular weight (kg/kmol) MW Ni site density of active nickel, 2.6 × 10−9 mol/cm2 (mol/cm2 ) site fraction of species k rate constant of reaction (depends on reaction) molar concentration of site species (gas species) [Y] (mol/cm2 (mol/cm3 )) stoichiometric coefficient, ( reactant, product) Subscripts i reactant or product chemical species reaction index j in flow in Out flow out forward reaction f r reverse reaction w reactor wall gas G MS molten salt
cloud passage. Moreover, the solar chemical receiver reactor to which the concentrated solar radiation is directed requires thermal uniformity [9]. One effective solution to overcome these drawbacks is to use intermediate working fluids (i.e. molten salts) that provide a constant-rate solar heat supply for stable solar-aided methane reforming [2,10]. Typical reformers employing molten salts as the heat transfer medium can be classified broadly as director indirect-contact reformers. The supplied heat in the indirectly heated tubular reactor (IHTR) can be transferred by heat exchange across an intervening wall separating the methane reforming reaction from the molten salt. The performance of such “double-walled” tubular absorbers/reformers with molten salt, in particular, molten alkali carbonate mixtures, was experimentally demonstrated with a 5-kW dish-type solar concentrator [11]. In contrast, the heat for the reforming reactions in a direct-contact bubble reactor (DCBR) [10,12] is supplied internally by direct contact with the three phases of the system: the reactant gas, the heat carrier molten salt and the solid catalyst.
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Many studies have reported that the Ni-Al2 O3 catalyst is deactivated as a result of coking [13], poisoning, Ni dissolution and sintering [14]. Our previous experimental observations [12] revealed the low catalytic activity of DCBR containing carbonatebased molten salts, in comparison with IHTR, as a consequence of liquid film formation during bubble–catalyst attachment. However, in terms of the heat transfer coefficients, and because of the absence of an intervening wall separating the processing fluids, DCBR has still the advantage of higher thermal efficiency over the traditional shell-and-tube heat exchanger of IHTR. In this study, a mathematical model has been used to examine the reaction performance of DCBR by comparison with IHTR. The motivation is to determine the most energy efficient reactor and configuration in terms of heat transfer and active surface area of the catalyst in the absence of deactivation effects. A certain available active surface area of the catalyst is required for DCBR to approach IHTR under specific heat transfer conditions. The conversion process of the methane dry reforming reaction was simulated for a single reactor to evaluate the performance of both reactors. A steady-state model is presented which takes into account variables including flow modes, inlet and outlet temperatures and the available catalytic activity. 2. Physical aspects of DCBR and IHTR The present study compares DCBR and IHTR theoretically. Schematic representations of both reactors are shown in Fig. 1. In DCBR, the reactant gas stream (a mixture of CH4 and CO2 ) is injected at the bottom of the reactor. The gas is dispersed as bubbles, moving as a counter flow for heat absorption by the thermal fluid (i.e. molten salt). The reactant bubbles come into direct contact with the suspended catalyst particles (i.e. Ni-Al2 O3 ) which is present throughout the thermal fluid. Consequently, the endothermic reaction proceeds across the available contact area, termed as the total catalytic active surface area (TCASA ). Detailed modeling and simulation data are reported in Ref. [12]. In a double-walled IHTR, the reactants enter the inner tube while the molten salt moves in the reverse direction through the outside tube. The inner tube is filled with a porous catalyst. A similar reactor was previously simulated using two overall reactions for optimizing a methane steam reformer [15]. However, the present study uses a previously reported, well-tested, micro-kinetics reaction model [12], developed to describe catalytic methane reforming by CO2 over Ni-based catalysts. The contents of the DCBR are assumed to be nearly spatially uniform because of the high diffusion rates and the forced turbulent mixing provided by the mechanical agitation of the fine bubbles generated throughout the thermal fluid. That is, the rate of conversion of reactants to products is controlled by chemical reaction rates and not by mixing processes. The isothermality and bubble flow of DCBR can be approximated by a perfectly stirred reactor model. In the case of IHTR, the temperature profile can play a decisive role in defining the appropriate model along the reactor. The IHTR model can predict non-isothermal regimes along the bed (assuming it is isothermal in the lateral dimension) by the convergence of a series-connection of perfectly stirred reactor models to a “plug flow reactor model”. 3. Reaction kinetics and rate expression A previously reported reaction kinetics mechanism, developed to describe methane dry reforming over nickel-based catalysts [12], was considered for numerical calculations in the DCBR. The mechanism, composed of 21 reversible elementary reactions, exhibited satisfactory results when tested using a fixed bed reactor, as well as with the DCBR. The numerical calculation of the
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Fig. 1. Physical domains for (a) the DCBR and (b) the IHTR models.
mechanism requires solving systems of nonlinear algebraic equations to compute the molar rate of production, s˙ i , for each species. The steady-state equations describing the well-mixed reactor can be classified as the mass conservation of the gas-phase species (i.e. CH4 , CO2 , Ar, H2 , CO, H2 O and O2 ) and the surface site conservation equation, which must hold true for each of the 13 surface species on the nickel sites including free nickel sites. The site species are CH4 (s), CO2 (s), H(s), CO(s), H2 O(s), CH3 (s), CH2 (s), CH(s), C(s), OH(s), HCO(s), O(s), and Ni(s). The production rates of the gas-phase and surface species, s˙ i , can be expressed as the sum of the reaction rates involving the species i in the reaction kinetics mechanism, calculated as the difference between the forward and reverse rates as follows:
J
s˙ i =
ij
kfj
j=1
I
[Yi ]
i=1
ij
I
− krj
[Yi ]
ij
(1)
i=1
For adsorption, desorption and surface reactions, the molar concentration [Yi ] of surface species i can be written in terms of the site coverage fraction as follows: [Yi ] = i Ni
(2)
3.1. Conservation equation for gas-phase species (steady-state) V (yi − yi,in ) = res TCASA s˙ i MWi
(7 equations)
(i = CH4 , CO2 , Ar, H2 , CO, H2 O and O2 )
(3)
concentration) must be equal to zero. Therefore, the steady-state equation for surface species is given by Eq. (1) at s˙ i = 0. To obtain a well-posed system of equations, one must introduce the additional constraints of the system. For example, that the sum of all surface site fractions, i , is one; this also applies to all mole fractions, yi , as described in the following equations:
i = 1
(5)
yi = 1
(6)
To solve the set of nonlinear equations, a modified Newton iteration method was implemented using numerical software (MatLab® ). There are 20 implicit equations associated with 20 unknowns. The initial estimates suggested the use of thermodynamics equilibrium compositions to ensure the “constrained” algorithm converges to the solution. Thermodynamics properties for all species considered were collected from the Chemkin thermodynamic database in the NASA chemical equilibrium program [16]. The chemical kinetics calculation is a prerequisite for the heat transfer model and the calculations are explained in the following section.
