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15GENERAL COMMENTS ABOUT THE EFFICIENCY OF THE IODINE-SULPHUR CYCLE COUPLED TO A HIGH TEMPERATURE GAS COOLED REACTOR
S. Goldstein, X. Vitart, J. M. Borgard
CEA/DEN/DPC-CEN Saclay 91191 Gif sur Yvette, France E-mail: stephen.goldstein@cea.fr
Scientific consultant DEN/DPC Dir Saclay Nuclear Center
Research activities: currently retired and scientific consultant of the Dep.
of physico-chemistry.
Research interests:
■ hydrogen production by water splitting: thermochemical cycles and elec-
trolysis;
hydrogen storage in metallic hydrides;
isotopic separation processes — course at Ecole Nationale Supérieure des
Techniques Avancées.
Educational background:
Goldstein Stephen
engineer graduated from École Polytechnique; MS of reactor physics.
Professional experience:
1980-1988 — head of the Heat Transfer and Fluid Flow Lab., Saclay;
■ 1988-1990 — head of the Process Evaluation Section for Uranium Enrichment, Saclay;
■ 1990-1993 — assistant of the head of the Dep. of Analytical Analysis and Physical Experimentation;
■ 1994-2001 — head of the Physico-Chemistry Lab.;
■ 2001-2003 — scientific assistant of the Physico-Chemistry Dep., in charge of hydrogen production processes;
■ December 2003 — Retired — scientific consultant. Publications:
■ few publications on the AVLIS isotopic separation process;
■ course on isotopic separation;
■ several publications on hydrogen production, mainly on the iodine-sulphur process.
The CEA has initiated an important R&D programme to evaluate the potentialities of the thermochemical cycles for massive production of hydrogen. In this paper, the iodine-sulphur (IS) cycle which is the most popular cycle is discussed from the point of view of the efficiency. This cycle is divided in three sections: the Bunsen section, the oxygen production section where the sulphuric acid is separated from the water, boiled and decomposed and the hydrogen production section where the hydriodic acid is separated from the HIX mixture and decomposed.
In a first step an upper bound of the efficiency of this cycle is estimated, following the method of Esteve and coll (Entropie. 1975. No. 61). The cycle is decomposed in elementary reactions (including separations and phases changes). Each reaction is supposed to be performed at a temperature as close as possible from the temperature at which the free Gibbs energy of the reaction vanishes. The reactants are brought from the ambient temperature to the reaction temperature and the products are brought back to the ambient temperature. The heat and work requirements of the reactions are then calculated assuming that the reactions are performed reversibly. On the way, the thermodynamic of the reactions data and their uncertainties are discussed. The upper bound is found to be 48 % for a maximum temperature of 1144 K delivered by the nuclear reactor.
In a second step, the efficiency is calculated on a more realistic basis. The heat request of the oxygen production section is taken in the literature while an energy and mass balance calculation of the hydrogen production section is performed with the chemical engineering code PROSIM. The hydriodic acid is separated and decomposed in a single reactive distiller. The liquid vapour equilibrium model used in the code is that of Neumann fitted on Knoche's total pressure measurements. It appears that the assumptions made concerning the internal energy recovery have a significant influence on the efficiency.
In conclusion, the R&D needs to improve this cycle are reviewed. The excesses of water and iodine in the Bunsen reaction must be reduced and membranes can be used to reduce the energy requirements provided they are enough selective and permeable. Finally there is a need to complete and improve the thermodynamic data of the HIX mixture.
Introduction
Among the large scale, cost effective and environmentally attractive hydrogen production processes, Sulfur Iodine (IS) thermochemical cycle [1] seems to be a quite promising one. This cycle ¡1 was originally studied in the 80's by General Atom-t ics Corp. (GA) [2] and can be split into the fol-f lowing reactions:
au
| (9I2); + (SO2)g + (16H2O); ^
I ^ (2HI + 10H2O + 8I2) + (1)
f + (H2SO4 + 4H2O) [120 °C];
§ L2 = (2HI + 10H2O + 8I2) ^
^ (2HI)g + (10H2O + 8I2), [230 °C];
(2HI)g ^ H2 + (12)/ [330 °C]; (3)
A = (H2SO4+4H2O) ^ (4)
^ (H2SO4) + (4H2O), [300 °C];
(H2SO4) ^ (H2SO4)g [360 °C]; (5)
(H2SO4)g ^ (SO3)g + (H2O)g [400 °C]; (6)
(SO3)g ^ (SO2)g + |O2 [870 °C]; (7)
The temperatures between brackets are approximate and depend upon the pressure which is
not necessarily uniform in the different parts of the cycle.
