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Physicochemical and thermodynamic investigation of the UT-3 hydrogen production cycle: A new technological assessment
International Journal of Hydrogen Energy 31 (2006) 906 – 918 www.elsevier.com/locate/ijhydene
Physicochemical and thermodynamic investigation of the UT-3 hydrogen production cycle: A new technological assessment F. Lemorta,∗ , C. Lafona , R. Dedryvèreb , D. Gonbeaub a CEA-Marcoule, BP 17171, 30207 Bagnols sur Cèze Cedex, France b LCTPCM UMR 5624 Helioparc Pau Pyrénées, 2 avenue Pierre Angot, 64053 Pau Cedex 9, France
Available online 23 September 2005
Abstract The previous study of the UT-3 thermochemical cycle was based on laboratory-scale tests to assess the feasibility of each step. In order to enclose the reactants in the experimental vessel the study was carried out on pellets that had to be prepared according to a specific protocol. It would appear difficult to design and operate a cost-effective industrial process using this technique, and it could be worthwhile to consider another approach. This paper identifies problem areas in the cycle as expressed below: (A) (B) (C) (D)
CaO + Br 2 → CaBr 2 + 21 O2 , CaBr 2 + H2 O → CaO + 2HBr, Fe3 O4 + 8HBr → 3FeBr 2 + 4H2 O + Br 2 , 3FeBr 2 + 4H2 O → Fe3 O4 + 6HBr + H2 .
A thermodynamic approach predicts that the first and third reactions will occur easily, but a physicochemical approach predicts some difficulties due to sintering of the solid reactants. With regard to the second and fourth reactions, thermodynamics would dictate operation at low pressure and high temperature, in contradiction with the volatility of the bromides. One solution could be to ensure superstoichiometric conditions, which assumes a large quantity of gas. Based on these considerations, and on the fact that all the steps of this thermodynamic cycle are gas–solid reactions, a fluidized bed process is proposed with continuous feed of the gaseous reactants. The efficiency of the process is discussed. 䉷 2005 International Association for Hydrogen Energy. Published by Elsevier Ltd. All rights reserved. Keywords: Hydrogen production; Thermodynamical cycle; UT-3 cycle; Technology
1. Description of the UT-3 cycle The UT-3 thermochemical cycle has been largely described in previous papers regarding the concept of the technology (MASCOT) that can be implemented in order to produce the hydrogen from the cycle [1,2] but also regarding the basic data that have to be considered for its assessment ∗ Corresponding author.
E-mail address: fl[email protected] (F. Lemort).
[3,4]. The first results obtained lead to improve the reactivity of the whole reactants [5,6] and to consider other primary heat source than nuclear power in order to offset the low efficiency of an industrial loop [7]. A recent study points out that with the concept of the MASCOT technology it could be very difficult to increase the efficiency of the loop [8]. The UT-3 cycle is based on two pairs of chemical reactions. The first two ensure the formation of hydrobromic acid releasing oxygen, and the other two ensure the reduction of water by a bromide.
0360-3199/$30.00 䉷 2005 International Association for Hydrogen Energy. Published by Elsevier Ltd. All rights reserved. doi:10.1016/j.ijhydene.2005.07.011
F. Lemort et al. / International Journal of Hydrogen Energy 31 (2006) 906 – 918
can be a limiting factor since the following must be taken into account:
Nomenclature reaction temperature, ◦ C reaction pressure, bar time, s free enthalpy of reaction, kJ mol−1 entropy of reaction, J mol−1 K −1 heat of reaction, kJ mol−1 variation of gaseous species, mol degree of reaction progress flow of gaseous species, mol s−1 quantity of gaseous reactants, mol quantity of solids reactants, mol
T P t G◦ S ◦ H ◦ vg qg Ng Ns (A) (B) (C) (D)
907
CaO + Br 2 → CaBr 2 + 21 O2 , CaBr 2 + H2 O → CaO + 2HBr, Fe3 O4 + 8HBr → 3FeBr 2 + 4H2 O + Br 2 , 3FeBr 2 + 4H2 O → Fe3 O4 + 6HBr + H2 .
The choice of temperature and pressure conditions depends on the physical and chemical properties of the reactants and the thermodynamics of the reactions. The physical and chemical data concern mainly the phase transitions and the possible formation of azeotropes. The available data on all the products of the above-described system are indicated in Table 1. The thermodynamic data [9] concerning all the reactions are indicated in Table 2, where vg represents the variation in the number of moles of the gaseous species. The temperatures were selected to obtain realistic operating conditions. 2. Feasibility analysis A relatively simple cycle feasibility analysis was performed using published data. Each of the cycle reactions was considered. Reaction A The data suggest a quantitative, thermodynamically unlimited reaction. As this is a gas–solid reaction, the kinetics
• the accumulation of reaction products on the surface of the grain could lead to a protective layer, • diffusion phenomena (diffusion of reactants through a film of gas and solid), • adsorption and desorption phenomena, • decreasing of the active surface. It would be preferable to use a fluidized bed to ensure thorough agitation of the system and thus enhance the diffusion phenomena. Moreover, the possible appearance of a specific Ca3 OBr 4 compound cannot be discounted; this compound appears in a chloride system in which the phase diagram [10] exhibits a eutectic melting point. Reaction B The data in Table 2 reveal a non-quantitative reaction favored by low pressures and high temperatures. The degree of reaction progress can be related as follows to the CaBr 2 consumption, the total pressure system P , and the equilibrium constant K: b b K K = ≈ , (1) a K + 4P 2a P where a represents the quantity of CaBr 2 and b the quantity of H2 O. The graphic representation of =f (b/a) as shown in Fig. 1 shows that satisfactory reaction progress can be expected at low pressure values. The figure shows that operation at ambient pressure is relatively unfavorable, even with excess water. Low pressure operation appears to be more advantageous, provided the volatilization of species is limited, as suggested by the data in Table 1. As for the previous reaction, allowance must be made for the possible formation of specific fully immiscible compounds generating eutectic mixtures with melting points that can be below the operating temperature. Reaction C Like reaction A, reaction C is quantitative and thus not limited thermodynamically. The remarks concerning the
Table 1 Reactant properties [9,10] Species
Phase/Transition
Specified compound
Cp
CaO Br 2 CaBr 2 H2 O
S S → L: 59.6 ◦ C S → L: 742 ◦ C L → G: 100 ◦ C
CaO–CaBr 2 eutectic (Ca3 OCl4 )
≈ 51.3 (627 ◦ C) ≈ 37.4 (427 ◦ C) 59.2 + 0.0819T − 0.0001T 2 + 6 × 10−8 T 3 30.2 + 0.0109T
HBr
G
Fe3 O4 FeBr 2 H2
S S → L: 691 ◦ C G
CaO–CaBr 2 eutectic (Ca3 OCl4 ) H2 O–HBr azeotrope T = 126 ◦ C 47% HBr H2 O–HBr azeotrope T = 126 ◦ C 47% HBr Fe3 O4 –FeBr 2 eutectic? Fe3 O4 –FeBr 2 eutectic?
Vapor pressure (bar)
4 × 10−7 (727 ◦ C)
29.505 − 0.0032T + 6 × 10−6 T 2 86.4 + 0.2086T 73.6 + 0.0223T
5 × 10−3 (627 ◦ C)
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F. Lemort et al. / International Journal of Hydrogen Energy 31 (2006) 906 – 918
Table 2 Thermodynamic data for the UT-3 cycle reactions [9] Reaction
T (◦ C)
G◦ (kJ mol−1 )
S ◦ (J mol−1 K−1 )
H ◦ (kJ mol−1 )
vg (mol)
(A) (B) (C) (D)
527 727 227 627
−40.9 104.0 −119.9 117.5
−41.8 107.4 −303.8 286.3
−74.3 211.4 −271.8 386.7
−0.5 +1.0 −3.0 +3.0
1.0 0.9 0.8 0.7
α
0.6 0.5
P = 1 bar
0.4
P = 5 mbar
0.3
P = 1 mbar
0.2
P = 0.1 mbar
0.1 0.0
0
10
20
30 nH2 O/n CaBr2
40
50
60
Fig. 1. Progress of reaction B versus the H2 O/CaBr 2 ratio at different total pressures.
64a 7 7 3(b + a)3 (b − 4/3a)
P 3 − K = 0, 4
1.2 1 0.8 α
kinetics of reaction A are also applicable here. The use of a fluidized bed at a temperature near 200 ◦ C could be interesting from this standpoint. The reaction is sufficiently exothermic that the use of a heat exchanger is recommended. Reaction D As for reaction B, this reaction is not quantitative and is favored by low pressures and high temperatures. It should be noted, however, that 627 ◦ C is not far from the melting temperature of FeBr 2 . Problems can be expected from sintering of the system, hence the unquestionable advantages of a stirred system to prevent or limit plugging. As with reaction B, the coefficient of reaction progress (chemical conversion) can be related to the equilibrium constant K and the total pressure P :
0.6 0.4 0.2 0
0
5
10
15
-Ln(P(bar))) Fig. 2. Progress of reaction D versus the total pressure.
relation (2) becomes (2)
where a represents the quantity of FeBr 2 and b the quantity of H2 O. Eq. (2) demonstrates that at a sufficiently low pressure for given material quantities, the reaction progress tends toward 1. For example, assuming stoichiometric proportions,
1399687 3(b + 4)3 (b − 4)4
P 3 − K = 0.
(3)
The graphic representation of =f (− ln P) (ln P has been chosen in order to facilitate the reading of the graphic) as shown in Fig. 2 indicates that satisfactory reaction progress can be expected at low pressure values. For example, a mean
F. Lemort et al. / International Journal of Hydrogen Energy 31 (2006) 906 – 918
pressure of 5 mbar ensures a progress of 0.6. As shown above for reaction B, excessively low pressures can be avoided by ensuring excess water, thus limiting the volatility of FeBr 2 which could otherwise be entirely volatilized.