4. Heat transfer equations and coefficients 4.1. DCBR
The gas-phase species equation (Eq. (3)) accounts for the change in mole fraction of species as the result of the catalytic reactions over the nickel surface which is characterized by the total active surface area TCASA .
The heat transfer in the DCBR is simulated as a perfectly stirred reactor model in a zero-dimensional homogenous system for a single-domain reactor space:
3.2. Conservation equation for surface-species (steady-state)
˙ G,in ˙ MS Cp,MS T = m m
i=1
d[Yi ] = s˙ i = 0 dt
(Hi s˙ i )
(7)
i=1
Eq. (7) accounts for the energy transport by the molten salt associated with the reaction heat plus the heat required to raise the gas temperature to the specified reaction temperature.
(13 equations)
(i = CH4 (s), CO2 (s), H(s), CO(s), H2 O(s), CH3 (s), CH2 (s), CH(s), C(s), OH(s), HCO(s), O(s), and Ni(s))
yi,in (Hi − Hi,in ) + TCASA
(4)
The equation of the surface site species at steady state (Eq. (4)) implies that the molar surface production rate of each surface species on the nickel sites (or the rate of change in the molar
4.2. IHTR The heat transfer in IHTR is simulated in one axial dimension for three different domains of the reactor space, as shown in Fig. 1 and described as follows.
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4.2.1. Molten salt space The temperature gradient of the hot molten salt with x-axial distance is given by the differential equation ˙ MS Cp,MS m
dTMS = Dt hMS (TMS − Tw ) dx
(8)
4.2.2. Reactor wall domain High thermal conductivity of the reactor wall is presumed; therefore, the temperature gradient across the reactor tube is negligible. Accordingly, the net energy convected by the molten salt into the wall domain is equal to the net energy transferred into the reactant gas domain by heat convection. Hence, the temperature of the wall is taken as the average of the molten salt and reactant gas temperatures: Tw =
hMS hG TMS + TG hMS + hG hMS + hG
(9)
4.2.3. Gas space Gas energy transport determines the amount of energy convected by the gas and the heat required for the catalytic reaction. The temperature gradient of the reactant and/or product gases with respect to x-axial distance is given by the differential equation ˙ G,in Cp,G m
dTG = Dt hG (Tw − TG ) − TCASAϕ (Hi s˙ i ) dx
(10)
i=1
The set of ordinary differential equations (Eqs. (8)–(10)) is numerically solved with boundary conditions at x = 0 for a reactor gas inlet temperature of 25 ◦ C, at a CH4 :CO2 :Ar molar ratio of 1:1:1 and an outlet molten salt temperature of 700 ◦ C. The gas flows in a tube diameter of 0.02 m with a volumetric flow rate of 100 L min−1 . A molten salt system with a specific heat capacity of 0.001842 kJ g−1 K−1 flows with a mass rate of 30 g s−1 . The same conditions were applied for the DCBR in Eq. (7) with an estimated residence time of 1.0 s. The employed convective heat transfer coefficient for the molten salt, hMS , was estimated to be approximately 300 W m−2 K−1 , depending on geometry. While, the value of hG was estimated for the turbulent gas flow inside a smooth tube using Gnielinski correlation [17], with gas transport properties, viscosity and thermal conductivity, calculated based on Chapman–Enskog model [18], and the Wilke mixing rule [19]. These values of convective heat transfer coefficients are presented in Fig. 2. The slight perturbation at 370 cm represents the initiation of the catalytic reaction. 5. Results and discussion The reaction performance of the reactor was analyzed using the conversion process of the methane dry reforming reaction. The methane conversion can be defined with respect to the outlet concentrations of the gas reaction products, without carbon deposition, as follows: Conversion (CH4 )% = 1 −
[CH4 ] ([CH4 ] + [CO2 ] + [CO]) × 0.5
(11)
5.1. Effect of flow mode in the IHTR The effect of the flow mode in the IHTR was considered to optimize the flow arrangement for achieving higher reactor performance. Simulations were performed for both the co-current and counter-current flow modes at the same flow rates and under the same reaction conditions. The predicted temperatures of the reformer gas, molten salt and tube wall are presented in Fig. 3(a) for co-current flow and
Fig. 2. Heat transfer coefficients used in the model between the molten salt and the outside reformer tube (hMS ) and between the reformer gas and the inside reformer tube (hG ).
