sulfur–iodine plant for large scale hydrogen production by nuclear power
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i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 3 5 ( 2 0 1 0 ) 4 0 0 2 – 4 0 1 4
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Sulfur–Iodine plant for large scale hydrogen production bynuclear power
Giovanni Cerri a,*, Coriolano Salvini a, Claudio Corgnale a, Ambra Giovannelli a,Daniel De Lorenzo Manzano b, Alfredo Orden Martinez b, Alain Le Duigou c,Jean-Marc Borgard c, Christine Mansilla c
a Department of Mechanical and Industrial Engineering, Universita degli Studi Roma Tre, Via Vasca Navale 79, 00146 Rome, Italyb Empresarios Agrupados Internacional, S.A., Magallanes 3, 28015 Madrid, Spainc Department of Physico Chemistry – Commissariat a l’Energie Atomique/Saclay – 91191 Gif-Sur-Yvette Cedex – France
a r t i c l e i n f o
Article history:
Received 6 May 2008
Received in revised form
11 January 2010
Accepted 15 January 2010
Available online 7 March 2010
Keywords:
HYTHEC
Hydrogen production
Thermochemical
Sulfur–Iodine cycle
Nuclear energy
Costs
* Corresponding author. Tel.: þ39 6 57333251E-mail address: cerri@uniroma3.it (G. Cer
0360-3199/$ – see front matter ª 2010 Profesdoi:10.1016/j.ijhydene.2010.01.066
a b s t r a c t
The Sulfur–Iodine (S_I) cycle, driven by nuclear power, seems to be one of the main
candidates to produce hydrogen on a large scale. A new S_I process flowsheet is proposed,
set up at CEA and simulated by ProSim code and, based on that, data and results on the
coupling of the thermochemical plant with a Very High Temperature Nuclear Reactor
(VHTR) are presented. The scale up to industrial level, the conceptual design and cost
estimation of the plant are then presented and discussed. In order to support a high
temperature aggressive environment, well established chemical engineering methods as
well as non traditional materials, devices and technologies have been selected. The
influence of the adopted technology on the H2 cost has also been investigated and is widely
discussed, comparing two different cases. An economic sensitivity analysis carried out by
varying the hydrogen production level is presented, showing that an optimum H2
production exists and, due to relevant heat recovery processes, the minimum cost is not
achieved for the maximum allowable H2 production rate. Finally an optimized layout for
the minimum cost plant, set up adopting the pinch technique, is presented leading to
a further reduction of H2 production costs.
ª 2010 Professor T. Nejat Veziroglu. Published by Elsevier Ltd. All rights reserved.
1. Introduction sources. Thus it becomes more and more important with time
Today, hydrogen is used for various purposes and is mainly
produced from fossil resources. In the long term, given the
prospect of an increasing energy demand (þ20% by 2020,
expected to double by 2030, with a possible threefold increase
by 2050), a lack of fossil resources and limitations on the
release of greenhouse gases, renewable and nuclear energy
sources will play a more and more relevant role [1–3]. H2 as an
intermediate artificial energy carrier can solve storage prob-
lems related to the use of nuclear and renewable energy
; fax: þ39 6 57333252.ri).sor T. Nejat Veziroglu. Pu
depletion of fossil energy resources, which intrinsically store
sun energy as chemical energy.
However, H2, being a very reactive element, is usually
present on the Earth only combined with other atoms, to form
molecules of different type. Water contains H2 at 11 wt% and
can represent a raw material which can be split, producing H2
and O2 (as byproduct), by supplying external power. Energy
stored in the hydrogen (fuel) can be released by an oxidation
(combustion) process that produces water again. By this
approach hydrogen will play a relevant role both in storage of
blished by Elsevier Ltd. All rights reserved.
Fig. 1 – The Sulfur–Iodine (S_I) cycle.
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 3 5 ( 2 0 1 0 ) 4 0 0 2 – 4 0 1 4 4003
energy and in power production as an artificial energy carrier.
The two processes that have the greatest likelihood of
successful massive hydrogen production from water are
electrolysis and thermochemical cycles [4–6]. The latter are
processes where water is decomposed into hydrogen and
oxygen via chemical reactions, using intermediate
compounds, which are recycled, and supplying directly
external power, that can be provided by either nuclear or solar
plant. Among the various thermochemical processes, the S_I
cycle, which is extensively being investigated all around the
world, represents one of the most promising approaches to
produce hydrogen by water splitting on a large scale. In
particular, to demonstrate the techno-economic feasibility of
the process matched to a nuclear plant, the scale up analysis,
plant conceptual design and economic assessment of the
process play a key role. In the United States, under the NERI
Program, various activities have been carried out to evaluate
the performance and economic potential of the S_I plant
driven by nuclear power for large scale hydrogen production
[5,7,8]. These analyses have been performed on the basis of
flowsheets developed by General Atomics Company, which
first proposed and studied the S_I cycle starting from the 70s
[9]. In Europe, the EU funded HYTHEC project aimed at
investigating the effective potential of the S_I cycle (along
with the Hybrid Sulfur cycle) for massive hydrogen production
[10,11]. University Roma Tre was tasked within this project
with performing techno-economic analyses for nuclear driven
S_I plant, based upon new chemical flowsheets developed by
Commissariat a l’Energie Atomique (CEA). After a first techno-
economic assessment for a baseline H2 production plant,
different sensitivity analyses have been carried out investi-
gating the influence of selected degrees of freedom (DOF’s) on
the capital investment and nuclear power costs. In particular,
differently from other works, the sensitivity of hydrogen cost
to the production level was evaluated without adopting cost
scaling rules but re-designing ex novo the plant for each
production rate. Moreover different optimized plant sche-
matic layouts were set up, with varying the H2 production
level, to see the effect on the overall cost. The sensitivity of the
H2 cost to selected heat transfer technologies was also
investigated, highlighting new and important results for the
S_I plant. The purpose of the present paper is to show and
discuss the results obtained from such activities, carried out
within HYTHEC. The influence of material choices on the
investment cost of selected components of the thermo-
chemical plant has been analyzed by CEA, outside the
HYTHEC project. This analysis results can be found else-
where [12].
