Modelling and experimental validation of an alkaline electrolysis cell for hydrogen production

Patrícia Alexandra Varela Pinto

Thesis to obtain the Master of Science Degree in

Mechanical Engineering

Supervisors: Prof. Ana Sofia Oliveira Henriques Moita

Dr. Rui Pedro da Costa Neto

Examination Committee

Chairperson: Prof. Edgar Caetano Fernandes

Supervisor: Dr. Rui Pedro da Costa Neto

Member of the Committee: Prof. Alda Maria Pereira Simões

January 2021

i

Dedication

To my father José Alberto who prematurely passed away while I was developing this thesis and never

got the chance to see me graduate. He surely was my biggest fan and my most avid supporter and

would be very proud to see what I have accomplished. A wholehearted thank you to you, dad.

ii

Acknowledgements

Undertaking this project during this year was quite challenging and the guidance and support I received

from many people were fundamental to the conclusion of this chapter of my life.

To my supervisors, Professor Rui Costa Neto and Professor Ana Moita, for their relentless guidance and

invaluable advices during this project.

To professor Fernanda Margarido, who kindly allowed the access to the laboratory for the realization of

this project.

To Dr. Nick Valckx from Agfa for providing me with Zirfon which was essential for the construction of the

lab-scale electrolyser.

A wholehearted thanks to my friends Ana Margarida Coelho, André Coelho, Inês Victorino and João

Testas who helped me make the best of my time at IST.

To my boyfriend Marc for all the support, the companionship, the optimism and the love he has never

failed to provide me with over the last 5 years.

To my cousins Liliane and João, my best friends of a lifetime, for always supporting and cheering my

victories from halfway across the globe.

To my brother Hugo who was always there for me during every step of the way.

To mother Helena for her endless love and support. Thank you for guiding me through the hardest times

of my journey at IST. Thank you for your appreciation and kind words when I needed them the most.

Without you, I would not be here today. We did it, mom!

To all of you, thank you again.

iii

Resumo

A produção de hidrogénio através de electrólise para armazenar o excedente de energia eléctrica

renovável é actualmente considerada uma das tecnologias mais promissoras para permitir a integração

adequada de energias renováveis na rede. Sendo as células alcalinas de electrólise a tecnologia

electrolítica com maior maturidade e o maior alcance comercial, o objectivo desta tese é construir um

modelo unidimensional em estado estacionário de uma célula de electrólise alcalina, modelando os

processos físicos envolvidos no processo de electrólise e implementando-o em MATLAB. O modelo

produzirá as curvas de polarização do electrolisador que permitirão não só simular o comportamento

do electrolisador sob certas condições de trabalho, mas também estudar a sensibilidade da curva de

polarização a alguns dos parâmetros dos electrolisadores e analisar a contribuição mais relevante para

o sobrepotencial da célula. Este modelo foi validado utilizando três conjuntos de dados experimentais:

dois conjuntos disponíveis na literatura e um conjunto de um electrolisador à escala de laboratório

construído para o efeito. Ao comparar os dados experimentais com os resultados da simulação

produzida pelo modelo, verificou-se que o modelo simula a realidade com muito boa precisão,

apresentando desvios inferiores a 1% na maioria dos casos.

Palavras-chave: Electrolizador Alcalino, Produção de Hidrogénio, Sobrepotenciais, Curvas de

Polarização, Simulação

iv

v

Abstract

The production of hydrogen through electrolysis to store surplus electric energy is currently considered

one of the most promising technologies to allow proper integration of renewable energies in the grid.

Being alkaline electrolysis cells the electrolytic technology with greater maturity and the larger

commercial outreach, the aim of this thesis is to build a one-dimensional steady state model of an alkaline

electrolysis cell by modelling the physical processes involved in the electrolysis process and implement

it in MATLAB. The model will produce the electrolyser’s polarization curves which will not only allow to

simulate the electrolyser’s behaviour under certain working conditions but also study the polarization

curve sensibility to some of the electrolysers parameters and analyse the most relevant contribution to

the cell’s overpotential. This model was validated using three sets of experimental data: two sets

available in the literature and one set from a lab-scale electrolyser built for this purpose. By comparing

the experimental data with the simulation results produced by the model, it was found that the model

simulates the reality with very good accuracy presenting deviations below 1% in most cases.

Keywords: Alkaline Electrolyser, Hydrogen Production, Overpotentials, Polarization Curves, System

Simulation

vi

vii

Contents

Dedication ________________________________________________________________ i

Acknowledgements ________________________________________________________ ii

Resumo _________________________________________________________________ iii

Abstract __________________________________________________________________ v

List of Figures _____________________________________________________________ x

List of Tables _____________________________________________________________ xii

Nomenclature ____________________________________________________________xiii

Glossary _________________________________________________________________ xv

1 Introduction ___________________________________________________________ 1

1.1. Motivation _________________________________________________________ 1

1.2. Objectives _________________________________________________________ 2

1.3. Thesis Outline ______________________________________________________ 3

2 Background ___________________________________________________________ 5

2.1. Renewable energy penetration on electrical networks ___________________ 5 2.1.1. Energy Storage on Electrical Networks _____________________________ 6

2.2. Hydrogen for Energy Storage ________________________________________ 8 2.2.1. Physical and Chemical Properties of Hydrogen ______________________ 9 2.2.2. Current uses and production ____________________________________ 10

2.3. Water Electrolysis __________________________________________________ 11 2.3.1. Fundamentals of water electrolysis _______________________________ 12 2.3.2. Types of electrolysers __________________________________________ 13

2.3.2.1. Alkaline Electrolysis Cell (AEC) _________________________________ 13 2.3.2.2. Proton Exchange Membrane Electrolysis Cell (PEMEC) ___________ 14 2.3.2.3. Solid Oxide Electrolysis Cell (SOEC) ____________________________ 16 2.3.2.4. Molten Carbonate Electrolysis Cell (MCEC) ______________________ 17

2.3.3. Water Electrolysis Powered by Intermittent Energy ___________________ 18 2.3.3.1. Power Electronics and Process Control _________________________ 18 2.3.3.2. Impacts on Dynamic Operation Requirements ____________________ 19 2.3.3.3. Impacts on Hydrogen Production Characteristics and Efficiency ____ 20 2.3.3.4. Impacts on Reliability and Durability and Solution Approaches ______ 23

2.4. Alkaline Water Electrolysis Fundamentals _____________________________ 24

viii

2.4.1. Conventional vs Zero-Gap _______________________________________ 25 2.4.2. Monopolar vs Bipolar ___________________________________________ 26 2.4.3. Electrolyte ____________________________________________________ 27 2.4.4. Electrodes and Electrocatalysts __________________________________ 28

2.4.4.1. Cathode ____________________________________________________ 29 2.4.4.2. Anode ______________________________________________________ 31

2.4.5. Diaphragms ___________________________________________________ 33 2.4.5.1. Inorganic Materials ___________________________________________ 34 2.4.5.2. Organic Materials __________________________________________ 37

3 Model Development ___________________________________________________ 40

3.1. Electrodes module _________________________________________________ 40 3.1.1. Cathode Module _______________________________________________ 40 3.1.2. Anode Module _________________________________________________ 42

3.2. Voltage module ____________________________________________________ 42 3.2.1. Reversible Potential ____________________________________________ 43 3.2.2. Activation overpotential _________________________________________ 43 3.2.3. Concentration overpotential _____________________________________ 45 3.2.4. Ohmic overpotential ____________________________________________ 45

3.2.4.1. Electrodes __________________________________________________ 47 3.2.4.2. Electrolyte ___________________________________________________ 47 3.2.4.3. Separator ___________________________________________________ 48

3.3. Auxiliary Modules __________________________________________________ 49 3.3.1. Diffusion ______________________________________________________ 49 3.3.2. Electrolyte Saturated Vapor Pressure _____________________________ 50

3.3.2.1. Potassium Hydroxide _________________________________________ 51 3.3.2.2. Sodium Hydroxide ___________________________________________ 51

3.3.3. Thermodynamic Water Activity ___________________________________ 51 3.3.3.1. Potassium Hydroxide _________________________________________ 51 3.3.3.2. Sodium Hydroxide ___________________________________________ 52

3.3.4. Electrolyte Electrical Conductivity ________________________________ 52 3.3.4.1. Potassium Hydroxide _________________________________________ 52 3.3.4.2. Sodium Hydroxide ___________________________________________ 52

