Thermodynamic analysis of hydrogen production via …...reforming (in this case methane reforming) by chemical-looping auto-thermal reforming, with water-gas shift and carbonation

Thermodynamic analysis of hydrogen production via Sorption

Enhanced Chemical-Looping Reforming

Pedro da Fonseca Ramos

Dissertação para obtenção do Grau de Mestre em

Engenharia Mecânica

Júri

Presidente: Professor Mário Manuel Gonçalves da Costa

Orientador: Doutor Rui Pedro da Costa Neto

Vogal: Professor João Luís Toste de Azevedo

Outubro de 2011

i

Abstract

Hydrogen production via Sorption-Enhanced Chemical-Looping reforming combines hydrocarbons

reforming (in this case methane reforming) by chemical-looping auto-thermal reforming, with water-

gas shift and carbonation reactions. Both H2 production rate and concentration depend on the chemical

equilibrium that is established in the fuel reactor and that is influenced by several parameters

(temperature, pressure, H2O/CH4 content and presence/absence of CaO).

This report focus its attention, in a first part, on understanding, through an Aspen Plus model, the

influence that those parameters have on the equilibrium composition, as well as in the thermal

optimization of the process, so that no heat demand occurs. In a second part, experiments were carried

out, in a bench-scale fluidized bed reactor, in order to demonstrate the feasibility of this process, as

well as the effects created by some of those previously mentioned parameters.

For the thermodynamically balanced system it was possible to produce a high purity H2 (> 95%) at

650°C and 5 atm, using a H2O/CH4 ratio of 2. At these conditions, the process efficiency was 77,8%

and the CO2 capture rate of 95,0%.

In the experimental part it was possible to demonstrate de advantages of mixing oxygen carrier

particles with CO2 sorbents, in order to enhance the H2 production specially at low temperatures (600

°C).

Keywords: Sorption-Enhanced; Chemical-Looping; methane reforming; simulation

ii

Resumo

A produção de hidrogénio através do mecanismo de reformação por Sorption-Enhanced Chemical-

Looping combina a reformação de hidrocarbonetos por chemical-looping com as reacções de

deslocamento de água e de carbonatação. A taxa de produção de hidrogénio e a concentração

dependem do equilíbrio químico que se estabelece no reactor do combustível e que é influenciado por

diversos parâmetros (temperatura, pressão, razãoH2O/CH4 e presença/ausência de CaO).

Este relatório foca a sua atenção, num primeira parte, através de um modelo desenvolvido no

programa Aspen Plus, no entendimento da influência que esses parâmetros têm na composição de

equilíbrio, assim como na optimização térmica, de modo a que não seja necessário fornecer calor ao

reactor externamente. Num segunda parte, experiências foram levadas a cabo num reactor de leito

fluidizado à escala laboratorial, com o intuito de demonstrar exequibilidade deste processo e os efeitos

criados pelos parâmetros mencionados acima.

Para o caso termodinamicamente integrado, foi possível produzir hidrogénio de elevado nível de

pureza (> 95 %) a 650 °C e 5 atm, usando uma razão H2O/CH4 de 2. Nestas condições, a eficiência do

processo foi 77,8 % (relativamente ao poder calorífico inferior do metano e do hidrogénio) e a taxa de

captura de CO2 de 95,0 %.

Na parte experimental foi possível demonstrar as vantagens de misturar as partículas transportadoras

de oxigénio com os sorventes de CO2, de modo a aumentar a produção a produção de hidrogénio

especialmente a temperaturas baixas (600 °C).

Palavras-Chave: Sorption-Enhanced; Chemical-Looping; reformação de metano; simulação

iii

Agradecimentos

Gostaria de deixar os meus agradecimentos a algumas pessoas que desempenharam um papel

importante ao longo do meu percurso académico que culminou com este trabalho.

Em primeiro lugar quero agradecer muito ao Professor Magnus Rydén pela orientação e o apoio dado

ao longo da realização do trabalho, demonstrando-se extremamente disponível para quaisquer

esclarecimentos ou para qualquer discussão de ideias.

Ao professor Henrik Leion, aos estudantes de doutoramento Mehdi Arjmand e Erik Jerndal e ao

estudande de mestrado Ali Hedayati pelo seu apoio durante o trabalho experimental.

Ao Dr. Rui Neto pela sua contribuição para a revisão do trabalho, tendo estado sempre disponível para

qualquer explicação sobre o melhor modo de apresentação dos diversos pontos do trabalho.

Aos meus Pais, pela educação e apoio prestado ao longo dos anos.

Aos meus Amigos, por todos os bons momentos que passamos juntos, diversão e descontracção,

adrenalina, estudo e trabalho.

iv

Contents

List of Figures .................................................................................................................................. vi

List of Tables .................................................................................................................................... ix

Notations .......................................................................................................................................... xi

1. Scope ..........................................................................................................................................1

2. Introduction ................................................................................................................................2

2.1. Hydrogen as a carbon free energy carrier .............................................................................2

2.2. Carbon Capture and Storage ................................................................................................2

2.3. Chemical-Looping Combustion ...........................................................................................5

2.4. Chemical-Looping Reforming .............................................................................................7

2.5. Steam Reforming .................................................................................................................8

2.6. Sorption-Enhanced Water-Gas Shift ....................................................................................9

3. Sorption-Enhanced Chemical-Looping Reforming..................................................................... 11

3.1. Description ........................................................................................................................ 11

3.2. Reactions and chemical equilibrium ................................................................................... 14

3.3. Process parameters ............................................................................................................ 17

3.4. Oxygen carriers ................................................................................................................. 19

3.5. Carbon Dioxide sorbents.................................................................................................... 19

3.6. Work Purpose .................................................................................................................... 20

4. Simulation methodology and parameters ................................................................................... 21

4.1. Methodology ..................................................................................................................... 21

4.2. Parameters ......................................................................................................................... 24

5. Simulation results and discussion .............................................................................................. 25

5.1. Steam and Calcium Oxide enhancing effect ....................................................................... 25

5.2. Pressure change effect ....................................................................................................... 27

5.3. Temperature variation effect .............................................................................................. 28

5.4. Thermal optimization......................................................................................................... 31

v

6. Experiments description and parameters .................................................................................... 36

7. Experiments results and discussion ............................................................................................ 41

8. Conclusions .............................................................................................................................. 47

9. Future Work .............................................................................................................................. 48

Appendix A – Experimental and cleaning procedure ......................................................................... 52

Appendix B – Experimental equipment specifications ....................................................................... 56

Appendix C – Experimental results for 5g N4MZ1400 with 10g sand................................................ 60

Appendix D – Experimental results for 5g N4MZ1400 with 10g CaO ............................................... 63

Appendix E – Experimental results for 15g N4MZ1400 with 30g sand .............................................. 66

vi

List of Figures

Figure 1 – Schematic description of Chemical-Looping Combustion ...................................................6

Figure 2 – Schematic description of Chemical-Looping Reforming .....................................................7

Figure 3 – Schematic description of Steam Reforming (without water-gas shift) ..................................8

Figure 4 – Schematic description of Sorption-Enhanced Water-Gas Shift ............................................9

Figure 5 - Equilibrium CO2 pressure with CaO/CaCO3 as a function of temperature (Harrison, D.,

2008). ............................................................................................................................................... 10

Figure 6 – Schematic description of Sorption-Enhanced Chemical-Looping ...................................... 11

Figure 7 – Model developed in Aspen Plus ........................................................................................ 22

Figure 8 – H2 dry concentration at 600°C and 1 atm in the reforming reactor for a CH4 flow of 1

kmol/s and for different conditions: ref – without H2O and CaO particles; CaO – with CaO particles

but without H2O; H2O/CH4 – with H2O at different H2O/CH4 ratios but without CaO particles; 2

H2O/CH4 + CaO – with a H2O/CH4 ratio of 2 and CaO particles. ....................................................... 26

Figure 9 – H2 dry concentration evolution with temperature with and without CaO particles at 5 atm in

the fuel reactor and for a CH4 flow of 1 kmol/s, with a NiO/CH4 ratio of 1 and with a H2O/CH4 ratio

of 2 ................................................................................................................................................... 30

Figure 10 – H2 molar flow evolution with temperature with and without CaO particles at 5 atm in the

fuel reactor and for a CH4 flow of 1 kmol/s, with a NiO/CH4 ratio of 1 and with a H2O/CH4 ratio of 2

......................................................................................................................................................... 31

Figure 11 –Composite Curves considering input streams at 600°C and almost null Heat Duties ......... 33

Figure 12 –Composite Curves considering input streams at 550°C, almost null Heat Duties for the

Calcination and Air reactors and a 650°C temperature in the fuel reactor ........................................... 35

Figure 13 – Reactor location inside the oven ..................................................................................... 36

Figure 14 – Schematic description of experimental setup................................................................... 37

Figure 15 – Laboratory workstation: a - oven with reactor inside; b – electronic equipment for cycle

switch; c – gases flow meters; d – gases piping system; e – cooling unit; f – general perspective; g –

console for flow regulator; h – steam generator; i – analyser; j – valves control. ................................ 38

vii

Figure 16 – Typical evoltion of O2 concentration at the reactor’s exit in a oxidation cycle ................. 39

Figure 17- H2 concentration for different temperatures from experiments carried out with a 5g of

N4MZ1400 particles and CaO particles or sand, for a CH4 flow of 200 ml/min with a H2O/CH4 ratio

of 1,8. ............................................................................................................................................... 42

Figure 18 – Evolution of the H2/(H2+CO2+CO) ratio for different temperatures from experiments

carried out with a 5g of N4MZ1400 particles and CaO particles or sand, for a CH4 flow of 200 ml/min

with a H2O/CH4 ratio of 1,8. ............................................................................................................. 43

Figure 19 – CH4, CO2 and CO concentration for different temperatures from experiments carried out

with a 5g of N4MZ1400 particles and CaO particles or sand, for a CH4 flow of 200 ml/min with a

H2O/CH4 ratio of 1,8. ........................................................................................................................ 44

Figure 20 – CH4, CO2 and CO concentration for different temperatures obtained from Aspen Plus

simulation at 1 atm in the fuel reactor, for a CH4 flow of 1 kmol/s, with a NiO/CH4 ratio of 1, with a

H2O/CH4 ratio of 1.8, with and without CaO particles ....................................................................... 44

Figure 21 - H2 concentration for different temperatures from experiments carried out with a 15g of

N4MZ1400 particles and sand, for a CH4 flow of 200 ml/min with a H2O/CH4 ratio of 1,8. ............... 46

Figure 22 - CH4, CO2 and CO concentration for different temperatures from experiments carried out

with a 15g of N4MZ1400 particles and sand, for a CH4 flow of 200 ml/min with a H2O/CH4 ratio of

1,8 .................................................................................................................................................... 46

Figure 23 – Experimental results for 5g N4MZ1400 with 10g sand at 600°C and 1 atm in the fuel

reactor, with a H2O/CH4 ratio of 2 ..................................................................................................... 61

Figure 24 - Experimental results for 5g N4MZ1400 with 10g sand at 650°C and 1 atm in the fuel


Figure 25– Experimental results for 5g N4MZ1400 with 10g sand at 700°C and 1 atm in the fuel




Figure 27– Experimental results for 5g N4MZ1400 with 10g CaO at 600°C and 1 atm in the fuel




viii



Figure 30 - – Experimental results for 5g N4MZ1400 with 10g CaO at 750°C and 1 atm in the fuel










ix

List of Tables

Table 1 – Species expected to be found in each material stream......................................................... 22

Table 2 – Legend of the several blocks and streams present in Figure6 .............................................. 23

Table 3 – Molar flows and H2 concentration at 600°C and 1 atm obtained in the reforming reactor for a

CH4 flow of 1 kmol/s and for different conditions: ref – without H2O and CaO particles; CaO – with

CaO particles but without H2O; H2O/CH4 – with H2O at different H2O/CH4 ratios but without CaO

particles; 2 H2O/CH4 + CaO – with a H2O/CH4 ratio of 2 and CaO particles. ..................................... 25

Table 4 - Molar flows and H2 concentrations at 600°C and different pressure values in the reforming

reactor for a CH4 flow of 1 kmol/s, with a NiO/CH4 ratio of 1 and with a H2O/CH4 ratio of 2 and CaO

particles. ........................................................................................................................................... 27

Table 5 - Molar flows and H2 concentrations at different temperatures and 5 atm in the fuel reactor for

a CH4 flow of 1 kmol/s, with a NiO/CH4 ratio of 1 and with a H2O/CH4 ratio of 2 and CaO particles.