in Fig. 3(b) for counter-current flow as a function of axial distance. The catalytic surface reactions are programmed to commence at the moment the gas temperature reaches 700 ◦ C; hence, the required tube length can be estimated (ca. 3.7 and 4.9 m) for the countercurrent and co-current flows, respectively. This pre-heating process is required for the IHTR to heat the reactant gas to 700 ◦ C. The counter-current exchanger requires a smaller reactor size because the temperature difference between the hot molten salt stream and the cold gas stream at the cold end of the exchanger is higher than for the co-current exchanger. This is practically feasible because a zero temperature difference would require an infinitely large heat exchanger [20]. Moreover, the comparative profiles, given in Fig. 4, reveal that the counter-current flow mode of the IHTR is more efficient, in terms of methane conversion, than the co-current flow mode. This is attributed to higher reformer gas temperature in the counter-current than in the co-current flow (740 and 700 ◦ C, respectively; Fig. 3). By considering a constant gas temperature in the DCBR of 700 ◦ C, the predicted gas temperatures in the DCBR and IHTR are plotted in Fig. 5. In the pre-heating zone (i.e. x < 5.5 m), the reactor performance order is DCBR > IHTR. Apart from the preheating process (i.e. at x > 5.5 m), the reactor performance order would be IHTRcounter-current > DCBR > IHTRco-current . This is attributed to the high reaction temperatures, which favor methane dry reforming. 5.2. Comparison between the DCBR and IHTR After pre-heating in the counter-current flow mode in the IHTR, the catalytic dry reforming reaction proceeds along the reactor length with a fixed bed catalyst surface area of 100 m2 per unit length of IHTR (i.e. TCASA ϕ = 100 m). It is predicted that the reforming reaction continues for 1.8 m at gas reaction temperatures less than 700 ◦ C because of the endothermicity of the reaction. Subsequently, the gas temperature approaches 740 ◦ C. In contrast, pre-heating in the DCBR is not a considerable process because the gas temperature instantly increased from room temperature to a reaction temperature of 700 ◦ C, showing a higher overall heat transfer coefficient compared with the IHTR. Accordingly, the catalytic reaction promptly starts and continues at a constant reaction temperature of 700 ◦ C. To examine the effect of the pre-heating process on the catalytic reactions in both the IHTR
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Fig. 4. Comparison of reactor performance for co-current and counter-current flow arrangements in the IHTR (the gas pre-heating process is not considered).
Fig. 5. Predicted gas temperatures for counter-current, co-current IHTR and DCBR. Fig. 3. Predicted temperature profiles for different flow arrangements: (a) cocurrent flow (b) counter-current flow.
and DCBR, the conversion process of dry reforming was simulated with respect to the available catalytic surface area TCASA . Fig. 6 shows the predicted reactor performance results. It is apparent that the performance of the DCBR is superior to that of the IHTR at low available catalytic activities (i.e. TCASA < 180 m2 ). This is attributed to the excellent heat transfer coefficient of the DCBR and the consequently higher thermal efficiency. A critical catalyst activity can be defined, termed TCASA ∗ , at which the performances of the DCBR and IHTR are equal under the specific reaction conditions. Therefore, Fig. 6 illustrates that when TCASA < TCASA ∗ , the DCBR will perform better than the IHTR in terms of reaction conversion and vice versa. The predicted reactant and product gas concentrations are depicted for both reactors in Fig. 7 as a function of catalyst surface area and/or axial distance. The mole fractions of H2 and CO formed in the synthesis gas production are higher in the DCBR than the IHTR when TCASA < TCASA ∗ .
Fig. 6. Comparison of reactor performance of the DCBR and IHTR.
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Table 1 Required reactor volume to evaluate the process intensification potential.
Total pre-heating tube length, m Volumea,b (without molten salt), cm3 Volumea,c (with molten salt), cm3 Process volume intensificationd , %
DCBR
IHTR
– 44 972 34.4%
5.5 1730 3892
The reactor volume is estimated at the critical state TCASA ∗ . The DCBR volume is the catalyst volume considering a 6%Ni-Al2 O3 catalyst (nickel active surface area, BETNi = 1.0 m2 /g) and the IHTR volume is the inner tube volume (Dia. 20 mm) of a concentric heat exchanger. c The molten salt volume is estimated for an alkali carbonate mixture system composed of Na2 CO3 , Li2 CO3 and K2 CO3 in a ratio 1:2:2, and considering the bubbly flow reactor of the DCBR to have a catalyst/salt ratio of 0.1. The IHTR volume is the outer tube volume (Dia. 30 mm) of a concentric heat exchanger. d Process volume intensification indicates the ratio of the eliminated volume using the DCBR in the combination system to the total process IHTR volume (i.e. x = 12 m) without using the DCBR. a
b
the counter-current flow of the IHTR, the better the pre-heating process. This reduces the required tube length because the gas temperature approaches the reaction temperature of 700 ◦ C more rapidly. Consequently, the process volume of the IHTR is decreased. The reduction of the percentage volume intensification is as a result of a reduction in the DCBR size TCASA ∗ , required to compensate the pre-heating process. This implies that high molten salt inlet temperatures discourage the catalytic reactions in the DCBR in comparison with the IHTR, and the significant role of DCBR will encounter in the gas pre-heating process to the reaction temperature. In contrast, in the IHTR, the methane conversion at x = 12 m increases as the inlet molten salt temperature increases.
5.3. Combination system of the DCBR and IHTR
Fig. 7. Predicted reformer gas concentrations of reactants and products (a) in the DCBR as a function of catalyst activity and (b) in the IHTR as a function of catalyst activity and axial distance.