2. S_I Thermochemical cycle background
The Sulfur–Iodine (S_I) cycle which is schematically sketched
in Fig. 1, was extensively studied by General Atomics
Company [9]. Japan has built a small pilot plant of this
process [13]. Accordingly, the S_I cycle seems to be one of the
best known, internationally leading candidates, as very
promising thermochemical options. The major chemical
reaction stages involved in the S_I cycle may be summarized
as follows:
R1 – 9I2þ SO2þ 16H2O / (2HIþ 10H2Oþ 8I2)þ (H2SO4þ 4H2O)
(1)
[120 �C] which represents the hydrogen iodide and sulfuric
acid production step;
R2 – 2 HI / H2þ I2 (2)
[220–330 �C] where the hydrogen iodide is split into H2
and I2;
R3 – H2SO4 / SO2þH2Oþ 1⁄2 O2 (3)
[850 �C] which represents the H2SO4 decomposition step.
Globally, the sum gives the decomposition of water into
hydrogen and oxygen:
R4. H2O / H2þ 1⁄2 O2 (4)
Fig. 2 – H2SO4 section flowsheet.
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 3 5 ( 2 0 1 0 ) 4 0 0 2 – 4 0 1 44004
The first reaction, called the Bunsen-reaction, proceeds
exothermically in the liquidphase and producestwoimmiscible
aqueous acid phases whose compositions are aqueous sulfuric
acid (light phase) and a mixture of hydrogen iodide, iodine and
water named HIx (heavy phase). These acids are concentrated
and the excesses of water and iodine are recycled in R1. HI and
H2SO4 are then decomposed according to reactions R2 and R3.
Only high temperature heat sources, such as nuclear VHTR or
solar energy concentrators, may be relevant for the reaction R3.
Reaction R2 is the HI decomposition step with a little endo-
thermic heat of reaction. Reaction R3 is the major endothermic
reaction releasing water, oxygen and sulfur dioxide. It takes
place in the vapor phase in a catalytic reactor at about 900 �C
(1173 K). The concentration by distillation of the two acids HI
and H2SO4 involves significant heat consumption, which has
a direct influence on the H2 costs. Among all options available
for the HIx section (e.g. extractive distillation using phosphoric
acid, electro-dialysis), the HYTHEC project has focused on the
reactive distillation concept as proposed by Knoche et al. [14].
Their approachallowsthissteptobedoneinasinglereactorso it
seems to have the highest efficiency potential. An improved
version has already been proposed in [15].
A new flowsheet has been set up within the HYTHEC
project, with the aim of achieving the minimum external heat
requirement with the highest possible easy in the arrange-
ment of the various sections. This flowsheet shall be pre-
sented and discussed in the following sections and will
represent the baseline concept to perform the scale up anal-
ysis, the equipment design and the economic evaluations.
3. The plant design
3.1. The adopted approach
A plant design can be carried out when the values of a set of
quantities (DOF’s) are established. All the other quantities
(unknowns) are obtained by solving the plant model, which
can be split into modules, each concerning a single compo-
nent or a group of components. In particular, in order to select
suitable plant design solutions for the specific case, the
following main aspects need to be taken into consideration:
� definition of the process in terms of required unit operations
and related process quantities (pressures, temperatures,
compositions, heat and work requirements, etc);
� coupling between the nuclear thermal source and the S_I
process;
� set up of the plant layout in order to find the most effective
arrangement to perform the required operations;
� selection of appropriate technology and materials for each
plant component;
� design and cost accounting of components;
� estimation of H2 production specific costs.
Conceptually, a plant model based upon all the above
aspects simultaneously can be established and the design
problem solved by adopting a global optimization procedure
aimed at the evaluation of DOF’s which minimize the objec-
tive function (in this case, hydrogen production costs). In
practice such an approach is really challenging due to the
complexity of the actual system and the great number of
DOF’s of different type involved. To simplify the problem, the
different aspects have been considered sequentially to carry
out the present analysis, adopting the following approach:
1) the thermochemical cycle chemical flowsheet and the
nuclear source coupling scheme have been established ‘‘ab
initio’’ and assumed as fixed constraints of the overall
problem; 2) selected DOF’s (e.g. materials, H2SO4 technologies,
etc.) have preliminarily been assigned and chosen based upon
available literature, knowledge, state of the art, etc; 3) the
design and economic evaluation of the S_I plant have been
carried out starting from a first plant, assumed as the baseline,
and, then, based upon results obtained, analyzing the influ-
ence of selected DOF’s (i.e. the H2 production rate level, the
heat transfer technologies adopted in the HIx section and the
Fig. 3 – HIx section flowsheet.
Table 1 – S_I cycle energy balances.
Duty (kJ/mol H2)
H2SO4 section heat duty 422
HIx section
heat duty
HIx Pinch¼ 10 �C 59
HIx Pinch¼ 20 �C 140
S_I plant
heat duty
HIx Pinch¼ 10 �C 481
HIx Pinch¼ 20 �C 562
S_I plant
electric dutya
129.9
a Without pressure drops.
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 3 5 ( 2 0 1 0 ) 4 0 0 2 – 4 0 1 4 4005
hydrogen iodide section internal heat recovery arrangement)
on the H2 cost, which represents the objective function to be
minimized.
3.2. S_I thermochemical cycle flowsheet
A new H2SO4 section flowsheet has been set up and shown in
Fig. 2 and modeled using ProSim chemical process simulator.
A sulfuric acid mixture at 57 wt% coming from the Bunsen
section enters (stream S101) at pressure of 0.08 bar and
temperature of 117 �C (390 K). To reach suitable conditions for
the decomposition, H2SO4 is concentrated up to 87 wt% by
a three flash distillation process (Flash 1, Flash2 and Flash3 in
Fig. 2) at different pressures (0.08 bar, 5.7 bar and 7 bar) up to
approximately 300 �C (573 K). The H2SO4 is then vaporized in
the E2 component at temperature of 412 �C (685 K) and pres-
sure of 7 bar, is dehydrated (R1) at temperature of 602 �C
(875 K) and pressure of 7 bar, with production of SO3 and H2O,
and then superheated (E3) up to 767 �C (1040 K). Finally it
enters the high temperature reactor R2 to be decomposed into
SO2, O2 and H2O (approximately 55% of SO3 is reacted) up to
temperatures of about 827 �C (1100 K). The high temperature
product exiting the R2 is cooled (S114–S116) down to temper-
ature of 300 �C (573 K) with recycling of the un-decomposed
SO3. Thus, the high temperature heat, coming from steam
condensation and SO3 and H2O recombination into H2SO4, can
be internally recovered. Water, SO2 and O2 are extracted from
this section at 5.7 bar and enter the Bunsen section for the
removal of oxygen as byproduct of the thermochemical cycle.