......................................................................................................................................................... 28

Table 6 - Molar flows and H2 concentrations at different temperatures and 5 atm in the fuel reactor for

a CH4 flow of 1 kmol/s, with a NiO/CH4 ratio of 1 and with a H2O/CH4 ratio of 2 and without CaO

particles. ........................................................................................................................................... 29

Table 7 – Temperatures and Heat Duties of the three reactors considering input streams at 600°C and

almost null Heat Duties ..................................................................................................................... 32

Table 8 – Molar flows and H2 concentration at 702°C and 5 atm in the fuel reactor for a CH4 flow of 1

kmol/s, with a NiO/CH4 ratio of 1, with a H2O/CH4 ratio of 2 and with CaO particles ....................... 33

Table 9 – Temperature and Heat Duties of the three reactors considering input streams at 550°C,

almost null Heat Duties for the Calcination and Air reactors and a 650°C temperature in the fuel

reactor .............................................................................................................................................. 34

Table 10 – Molar flows and H2 concentration at 650°C and 5 atm in the fuel reactor for a CH4 flow of

1 kmol/s, with a NiO/CH4 ratio of 1, with a H2O/CH4 ratio of 2 and with CaO particles .................... 34

Table 11 – Experiment results from the 5g of N4MZ1400 particles experimental set with sand, for a

CH4 flow of 200 ml/min with a H2O/CH4 ratio of 1,8 ........................................................................ 41

Table 12 - Experiment results from the 5g of N4MZ1400 particles experimental set with CaO, for a


x

Table 13 - Experiment results from the 15g of N4MZ1400 particles experimental set with sand, for a


xi

Notations

Latin symbols

Formula

symbol

Unit Definition

U [J] Internal energy

H [J] Enthalpy

Q [J] Heat

p [atm] Pressure

V [m3] Volume

T [°C] or[K] Temperature

S [J/K] Entropy

deS [J/K] External entropy

diS [J/K] Internal entropy

G [J] Gibbs free energy

Ro [J/(kmol K)] Ideal gas universal constant

n [kmol] Number of moles

� [J/kmol] Molar Gibbs energy

Kp [-] Equilibrium constant

Ea [kJ/mol] Activation Energy

A [kmol bar1/2 kgcat

-1 h-1] or

[kmol bar-1 kgcat-1 h-1]

Pre-exponential factor

k [kmol bar1/2 kgcat

-1 h-1] or

[kmol bar-1 kgcat-1 h-1]

Rate constant

xii

Indices

Index Definition

( )o Standard conditions

( )Ref Reference conditions

( )i Chemical specie i

Abbreviations

Abbreviation Definition

CCS Carbon Capture and Storage

CLC Chemical-Looping Combustion

CLR Chemical-Looping Reforming

SE Sorption-Enhanced

WGS Water-Gas shift

SECLR Sorption-Enhanced Chemical-Looping Reforming

IPCC Intergovernmental Panel for Climate Change

PC Pulverized Coal

NGCC Natural Gas Combined Cycle

IGCC Integrated gasification Combined Cycle

MEA Monoethanolamine

N4MZ1400 40% Ni supported on Mg-ZrO2

xiii

Chemical Abbreviations

Molecular Formula Name

CnHm Hydrocarbon

CH4 Methane

H2 Hydrogen

CO2 Carbon Dioxide

CO Carbon Monoxide

H2O Water

NOx Nitrogen Oxides

NiO Nickel Oxide

Ni Nickel

CaO Calcium Oxide

CaCO3 Calcium Carbonate

Me Metal or Metal Oxide with lower oxygen content

MeO Metal Oxide

N2 Nitrogen

O2 Oxygen

C Carbon (coke)

1

1. Scope

Hydrogen production is, presently, based on hydrocarbons reforming, mainly natural gas. Most

processes use steam reforming as a way to produce H2 (Hydrogen) mixed with CO2 (Carbon

Dioxide) and CO (Carbon Monoxide). Thus, further complex gas-separation systems are needed, so

that a high purity stream of H2 can be obtained. Moreover, steam reforming is a very energy

demanding process, due to the endothermic behaviour of steam reforming reaction.

Sorption-Enhanced Chemical-Looping Reforming (SECLR) uses methane partial oxidation for

hydrogen production, so lower energy demands are required, and it has a good potential for CO2

capture, since the production of H2 and CO2 occurs in two separate streams, avoiding complex

separation systems. Nevertheless, during the process, the heat released and heat consumption occur

in different reactors and therefore, transport of sensible heat between the three reactors is required.

This work focuses its attention on the thermodynamic analysis of SECLR process, based on Aspen

Plus modelling software. The optimization of system parameters, such as temperature, pressure,

steam/carbon ratio content of input stream, type and flow of solid particles and heat exchangers

network, is essential, in order to avoid external heat demand.

2

2. Introduction

Over the last years, concerns regarding CO2 (Carbon Dioxide) emissions from fossil fuel combustion

and their impact on Earth’s climate have been growing, not only among the scientific community,

but also in the general population. Therefore, efforts for reducing these emissions led to an increase

in new technologies research, not only in new ways for converting fossil fuels that would allow CO2

capture and storage (CCS), but also in alternative fuels that could replace traditional fuels in thermal

engines and boilers, for instance, with lower contribution to greenhouse gases emissions.

2.1. Hydrogen as a carbon free energy carrier

One of those alternative fuels is H2 (Hydrogen), which is already considered an important raw

material (Ramachandran, R. et al., 1998) by the chemical industry for production of fertilizers

(ammonia) and methanol, and by the refinery industry for upgrading some products. Its importance

tends to grow as the world starts to look at H2 as a free carbon energy carrier, so that from the

combustion of H2 with O2 (Oxygen) results just water vapour (H2O) (2.1).

�� + 1 2 �� → �� (2.1)

In fact, water vapour has the highest contribution to the natural greenhouse effect. However, due to

the natural hydrological cycle, its atmospheric concentration is rather stable.

Besides H2 advantage of being a CO2 non-emitting fuel when is burnt, some of the technologies for

energy conversion that can be used with H2, like fuel cells, have higher efficiencies than those used

for common fuels (Carrette, L. et al., 2001). Nevertheless, the processes for large scale production of

H2 are still based in fossil hydrocarbons reforming (Holladay, J. et al., 2009) (Kothari, R. et al.,

2008) (Rostrup-Nielsen, J. et al., 2002), which contribute to CO2 emissions.

2.2. Carbon Capture and Storage

CO2 is believed to be the main pollutant responsible for anthropogenic climate change, through the

greenhouse effect (Lashof, D. et al., 1990). Hence, efforts to mitigate its world impact have led to the

growing importance and investment in CCS technologies. These technologies involve the capture,

transport and storage of CO2, from large point sources, such as power plants and CO2-emiting

industries. However, many of these technologies are not in a mature market stage and can be found

3

in different implementation levels like research phase, demonstration phase or economically feasible

under specific conditions (IPCC, 2005). SECLR is a process still in research phase.

The introduction of an appropriate tax on emitted CO2 tonnes would potentiate the use of these

technologies, because many of them are only feasible under specific conditions. Several authors have

developed models that describe the degree of implementation of these technologies with the level of

taxes introduced in the system (Biggs, S. et al, 2000) (McFarland, J. et al., 2003). Nowadays, the

value for this tax would be around $50/ton of CO2 (McFarland, J. et al, 2004), value that tends to

increase gradually along the years and depends on the power generation technology used (NGCC,

IGCC, PC).

CO2 Capture

Nowadays, there are four main processes and systems for CO2 capture: post-combustion, pre-

combustion, oxy-fuel and some industrial processes.

Post-combustion systems

Post-combustion systems consist in CO2 separation from flue gases produced in a standard

combustion reaction with air, through the use of an organic solvent like monoethanolamine (MEA).

This system consists in an absorption/stripping process and it is used to recover the small fraction

(typically 3 - 15% by volume) of CO2 present in exhaust gases that are diluted with nitrogen from

air. Although, this technology can be applied to a modern pulverized coal (PC) power plant or to a

natural gas combine cycle (NGCC) power plant, it is only economically feasible under specific

conditions, such as a favourable tax policy (IPCC, 2005) (Kanniche, M. et al., 2010).

Pre-combustion systems

In pre-combustion, fuel is used in a gasifier or a reforming reactor in order to produce syngas or H2

that later can be used as a fuel in other equipment. As long as CO2 is captured and prevented from

reaching the atmosphere, H2 can be used as a carbon free energy carrier. CO2 concentrations and its

pressure in the gas stream are higher, which makes the physical or chemical separation process

easier. Nevertheless, the initial gasification or reform is complex and expensive. Pre-combustion

technologies are also economically feasible under specific conditions and they are suitable for power

plants that employ integrated gasification combined cycle (IGCC) technology (IPCC, 2005)

(Kanniche, M. et al., 2010).

4

Oxy-fuel Systems

Oxy-fuel combustion technology uses oxygen, instead of air, in the fuel combustion. Thus, flue gases

are mainly composed by CO2 and H2O that can be easily separated by cooling and compression,

retrieving a high CO2 concentration stream suitable for capture and storage. Although, oxy-fuel

systems do not have the need for complex separation procedures after combustion, it needs oxygen

separation from air in initial stages with consequent increase in energy requirements. These systems

are in demonstration phase regarding its application as a method for CO2 capturing in boilers.

However, its use in gas turbines is still in research phase (IPCC, 2005) (Kanniche, M. et al., 2010).

Chemical-looping processes, where SECLR is included, are sometimes considered oxy-fuel systems,

but here complete or partial oxidation of the fuel occurs when the fuel reacts with oxygen that is

supplied through a metal oxide, instead of high purity oxygen previously separated from air.

Industrial processes

Carbon capture in industrial processes is the only technology in mature market stage. It is used in

multiple industrial processes, like ammonia production or natural gas separation from CO2, in which

CO2 is a sub product and instead of being released to the atmosphere it can be stored (IPCC, 2005).

CO2 Transportation

There are two ways of transporting the captured CO2 to the storage sites. It can be either transported

by pipelines or by ship. Pipeline technology operates as a mature market and it is already in place to

transport large amounts of CO2 from one place to another in the USA, for instance. However, the

shipping of CO2 is still only feasible under specific conditions. This type of transport is only used in

small scale due to limited demand (IPCC, 2005) (Svensson, R. et al, 2004).

CO2 Storage

There are four main ways for CO2 storage: geological storage, ocean storage, mineral carbonation

and industrial uses (IPCC, 2005) (Blunt, M., 2010).

The first one consists in CO2 injection into geological formations bellow the earth’s surface. This can

be performed, for example, in oil or gas reservoirs, deep saline aquifer or unminable coal beds. As an

enhanced oil recovery technic, the use of CO2 is already a mature market technology. However,

when regarded as a CO2 storage technic, it is only economically feasible under specific conditions.

5

Ocean storage can be either by deposition as a lake type or by dissolution. In both cases, CO2 would

be carried out by ship or pipeline to the storage site. Both technologies are still in research phase and

there have been only small-scale field experiments (Adams, E. et al, 1998) (Herzog, H., 1998). The

environmental impact of this storage alternative is a major concern among the scientific community,

because CO2 dissolution would decrease ocean’s pH with significant biological impact (Haugan, P.,

1997).

The third way involves a chemical reaction between CO2 and alkaline or alkaline-earth oxides. The

process produces silica and carbonates that remain stable for millions of years, avoiding the CO2

release. These final products could then be stored in mines, for instance. Nevertheless, this

technology is still in research phase or in demonstration phase.

Industrial uses of CO2 already operates as a mature market, but it has a small potential, mainly due to

the short periods in which CO2 is retained, and in many processes the net lifecycle emissions are not

reduced.

2.3. Chemical-Looping Combustion

Chemical-looping combustion (Figure1) is a novel process for combustion of different fuels in

which combustion reactions occur in two separate reactors, typically two fluidized bed reactors,

without mixing the fuel and the oxidizing agent, normally air (Lyngfelt, A. et al., 2008). Oxygen

transfer from one reactor to the other is performed by a solid oxygen carrier, typically a metal oxide

(MeO).

In one reactor, fuel is oxidized by the solid oxygen carrier, in the fuel reactor, which results in CO2

and H2O, at the same time, MeO is reduced to metal or a metal oxide with a lower oxygen content

(Me) by the fuel, according to reaction (2.2). On the other reactor, the Me produced is oxidized by O2

to its initial state, in the air reactor, following reaction (2.3).