Furthermore, the required reactor volumes to achieve equal performance of both reactors at TCASA ∗ are provided in Table 1. Comparison with the IHTR shows that a significant volume reduction is possible. For the reactors studied, process intensification is achieved when heat transfer to the gas side is dominant. Process volume intensification estimates the reduction percentage of the IHTR process volume when utilizing the DCBR. This is because the DCBR can replace the total pre-heating process of the IHTR. Specifically, in the IHTR, a 3.7-m tube is used to heat the gas temperature to the same reaction temperature as in the DCBR (700 ◦ C). This is in addition to the catalytic bed portion of the tube (1.8 m) required to achieve an equivalent performance at TCASA ∗ . The results suggest that the DCBR is the more suitable reactor for process intensification with an effective total volume reduction of 34%. Simulation results, given in Fig. 8, demonstrate that the inlet molten salt temperature controls the IHTR performance. This is in contrast to the DCBR, the performance of which is controlled by the outlet molten salt temperature, as reported in Ref. [12]. Table 2 shows that the higher the temperature of the molten salt entering
The maximum temperature a solar field can achieve depends on how concentrated the sunlight is. High-quality concentrators, however, are very expensive. Therefore, ideally, we are looking to miniaturize the entire system with minimum impact on the conversion process, increase thermal efficiency, and reduce the cost as far as possible. Despite the problem of the low inlet molten salt temperature that reduces the IHTR performance, the combination system of the DCBR and IHTR is still superior in terms of process intensification with higher thermal efficiency. Furthermore, there have been efforts to circumvent the high temperature constraints by the use of a hydrogen-selective membrane installed in the catalyst bed [21]. The membrane continuously removes hydrogen from the reformer gas stream, thereby increasing the conversion at lower temperature. The integrated system of the DCBR and the countercurrent flow mode of the IHTR are depicted in Fig. 9, where the process data are extracted from Fig. 3(b) and Fig. 6.
Fig. 8. Predicted methane conversion in counter-current IHTR at x = 12 m as a function of inlet molten salt temperature.
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Table 2 Effect of inlet molten salt temperature on reactor performance. Inlet molten salt temperature (◦ C)
Critical TCASA ∗ (m2 )
847 767 743 732
33 99 180 264
Total pre-heating length (m) Pre-heating length
Reactor Bed length
2.65 3.30 3.71 4.00
0.33 0.99 1.80 2.64
Volume intensification (%)
CH4 conversion (% @ TCASA ∗ )
22.7 29.4 34.4 38.6
15 25 32 36
Fig. 9. Schematic representation of the combination system of the DCBR and IHTR.
6. Conclusion
Acknowledgment
Owing to the potentially major role of solar-aided methane dry reforming in the chemical industry, two atmospheric pressure reactors, the DCBR and the IHTR, were compared theoretically. The reactor performance was numerically estimated to evaluate the effect of heat transfer on the methane conversion process, along with a validated catalytic reactions mechanism, and considering the effective variables like flow modes, inlet and outlet temperatures of a thermal fluid and the available catalytic active surface area. The studied reactors’ performances were determined to follow the order: IHTRcounter-current > DCBR > IHTRco-current . Nevertheless, the performance of the DCBR was superior to the IHTR at low available catalytic surface areas (i.e. when TCASA < TCASA ∗ ), which was attributed to the ideally heat transfer coefficient of the DCBR, and the consequently higher thermal efficiency. Therefore, the DCBR can replace the pre-heating process required in the counter-current flow mode of the IHTR. This suggests that the DCBR is a suitable reactor for process volume intensification. The relative efficiency of the IHTR is increased at higher inlet molten salt temperatures; however, operation under such conditions is not cost-effective. A better understanding of the limits of the solar-aided methane reforming process has resulted in a more economical system, namely, the combination system of the DCBR and the counter-current flow mode of the IHTR.
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Contents lists available at ScienceDirect
Chemical Engineering and Processing: Process Intensification journal homepage: www.elsevier.com/locate/cep
Comparison of the performance of a direct-contact bubble reactor and an indirectly heated tubular reactor for solar-aided methane dry reforming employing molten salt Khalid Al-Ali ∗ , Satoshi Kodama, Hidetoshi Sekiguchi Department of Chemical Engineering, Tokyo Institute of Technology, 2-12-1 Ookayama, Meguro-ku, Tokyo 152-8552, Japan
a r t i c l e
i n f o
Article history: Received 13 June 2014 Received in revised form 21 July 2014 Accepted 24 July 2014 Available online 4 August 2014 Keywords: Dry reforming Direct contact bubble reactor Tubular reactor Heat transfer Simulation
a b s t r a c t A theoretical approach is presented for the comparison of two different atmospheric pressure reactors—a direct-contact bubble reactor (DCBR) and an indirectly heated tubular reactor (IHTR)—to evaluate the reactor performance in terms of heat transfer and available catalytic active surface area. The model considers the catalytic endothermic reactions of methane dry reforming that proceeds in both reactors by employing molten salts at elevated temperatures (700–900 ◦ C) in the absence of catalyst deactivation effects. The methane conversion process is simulated for a single reactor using both a reaction kinetics model and a heat transfer model. A well-tested reaction kinetics model, which showed an acceptable agreement with the empirical observations, was implemented to describe the methane dry reforming. In DCBR, the heat is internally transferred by direct contact with the three phases of the system: the reactant gas bubbles, the heat carrier molten salts and the solid catalyst (Ni-Al2 O3 ). In contrast, the supplied heat in the conventional shell-and-tube heat exchanger of the IHTR is transferred across an intervening wall. The results suggest a combination system of DCBR and IHTR would be a suitable configuration for process intensification associated with higher thermal efficiency and cost reduction. © 2014 Elsevier B.V. All rights reserved.