To model the H2SO4 mixtures, the Engels and Bosen
approach has been adopted [16]. H2SO4 and H2O are bound
together in complexes and are modeled by the NRTL ther-
modynamic model. Further details on the chemical model
adopted are reported elsewhere [17,10].
The H2SO4 concentration process requires a duty of 242 kJ/
mol H2, supplied by internal heat recovery. Approximately
150 kJ/mol H2 are provided by the cooling of the product
stream of the SO3 decomposition reactor (S114–S116), at high
temperature. The remainder comes from the cooling of the
mixture feeding the Bunsen section (E4 in Fig. 2) at low
temperatures (approximately 130 �C, 403 K). The high
temperature H2SO4 subsection (streams S110–S111–S112–
S113–S114) duty is 539 kJ/mol H2 and is supplied for approxi-
mately 78% by the external source (422 kJ/mol H2), while the
remainder is internally recovered by the cooling of the prod-
ucts of the R2 reactor.
The HIx section is handled by a reactive distillation process
as proposed in [15]. The system has been modeled using Pro-
Sim chemical process simulator. A mixture of HI in excess of
water (5.1 mol H2O/mol HI) and I2 (3.9 mol I2/mol HI) comes
from the Bunsen section at 2 bar and 120 �C (393 K). To operate
the reactive distillation process the mixture entering the
process (S201 in Fig. 3) needs pumping up to 50 bar and
heating up to approximately 300 �C (573 K). The decomposi-
tion of HI into H2 and I2 takes place along with the distillation
process in a single column (HC). 25 stages are needed to reach
an H2 production at 10 mol%, with a molar reflux ratio of 2.2,
a 100% vapor distillate mixture and a bottom liquid
Fig. 4 – S_I Cycle coupling to a Nuclear Reactor.
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 3 5 ( 2 0 1 0 ) 4 0 0 2 – 4 0 1 44006
concentration of 49.55 mol/mol H2. The flashes HF1 and HF2
(Fig. 3) are needed to recover the steam condensing heat as
well as to purify the H2 produced inside the column. They
operate at pressure of 50 bar and temperatures of 276 �C
(549 K) and 120 �C (393 K) respectively. H2 is extracted from the
section (S212), while I2 and the un-decomposed HI are recycled
into the Bunsen section at temperature of 120 �C (393 K) and
pressure of 50 bar (streams S209 and S211).
Neumann’s thermodynamic model has been adopted to
model the H2O–HI–I2–H2 reactive system [18]. The NRTL
activity coefficient model has been used to take into account
the non ideality of the mixture and binary interaction
parameters have been estimated from data available in [19] for
the mixture. More details about the adopted model can be
found in [17,10].
To enhance the heat recovery a steam heat pump has been
adopted to transfer the heat available from the column
condenser (276 �C, 549 K) to the column reboiler (311 �C,
584 K). Steam pressure for the heat pump water vaporizer is
55 bar and 106 bar for the steam condenser. The heat pump
electric requirement is 113.8 kJ/mol H2.
The overall HIx section thermal duty, needed to heat the
mixture feeding the HIx section up to 300 �C (573 K), is 1460 kJ/
mol H2. Adopting a temperature pinch of 10 �C in the HIx
section, approximately 93% of the total duty can be recovered
internally by the cooling of the column products and the
remainder (approximately 59 kJ/mol H2) needs supplying
externally. When a temperature pinch of 20 �C is assumed, the
external heat duty is about 140 kJ/mol H2, representing
approximately 13% of the total thermal duty, with the
remainder (87% of the total duty) to be internally recovered.
The S_I cycle energy balance is summarized in Table 1,
considering the two different pinches (10 �C and 20 �C) for the
HIx section. For the first case the HIx section thermal
consumption represents approximately 12% of the overall S_I
plant thermal duty, while it becomes about 29% of the overall
thermal duty when a pinch of 20 �C is assumed. An electric
power of almost 130 kJ/mol H2 is required by the current S_I
process, without accounting for pressure drops inside the
equipment, which will be considered separately in the design
of the plant.
The Bunsen section flowsheet has been adopted from the
General Atomics report [5].
3.3. S_I cycle and nuclear source coupling
The coupling of the S_I cycle to a Nuclear Reactor has been
studied as part of the work performed inside the HYTHEC
project [10,20]. Due to the high temperatures needed for the
H2SO4 decomposition, the best connection option is with
a VHTR. Fig. 4 shows a connection scheme as an example
between HYTHEC and the European Project RAPHAEL. This
scheme represents a self-sustainable plant concept, in which,
in addition to the heat supply to the S_I cycle, the electrical
demand of the internal consumers is provided by the same
nuclear reactor. The high temperature heat from the reactor is
recovered in an Intermediate Heat eXchanger (IHX), which
provides heat to a secondary loop that interacts with the S_I
cycle components, improving thermal recovery. The heat goes
partially to the S_I cycle and partially to a Brayton helium
cycle for an electricity production that equals the consump-
tion of the thermochemical cycle and of the overall system
auxiliaries.
The adopted scheme is flexible allowing the electric power
rate and the H2 production rate to be changed as a function of
helium by-passed to the gas turbine or to the S_I cycle.
A single 600 MWth VHTR has been selected as the baseline
reactor concept to deliver the needed thermal (and electric)
Fig. 5 – Simplified H2SO4 section plant layout. The external thermal source feeds R1, E3, R2 and, partially, E2.
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 3 5 ( 2 0 1 0 ) 4 0 0 2 – 4 0 1 4 4007
power to the plant. On the basis of simulations carried out by
Empresarios Agrupados a pinch temperature of 37 �C has been
assumed for the IHX, with the primary helium flow cooled
from 950 �C (1223 K) to 387 �C (660 K), so as to enter the VHTR
at 400 �C (673 K), and the secondary helium flow heated from
350 �C (623 K) to 890 �C (1163 K).