The first reaction, called Bunsen reaction, proceeds exothermically in liquid phase and produces two immiscible aqueous acid phases which compositions are indicated between brackets: L1 phase which is aqueous sulfuric acid and L2 phase which is a mixture of hydrogen iodide, iodine and water named HIX. In the second reaction, HI is separated from L2. This separation is the most critical phase of the cycle. Several separation processes have been proposed but today, none is demonstrated and much remains to do. Reaction (3) is the thermal decomposition of HI. Knoche [3] proposed to perform reactions (2) and (3) in the same reactive distillation column. Reaction (4) is the separation of L1 in H2SO4 and H2O. Up to now only distillation has been proposed for this separation. Several distillation flowsheets are found, ranging from the simple column [4] to multi effect arrangements [5, 6, 9]. This step is energy consuming due to the large number of water moles to evaporate. Reactions (5) to (7) proceed in gas phase and produce H2O, SO2 and oxygen. These gases are cooled down before to bubble in the Bunsen reactor to separate oxygen from SO2 and H2O. Due to the fact that reaction (7) is incomplete, a residual quantity of SO3 is found in the hot gases at the outlet of reaction (7). This SO3 is recombined to H2O in a reactor where the reverse
Section II
Fig. 1. Sketch of the IS-cycle
of reaction (6) is performed and the produced diluted H2SO4 is recycled in reaction (4). US teams are currently working on ceramic membrane separation as an alternative process to separate oxygen at high temperature. It has the double advantage to shift the SO3 decomposition reaction and to increase the partial pressure of SO2 in the Bunsen reactor. Besides it reduces the amount of H2SO4 recycled. The whole cycle can be divided into three sections according to the GA nomenclature [2], we shall call section I reaction (1), section II reactions (4) to (7) and section III reactions (2) and (3). A sketch of the cycle is shown fig. 1.
In this paper, we would like first to find an upper bound of the efficiency of this cycle. Therefore, in next paragraph, we assumed that the energy requirement of the reactions (1) to (7) is the minimum energy prescribed by the thermodynamics, and that the decomposition reaction (7) is complete, therefore no H2SO4 is recycled. In paragraph III, we review the published efficiency calculations and present our best assessment of section III. A more realistic efficiency of the whole cycle is then deduced. Finally we conclude with the R&D needs to bring this cycle to an industrial level.
Efficiency bound
We define the thermal efficiency as the ratio of the enthalpy of the hydrogen and oxygen recombination reaction at ambient temperature and
pressure AH0O (Ta) = 286 kJ/mole to the total heat
requirement of the cycle:
AHH2O (T ) W '
Q+-
Q
> AH - T AS ;
T
1 —a
T
Q - Qa =Ah .
All the energy input is heat, and in the reversible case, the turbine is supposed to have a Carnot efficiency, which is too optimistic. Consequently we shall use the first formulation. Hence, the heat requirement is — + (AH - AG) if AH -AG > 0 and
AG
-if AH - AG < 0. In the latter case, the reaction
%
is exothermic and the heat of reaction is lost, or can be recovered for internal use if T is high enough.
Thermodynamic data of the reactions. The optimum pressures of the three sections of the cycle are not equal due to thermodynamic reasons. The temperature of section I must be higher than the melting point of iodine but not too high to avoid the occurrence of the reverse reaction. Therefore, we assumed that this reaction (1) is performed at 400 K and 2 bar. The pressure of section II must be low, due to the fact that the number of gaseous moles is increasing, therefore, the low pressure shifts the reactions towards completion. We have chosen 1 bar for this section. Finally the best pressure for section III is 50 bars. The more the pressure of this section is high, easier is the distillation of HI, because, the position of the azeotrope depends upon the pressure. 50 bars seems a reasonably high value. Of course, pumping power is needed to adjust the pressures. They will be taken into account in the energy balance.
The Bunsen reaction (T = 400 K, P = 2 bars) can be split into four elementary reactions
Q and W are the heat and work requirements and nr the efficiency of the conversion system. We used for the numerical calculations nr = 0.5, which corresponds to the efficiency of the GTMHR.