1. Introduction There is currently great interest in the development of solar fuels as concentrated energy carriers with long-term storage capacity. Solar fuels, such as hydrogen, are expected to become an important contributor to our future energy systems. At the present time, one of the most promising processes for producing hydrogen at a reasonable cost is hydrocarbon reforming, especially from methane. Methane dry reforming is particularly appealing because both methane and carbon dioxide are relatively inexpensive, naturally abundant, and readily available. The highly endothermic reaction of methane dry reforming, which requires substantial energy input, can be rendered by solar-aided, thereby offering, in principle, a real possibility to lower the cost for introducing renewable hydrogen production technologies to the market by a combination of fossil fuels and solar energy [1,2]. This research field is currently growing very rapidly. Many studies of catalytic dry reforming have been reported. Nickel-based and noble metal catalysts have been used for the
∗ Corresponding author. Tel.: +81 3 5734 2110; fax: +81 3 5734 2110. E-mail addresses: [email protected], [email protected] (K. Al-Ali). http://dx.doi.org/10.1016/j.cep.2014.07.005 0255-2701/© 2014 Elsevier B.V. All rights reserved.
reaction because they are relatively cheap and stable, respectively [3,4]. In conventional fired tubular reactors, the required heat for the reforming reaction is transferred indirectly by the flue gas that flows through the annulus of a concentric-type heat exchanger. The high endothermicity of the reforming reaction and limited wall heat transfer coefficient necessitate considerable energy input for such reactors [5]. Additionally, the reformer tubes are subjected to very large stresses because they operate at high temperatures [6]. In contrast to fired tubular and exchange reforming, autothermal reforming, where the heat for the reforming is supplied internally by combustion of part of the reactants, combines exothermic oxidation and endothermic reforming reactions in a single reactor [7]. Although the autothermal processes are efficient, they are not flexible in terms of the choice of fuel and the cost of the supplied heat is high in comparison with free, abundant solar energy. Moreover, the cost of the air separation unit can account for up to 40% of the total cost of a syngas plant [2]. Accordingly, solar energy has been exploited in many investigations addressing the practicality of solar-aided methane reformers. However, the intermittent nature of solar energy, its low spatial density and inconvenient distribution are disadvantages compared with conventional resources [8]. Catalytic methane reforming with CO2 requires stable operation despite the fluctuations caused by
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Nomenclature Acronyms DCBR direct contact bubble reactor IHTR indirectly heated tubular reactor Latin and Greek s˙ molar surface rate of production by surface kinetic reaction per unit surface area (mol/cm2 s) ˙ m mass flow rate (kg/s) y mole (or mass) fraction molar enthalpy (kJ/mol) H Cp specific heat capacity (kJ/kg K) T temperature (K or ◦ C) TCASA total catalytic active surface area available in a system (DCBR) (cm2 ) TCASA ϕ total catalytic active surface area per unit reactor length (IHTR) (cm) TCASA ∗ critical state of a catalyst at which performance of DCBR = IHTR (cm2 ) residence time (s) res gas density (kg/cm2 ) V gas volume in reactor (cm3 ) reactor diameter of inner tube (m) Dt convective heat transfer coefficient (W/m2 K) h molecular weight (kg/kmol) MW Ni site density of active nickel, 2.6 × 10−9 mol/cm2 (mol/cm2 ) site fraction of species k rate constant of reaction (depends on reaction) molar concentration of site species (gas species) [Y] (mol/cm2 (mol/cm3 )) stoichiometric coefficient, ( reactant, product) Subscripts i reactant or product chemical species reaction index j in flow in Out flow out forward reaction f r reverse reaction w reactor wall gas G MS molten salt
cloud passage. Moreover, the solar chemical receiver reactor to which the concentrated solar radiation is directed requires thermal uniformity [9]. One effective solution to overcome these drawbacks is to use intermediate working fluids (i.e. molten salts) that provide a constant-rate solar heat supply for stable solar-aided methane reforming [2,10]. Typical reformers employing molten salts as the heat transfer medium can be classified broadly as director indirect-contact reformers. The supplied heat in the indirectly heated tubular reactor (IHTR) can be transferred by heat exchange across an intervening wall separating the methane reforming reaction from the molten salt. The performance of such “double-walled” tubular absorbers/reformers with molten salt, in particular, molten alkali carbonate mixtures, was experimentally demonstrated with a 5-kW dish-type solar concentrator [11]. In contrast, the heat for the reforming reactions in a direct-contact bubble reactor (DCBR) [10,12] is supplied internally by direct contact with the three phases of the system: the reactant gas, the heat carrier molten salt and the solid catalyst.
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Many studies have reported that the Ni-Al2 O3 catalyst is deactivated as a result of coking [13], poisoning, Ni dissolution and sintering [14]. Our previous experimental observations [12] revealed the low catalytic activity of DCBR containing carbonatebased molten salts, in comparison with IHTR, as a consequence of liquid film formation during bubble–catalyst attachment. However, in terms of the heat transfer coefficients, and because of the absence of an intervening wall separating the processing fluids, DCBR has still the advantage of higher thermal efficiency over the traditional shell-and-tube heat exchanger of IHTR. In this study, a mathematical model has been used to examine the reaction performance of DCBR by comparison with IHTR. The motivation is to determine the most energy efficient reactor and configuration in terms of heat transfer and active surface area of the catalyst in the absence of deactivation effects. A certain available active surface area of the catalyst is required for DCBR to approach IHTR under specific heat transfer conditions. The conversion process of the methane dry reforming reaction was simulated for a single reactor to evaluate the performance of both reactors. A steady-state model is presented which takes into account variables including flow modes, inlet and outlet temperatures and the available catalytic activity. 2. Physical aspects of DCBR and IHTR The present study compares DCBR and IHTR theoretically. Schematic representations of both reactors are shown in Fig. 1. In DCBR, the reactant gas stream (a mixture of CH4 and CO2 ) is injected at the bottom of the reactor. The gas is dispersed as bubbles, moving as a counter flow for heat absorption by the thermal fluid (i.e. molten salt). The reactant bubbles come into direct contact with the suspended catalyst particles (i.e. Ni-Al2 O3 ) which is present throughout the thermal fluid. Consequently, the endothermic reaction proceeds across the available contact area, termed as the total catalytic active surface area (TCASA ). Detailed modeling and simulation data are reported in Ref. [12]. In a double-walled IHTR, the reactants enter the inner tube while the molten salt moves in the reverse direction through the outside tube. The inner tube is filled with a porous catalyst. A similar reactor was previously simulated using two overall reactions for optimizing a methane steam reformer [15]. However, the present study uses a previously reported, well-tested, micro-kinetics reaction model [12], developed to describe catalytic methane reforming by CO2 over Ni-based catalysts. The contents of the DCBR are assumed to be nearly spatially uniform because of the high diffusion rates and the forced turbulent mixing provided by the mechanical agitation of the fine bubbles generated throughout the thermal fluid. That is, the rate of conversion of reactants to products is controlled by chemical reaction rates and not by mixing processes. The isothermality and bubble flow of DCBR can be approximated by a perfectly stirred reactor model. In the case of IHTR, the temperature profile can play a decisive role in defining the appropriate model along the reactor. The IHTR model can predict non-isothermal regimes along the bed (assuming it is isothermal in the lateral dimension) by the convergence of a series-connection of perfectly stirred reactor models to a “plug flow reactor model”. 3. Reaction kinetics and rate expression A previously reported reaction kinetics mechanism, developed to describe methane dry reforming over nickel-based catalysts [12], was considered for numerical calculations in the DCBR. The mechanism, composed of 21 reversible elementary reactions, exhibited satisfactory results when tested using a fixed bed reactor, as well as with the DCBR. The numerical calculation of the
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Fig. 1. Physical domains for (a) the DCBR and (b) the IHTR models.