The secondary helium feeds the Brayton power cycle at
a pressure of 50 bar with inlet temperature of 890 �C (1163 K)
and outlet temperature of 350 �C (623 K). It flows through
a high pressure gas turbine, which provides the required
power to the high pressure compressor, then flows through
a low pressure turbine that moves the low pressure
compressor and finally it expands in the power turbine, to
generate electrical power. The thermal-electric efficiency of
the power plant has been estimated by Empresarios Agrupa-
dos in 48%.
The secondary loop helium also delivers the heat needed to
the S_I process endothermic sections (H2SO4 section at high
temperatures and HIx section at low temperatures) from
890 �C (1163 K) down to 442 �C (715 K). In particular 422 kJ/mol
H2 are required by the H2SO4 section to vaporize and decom-
pose the acid, while the heat supplied to the HIx section
depends on the H2 production level, as discussed in the other
sections.
On the basis of information exchanged with the RAPHAEL
project [20] and evaluations carried out by Empresarios
Agrupados [20], a cooler has been located at the exit of the S_I
plant cooling down the helium flow from 442 �C (715 K) to
329 �C (602 K).1 By this approach helium flow reaches suitable
conditions to feed the secondary loop compressor, which
increases helium temperature up to 350 �C (623 K) and pres-
sure up to 50 bar.
The primary He circuit pressure drop has been estimated
by Empresarios Agrupados, as part of the HYTHEC project
work, equal to approximately 2 bar while the secondary He
circuit pressure drop has been estimated in approximately
3 bar. Helium recirculation in the primary circuit requires
1 The cooler also represents a passive safety system in case ofthermal turbulence induced by partial losses occurred in thethermochemical plant [20].
a specific power of approximately 67 kW/kg He, while for the
secondary loop a specific compression power of about 105 kW/
kg He is required.
3.4. The baseline schematic plant layout, technologiesand materials
The H2SO4 section schematic layout (Fig. 5) has been estab-
lished according to the corresponding flowsheet shown in
Fig. 2. The first two flash units (Flash1 and Flash2) have been
arranged as kettle pool boilers, while the third flash (Flash3) as
an adiabatic reactive vapour–liquid separator tank. As shown
in Fig. 2, the H2SO4 vaporizer (E2) is split in two parallel units,
with the needed heat supplied by the product stream of the
reactor R2 and by helium respectively. The major part of
H2SO4 is then dehydrated in the R1 reactor. The remainder is
decomposed and heated in the reactive heat exchanger E3.
H2SO4 is practically all decomposed into SO3 before entering
R2. The reactor R1, as well as the heat exchanger E3, has been
designed as shell and tube heat exchangers, as the dehydra-
tion reaction has reasonably been assumed instantaneous,
without the need for catalysts. The SO3 decomposer has been
arranged as a tubular reactor with reactants flowing inside the
tubes internally coated with catalysts. Both tubes and shell of
this reactor have been assumed made of Incoloy. Ceramic
materials (SiC) have been selected as the constitutive heat
exchanger tube materials, since SiC tubes show an excellent
thermal conductivity and seem to have a good behaviour also
in highly aggressive environments. Tanks, as well as heat
exchanger shells, have been assumed made of CS, internally
covered with acid brick liners (ABL) that act as corrosion
resistant materials.
The HIx section schematic layout has been set up accord-
ing to the correspondent HIx flowsheet and is depicted in
Fig. 6. The entering flow is first heated (HE1–HE3 and HE1–HE4)
by the heat available from the products of the first flash unit
(HF1) and is further heated up by the heat from liquid product
of the distillation column (HE1–HE2). To reach suitable ther-
mochemical conditions for the distillation process, addition
heat is supplied by external thermal source (Helium) (HE1-He).
All the HIx section heat exchangers have been assumed of
Fig. 6 – Simplified HIx section plant layout. The external source partially feeds HE1.
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 3 5 ( 2 0 1 0 ) 4 0 0 2 – 4 0 1 44008
shell and tube type. HF1 has been designed as a shell and tube
condenser followed by an accumulator-separator tank while
HF2 as a vapour–liquid separator tank, with gaseous hydrogen
separation from the liquid mixture. Distillation column sieve
plates have been selected to be made of Monel, coated with
catalysts on the surface. Due to highly aggressive environ-
ment, likewise the H2SO4 section, SiC has been chosen as the
baseline material for the HIx section heat exchanger tubes,
while pressure vessels and heat exchanger shells have been
assumed made of Carbon Steel, with ABL as the internal clad.
2 hEL being assumed equal to 48%.
3.5. Sizing of components and cost evaluation of thebaseline S_I plant
Tanks (accumulator, vapour–liquid separators, etc.) have been
sized using the residence time method (also called liquid
phase surge time) [21]. Shell and tube heat exchangers have
been designed by the LMTD method according to TEMA
standards. Various issues (e.g. pressure drops, vibration and
resonance) have been considered and solved by adopting
suitable types of baffles, varying inlet and outlet nozzle sizes,
tubes diameters or tube pitches, etc. The distillation has been
sized using the nested design method and the Fair flooding
calculation approach [22]. The high temperature reactor R2,
for the SO3 decomposition into SO2, has been designed as
a tubular reactor. A 1D plug flow-based model, constituted by
23 plug flow reactors in series, has been set up and solved by
CHEMKIN code, considering kinetics reported in [23]. Pumps
and compressors have been sized according to standard
engineering methods [24].
A hydrogen production rate of 633 mol H2/s could be ach-
ieved by the HYTHEC S_I process coupled to the 600 MWth
VHTR in the self sustaining arrangement. Within the HYTHEC
project, this value was assumed as the initial baseline
production level.