Let AH, AS and AG be the enthalpy, entropy and Gibbs free enthalpy of a chemical reaction at a given temperature T, the heat Q and work W requirements are:
Q < T AS ; W >AH-TAS = AG; W + Q = AH.
The signs " <, > " are replaced by "=" when the reactions are reversible.
If the turbine which produces W in included in the system, then we must consider the heat Qa rejected to the ambient temperature Ta. In this case:
W = 0;
(SO2)g + (2H2O) ^ (^БОД + H2, AH = 53 kJ/mole, AG = 95kJ/mole;
(U + H2 ^ (2HI)g, AH = 34 kJ/mole, AG = -11 kJ/mole;
(H2SO4X + (4H2O) ^ (H2SO4 + 4H2O), AH = -58 kJ/mole, AG = -66 kJ/mole; (2HI)g + 8I2 + (10H2O) ^ (2HI + 10H2O) + 8I2 AH = -122 kJ/mole, AG = -121 kJ/mole;
(8) (9) (10) (11)
(8) and (9) are endothermic and their heats of reaction can be computed with standard thermody-namic tables [7]. (10) and (11) stand for the exothermic dissolutions of H2SO4 and HI in water. Their heats of solution are computed using Engels strong acid model [16]. In (11) iodine is supposed to have no influence on the heat of solution. This assumption is probably wrong because evidence of ion solvation with iodine molecules has been demonstrated in several papers. Calabrese V. [8] for example, shows experimentally that I- ions exist in dilute solutions and theoretically that I2xH+ should exist as well. Unfortunately, there is no direct measurement of the heat of the Bunsen reaction. We shall then assume that we can add the heats of (4) to (7), which gives -93 kJ/mole.
The data for the other reactions are found for example in [7].
Given the fact that the sulfuric acid boils at 612 K at 1 bar and that the heat of vaporization is 58 kJ/mole, we can draw the Q/T diagram of section II (fig. 5).
We assumed that the maximum helium temperature was 1144 K and the return to the reactor temperature was 760 K. The pitch point was taken equal to 17 °C.
We have now all the inputs to calculate the total heat request. Section II
Reaction (4): the temperature has a slight influence on the heat requirement of this reaction. Therefore, we supposed that this reaction is the reverse of reaction (10). Hence, AH = 58 kJ/mole and AG = 66 kJ/mole. The reaction can be performed with 66 kJ of mechanical work and releases 8 kJ heat. The heat demand is then: 66/0.5 = = 132 kJ/mole.
Reaction (5): heat demand 58 kJ/mole.
Reaction (6): according to fig. 3 and 5, the reaction is performed at AG = 0, therefore, the heat demand is TAS = 94.3 kJ/mole.
Reaction (7): due to the pitch of 17 °C, the reaction occurs with a slightly positive AG i. e.: AG = 4.5 kJ/mole and AH = 97.6. The heat request is then: (97.6 - 4.5) + 4.5/0.5 = 102 kJ/mole.
The hot gases at the outlet of reaction (7) are supposed to exchange exactly their heat with the gases at the outlet of the Bunsen reaction. Hence, the total amount of heat request for this section is 132 + 58 + 94.3 + 102 = 386 kJ/mole. Section III
Reaction (2) is the reverse of reaction (11), but occurs at higher temperature and pressure. At 500 K and 50 bars, we have AH = 66 kJ/mole and AG = 77 kJ/mole. Due to the fact that the mixing enthalpy depends upon the temperature, we cannot assume that the specific heats of the reactants and products are the same. Therefore we shall take AH = 122 kJ/mole (opposite of reaction (11)) and AG = 77 kJ/mole. Hence the heat requirement is: (122 - 77) + 77/0.5 = 199 kJ/mole.
Reaction (3): according to fig. 4, AG =12 kJ/ mole, AH = 24 kJ/mole, therefore the heat is request is 12/0.5 = 24 kJ/mole.
The two reactions occurring in the same reactor, we added the AG and AH, which results in AG = 89 kJ/mole and AH = 98 kJ/mole. The heat requirement is then 89/0.5 + 98 - 89 = 187 kJ/mole. Pumping power
The pumping power (mechanical efficiency of 0.75) to raise the pressure of SO2 from 1 to 2 bars RT ln(2)
is
0.75
= 3 kW and the pumping power of sec-
tion III (see next paragraph) is 5.3 kW. The additional heat due to the pumping power is then (3 + 5.3)/0.5 = 17 kW.