909
decided to perform the calculation for stoichiometric reaction conditions—i.e. for a system comprising one mole of CaBr 2 and one mole of water—at temperatures ranging from 300 to 1000 K. The result shows clearly a total chemical conversion in the temperature range. The simulation shows that this bromination reaction should not raise any difficulties. The fact that the molar volume of CaBr 2 is about 3.6 times greater than the molar volume of CaO [3] indicates 70% expansion of the system that must be taken into account. Reaction B The initial evaluation showed that this reaction must occur at low pressure with excess water. We simulated the evolution of a system initially comprising one mole of CaBr 2 for 40 moles of water in a vessel pressurized to 5 mbar, with the result shown in Fig. 3. Unlike the previous results, hydrolysis of CaBr 2 leaves very little leeway for the process. Ideally, the process temperature would be 1100 K, which corresponds to the best chemical conversion. However, the melting point of CaBr 2 means the process temperature must remain below 1000 K, i.e. with a reaction progress of about 0.5 under the conditions described above. Reaction C The results (Table 2) show that the reaction is feasible. The simulation performed under standard conditions with stoichiometric quantities reveals that many species can appear depending on the working temperature. Fig. 4 shows the system evolution according to the temperature.
3. Numerical simulation The first step toward assessing the cycle feasibility is to consider the system variations versus temperature for specified material quantities. This can be done on the basis of the conclusions of the preceding section (operating pressure, superstoichiometry, etc.) by evaluating systems at equilibrium initially containing all the reactive species (solid reactants Ns and gaseous reactants Ng ). During a time interval t the actual process will be subject to a flow of gaseous species qg such that T Ng = qg dt. (4) 0
The calculation is free enthalpy minimization of a system initially defined (chemical composition and pressure) with respect to all the possible compositions compatible with the conservation of mass. The Factsage thermochemical database [9] and its computation software were used for this purpose. Reaction A The assessment in the preceding section showed that this reaction is theoretically highly feasible. We therefore 2.0 CaBr2+ 40 H2O P = 5mbar
HBr
Quantity (mol)
1.6
1.2
CaBr2 (H2O)6 CaBr2(s)
CaO(s)
0.8
0.4 Br (g)
0 300
400
500
600
700
800
900
CaBr2(l)
H2
1000
1100
T (K) Fig. 3. CaBr 2 + 40H2 O system response versus temperature at 5 mbar.
1200
910
F. Lemort et al. / International Journal of Hydrogen Energy 31 (2006) 906 – 918
4.8 Fe3O4+ 8HBr
4.4 4.0
HBr
P=1bar
H2O(g)
H2O(l)
Quantity (mol)
3.6 3.2
FeBr2 (s1)
2.8 2.4 2.0
FeBr3 (s)
FeBr2 (s2)
1.6 1.2
FeBr2( g) FeBr2 (liq)
Br2 Fe2O3 (s)
0.8
Fe2O4 (s)
0.4 0 300
400
500
600
700
800 T (K)
900
1000
1100
1200
Fig. 4. Fe3 O4 + 8HBr system response versus temperature at 1 bar.
6.0 5.6
3FeBr2 + 40H2O
4.2
P=1bar
HBr
4.8 4.4 Quantity (mol)
4.0 3.6 3.2
FeBr2 (s1)
FeBr2 (s2)
2.8 2.4 2.0 1.6 H2
1.2 0.8
FeBr2( g)
0.4 0 300
Fe3O4 (s)
400
500
600
700
800
900
1000
1100
1200
T (K) Fig. 5. 3FeBr 2 + 40H2 O system response versus temperature at 1 bar.
The optimum equilibrium occurs at 400 K. At higher temperatures, oxidizing hydrolysis of FeBr 2 begins to form Fe2 O3 and HBr. This reaction is liable to limit the cycle efficiency but does not prevent operation since Fe2 O3 is easily reduced to FeBr 2 by HBr. Nevertheless, it is advisable to maintain the temperature as close as possible to 400 K.
Reaction D Eq. (3) shows that the most advantageous conditions are with excess water (rather than under low pressure, which would be less efficient and more energy-intensive). Fig. 5 shows the calculated result. With 10-fold excess water, the maximum reaction progress is obtained at 900 K. However,
F. Lemort et al. / International Journal of Hydrogen Energy 31 (2006) 906 – 918
911
N2, H2
N2 Water
Tubular furnace
CaBr2 or FeBr2
Flask of gard
Water (HBr trap)
Fig. 6. Experimental setup for evaluating hydrolysis steps.
as observed earlier, the vapor pressure of FeBr 2 leads to appreciable volatilization of the product, making it preferable not to exceed 800 K for this step.
4. Experimental assessment
recovered powder. These analyses give the composition of the surface of the grains and then the analysis of the products of reaction. The chemical conversion is given by the analysis of the water contained in the HBr trap. The XPS analyzer was connected to the glove box, allowing the analyses to be carried out only a few minutes after the end of the test to avoid degradation of the reaction products.
4.1. Material and method 4.2. Experimental results A few experimental tests were performed to check some of the points mentioned above. The tests concerned the most limiting steps, i.e. the hydrolysis of CaBr 2 and FeBr 2 , and mainly involved estimating the reaction progress for these two steps. A relatively simple experimental setup was used, in which CaBr 2 or FeBr 2 powder was swept by steam carried in a nitrogen stream. The experimental device is shown in Fig. 6. The tubular furnace was made of silica, and the compounds were placed in an alumina boat (a silica boat could not be used because of the formation of silicates). The quantity of HBr dissolved in the trap gives the rate of chemical conversion. The products were kept and handled in a glove box under anhydrous nitrogen atmosphere. The test procedure was as follows: • 2 g ± 0.01 of bromide (CaBr 2 or FeBr 2 ) were placed in the sample boat. • The boat was inserted in the silica tube inside the tubular furnace in a dry nitrogen stream. • The nitrogen flow was maintained for 1 h at room temperature to eliminate all the oxygen. • The furnace was heated to the test temperature (725 ◦ C for CaBr 2 , 600 ◦ C for FeBr 2 ). • The nitrogen stream was then charged with steam by sparging and was maintained at the test temperature for hours. • The test was terminated, and the sample transferred to the glove box. Qualitative analysis was then performed by X-ray photoelectron spectroscopy (XPS) and by X-ray diffraction on the
(1) CaBr 2 → CaO The main conclusions after analyzing the residues were the following: • Major sintering of the solid mass left a compact compound at the bottom of the crucible. • The reaction progress reached about 11%. (2) FeBr 2 → Fe3 O4 The main conclusions after analyzing the residues were the following: • The product was volatile with recondensation of reaction products at various points in the sample boat. • The reaction progress reached 87%. • The major reaction product was Fe2 O3 . 4.3. Discussion The most important phenomenon observed during these two experiments was significant sintering in both cases. Sintering modifies the reactive surface finish and probably accounts for the decreased reactivity observed by other authors [3,6] during kinetic tests on CaO pellets incorporated in a CaTiO3 matrix. Aihara et al. [3] demonstrated that the reactivity diminishes as the temperature rises; this is likely attributable to reactant melting problems. The second feature observed was the partial progress of the hydrolysis reactions, the least favorable being only 11% for CaBr 2 . This was much lower than the value measured by Aihara et al. [3] in an experimental setup in which the
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CaBr2 + H2O
3FeBr2 + 4H2O
CaO + 2HBr
Fe3O4 + 6HBr + H2 HBr
Ca reactor unit
Fe reactor unit
O2
H2
CaO + Br2
Fe3O4 + 8HBr
CaBr2 + 1/2O2
3FeBr2 + 4H2O + Br2
Fig. 7. UT-3 cycle flowsheet for the Mascot pilot facility.
gas passed through the sample; the difference may be attributable to the fact that the gas stream flowed over the sample in our experiment. The limited reaction progress is in any event related to passivation of the reaction interface, where the CaO or Fe2 O3 layer becomes thick enough to limit diffusion of the reactants or gaseous reaction products. The satisfactory progress (87%) observed for FeBr 2 is attributable to partial volatilization of FeBr 2 resulting in a gas–gas reaction with favorable kinetics. The third phenomenon observed was the preferential formation of hematite Fe2 O3 . This result is rather unusual for this kind of reaction since the hydrogen potential is high enough to keep the iron under the magnetite form. This result that is not in accordance with previous results obtained by other authors nor with the results of calculation showed in Fig. 5 could be explained by the reactive system which is always far from the equilibrium and by the experimental method that is different from the other and may favor the formation of hematite. Furthermore, the calculation of equilibrium constant at 900 K for the reaction modified reaction D giving the hematite Fe2 O3 (and described below) shows a value 38 times lower than for the production of magnetite. This difference is not very significant and both the reactions may be competitive. In this case, the experimental condition could favor one or the other way. (E) 2FeBr 2 + 3H2 O → Fe2 O3 + 4HBr + H2 . This is a very significant observation that changes the frame of the cycle since the amount of released hydrogen versus the amount of iron reactant is higher. If this observation is true, the efficiency of the cycle could be increased. This result that could be connected with the experimental condition has to be checked before going further on this way. The amount of hydrogen evolved in this reaction has to be very carefully analyzed.