3.4. Cell Efficiencies ___________________________________________________ 53

4 Experimental Approach ________________________________________________ 55

4.1. Construction of the Lab-Scale Electrolyser ____________________________ 55 4.1.1. Experimental Setup and Procedure _______________________________ 57

4.1.1.1. Experimental Procedure ______________________________________ 59

5 Results and Discussion _________________________________________________ 60

ix

5.1. Model Validation with Data from the Literature _________________________ 60 5.1.1. HRI Electrolyser ________________________________________________ 60 5.1.2. PHOEBUS Electrolyser _________________________________________ 63

5.2. Model Validation with Data from a Lab-scale Electrolyser ________________ 66 5.2.1. Physical Parameters and Model Validation _________________________ 66

6 Conclusions __________________________________________________________ 73

6.1. Achievements _____________________________________________________ 73

6.2. Future Work ______________________________________________________ 74

Bibliography ______________________________________________________________ 75

Appendix A ______________________________________________________________ 81

A1. HRI Electrolyser Experimental Data ____________________________________ 81

A2. PHOEBUS Electrolyser Experimental Data ______________________________ 82

A3. Lab-Scale Electrolyser Experimental Data _______________________________ 84

Appendix B ______________________________________________________________ 85

B1. HRI Electrolyser Polarization Curves ____________________________________ 85

B2. PHOEBUS Electrolyser Polarization Curves _____________________________ 86

B3. Lab-scale Electrolyser Polarization Curves ______________________________ 87

x

List of Figures

Figure 1 - Current Uses Of Hydrogen Worldwide adapted from [12]10

Figure 2 - Current Production of Hydrogen Worldwide adapted from [10]. _______________ 11

Figure 3 – ΔG(T), ΔH(T), and TΔS(T) of the water-splitting reaction at P=1 bar [2]. _______ 13

Figure 4 - Conventional Cell Design vs Zero-Gap Cell Design [17]______________________ 26

Figure 5 - Monopolar (a) and Bipolar (b) electrolyser configurations [27] ________________ 27

Figure 6 - KOH and NaOH Specific Conductivity at 60°C and 80°C (based on [23]) _______ 28

Figure 7 - Potential for hydrogen production as a function of temperature [22] ___________ 53

Figure 8 - Lab-scale Electrolyser 3D Model _________________________________________ 56

Figure 9 - 3D Model Detail of the Electrical Connection Mechanisms ___________________ 56

Figure 10 - Connecters, Inlets and Outlets (Lab-scale Electrolyser Top View) ____________ 57

Figure 11 - Connections, Inlets and Outlets _________________________________________ 57

Figure 12 - Schematic Representation of the Experimental Setup ______________________ 58

Figure 13 - Experimental Setup ___________________________________________________ 58

Figure 14 - Memmert U40 Incubator _______________________________________________ 58

Figure 15 - Zirfon® Separator areal ionic resistance in 30 wt.% KOH: experimental data by

Agfa [77] and fitting curve. ________________________________________________________ 62

Figure 16 – HRI electrolyser polarization curves at different temperatures with corresponding

experimental data [23]. ___________________________________________________________ 63

Figure 17 – PHOEBUS electrolyser polarization curves at different temperatures with

corresponding experimental data [104]. ____________________________________________ 65

Figure 18 - Difference between assuming a zero-gap configuration and assuming a 1 mm gap

_______________________________________________________________________________ 67

Figure 19 - Zirfon® Ionic Resistance in 20 wt.% NaOH vs Zirfon® Ionic Resistance in 30 wt.%

KOH __________________________________________________________________________ 68

Figure 20 - Comparison between several limiting current densities at a working temperature

of 60 ºC _______________________________________________________________________ 68

xi

Figure 21 - Influence of the charge transfer coefficient on the Polarization Curve _________ 69

Figure 22 - Anodic and Cathodic Contribution to the Activation Overpotential at 60 ºC ____ 69

Figure 23 – Lab-scale electrolyser polarization curves at different temperatures with

corresponding experimental data. _________________________________________________ 70

Figure 24 - Overpotential Contribution to the total Cell Voltage ________________________ 71

Figure 25 - Contributions to the Ohmic Overpotential at 60 ºC _________________________ 71

Figure 26 - Contributions to the Anodic and Cathodic Activation Overpotentials at 60 ºC __ 72

xii

List of Tables

Table 1 - Characteristics of Energy Storage Technologies [7] ___________________________ 7

Table 2 - Selected physical Properties of Hydrogen [2], [4] _____________________________ 9

Table 3 - Specific Energy and Energy Density of different fuel types [4], [11] _____________ 10

Table 4 – Summary of cathode electrocatalyst performance __________________________ 31

Table 5 - Summary of anode electrocatalyst performance ____________________________ 33

Table 6 - Physical Properties of pure NiO-diaphragm [68] _____________________________ 36

Table 7 - Zirfon® PERL properties [74], [75], [77] ____________________________________ 38

Table 8 - HRI Electrolyser information [6], [23], [74], [82] ______________________________ 61

Table 9 - Fitted parameters for the HRI electrolyser's anode and cathode. _______________ 61

Table 10 - Maximum and average deviations between HRI electrolyser experimental data and

the model at each temperature. ___________________________________________________ 63

Table 11 - PHOEBUS electrolyser information [5], [6], [23], [103] _______________________ 64

Table 12 – Fitted Parameters for the PHOEBUS electrolyser’s anode and cathode. _______ 65

Table 13 – Maximum and average deviations between the PHOEBUS electrolyser

experimental data and the model at each temperature. _______________________________ 65

Table 14 - Lab-scale Electrolyser summarized information ____________________________ 66

Table 15 - Fitted Parameters for the Lab-scale electrolyser anode and cathode. _________ 70

Table 16 - Maximum and average deviations between the lab-scale electrolyser experimental

data and the model at each temperature. ___________________________________________ 70

xiii

Nomenclature

Latin Symbols

𝑨 Tafel Slope (mV/dec)

𝑨𝒆 Electrode Geometric Area (cm2)

𝑨𝒔 Separator Geometric Area (cm2)

𝒂 Thermodynamic Activity (bar)

𝑪 Concentration (mol/dm3)

𝑫 Diffusion Coefficient

𝒅 Density (kg/m3)

𝑭 Faraday Constant (96 485.3329 s.A/mol)

𝑯 Enthalpy (kJ/mol)

𝑰 Current (A)

𝒊 Current Density (A/cm2)

𝒊𝟎 Exchange Current Density (A/cm2)

𝒊𝒍𝒊𝒎 Limiting Current Density (A/cm2)

𝒍𝒂𝒏 𝒄 Distance between the anode channel and the catalyst (cm)

𝒍𝒄𝒂𝒕 𝒄 Distance between the anode cathode and the catalyst (cm)

𝒍𝒂𝒏 𝒔 Distance between the anode and the separator (cm)

𝒍𝒄𝒂𝒕 𝒔 Distance between the cathode and the separator (cm)

𝑴 Molecular Weight (kg/kmol)

Molar Flux (mol/s)

𝑵𝒄𝒆𝒍𝒍𝒔 Number of Electrolytic Cells in the Electrolyser

Molar Flow Rate (mol/s.cm2)

𝑷 Total Pressure (Pa)

𝒑 Partial Pressure (Pa)

𝑹 Resistance (Ω)

𝑹𝒈𝒂𝒔 Universal Gas Constant (8.3144 J/mol.K)

𝒓 Mean Pore Radius (m)

𝑺 Entropy (kJ/mol K)

𝑽 Potential (V)

𝑽𝒓𝒆𝒗 Reversible Potential (V)

𝑻 Temperature (K)

𝑻𝒓𝒆𝒇 Reference Temperature (K)

𝒀 Molar Fraction

xiv

Greek Symbols

𝜶 Charge Transfer Coefficient

𝜷 Bubble Zone Thickness (cm)

𝚫 Variation

𝜹𝒂𝒏 Anode Thickness (cm)

𝜹𝒄𝒂𝒕 Cathode Thickness (cm)

𝜹𝒔 Separator Thickness (cm)

𝜺 Lennard-Jones Energy

𝝐𝒂𝒏 Anode Porosity

𝝐𝒄𝒂𝒕 Cathode Porosity

𝝐𝒔 Separator Porosity

𝜼 Overpotential (V)

𝜼𝒂𝒄𝒕 Activation Overpotential (V)

𝜼𝒄𝒐𝒏 Concentration Overpotential (V)