Fig. 2. Thermodynamic data for reaction (7), P = 1 bar
Fig. 3. Thermodynamic data for reaction (6), P =1 bar
Fig. 4. Thermodynamic data for reaction (3)
kJ/mole
Fig. 5. QT-diagram of section II
According to these figures, an upper bound
286
for the efficiency is nh =
386 +187 +17
= 0.485.
Best estimate of the efficiency
Many estimations are found in the literature. Either for the whole cycle or for sections II or III taken separately. The first efficiency estimation was published by GA [2] in 1981. H2SO4 was separated from water by a constant pressure multi effect distillation and HI was separated from HIx by using phosphoric acid which had to be regenerated. The heat demand of section II was 460 kJ/mole H2 and of section III, 148 kJ/mole. The total efficiency was then 286/(421 +T187) = 0.47. The value for section III seems low.
Fig. 6. Influence of L on the maximum bubble temperatures
Other papers deal with the heat demand of section II. In a recent work GA [9] using a more accurate thermodynamic model, estimated the heat demand at 420 kJ/mole. The pressure was decreased from 35 bar to 0.07 bar. The reason of this variation is to recover the maximum heat for internal reuse in the first stages of the distillation. Ozturk [5] proposed a flowsheet where the produced oxygen is used to boil by direct contact the sulfuric acid. The heat demand was 441.5 kJ/mole. In an alternative flowsheet, Schepers [6] proposed an increasing pressure multi effect distillation process. The heat demand could be decreased to 389 kJ/mole. Finally the author [4] found that with a simple distillation column a heat demand was 520 kJ/mole. The conclusion of these results
Fig. 7. Ternary diagram of the system H2O-HI-I2
is that there should a compromise between efficiency and complexity and therefore cost of this section and that the optimum design is not necessarily the one with the minimum heat request. In the following we shall take the GA value of 420 kJ/mole as a best estimate value.
Much less results are available for section III, due to the lack of thermodynam-ic data of the HI mixture.
x
Roth and Knoche [10] calculated a reactive distillation column and external circuits and found an overall heat request of 237 kJ/mole. Unfortunately this calculation could not be reproduced due to a mismatch in the liquid and vapor flows in the column. Anyway, this value seems very low. The Japanese teams use electrodialysis to concentrate HI in HI before to distillate the enriched HI ,
x X7
but their publications are not yet available.
The calculation of the reactive distillation of section III is very difficult because of the complex behavior of the vapor-liquid equilibrium of HIx. In the next section we describe briefly the model used and give preliminary results concerning the heat requirement of this section.
Available experimental data. Experimental data are available for the binary systems I2 HI-H2O and for the ternary HI-I2-H2O
The I2-H2O system has been studied by Kracek [11] in 1931. He measured the solubility of iodine in water and put in evidence a miscibility gap lying between the solid-liquid equilibrium point at 112.3 °C (very close to the melting point of iodine 113.7 °C) and an upper temperature of approximately 280 °C. At 112.3 °C, the light aqueous liquid contains a very small amount of iodine (0.05 mole %) and the heavy liquid 98 mole % of iodine.
Total pressures of HI-H2O mixtures have been measured by Wüster [12] and the mixing enthalpies by Vanderzee [13]. This system exhibits an azeotrope whose precise location depends on temperature and pressure. At ambient temperature, the azeotrope corresponds to a molar frac-
'2-H2O,
tion of HI of almost 15 %. For HI concentrations higher than the azeotrope, the vapor phase is very rich in HI, and for high temperatures (>200 °C), HI might be dissociated in the vapor phase in H2 and I2. The kinetics of this reaction is in principle very slow but possible catalytic effects or hydrogen production in the liquid phase might appear as discussed by Berndhauser [14]. We shall assume in this paper that the decomposition reaction of HI takes place only in the vapor phase. These experimental results have been used by Engels [16] to fit the binary interaction parameters of his model, using the Wilson model for the calculation of the activity coefficients.
Detailed total pressure measurements for the HI-H2O-I2 ternary system for a [HI]/[H2O] ratio up to 0.19 and various iodine concentrations are found in Neumann [15]. A synthesis of these results has been published by Engels and Knoche [17]. Experimental location of borderlines of the ternary phase diagram is found in [2]. From these results, it appears that the liquid phase exhibits two miscibility gaps which are the extensions in the ternary diagram of those found for the binaries I2-H2O and HI-H2O. The I2-H2O miscibility gap tends to disappear rapidly with increasing HI con-
centration probably because of the formation of polyiodides ions like I- or I2H +.