5. Process considerations The process implementing the UT-3 cycle is based on the use of four series-connected reversible stationary beds in which the reactants are incorporated in a titanate matrix for calcium and a silica matrix for iron [2]. Fig. 7 shows how the reactors are arranged [1]. The Mascot process developed in Japan [1] has several advantages: • No transport of solids. • Reactant surface condition conserved by inclusion in matrices [2]. • Simple process design obtained by flow reversal. Nevertheless, the process has the double disadvantage of requiring a reagent preparation unit and increasing the quantity of material in the process by adding an inert phase representing 33–50% of the reactive load [3]. A material balance was determined to predict the operation of a unit capable of generating 30 000 Nm3 h−1 (i.e. about 90 MWth). Table 3 indicates the minimum masses of reactants (intended to react within 1 h) and inert phases required in the reactors, together with the gas mass and volume requirements. The worst-case situation with a 2:1 mass ratio between the reactive and inert materials was considered [3]. In the above mass balance, inert material (525 t) accounts for 25 wt% of the total solid mass. Eliminating the use of reactant matrices would appreciably diminish the total product mass while also reducing the number of individual operations necessary for sustained operation of a production unit (as it would be necessary to remove the reactant matrices periodically from the reactors for regeneration). In fact, although theoretically reactants are incorporated in matrices to ensure durability, they will not last indefinitely because
F. Lemort et al. / International Journal of Hydrogen Energy 31 (2006) 906 – 918 Table 3 Hourly reactive material mass necessary 30 000 Nm3 h−1 of H2 with the Mascot process Material
Mass (metric tons)
CaO CaBr 2 CaTiO3 Fe3 O4 FeBr 2 SiO2 Br 2 HBr H2 O Total
75 268 364 310 866 161 214 867 96 3221
to
produce
Volume (Nm3 )
30 240 120 390
– – – – – – 000 000 000 000
of cyclic expansion and contraction of the reactive agglomerates [3]. The resulting mechanical stresses will require periodic regeneration of the titanate (calcium unit) or silicate (iron unit) matrices. Moreover, on the calcium side of the process, decomposition of the CaTiO3 is likely by HBr diffusion through the layer that has already reacted, according to the following balance: (F) CaTiO3 + 2HBr → CaBr 2 + TiO2 + H2 O = 560.
K(800 K)
Even if the matrix is produced with a very small specific surface area, bromination of the titanate will result in decohesion of the “inert” matrix phase and the matrix will lose its mechanical properties. In order to limit the mass of the constituent materials and the matrix preparation operations which require several steps if the alkoxide process is used [3], it could be advantageous to use fluidized bed technology in which the calcium unit is coupled in a single reactor like the iron unit. The reactor design could take into consideration the properties of the reactants and reaction products.
913
tion is ensured first by the gas flow that is enough to expand the solid reactive bed in the upflow portion and then by the gravity in the downflow zone into which the bromination of calcium oxide will occur. The volume difference between the two zones of the asymmetric torus will ensure a ratio of about 50:1 between the solid particle residence times, due not only to the gas velocity but also to expansion of the fluidized bed [3]. The gas flow direction is controlled by two blowers ensuring not only oxygen extraction at the bottom of the unit but also extraction of the HBr/H2 O mixture at the top; the two reactor zones are separated by non-return valves that allow solid particles to flow in only one direction. As the purpose of the calcium reaction cycle is to produce HBr, it could be highly simplified by focusing solar radiation on a reactor dedicated to the simple reaction: (G) H2 O + Br 2 → 2HBr + 21 O2 . In this case the UV radiation is used in order to increase the kinetics and the chemical conversion of this reaction by dissociating directly Br 2 homolytically or heterolytically [11,12] and to lower the amount of the heat that is required. 5.2. Iron reactor Tests showed that the FeBr 2 hydrolysis reaction progress is much greater for any given reaction time; this may be attributable to the slower CaBr 2 hydrolysis kinetics. These results were confirmed by Yoshida et al. [11], who reported that the hydrolysis rate is about 20 times higher for FeBr 2 than for CaBr 2 . The technology used must take this result into account, but could also allow for the differential properties of FeBr 2 and Fe3 O4 . The most important of these are the molar volumes, which vary in the same way as those of the calcium reaction zone, and the magnetic properties of the two compounds: Fe3 O4 is ferromagnetic with a susceptibility of about 106 , whereas FeBr 2 is paramagnetic with a susceptibility of about 10−2 [14]. These properties could thus be used to separate the two reaction products within a single reactor [15].
5.1. Calcium reactor 5.3. Process description Two reactions occur in this unit; the first (A) is quantitative and relatively rapid (assuming a suitable reaction surface condition), while the second (B) is much less favorable from a thermodynamic and kinetic standpoint [3,5]. It must also allow for a substantial volume variation of the process compounds: the molar volume of CaBr 2 is about 3.6 times greater than that of CaO [3]. Finally, Br 2 dissociates either homolytically under the effect of near-UV visible electromagnetic radiation [11] or heterolytically when higher-energy wavelengths are used [12]. A vertical nonsymmetric toroidal reactor [13] could be used in which hydrolysis could be carried out in the upflow portion, so the grain size reduction arising from striction would ensure that nearly all the particles in the downflow zone would have already finished reacting. The solid circula-
Fig. 8 illustrates a flowsheet for a facility capable of producing 30 000 Nm3 h−1 of hydrogen. The given pressures have been estimated (in bars) in the reactors, and upstream and downstream from each unit. It is assumed here that a membrane causes a pressure drop of about 2000 Pa. The calcium process zone consists mainly of a circulating fluidized bed reactor with asymmetric toroidal geometry in which CaBr 2 is hydrolyzed in the riser and CaO is brominated in the downer. Based on a maximum reaction progress of 70% [3], the reactor contains 140 t of calcium distributed between CaBr 2 and CaO. The reactor can be designed with reference to a few examples implementing circulating fluidized beds [14] by assuming the limiting step is the hydrolysis of CaBr 2 . The proposed flowsheet recommends low
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Fig. 8. Simplified diagram of a process implementing the UT-3 cycle.
pressure to facilitate the hydrolysis step. If this approach is too energy-intensive, this step could be enhanced by UV radiation [11] with both reactions operating at atmospheric pressure. The design basis criterion would be the reactant residence time in the reactor. According to the results reported by Yoshida et al. [4], the residence time for CaBr 2 should be about 1 h in a stationary bed configuration. The use of a fluidized bed reactor to stir the system could shorten this time by half. The residence time for CaO bromination should be 10 min. As indicated in Fig. 1, the temperatures can be identical at 650 ◦ C. Reactor design proposals can be based on the data consolidated in Table 4. The gas flow rate will fluidize the beds (expansion in the first case, contraction in the second), depending in fact on the particle size used in the reactor. Although the working temperatures were lowered to limit degradation of the reactant particles, provision must be made for calibrating the particles. This could be done using a pan mill calibrated for the required particle size. Installing two mills in the downflow zone of the torus would not only recalibrate the particles but also separate the two zones and thus prevent any mixing of their atmospheres.
The same principle could be used for the reactor in the second zone except that applying a permanent field could facilitate magnetite expansion in the upper reactor zone and thus its transfer to the downflow leg. The temperatures must be more closely controlled than in the calcium reactor zone to avoid the formation of undesirable compounds as shown in Figs. 3 and 4 (the superstoichiometry actually taken into account will be much less than 10 to avoid lowering the overall cycle efficiency). The hydrolysis temperature was set at 600 ◦ C to limit the quantity of FeBr 2 entering the gas phase (the possible volatilized fraction will be recovered by condensation at high temperature (500 ◦ C) at the level of the off-gas filter). The hydrobromination temperature is limited to 130 ◦ C to prevent reoxidation of FeBr 2 by Br 2 /H2 O. The proposed geometric characteristics are indicated in Table 5, once again allowing for a 70% limit on the reaction progress. The off-gas treatment proposed in Fig. 7 assumes the calcium zone includes a membrane separating the HBr/H2 O mixture; the water purity must be as high as possible to ensure hydrolysis of the CaBr 2 . Considering that FeBr 2 hydrolysis is less sensitive to the presence of HBr, it is proposed that HBr be separated from its azeotropic mixture by fractionated condensation. If this is unsuitable, a membrane
F. Lemort et al. / International Journal of Hydrogen Energy 31 (2006) 906 – 918
915
Table 4 Dimensional characteristics of a torus reactor in the calcium process zone Reaction zone
Upflow
Downflow
Reaction Residence time Reactant mass Reactant volume (compact) Expansion factor Reactor volume Temperature Type of bed Gas velocity Zone diameter Zone height
CaBr 2 → CaO 30 min 134 t 40 m3 4 160 m3 650 ◦ C Fluidized (CaO particles expand) 2 m s−1 4.1 m 12 m
CaO → CaBr 2 10 min 16 t 5 m3 – 217 m3 650 ◦ C Fluidized (CaBr 2 particles settle) 1.5 m s−1 4.8 m 12 m
Table 5 Dimensional characteristics of a torus reactor in the iron process zone Reaction zone
Upflow
Downflow
Reaction Residence time Reactant mass Reactant volume (compact) Expansion factor Reactor volume Temperature Type of bed Gas velocity Zone diameter Zone height
FeBr 2 → Fe3 O4 30 min 563 t 122 m3 4 488 m3 600 ◦ C Fluidized (Fe3 O4 particles expand or migrate magnetically) 2.5 m s−1 10 m 6m
Fe3 O4 → FeBr 2 10 min 67 t 13 m3 – 292 m3 130 ◦ C Fluidized (FeBr 2 particles settle) 2 m s−1 8m 6m
system could be used. Like the material quantities indicated in the dimensional tables, the gas flows are based on the production of 30 000 m3 h−1 of hydrogen. The water masses X indicated in the diagram correspond to the possible superstoichiometry necessary to ensure satisfactory progress of the hydrolysis reactions. The bromine that does not react during the CaO bromination step is separated during cooling of the gas prior to storage of the oxygen produced. In the iron zone, separation can be obtained by condensation or possibly using a specific membrane at high temperature. Contact with the heat source (not shown in the diagram) will be made in the torus upflow legs, for example by fluid flowing in a jacket. The values given in Tables 4 and 5 show clearly that the facilities can be huge for the production of 30 000 Nm3 h−1 of hydrogen. The process proposed in Fig. 8 can be divided into smaller unit in series in order to facilitate the handling or maybe the running of the process.