�� + �2� + 1 2 �� → � �� + 1 2 �� + �2� + 1 2 �� (2.2)

�� + 1 2 �� → �� (2.3)

6

Figure 1 – Schematic description of Chemical-Looping Combustion

The oxygen carrier has another important role. It is responsible to assure the transport of sensible

heat from de air reactor to the fuel reactor (Rydén, M. et al, 2009). Reaction (2.3) is always strongly

exothermic while reaction (2.2) is endothermic, for most oxygen carriers and fuels. The net energy

released by the two reactors is the same as for the ordinary combustion reaction (2.4), whose

equation can be obtained by combining reactions (2.2) and (2.3).

�� + �� + 1 4 �� → � �� + 1 2 �� (2.4)

Flue gases from fuel reactor contain only CO2 and H2O, thus, in order to obtain a stream of CO2 with

a high level of purity, they go through a cooling process in a condenser. This separate stream of CO2

can then be easily captured and stored, which makes CLC a very interesting technology for electrical

power generation with CO2 capture in NGCC or PC power plants.

Thermal NOx emissions are also minimized, which is another important feature of this process.

Besides the fact that fuel is oxidized in a different reactor from the one in which oxygen carrieres are

re-oxidized, the maximum temperature that might be reached in some conditions in the air reactor is

around 1200°C, and thermal NOx formation occurs for higher temperatures (Ishida, M. et al., 1996)

(Hossain, M. et al., 2008).

FUEL

REACTOR

AIR

REACTOR

Fuel

CnHm

Air

N2, O2

Carbon Dioxide

and Steam

CO2, H2O Nitrogen

N2

Me (s)

MeO (s)

7

2.4. Chemical-Looping Reforming

Chemical-looping reforming (CLR) (Figure 2) follows similar principles to CLC. Though, instead of

heat, the product desired is a mixture of CO (Carbon Monoxide) and H2, a mix typically referred to

as synthesis gas. This can be achieved by partially oxidizing the fuel, reaction (2.5), which occurs at

under-stoichiometric conditions, i.e. if insufficient oxygen is added to the air reactor to completely

oxidize the fuel added to the fuel reactor (Rydén, M. et al, 2009).

�� + �� → � �� + 1 2 �� + �� (2.5)

Figure 2 – Schematic description of Chemical-Looping Reforming

Either steam or CO2 could be added to the fuel to promote the steam reforming reaction (2.6),

increasing the H2/CO ratio, or the CO2 reforming reaction (2.7), giving more importance to CO

content of the syngas, respectively (Rydén, M. et al., 2006).

�� + �� → �� + 1 2 �� + � �� (2.6)

�� + � �� → 1 2 �� + 2� �� (2.7)

FUEL

REACTOR

AIR

REACTOR

Fuel

CnHm

Air

N2, O2

Synthesis Gas

CO, H2

Nitrogen

N2

Me (s)

MeO (s)

8

Besides steam reforming and CO2 reforming reactions, also the water-gas shift (WGS) reaction takes

part in this equilibrium and is described by reaction (2.8).

�� +�� ⇌ �� + �� (2.8)

H2 production via WGS reaction is favoured at low temperatures, because it is an exothermic

reaction. As the temperature inside the fuel reactor is high, an excess of steam should be added to

this reactor, in order to counterbalance the negative effect of the higher temperature, so that, exhaust

gases from the fuel reactor would contain CO2, H2 and H2O, in large concentrations, but also lower

amounts of CO. Therefore, if the process is aiming for H2 production, this stream would be subjected

to a cooling stage to withdraw H2O content and, for instance, to several physical or chemical

adsorption/absorption separation processes for decreasing CO2 content, which leads to a higher H2

purity, but decreases overall efficiency of CLR.

2.5. Steam Reforming

Hydrogen production processes mostly used nowadays are based in steam reforming (2.9), combined

with the water-gas shift reaction (2.8). In fact, steam reforming and water-gas shift reaction do not

occur seperately, since both reaction use steam as a reactant, and therefore are competing with each

other. Equation (2.9) and Figure 3 describe only the steam reforming reaction and not the

combination of both. The water-gas shift reaction is explained in the next subsection.

�� + �� → �� + 1 2 �� + � �� (2.9)

Figure 3 – Schematic description of Steam Reforming (without water-gas shift)

Typically in these processes, the hydrocarbon feedstock is initially preheated before it undergoes a

pre-treatment stage, in order to remove traces of sulphur compounds.

After pre-treatment, the feed is mixed with steam at optimized steam/carbon ratio and superheated.

The mixture is then converted in tubes filled with a nickel catalyst. Steam reforming reaction is

endothermic, so it demands external heating to maintain the process temperature, which is usually

FUEL

REACTOR

Fuel and Steam

CnHm, H2O

Synthesis Gas

CO, H2

9

supplied by burning fuel in a furnace surrounding the tubes (Dupont, V. et al., 2008) (Kumar, R. et

al, 1999).

Synthesis gas that comes out of the reformer contains H2, CO2, CO, CH4 and steam, and goes

through the water-gas shift reactor with the purpose of increasing H2 and CO2 content of the syngas,

by reacting CO with steam.

Finally, the resulting syngas is subjected to alternate stages of PSA (Pressure Swing Adsorption) or

to absorption separation with MEA to separate hydrogen from the other components.

2.6. Sorption-Enhanced Water-Gas Shift

There are two ways to enhance, via selective sorption, the production of H2 through the WGS

reaction equilibrium. Either H2 is separated from the main process stream or CO2 is captured and

then released in a different stream (Figure 4). In the first case, a selective permeate membrane is used

to remove H2 from syngas into a separate stream, and a rich CO2 flue gas is produced as a

consequence (Barelli, L. et al., 2008).

For the CO2 separation, a CO2 sorbent could be added. The mostly used sorbent for SE is CaO

(Calcium Oxide), which can be obtained by calcination of limestone at a high temperature (at least

900°C @ 1 atm) (Harrison, D., 2008). For this sorbent, calcium carbonate (CaCO3) is formed after

its reaction with CO2, according with equation (2.10).

�� + �� ⇌ � �� (2.10)

Figure 4 – Schematic description of Sorption-Enhanced Water-Gas Shift

CaCO3 (s)

CaO (s)

Synthesis Gas

CO, H2

Hydrogen

H2

Carbon Dioxide

CO2

FUEL

REACTOR

CALCINATION

REACTOR

Steam

H2O

10

As a result, H2 production and CO consumption will increase, according with the water-gas shift

reaction (2.8), increasing the purity of H2 produced. The spent sorbent would then be regenerated at a

different reactor by heating and calcination, releasing the captured CO2. This process allows the

production of two separate streams, one of high purity H2 and the other of CO2, ready for capture and

storage.

The shift from carbonation to calcination is performed by temperature change as it can be seen in

Figure 5.

Figure 5 - Equilibrium CO2 pressure with CaO/CaCO3 as a function of temperature (Harrison, D., 2008).

11

3. Sorption-Enhanced Chemical-Looping Reforming

Sorption-enhanced chemical looping reforming (SECLR) (Figure 6) combines fuel partial oxidation

that takes place in CLR with CO2 capture and then release in a separate stream from SE Water-Gas

Shift. In this technology as demonstrated in Figure 6 there is a combination of three different

reactors: reforming reactor, calcination reactor and air reactor.

Figure 6 – Schematic description of Sorption-Enhanced Chemical-Looping

SECLR is a rather novel concept but it has been studied by a few researchers with promising results.

Pimenidou et al that had run already tests of chemical-looping reforming with waste cooking oil

(Pimenidou, P. et al., 2010a), improved the process, with the addition of CO2 sorbents (dolomite), to

sorption-enhance chemical-looping reforming (Pimenidou, P. et al., 2010b). They were able to

produce a high purity hydrogen (>95%) at 600°C and 1 atm, using a H2O/CH4 ratio of 4 and with a

100% carbonation efficiency in initial cycles.

3.1. Description

According to what was mentioned in the previous subsection, SECLR process can be considered an

upgrade of CLR process, through the introduction of a third reactor denominated the calcination

reactor.

As in CLR, fuel is introduced in the Reforming reactor (also called fuel reactor) mixed with steam, at

the same time two different types of particles are introduced: oxygen carriers and CO2 sorbents. Fuel

FUEL

REACTOR

AIR

REACTOR

CALCINATION

REACTOR

Fuel

CH4, H2O

Air

N2, O2

Hydrogen

H2

Carbon Dioxide

CO2

Nitrogen

N2

CaCO3 (s)

Ni (s)

CaO (s)

Ni (s)

CaO (s)

NiO (s)

CaO (s)

NiO (s)

12

starts by being partially oxidized by the oxygen carriers (3.1) producing CO and H2. Then, CO reacts

with steam increasing the H2 yield and producing CO2. At this time, the CO2 produced is captured by

sorbents.

In this work, SECLR will be used for methane reforming with NiO, as an oxygen carrier, and CaO,

as a CO2 sorbent. Hence, the general equation for the process can be specified for this hydrocarbon.

�� + �� → �� + 2�� + �� (3.1)

∆�� ! = 211$%/�'( (Rydén, M. et al., 2006)

In fact, not only partial oxidation reaction occurs in the fuel reactor. As fuel is introduced mixed with

steam, there is a combination of partial oxidation (3.1), steam reforming (3.2) and water-gas shift

(3.3) reactions.

�� + �� ⇌ 3�� + �� (3.2)

∆�*�+! = 224,0$%/�'( (Xu, J et al, 1989)

./ = 240$%/�'( 0 = 4,225 × 10�3$�'( ∙ 5�6� � ∙ $�7/89� ∙ ℎ9� (Xu, J et al, 1989)

$ = 251,994$�'( ∙ 5�6� � ∙ $�7/89� ∙ ℎ9�

�� +�� ⇌ �� + �� (3.3)

∆�*�+! = −37,3$%/�'( (Xu, J et al, 1989)

./ = 67,13$%/�'( 0 = 1,955 × 10?$�'( ∙ 5�69� ∙ $�7/89� ∙ ℎ9� (Xu, J et al, 1989)

$ = 390,986$�'( ∙ 5�69� ∙ $�7/89� ∙ ℎ9�

In this work, for the modelling part, it was considered that the outlet flow from the fuel reactor had

reached the chemical equilibrium, but in most cases, the residence time is not enough. So, the final

concentrations depend on the rates of the reactions, as it happened in the experimental part.

Reaction rate of equations (3.2) and (3.3) can be considered proportional to reactions’ rate coefficient

(k) which can be determined by the Arrhenius law. For this model, both steam reforming and water-

13

gas shift reactions have rate coefficient of the same order of magnitude, but the water-gas shift

reaction is faster.

Both partial oxidation reaction (3.1) and steam reforming (3.2) reaction are endothermic, which

means that are favourable at higher temperatures, while water-gas shift reaction (3.3) is exothermic

and therefore, favourable at lower temperatures.

The final composition that comes out of this reactor is given by the equilibrium of these three

reactions and the carbonation reaction (3.4).

�� + �� ⇌ � �� (3.4)

∆��*+! = −179$%/�'( (Lackner, K. et al, 1995)

This carbonation reaction is, like the WGS reaction, an exothermic process.

In the next stage, the particles, after being separated from the gaseous species, go to the calcination

reactor. In this reactor, CaCO3 particles that were originated by the carbonation reaction are calcined,

releasing the CO2 previously captured in a separate stream. This calcination reaction (3.5) is

endothermic and occurs at higher temperatures (at least 900°C @ 1 atm) then those present in the

first reactor, and the different temperatures from the fuel reactor to the calcination reactor are

responsible for this capture/release mechanism. The CO2 stream produced is a high purity stream

with potential for sequestration because only CO2 is released.

� �� ⇌ �� + �� (3.5)

∆��*+! = 179$%/�'( (Zeman, F. et al, 2004)

Once again, particles are separate from the gaseous species and enter the air reactor, where Ni

particles reduced in the fuel reactor are re-oxidized.

�� + 1 2 �� → �� (3.6)

∆�� ! = −234$%/�'( (Rydén, M. et al., 2006)

This oxidation reaction (3.6) is the most exothermic reaction in the all process. This is the main

reason why particles in the system are also used for sensible heat transport from this reactor to the

others, and why the air reactor is typically at a higher temperature than the others.

14

Since the calcination reactor has a heat demand, the particles that leave the air reactor have to go

through this reactor and afterwards, they are divided in different fractions into the fuel and the air

reactors.

3.2. Reactions and chemical equilibrium

Chemical reactions do not always develop in only one direction, as well as reactants are not always

fully consumed according with the equation stoichiometry. Instead, there is a competition between

forward and backward reactions until a chemical equilibrium is reached, in which both have the

same rate (Thunman, H., 2008) (Coelho, P. and Costa M., 2007). After the chemical equilibrium has

been reached, components concentrations do not change until there is any kind of disturbance in the

system like temperature, pressure or composition change.