Neumann's thermodynamic model. The ther-modynamic model for the H2O-HI-I2-H2 reactive-liquid-liquid-vapor system proposed by Neumann is based on the following assumptions:
■ hydrogen is only present in the gas phase;
■ the vapor phase is ideal, despite high pressures;
■ the following solvation equation is taken into account in the liquid phase [15]:
mH2O + HI ^ [(mH2O, H +) +1- ] with m = 5;
■ HI decomposition 2HI ^ H2 +12 takes place in the gas phase only.
The NRTL activity coefficient model is used to take into account the non ideality of the liquid phase and the binary interaction parameters (including solvent-complexe) have been estimated by Neumann [15] from the experimental data of Engels and Knoche [17] for the H2O-HI-I2-H2 system. The entire thermodynamical model includes solvation equilibrium, mass balances, HI decomposition equilibrium in vapor phase and vapor-liquid phase equilibria. With such model, the phase diagrams can be calculated with confidence on the
The pump P rises the pressure at 50 bars and the valve V brings it back to 2 bars. The results of the heat and mass balance calculation are the following:
Inlet: T = 393 K; molar flow rate 20 moles/s (H2O/HI/I2 = 0.51/0.1/0.39) pressure 2 bars.
Pump P: electrical power 5.3 (mechanical efficiency 0.75).
Heat exchanger E1 exit temperature: 575.4 K (bubble temperature of the mixture); heat input 482 kW.
The column has 25 stages, feed is at stage 18 ( stage No. 1 is the upper stage). The condenser Ec is at 548 K and the heat output is 82 kW. The boiler is at 580 K and the heat input is 111kW.
The molar flow rate at the outlet of the condenser is 1 mole/s and the composition is: H2O /HI/ I2/H2 = 0.84/0.0186/0.00003/0.14.
The molar flow rate at the outlet of the boiler is 19 mole/s and the composition is H2O /HI/ I2 = 9.36/1.70/7.94.
Heat exchanger E2 cools the residue at the same temperature as the condenser: 548 K. Heat output: 76 kW.
The liquid vapor separator S condenses part of the water and HI of the distillate. This additional condensation occurs because the composition of the distillate and the residue are different. Heat output: 18 kW.
Heat exchanger E3 cools down the residue to 393 K to be fed in the Bunsen reactor. Heat output: 402 kW.
Heat exchanger E4 cools down the distillate to 393 K. Heat output: 11 kW.
The compositions and temperatures at the outlets are:
exit 1: T = 393 K, H2O /HI/ I2/H2 = 0.19/0.008/0.027/0.141; exit 2: T = 393 K, H2O /HI/ I2 = 10/1.71/7.91.
left-hand side of the binary azeotropic H2O-HI or for low iodine content. Uncertainties remain for high iodine contents (>20 %) and high temperatures ((>270 °C).
The main features of this model, have been published in [4]. We recall here the behavior of the ternary mixture at 22 bars, which is the pressure chosen by Roth and Knoche [3] for their reactive distillation calculation.
Ternary system HI-I2-H2O. A fundamental aspect for the reactive distillation design is the influence of I2 concentration on the maximum bubble temperature of the ternary system H2O-HI-I2. The calculated results are presented on fig. 6. Each curve corresponds to the bubble curve of a ternary system H2O-HI-I2 with constant I2 mole fraction varying in the interval: 1 (pure I2), 0 (binary H2O-HI). Thus the last curve on the right is the bubble curve of the binary H2O-HI with a maximum azeotrope at XHI = 0.133. The crest line, locus of the maximum bubble points, joins the pure I2 point (T = 359.17 °C) to the binary azeotrope H2O-HI (T = 245.13 °C).
Fig. 7 shows the ternary diagram of the system H2O-HI-I2 at 22 bar, limited to the region of interest for the reactive distillation design. We have notably reported:
■ the singular points:
— pure I2 (origin) and pure H2O;
— the binary azeotrope H2O-HI;
— the binary heteroazeotrope H2O-I2.
■ the distillation frontier that is the projection of the crest showed on fig. 7. Note that a linear approximation of this frontier isn't correct.
■ the liquid-liquid — vapor domain for low HI content.