The power balance of the overall system is expressed by
the following relationship:
QTH ¼ QTH;S þ QTH;HI þ QEL=hELþQTH;Cool (5)
with QTH being the nuclear power (equal to 600 MWth), QTH,S
the thermal power delivered to the H2SO4 section, QEL the
electricity needed for the S_I plant and for the He recirculation
auxiliaries, hEL (equal to 48%) the thermal-electric efficiency of
the Helium Brayton power cycle (Power Island), QTH,Cool the
thermal power rejected externally by the cooler at the exit of
the S_I plant and QTH,HI the thermal power available for the
HIx section. For the baseline H2 production level (i.e. 633 mol
H2/s), temperature vs. thermal power profiles for the S_I plant
equipment interfaced to the He secondary loop are shown in
Fig. 7. To reach the selected H2 production rate, QTH,S repre-
sents almost 45% of the VHTR power (267.7 MWth) and a pinch
of approximately 20 �C in the high temperature SO3 decom-
position reactor is achieved. The required electric power (QEL),
in the self sustaining arrangement (110 MWe) represents
approximately 38% of the total nuclear thermal power2 with
more than 75% needed for the HIx section (the heat pump
requires approximately 72 MWe) and the remainder due to the
other plant sections and to the helium recirculation. A further
11% of the VHTR thermal power (66 MWth) is rejected by the
cooler (QTH,Cool). As a consequence, 6% of the overall nuclear
power (37.3 MWth) is available to feed the HIx section (QTH,HI).
This implies that a thermal power approximately 1.4 times
the VHTR one (850 MWth) needs to be recovered within the
HIx section with a pinch temperature of about 10 �C. Conse-
quently huge areas for heat transfer devices are needed: in
particular the HE1–HE3 heat exchanger, which requires about
169,000 m2 to exchange approximately 660 MWth (with an
LMTD of about 10 �C), is the most expensive component of the
overall plant. Regarding the H2SO4 section, the most dis-
tinguishing component is the high temperature SO3 decom-
position reactor. Approximately 12% of the VHTR power
(71.5 MWth) is required to decompose SO3 into SO2, with
a heat transfer area of almost 5000 m2. The device has been
designed considering 4 shell and tube units, with each tube
length equal to approximately 9 m.
Based upon good engineering practices and constraints on
sizes of heat transfer devices, the H2SO4 concentration section
has been arranged as one line, while the high temperature
subsection (i.e. dehydration and SO3 decomposition) has been
arranged on two parallel lines. Regarding the HIx section, the
equipment has been arranged as six parallel lines, with each
heat pump equipped with two compressors.
To evaluate costs of the S_I plant equipment a factored
approach has been adopted [24–26] and 2005 year euro (V) has
been used as the currency to assess the plant costs.
The equipment base costs (also called FOB costs) have been
evaluated taking into consideration suitable factors to
account for the actual temperature, pressure, material and
specific design. The basic equipment costs have been modified
through the use of ‘‘adders’’ accounting for piping, concrete,
instruments, etc, for each component. Further additional
costs (labour and indirect costs) have been added by
Fig. 7 – Temperature vs thermal power profiles for the baseline production (633 mol H2/s) S_I cycle coupled to a 600 MWth
VHTR. The equipment interfaced to the Helium secondary loop (H2SO4 section: E2, R1, E3 and R2; HIx section: HE1) is shown.
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 3 5 ( 2 0 1 0 ) 4 0 0 2 – 4 0 1 4 4009
introducing appropriate cost factors and, finally, total equip-
ment installed costs have been estimated. To evaluate
chemical inventory cost (basically iodine in the HIx section) an
iodine price of 15 V/kg (18 $/kg) has been assumed. The total
capital investment cost has been calculated introducing
standard factors to take further indirect costs (e.g. contin-
gency, fees, auxiliary facilities, interest during construction)
into account. These calculations have been performed taking
both information available in literature [24,5] and by means of
‘‘ad hoc’’ developed databases, accounting for costs of non
traditional devices and materials (e.g. SiC, Internal Liners, etc.)
[27]. In Fig. 8 installed costs of the baseline production S_I
plant are reported, highlighting the percent contribution of
each part. The HIx section is the most expensive one, with an
installed cost of approximately 620 MV (744 M$), contributing
for 85% to total installed costs. In particular the internal
recovery heat exchangers are responsible for approximately
65% of the overall section cost. Regarding the H2SO4 section
installed cost, which has been estimated in approximately
84 MV (101 M$), this represents 12% of the total installed costs
and is strongly affected by the high temperature R2 reactor,
which accounts for more than 45% of the section cost. The
Fig. 8 – Installed costs for the baseline production (633 mol
H2/s) S_I plant, with the percent influence of each section.
total investment charge (approximately 970 MV, or 1164 M$) is
also considerably affected by the cost of chemicals, which
reaches almost 20% of the HIx section installed cost, with the
majority of the iodine within the internal recovery heat
exchangers and the distillation column.
Total operating cost has been assessed based upon an
annual operation scenario which accounts for operating
labour costs, maintenance and repairs costs, operating
supplies costs, administrative costs, nuclear power costs and
chemical losses (iodine losses). These quantities have been
estimated on the basis of traditional chemical plant values
and data available in literature for thermochemical
hydrogen production plants [5]. A nuclear power cost of
1.20 cV/kWhth (1.44 c$/kWhth) has been assumed on the
basis of evaluations carried out by CEA in conjunction with
University Roma Tre. Along with nuclear power cost, main-
tenance & repairs and taxes are the most relevant cost items,
reasonably assumed as 4% and 2% of the fixed capital
investment respectively.
Fig. 9 – Hydrogen specific cost (Cs) for different H2
production level S_I plants, with influence of Capital
investment, costs of Maintenance and Taxes (M&T) and
VHTR Power cost.
Fig. 10 – Hydrogen specific cost (Cs) comparison between
the baseline S_I plant (633 mol H2/s) and the minimum cost
S_I plant (540 mol H2/s). Fig. 11 – Hydrogen specific cost (Cs) comparison between
traditional plain tube heat exchangers (Bare tubes)-based
S_I plant and HPHE-based S_I plant.
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 3 5 ( 2 0 1 0 ) 4 0 0 2 – 4 0 1 44010
The H2 production cost (Cs) has been evaluated according to (6):
Cs ¼Sn
k¼1TOCkð1þ iaÞ�kþTCI
Snk¼1mH2ð1þ iaÞ�k
(6)
with TOC (MV/y) being the total operating cost per year k, ia
being the discount rate, TCI (MV) being the total capital
investment cost, mH2 (kg/y) being the hydrogen produced in
the kth year and n (y) representing the life (years) of the plant.