6. Efficiency The efficiency of the UT-3 using this technology has not yet been assessed in detail, but the issues have been exam-
ined on the basis of work by Sakurai et al. [7] and by Teo et al. [8] concerning the flowsheet proposed in Fig. 8. The efficiency may be simply defined as follows: =
Q(H2 ) + Qe , i Qi
(5)
recombination, Qe where Q(H2 ) is the heat of hydrogen is the excess thermal energy and Qi is the total thermal energy supplied to the process, including the energy recombined as electric power with 35% efficiency to operate the process equipment. As the excess energy Qe is difficult to estimate accurately, we chose to assign a highly unfavorable zero value to this parameter. The calculation was performed here for the excess water X1 and X2 in the calcium and iron reactors, respectively. These values affect the quantity of material circulating between the reactors, as well as the heat flux necessary to heat the gas streams to the working temperature. Contrary to the adiabatic approach by Sakurai et al. [7], we preferred to use isothermal reactors to avoid raising the reactant temperatures and thereby encountering difficulties related to sintering of the materials and the formation of undesirable reaction products such as Fe2 O3 in the downflow leg of the iron reactor, which would require separation of the HBr produced. The reaction isothermicity
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Table 6 Estimated UT-3 cycle power requirement Item
T (◦ C)
Inlet pressure
Outlet pressure
Power (W)
A B
650 650
0.1 0.18
18 0.2
C
150
P6
1
D
600
P2
1
E F
100 650
P5 0.13
498 123 394(25 + X1 ) 20 000 + 294X2 703 364 ln 24 000 + 22.05X 2 (100 + 1.47X2 )2 7258(100 + 1.47X2 ) ln 2.16X22 + 70.56X2 + 4800 77 528 ln[25(100 + 1.47X2 )] 114 693
25 0.2
Table 7 Heat exchanger energy balance Heat exchanger
Inlet temperature (◦ C)
Outlet temperature (◦ C)
Power (W)
Recoverable fraction (%)
E1 E2 E3 E4 E5 E6
150 650 130 100 100 650
650 150 600 150 126 25
462 500 737 500 1598 000 + 23 491X2 221 250 1516 800 + 21 690X2 268 750
– 50 – – 40 30
1225 655 + 14 815X2
Total
35 30 Efficiency (%)
can be ensured by the use of a jacketed reactor to supply heat to the upflow legs (endothermic reactions) and to remove heat from the downflow legs (exothermic reactions). The recovered heat can be used in part to vaporize 24 t h−1 of water. Although the flowsheet shown in Fig. 8 corresponds to the production of 30 000 Nm3 h−1 of hydrogen, the efficiency was estimated for a 25 mol s−1 hydrogen production cycle as in the article by Sakurai et al. [7]. In this case, the power available from the hydrogen produced Q(H2 ) is 6.05 MW. The heat input required to ensure reaction isothermicity is 7.7 MW. The power requirements for the blowers were estimated in terms of the excess water X1 and X2 using a set of isothermal transformations. Table 6 indicates the power exchanged in the heat exchangers Ei . The power necessary to operate the calibrators must be added to the balance given in Table 7. It can be roughly estimated at about 200 kW per reactor. The efficiency is calculated using relation (4). It is interesting to perform the calculation with the data established for the process flowsheet (Fig. 8), but also to consider the use of two membranes rather than fractionated condensation to separate the H2 /HBr/H2 O gases. Using membranes would avoid the need for a cooling/heating cycle to separate these species. The calculation assumes the process equipment (except for the heat exchangers) is powered by electricity obtained with 35% efficiency from the waste heat. Considering the pres-
25 Membrane 20 Condensation 15 10
0
5 10 15 20 Stoichiometry 4H2O/3FeBr2
25
Fig. 9. Cycle efficiency versus water stoichiometry in the iron reactor.
sure in the calcium reactor is 0.2 bar and not 5 mbar, the stoichiometric ratio must be greater than 40 as indicated by the results shown in Fig. 3. A ratio of 250 yields similar results; we adopted this value since the excess water inventory in the second reactor determines the overall efficiency. Fig. 9 shows that the water quantities have a major effect on the overall cycle efficiency in the case of fractionated condensation. The efficiency is about 15% with a super-
F. Lemort et al. / International Journal of Hydrogen Energy 31 (2006) 906 – 918
stoichiometry of 10 (to obtain results comparable to those shown in Fig. 5). The efficiency rises to 22.6% with the use of membrane technology. This is consistent with the conclusions reported by Teo et al. [8], and highlights the need to develop efficient membranes with a low pressure drop to allow thermochemical cycles to operate under optimum conditions. In the present case, the efficiency is not diminished by flow reversibility and the resulting temperature changes in the solid masses. In addition, fluidized beds appreciably diminish the pressure drop compared with a stationary bed and thus also limit the energy requirements. Finally, eliminating the regeneration steps for reactants embedded in matrices certainly allows appreciable energy savings, although the amount has not been estimated to date. We chose not to perform calculation regarding a UT3 cycle involving the production of hematite Fe2 O3 because it is first necessary to check the experimental results.
7. Conclusion A physicochemical evaluation of the UT-3 cycle based on current trends in the literature reveals a number of difficulties. First of all, the systems implemented can lead to the formation of mixtures that converge toward low melting-point eutectic compositions or azeotropes that complicate process operation. The occasionally unfavorable thermochemistry highlights the need for modifying some operating parameters, including the reactant stoichiometry or the working pressure. Recirculation of gaseous reactants can be advantageous in this regard. The present study together with a few experimental verifications has identified a new process flowsheet implementing two asymmetric torus reactors with fluidized beds of solid reactants in each leg; the solid particles transit from one leg to the next via a recalibration system according to the degree of reaction progress. This technology has several advantages, the first of which is that it avoids a probably energy-intensive reactant preparation step in a stationary bed by incorporating them in titanate or silicate matrices for use with calcium or iron. The second advantage is that it improves the reaction kinetics while ensuring continuous operation of both reactors, which in this case do not have to be reversible and therefore subject to major efficiency losses. The calculated process efficiency is about 15% if the gas separation is partly ensured by fractionated condensation; the use of dedicated membranes increases the mean efficiency to 22.5% and highlights the importance of developing this type of separation technique. This assessment shows the high necessity to make significant advances in improving the efficiency of membrane technology to avoid the energy loss in the separation using condensation. Advances have also to be made in the field of the process technology in order to perform the assessment of the proposed technology.
917
Acknowledgements The authors are grateful to the referee whose remarks allowed improvements in the presentation of their work.
Appendix Calculation of for the reactions B and D: Reaction B CaBr 2 +H2 O t =0 a b t= ∝ a(1 − ) b − a K=
=
→ CaO+ 0 a
2HBr, 0 2a,
2 PHBr 4a 2 2 = 2 P, PH2 O b − a 2 2
b a
K . K + 4P
Reaction D 3FeBr 2 + 4H2 O → Fe3 O4 + 6HBr + H2 , t =0 a b 0 0 0 t= ∝ a(1 − ) b − 4/3a 1/3a 2a 1/3a, P 6 PH2 64a 6 6 a (a + b)4 3 = P , K = HBr 4 3(a + b)7 (b − 4/3)4 PH O 2
=
64a 7 7 3(b + a)3 (b − 4/3)4
P 3.
References [1] Yoshioka H, Nakayama T, Kameyama H, Yoshida K. Operation of a bench-scale plant for hydrogen production by the UT-3 cycle. Hydrogen energy progress V: proceedings of the fifth World hydrogen energy conference. Toronto, Canada, 15–20 July 1984. p. 413–20. [2] Kameyama H, Tomino Y, Sato T, Amir R, Orihara A, Aihara M, Yoshida K. Process simulation of mascot plant using the UT-3 thermochemical cycle for hydrogen production. Int J Hydrogen Energy 1988;14(5):323–30. [3] Aihara M, Umida H, Tsutsumi A, Yoshida K. Kinetic study of UT-3 thermochemical hydrogen production process. Int J Hydrogen Energy 1990;15(1):7–11. [4] Yoshida K, Kameyama H, Aochi T, Nobue M, Aihara M, Amir R, Kondo H, Sato T, Tadokoro Y, Yamaguchi T, Sakai N. A simulation study of the UT-3 thermochemical hydrogen production process. Int J Hydrogen Energy 1990;15(3): 171–8. [5] Sakurai M, Tsutsumi A, Yoshida K. Improvement of Ca-pellet reactivity in the UT-3 thermochemical hydrogen production cycle. Int J Hydrogen Energy 1995;20(4):297–301.
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[6] Sato T, Sakurai M, Matsumura Y, Tsutsumi A, Yoshida K. Kinetics and mechanism of hydrolysis of a Ca pellet in the UT3 thermochemical cycle for hydrogen production. Hydrogen energy progress XII: proceedings of the 12th World hydrogen energy conference, Buenos Aires, Argentina, 21–26 June 1998. p. 561-69. [7] Sakurai M, Bilgen E, Tsutsumi A, Yoshida K. Solar UT-3 thermochemical cycle for hydrogen production. Solar Energy 1996;57(1):51–8. [8] Teo ED, Brandon NP, Vos E, Kramer GJ. A critical pathway energy efficiency analysis of the thermochemical UT-3 cycle. Int J Hydrogen Energy 2005;30(5):559–64. [9] FACTSAGE database, Thermofact, Version 5.1-2002. [10] Wenz DA, Johnson I, Wolson RD. J Chem Eng Data 1969;14(2):252.
[11] Lindeman TG, Wiesenfelda JR. Photodissociation of Br 2 in the visible continuum. J Chem Phys 1979;70(6): 2882–8. [12] Maazouz M, Maazouz PL, Jacobs DC. Anomalous chargetransfer behavior in the scattering of hyperthermal Br+(3P2) on Pt(111). J Chem Phys 2002;117(24):10917–20. [13] Benkhelifa H, Legrand J, Legentilhomme P, Montillet A. Study of the hydrodynamic behavior of the batch and continuous torus reactor un laminar and turbulent flow regimes by means of tracer methods. Chem Eng Sci 2000;55(10):1871–82. [14] David R Lide, editor. CRC handbook of chemistry and physics. 79th ed. Boca Raton, FL: CRC Press; 1998. [15] Song MC, Lee KJ. A study on a magnetic separation of radioactive corrosion products from NPP using permanent magnets. Nuclear Eng and Design 2004;229:101–11.