𝜼𝒐𝒉𝒎 Ohmic Overpotential (V)

𝚯 Bubble Coverage

𝚱 Specific Conductivity (S/cm)

𝜿 Temperature Coefficient of Resistivity (K-1)

𝝀 Mean Free Path (m)

𝝆 Resistivity (Ω/cm)

𝝈 Mean Molecular Radii (Å)

𝝉𝒂𝒏 Anode Tortuosity

𝝉𝒄𝒂𝒕 Cathode Tortuosity

𝝉𝒔 Separator Tortuosity

𝝋 Volume Fraction

𝝌 Molar Fraction

𝛀𝑫 Dimensionless Diffusion Collision Integral

𝝎𝒔 Wettability Factor

xv

Glossary

AEC Alkaline Electrolysis Cell

BoP Balance of Plant

DC Direct Current

HER Hydrogen Evolution Reaction

HRI Hydrogen Research Institute

IEA International Energy Agency

MCEC Molten Carbonate Electrolysis Cell

MIE Minimum Ignition Energy

NTP Normal Temperature and Pressure

OER Oxygen Evolution Reaction

ORR Oxygen Reduction Reaction

P2H Power to Hydrogen

PEM Proton Exchange Membrane

PEMEC Proton Exchange Membrane Electrolysis Cell

PHES Pumped Hydro Energy Storage

PSU Polysulfone

PTFE Polytetrafluoroethylene

PVP Polyvinylpyrrolidone

RES Renewable Energy Sources

RHE Reversible Hydrogen Electrode

SMES Superconducting Magnetic Energy Storage

SNG Synthetic Natural Gas

SOEC Solid Oxide Electrolysis Cell

SPE Solid Polymer Electrolyte

STP Standard Temperature and Pressure

xvi

1

1 Introduction

1.1. Motivation

The energy sector is the biggest source of greenhouse gas emissions. According to the International

Energy Agency (IEA) [1], in 2018, global energy consumption grew by 2.3%, almost twice the average

growth rate since 2010. Fossil fuels were responsible for almost 70% of the global energy consumption

increase for the second consecutive year. Despite the significant growth of renewables, it was not

enough to compensate for the increased world electricity demand. The increasingly high energy demand

was responsible for a rise of 1.7% in the CO2 emissions to a historical value of 33.1 Gt CO2. This

represents the highest growth rate since 2013, 70% higher than the average increase since 2010, being

the power sector responsible for nearly two-thirds of emissions growth. While Germany, Japan, Mexico,

France, and the United Kingdom reduced their CO2 emissions, countries like China, India, and the United

States accounted for 85% of the net increase. In 2018, the global annual average concentration of CO2in

the atmosphere was 407.4 ppm, an increase of 2.4 ppm when compared to 2017. The constant increase

in these concentrations is very worrisome because it represents a major increase from pre-industrial

levels which were below 280 ppm. [1]

In order to avoid major anthropogenic interference in the climate system and the Earth’s ecosystems, it

is mandatory that the world shifts to cleaner energy sources that emit significantly less pollutants.

Renewable energies play a fundamental role in achieving those goals however, renewable energies are

intermittent in time and space and the energy production from such sources is difficult to forecast with

a good level of accuracy [2]. To maintain the balance in the electrical network, the total electric energy

injected into the grid must always be equal to the energy consumption. Not being able to forecast the

energy production with enough accuracy may introduce a difference between energy production and

consumption, which results in a variation of the network frequency from the nominal frequency. This

frequency deviation can cause damage to electrical devices since they are designed to work at the

nominal frequency [2].

To achieve proper integration of renewable energy sources (RES) on the electrical network, several

solutions have been identified. The most promising solutions can be categorized into actions on the

structure of electrical networks, management of power sources and electrical loads and introduction of

energy storage systems.

The introduction of energy storage systems in the electrical network to maintain the network’s balance

consists of the introduction of energy storage systems to absorb the excess available power when the

supply surpasses the demand and redeliver it when the power supply cannot satisfy the demand. This

solution appears to be the most advantageous one since it does not depend on political conditions and

2

does not imply losing part of the potential energy that could be produced on a given site. Currently,

there is a wide set of energy storage technologies available however not all of them are suitable for this

kind of application. Technologies like supercapacitors and flywheels can be used to store reduced

amounts of power during short periods of time and redeliver it quickly which makes them unsuitable for

the purpose in question. Batteries, for example, may seem a good option at first, but their low lifespan,

high cost, low energy density and high risk of environmental contamination pose major setbacks for its

widespread implementation as a renewable energy storage system. Nowadays, the most viable

technology for the storage of large amounts of renewable energy is pumped hydro energy storage

(PHES) yet this technology has big disadvantages related to its mandatory need for specific geographical

conditions and the need for good geopolitical conditions [3]. Currently, hydrogen is considered the most

promising technology for future large-scale energy storage since it is the most abundant element in the

Universe, it has the highest energy density per unit mass, it is sustainable and non-toxic, and it can be

produced from both renewable and non-renewable energy sources. The concept of using hydrogen as

an energy carrier dated to over two centuries ago but re-emerged during the 1970s energy crisis,

undergoing significant technological advances in the 1980s [4].

Given this picture, the use of hydrogen as an energy storage means to store the surplus energy

produced from renewables to be used when the energy demand surpasses the energy production is a

crucial step forward in the integration of renewable energy on the electrical network. Besides, hydrogen

can also be used as a clean fuel with water being the only exhaust product from its conversion back to

energy.

Being alkaline water electrolysis a very well-established technology it is seen as one of the most

promising methods to produce hydrogen from the surplus renewable electric energy. However, some

issues still need to be overcome. The physical modelling of electrolysis cells as well as the research for

new materials to increase the cell efficiency are some of the most important research topics in AEC

technology [2], [5], [6].

1.2. Objectives

In line with the arguments presented in the motivation, the aim of this thesis is to build a steady state

model of an alkaline electrolysis cell by modelling the physical processes involved in the electrolysis

process and implement it in MATLAB. The model will produce the electrolyser’s polarization curves

which will not only allow to simulate the electrolyser’s behaviour under certain working conditions but

also study the polarization curve sensibility to some of the electrolysers parameters and analyse the

most relevant contribution to the cell’s overpotential.

3

1.3. Thesis Outline

This thesis is divided into six main sections, including the present Introduction which provides the

motivation for this work, the main objectives and the thesis outline, Background, Model Development,

Experimental Approach, Results and Discussion, and Conclusions.

In the Background, a review of several topics surrounding the production of hydrogen from renewable

sources is done, followed by a state of the art on the main topic related to alkaline water electrolysis.

In the third section, a detailed description of the physics model along with its modules and submodules

is presented. This section also includes a subsection with auxiliary modules that are not crucial to the

model itself but are important to compute the influencing parameters that important in the model.

The fourth section is dedicated to the detailed description of the construction of a lab-scale electrolyser

built with the purpose of validating the physics model and the corresponding experimental procedure

used to collect the experimental data to compare with the model.

In the fifth section, the model validation is performed. This validation uses three sets of experimental

data: two reported in the literature and one set of experimental data obtained from a lab-scale

electrolyser built for this purpose. In this chapter, the results produced by the model are displayed and

analysed, and a comparison between the experimental data and the data computed by the model is

made.

This report finally ends with the sixth section where the most important conclusions withdrawn from this

work are reported and briefly discussed. This final section ends with some recommendations for future

work.

4

5

2 Background

This chapter provides the main concepts that are required to assimilate and understand in order to be

able to follow the analysis of the results.

In the first section, an assessment of the setbacks and possible solutions for the integration of renewable

energies on electrical networks is performed. The second section provides an overview of the use of

hydrogen for energy storage and the current worldwide hydrogen production and uses. The third

subsection is dedicated to a state-of-the-art review of water electrolysis comprising the main types of

electrolysers, issues related to powering water electrolysis with renewable energies. The fourth

subsection is dedicated to the fundamentals of the main topic of this thesis: Alkaline Electrolysis Cells.

2.1. Renewable energy penetration on electrical networks

Problems related to climate change, environmental pollution and the security of energy supply have led

countries to reconsider their energy production means and to analyze their energy dependence on other

countries. Shifting the country’s energy production from fossil fuels to RES is currently considered the

best solution to these problems. However, the intermittency of RES and our limited capacity to forecast

their behaviour raises some issues regarding the broad implementation of renewable energies as a

replacement for fossil fuel-derived energies. These challenges become especially relevant when the

share of renewable power production becomes relevant in the global mix production. The maximum

share of intermittent sources in a certain global mix strongly depends on the characteristics of the

network, the other production means, and the meteorological conditions. Usually, without case-specific

assessment, an upper value of 30% is currently taken as a limit.

