Sulfur–Iodine plant for large scale hydrogen production by nuclear power

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: [email protected] (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.

mailto:[email protected]

http://www.elsevier.com/locate/he

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.


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


S_I plant

heat duty



S_I plant

electric dutya

129.9

a Without pressure drops.


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.


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.


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.


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.


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.


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


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.


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.


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


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)


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.


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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Sulfur–Iodine plant for large scale hydrogen production by nuclear power