Physicochemical and thermodynamic investigation of the UT-3 hydrogen production cycle: A new technological assessment F. Lemorta,∗ , C. Lafona , R. Dedryvèreb , D. Gonbeaub a CEA-Marcoule, BP 17171, 30207 Bagnols sur Cèze Cedex, France b LCTPCM UMR 5624 Helioparc Pau Pyrénées, 2 avenue Pierre Angot, 64053 Pau Cedex 9, France
Available online 23 September 2005
Abstract The previous study of the UT-3 thermochemical cycle was based on laboratory-scale tests to assess the feasibility of each step. In order to enclose the reactants in the experimental vessel the study was carried out on pellets that had to be prepared according to a specific protocol. It would appear difficult to design and operate a cost-effective industrial process using this technique, and it could be worthwhile to consider another approach. This paper identifies problem areas in the cycle as expressed below: (A) (B) (C) (D)
CaO + Br 2 → CaBr 2 + 21 O2 , CaBr 2 + H2 O → CaO + 2HBr, Fe3 O4 + 8HBr → 3FeBr 2 + 4H2 O + Br 2 , 3FeBr 2 + 4H2 O → Fe3 O4 + 6HBr + H2 .
A thermodynamic approach predicts that the first and third reactions will occur easily, but a physicochemical approach predicts some difficulties due to sintering of the solid reactants. With regard to the second and fourth reactions, thermodynamics would dictate operation at low pressure and high temperature, in contradiction with the volatility of the bromides. One solution could be to ensure superstoichiometric conditions, which assumes a large quantity of gas. Based on these considerations, and on the fact that all the steps of this thermodynamic cycle are gas–solid reactions, a fluidized bed process is proposed with continuous feed of the gaseous reactants. The efficiency of the process is discussed. 䉷 2005 International Association for Hydrogen Energy. Published by Elsevier Ltd. All rights reserved. Keywords: Hydrogen production; Thermodynamical cycle; UT-3 cycle; Technology
1. Description of the UT-3 cycle The UT-3 thermochemical cycle has been largely described in previous papers regarding the concept of the technology (MASCOT) that can be implemented in order to produce the hydrogen from the cycle [1,2] but also regarding the basic data that have to be considered for its assessment ∗ Corresponding author.
E-mail address: fl[email protected] (F. Lemort).
[3,4]. The first results obtained lead to improve the reactivity of the whole reactants [5,6] and to consider other primary heat source than nuclear power in order to offset the low efficiency of an industrial loop [7]. A recent study points out that with the concept of the MASCOT technology it could be very difficult to increase the efficiency of the loop [8]. The UT-3 cycle is based on two pairs of chemical reactions. The first two ensure the formation of hydrobromic acid releasing oxygen, and the other two ensure the reduction of water by a bromide.
0360-3199/$30.00 䉷 2005 International Association for Hydrogen Energy. Published by Elsevier Ltd. All rights reserved. doi:10.1016/j.ijhydene.2005.07.011
F. Lemort et al. / International Journal of Hydrogen Energy 31 (2006) 906 – 918
can be a limiting factor since the following must be taken into account:
Nomenclature reaction temperature, ◦ C reaction pressure, bar time, s free enthalpy of reaction, kJ mol−1 entropy of reaction, J mol−1 K −1 heat of reaction, kJ mol−1 variation of gaseous species, mol degree of reaction progress flow of gaseous species, mol s−1 quantity of gaseous reactants, mol quantity of solids reactants, mol
T P t G◦ S ◦ H ◦ vg qg Ng Ns (A) (B) (C) (D)
907
CaO + Br 2 → CaBr 2 + 21 O2 , CaBr 2 + H2 O → CaO + 2HBr, Fe3 O4 + 8HBr → 3FeBr 2 + 4H2 O + Br 2 , 3FeBr 2 + 4H2 O → Fe3 O4 + 6HBr + H2 .
The choice of temperature and pressure conditions depends on the physical and chemical properties of the reactants and the thermodynamics of the reactions. The physical and chemical data concern mainly the phase transitions and the possible formation of azeotropes. The available data on all the products of the above-described system are indicated in Table 1. The thermodynamic data [9] concerning all the reactions are indicated in Table 2, where vg represents the variation in the number of moles of the gaseous species. The temperatures were selected to obtain realistic operating conditions. 2. Feasibility analysis A relatively simple cycle feasibility analysis was performed using published data. Each of the cycle reactions was considered. Reaction A The data suggest a quantitative, thermodynamically unlimited reaction. As this is a gas–solid reaction, the kinetics
• the accumulation of reaction products on the surface of the grain could lead to a protective layer, • diffusion phenomena (diffusion of reactants through a film of gas and solid), • adsorption and desorption phenomena, • decreasing of the active surface. It would be preferable to use a fluidized bed to ensure thorough agitation of the system and thus enhance the diffusion phenomena. Moreover, the possible appearance of a specific Ca3 OBr 4 compound cannot be discounted; this compound appears in a chloride system in which the phase diagram [10] exhibits a eutectic melting point. Reaction B The data in Table 2 reveal a non-quantitative reaction favored by low pressures and high temperatures. The degree of reaction progress can be related as follows to the CaBr 2 consumption, the total pressure system P , and the equilibrium constant K: b b K K = ≈ , (1) a K + 4P 2a P where a represents the quantity of CaBr 2 and b the quantity of H2 O. The graphic representation of =f (b/a) as shown in Fig. 1 shows that satisfactory reaction progress can be expected at low pressure values. The figure shows that operation at ambient pressure is relatively unfavorable, even with excess water. Low pressure operation appears to be more advantageous, provided the volatilization of species is limited, as suggested by the data in Table 1. As for the previous reaction, allowance must be made for the possible formation of specific fully immiscible compounds generating eutectic mixtures with melting points that can be below the operating temperature. Reaction C Like reaction A, reaction C is quantitative and thus not limited thermodynamically. The remarks concerning the
Table 1 Reactant properties [9,10] Species
Phase/Transition
Specified compound
Cp
CaO Br 2 CaBr 2 H2 O
S S → L: 59.6 ◦ C S → L: 742 ◦ C L → G: 100 ◦ C
CaO–CaBr 2 eutectic (Ca3 OCl4 )
≈ 51.3 (627 ◦ C) ≈ 37.4 (427 ◦ C) 59.2 + 0.0819T − 0.0001T 2 + 6 × 10−8 T 3 30.2 + 0.0109T
HBr
G
Fe3 O4 FeBr 2 H2
S S → L: 691 ◦ C G
CaO–CaBr 2 eutectic (Ca3 OCl4 ) H2 O–HBr azeotrope T = 126 ◦ C 47% HBr H2 O–HBr azeotrope T = 126 ◦ C 47% HBr Fe3 O4 –FeBr 2 eutectic? Fe3 O4 –FeBr 2 eutectic?
Vapor pressure (bar)
4 × 10−7 (727 ◦ C)
29.505 − 0.0032T + 6 × 10−6 T 2 86.4 + 0.2086T 73.6 + 0.0223T
5 × 10−3 (627 ◦ C)
908
F. Lemort et al. / International Journal of Hydrogen Energy 31 (2006) 906 – 918
Table 2 Thermodynamic data for the UT-3 cycle reactions [9] Reaction
T (◦ C)
G◦ (kJ mol−1 )
S ◦ (J mol−1 K−1 )
H ◦ (kJ mol−1 )
vg (mol)
(A) (B) (C) (D)
527 727 227 627
−40.9 104.0 −119.9 117.5
−41.8 107.4 −303.8 286.3
−74.3 211.4 −271.8 386.7
−0.5 +1.0 −3.0 +3.0
1.0 0.9 0.8 0.7
α
0.6 0.5
P = 1 bar
0.4
P = 5 mbar
0.3
P = 1 mbar
0.2
P = 0.1 mbar
0.1 0.0
0
10
20
30 nH2 O/n CaBr2
40
50
60
Fig. 1. Progress of reaction B versus the H2 O/CaBr 2 ratio at different total pressures.
64a 7 7 3(b + a)3 (b − 4/3a)
P 3 − K = 0, 4
1.2 1 0.8 α
kinetics of reaction A are also applicable here. The use of a fluidized bed at a temperature near 200 ◦ C could be interesting from this standpoint. The reaction is sufficiently exothermic that the use of a heat exchanger is recommended. Reaction D As for reaction B, this reaction is not quantitative and is favored by low pressures and high temperatures. It should be noted, however, that 627 ◦ C is not far from the melting temperature of FeBr 2 . Problems can be expected from sintering of the system, hence the unquestionable advantages of a stirred system to prevent or limit plugging. As with reaction B, the coefficient of reaction progress (chemical conversion) can be related to the equilibrium constant K and the total pressure P :
0.6 0.4 0.2 0
0
5
10
15
-Ln(P(bar))) Fig. 2. Progress of reaction D versus the total pressure.
relation (2) becomes (2)
where a represents the quantity of FeBr 2 and b the quantity of H2 O. Eq. (2) demonstrates that at a sufficiently low pressure for given material quantities, the reaction progress tends toward 1. For example, assuming stoichiometric proportions,
1399687 3(b + 4)3 (b − 4)4
P 3 − K = 0.
(3)
The graphic representation of =f (− ln P) (ln P has been chosen in order to facilitate the reading of the graphic) as shown in Fig. 2 indicates that satisfactory reaction progress can be expected at low pressure values. For example, a mean
F. Lemort et al. / International Journal of Hydrogen Energy 31 (2006) 906 – 918
pressure of 5 mbar ensures a progress of 0.6. As shown above for reaction B, excessively low pressures can be avoided by ensuring excess water, thus limiting the volatility of FeBr 2 which could otherwise be entirely volatilized.