A plant availability of 90%, along with a plant lifetime of 30
years and a discount rate of 0.06 have been assumed on the
basis of evaluations carried out by CEA in conjunction with
University Roma Tre.
A specific (perkg ofH2) hydrogen cost of approximately 5.3 V/
kg H2 (6.4 $/kg H2) has been assessed for the baseline production
level plant. The specific capital investment represents approx-
imately 37% of the overall costs and is the most relevant item,
while cost due to the VHTR power influences the H2 cost at 30%.
The remaining is mainly due to maintenance and taxes.
4. Techno-economic sensitivity analysis
The previous economic assessment has highlighted that for
high production levels (close to 633 mol H2/s) the capital
Table 2 – Cost comparison (installed costs andinvestment cost) between the baseline S_I plant (633 molH2/s) and the minimum cost plant (540 mol H2/s).
Baseline plant(633 mol H2/s)
Minimum costplant (540mol H2/s)
H2SO4 section installed
costs [MV (M$)]
84 (101) 66 (79)
Bunsen section installed costs
[MV (M$)]
20 (24) 18 (22)
HIx section installed costs
[MV (M$)]
624 (749) 268 (322)
Total capital investment
[MV (M$)]
970 (1164) 470 (564)
investment represents the main contribution to the H2 cost,
which reaches values of the order of 5–6 V/kg H2 (6–7.2 $/kg H2).
In order to explore ways to decrease the influence of the capital
investment on the H2 production cost (i.e. to reduce the
hydrogen cost), a parametric analysis has been carried out by
varying the maximum H2 production level. The overall available
thermal power QTH has been kept at the initial value (600 MWth)
and the self sustaining plant concept has been assumed.
With reference to equation (5), the lower the hydrogen
produced from the plant, the lower QTH,S is, fixing the specific
H2SO4 section external duty. QEL decreases too, being the
electrical power demand roughly linearly dependent on the H2
production rate. As a result a higher amount of thermal power
is available to feed the HIx section. An augmented QTH,HI value
implies a reduction of the thermal power to be recovered and
a raise of the pinch temperature for the HIx section heat
recovery devices. This leads to reduced heat transfer areas
and, consequently, relevant cost benefits are expected to be
achieved, with a reduction of specific capital investment
costs. Conversely, lowering the production level the specific
Fig. 12 – Hydrogen specific cost (Cs) comparison between
traditional plain tube heat exchangers (Bare tubes)-based
S_I plant and HPHE-based S_I plant for the minimum cost
production (540 mol H2/s).
Fig. 13 – Heating and cooling profiles for the 540 mol H2/s HIx section new arrangement, obtained by applying the pinch
technique.
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 3 5 ( 2 0 1 0 ) 4 0 0 2 – 4 0 1 4 4011
heat consumption raises with increased nuclear power
specific costs.
Four different S_I plants, which can achieve production of
410 mol H2/s, 540 mol H2/s, 570 mol H2/s and 650 mol H2/s
respectively and matched to the same 600 MWth VHTR in the
self sustaining arrangement, have been analyzed and
compared to assess the H2 cost sensitivity to the production
level. Due to the particular type of process and accounting for
results obtained from the baseline case study, each S_I plant
has been designed ex novo, without adopting cost scaling
factors (such as the 0.6 power rule). To perform the analysis,
the design arrangements, rules and choices described above
have been adopted again. The coupling scheme between the
S_I and the VHTR plant has been adapted to deliver the needed
thermal (and electric) power to the thermochemical plant,
accounting for the different production levels.
Assuming rules and assumptions established and
described previously, the economic analysis has been carried
out for each S_I plant and the costs obtained have been
compared to the 633 mol H2/s plant ones.
In Fig. 9 the specific H2 production costs (Cs) are shown for
the selected plants, highlighting the influence of the main
items: 1) specific costs due to S_I capital investment (Capital
Investment); 2) specific costs due to lifetime operating costs,
basically maintenance and repairs and taxes (M&T); 3) specific
costs related to the VHTR thermal power (VHTR Power). The
figure shows that the minimum cost has not been achieved for
the maximum allowable H2 production level plant (corre-
sponding to the maximum efficiency). The minimum cost
(4.2 V/kg H2 or 5.0 $/kg H2) has been assessed corresponding to
a production of 540 mol H2/s, with a cost reduction of more
than 20% relative to the baseline plant. Moreover, the cost
profile for lower H2 production level plants is almost flat
(Fig. 9): the increasing specific cost due to the needed VHTR
power is balanced by the decreasing capital and lifetime
operating costs. Comparing the minimum cost plant (540 mol
H2/s) to the baseline one, two main aspects should be high-
lighted. The first one regards the reduction of approximately
20% of the themochemical plant electric consumption, due to
the lower H2 production level, with an increased thermal
power available for the HIx section. As shown in Fig. 10, the
VHTR power cost (1.9 V/kg H2 or 2.3 $/kg H2) influences the
overall H2 cost for approximately 45.8%, while this item
represents 30.4% of the overall cost for the baseline plant. The
second aspect, which is a consequence of the first one,
concerns the decrease of more than 20% of the thermal power
to be recovered inside the HIx section (about 685 MWth) with
an increased LMTD of 25 �C. Consequently, as shown in Table
2, the 540 mol H2/s plant installed costs are dramatically lower
than the baseline plant ones, especially for the HIx section
with a cost decrease of approximately 57%. However the
influence of the hydrogen iodide section, representing more
than 75% of the total installed cost, is still significant. The
effects on inventory cost are particularly relevant too, even if
a reduction of more than 50% compared to the baseline plant
is evidenced. As shown in Fig. 10, the specific capital invest-
ment cost (1.2 V/kg H2 or 1.4 $/kg H2) represents only 28.4% of
the overall H2 production cost of the 540 mol H2/s plant, while
it is more than 37% of the H2 cost for the baseline case.
Operating costs (Fig. 10) affect the overall production cost
at 25.7%, while they represent 32.5% of the baseline plant cost.