mechanism requires solving systems of nonlinear algebraic equations to compute the molar rate of production, s˙ i , for each species. The steady-state equations describing the well-mixed reactor can be classified as the mass conservation of the gas-phase species (i.e. CH4 , CO2 , Ar, H2 , CO, H2 O and O2 ) and the surface site conservation equation, which must hold true for each of the 13 surface species on the nickel sites including free nickel sites. The site species are CH4 (s), CO2 (s), H(s), CO(s), H2 O(s), CH3 (s), CH2 (s), CH(s), C(s), OH(s), HCO(s), O(s), and Ni(s). The production rates of the gas-phase and surface species, s˙ i , can be expressed as the sum of the reaction rates involving the species i in the reaction kinetics mechanism, calculated as the difference between the forward and reverse rates as follows:
J
s˙ i =
ij
kfj
j=1
I
[Yi ]
i=1
ij
I
− krj
[Yi ]
ij
(1)
i=1
For adsorption, desorption and surface reactions, the molar concentration [Yi ] of surface species i can be written in terms of the site coverage fraction as follows: [Yi ] = i Ni
(2)
3.1. Conservation equation for gas-phase species (steady-state) V (yi − yi,in ) = res TCASA s˙ i MWi
(7 equations)
(i = CH4 , CO2 , Ar, H2 , CO, H2 O and O2 )
(3)
concentration) must be equal to zero. Therefore, the steady-state equation for surface species is given by Eq. (1) at s˙ i = 0. To obtain a well-posed system of equations, one must introduce the additional constraints of the system. For example, that the sum of all surface site fractions, i , is one; this also applies to all mole fractions, yi , as described in the following equations:
i = 1
(5)
yi = 1
(6)
To solve the set of nonlinear equations, a modified Newton iteration method was implemented using numerical software (MatLab® ). There are 20 implicit equations associated with 20 unknowns. The initial estimates suggested the use of thermodynamics equilibrium compositions to ensure the “constrained” algorithm converges to the solution. Thermodynamics properties for all species considered were collected from the Chemkin thermodynamic database in the NASA chemical equilibrium program [16]. The chemical kinetics calculation is a prerequisite for the heat transfer model and the calculations are explained in the following section.
4. Heat transfer equations and coefficients 4.1. DCBR
The gas-phase species equation (Eq. (3)) accounts for the change in mole fraction of species as the result of the catalytic reactions over the nickel surface which is characterized by the total active surface area TCASA .
The heat transfer in the DCBR is simulated as a perfectly stirred reactor model in a zero-dimensional homogenous system for a single-domain reactor space:
3.2. Conservation equation for surface-species (steady-state)
˙ G,in ˙ MS Cp,MS T = m m
i=1
d[Yi ] = s˙ i = 0 dt
(Hi s˙ i )
(7)
i=1
Eq. (7) accounts for the energy transport by the molten salt associated with the reaction heat plus the heat required to raise the gas temperature to the specified reaction temperature.
(13 equations)
(i = CH4 (s), CO2 (s), H(s), CO(s), H2 O(s), CH3 (s), CH2 (s), CH(s), C(s), OH(s), HCO(s), O(s), and Ni(s))
yi,in (Hi − Hi,in ) + TCASA
(4)
The equation of the surface site species at steady state (Eq. (4)) implies that the molar surface production rate of each surface species on the nickel sites (or the rate of change in the molar
4.2. IHTR The heat transfer in IHTR is simulated in one axial dimension for three different domains of the reactor space, as shown in Fig. 1 and described as follows.