In order to maintain the electrical network balanced, at each moment, the total electricity production

injected into the grid must be equal to the total electricity consumption. When the power production and

the power consumption do not match, the network frequency deviates from its nominal frequency (50Hz

or 60Hz) and it keeps deviating until de equilibrium is re-established. Operating electrical devices at

different frequencies leads to premature fatigue and failure. Thus, maintaining the network frequency at

values very close to its nominal value is crucial and is one of the quality requirements of the supplied

electrical energy. To maintain the frequency around its nominal value, the production levels must follow

the consumption at each instant which requires power production units whose outputs can be totally

controllable allowing a rapid power production increase or decrease according to the load level.

Intermittent power production does not follow these requirements therefore, its proper integration in the

electrical grid is strongly dependent on the ability of the other production units of the global mix to handle

this lack of flexibility.

6

Increasing the injection of renewable power into the electrical network reduces operation flexibility which

is crucial to ensure the reliability of the electrical system. The most promising solutions for the flexibility

problem of intermittent power production can be categorized into (1) actions on the structure of electrical

networks, (2) management of power sources and (3) electrical loads and introduction of energy storage

systems.

The first solution consists in reinforcing and further developing the electrical grid in order to mutualize

renewable energy production over a wide region, such as Europe, covered by different meteorological

conditions. Within wide regions, when renewable power production is not possible in one location, it may

be possible somewhere else. Mutualizing these sources can only be achieved by reinforcing existing

grid infrastructures to increase the capacity of transmission lines and increasing interconnection

capacities between countries using high voltage direct current (HVDC) lines. The concept of largely

interconnected networks is often designated by “super grid”. These types of measures require political

and economic agreements between several countries which, in some cases, may be difficult to reach.

Management of power sources and electrical loads is another way of increasing operation flexibility. It

consists in changing the operating point of PV modules and wind turbines to intentionally reduce the

amount of power delivered in order to match the power demand. Power production from RES is generally

operated to deliver the maximum amount of power it can provide based on the site’s meteorological

conditions. Changing the operating point of the devices allows at each moment to increase or decrease

the production according to the fluctuations of the load however, this implies losing part of the potential

energy of a given site.

The introduction of energy storage systems in the electrical network to maintain the network’s balance

consists of the introduction of energy storage systems to absorb the excess available power when the

supply surpasses the demand and redeliver it when the power supply cannot satisfy the demand. This

solution appears to be the most advantageous one since it does not depend on political conditions and

does not imply losing part of the potential energy that could be produced on a given site.

2.1.1. Energy Storage on Electrical Networks

Energy storage is seen as a great means to increase the penetration level of intermittent renewables

into the grid. Energy storage systems can play several roles in the integration of renewable energy in

the electrical grid being peak production shaving, production smoothing, and production shifting the

three main ones.

Peak production shaving consists in storing the energy produced above a certain threshold allowing to

keep PV modules and wind turbines always working at their optimal production points maximizing the

power production.

Production smoothing consists in using a storage system in charge or discharge, alternately, to avoid

fast fluctuating wind or PV power injection into the grid to relieve other production units of their obligation

to provide power reserve.

7

Production shifting consists in storing the energy produced at a given moment in time to discharge this

energy later. This function would be particularly useful in energy mixes with large shares of renewable

energy allowing, for example, the storage of PV energy during summer when the PV production is

maximal to be used during winter when less PV energy is available.

Electricity itself cannot be stored and therefore, it needs to be converted into another form of energy so

that it can be used later. Storing electricity requires an energy carrier to hold part of the energy content

present in electrical energy. Presently, the most used energy carriers are electricity itself, natural gas,

and oil-derived products. The extensive use of natural gas and oil-derived products as energy carriers

has brought the world severe consequences, and for this reason, shifting to cleaner energy carriers is

mandatory.

All energy storage technologies (batteries, PHES, CAES, etc.) use energy carriers. Given the diversity

of storage technologies available on the market and in development, the practical question of choosing

the appropriate technology for a given storage application emerges. Energy capacity, maximal charge

and discharge powers, self-discharge, efficiency, aging (calendar and cycling), costs, specific energy,

specific power, response time, grey energy, and technology maturity are some of the most important

selection criteria that must be taken into account when choosing the appropriate storage technology for

a given application [2].

Table 1 - Characteristics of Energy Storage Technologies [7]

Tech. Power Rating (MW)

Storage Duration (h)

Cycling/ Lifetime

Self-discharge (%)

Energy Density (MWh/l)

Power Density (W/l)

Efficiency (%)

Response time

Super-Capacitor

0.01-1 ms-min 10,000-100,000

20-40 10-20 40,000-120,000

80-98 10-20 ms

SMES 0.1-1 ms-min 100,000 10-15 ~6 1,000-4,000

80-95 <100 ms

PHES 100-1,000 4-12 h 30-60 years ~0 0.2-2 0.1-0.2 70-85 sec-min

CAES 10-1,000 2-30 h 20-40 years ~0 2-6 0.2-0.6 40-75 sec-min

Flywheels 0.001-1 sec-hours 20,000-100,000

1.3-100 20-80 5,000 70-95 10-20 ms

NaS battery

10-100 1 min-8 h 2,500-4,400 0.05-20 150-300 120-160 70-90 10-20 ms

Li-ion battery

0.1-100 1 min-8 h 1,000-10,000

0.1-0.3 200-400 1,300-10,000

85-98 10-20 ms

Flow battery

0.1-100 1-10h 12,000-14,000

0.2 20-70 0.5-2 60-85 10-20 ms

Hydrogen 0.01-1,000 min-weeks

5-30 years 0-4 6001 0.2-20 25-45 sec-min

SNG 50-1,000 hours-weeks

30 years ~0 1,8001 0.2-2 25-50 sec-min

1 Energy density at 200 bar

8

Pumped hydro energy storages (PHES) and lead-acid batteries have been the technologies of choice

for decades. Currently, PHES is the most used energy storage technology worldwide accounting for

around 99.3% of the total storage capacity [2], [8], [9] (Table 1). Battery technologies and compressed

air energy storage (CAES) account for the most part of the remaining share representing 0.36% and

0.31%, respectively [2], [8].

Electricity and hydrogen, if produced from renewable energies, are a great alternative to natural gas and

oil-derived products since both of them can also be transported and are not polluting. The development

of hydrogen technologies constitutes a big step forward in the decarbonization of the energy sector

allowing the storage and transport of clean electrical energy.

2.2. Hydrogen for Energy Storage

Hydrogen can be used to store available electric energy through electrolysis in the so-called Power-to-

Hydrogen process (P2H). The electrolysis is an electricity-powered process that consists in breaking the

water molecule into molecular hydrogen and elemental oxygen. The hydrogen produced from

electrolysis is then stored either in gaseous, liquid, or solid state to be used later. The produced oxygen

is usually released to the atmosphere however it may also be stored for later use in several applications.

Electrolysis and hydrogen have the great advantage of allowing the separation between charging power

and stored energy, unlike some of the other energy storage technologies. In the case of battery

technologies, for example, each battery has a maximum charging power and a maximum energy

capacity which means that to increase the storage capacity, the number of batteries must be increased.

Since electrolysis and hydrogen rely on different objects, charging power (thought the electrolyser size)

and stored energy (through the hydrogen storage size) can be decoupled [2].

Unlike battery technologies or super-capacitors, hydrogen tanks do not suffer from self-discharge [2],

[7]. Furthermore, hydrogen in its gaseous state can benefit from the existing gas transportation and

distribution networks.

Hydrogen produced from renewable energies can be used to produce electrical energy, mechanical

energy, and thermal energy. The use of hydrogen directly in a fuel cell, which can be perceived as the

discharge phase, produces electricity and heat. The combination of an electrolyser, hydrogen storage

devices, and a fuel cell constitutes a fully reversible energy storage device and it is often named the

hydrogen chain [2]. Hydrogen can also be used in a motor or turbine to produce thermal and mechanical

energy which in turn can drive an alternator to produce electrical energy. Hydrogen produced from

renewable energies can also be combined with carbon dioxide to produce synthetic natural gas and

other synthetic liquid hydrocarbons such as diesel or kerosene [2], [10].