The investment cost has been demonstrated to be influ-
enced strongly by the characteristics and performance of the
HIx section heat transfer devices for both plants compared.
On the basis of the results achieved, a second parametric
analysis has been carried out, to investigate the H2 production
cost sensitivity to the adopted heat exchange technologies. To
enhance heat transfer coefficients, High Performance Heat
Exchangers (HPHE) have been introduced with plain tubes
replaced by corrugated ones. In general, by this approach, heat
transfer coefficients can achieve values two or three times
higher than traditional plain tubes ones [28], despite a higher
cost per area and higher pressure drops [28].
Four S_I plants, producing 410 mol H2/s, 540 mol H2/s,
570 mol H2/s and 633 mol H2/s respectively, have been
designed (each plant has been sized ex novo, without using
cost scaling rules) adopting HPHE’s for selected components of
the HIx section, where working temperatures and conditions
are suitable to use the HPHE technology. The rules and
arrangements previously described have been adopted to
carry out the design of each plant matched to the 600 MWth
VHTR in the self sustaining arrangement.
Fig. 14 – New HIx section simplified layout optimized for the 540 mol H2/s plant, based on results obtained by pinch
analysis.
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 3 5 ( 2 0 1 0 ) 4 0 0 2 – 4 0 1 44012
HPHE installed costs have been assessed on the basis of
data available from manufacturers and suitably integrated in
the cost database adopted by University Roma Tre within the
HYTHEC project. Lifetime costs have been evaluated consid-
ering rules and assumptions described above and, based upon
that, the economic analysis has been carried out for each
plant and costs have been compared to the traditional plain
tube heat exchangers-based plants.
In Fig. 11 specific cost values assessed for the HPHE-based
plants are reported together with those already given in Fig. 9
for traditional heat exchangers-based plants (Bare Tubes). The
minimum cost production rate basically does not vary intro-
ducing the HPHE technology (540 mol H2/s). Such a cost has
been assessed in 3.6 V/kg H2 (4.3 $/kg H2) with a reduction of
more than 14% relative to the correspondent Bare Tubes plant
and of 40% compared to the initial baseline case (633 mol H2/s
with bare tubes). Moreover, a reduction of costs along all the
production range is observed introducing the new technology,
especially at high H2 production levels where the effect of the
investment cost is dramatic. A more detailed cost comparison
between the minimum cost plants, adopting different tech-
nologies, is shown in Fig. 12. The introduction of the corru-
gated tubes-based technology doesn’t modify VHTR power
costs (1.9 V/kg H2 or 2.3 $/kg H2) with the overall pressure
drops inside the HPHE’s being approximately the same of Bare
Tubes heat exchangers.3 Consequently, for the HPHE plant,
the overall H2 cost is influenced by VHTR power for more than
53%, while it represents 45.8% of the overall specific cost for
the traditional Bare Tubes plant. Concerning the specific
investment cost, due to reduced heat transfer areas, this item
represents only 22.5% of the overall H2 specific cost of the
HPHE plant.
This is due to the fact that the investment cost resulted in
approximately 350 MV (420 M$), with a reduction of more than
25% compared to the Bare Tubes plant. The HIx section
represents again the most expensive plant section. Its
installed cost (192 MV or 230 M$), which constitutes almost
55% of the capital investment, is more than 28% lower than
the Bare Tubes-based plant correspondent value. Likewise,
a reduction of more than 50% of the inventory cost has been
assessed.
3 Compared to Bare Tubes heat exchangers HPHE’s show higherpressure drops per area. However, the reduction of the overallheat transfer area makes the overall HPHE pressure drops beapproximately the same of Bare Tubes ones.
The previous analysis has highlighted that, for the
minimum cost plants, a high external cold utility thermal
power is needed for the HIx section and the specific H2 cost is
strongly affected by the VHTR power cost. Consequently
a new optimized layout for the 540 mol H2/s plant HIx section
has been developed, using the pinch technique, with the aim
of reducing the heat pump electric consumption and
increasing the HIx section internal heat recovery. Based upon
previous results, the new layout has been set up keeping the
LMTD value higher than 10 �C for all the heat exchangers and,
in particular, assuring an LMTD approximately equal to 25 �C
for the HE1–HE3 device.4
Heating and cooling curves for the new arrangement are
reported in Fig. 13, plotting the temperatures on the ordinate
and the relative enthalpy on the abscissa. REBO profile indi-
cates the thermal power needed to the reactive distillation
column reboiler. COND refers to the thermal power available
from the column condenser and internally recovered in the
HIx section, while E2, E3, E4 and E1 refer to HE2, HE3, HE4 and
HE1 respectively. By the new arrangement the reboiler is fed
by helium thermal power for approximately 16% (about
81 MWth) of the total duty with the remainder (424.4 MWth)
provided by the heat pump. Thus a part of the column
condenser power can be recovered (COND) to heat up directly
the HIx mixture feeding the column, with the remainder
feeding the reboiler, by the heat pump. By this approach the
external cold utility duty can be reduced and the heat pump
consumption decreased.
A new schematic layout has been set up and shown in
Fig. 14. Compared to the previous scheme (Fig. 6), which was
optimized for higher production levels, the new design is
characterized by the following distinguishing arrangement.
Based on the results obtained by the pinch technique,
a parallel arrangement has been adopted for both the column
condenser, recovering the available thermal power (COND) in
the E1–E2 heat exchanger, and the reboiler, with a part of the
thermal power supplied by the heat pump and the remainder
by the external source (helium).
The plant design and the economic evaluations have been
carried out, assuming the same quantities, data and boundary
conditions of the previous calculations and adopting the same
4 Three different layouts have been set up and analyzed, withpinch temperature differences of 7 �C, 10 �C and 15 �C respec-tively. On the basis of results obtained, 7 �C has been assumed asthe baseline pinch temperature difference, since it respects theassumed constraints.
Fig. 15 – Hydrogen specific cost (Cs) comparison among
traditional plain tube heat exchangers (Bare tubes)-based
S_I plant, HPHE-based S_I plant and HPHE-based S_I plant
with new pinch-based HIx section layout for the minimum
cost production (540 mol H2/s).
Table 3 – Installed costs and investment cost comparisonamong the three 540 mol H2/s S_I plants.