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4.2.1. Molten salt space The temperature gradient of the hot molten salt with x-axial distance is given by the differential equation ˙ MS Cp,MS m
dTMS = Dt hMS (TMS − Tw ) dx
(8)
4.2.2. Reactor wall domain High thermal conductivity of the reactor wall is presumed; therefore, the temperature gradient across the reactor tube is negligible. Accordingly, the net energy convected by the molten salt into the wall domain is equal to the net energy transferred into the reactant gas domain by heat convection. Hence, the temperature of the wall is taken as the average of the molten salt and reactant gas temperatures: Tw =
hMS hG TMS + TG hMS + hG hMS + hG
(9)
4.2.3. Gas space Gas energy transport determines the amount of energy convected by the gas and the heat required for the catalytic reaction. The temperature gradient of the reactant and/or product gases with respect to x-axial distance is given by the differential equation ˙ G,in Cp,G m
dTG = Dt hG (Tw − TG ) − TCASAϕ (Hi s˙ i ) dx
(10)
i=1
The set of ordinary differential equations (Eqs. (8)–(10)) is numerically solved with boundary conditions at x = 0 for a reactor gas inlet temperature of 25 ◦ C, at a CH4 :CO2 :Ar molar ratio of 1:1:1 and an outlet molten salt temperature of 700 ◦ C. The gas flows in a tube diameter of 0.02 m with a volumetric flow rate of 100 L min−1 . A molten salt system with a specific heat capacity of 0.001842 kJ g−1 K−1 flows with a mass rate of 30 g s−1 . The same conditions were applied for the DCBR in Eq. (7) with an estimated residence time of 1.0 s. The employed convective heat transfer coefficient for the molten salt, hMS , was estimated to be approximately 300 W m−2 K−1 , depending on geometry. While, the value of hG was estimated for the turbulent gas flow inside a smooth tube using Gnielinski correlation [17], with gas transport properties, viscosity and thermal conductivity, calculated based on Chapman–Enskog model [18], and the Wilke mixing rule [19]. These values of convective heat transfer coefficients are presented in Fig. 2. The slight perturbation at 370 cm represents the initiation of the catalytic reaction. 5. Results and discussion The reaction performance of the reactor was analyzed using the conversion process of the methane dry reforming reaction. The methane conversion can be defined with respect to the outlet concentrations of the gas reaction products, without carbon deposition, as follows: Conversion (CH4 )% = 1 −
[CH4 ] ([CH4 ] + [CO2 ] + [CO]) × 0.5
(11)
5.1. Effect of flow mode in the IHTR The effect of the flow mode in the IHTR was considered to optimize the flow arrangement for achieving higher reactor performance. Simulations were performed for both the co-current and counter-current flow modes at the same flow rates and under the same reaction conditions. The predicted temperatures of the reformer gas, molten salt and tube wall are presented in Fig. 3(a) for co-current flow and
Fig. 2. Heat transfer coefficients used in the model between the molten salt and the outside reformer tube (hMS ) and between the reformer gas and the inside reformer tube (hG ).
in Fig. 3(b) for counter-current flow as a function of axial distance. The catalytic surface reactions are programmed to commence at the moment the gas temperature reaches 700 ◦ C; hence, the required tube length can be estimated (ca. 3.7 and 4.9 m) for the countercurrent and co-current flows, respectively. This pre-heating process is required for the IHTR to heat the reactant gas to 700 ◦ C. The counter-current exchanger requires a smaller reactor size because the temperature difference between the hot molten salt stream and the cold gas stream at the cold end of the exchanger is higher than for the co-current exchanger. This is practically feasible because a zero temperature difference would require an infinitely large heat exchanger [20]. Moreover, the comparative profiles, given in Fig. 4, reveal that the counter-current flow mode of the IHTR is more efficient, in terms of methane conversion, than the co-current flow mode. This is attributed to higher reformer gas temperature in the counter-current than in the co-current flow (740 and 700 ◦ C, respectively; Fig. 3). By considering a constant gas temperature in the DCBR of 700 ◦ C, the predicted gas temperatures in the DCBR and IHTR are plotted in Fig. 5. In the pre-heating zone (i.e. x < 5.5 m), the reactor performance order is DCBR > IHTR. Apart from the preheating process (i.e. at x > 5.5 m), the reactor performance order would be IHTRcounter-current > DCBR > IHTRco-current . This is attributed to the high reaction temperatures, which favor methane dry reforming. 5.2. Comparison between the DCBR and IHTR After pre-heating in the counter-current flow mode in the IHTR, the catalytic dry reforming reaction proceeds along the reactor length with a fixed bed catalyst surface area of 100 m2 per unit length of IHTR (i.e. TCASA ϕ = 100 m). It is predicted that the reforming reaction continues for 1.8 m at gas reaction temperatures less than 700 ◦ C because of the endothermicity of the reaction. Subsequently, the gas temperature approaches 740 ◦ C. In contrast, pre-heating in the DCBR is not a considerable process because the gas temperature instantly increased from room temperature to a reaction temperature of 700 ◦ C, showing a higher overall heat transfer coefficient compared with the IHTR. Accordingly, the catalytic reaction promptly starts and continues at a constant reaction temperature of 700 ◦ C. To examine the effect of the pre-heating process on the catalytic reactions in both the IHTR
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Fig. 4. Comparison of reactor performance for co-current and counter-current flow arrangements in the IHTR (the gas pre-heating process is not considered).
Fig. 5. Predicted gas temperatures for counter-current, co-current IHTR and DCBR. Fig. 3. Predicted temperature profiles for different flow arrangements: (a) cocurrent flow (b) counter-current flow.
and DCBR, the conversion process of dry reforming was simulated with respect to the available catalytic surface area TCASA . Fig. 6 shows the predicted reactor performance results. It is apparent that the performance of the DCBR is superior to that of the IHTR at low available catalytic activities (i.e. TCASA < 180 m2 ). This is attributed to the excellent heat transfer coefficient of the DCBR and the consequently higher thermal efficiency. A critical catalyst activity can be defined, termed TCASA ∗ , at which the performances of the DCBR and IHTR are equal under the specific reaction conditions. Therefore, Fig. 6 illustrates that when TCASA < TCASA ∗ , the DCBR will perform better than the IHTR in terms of reaction conversion and vice versa. The predicted reactant and product gas concentrations are depicted for both reactors in Fig. 7 as a function of catalyst surface area and/or axial distance. The mole fractions of H2 and CO formed in the synthesis gas production are higher in the DCBR than the IHTR when TCASA < TCASA ∗ .
Fig. 6. Comparison of reactor performance of the DCBR and IHTR.