The decoupling between charging power and stored energy also applies to the production of electricity

from hydrogen since the power produced in a fuel cell or hydrogen turbine (discharging power) is also

totally independent of the amount of energy stored in a hydrogen tank. Consequently, the three main

9

components of the hydrogen chain (electrolyser, hydrogen storage system, and fuel cell) can be

dimensioned separately [2].

The major disadvantage of the hydrogen chain is the relatively low overall efficiency in comparison to

other technologies however, in most applications, efficiency is not the only and most important selection

criteria. Criterion such as self-discharge, environmental footprint, lifetime, reliability, and costs are also

very relevant in the overall evaluation of energy storage technologies.

2.2.1. Physical and Chemical Properties of Hydrogen

Hydrogen is a colourless, odourless, tasteless gas. The fact that hydrogen is the gas with the lowest

density (0.08999 g/L at STP; 0.08375 g/L at NTP) provides it with the ability to quickly disperse and

ascend to the upper atmosphere. Hydrogen changes from liquid to gas at -252.8ºC (~20.35K) and from

liquid to solid at -259.2ºC (~13.95K) at atmospheric pressure [10] (Table 2).

Molecular hydrogen dissociates into atomic hydrogen with a dissociation energy of approximately 435

kJ/mol [10].

Table 2 - Selected physical Properties of Hydrogen [2], [4]

Parameter Value Unit

Molecular weight 2.016 mol

Melting point 13.96 K

Boiling point at 1atm 14.0 K

Density solid at 4.2 K 89 g dm-3

Density liquid at 20.4 K 71 g dm-3

Gas density at 0ºC and 1 atm 0.0899 g dm-3

Gas thermal conductivity at 25ºC 0.185 W m-1 ºC-1

Gas dynamic viscosity at 25ºC and 1 atm 8.9×10-6 kg m-1 s-1

Autoignition temperature 858 K

Flammability in oxygen 4-94 %

Flammability in air 4-74 %

Hydrogen has the highest energy content per unit of mass (LHV≈120 MJ/kg, HHV≈142MJ/kg) (Table 3)

and the highest flammability limit range when compared to all the commonly used fuels [2], [10]. Although

its energy content per unit mass is the highest, its energy content per unit of volume is rather small

which gives rise to several issues concerning storage since the volume needed to store a certain amount

of energy in the form of hydrogen gas is much bigger than for conventional fuels. When mixed with

chlorine or air, hydrogen can spontaneously explode by spark, heat, or sunlight. The minimum ignition

energy (MIE) for hydrogen is very low being 0.017MJ in air and 0.0012MJ in oxygen at ambient

conditions. These values are much lower than MIE values for commonly used fuels which usually range

from 0.1 to 0.3MJ in air at ambient conditions [2]. The fact that hydrogen is an extremely flammable gas

allied to its high energy-to-weight ratio (about three times higher than gasoline, diesel, or kerosene)

10

makes it hazardous to handle. However, if kept in well-designed containers in properly ventilated areas,

the risk of reaching this limit is very low. Furthermore, hydrogen’s autoignition temperature is rather high

(858K) when compared to conventional fuels (~640K for ethanol, ~520K for gasoline, ~480K for diesel)

[2], [10].

Table 3 - Specific Energy and Energy Density of different fuel types [4], [11]

Fuel Types Specific Energy (MJ/kg) Energy density (MJ/L)

Diesel 45.6 38.6

Gasoline 46.4 34.2

Kerosene 42.8 33

LPG (propane) 49.6 25.3

Natural Gas 53.6 -

Methane 55.6 -

Ethanol 29.7 23.4

Liquid H2 (690 bar and 15ºC) 141.77 5.323

Hydrogen Gas 141.88 0.01188

2.2.2. Current uses and production

Nowadays, industrial plants produce around 60 million tons of hydrogen every year with an estimated

yearly demand increase of about 6% [11]. More than half of the world hydrogen production goes into

making ammonia (53%), a major component of fertilizers. The next biggest share of the world production

(20%) is used in the refining of crude oils. Around 7% goes into the production of methanol and the

remaining 20% are used in several other applications such as the pharmaceutical industry, metallurgy,

glass industry, electronics fabrication, cooling of thermal generators, food industry, high-temperature

flames, etc. [12] (Figure 1).

Figure 1 - Current Uses Of Hydrogen Worldwide adapted from [12]

Hydrogen gas is currently produced from a wide variety of primary sources such as natural gas, heavy

oil, water, and coal. Large-scale production of hydrogen gas currently relies on steam reforming from

methane/natural gas representing almost 50% of the world's hydrogen production. Less representative

11

methods of production comprise higher hydrocarbon reforming from refinery-chemical industrial off-

gases (30%), coal gasification (18%), water electrolysis (4%), and a very low fraction (~0.1%) from other

sources such as photodissociation, biomass gasification, direct thermal and catalytic splitting of water

and biological processes which are currently a hot research topic [10] (Figure 2).

Figure 2 - Current Production of Hydrogen Worldwide adapted from [10].

Aside from not being sustainable, hydrogen production from methane and other fossil fuels fails to

provide a solution to reduce greenhouse gas emissions. It is clear at this point that hydrogen production

can be environmentally friendly if we can assure that the primary resource used to produce hydrogen is

renewable. Currently, biomass approaches and water electrolysis are the primary sources of renewable

hydrogen [11].

Biomass, a product of photosynthesis, poses an alternative to fossil fuels for clean hydrogen production

as CO2-neutral production can be achieved through several processes such as biomass gasification,

pyrolysis of bio-oils, steam reforming of biomass delivered higher alkanes and alcohols, and aqueous

phase reforming of oxygenated hydrocarbons.

Solar energy is also seen as an almost ideal means for environmentally friendly hydrogen production

since it is a very abundant energy resource that may allow the production of hydrogen from water

through several different processes. Solar energy can be used in the form of heat in thermochemical

water splitting; in the form of light in photoelectrochemical or photocatalytic water splitting; or after

converted into electricity to power water electrolysis.

Hydrogen production through water electrolysis is a very promising method being considered by some

authors the only practical method for hydrogen large-scale production from renewable energy resources

such as solar photovoltaic, hydroelectric power, and wind energy.

2.3. Water Electrolysis

Water electrolysis consists in the decomposition of water into oxygen and hydrogen gas due to the

passage of a direct electric current (DC). It is a very mature and widely available technology, currently

12

responsible for around 4% of the world hydrogen production, that can be used to produce green

hydrogen if the electricity used to power the system can be assured to be produced from renewable

energy sources (solar, wind, hydro, etc.). Hydrogen production capacities range from a few cm3/h to

hundreds of m3/h depending on the characteristics of the electrolyser and operating conditions. In

general, water electrolysers have relatively good efficiencies (~60%), relatively low energy consumption

(~4 kWh/Nm3 H2) and are able to produce highly pure hydrogen since the produced gases are physically

separated by a membrane during their evolution at the electrodes. Water electrolysers operate with very

few moving parts, are usually compact and their maintenance costs are low. The production costs,

however, strongly depend on the price of the electricity used for hydrogen production which can be

minimized if it is powered by surplus electricity produced during off-peak hours. Despite being a well-

established technology, widespread hydrogen production from water electrolysis faces some

challenges. One of the biggest challenges is to improve overall efficiencies in order to reduce energy

consumption as well as reducing maintenance and cell materials costs to reduce hydrogen production

costs. Also, although electrolysis cells have very long lifespans (15 to 25 years), the intermittency of RES

might lead the cell to work under unfavourable conditions compromising its reliability and durability.

2.3.1. Fundamentals of water electrolysis

When supplied with the right amount of energy, liquid water can be dissociated into its elemental

components: molecular hydrogen and oxygen (Eq. (1)). At atmospheric pressure and ambient

temperature, water is liquid and H2 and O2 are gaseous.

H O (l) → H (g) +1

2O (g) (1)

The water-splitting reaction has a favourable entropic contribution since it leads to the formation of 1.5

mole of gaseous species however, the reaction is strongly endothermic and, consequently, the Gibbs

free energy change is positive, and the reaction is non-spontaneous [2].

The enthalpy (ΔH) needed to split one mole of water into hydrogen and oxygen is almost constant over

the practical temperature range (up to ~100ºC). The entropy change (ΔS) is also approximately constant

and positive resulting in an increase of the entropic contribution (T ΔS) with the increasing temperature.