Besides this effect, we also have to take into account that there might be several reactions occurring

at the same time that interfere with each other’s equilibrium.

The equilibrium condition is given by the combination of 1st (3.7) and 2nd (3.8) laws of

thermodynamics.

AB = AC − DAE (3.7)

AF = AGF + AHF = ACI + AHF (3.8)

AB = IAF − DAE − IAHF (3.9)

As in combustion usually neither the internal energy (U) of the system nor the entropy (S) is kept

constant, it’s preferable to use the Gibbs free energy function, for processes at constant pressure (p)

and temperature (T).

J = � − IF (3.10)

Considering the differential expression of function G and replacing H by � = B+ DE

AJ = A� − IAF − FAI (3.11)

AJ = �AB + DAE + EAD�− IAF − FAI (3.12)

15

AJ = IAF − DAE − IAHF + DAE + EAD − IAF − FAI (3.13)

AJ = EAD − FAI − IAHF (3.14)

At chemical equilibrium the system is considered reversible (IAHF = 0), because both forward and

backward reactions have the same rate that maintains constant the several species concentrations.

Moreover, temperature is kept constant (AI = 0), which simplifies the previous equation.

AJ = EAD (3.15)

If the ideal gas law is introduced in the previous equations and then it is integrated, considering the

reference pressure p° = 1 atm.

AJ = �KLI ADD ⟺ AJ = �KLIA�(� D� (3.16)

J�I, D� = JL�I� + �KLI (� ND DL O (3.17)

Assuming an ideal mixture of gases,

J�I, D� =P�H�H�I, D��

HQ�=P�H R�HL�I� + KLI (� NDH DL OS

�

HQ� (3.18)

For a given specie, Gibbs energy at reference pressure can be determined by the formation Gibbs

energy at reference temperature and the sensible Gibbs energy.

�HL�I� = �H,TL �IUGT� + ∆�H,VL �I� = �H,TL �IUGT� +W AX

XYZ[�HL (3.19)

In equilibrium, where Gibbs energy is minimum, the final expression can be reach for each chemical

reaction or for a complete set of species with several reactions.

AJ�I, D� = 0 ⟹ PA�H R�HL�I� + KLI (� NDH DL OS�

HQ�= 0 (3.20)

−∆JL = KLI (�]^ (3.21)

16

In the fuel reactor, it is not possible to reach a complete conversion of methane due to the chemical

equilibrium between all species: products and reactants. Nevertheless, all oxygen, present in the NiO

particles, is consumed which can be modelled by assuming that reaction (3.1) occurs in a full extent

and that reactions (3.2) and (3.3) are those that explain equilibrium composition variations.

There are five gaseous species that may be present in large concentration in the products equilibrium

composition: CH4, H2O, H2, CO2 and CO. Besides these five species, there are other components that

may also be present, such as light hydrocarbons, but in very small concentrations, and because of

that their concentration is neglected.

The global equation (3.22), that follows, describes the reactions that happen in the fuel reactor.

�� + �� + �� →→ �_ a �� + � bc�� + � b�� + �_c �� + �_cb �� + �� (3.22)

� = �� (3.23)

The final composition of these five components can be determined by three molar balance equations

to the three different chemical elements involved (C, H and O), together with two chemical

equilibrium equations, that result from the combination of expression (3.20) with two linearly

independent reactions equations that characterize equilibrium stage inside that reactor.

Molar balance:

C: 1 = �_ a + �_c + �_cb (3.24)

H: 4 + 2� = 4�_à + 2� bc + 2� b (3.25)

O: m+1 = nfbg + nhg + 2nhgb (3.26)

Chemical equilibrium:

�� +�� ⇌ �� + �� (3.27)

]^ = îjb ∙^kb^kbj ∙îj

�log� ]^ = −5.018@�Iq� (3.28)

17

�� + �� ⇌ 3�� + �� (3.29)

]^ = îj ∙^kbrîka ∙^kbj

�log� ]^ = −3,027@�Iq� (3.30)

Carbon (or coke) deposition is normally undesirable in these types of processes. Carbon (C)

formation reaction can be modelled by several chemical equations (Trimm, D., 1997) (Trimm, D.,

1999).

�� ⇌ �� + 2�� (3.31)

∆��*+! = 75,2$%/�'( (Clarke, S. et al., 1997)

2 �� ⇌ �� + �� (3.32)

∆��*+! = 86,2$%/�'( (Clarke, S. et al., 1997)

The first reaction (3.31), known as methane pyrolysis or methane thermal cracking, is favoured only

at high temperatures (above 800°C), which are not expected to be reach inside the fuel reactor, while

the second one (3.32) is the Boudouard reaction that occurs for lower temperatures (lower than

750°C-700°C at 1 atm). A third chemical equation (3.33) is also commonly used to describe carbon

consumption process, but in fact, it results from the combination of Boudouard and WGS reactions.

�� + �� ⇌ �� + �� (3.33)

Based in this last reaction (3.33), it is easy to understand why there is carbon consumption in the fuel

reactor. In this reactor there is normally an excess of steam which contributes to prevent carbon

deposition.

3.3. Process parameters

Temperature, pressure and H2O/CH4 ratio are the three parameters that change the equilibrium

composition. Nevertheless, these parameters are competing with each other and produce opposite

effects according with the modelling equations. These effects can be somehow predicted according

with Le Chatelier's principle.

The temperature has the most influence in equilibrium composition. An increase in the fuel reactor

temperature would result in a higher methane conversion and because of that, H2 production rate may

18

also increase until a certain temperature, after which starts to decrease due to the lower extent of the

WGS reaction that is favourable by lower temperatures. However, H2 purity level is always

decreasing as a consequence of the reduction of equilibrium carbonation capacity, together with the

previous effect of WGS reaction.

In the calcination reactor, temperature determines whether calcination of CaCO3 takes place or not.

Hence, it is only important to assure the minimum temperature possible for this reaction to occur.

In the air reactor, oxidation reaction is highly exothermic. Thus, the only constrain is that

temperature should be higher enough than the previous reactor, so that sensible heat can be

transported by particles from this reactor to the others, but taking into consideration material

restrictions at high temperatures.

The changes in pressure have no influence in the WGS reaction, since both reactants and products

have the same molar proportions. If the pressure is increased, steam reforming reaction and partial

oxidation reaction are inhibited, due to the volumetric expansion from reactants to products, i.e., in

the products of these two there are more moles than in the reactants. However, the CO2 capture is

improved.

Regarding H2 production, it is preferable to run the process at atmospheric pressure, but other

advantages of using pressurized process must be taken into account. With a pressurized system, the

amount of H2 produced (molar flow) for the same volumetric flow is proportionally higher and there

is no need for so high decompression and (re-)compression systems that reduces the general

efficiency of the process.

The initial concentrations or the inlet mass flow of reactants have also implications on the

establishment of chemical equilibrium. Although, the proportion between particles and methane

introduced in the reactor is kept the same (NiO/CH4=1), the ratio steam/methane (H2O/CH4) can

change. When steam is introduced in the system, methane conversion increases and CO is consumed

through the WGS reaction. As a consequence, H2 production rate increases as well as the H2 purity

of the outlet stream.

Not all steam that is introduced in the reactor is consumed, and a balance must be made between the

H2 production enhancing effect created by the steam and the energy needed to produce that amount

of steam, in order to determine the optimal operating conditions.

19

3.4. Oxygen carriers

In SECLR there are two types of particles involved: oxygen carriers and CO2 sorbents. Oxygen

carriers are used to transport oxygen from the air reactor to the fuel reactor, so that there is no mixing

between fuel and air. CO2 sorbents are used in the fuel reactor to capture CO2 (carbonation reaction)

and to release it in the calcination reactor as a separate stream. This second particles type is

explained in more detail in the next sub-section.

Oxygen carriers are typically a metal oxide and are characterized by several aspects: reactivity,

tendency for agglomeration and resistance to attrition and fragmentation (strength). These are

influenced by the high heat treatment temperatures, active element and supporting material.

Furthermore, these particles should have a low production cost, as well as being environmentally

friendly and non-toxic.

A wide range of oxygen carriers have been tested by several research groups (Cho, P. et al., 2004)

(Adánez, J. et al., 2004). Not only different active components, but also with different supporting

materials. The most common active materials are nickel, copper, iron and manganese oxides. Among

these, nickel-based oxygen carriers are characterized by their high reactivity and capability to avoid

agglomeration, especially if the strength of the particle is increased.

When it comes to CLR, the range of options for oxygen carriers is quite narrow, because for these

applications, the catalytic properties of the particle are extremely important. This good catalytic

behaviour can be found in Ni particles (Mattisson, T. et al, 2006).

The oxygen carrier used during the experimental part consisted of a 40% NiO supported on Mg-ZrO2

(N4MZ1400) that had been tested already for CLC and CLR experiments, and indicate that it could

work well for these applications (Rydén, M. et al, 2009).

3.5. Carbon Dioxide sorbents

Carbon Dioxide sorbents are the particles responsible by the sorption-enhanced mechanism, in which

molecules of CO2 are captured from products, enhancing hydrogen production via water gas shift

reaction. Calcium oxide is a strong candidate as sorbent, because of its availability, low price and

high reactivity in early stages. The reactivity of the sorbent tents to decrease in following cycles,

though, and therefore is necessary to reactivate periodically the sorbent absorption capacity.

Research is being carried out aiming for improvement of the sorbents durability at lower calcination

temperature and higher absorption capacity (Albrecht, K. et al., 2008).

20

Besides Ca-based type, there are also Hydrotalcite (HTC) sorbents under development that can either

be found as a secondary mineral throughout the world or be synthesized. There are as well others

synthetic sorbents under investigation, like potassium, sodium, magnesium or mixed metal oxide of

lithium and sodium (Harrison, D., 2008).

3.6. Work Purpose

The purpose of this work is the study and understanding of the potential of sorption-enhanced

chemical-pooping reforming for hydrogen production with inherent CO2 capture.

The objective of the work is the development of a thermodynamic model of this process in a

commercial software, such as Aspen Plus.

This model will allow the understanding on the influence of steam and CaO presence in the

reforming reactor comparatively to a standard chemical-looping reforming process without them.

Afterwards, it will permit the comprehension on the influence of thermodynamic parameters

(temperature and pressure) in equilibrium concentrations.

The thermodynamic balance of the process is another objective for this model, in order to determine

at what conditions the process can be balanced, avoiding external heat demands and maximizing H2

production and concentration.

Finally, this concept will be tested in a bench-scale fluidize bed reactor, in a laboratory facility, using

a mixture of NiO and CaO particles and CH4 as a fuel. With these experiments is intended to

demonstrate that the combination of methane reforming and CaO carbonation reactions is possible

with improving results.

21

4. Simulation methodology and parameters

4.1. Methodology

The modelling and simulation of SECLR process was made with the support of commercial software

Aspen Plus. This software allows process modelling based on a diagram of blocks that represent the

equipment, reactors and particles separators, which are connected by streams. However, several

assumptions and simplifications must be taken into account.

For modelling each reactor, it was considered a RGibbs block, that calculates the equilibrium

composition of specified components, based in the minimum Gibbs free energy, according with input

streams constituents (reactants) and process operating conditions: temperature and pressure. Hence,

it was assumed that residence time inside the reactor is enough, so that all chemical reactions reach

equilibrium stage.

After the reactors there is always a separator, typically a cyclone separator, whose function is to

perform the separation of the mixed streams that come out of the reactors into two separate streams,

purely solid and gaseous. In this case, an ideal cyclone separator with 100% effectiveness was used,

which corresponds to the SSplit block in Aspen Plus.

Regarding the several chemical species taken into consideration, it was assumed that the only species

that could be present in the equilibrium were CH4, H2O, H2, CO2, CO, C and O2. All other products

that could eventually be formed with the same base elements (C, O, H) were neglected, since their

formation is not expected to occur for these operating conditions. In the case of C and O2, their

concentration in the equilibrium composition was always negligible, as expected.

The raw material for this process is methane, which is the main constituent of natural gas. Although

there are other substances present in natural gas, it was considered a pure methane inlet stream.

The simulation carried out did not take into consideration any changes in the oxygen carriers and

CO2 sorbents reactivity or mechanical strength and that their properties do not depend of any

physical characteristic like porosity or granulometry.

The different particles (NiO and CaO) circulate within the system from one reactor to another.