909
decided to perform the calculation for stoichiometric reaction conditions—i.e. for a system comprising one mole of CaBr 2 and one mole of water—at temperatures ranging from 300 to 1000 K. The result shows clearly a total chemical conversion in the temperature range. The simulation shows that this bromination reaction should not raise any difficulties. The fact that the molar volume of CaBr 2 is about 3.6 times greater than the molar volume of CaO [3] indicates 70% expansion of the system that must be taken into account. Reaction B The initial evaluation showed that this reaction must occur at low pressure with excess water. We simulated the evolution of a system initially comprising one mole of CaBr 2 for 40 moles of water in a vessel pressurized to 5 mbar, with the result shown in Fig. 3. Unlike the previous results, hydrolysis of CaBr 2 leaves very little leeway for the process. Ideally, the process temperature would be 1100 K, which corresponds to the best chemical conversion. However, the melting point of CaBr 2 means the process temperature must remain below 1000 K, i.e. with a reaction progress of about 0.5 under the conditions described above. Reaction C The results (Table 2) show that the reaction is feasible. The simulation performed under standard conditions with stoichiometric quantities reveals that many species can appear depending on the working temperature. Fig. 4 shows the system evolution according to the temperature.
3. Numerical simulation The first step toward assessing the cycle feasibility is to consider the system variations versus temperature for specified material quantities. This can be done on the basis of the conclusions of the preceding section (operating pressure, superstoichiometry, etc.) by evaluating systems at equilibrium initially containing all the reactive species (solid reactants Ns and gaseous reactants Ng ). During a time interval t the actual process will be subject to a flow of gaseous species qg such that T Ng = qg dt. (4) 0
The calculation is free enthalpy minimization of a system initially defined (chemical composition and pressure) with respect to all the possible compositions compatible with the conservation of mass. The Factsage thermochemical database [9] and its computation software were used for this purpose. Reaction A The assessment in the preceding section showed that this reaction is theoretically highly feasible. We therefore 2.0 CaBr2+ 40 H2O P = 5mbar
HBr
Quantity (mol)
1.6
1.2
CaBr2 (H2O)6 CaBr2(s)
CaO(s)
0.8
0.4 Br (g)
0 300
400
500
600
700
800
900
CaBr2(l)
H2
1000
1100
T (K) Fig. 3. CaBr 2 + 40H2 O system response versus temperature at 5 mbar.
1200
910
F. Lemort et al. / International Journal of Hydrogen Energy 31 (2006) 906 – 918
4.8 Fe3O4+ 8HBr
4.4 4.0
HBr
P=1bar
H2O(g)
H2O(l)
Quantity (mol)
3.6 3.2
FeBr2 (s1)
2.8 2.4 2.0
FeBr3 (s)
FeBr2 (s2)
1.6 1.2
FeBr2( g) FeBr2 (liq)
Br2 Fe2O3 (s)
0.8
Fe2O4 (s)
0.4 0 300
400
500
600
700
800 T (K)
900
1000
1100
1200
Fig. 4. Fe3 O4 + 8HBr system response versus temperature at 1 bar.
6.0 5.6
3FeBr2 + 40H2O
4.2
P=1bar
HBr
4.8 4.4 Quantity (mol)
4.0 3.6 3.2
FeBr2 (s1)
FeBr2 (s2)
2.8 2.4 2.0 1.6 H2
1.2 0.8
FeBr2( g)
0.4 0 300
Fe3O4 (s)
400
500
600
700
800
900
1000
1100
1200
T (K) Fig. 5. 3FeBr 2 + 40H2 O system response versus temperature at 1 bar.
The optimum equilibrium occurs at 400 K. At higher temperatures, oxidizing hydrolysis of FeBr 2 begins to form Fe2 O3 and HBr. This reaction is liable to limit the cycle efficiency but does not prevent operation since Fe2 O3 is easily reduced to FeBr 2 by HBr. Nevertheless, it is advisable to maintain the temperature as close as possible to 400 K.
Reaction D Eq. (3) shows that the most advantageous conditions are with excess water (rather than under low pressure, which would be less efficient and more energy-intensive). Fig. 5 shows the calculated result. With 10-fold excess water, the maximum reaction progress is obtained at 900 K. However,
F. Lemort et al. / International Journal of Hydrogen Energy 31 (2006) 906 – 918
911
N2, H2
N2 Water
Tubular furnace
CaBr2 or FeBr2
Flask of gard
Water (HBr trap)
Fig. 6. Experimental setup for evaluating hydrolysis steps.
as observed earlier, the vapor pressure of FeBr 2 leads to appreciable volatilization of the product, making it preferable not to exceed 800 K for this step.
4. Experimental assessment
recovered powder. These analyses give the composition of the surface of the grains and then the analysis of the products of reaction. The chemical conversion is given by the analysis of the water contained in the HBr trap. The XPS analyzer was connected to the glove box, allowing the analyses to be carried out only a few minutes after the end of the test to avoid degradation of the reaction products.
4.1. Material and method 4.2. Experimental results A few experimental tests were performed to check some of the points mentioned above. The tests concerned the most limiting steps, i.e. the hydrolysis of CaBr 2 and FeBr 2 , and mainly involved estimating the reaction progress for these two steps. A relatively simple experimental setup was used, in which CaBr 2 or FeBr 2 powder was swept by steam carried in a nitrogen stream. The experimental device is shown in Fig. 6. The tubular furnace was made of silica, and the compounds were placed in an alumina boat (a silica boat could not be used because of the formation of silicates). The quantity of HBr dissolved in the trap gives the rate of chemical conversion. The products were kept and handled in a glove box under anhydrous nitrogen atmosphere. The test procedure was as follows: • 2 g ± 0.01 of bromide (CaBr 2 or FeBr 2 ) were placed in the sample boat. • The boat was inserted in the silica tube inside the tubular furnace in a dry nitrogen stream. • The nitrogen flow was maintained for 1 h at room temperature to eliminate all the oxygen. • The furnace was heated to the test temperature (725 ◦ C for CaBr 2 , 600 ◦ C for FeBr 2 ). • The nitrogen stream was then charged with steam by sparging and was maintained at the test temperature for hours. • The test was terminated, and the sample transferred to the glove box. Qualitative analysis was then performed by X-ray photoelectron spectroscopy (XPS) and by X-ray diffraction on the
(1) CaBr 2 → CaO The main conclusions after analyzing the residues were the following: • Major sintering of the solid mass left a compact compound at the bottom of the crucible. • The reaction progress reached about 11%. (2) FeBr 2 → Fe3 O4 The main conclusions after analyzing the residues were the following: • The product was volatile with recondensation of reaction products at various points in the sample boat. • The reaction progress reached 87%. • The major reaction product was Fe2 O3 . 4.3. Discussion The most important phenomenon observed during these two experiments was significant sintering in both cases. Sintering modifies the reactive surface finish and probably accounts for the decreased reactivity observed by other authors [3,6] during kinetic tests on CaO pellets incorporated in a CaTiO3 matrix. Aihara et al. [3] demonstrated that the reactivity diminishes as the temperature rises; this is likely attributable to reactant melting problems. The second feature observed was the partial progress of the hydrolysis reactions, the least favorable being only 11% for CaBr 2 . This was much lower than the value measured by Aihara et al. [3] in an experimental setup in which the
912
F. Lemort et al. / International Journal of Hydrogen Energy 31 (2006) 906 – 918
CaBr2 + H2O
3FeBr2 + 4H2O
CaO + 2HBr
Fe3O4 + 6HBr + H2 HBr
Ca reactor unit
Fe reactor unit
O2
H2
CaO + Br2
Fe3O4 + 8HBr
CaBr2 + 1/2O2
3FeBr2 + 4H2O + Br2
Fig. 7. UT-3 cycle flowsheet for the Mascot pilot facility.
gas passed through the sample; the difference may be attributable to the fact that the gas stream flowed over the sample in our experiment. The limited reaction progress is in any event related to passivation of the reaction interface, where the CaO or Fe2 O3 layer becomes thick enough to limit diffusion of the reactants or gaseous reaction products. The satisfactory progress (87%) observed for FeBr 2 is attributable to partial volatilization of FeBr 2 resulting in a gas–gas reaction with favorable kinetics. The third phenomenon observed was the preferential formation of hematite Fe2 O3 . This result is rather unusual for this kind of reaction since the hydrogen potential is high enough to keep the iron under the magnetite form. This result that is not in accordance with previous results obtained by other authors nor with the results of calculation showed in Fig. 5 could be explained by the reactive system which is always far from the equilibrium and by the experimental method that is different from the other and may favor the formation of hematite. Furthermore, the calculation of equilibrium constant at 900 K for the reaction modified reaction D giving the hematite Fe2 O3 (and described below) shows a value 38 times lower than for the production of magnetite. This difference is not very significant and both the reactions may be competitive. In this case, the experimental condition could favor one or the other way. (E) 2FeBr 2 + 3H2 O → Fe2 O3 + 4HBr + H2 . This is a very significant observation that changes the frame of the cycle since the amount of released hydrogen versus the amount of iron reactant is higher. If this observation is true, the efficiency of the cycle could be increased. This result that could be connected with the experimental condition has to be checked before going further on this way. The amount of hydrogen evolved in this reaction has to be very carefully analyzed.