The Gibbs free energy change (ΔG=ΔH-TΔS) is also positive but decreases with temperature since the

total energy is almost constant and the entropic contribution increases (Eq. (2), Figure 3).

ΔG ° H O (l) = ΔH ° H O (l) − T ∙ ΔS ° H O (l) = +237.22 kJ mol (2)

ΔH ° H O (l) = +285.840 kJ mol

ΔS ° H O (l) = 163.15 J mol K

13

Figure 3 – ΔG(T), ΔH(T), and TΔS(T) of the water-splitting reaction at P=1 bar [2].

Discontinuities observed at 100°C and 1 bar on ΔH and T ⋅ ΔS are due to water vaporization being the

magnitude of the enthalpy discontinuity equal to the enthalpy of water vaporization (+45 kJ/mol) and

equal to the decrease of the entropic contribution. The slope difference of T ⋅ ΔS(T), before and after

100°C (at 1 bar), reflects the entropy change after water vaporization.

It is clear that a higher operating temperature favours the dissociation of water since it allows the

decrease of the electrolysis voltage. At ambient temperature, around 15% of the total energy required

to split 1 mole of water comes from heat, and the remaining 85% comes from electricity [2]. When the

operating temperature is increased, the share of energy coming from heat increases, and the share of

energy coming from electricity decreases making high-temperature electrolysis an interesting topic

when heat is available since it is usually much cheaper than electricity.

2.3.2. Types of electrolysers

There are several types of electrolysers which can be divided into four main categories based upon the

electrolytes and the operating temperatures: Alkaline Electrolysis Cell (AEC), Proton Exchange

Membrane Electrolysis Cell (PEMEC), Solid Oxides Electrolysis Cell (SOEC), and Molten Carbonate

Electrolysis Cell (MCEC).

2.3.2.1. Alkaline Electrolysis Cell (AEC)

The electrolysis phenomenon was discovered in 1789 by Paets van Troostwijk and Diemann. Since then,

alkaline electrolysis became a well-matured technology for hydrogen production up to the megawatt

range [13]. Currently, AEC is the electrolytic technology with greater maturity and the larger commercial

outreach being the most widely used for hydrogen production. For this reason, this project will be

focused on modelling an AEC.

14

An alkaline electrolysis cell consists of an anode, a cathode, an alkaline electrolyte, and a power supply.

Electrodes are must be corrosion resistant, have good electrical conductivity and catalytic properties,

allowing better electrochemical transfer. They are usually made of nickel-based materials due to its

stability, high activity, high availability and its low cost. Furthermore, these electrodes must be porous to

allow the circulation of water, produced gases, electrons, and ions. Electrocatalysts are normally added

to the electrode base materials in order to enhance is electrochemical properties. The electrolyte is

usually an aqueous solution of potassium hydroxide (KOH) or sodium hydroxide at a concentration of

20-30 wt. % [3]. The diaphragm is a solid porous material that allows the transport of hydroxyl ions

between the electrodes while having a very low permeability to oxygen and hydrogen. The diaphragm

is placed between the electrodes to effectively separate the produced gases in order to avoid the

recombination of H2 and O2 into water and to avoid mix-ups that could lead to safety hazards and low

faradaic efficiencies [2].

The biggest advantages of alkaline water electrolysis are the fact that AEC can be made from abundant

and inexpensive materials and that electrolysis in alkali medium results in lower overpotentials [2].

Cheaper cell components and lower energy consumption lead to the possibility of cheaper hydrogen

production which is crucial to be able to compete with the prices of fossil hydrogen.

Three major problems are often associated with AEC: low partial load range, limited current density, and

low operating pressures. As the diaphragm does not completely prevent the product gas cross-diffusion,

small quantities of O2 and H2 will cross over. When oxygen crosses the membrane and reaches the

cathode chamber, it will be combined with the hydrogen produced at the cathode and catalyzed back

to water which leads to a reduction of the electrolyser’s efficiency [13]. Moreover, hydrogen may also

be diffused to the oxygen evolution chamber reducing the electrolyser’s efficiency and safety. This

phenomenon is particularly severe at low loads (<40%) which results in a significant decrease in the

oxygen production rate, thus drastically increasing the hydrogen concentration to undesirable and

dangerous levels prone to reaching the lower explosion limit for hydrogen (~4 mol% of hydrogen). The

second setback for AEC is the low maximum achievable current density (100 mA/cm2 - 350 mA/cm2)

[13] which is highly influenced by the high ohmic losses across the liquid electrolyte and the diaphragm.

The third issue is the inability to operate at high pressures mainly due to the liquid electrolyte which

leads to bulky stack design configurations and high compression needs for the produced hydrogen.

2.3.2.2. Proton Exchange Membrane Electrolysis Cell (PEMEC)

Proton Exchange Membrane Electrolysis Cells contain several components. At the centre, there is a thin

membrane of a proton-conduction polymer electrolyte (about 0.2mm thick) with two porous

electrocatalytic layers on each side (electrodes). This electrochemically active central component is the

membrane-electrodes assembly (MEA). The MEA is usually clamped by two porous current distributors

made of sintered titanium particles [2]. The contact points between the current collectors and the

electrocatalyst layers are critical since the average distance between contact points must be sufficiently

15

small to obtain a good distribution of current lines. In addition, two hollow bipolar plates are used to

provide electric current to the cell and separate two adjacent cells in an electrolyser. In these systems,

there is no liquid electrolyte, and its only fluid is deionized water and so, channels in the bipolar plates

are used to carry feed water to the anode and to collect liquid-gas mixtures in each cell compartment.

The central thin polymer electrolyte membrane is responsible for providing high proton conductivity

carrying solvated protons from the anode to the cathode, for guaranteeing low gas crossover assuring

the separation of the electrolysis products in order to prevent spontaneous recombination of oxygen

and hydrogen into water and allowing for high-pressure operation. The most popular material used as

solid polymer electrolyte (SPE) is Nafion® which is a sulfonated tetrafluoroethylene-based

fluoropolymer–copolymer developed in the late 1960s. The organic fluorinated backbone has the double

purpose of fully ionizing the sulphonic acid groups since the acidity of the hydrated Nafion® membrane

is similar to 1M aqueous sulphuric acid solutions, and providing adequate chemical stability especially

on the anodic side due to severe oxidizing conditions[2].

A direct current (DC) is supplied to the electrodes to power the water-splitting reaction where liquid

water is split into gaseous oxygen and protons at the anode. Due to the created electrical field set across

the cell, solvated protons migrate to the cathode through the solid polymer electrolyte membrane. In the

cathode, these protons are desolvated and reduced into molecular hydrogen. These cells can operate

at high current densities (>2000mA cm-2) to increase the hydrogen production rates [13]. The possibility

of working at high current densities reduces the operational costs and potentially the overall costs of

electrolysis. Maximum achievable current densities are usually limited by ohmic losses however higher

current densities can be achieved with a thin membrane that provides good proton conductivity (0.1 ±

0.02 S/cm)[13].

The low gas crossover rate of the polymer electrolyte membrane not only allows to produce highly pure

hydrogen but also allows for the PEM electrolyser to work under a wide range of power input covering

practically the entire power density range (10-100%) due to the quick response of proton transport

mechanism across the membrane which is not delayed by inertia as in liquid electrolytes[13].

A solid electrolyte gives rise to compact system designs with good structural properties allowing to

achieve high operational pressures which deliver hydrogen at higher pressures reducing the amount of

energy needed to further compress and store it and since it reduces the volume of the gaseous phase

at the electrodes, the product gas removal is also significantly improved [2]. Despite the many

advantages of high-pressure PEM electrolysis, the use of high pressures gives rise to the cross-

permeation phenomenon. Pressures above 100 bar will require thicker membranes and internal gas

recombiners to maintain the critical concentrations under the safety threshold of about 4 vol% of H2 in

O2.

Due to the corrosive acidic regime provided by the proton exchange membrane (pH~2) and the high

applied overvoltage (~2V) especially at high current densities, highly corrosion-resistant materials are

required not only for the used electrocatalysts but also for the current collectors and the separator plates.

16

Platinum group metals such as Pt, Ir, and Ru are usually used as catalysts and are the only ones that

have been tester commercially for being stable and having and acceptable ionic conductivity while

current collectors and separator plates are commonly made of titanium-based materials [2]. The fact that

these are scarce and expensive materials significantly increases the cost of PEMEC.