However, in order to define a specific flow of particles in the closed system, there is a gap between

the air reactor and the calcination reactor, in which particles that come out of the air reactor are fully

calcined and oxidized, i.e., only CaO and NiO particles leave this reactor. Thus, an extra input

stream is then created that has exactly the same composition as the one that leaves the air reactor.

22

Besides this stream, there is a Heater block (B7), which by an iterative process, establishes to this

extra stream the same temperature and pressure that exist inside the air reactor. Both the Heater

block and the auxiliary stream were used just for modeling purposes, as an artifice.

Figure 7 represents the model developed in Asplen Plus with the main components that enter and

leave the three reactors.

Figure 7 – Model developed in Aspen Plus

In Table 1, the several species that are expected to be found in each material stream are indicated,

and Table 2 presents the legend of Figure 7, in which are specified all components present in the

developed model

Table 1 – Species expected to be found in each material stream

Stream Species Stream Species

2, 9 Ni, NiO, CaO 6, 13, 14 NiO, CaO

3 H2, CH4, H2O, CO, CO2

Ni, CaO, CaCO3

7 Ni, NiO, CaO, CO2

11 CaO, NiO, O2, N2

5 Ni, CaO, CaCO3

CH4 (g)

H2O (g)

O2 (g)

N2 (g)

NiO (s)

CaO (s)

High purity H2 (g)

High purity CO2 (g)

Mainly N2 (g)

23

Table 2 – Legend of the several blocks and streams present in Figure6

Block

name

Type Description

B1 RGibbs Fuel (or Reforming) Reactor

B2 Splitter Splits the input stream (3) into two separate streams, one gaseous

(4) and other solid (5)

B3 RGibbs Calcination Reactor

B4 Splitter Splits the input stream (7) into three separate streams, one

gaseous (8) and two solid (2) (9), according with the fraction

specified

B5 RGibbs Air Reactor

B6 Splitter Splits the input stream (11) into two separate streams, one

gaseous (12) and other solid (13)

B7 Heater Auxiliary heater the establishes to the ausiliary stream (14) the

same conditions (temperature and pressure) of stream (13)

1

Material Stream

Input stream of the Fuel reactor, that normally contains CH4 and

H2O

4 Output stream enriched in H2 (CH4, H2O, CO, CO2)

8 Output stream of CO2

10 Input stream of air for the air reactor

12 Output stream of oxygen depleted air (N2)

13 Output stream of particles fully calcined and oxidized

14 Auxiliary input stream, with the same composition as stream (13),

that establishes the particles circulation within the reactors system

2, 3, 5, 6,

7, 9, and

11

Material Stream Streams that establish the connection between all other blocks

24

4.2. Parameters

There are several parameters that can be adjusted in the simulation model. For the reactors, there are

two out of three process conditions that must be specified: temperature, pressure and heat duty. The

temperature of each reactor can be set to different values according with the reactor’s purpose

(reforming, calcination or oxidation). Although the pressure of the three reactors can also be set to

different values, it does not make sense in doing it so, because the model represents a three fluidized

bed system and they operate normally at the same pressure. Heat duty is the third parameter that can

be adjusted, depending on whether the reactors have the need for a specific external heat demand or

release, or on the contrary, they are supposed to work without any external heat demand.

Regarding input streams, besides their relative composition, it must be specified the pressure and the

temperature of their constituents, because these properties will influence the conditions of the reactor

(temperature or heat duty).

The final parameter that can be changed in this model is the ratio of particles that go from the

calcination reactor to the fuel reactor or to the air reactor. This last parameter, together with the

specified particles flow, determines the amount of sensible heat that is transported by these particles

from one reactor to the other. Hence, it ends up by having a crucial effect on the reactors’ operating

conditions.

25

5. Simulation results and discussion

The results presented in the next three subsections (5.1, 5.2, 5.3) intend to demonstrate the effect that

all those process parameters described above have in the equilibrium concentrations of the fuel

reactor. In the fourth subsection (5.4), the thermal optimization of the process is explained.

Finally, in the last subsection (5.5), results are presented that will allow the comparison with

experimental results, presented in the next chapter (6).

5.1. Steam and Calcium Oxide enhancing effect

The simplest mechanism of hydrogen production via partial oxidation does not involve either the

introduction of steam or the presence of CaO particles for CO2 sorption. In these conditions, and for

a specific temperature and pressure, hydrogen production rate and its purity were very low.

Considering as a reference, the reforming without either the introduction of steam or CaO particles,

for a pressure of 1 atm and a temperature of 600°C, it was possible to understand the enhancing

effect that both steam and CO2 sorbents had in the methane reforming.

Table 3 presents the results reached through Aspen Plus modelling in the reforming reactor for five

different condition sets. The first one is the denominated reference that was mentioned above, and it

is followed by the results obtained for similar conditions but in the presence of CaO. H2O effect on

the system’s equilibrium is demonstrated by the third and fourth set, for two different H2O/CH4

ratios. Finally, results for the combination of the two effects are presented.

Table 3 – Molar flows and H2 concentration at 600°C and 1 atm obtained in the reforming reactor for a CH4 flow of 1

kmol/s and for different conditions: ref – without H2O and CaO particles; CaO – with CaO particles but without H2O;

H2O/CH4 – with H2O at different H2O/CH4 ratios but without CaO particles; 2 H2O/CH4 + CaO – with a H2O/CH4 ratio of

2 and CaO particles.

Molar Flow [kmol/s] H2 Dry

Concentration H2 CH4 H2O CO CO2 CaCO3

Reference 0,884 0,4466 0,2232 0,3300 0,2234 - 0,4691

CaO 0,932 0,506 0,0548 0,0418 0,0066 0,4452 0,627

1 H2O/ CH4 1,637 0,2629 0,837 0,3111 0,4260 - 0,621

2 H2O/ CH4 2,118 0,1497 1,583 0,2832 0,567 - 0,679

2 H2O/ CH4 + CaO 2,820 0,0407 1,098 0,0164 0,0171 0,926 0,974

26

In all sets, the inlet flow of methane was 1 kmol/s and the NiO/CH4 ratio of 1, as well. In neither of

them it was taken into account any limitation in the heat duty of the reactors.

When steam was added to the fuel reactor, H2 production rate increased substantially, depending on

the H2O/CH4 ratio. With a ratio of 1, the molar flow of H2 that came out of the reactor increased

85,3%, while with a ratio of 2 it enhanced 139,7%. This large variation was related with the H2O

reaction either via steam reforming or water-gas shift that consequently led to the decrease of

unreformed CH4 equilibrium fraction and the increase of CO2 molar flow. From the combination of

these two reactions, CO decreased.

The presence of CaO particles created a different effect in the equilibrium. In this case, there was

only a small increase of 5,43% in the H2 production rate comparatively with the reference value. In

fact, there was even a small increase in the unconverted methane ratio. Despite all of this, dry H2

purity level of the gas that left the reactor was very similar to the one reached with steam.

Figure 8 allows the comparison of the H2 dry concentration between the several conditions described

above.

Figure 8 – H2 dry concentration at 600°C and 1 atm in the reforming reactor for a CH4 flow of 1 kmol/s and for different

conditions: ref – without H2O and CaO particles; CaO – with CaO particles but without H2O; H2O/CH4 – with H2O at

different H2O/CH4 ratios but without CaO particles; 2 H2O/CH4 + CaO – with a H2O/CH4 ratio of 2 and CaO particles.

From the combination of both effects (steam and CaO particles) it was possible to increase from

0,884 kmol/s of H2 leaving the reactor, with a dry concentration of 46,91%, to a flow of 2,820 kmol/s

with a dry purity level of 97,4%. In this situation, H2 production rate and methane conversion ratio

ref

CaO 1 H2O/CH4

2 H2O/CH4

2 H2O/CH4 +

CaO

0

0,1

0,2

0,3

0,4

0,5

0,6

0,7

0,8

0,9

1

H2

con

cen

tra

tio

n [

% (

V/V

)]

27

were enhanced in a great extent due to the steam, at the same time the CO2 was removed by CaO

particles, which led to the CO reduction as well, allowing a very high purity.

The major advantage existed went these two effects were combined. This allowed the highest H2

production rate as well as the highest H2 purity level for the five tested cases.

5.2. Pressure change effect

The pressure is a very important parameter, since it influences not only the equilibrium to some

extent, but also the whole system dimensioning and efficiency. If the only interest is on the optimal

equilibrium conditions, the value of the pressure in the reactors should be low. However, it should be

taken into consideration that methane needed for this application is available pressurized, and that

the resulting products – H2 and CO2 – have to be at high pressures so that they can be stored. Hence,

by increasing the system pressure conditions, the work needed for the compression stage of the

products would be much lower, increasing process global efficiency.

A pressurized system also allows a proportionally higher production rate (molar flow) for the same

components dimensions comparatively with an atmospheric system, i.e. for the same volumetric

flow.

The values presented in Table 4 were determined in the reforming reactor with a temperature of

600°C and a H2O/CH4 ratio of 2. Like in all cases, the CH4 flow was 1kmol/s and a NiO/CH4 ratio of

1. In this case, the only interest was to understand the pressure influence in the equilibrium

composition.

Table 4 - Molar flows and H2 concentrations at 600°C and different pressure values in the reforming reactor for a CH4 flow

of 1 kmol/s, with a NiO/CH4 ratio of 1 and with a H2O/CH4 ratio of 2 and CaO particles.

Pressure

[atm]



1 2,820 0,0407 1,098 0,0164 0,0171 0,926 0,974

5 2,548 0,1121 1,227 0,0026 0,0033 0,882 0,956

7 2,471 0,1315 1,266 0,0017 0,0024 0,864 0,948

10 2,383 0,1536 1,309 0,0011 0,0016 0,844 0,938

20 2,198 0,2000 1,401 0,0005 0,0008 0,799 0,916

28

For increasing pressures, CH4 conversion was reduced due to the volumetric increase, inherent to the

reactions in hand, that is hampered at pressurized conditions. Consequently, CO flow and H2

production and concentration were also reduced. Only CO2 sequestration by particles was enhanced,

although the total amount of CO2 produced was also reduced.

5.3. Temperature variation effect

The temperature variation simulation was carried out for a H2O/CH4 ratio of 2, in a pressurized

process at 5 atm a methane flow of 1 kmol/s and a NiO/CH4 ratio of 1. Both cases, with and without

CaO particles, were tested, so that it was possible to understand the influence of these particles, when

the reforming reactor’s temperature changes. The pressure of 5 atm was chosen in order to assure the

pressure effect on these results, but without any specific reason for this particular value.

Table 5 presents the results obtained for the case with CaO particles and at different temperatures.

Table 5 - Molar flows and H2 concentrations at different temperatures and 5 atm in the fuel reactor for a CH4 flow of 1

kmol/s, with a NiO/CH4 ratio of 1 and with a H2O/CH4 ratio of 2 and CaO particles.

Temperature

[°C]



600 2,548 0,1121 1,227 0,0026 0,0033 0,882 0,956

625 2,558 0,1088 1,224 0,0058 0,0065 0,879 0,955

650 2,567 0,1049 1,223 0,0127 0,0124 0,870 0,952

675 2,574 0,0997 1,226 0,0263 0,0227 0,851 0,945

700 2,577 0,0924 1,238 0,0524 0,0407 0,814 0,933

725 2,572 0,0819 1,264 0,0999 0,0713 0,747 0,910

750 2,550 0,0669 1,316 0,1821 0,1232 0,628 0,872

775 2,493 0,0480 1,411 0,3145 0,2117 0,4258 0,813

800 2,378 0,0284 1,566 0,5088 0,3638 0,0990 0,725

Temperature has a major influence in reactions extension. When this parameter rises, partial

oxidation and steam reforming reactions are intensified, what leads to a higher methane conversion

rate. When the temperature went from 600°C to 750°C, unreformed methane fraction decreased from

29

11,21% to 6,69%, in presence of CaO particles. As a consequence, H2 molar flow tended to increase,

but it only happened until a certain temperature (approximately 700°C) and slowly, because at the

same time, the water-gas shift reaction, that is favoured at low temperatures (typically 200°C –

300°C), had a lower reaction rate. For the same reason, the water content started to increase slowly

(at 675°C).

Due also to the increased temperature, particles ability for CO2 capture decreased. Because of this,

and despite the lower amount of CO2 formed, the amount of CO2 in the outlet stream tended to

increase.

Besides CO2, the CO content also increased, because of a greater methane conversion and a

reduction of the reaction between CO and steam by the WGS.

In order to enable the comparison between the previous results, Table 6 is used to demonstrate the

flow values computed for the absence of CaO particles.