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Table 1 Required reactor volume to evaluate the process intensification potential.
Total pre-heating tube length, m Volumea,b (without molten salt), cm3 Volumea,c (with molten salt), cm3 Process volume intensificationd , %
DCBR
IHTR
– 44 972 34.4%
5.5 1730 3892
The reactor volume is estimated at the critical state TCASA ∗ . The DCBR volume is the catalyst volume considering a 6%Ni-Al2 O3 catalyst (nickel active surface area, BETNi = 1.0 m2 /g) and the IHTR volume is the inner tube volume (Dia. 20 mm) of a concentric heat exchanger. c The molten salt volume is estimated for an alkali carbonate mixture system composed of Na2 CO3 , Li2 CO3 and K2 CO3 in a ratio 1:2:2, and considering the bubbly flow reactor of the DCBR to have a catalyst/salt ratio of 0.1. The IHTR volume is the outer tube volume (Dia. 30 mm) of a concentric heat exchanger. d Process volume intensification indicates the ratio of the eliminated volume using the DCBR in the combination system to the total process IHTR volume (i.e. x = 12 m) without using the DCBR. a
b
the counter-current flow of the IHTR, the better the pre-heating process. This reduces the required tube length because the gas temperature approaches the reaction temperature of 700 ◦ C more rapidly. Consequently, the process volume of the IHTR is decreased. The reduction of the percentage volume intensification is as a result of a reduction in the DCBR size TCASA ∗ , required to compensate the pre-heating process. This implies that high molten salt inlet temperatures discourage the catalytic reactions in the DCBR in comparison with the IHTR, and the significant role of DCBR will encounter in the gas pre-heating process to the reaction temperature. In contrast, in the IHTR, the methane conversion at x = 12 m increases as the inlet molten salt temperature increases.
5.3. Combination system of the DCBR and IHTR
Fig. 7. Predicted reformer gas concentrations of reactants and products (a) in the DCBR as a function of catalyst activity and (b) in the IHTR as a function of catalyst activity and axial distance.
Furthermore, the required reactor volumes to achieve equal performance of both reactors at TCASA ∗ are provided in Table 1. Comparison with the IHTR shows that a significant volume reduction is possible. For the reactors studied, process intensification is achieved when heat transfer to the gas side is dominant. Process volume intensification estimates the reduction percentage of the IHTR process volume when utilizing the DCBR. This is because the DCBR can replace the total pre-heating process of the IHTR. Specifically, in the IHTR, a 3.7-m tube is used to heat the gas temperature to the same reaction temperature as in the DCBR (700 ◦ C). This is in addition to the catalytic bed portion of the tube (1.8 m) required to achieve an equivalent performance at TCASA ∗ . The results suggest that the DCBR is the more suitable reactor for process intensification with an effective total volume reduction of 34%. Simulation results, given in Fig. 8, demonstrate that the inlet molten salt temperature controls the IHTR performance. This is in contrast to the DCBR, the performance of which is controlled by the outlet molten salt temperature, as reported in Ref. [12]. Table 2 shows that the higher the temperature of the molten salt entering
The maximum temperature a solar field can achieve depends on how concentrated the sunlight is. High-quality concentrators, however, are very expensive. Therefore, ideally, we are looking to miniaturize the entire system with minimum impact on the conversion process, increase thermal efficiency, and reduce the cost as far as possible. Despite the problem of the low inlet molten salt temperature that reduces the IHTR performance, the combination system of the DCBR and IHTR is still superior in terms of process intensification with higher thermal efficiency. Furthermore, there have been efforts to circumvent the high temperature constraints by the use of a hydrogen-selective membrane installed in the catalyst bed [21]. The membrane continuously removes hydrogen from the reformer gas stream, thereby increasing the conversion at lower temperature. The integrated system of the DCBR and the countercurrent flow mode of the IHTR are depicted in Fig. 9, where the process data are extracted from Fig. 3(b) and Fig. 6.
Fig. 8. Predicted methane conversion in counter-current IHTR at x = 12 m as a function of inlet molten salt temperature.
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Table 2 Effect of inlet molten salt temperature on reactor performance. Inlet molten salt temperature (◦ C)
Critical TCASA ∗ (m2 )
847 767 743 732
33 99 180 264
Total pre-heating length (m) Pre-heating length
Reactor Bed length
2.65 3.30 3.71 4.00
0.33 0.99 1.80 2.64
Volume intensification (%)
CH4 conversion (% @ TCASA ∗ )
22.7 29.4 34.4 38.6
15 25 32 36
Fig. 9. Schematic representation of the combination system of the DCBR and IHTR.
6. Conclusion
Acknowledgment
Owing to the potentially major role of solar-aided methane dry reforming in the chemical industry, two atmospheric pressure reactors, the DCBR and the IHTR, were compared theoretically. The reactor performance was numerically estimated to evaluate the effect of heat transfer on the methane conversion process, along with a validated catalytic reactions mechanism, and considering the effective variables like flow modes, inlet and outlet temperatures of a thermal fluid and the available catalytic active surface area. The studied reactors’ performances were determined to follow the order: IHTRcounter-current > DCBR > IHTRco-current . Nevertheless, the performance of the DCBR was superior to the IHTR at low available catalytic surface areas (i.e. when TCASA < TCASA ∗ ), which was attributed to the ideally heat transfer coefficient of the DCBR, and the consequently higher thermal efficiency. Therefore, the DCBR can replace the pre-heating process required in the counter-current flow mode of the IHTR. This suggests that the DCBR is a suitable reactor for process volume intensification. The relative efficiency of the IHTR is increased at higher inlet molten salt temperatures; however, operation under such conditions is not cost-effective. A better understanding of the limits of the solar-aided methane reforming process has resulted in a more economical system, namely, the combination system of the DCBR and the counter-current flow mode of the IHTR.
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