5. Process considerations The process implementing the UT-3 cycle is based on the use of four series-connected reversible stationary beds in which the reactants are incorporated in a titanate matrix for calcium and a silica matrix for iron [2]. Fig. 7 shows how the reactors are arranged [1]. The Mascot process developed in Japan [1] has several advantages: • No transport of solids. • Reactant surface condition conserved by inclusion in matrices [2]. • Simple process design obtained by flow reversal. Nevertheless, the process has the double disadvantage of requiring a reagent preparation unit and increasing the quantity of material in the process by adding an inert phase representing 33–50% of the reactive load [3]. A material balance was determined to predict the operation of a unit capable of generating 30 000 Nm3 h−1 (i.e. about 90 MWth). Table 3 indicates the minimum masses of reactants (intended to react within 1 h) and inert phases required in the reactors, together with the gas mass and volume requirements. The worst-case situation with a 2:1 mass ratio between the reactive and inert materials was considered [3]. In the above mass balance, inert material (525 t) accounts for 25 wt% of the total solid mass. Eliminating the use of reactant matrices would appreciably diminish the total product mass while also reducing the number of individual operations necessary for sustained operation of a production unit (as it would be necessary to remove the reactant matrices periodically from the reactors for regeneration). In fact, although theoretically reactants are incorporated in matrices to ensure durability, they will not last indefinitely because
F. Lemort et al. / International Journal of Hydrogen Energy 31 (2006) 906 – 918 Table 3 Hourly reactive material mass necessary 30 000 Nm3 h−1 of H2 with the Mascot process Material
Mass (metric tons)
CaO CaBr 2 CaTiO3 Fe3 O4 FeBr 2 SiO2 Br 2 HBr H2 O Total
75 268 364 310 866 161 214 867 96 3221
to
produce
Volume (Nm3 )
30 240 120 390
– – – – – – 000 000 000 000
of cyclic expansion and contraction of the reactive agglomerates [3]. The resulting mechanical stresses will require periodic regeneration of the titanate (calcium unit) or silicate (iron unit) matrices. Moreover, on the calcium side of the process, decomposition of the CaTiO3 is likely by HBr diffusion through the layer that has already reacted, according to the following balance: (F) CaTiO3 + 2HBr → CaBr 2 + TiO2 + H2 O = 560.
K(800 K)
Even if the matrix is produced with a very small specific surface area, bromination of the titanate will result in decohesion of the “inert” matrix phase and the matrix will lose its mechanical properties. In order to limit the mass of the constituent materials and the matrix preparation operations which require several steps if the alkoxide process is used [3], it could be advantageous to use fluidized bed technology in which the calcium unit is coupled in a single reactor like the iron unit. The reactor design could take into consideration the properties of the reactants and reaction products.
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tion is ensured first by the gas flow that is enough to expand the solid reactive bed in the upflow portion and then by the gravity in the downflow zone into which the bromination of calcium oxide will occur. The volume difference between the two zones of the asymmetric torus will ensure a ratio of about 50:1 between the solid particle residence times, due not only to the gas velocity but also to expansion of the fluidized bed [3]. The gas flow direction is controlled by two blowers ensuring not only oxygen extraction at the bottom of the unit but also extraction of the HBr/H2 O mixture at the top; the two reactor zones are separated by non-return valves that allow solid particles to flow in only one direction. As the purpose of the calcium reaction cycle is to produce HBr, it could be highly simplified by focusing solar radiation on a reactor dedicated to the simple reaction: (G) H2 O + Br 2 → 2HBr + 21 O2 . In this case the UV radiation is used in order to increase the kinetics and the chemical conversion of this reaction by dissociating directly Br 2 homolytically or heterolytically [11,12] and to lower the amount of the heat that is required. 5.2. Iron reactor Tests showed that the FeBr 2 hydrolysis reaction progress is much greater for any given reaction time; this may be attributable to the slower CaBr 2 hydrolysis kinetics. These results were confirmed by Yoshida et al. [11], who reported that the hydrolysis rate is about 20 times higher for FeBr 2 than for CaBr 2 . The technology used must take this result into account, but could also allow for the differential properties of FeBr 2 and Fe3 O4 . The most important of these are the molar volumes, which vary in the same way as those of the calcium reaction zone, and the magnetic properties of the two compounds: Fe3 O4 is ferromagnetic with a susceptibility of about 106 , whereas FeBr 2 is paramagnetic with a susceptibility of about 10−2 [14]. These properties could thus be used to separate the two reaction products within a single reactor [15].
5.1. Calcium reactor 5.3. Process description Two reactions occur in this unit; the first (A) is quantitative and relatively rapid (assuming a suitable reaction surface condition), while the second (B) is much less favorable from a thermodynamic and kinetic standpoint [3,5]. It must also allow for a substantial volume variation of the process compounds: the molar volume of CaBr 2 is about 3.6 times greater than that of CaO [3]. Finally, Br 2 dissociates either homolytically under the effect of near-UV visible electromagnetic radiation [11] or heterolytically when higher-energy wavelengths are used [12]. A vertical nonsymmetric toroidal reactor [13] could be used in which hydrolysis could be carried out in the upflow portion, so the grain size reduction arising from striction would ensure that nearly all the particles in the downflow zone would have already finished reacting. The solid circula-
Fig. 8 illustrates a flowsheet for a facility capable of producing 30 000 Nm3 h−1 of hydrogen. The given pressures have been estimated (in bars) in the reactors, and upstream and downstream from each unit. It is assumed here that a membrane causes a pressure drop of about 2000 Pa. The calcium process zone consists mainly of a circulating fluidized bed reactor with asymmetric toroidal geometry in which CaBr 2 is hydrolyzed in the riser and CaO is brominated in the downer. Based on a maximum reaction progress of 70% [3], the reactor contains 140 t of calcium distributed between CaBr 2 and CaO. The reactor can be designed with reference to a few examples implementing circulating fluidized beds [14] by assuming the limiting step is the hydrolysis of CaBr 2 . The proposed flowsheet recommends low
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Fig. 8. Simplified diagram of a process implementing the UT-3 cycle.
pressure to facilitate the hydrolysis step. If this approach is too energy-intensive, this step could be enhanced by UV radiation [11] with both reactions operating at atmospheric pressure. The design basis criterion would be the reactant residence time in the reactor. According to the results reported by Yoshida et al. [4], the residence time for CaBr 2 should be about 1 h in a stationary bed configuration. The use of a fluidized bed reactor to stir the system could shorten this time by half. The residence time for CaO bromination should be 10 min. As indicated in Fig. 1, the temperatures can be identical at 650 ◦ C. Reactor design proposals can be based on the data consolidated in Table 4. The gas flow rate will fluidize the beds (expansion in the first case, contraction in the second), depending in fact on the particle size used in the reactor. Although the working temperatures were lowered to limit degradation of the reactant particles, provision must be made for calibrating the particles. This could be done using a pan mill calibrated for the required particle size. Installing two mills in the downflow zone of the torus would not only recalibrate the particles but also separate the two zones and thus prevent any mixing of their atmospheres.
The same principle could be used for the reactor in the second zone except that applying a permanent field could facilitate magnetite expansion in the upper reactor zone and thus its transfer to the downflow leg. The temperatures must be more closely controlled than in the calcium reactor zone to avoid the formation of undesirable compounds as shown in Figs. 3 and 4 (the superstoichiometry actually taken into account will be much less than 10 to avoid lowering the overall cycle efficiency). The hydrolysis temperature was set at 600 ◦ C to limit the quantity of FeBr 2 entering the gas phase (the possible volatilized fraction will be recovered by condensation at high temperature (500 ◦ C) at the level of the off-gas filter). The hydrobromination temperature is limited to 130 ◦ C to prevent reoxidation of FeBr 2 by Br 2 /H2 O. The proposed geometric characteristics are indicated in Table 5, once again allowing for a 70% limit on the reaction progress. The off-gas treatment proposed in Fig. 7 assumes the calcium zone includes a membrane separating the HBr/H2 O mixture; the water purity must be as high as possible to ensure hydrolysis of the CaBr 2 . Considering that FeBr 2 hydrolysis is less sensitive to the presence of HBr, it is proposed that HBr be separated from its azeotropic mixture by fractionated condensation. If this is unsuitable, a membrane
F. Lemort et al. / International Journal of Hydrogen Energy 31 (2006) 906 – 918
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Table 4 Dimensional characteristics of a torus reactor in the calcium process zone Reaction zone
Upflow
Downflow
Reaction Residence time Reactant mass Reactant volume (compact) Expansion factor Reactor volume Temperature Type of bed Gas velocity Zone diameter Zone height
CaBr 2 → CaO 30 min 134 t 40 m3 4 160 m3 650 ◦ C Fluidized (CaO particles expand) 2 m s−1 4.1 m 12 m
CaO → CaBr 2 10 min 16 t 5 m3 – 217 m3 650 ◦ C Fluidized (CaBr 2 particles settle) 1.5 m s−1 4.8 m 12 m
Table 5 Dimensional characteristics of a torus reactor in the iron process zone Reaction zone
Upflow
Downflow
Reaction Residence time Reactant mass Reactant volume (compact) Expansion factor Reactor volume Temperature Type of bed Gas velocity Zone diameter Zone height
FeBr 2 → Fe3 O4 30 min 563 t 122 m3 4 488 m3 600 ◦ C Fluidized (Fe3 O4 particles expand or migrate magnetically) 2.5 m s−1 10 m 6m
Fe3 O4 → FeBr 2 10 min 67 t 13 m3 – 292 m3 130 ◦ C Fluidized (FeBr 2 particles settle) 2 m s−1 8m 6m
system could be used. Like the material quantities indicated in the dimensional tables, the gas flows are based on the production of 30 000 m3 h−1 of hydrogen. The water masses X indicated in the diagram correspond to the possible superstoichiometry necessary to ensure satisfactory progress of the hydrolysis reactions. The bromine that does not react during the CaO bromination step is separated during cooling of the gas prior to storage of the oxygen produced. In the iron zone, separation can be obtained by condensation or possibly using a specific membrane at high temperature. Contact with the heat source (not shown in the diagram) will be made in the torus upflow legs, for example by fluid flowing in a jacket. The values given in Tables 4 and 5 show clearly that the facilities can be huge for the production of 30 000 Nm3 h−1 of hydrogen. The process proposed in Fig. 8 can be divided into smaller unit in series in order to facilitate the handling or maybe the running of the process.