Table 6 - Molar flows and H2 concentrations at different temperatures and 5 atm in the fuel reactor for a CH4 flow of 1

kmol/s, with a NiO/CH4 ratio of 1 and with a H2O/CH4 ratio of 2 and without CaO particles.

Temperature

[°C]



600 1,302 0,3937 1,911 0,1232 0,4831 - 0,566

625 1,484 0,3363 1,843 0,1706 0,4932 - 0,598

650 1,668 0,2762 1,780 0,2274 0,4965 - 0,625

675 1,844 0,2162 1,723 0,2910 0,4928 - 0,648

700 2,003 0,1599 1,677 0,3570 0,4830 - 0,667

725 2,135 0,1112 1,642 0,4200 0,4688 - 0,681

750 2,234 0,0728 1,621 0,4753 0,4519 - 0,691

775 2,298 0,0454 1,611 0,520 0,4341 - 0,697

800 2,335 0,0274 1,611 0,556 0,4166 - 0,700

In the absence of CaO particles, the percentage of unreformed methane decreased more sharply when

the temperature went from 600°C to 750°C, from 39,37% to 7,28%, and therefore, the flow variation

of the other components was different from the previous case.

30

In this situation, the higher variation of methane conversion rate overcame the variation in the WGS

reaction rate. Thus, the flow of H2 and CO was continuously increasing, while the H2O content was

decreasing.

Only the CO2 that is formed just by the WGS reaction showed a reduction of its production rate after

reaching a maximum flow for 650°C, caused by the increase in the steam reforming reaction extent.

Another important aspect was the understanding of the effect that the presence of CO2 sorbent

particles had in the stability of H2 flow, as well as, in the increase of its concentration for lower

temperatures than 800°C, approximately. For higher temperatures, the carbonation reaction is

deprecated comparatively to the calcination reaction, and therefore there was no difference between

the two cases.

These effects can be better perceived in curves from Figures 9 and 10, that were plotted based on

values from Tables 5 and 6.

Figure 9 – H2 dry concentration evolution with temperature with and without CaO particles at 5 atm in the fuel reactor and

for a CH4 flow of 1 kmol/s, with a NiO/CH4 ratio of 1 and with a H2O/CH4 ratio of 2

Regarding H2 dry concentration in presence of CaO particles, the outlet stream increased from 72,5%

at 800°C to 95,6% at 600°C, while without their presence, it decreased from 70,0% to 56,6%, for the

same temperature variation.

0

0,1

0,2

0,3

0,4

0,5

0,6

0,7

0,8

0,9

1

600 650 700 750 800

H2

Vo

lum

etr

ic c

on

cen

tra

tio

n [

%]

Temperature [°C]

With CaO

Without CaO

31

Figure 10 – H2 molar flow evolution with temperature with and without CaO particles at 5 atm in the fuel reactor and for a

CH4 flow of 1 kmol/s, with a NiO/CH4 ratio of 1 and with a H2O/CH4 ratio of 2

Relatively to H2 flow, it had in presence of CaO a slight increase of 7,15%, reaching a maximum of

8,37% at 700°C. Without CaO particles, there was a sharp descent of 44,2% from 2,335 kmol/s to

1,302 kmol/s. At 600°C, it was possible to reach a H2 flow 1,95 times higher with the sorption-

enhanced mechanism.

For a SECLR system operating at 5 atm, with a H2O/CH4 ratio of 2 and without any restrictions in

the reactors’ heat duty, the best operating conditions are between 600°C and 650°C which allow H2

concentrations in the outlet stream of over 95% for similar molar flows reached at a higher

temperature (700°C). Furthermore, if unconverted methane could be easily retrieved from this outlet

stream and reused, it would even result in lower CH4 inlet flow for the same results.

5.4. Thermal optimization

The three reactors have different heat requirements due to the endothermic and exothermic reactions

that take place in each one. In one hand, the air reactor usually releases heat, because the oxidation of

Ni particles into NiO is an exothermic process. On the other hand, calcination reaction is

endothermic and therefore, a heat supply is required for the release of the previously captured CO2 in

the calcination reactor. Furthermore, in the fuel reactor, there is also an heat demand, because of the

partial oxidation and steam reforming, which are both endothermic in this case, despite the

exothermic behaviour of the WGS and carbonation reactions.

0

0,5

1

1,5

2

2,5

3

600 650 700 750 800

H2

Mo

lar

Flo

w [

km

ol/

s]

Temperature [°C]

With CaO

Without CaO

32

For a self-sufficient process, the heat release from the air reactor should be enough to compensate the

heat demand from the other reactors. Additionally, in order to avoid complex systems for this heat

exchange between reactors, the flow of particles from one reactor to another should be optimized in a

way, so that no extra heat is required.

For this optimization, it was also taken into consideration the possible pre-heating of inlet streams

through a heat exchange with outlet streams. It was assumed that all streams were pre-heated at the

same temperature, and therefore, it was possible with a composite curves analysis to determine the

amount of heat exchanged between these hot and cold streams.

The method of setting the pre-heating temperature was an iterative process. Firstly, inlet streams

were assumed to be pre-heated to certain temperature and based on this assumption, it was

determined the temperature of the reactors, considering a null heat duty on those reactors.

Afterwards, a composite curves analysis was performed, in order to check that the assumed pre-

heating temperature for the inlet streams was acceptable.

The results from Table 7 were achieved with a pre-heating temperature of 600 °C.

The temperature in the calcination reactor (Table 7) was well above the limit temperature for the

calcination reaction to occur, which at 5 atm is approximately 967°C and thus the SE mechanism is

assured.

The temperatures of the three reactors as well as their heat duty are presented in Table 7. The heat

duty is the energy necessary to supply to or to extract from the reactors for assuring the specified

temperature.

Table 7 – Temperatures and Heat Duties of the three reactors considering input streams at 600°C and almost null Heat

Duties

Reactor Temperature [°C] Heat Duty [MW]

Fuel Reactor 702 0,055

Calcination reactor 1071 -0,52

Air Reactor 1190 0,388

The values of the residual heat duty present in Table 7 for those reactors’ temperatures is very low

comparatively to the values involved in this heat exchange process (< 0,5%) and hence it can be

neglected.

33

In Figure 11 the composite curves that correspond to a hot and cold stream was plotted, and from

which was possible to determine the amount of sensible heat that is possible to transfer between

them.

Figure 11 –Composite Curves considering input streams at 600°C and almost null Heat Duties

Based on this analysis, from Figure 11, it was possible to transfer, approximately, 205MW of heat

from hot streams to cold streams, through a heat exchanger network, for each kmol of CH4 that

entered the fuel reactor.

The composition of the stream that leaves the reforming reactor at 702°C is shown in Table 8.

Table 8 – Molar flows and H2 concentration at 702°C and 5 atm in the fuel reactor for a CH4 flow of 1 kmol/s, with a

NiO/CH4 ratio of 1, with a H2O/CH4 ratio of 2 and with CaO particles

Molar flow [kmol/s] H2 Dry


2,577 0,0917 1,239 0,0552 0,0426 0,810 0,931

0

200

400

600

800

1000

1200

1400

0 50 100 150 200 250 300

Te

mp

era

ture

[°C

]

Power [MW]

cold stream hot stream

205 MW

34

For the first optimization case, the process efficiency was 77,8%, regarding the conversion of CH4 in

H2 (LHVCH4 = 802MJ/kmol; LHVH2 = 242MJ/kmol), and at the same time, the CO2 capture rate was

85,0%.

With a temperature of nearly 700°C, the H2 produced in the fuel reactor had a purity level lower than

95%. Hence, it was possible to reduce the fuel reactor´s temperature to 650°C, which lead to an

increase in H2 concentration. However, pre-heating temperature had to be reduced due to lower

temperatures in all reactors.

In a similar way, results were recalculated considering a temperature inside the reforming reactor of

650°C and these results are presented in Table 9.

Table 9 – Temperature and Heat Duties of the three reactors considering input streams at 550°C, almost null Heat Duties

for the Calcination and Air reactors and a 650°C temperature in the fuel reactor

Reactor Temperature [°C] Heat Duty [MW]

Fuel Reactor 650 -11,10

Calcination reactor 977 -0,229

Air Reactor 1076 0,177

After adjusting the pre-heating temperature to 550°C, it was possible to optimize the process, with

almost the same amount of H2 produced and with a concentration higher than 95%. For these

conditions, it was necessary to increase the particles flow within the system, so that the temperature

of the calcination reactor was high enough to allow the calcination of CaCO3 particles. As a

consequence, the temperature of the air reactor suffered a decrease.

Like in Table 8, Table 10 presents the composition of the stream that leaves the reforming reactor at

650°C.

Table 10 – Molar flows and H2 concentration at 650°C and 5 atm in the fuel reactor for a CH4 flow of 1 kmol/s, with a

NiO/CH4 ratio of 1, with a H2O/CH4 ratio of 2 and with CaO particles

Molar flow [kmol/s] H2 Dry


2,567 0,1049 1,223 0,0127 0,0124 0,870 0,952

At 650°C, the process efficiency had a slight decrease to 77,4%, while the CO2 capture rate increased

to 98,6%.

35

Figure 12 presents a plot of the composite curves that correspond to a hot and cold stream, and from

which was possible to determine the amount of sensible heat that is possible to transfer between

them.

Figure 12 –Composite Curves considering input streams at 550°C, almost null Heat Duties for the Calcination and Air

reactors and a 650°C temperature in the fuel reactor

For this second case, based on Figure 12, it was possible to transfer, approximately, 194MW of heat

from hot streams to cold streams, through a heat exchanger network, for each kmol of CH4 that

entered the fuel reactor.

As there was still a heat release of 11,10 MW in the fuel reactor (Table 8), which is equivalent to

5,7% of the energy involved in the heat exchanged, the H2O/CH4 ratio could be increased.

Nevertheless, this increase could only be from 2 to 2,1, decreasing this heat release to 9,68MW,

since for higher values the temperature in calcination reactor was no longer sufficient for the

calcination reaction.

0

200

400

600

800

1000

1200

0 50 100 150 200 250 300

Te

mp

era

ture

[°C

]

Power [MW]

cold stream hot stream

194 MW

36

6. Experiments description and parameters

The experiments had the main objective to demonstrate, that in fact, this enhancing effect, created by

the addition of Calcium Oxide particles to the standard CLR process, improves the purity of the

hydrogen produced and allows the creation of two separate streams of H2 and CO2. These

experiments were carried out in a bench-scale fluidized bed reactor system at atmospheric pressure

with a bed mass constituted by 1/3 of 40% Ni particles (N4MZ1400) and 2/3 of sand or CaO

particles (limestone, calcined at 1000°C). The quartz reactor used had 820mm long and 22mm wide

with a porous plate located at 370mm from the bottom.

Figure 13 shows the reactor location inside the oven, as well as the two thermocouples introduced

inside.

Figure 13 – Reactor location inside the oven

The experimental procedure is divided in six different stages: setting up the reactor, calibration, start

up, experiments, turn off and cleaning. The procedures for each stage are presented in detailed in

Appendix A.

In set up phase, particles are weighted and introduced inside the reactor. The reactor is then

assembled, attached to the oven and remaining connections and thermocouples are introduced. The

two thermocouples were placed over and above the porous plate inside the reactor where the bed

material is introduced. Before the reactor is insulated, a leakage test is performed to make sure that

all junctions are well tight.

Quartz

Reactor

Thermocouple

Thermocouple

37

During calibration stage, firstly pure nitrogen is used for a zero calibration and hereafter, two

different gases are used for a span calibration, one for CH4, H2, CO and CO2 sensors, and another for

O2 sensor. The zero calibration together with the span calibration were used in order to establish a

linear relation between the sensor output signal and the corresponding concentration value.

In start-up, all other equipment (oven, cooler, heating bands and steam generator) are turned on as

well as the cycle controller window on the computer. Before starting the experiments, a steam

calibration is performed in a reduction cycle.

For the experimental phase, all flows should be set to the correct value, as well as the oven

temperature, and the results ought to be monitored through the control window on the computer.

After all experiments, all equipment and gas flows should be shut down by a specific order and

particles are to be removed to a specific container.

Finally, the reactor must be cleaned properly and before being stored.

Figure 14 shows a schematic description of experimental setup and Figure 15 presents the laboratory

workstation

Figure 14 – Schematic description of experimental setup

38

a b

c d

e f

g

h

i

j

Figure 15 – Laboratory workstation: a - oven with reactor inside; b – electronic equipment for cycle switch; c – gases flow meters; d – gases piping system; e – cooling unit; f – general perspective; g – console for flow regulator; h – steam generator; i – analyser; j – valves control.