540 mol H2/sBare tubes
540 molH2/s HPHE
540 mol H2/sHPHE Pinch
H2SO4 section
installed
costs [MV (M$)]
66 (79) 66 (79) 66 (79)
Bunsen section
installed
costs [MV (M$)]
18 (22) 18 (22) 18 (22)
HIx section installed
costs [MV (M$)]
268 (322) 192 (230) 177 (212)
Total installed costs
[MV (M$)]
352 (422) 276 (331) 261 (313)
Total investment
costs [MV (M$)]
470 (564) 351 (421) 329 (395)
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 3 5 ( 2 0 1 0 ) 4 0 0 2 – 4 0 1 4 4013
rules and criteria of the previous analyses. Constitutive
materials have been selected following the same criteria
described above, adopting the HPHE technology for the HIx
section heat transfer devices and bare tube heat exchangers
for the H2SO4 section, with SiC as the baseline tube material.
Both investment and lifetime costs have been assessed
following criteria and rules described previously in the paper.
In Table 3 a comparison among investment costs obtained
for the three different plants producing 540 mol H2/s is shown.
In particular, the installed cost of the HIx section of the HPHE
Pinch plant is approximately 8% lower than the HPHE plant
one. This is mainly due to a reduction of approximately 11% of
the overall heat transfer area of the reactive distillation
column reboiler, passing from approximately 18 000 m2 to
about 16 000 m2 for the new HIx layout. As a consequence, the
plant investment cost resulted in a reduction of more than 6%
compared to the HPHE plant, also due to a decreased chemical
inventory cost of almost 16%.
The new arrangement has allowed a reduction of the
electric power needed to the heat pump compression unit
(equal to approximately 53 MWe, representing 84% of the
overall plant electric requirement) of more than 20% relative
to the old arrangement-based plant. As a consequence
a decrease of VHTR power specific cost of more than 4% (from
1.89 V/kg H2 or 2.27 $/kg H2 to 1.82 V/kg H2 or 2.18 $/kg H2) has
been assessed. As shown in Fig. 15, the nuclear power cost
affects the overall H2 cost of the pinch rearranged plant for
approximately 55%, while the specific capital investment
(0.79 V/kg H2 or 0.95 $/kg H2), influences the H2 cost for 24%,
with a specific cost reduction of almost 6% compared to the
HPHE plant.
5. Summary and conclusions
A conceptual design and economic analysis for an S_I ther-
mochemical process matched to a 600 MWth VHTR plant have
been carried out as part of the work within the EU funded
HYTHEC project.
A new flowsheet has been set up at CEA for both the H2SO4
and the HIx sections aiming at achieving low external power
requirements with easy arrangements. Three flashes in series
working at increasing pressures have been adopted to realize
the sulfuric acid concentration, with the H2SO4 decomposition
operated at 7 bar. Regarding the HIx section, the reactive
distillation approach has been adopted to concentrate the
feeding HI mixture and to decompose it into H2 and I2 in
a single component, introducing a steam heat pump to
enhance the internal heat recovery.
To evaluate the potential of the nuclear driven S_I process
on a large scale, the coupling with a nuclear reactor
(600 MWth VHTR) has been studied and set up by Empresarios
Agrupados considering evaluations and results shared with
the EU RAPHAEL project and assuming the self sustaining
arrangement as the baseline concept.
A hydrogen production rate of 633 mol H2/s has been ach-
ieved by the nuclear driven S_I plant and this level was
assumed as the baseline (first) case. The thermochemical
plant has been scaled up and designed choosing adequate
(and often non traditional) solutions in terms of materials and
technologies. In particular the H2SO4 section high tempera-
ture reactor has been designed based on the shell and tube
concept with reactants flowing inside tubes internally coated
with catalysts. After an examination of materials adequate to
be adopted in the process, SiC has been selected as the
constitutive material of heat transfer device tubes for the
sulfuric acid section and the hydrogen iodide section. Shells,
tanks and columns have been assumed made of CS with
internal liners as the material capable to resist such highly
aggressive environments.
The economic analysis of the plant has been carried out
assessing the plant installed cost by building adequate data-
bases, accounting for non traditional equipment and mate-
rials. Plant lifetime costs have been evaluated in line with
information available in literature for these processes and
considering typical scenarios for chemical plants. A H2
production cost of 5.3 V/kg H2 (6.4 $/kg H2) has been assessed
for the baseline S_I plant. To reduce the cost, based upon
results achieved, a first parametric analysis has been carried
out varying the H2 production level, keeping the same nuclear
reactor thermal power with the self sustaining arrangement.
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 3 5 ( 2 0 1 0 ) 4 0 0 2 – 4 0 1 44014
The thermochemical plant has been designed ex novo for
each different hydrogen production level, without adopting
cost scaling rules. Results showed that the best economic
solution has not been achieved for the higher H2 production
level plants. Mainly due to a strong reduction of the capital
investment, the lowest cost was reached for a plant designed
to produce 540 mol H2/s. On the basis of results achieved,
a second parametric analysis has been carried out to evaluate
the influence of the adopted technologies on the H2 production
cost. In particular HPHE-based heat exchangers have been
introduced as the HIx section heat transfer devices, replacing
the traditional plain tubes-based heat exchangers. A relevant
decrease of the final cost has been observed adopting the HPHE
technology, with an important reduction of plant installed
costs and chemical inventory costs. This analysis has high-
lighted that the H2 cost is strongly affected by the choice of
technology, in particular regarding the heat transfer processes.
Finally the HIx section layout has been rearranged and
optimized for lower H2 production rates (i.e. the minimum
cost hydrogen production levels) by adopting the pinch tech-
nique, aiming at reducing the electric requirement and the
cold utility heat duty. Though a more complex arrangement,
a further cost reduction was achieved, resulting in a H2
production cost of 3.3 V/kg H2 (4 $/kg H2).
The present analysis has highlighted which key points
need to be addressed to reduce the H2 cost further. In partic-
ular future activities are going to be focused on the possibility
of using less expensive materials, adopting different tech-
nologies for heat transfer process (e.g. compact heat
exchangers) and improving flowsheet arrangements for both
the H2SO4 and HIx sections, to reduce thermal as well as
electrical power demand.
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