6. Efficiency The efficiency of the UT-3 using this technology has not yet been assessed in detail, but the issues have been exam-
ined on the basis of work by Sakurai et al. [7] and by Teo et al. [8] concerning the flowsheet proposed in Fig. 8. The efficiency may be simply defined as follows: =
Q(H2 ) + Qe , i Qi
(5)
recombination, Qe where Q(H2 ) is the heat of hydrogen is the excess thermal energy and Qi is the total thermal energy supplied to the process, including the energy recombined as electric power with 35% efficiency to operate the process equipment. As the excess energy Qe is difficult to estimate accurately, we chose to assign a highly unfavorable zero value to this parameter. The calculation was performed here for the excess water X1 and X2 in the calcium and iron reactors, respectively. These values affect the quantity of material circulating between the reactors, as well as the heat flux necessary to heat the gas streams to the working temperature. Contrary to the adiabatic approach by Sakurai et al. [7], we preferred to use isothermal reactors to avoid raising the reactant temperatures and thereby encountering difficulties related to sintering of the materials and the formation of undesirable reaction products such as Fe2 O3 in the downflow leg of the iron reactor, which would require separation of the HBr produced. The reaction isothermicity
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Table 6 Estimated UT-3 cycle power requirement Item
T (◦ C)
Inlet pressure
Outlet pressure
Power (W)
A B
650 650
0.1 0.18
18 0.2
C
150
P6
1
D
600
P2
1
E F
100 650
P5 0.13
498 123 394(25 + X1 ) 20 000 + 294X2 703 364 ln 24 000 + 22.05X 2 (100 + 1.47X2 )2 7258(100 + 1.47X2 ) ln 2.16X22 + 70.56X2 + 4800 77 528 ln[25(100 + 1.47X2 )] 114 693
25 0.2
Table 7 Heat exchanger energy balance Heat exchanger
Inlet temperature (◦ C)
Outlet temperature (◦ C)
Power (W)
Recoverable fraction (%)
E1 E2 E3 E4 E5 E6
150 650 130 100 100 650
650 150 600 150 126 25
462 500 737 500 1598 000 + 23 491X2 221 250 1516 800 + 21 690X2 268 750
– 50 – – 40 30
1225 655 + 14 815X2
Total
35 30 Efficiency (%)
can be ensured by the use of a jacketed reactor to supply heat to the upflow legs (endothermic reactions) and to remove heat from the downflow legs (exothermic reactions). The recovered heat can be used in part to vaporize 24 t h−1 of water. Although the flowsheet shown in Fig. 8 corresponds to the production of 30 000 Nm3 h−1 of hydrogen, the efficiency was estimated for a 25 mol s−1 hydrogen production cycle as in the article by Sakurai et al. [7]. In this case, the power available from the hydrogen produced Q(H2 ) is 6.05 MW. The heat input required to ensure reaction isothermicity is 7.7 MW. The power requirements for the blowers were estimated in terms of the excess water X1 and X2 using a set of isothermal transformations. Table 6 indicates the power exchanged in the heat exchangers Ei . The power necessary to operate the calibrators must be added to the balance given in Table 7. It can be roughly estimated at about 200 kW per reactor. The efficiency is calculated using relation (4). It is interesting to perform the calculation with the data established for the process flowsheet (Fig. 8), but also to consider the use of two membranes rather than fractionated condensation to separate the H2 /HBr/H2 O gases. Using membranes would avoid the need for a cooling/heating cycle to separate these species. The calculation assumes the process equipment (except for the heat exchangers) is powered by electricity obtained with 35% efficiency from the waste heat. Considering the pres-
25 Membrane 20 Condensation 15 10
0
5 10 15 20 Stoichiometry 4H2O/3FeBr2
25
Fig. 9. Cycle efficiency versus water stoichiometry in the iron reactor.
sure in the calcium reactor is 0.2 bar and not 5 mbar, the stoichiometric ratio must be greater than 40 as indicated by the results shown in Fig. 3. A ratio of 250 yields similar results; we adopted this value since the excess water inventory in the second reactor determines the overall efficiency. Fig. 9 shows that the water quantities have a major effect on the overall cycle efficiency in the case of fractionated condensation. The efficiency is about 15% with a super-
F. Lemort et al. / International Journal of Hydrogen Energy 31 (2006) 906 – 918
stoichiometry of 10 (to obtain results comparable to those shown in Fig. 5). The efficiency rises to 22.6% with the use of membrane technology. This is consistent with the conclusions reported by Teo et al. [8], and highlights the need to develop efficient membranes with a low pressure drop to allow thermochemical cycles to operate under optimum conditions. In the present case, the efficiency is not diminished by flow reversibility and the resulting temperature changes in the solid masses. In addition, fluidized beds appreciably diminish the pressure drop compared with a stationary bed and thus also limit the energy requirements. Finally, eliminating the regeneration steps for reactants embedded in matrices certainly allows appreciable energy savings, although the amount has not been estimated to date. We chose not to perform calculation regarding a UT3 cycle involving the production of hematite Fe2 O3 because it is first necessary to check the experimental results.
7. Conclusion A physicochemical evaluation of the UT-3 cycle based on current trends in the literature reveals a number of difficulties. First of all, the systems implemented can lead to the formation of mixtures that converge toward low melting-point eutectic compositions or azeotropes that complicate process operation. The occasionally unfavorable thermochemistry highlights the need for modifying some operating parameters, including the reactant stoichiometry or the working pressure. Recirculation of gaseous reactants can be advantageous in this regard. The present study together with a few experimental verifications has identified a new process flowsheet implementing two asymmetric torus reactors with fluidized beds of solid reactants in each leg; the solid particles transit from one leg to the next via a recalibration system according to the degree of reaction progress. This technology has several advantages, the first of which is that it avoids a probably energy-intensive reactant preparation step in a stationary bed by incorporating them in titanate or silicate matrices for use with calcium or iron. The second advantage is that it improves the reaction kinetics while ensuring continuous operation of both reactors, which in this case do not have to be reversible and therefore subject to major efficiency losses. The calculated process efficiency is about 15% if the gas separation is partly ensured by fractionated condensation; the use of dedicated membranes increases the mean efficiency to 22.5% and highlights the importance of developing this type of separation technique. This assessment shows the high necessity to make significant advances in improving the efficiency of membrane technology to avoid the energy loss in the separation using condensation. Advances have also to be made in the field of the process technology in order to perform the assessment of the proposed technology.
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Acknowledgements The authors are grateful to the referee whose remarks allowed improvements in the presentation of their work.
Appendix Calculation of for the reactions B and D: Reaction B CaBr 2 +H2 O t =0 a b t= ∝ a(1 − ) b − a K=
=
→ CaO+ 0 a
2HBr, 0 2a,
2 PHBr 4a 2 2 = 2 P, PH2 O b − a 2 2
b a
K . K + 4P
Reaction D 3FeBr 2 + 4H2 O → Fe3 O4 + 6HBr + H2 , t =0 a b 0 0 0 t= ∝ a(1 − ) b − 4/3a 1/3a 2a 1/3a, P 6 PH2 64a 6 6 a (a + b)4 3 = P , K = HBr 4 3(a + b)7 (b − 4/3)4 PH O 2
=
64a 7 7 3(b + a)3 (b − 4/3)4
P 3.
References [1] Yoshioka H, Nakayama T, Kameyama H, Yoshida K. Operation of a bench-scale plant for hydrogen production by the UT-3 cycle. Hydrogen energy progress V: proceedings of the fifth World hydrogen energy conference. Toronto, Canada, 15–20 July 1984. p. 413–20. [2] Kameyama H, Tomino Y, Sato T, Amir R, Orihara A, Aihara M, Yoshida K. Process simulation of mascot plant using the UT-3 thermochemical cycle for hydrogen production. Int J Hydrogen Energy 1988;14(5):323–30. [3] Aihara M, Umida H, Tsutsumi A, Yoshida K. Kinetic study of UT-3 thermochemical hydrogen production process. Int J Hydrogen Energy 1990;15(1):7–11. [4] Yoshida K, Kameyama H, Aochi T, Nobue M, Aihara M, Amir R, Kondo H, Sato T, Tadokoro Y, Yamaguchi T, Sakai N. A simulation study of the UT-3 thermochemical hydrogen production process. Int J Hydrogen Energy 1990;15(3): 171–8. [5] Sakurai M, Tsutsumi A, Yoshida K. Improvement of Ca-pellet reactivity in the UT-3 thermochemical hydrogen production cycle. Int J Hydrogen Energy 1995;20(4):297–301.
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F. Lemort et al. / International Journal of Hydrogen Energy 31 (2006) 906 – 918
[6] Sato T, Sakurai M, Matsumura Y, Tsutsumi A, Yoshida K. Kinetics and mechanism of hydrolysis of a Ca pellet in the UT3 thermochemical cycle for hydrogen production. Hydrogen energy progress XII: proceedings of the 12th World hydrogen energy conference, Buenos Aires, Argentina, 21–26 June 1998. p. 561-69. [7] Sakurai M, Bilgen E, Tsutsumi A, Yoshida K. Solar UT-3 thermochemical cycle for hydrogen production. Solar Energy 1996;57(1):51–8. [8] Teo ED, Brandon NP, Vos E, Kramer GJ. A critical pathway energy efficiency analysis of the thermochemical UT-3 cycle. Int J Hydrogen Energy 2005;30(5):559–64. [9] FACTSAGE database, Thermofact, Version 5.1-2002. [10] Wenz DA, Johnson I, Wolson RD. J Chem Eng Data 1969;14(2):252.
[11] Lindeman TG, Wiesenfelda JR. Photodissociation of Br 2 in the visible continuum. J Chem Phys 1979;70(6): 2882–8. [12] Maazouz M, Maazouz PL, Jacobs DC. Anomalous chargetransfer behavior in the scattering of hyperthermal Br+(3P2) on Pt(111). J Chem Phys 2002;117(24):10917–20. [13] Benkhelifa H, Legrand J, Legentilhomme P, Montillet A. Study of the hydrodynamic behavior of the batch and continuous torus reactor un laminar and turbulent flow regimes by means of tracer methods. Chem Eng Sci 2000;55(10):1871–82. [14] David R Lide, editor. CRC handbook of chemistry and physics. 79th ed. Boca Raton, FL: CRC Press; 1998. [15] Song MC, Lee KJ. A study on a magnetic separation of radioactive corrosion products from NPP using permanent magnets. Nuclear Eng and Design 2004;229:101–11.