39

Experiments consisted in a sequence of oxidation, reduction and inert cycles, so that it was possible

to simulate the conditions for the air reactor, during oxidation, and for the fuel reactor, during

reduction. Inert cycles (N2) were used with a flow of 600ml/min to flush the reactor, in order to avoid

the mixture between the gases from oxidation cycle and the gases from reduction cycle.

During oxidation cycles, a mixture of pure nitrogen and a 20,8% of oxygen in nitrogen flowed

through the reactor until all oxygen carrier particles were fully oxidized with a volumetric flow of

900ml/min. As a criterion it was considered that oxygen carrier particles were fully oxidized when

oxygen concentration at the reactor’s exit was greater than 95% of the inlet value.

Figure 16 presents the typical evolution of O2 concentration in the oxidation cycle of Ni particles.

Figure 16 – Typical evoltion of O2 concentration at the reactor’s exit in a oxidation cycle

Initially in the oxidation cycle, the O2 concentration at the exit of the reactor was approximately zero

while most of the Ni particles were being re-oxidized, and after a certain stage, the concentration had

a rapid increase tending for the inlet concentration (in this case 5%), revealing that most particles

were already in NiO form.

During reduction cycles, it was introduced in the reactor methane and steam at a specific H2O/CH4

ratio with a CH4 flow of 200 ml/min. The duration of this cycle depended on the amount of Ni

particles introduced inside the reactor and methane flow rate.

The quantity of CaO particles introduced inside the reactor determines the CO2 absorption capacity

of the bed. Thus, after a few number of cycles the reactor had to be heated up at high temperatures

(800°C-900°C) so that sorbent particles were regenerated. This was performed in an inert cycle and

simulated the conditions inside the calcination reactor. The Oxygen carrier used was meant to work

at high temperatures (900°C) and therefore, sintering problems was not expected. As the objective

0

1

2

3

4

5

7200 7400 7600 7800 8000

Co

nce

ntr

açã

o [

% (

V/V

)]

Time reference [s]

Concentração de O2

40

was not the study of the particles performance, no further tests were conducted, but visually there

was not any significant change in the particles dimensions. CO2 sorbent (CaO) was previously

calcined at 1400°C, so it was adequate to work at 900°C. Like the oxygen carriers, there was not any

visual change in its dimensions.

Before starting the actual experiments at lower temperatures (600°C-750°C), several cycles had to be

run at higher temperatures (800°C-900°C) in order to activate the oxygen carrier particles and to

prevent the CO2 sorbent particles from capturing the CO2 released in those cycles.

When establishing comparisons between experimental results and the simulation results, one must be

aware that the gas mixture that left the reactor was not in equilibrium, contrary to what happed in the

simulation results. Neither the residence time of the gases inside the reactor was not enough, nor the

reactor was well stirred, so that equilibrium could be reached.

41

7. Experiments results and discussion

The first two sets analysed consisted in a 15g bed mass, composed by 5g of a 40% Ni particles and

10g of sand or 10g of CaO particles. For these two sets, it was used a flow, during oxidation, with

5% oxygen in nitrogen and a H2O/CH4 ratio of 1,8. There were some difficulties, though, in assuring

a stable steam flow by the steam generator. Nevertheless, reduction cycles were conducted for four

temperature conditions (600°C, 650°C, 700°C and 750°C).

Other difficulties were found in the oscillating values measured by the gas analyser, which only

allowed an estimate of the gases’ concentration that came out of the reactor, as it can be seen in

Appendix C and Appendix D. Typically, errors caused by the accuracy of instruments are smaller

than errors inherent to the whole process (interferences, answering time, gas leakage, pulsing steam

flow, etc.). For a standard measurement of one gas concentration in N2, the analyser could be easily

calibrated with a very high accuracy.

During reduction, the only gases that may have left the reactor in significant concentrations were H2,

CH4, CO2 and CO, but the sum of its measured concentration did not reach 100%. So, despite the

sum did not reach 100%, it was made the assumption that these measured concentrations were in the

right proportion with each other and therefore, the real concentration could be determined. Hence,

each measured value by the analyser was divided by the sum of all concentration values. This

assumption would need a more deep reflection, in order to access what is the reason for this effect in

the measured concentrations.

In Table 11 and Table 12 are presented the concentration results for the several species from the

experiments carried out with 5g of N4MZ1400 particles, with sand (instead of CaO) or with CaO

particles, respectively.

Table 11 – Experiment results from the 5g of N4MZ1400 particles experimental set with sand, for a CH4 flow of 200

ml/min with a H2O/CH4 ratio of 1,8

Temperature

[°C]

Measured Concentration [%] Corrected Concentration [%]

H2 CH4 CO2 CO ∑ H2 CH4 CO2 CO

600 56 16 14 4 90 62,2 17, 8 15, 6 4,4

650 60 13 13 6 92 65,2 14,1 14,1 6,5

700 61 9 12 11 93 65,6 9,7 12,9 11,8

750 58 8 7 16 89 65,2 9,0 7, 9 18,0

42

Table 12 - Experiment results from the 5g of N4MZ1400 particles experimental set with CaO, for a CH4 flow of 200


Temperature

[°C]

Measured Concentration [%] Corrected Concentration [%]


600 68 12 0 1 81 84,0 14,8 0,0 1,2

650 68 13 1 4 86 79,1 15,1 1,2 4,6

700 63 14 3 5 85 74,1 16,5 3,5 5,9

750 60 6 6 16 88 68,2 6,8 6,8 18, 2

Despite the rough estimate of gases concentration it was possible to demonstrate the enhancing effect

that results from the addition of CaO particles, as well as the tendency for gases concentration

evolution with temperature.

Figure 17 represents, graphically, H2 concentration evolution with temperature from experiments

carried out with 5g of N4MZ1400 particles, both with and without CaO, so that it is easier to

visualize the sorption-enhanced effect.

Figure 17- H2 concentration for different temperatures from experiments carried out with a 5g of N4MZ1400 particles and

CaO particles or sand, for a CH4 flow of 200 ml/min with a H2O/CH4 ratio of 1,8.

At 600°C, there was a 35% increase in H2 concentration, when CaO particles were used. This

enhancement decreased as reduction cycles were conducted at higher temperatures. Although this

tendency was expected, H2 concentration values were lower than predictable, because they do not

60

70

80

90

100

600 650 700 750

Vo

lum

etr

ic c

on

cen

tra

tio

n [

%]

Temperature [°C]

H2 with CaO

H2 with sand

43

correspond to equilibrium concentrations. In these experiments, neither the residence time inside the

reactor nor the reaction rate was sufficient to reach equilibrium stage. Therefore, unreformed CH4

was considerably high which contributed to lower H2 concentrations.

With the purpose of attenuate the effect of unreformed CH4 in the H2 concentration, the ratio

H2/(H2+CO2+CO) was used in Figure 18 instead of the actual H2 concentration. The lines represent

the H2 equilibrium concentration determined from the Aspen Plus simulation model for a CH4 flow

of 1 kmol/s, with a H2O/CH4 ratio of 1,8 and at 1 atm.

Figure 18 – Evolution of the H2/(H2+CO2+CO) ratio for different temperatures from experiments carried out with a 5g of

N4MZ1400 particles and CaO particles or sand, for a CH4 flow of 200 ml/min with a H2O/CH4 ratio of 1,8.

Once for SECLR process, the amount of unreformed CH4 in equilibrium is rather small, this ratio

approaches the expected H2 equilibrium concentrations.

Nevertheless, for CLR process, the ratio H/C at equilibrium was lower than the previous process, i.e.

the quantity of unreformed CH4 at equilibrium was higher. Therefore, the ratio H/C for the

experimental results, without considering CH4, was higher than the one at equilibrium, and as a

consequence, the value of H2/(H2+CO2+CO) was also higher.

Figure 19 represents, graphically, the CH4, CO2 and CO concentration evolutions with temperature

from the experiments carried out with 5g of N4MZ1400 particles, both with and without CaO, so

that it is easier visualize the sorption-enhanced effect.

60

70

80

90

100

600 650 700 750

Vo

lum

etr

ic c

on

cen

tra

tio

n [

%]

Temperature [°C]

H2 with CaO (equilibrium) H2 with sand (equilibrium)

H2/(H2+CO+CO2) with CaO H2/(H2+CO+CO2) with sand

44

Figure 20 describes the CH4, CO2 and CO concentration expected from the Aspen Plus simulation

model for a CH4 flow of 1kmol/s, with a H2O/CH4 ratio of 1,8 and at 1 atm.

Figure 19 – CH4, CO2 and CO concentration for different temperatures from experiments carried out with a 5g of

N4MZ1400 particles and CaO particles or sand, for a CH4 flow of 200 ml/min with a H2O/CH4 ratio of 1,8.

Figure 20 – CH4, CO2 and CO concentration for different temperatures obtained from Aspen Plus simulation at 1 atm in

the fuel reactor, for a CH4 flow of 1 kmol/s, with a NiO/CH4 ratio of 1, with a H2O/CH4 ratio of 1.8, with and without CaO

particles

0

10

20

600 650 700 750

Vo

lum

etr

ic c

on

cen

tra

tio

n [

%]

Temperature [°C]

CH4 with sand

CO2 with sand

CO with sand

CH4 with CaO

CO2 with CaO

CO with CaO

0

10

20

600 650 700 750

Vo

lum

etr

ic c

on

cen

tra

tio

n [

%]

Temperature [°C]

CH4 with CaO

CO2 with CaO

CO with CaO

CH4 with sand

CO2 with sand

CO with sand

45

The evolution demonstrated by the concentration curves for CO2 and CO was expected, since in CaO

presence, CO2 concentration was lower and CO concentration was higher than the ones reached in

absence of these particles. However, and contrary to what was expected, CH4 concentration was

higher for 650°C and 700°C in the experiments carried out with CaO instead of sand. These

unexpected results are possibly related with the uncertainty associated with those measures and also

with the fact that these chemical species were not in equilibrium yet when they left the reactor and

were analysed.

An attempt to improve the previous results, led to two other sets of experiments, in which the bed

mass was multiplied by a factor of 3, i.e. inside the reactor there was a particles sample of 15g of

40% Ni particles and 30g of CaO or 30g of sand. In order to reduce the oxidation time of these

particles, the concentration of oxygen in the flow during oxidation was increased to 10%. These

experiments were conducted with the same H2O/CH4 ratio and for the same temperatures.

Table 13 presents the concentration results for the several species from the experiments carried out

with 15g of N4MZ1400 particles with sand (instead of CaO).

Table 13 - Experiment results from the 15g of N4MZ1400 particles experimental set with sand, for a CH4 flow of 200


Temperature

[°C]

Measured Concentration [%]

Dry Concentration [%]


600 58 14 10 8 90 64,4 15, 6 11,1 8, 9

650 62 7 8 13 90 68, 9 7, 8 8, 9 14,4

700 64 3 6 16 89 71,9 3,4 6,7 18,0

750 66 1 5 18 90 73,3 1, 5, 6 20,0

On one hand, the gases concentration had smaller fluctuations, which allowed more precise readings,

and on the other, it was possible to reach a higher methane conversion and, consequently, the results

were closer to the ones obtained via Aspen Plus.

Figure 21 represents, graphically, the H2 concentration evolution with temperature from the

experiments carried out with 15g of N4MZ1400 particles.

46

Figure 21 - H2 concentration for different temperatures from experiments carried out with a 15g of N4MZ1400 particles

and sand, for a CH4 flow of 200 ml/min with a H2O/CH4 ratio of 1,8.

Figure 22 represents, graphically, the CH4, CO2 and CO concentration evolution with temperature

from the experiments carried out with 15g of N4MZ1400 particles.

Figure 22 - CH4, CO2 and CO concentration for different temperatures from experiments carried out with a 15g of

N4MZ1400 particles and sand, for a CH4 flow of 200 ml/min with a H2O/CH4 ratio of 1,8

In spite of the improvement on the results from the experiment carried out with sand, for the

experiments with CaO particles the results were inconclusive.

60

70

80

90

100

600 650 700 750

Vo

lum

etr

ic c

on

cen

tra

tio

n [

%]

Temperature [°C]

H2 with sand

0

10

20

600 650 700 750

Vo

lum

etr

ic c

on

cen

tra

tio

n [

%]

Temperature [°C]

CH4 with sand

CO2 with sand

CO with sand

47

8. Conclusions

SECLR was found to be a promising process for H2 production from hydrocarbons reforming, due to

high efficiencies of the conversion process, high purity of H2 outlet stream from the fuel reactor and

high CO2 capture rates. From the thermodynamic analyses several conclusions could be drawn:

• There is a major advantage in combining H2O and CaO in the standard CLR and these have

a complementary effect. H2O enhances the CH4 reforming and H2 production via steam

reforming and WGS reactions, while CaO increases H2 production and purity via WGS by

capturing the CO2.