This diagram gives information on feasible distillation paths at 22 bar. Thus, as mentioned by Roth and Knoche [3], from the feed composition (HI/H2O/I2 = 0.1/0.51/0.39), we could obtain pure iodine as residue and, by dissociation of HI, a distillate with H2, the feed point being just on the right of the distillation frontier. The straight line in fig. 6 represents the mass balance and gives the composition of the distillate we could obtain without chemical reaction (HI/H2O = 0.17/0.83). As a matter of fact we see that the number of water moles to evaporate is significant. Finally, we present in fig. 7 the liquid profile obtained by simulation of a reactive distillation column, with HI dissociation in the vapor phase. This profile follows very closely the distillation frontier. These preliminary simulation and thermodynamics results give the basis to the optimal design of the reactive distillation process [17].
Reactive distillation of HIX. The above Neumann's model has been implemented in the chemical engineering code PROSIM and we proceeded by trial and error to converge on the following flowsheet.
The heat released by E2, E3 and S can be used for E1. The amount of heat recovered by E1 de-
pends on the temperature difference assumed between the primary and the secondary circuit. For example, if the temperature difference is 10 °C, then the heat can be recovered in E1 from 570 K to 393 K i. e.: 482(570 - 393)/(575 - 393) = 469 kW. The amount of heat available in E2 and E3 is 76 + 401 = 478 kW which is enough. Consequently, ^ it remains 482 - 469 =13 kW to supply at a tempe- * rature comprised between 570-575 K. In a simi- | lar way, if the temperature difference is 20 °C, ^ the heat to supply becomes 40 kW to be supplied J between 560-575 K.
Due to the fact that the temperature differ- | ence between the condenser and the boiler is small s (32.5 °C), a heat pump can be used to raise the g temperature of the heat released at the condenser q and to use it for the boiler. We selected water as the working fluid and tabulated the performances of the heat pump as follows.
For one mole of water circulating in the heat pump, the heat request at its boiler (which is the condenser of the column) is qb = -0.0868Te + 52.44, where Te (°C) is the inlet temperature. The electric power Pc for an ideal compressor is Pc = = (-0.000283T + 0.131)AT, where AT is the difference temperature between boiler and condenser. Finally the heat recovered at its condenser (which is the boiler of the column) is qc = = (0.09325 - 0.000524Te)AT + 52.9 - 0.0833T.
If T and T are the condenser and boiler
c b
temperatures of the column, we assumed a temperature difference of 5 °C between primary and secondary circuits. Therefore, Te = Tc - 5 and AT = T - T + 10.
bc
The separator S being at the same temperature as the condenser, we applied the previous correlations to sum of the heat available at the condenser (82 kW) and at the separator S (18 kW). The application is straightforward: Te = 270 °C, AT = 580 -- 548 + 10 = 42 °C, qb = 29 kW, Pc = 2.29 kW, qc = 28.4 kW. Therefore if 100/29 = 3.45 moles/s circulate in the heat pump they can transfer 28.4 • 3.45 = 98 kW to the boiler of the column. Then it remains 111-98 = 13 kW to supply as heat. The work of the ideal compressor is 2.29 • 3.45 = 7.9 kW. If we assume a mechanical efficiency for the compressor of 0.75, the corresponding heat demand is 7.9/0.5/0.75 = 21 kW. The total heat demand for one mole of H2 produced is finally: (21 + 13 + 13 + 5.3/0.5)/0.141 = 408 kW if the temperature difference is 10 °C and (21 + 13 + 40 + 5.3/0.5)/0.141 = 600 kW if the temperature difference is 20 °C. In both cases the heat content of the gaseous stream at the outlet of the separator S is lost. It amounts 11 kW.
It can be noticed that in this flowsheet, the number of moles/s at the entry is exactly that at the exit of the Bunsen reaction. The H2 production being only 0.141 moles/s, it means that an important quantity of materials must be recirculated between sections I and III. Typically,
20/0.141 = 142 moles/s must enter section III. The head losses of such a mass flow rate have not been taken into account.
The best estimate of the efficiency with today's knowledge and a temperature difference of 10 °C for the heat recovery between E1 and E2, ^ E3 is then 286/(408 + 420) = 0.35. If the tempe* rature difference is 20 °C, the efficiency becomes J 286/(600+ 420) = 0.28.