2.3.2.3. Solid Oxide Electrolysis Cell (SOEC)

Solid Oxide Electrolysis Cells are built in a sandwich structure where a dense electrolyte is placed

between two porous electrodes (anode and cathode).

In this type of electrolyser, the cathode is fed with a stream of water vapor that is reduced to H2, then

dissociated oxygen ions (oxide ions) are transported through the dense electrolyte driven by an external

voltage that must be higher than the Nernst potential and finally, in the anode oxygen ions are oxidized

to molecular oxygen along with electron releasing [14].

To assure the efficiency of the SOEC, the electrode materials should provide sufficient active sites for

electrochemical reactions as well as provide paths for reactants, products, and electrons transport. Ni-

based cermet is a commonly used cathode material due to its high catalytic activity and its chemical and

thermal compatibility with the most used electrolytes. The anode most investigated materials are

perovskites, Ruddlesden-Popper (RP) phase oxides, and double-perovskite oxides [14]. The electrolyte

is responsible for the transport of ions from the cathode to the anode being the major responsible for

the cell’s ohmic resistance. For the electrolyte, the materials of choice comprise yttrium stabilized

zirconia, gadolinium doped ceria and strontium, and magnesium doped lanthanum gallate [14].

These types of electrolysers operate at very high temperatures, usually between 700°C and 1200°C [11],

[14], which is thermodynamically favourable since part of the energy needed to split water molecules

can be supplied as thermal heat reducing the amount of electric energy needed. By lowering the

activation barrier, the H2 production rate can be greatly increased. Furthermore, considering the

significantly lower cost of heat compared to the cost of electricity, high-temperature electrolysis is

economically favoured.

SOEC features a number of advantages for the H2 production [14]: (1) the process is very precise and

can be easily controlled by managing the electrode potentials and reaction temperatures; (2) solid oxide

electrolysis systems are compact, modular, and scalable; (3) SOEC can be powered by renewable

energy just like any other type of electrolyser.

Despite being a very promising technology due to its high efficiency and low cost, SOEC are still in its

early stages of development compared to AEC or PEMEC. The durability of ceramics at high

temperatures and long-term operation is a big setback to this technology[14]. In fact, the overall

degradation of SOEC is a significant barrier for its commercial viability. Lab-scale studies are being

conducted in order to develop novel, improved, low cost, and highly durable materials and to develop

manufacturing and integration processes for the construction of efficient and durable SOEC.

17

2.3.2.4. Molten Carbonate Electrolysis Cell (MCEC)

Molten Carbonate Electrolysis Cells are the most recently developed electrolysers for hydrogen

production[11]. MCEC are high-temperature electrolysis devices that operate at temperatures ranging

from 600ºC to 700ºC which allows the reduction of the applied voltage due to favourable thermodynamic

and kinetic conditions. When compared to SOEC, MCEC have lower operating temperature meaning

that they need less thermal energy to reach operating conditions[15].

In these cells, the ionic species are carbonate ions. The cathode is made of porous nickel often alloyed

with Cr and/or Al to prevent sintering and mechanical creep. The anode is made of porous lithiated

nickel oxide usually obtained from in-situ oxidation and lithiation of sintered nickel. Between the two

electrodes, there’s a highly porous γ-LiAlO2 matrix with pores smaller than 0.1µm that retains the

electrolyte, allows the flow of the carbonate ions from the cathode to the anode, and assures the

separating between the inlet gas and oxidant gases. In order the prevent gas crossover, the porous

matrix should be completely filled with electrolyte. The electrolyte consists of a eutectic mixture of lithium

and potassium carbonate (Li/K)2CO3 or lithium and sodium carbonate (Li/Na)2CO3 [16]. At the operating

temperatures, the carbonate salt mixture is liquid and has a high conductivity. The electrolyte distribution

is the cell components is balanced by capillary pressures. Both electrodes have a biporous structure in

which the micropores are filled with electrolyte and the macropores are used as gas channels.

The water electrolysis takes place at the cathode according to the following reaction (Eq. (3)) [15], [16]:

H O + CO + 2e → H + CO (3)

Unlike other electrolysis cells, in MCEC the inlet gas must contain CO2 along with water since CO2 is a

reactant to produce carbonate ions. At the anode, the electrolysis of carbonate ions produces oxygen

and carbon dioxide (Eq. (4)) [15], [16].

CO → 1

2 O + CO + 2e (4)

Due to the presence of CO2 in the inlet gas, direct CO2 electrolysis might happen at the cathode giving

rise to the production of CO however, as the kinetics of this reaction is much slower than water

electrolysis at nickel-based electrodes, the parasitic production of CO isn’t worrisome.

At first sight, this may not seem a very attractive technology due to the need for CO2 in the inlet gas and

the release of CO2 at the anode however, looking at the overall reaction, it’s clear that for each mole of

CO2 entering the electrolyser as inlet gas, one mole of CO2 is produced at the anode which means that

the release of carbon dioxide to the atmosphere can be avoided if it is recirculated inside the cell (Eq.

(5)) [15], [16].

H O + CO , → 1

2 O + H + CO , (5)

18

2.3.3. Water Electrolysis Powered by Intermittent Energy

The connection of water electrolysers to renewable energies demands the implementation of specific

system design and operating modes due to the intermittent nature of the power sources.

A commercial electrolysis system is usually composed of three main blocks:

1. The electrolysis stack which can be composed of one or several electrolysis systems. It is the main

component of the electrolysis system since it is where water is split into oxygen and hydrogen.

2. The control system which comprises the electric powering systems and the systems that control

the whole process. Several types of control systems can be found in commercial electrolysis

systems however, in most cases, it is constituted by a master control and some specific controls

for each electrolysis stack.

3. The balance of plant (BoP) which comprises the rest of the elements of the electrolysis system:

water pumps, gas-liquid separation units, water and gas purification systems, gas compressors,

cooling systems, instrumentation items, and piping items.

All the electrolysis system components are susceptible to be affected by the power intermittency and

the associated operating modes thus, proper integration of renewable power sources requires the

identification of technical concerns caused by power intermittency and renewable power source

particularities.

2.3.3.1. Power Electronics and Process Control

An electrolysis system involves both AC-powered and DC-powered components. While pumps, fans,

compressors, cooling systems, and control system parts usually operate on alternating current (AC), the

electrolysis process requires a direct current (DC) to split water molecules into oxygen and hydrogen.

The fact that most electrical networks use AC introduces the need for power transformation from AC to

DC if the electrolyser is to be powered by electrical energy coming from the grid. If the electrolyser is to

be powered by renewable electricity directly, power transformation is also needed since each renewable

power source has specific current and voltage characteristics that may not match the electrolysis system

needs. These setbacks can, in most cases, be solved by power converters but every power conversion

implies a loss of efficiency for the global system. The existence of AC and DC-powered equipment

requires the use of one or several power converters and the choice of the type of current for the main

bus (AC or DC). Given the existence of countless possible configurations, the standardization of

electrolysis systems is complicated, and the costs of the overall system may vary depending on the

chosen configuration. All the equipment of the system should be taken into account in the power

electronics design to ensure the quality of power provided to the loads and minimize losses. Since every

power transformation is a source of inefficiency for the global system, the number of transformations

should be limited, and a careful selection of power converters should be made. The efficiency of the AC-

DC conversion process in power converters is usually high (~90%) at nominal power yet it falls

19

significantly when working at partial load. Renewable energy-powered electrolysis systems are expected

to be working at partial load during most of their operational time making the selection of the power

converters even more relevant.

System process control must also be suitable for intermittent operation. When the electrolysis system is

directly connected to an intermittent power source, it must be controlled by applying adequate and

available current through the stacks or applying appropriate voltage across the stacks. The choice

between current and voltage control depends on the electrolysis system's overall architecture and on

its main function. In electrolysis systems directly powered by fluctuating power sources, the control loop

must be able to measure the available power and process that information in order to infer the current

to apply and keep the balance between energy production and energy consumption in the grid. A fast-

operating process control system is required to ensure rapid power transience.

The use of intermittent power sources with electrolysis systems requires the anticipation of high-

frequency power and voltage transience and power failures which implies the implementation of an

uninterruptible power system (UPS) to assure safe shutdown of the system and the use of battery banks

to handle power transience faster than what the electrolyser system can handle.