• From a thermodynamic point of view, low pressures would be better for H2 production.

However, pressurized conditions would be favourable for an industrial application for two

practical and economic reasons: higher overall efficiency, taking into consideration the

necessary compression work for CO2 and H2, and reasonable plant size for a typical H2

production rate.

• In terms of temperature, the best results were obtained for lower temperatures (600 °C - 650

°C) than the ones normally used in CLR.

• The process efficiency and CO2 capture rate is enhanced by the CaO presence. It was

possible to reach 77,8 % efficiency, regarding CH4 conversion to H2, and 95,0 % of CO2

capture.

• The process can be optimized in order to avoid any heat demands by the reactors, without

significant losses in H2 production and purity.

• It was possible to demonstrate this enhancing effect in the experiments carried out in the

bench-scale fluidize bed reactor. Despite the fact that the oxygen carriers used were not

specifically design to work at such low temperatures (600 °C – 700 °C) and therefore, there

performance could be improved, and that the residence time inside the reactor was not

enough so that chemical equilibrium could be reached.

48

9. Future Work

There are several aspects that can be the target of a future research and experiments. Oxygen carriers

are design and manufactured for particular working conditions, in this case, it was initially tested

three different oxygen carriers (N2AM1400 – 20% NiO supported on MgAl2O4, N4MZ1400 - 40%

NiO supported on Mg-ZrO2 and N4NA1400 – 40% NiO supported on NiAl2O3) all of them meant to

work at higher temperatures (900 °C – 950 °C). The first type tested, did not work for the lower

temperatures (600 °C), and, therefore, was excluded. The other two types worked at 600 °C, but the

N4MZ1400 had better results. Even though, the oxygen carrier properties could be improved since

all of them were not meant for this low temperatures.

As it was mentioned in sub-section 3.5, there are several types of CO2 sorbents under research, so

different particles for CO2 capture could be tested, comparing the effect of each one in equilibrium

composition at a specific temperature.

In this work there was no study on particles properties degradation with repeating cycles. There was

only a visual analysis making sure that there were no agglomeration signs in the particles. Further

tests, could be carried out with the objective of studying the particles properties evolution at low

temperatures as well as any kind of interaction between them that might exist.

New experiments could be conducted with a bigger bed mass, in order to increase the CH4

conversion, getting closer to the equilibrium composition.

Finally, the thermodynamic model could be improved by introducing a kinetic model for the

reactions in a different type of reactor. This model could have been developed if there were more

reliable experimental results.

49

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52

Appendix A – Experimental and cleaning procedure

53

Preparing the reactor

• Weight sample, (15g, 40g, etc.) introduce it to the reactor and measure height. Attach the

reactor inside the oven, attach thermocouples. Use silicone to avoid leakages.

• Turn on the computer… Start program, click “starta mätning”.

• Open nitrogen tube, open N2 in ceiling, turn up the flow (wheel #8). K2 to analyzer, K3 to

K7, K7 to reactor. Make sure flow is OK. Insulate to oven and above oven with wool (care

not to let it touch your skin). Insulate open surfaces and wool with aluminum foil.

Calibration

• Turn on the CO-measurement device (yellow).

• Use the nitrogen flow you will use during experiment. K2 to K7, K7 to analyzer. All

concentration should be 0. Zero calibrate all lines:

Main F3, Measure, Channel, until CH4, Calib, Zero Calibration

Measure, Channel until CO2, Calib, Zero Calibration

Measure, Channel until CO, Calib, Zero Calibration

Measure, Channel until H2, Calib, Zero Calibration

Measure, Channel until O2, Calib, Zero Calibration

• Turn down N2 flow to 0. Open gas tubes with 4,99% O2 and CH4,CO2,CO and H2. Turn up

calibration gases flow, use the flow you will use during experiment. Turn K1 to ON and

“blå” to CH4, CO2, CO. K2 still to vent. CO2 should be 40%, CO and CH4 20% and H2

19,9%. Calibrate each line by Span calibration:

Measure, Channel until CH4, Clib., Span Calibration

Measure, Channel until CO2, Clib., Span Calibration

Measure, Channel until CO, Clib., Span Calibration

Measure, Channel until H2, Clib., Span Calibration

Reduce calibration gas from the tube.

54

• Turn up the 4,99% O2 flow to a flow you will be using during the experiment. Turn “blå” to

O2. Calibrate O2 line:

Measure, Channel until O2, Calib., Span Calibration

Reduce O2 from the tube, K1 to off, K2 to analyzer. K7 to reactor.

Start up

• Name files (2, pressure file too)

• Turn on Air tube, set the flow you are going to use for Air and N2(ox).

• Open the “ventil tider”, put a long time (99999) to Oxidation and start.

• Start heating by turning on the power for the oven and heating bands, the cooling will also

start. Choose temperature on oven.

• K3 to Inert, “Starta logging” in program (2, pressure as well).

• Control files are logging when red “5min” is shown.

• Turn on remaining tubes you are going to use, make sure yellow CO measurement device in

on.

• When desired temperature is reached, rejoice! When you start, set the oxidation time to some

time soon (like10 seconds forward in time) and press enter.

• Turn up reduction flow.

Turn off

• Turn off during an oxidation (or inert) phase, reduce fuel flow.

• Turn off yellow CO measurement device.

• Turn off oven and carefully open it a little using gloves.

• “Stoppa logging”, (2 files) save data on USB.

• Turn off cooling and heat bands.

• When cooled down (to at least 300degrees), reduce flow.

• Turn off program and computer.

• When cooled, take out reactor, measure the height, weight and save sample.

• Clean reactor

55

Cleaning the reactor

• Make sure that you have removed the sample from the reactor main body and weight it;

• Start by uncoupling the reactor upper and lower parts from the main part;

• Remove the silicone used to avoid leakages in the connections. Frist clean it with some

paper and then use a solvent for the remaining’s;

• Clean with water the glass stick of the lower and upper parts and dry them with a paper;

• For the main part start by washing it just water a few times. Then, use water with soap and a

brush to reach the interior walls. Be careful, not to damage the reactor’s bed. Wash it again

with water to remove the soap from the interior;

• Take the reactor to the hotte and make sure the reactor bottom end is closed with a glass

stopper;

• Pour a bit of hydrochloric acid (≈50 ml) inside the reactor, cover the upper opening with a

cover connected to a tube and introduce the other end of the tube inside a recipient (beaker)

with water. Turn on the heater to boil the acid.

• When the acid is boiling, turn off the heater and remove the tube from the water, to avoid

suction of water to the interior of the reactor as the acid cools down;

• Pour the acid into a recipient and wash it with water inside the hotte;

• Wash the reactor with water from the water the water company and then with distilled water.

• Dry it out with compressed air before store the reactor.

56

Appendix B – Experimental equipment specifications

57

Equipment Brand Model Accuracy Obs.

Quartz reactor Custom Made -

820 mm long

22 mm wide

porous plate at 370 mm

from the bottom

Oven ElectroHeat Custom Made - -

Cooler M&C Peltier coolers

ECP1000 - -

Heating bands

system Termonic Serie 16150 -

1,5 kW

-15°C…+150°C

58

Mass Flow

regulation

console

unknown unknown unknown -

Mass Flow

controller Brooks 5800S Series ± 0,7% Several working ranges

Cycle regulation Parker

131V5406-

2995-

481865C1

-

Direct operated (3 way

corrosion resistant)

solenoid valve

Steam Generator Cellkraft E-1000 -

100% steam

0…18 g/min

≈150°C

1 atm

59

Valves Swagelok Ball - 2 or 3 way valves

Thermocouples -

K Type unknown unknown ± 2,2°C or 0,75% -

Analyser Rosemount

Analytical NGA 2000

H2: ≤ 2% Thermal conductivity

CH4: ≤ 1% Infrared spectra

CO2: ≤ 1% Infrared spectra

CO: ≤ 1% Infrared spectra

O2: ≤ 1% Para-magnetism

Typically, errors caused by the accuracy of instruments are smaller than errors inherent to the whole process (interferences, answering time, gas leakage,

pulsing steam flow, etc.). For a standard measurement of one gas concentration in N2, the analyser could be easily calibrated with a very high accuracy.

60

Appendix C – Experimental results for 5g N4MZ1400 with 10g sand

61

Figure 23 – Experimental results for 5g N4MZ1400 with 10g sand at 600°C and 1 atm in the fuel reactor, with a H2O/CH4

ratio of 2

Figure 24 - Experimental results for 5g N4MZ1400 with 10g sand at 650°C and 1 atm in the fuel reactor, with a H2O/CH4

ratio of 2

0

10

20

30

40

50

60

70

17500 17750 18000 18250 18500 18750 19000

Vo

lum

etr

ic c

on

cen

tra

tio

n [

%]

Time reference [s]

H2

CO

CO2

CH4

0

10

20

30

40

50

60

70

14500 14750 15000 15250 15500 15750 16000 16250 16500

Vo

lum

etr

ic c

on

cen

tra

tio

n [

%]

Time reference [s]

H2

CO

CO2

CH4

62

Figure 25– Experimental results for 5g N4MZ1400 with 10g sand at 700°C and 1 atm in the fuel reactor, with a H2O/CH4

ratio of 2


ratio of 2

0

10

20

30

40

50

60

70

11400 11600 11800 12000 12200 12400 12600 12800 13000

Vo

lum

etr

ic c

on

cen

tra

tio

n [

%]

Time reference [s]

CO2

CO

CH4

H2

0

10

20

30

40

50

60

70

8800 9300 9800 10300

Vo

lum

etr

ic c

on

cen

tra

tio

n [

%]

Time reference [s]

CO2

CO

CH4

H2

63

Appendix D – Experimental results for 5g N4MZ1400 with 10g CaO

64

Figure 27– Experimental results for 5g N4MZ1400 with 10g CaO at 600°C and 1 atm in the fuel reactor, with a H2O/CH4

ratio of 2


ratio of 2

0

10

20

30

40

50

60

70

80

7000 7250 7500 7750 8000 8250

Vo

lum

etr

ic c

on

cen

tra

tio

n [

%]

Time reference [s]

H2

CO

CO2

CH4

0

10

20

30

40

50

60

70

80

9200 9450 9700 9950 10200 10450

Vo

lum

etr

ic c

on

cen

tra

tio

n [

%]

Time reference [s]

H2

CO

CO2

CH4

65


ratio of 2

Figure 30 - – Experimental results for 5g N4MZ1400 with 10g CaO at 750°C and 1 atm in the fuel reactor, with a

H2O/CH4 ratio of 2

0

10

20

30

40

50

60

70

13100 13600 14100 14600 15100

Vo

lum

etr

ic c

on

cen

tra

tio

n [

%]

Time reference [s]

CO

CO2

CH4

H2

0

10

20

30

40

50

60

70

15400 15600 15800 16000 16200 16400 16600 16800

Vo

lum

etr

ic c

on

cen

tra

tio

n [

%]

Time reference [s]

CO

CO2

CH4

H2

66

Appendix E – Experimental results for 15g N4MZ1400 with 30g

sand

67


ratio of 2


ratio of 2

0

10

20

30

40

50

60

70

80

90

100

18800 19300 19800 20300 20800 21300

Vo

lum

etr

ic c

on

cen

tra

tio

n [

%]

Time reference [s]

CO

CO2

CH4

H2

0

10

20

30

40

50

60

70

80

90

100

15400 15900 16400 16900 17400 17900

Vo

lum

etr

ic c

on

cen

tra

tio

n [

%]

Time reference [s]

CO

CO2

CH4

H2

68


ratio of 2


ratio of 2

0

10

20

30

40

50

60

70

80

90

100

11900 12400 12900 13400 13900 14400

Vo

lum

etr

ic c

on

cen

tra

tio

n [

%]

Time reference [s]

CO

CO2

CH4

H2

0

10

20

30

40

50

60

70

80

90

100

8200 8700 9200 9700 10200 10700

Vo

lum

etr

ic c

on

cen

tra

tio

n [

%]

Time reference [s]

CO

CO2

CH4

H2

Documents

Thermodynamic analysis of hydrogen production via …...reforming (in this case methane reforming) by chemical-looping auto-thermal reforming, with water-gas shift and carbonation