■t; Conclusion - R&D needs. With the current
u
| reactions (1) to (7), widely adopted for the iodine
^ sulfur cycle, we found an upper bound of the
1 efficiency of 0.48. This value has been found by
au
& calculating the heat and work requirements of the g reactions, assuming that they were reversible. © A best estimate could be calculated as well by taking the latest GA value for the heat request of section II together with a reactive distillation calculation of section III. We found a value comprised between 0.28 and 0.34, depending on the temperature difference assumed for the heat recovery between E2 + E3 and E1. These values do not take into account the head losses due to the large material circulation.
Many values of the heat requirement of section II are found in the literature. The problem for this section is to design the flowsheet in order to optimize the production cost. Compromises have to be found between the complexity and sizes of the chemical reactors and the efficiency. The main difficulty for section III is the calculation of the reactive distillation column due to the complexity of the thermodynamic model. Anyway, some breakthrough is needed to improve its efficiency. The Japanese are possibly on the right way with the electrodialysis.
Many improvements can be imagined to rise these both figures.
Reduction of the amounts of iodine and water in the Bunsen reaction. We have seen that these excesses of iodine and water burden heavily the heat balances of both sections II and III.
The use of membranes at various steps of the cycle : to separate SO2 from O2, to separate liquid H2O from H2SO4, to separate HI from HIX and finally to separate H2 from (HI + H2O)(g). Many teams, in particular in Japan, currently work actively on these topics.
Finally the HIX section still represents a real chemical engineering challenge. Additional ther-modynamic measurements are necessary and the best distillation design has to be found. Further experiments are planned, to be performed by the Commissariat a l'Energie Atomique (CEA), in order to improve the database of the ternary mixture, more specifically at high iodine contents and to measure the vapor partial pressures as well.
References
1. Funk J. E. Thermochemical hydrogen production: past and present // Int. J. of Hydrogen Energy. 2001. Vol. 26. P. 185-190.
2. Norman J. H., Besenbruch G. E., Brown L. C. et al. Thermochemical water splitting cycle. Bench scale investigations and process engineering -DOE/ET/26225, 1981.
3. Roth M., Knoche K. F. Thermochemical water splitting through direct HI decomposition from H2O-HI-I2-solutions // Int. J. Hydrogen Energy. 1989. Vol. 14, No. 8. P. 545-549.
4. Goldstein S., Borgard J. M., Joulia X., Guittard P. Thermodynamic aspects of the iodine-sulphur thermochemical cycle // AIChE Spring National meeting, New Orleans, April 2003.
5. Öztürk I. T. et al. // Trans. IChE. Vol. 72. Part A. March 1994.
6. Knoche K. F., Schepers H., Hesselmann K. WHEC 1984. [4] Engels H., Knoche K. F. Vapor pressures of the system HI-H2O-I2 and H2 // Int. J. Hydrogen Energy. 1986. Vol. 11, No. 11. P. 703-707.
7. Barin I., Knache O. Thermodynamic properties of inorganic substances - Springer Verlag ed.
8. Calabrese V. Polyiodine and polyiodide Species in an aqueous solution of iodine + KI: Theoretical and experimental studies // J. Phys. Chem. A. 2000. Vol. 104. P. 1287-1292.
9. Brown L. C. at al. High efficiency generation of hydrogen fuels using nuclear power. GA-A24285, June 2003.
10. Roth M., Knoche K. F. // Int J. of Hydrogen Energy. 1989. Vol. 14, No. 8. P. 545-549.
11. Kracek F. C. Solubilities in the system water-iodine to 200 °C // J. Phys. Chem. 1931. Vol. 35. P. 417.
12. Wüster G. Dissertation, RWTH, Aachen, 1979.
13. Vanderzee C., Gier L. The enthalpy of solution of gazeous hydrogen iodide in water, and the relative apparent molar enthalpies of hydri-odic acid // J. Chem. Thermodyn. 1974. Vol. 6. P.441.
14. Berndhauser C., Knoche K. F. Experimental investigations of thermal HI decomposition from H2O-HI-I2-solutions // Int. J. Hydrogen Energy. 1994. Vol. 19, No. 3. P. 239-244.
15. Neumann, Diplomaufgabe, RWTH Aachen, 28.01.1987.
16. Engels H. Phase equilibria and phase diagrams of electrolytes // Chem. Data Series, Dechema ed.
17. Engels H., Knoche K. F. // Int. J. Hydrogen Energy. 1986. Vol. 11, No. 11. P. 703-707.