The integration of electrolysis systems with renewable power sources requires complex power adapting

systems and control schemes. The design of such systems must be done very carefully to minimize the

losses associated with power conversion and guarantee the quality of the power provided to the stacks.

2.3.3.2. Impacts on Dynamic Operation Requirements

Coupling electrolysis systems with intermittent power sources imposes not only current and voltage

adaptation but also severe dynamic constraints.

A RES powered electrolysis system must be able to handle recurrent start and stop requests due to

sudden loss of power caused by the variations in power production. Besides, if the electrolysis system

operates only when excess renewable power is available it must be able to start up very rapidly to make

the maximum use of the available energy. This is frequently not compatible with the start/stop procedure

of electrolysis systems which commonly lasts for several tens of minutes depending on the technology

in question and on the size of the equipment. The long start up and shut down times are usually due to

technical requirements (system pressurization, verification of components at start-up, etc.) or safety

precautions (slow current increase, leak testing, flow rate verification, safety sensors warm up time, etc.).

Operating and electrolysis system um renewable power requires the system’s ability to work on a wide

operating range, ideally from 0% to 100%. The system should be able to follow the exact profile of the

RES completely adapting its operation to the available power including at low duty. Most of the

commercially available electrolysers have restricted operating ranges especially due to the increased

gas crossover phenomenon that takes place at low loads. This phenomenon may lead to the mixing of

hydrogen and oxygen which can result in a hazardous mixture.

20

Furthermore, the systems must be able to handle fast power transience and maintain an appropriate

response time over the whole operating range. This is commonly dealt with by the control scheme and

the control loop. Response times are influenced by several parameters such as electric behaviour, power

converter technology, signal information transfer, thermal capacity, electrochemical phenomenon and

stack design.

It is important to stress that all the electrolysis system components should be able to handle dynamic

operation since all of them have a specific dynamic response and they all must be suitable for intermittent

operation and dynamic control.

2.3.3.3. Impacts on Hydrogen Production Characteristics and Efficiency

The performance of an electrolysis system powered by intermittent renewable power is always lower

than the optimal efficiency of the system as the system spends most of its operation time working at

partial load. The actual performance of the system is strongly influenced by parameters like pressure,

temperature, and operating power range which are all linked to the dynamic operation of the system.

The hydrogen flow, pressure, temperature, and purity can also be used to characterize the hydrogen

production of a given water electrolysis system. Hydrogen output flow rate is closely linked to the power

available for electrolysis however, pressure, temperature, and purity are also affected by the intermittent

operation of the system.

Temperature

Higher operating temperatures allow higher process efficiencies and so most electrolysers are designed

to operate at an optimal temperature which is often a compromise between technical feasibility,

economics, performance, and durability [2]. Frequent start and stop along with intermittent and variable

power operation impact directly the system operating temperature and the produced hydrogen

temperature. The thermal behaviour of an electrolysis system depends strongly on the heat generated

by the stacks and on the heat losses due to the external environment and auxiliary cooling [11]. The

amount of heat generated by the electrolysis stacks depends essentially on the Joule effect occurring in

the electrolyte. The heat generated by the Joule effect is a function of the current running through the

stacks which varies widely in intermittent operation. To avoid the overheating of the electrolysis stacks,

auxiliary cooling equipment is usually included in the electrolysis system. The cooling system is often

designed to maintain the system at an optimal temperature at the maximum operating current however,

at low current setpoint, the heat generation is much lower than at high setpoint leading to an operating

temperature below the optimal one. In most cases, the variation in operating temperature does not

influence the quality of the produced hydrogen since most systems have a downstream purification unit,

nonetheless, operating at a temperature different from the design temperature causes a reduction in the

electrolyser performance.

21

The efficiency of the production stacks is usually higher at current densities below the nominal current

density because the heat losses by Joule effect are reduced. Nonetheless, the overall system efficiency

is always higher when operating at nominal power. At low loads, the system’s specific consumption is

three to four times higher than at maximum load as the specific consumption of all complementary

systems increases sharply. In a large-scale electrolysis system, the auxiliary systems power demand

usually accounts for 10% of the total power consumption operating a nominal setpoint but, for small-

scale systems operating at low load, this value can be above 80% [2].

Pressure

The variation of the hydrogen flow rate due to the power fluctuations affects the pressure of the

electrolysis system. To avoid pressure variations due to the hydrogen flow rate variations, the produced

gas would have to be instantly absorbed by the downstream elements which is not possible due to the

fluid dynamics of the elements of the system. Pressure variations in the system can be minimized using

a backpressure regulator or downstream regulation elements yet, a varying flow rate results inevitably

in pressure fluctuations.

Thermodynamically, the lower the operating pressure, the better the efficiency of the electrochemical

reaction. Nonetheless, at high pressure, smaller bubbles are produced facilitating the access of the

electrolyte to the reaction sites which improves the process efficiency. Moreover, as hydrogen

production and usage do not occur simultaneously, hydrogen storage is required. Hydrogen is usually

stored in high-pressure tanks at pressures between 10 and 75 MPa [4]. Due to the very low volumetric

mass density of hydrogen, its pressurization to such high pressures is not only difficult but also

energetically costly. For these reasons, the working pressure of the overall system is very important for

the efficiency of the global system. Even though pressurized electrolysis systems are more expensive to

implement, high-pressure electrolysis at 10 to 30 bar is usually preferred since it reduces the amount of

energy needed to pressurize hydrogen to its storage pressure [2], [4], [17]. The need for pressurization

is common to all electrolysis systems whether or not they are powered by intermittent energy however,

coupling high-pressure electrolysis systems with RES raises some challenges that may reduce the

efficiency of the overall system. Pressurizing the stack before starting to produce hydrogen consumes

energy decreasing the efficiency and increases the system’s response time (due to the time needed to

pressurize the cell at start-up) which is not desirable in renewable energy storage. Additionally, when

the system is shut down due to lack of power to operate, the stack needs to be depressurized for safety

reasons which also introduces energy losses, and therefore, penalizes efficiency.

Gas Purity

The purity of the gases produced is also affected by the power intermittency. In low duty operation, the

permeation of gases through the membrane increases, increasing the content of oxygen in the hydrogen

stream and the content of hydrogen in the oxygen stream [2], [17]–[19]. Additionally, during the stand-

22

by period induced by intermittent operation, there is also gas permeation through the membrane,

especially in high-pressure electrolysers. This can lead to the formation of a highly inflammable mixture.

This phenomenon is of special relevance in alkaline electrolysers. Ursùa et al. [20] experimentally

demonstrated that the oxygen in the hydrogen stream quickly rises when the current decreases. Tests

performed on an alkaline electrolyser at 55ºC revealed that the amount of hydrogen in oxygen can be

up to 5 times higher in low duty operation than in full-power operation. To avoid the formation of a

hazardous mixture inside the system, manufacturers usually establish a lower operating limit which is

generally set between 10 and 25% of the nominal production capacity [18]. Below this limit, the cell must

be shut down and the gases must be evacuated [2], [18]. Additionally, safety sensors continuous monitor

the composition of the gas streams triggering an alarm and shutting down the system whenever the

hydrogen content in the oxygen stream reaches 1.5 to 2 vol.% [18]. The inability to operate under this

limit introduces energy losses and decreases the cell efficiency.

Gas contamination at low current densities can be reduced by changing the way the electrolyte is

handled. In general, the cell electrolyte needs to be recirculated to avoid variations in the electrolyte

concentration and increase efficiency. In the cathodic side, water is consumed increasing the electrolyte

concentration while on the anodic side, water is produced leading to a decrease in electrolyte

concentration. Thus, the electrolyte needs to be recirculated in order to maintain the electrolyte

concentration as uniform as possible keeping the conductivity at optimal values. As the product gases

are soluble in the electrolyte, mixing of the anodic electrolyte with the cathodic one causes losses and

increases gas contamination. Partly separated electrolyte cycles with an equalization line for liquid level

balancing can be used instead of mixed electrolyte cycles. Trinke et al. [19] studied anodic gas impurities

in alkaline electrolysers at current densities up to 700 mA/cm2 at 1 bar, 10 bar, and 20 bar with mixed

and separated electrolyte cycles. In both separated and mixed electrolyte cycles, the gas impurity

decreases with increasing current density and increases with increasing pressure. The contamination

flux stays approximately constant with varying current densities however, as the amount of produced

gas decreases with decreasing current density, the percentage of contaminant gas increases [17], [19],

[21]. With increasing pressure, the amount of product gas dissolved in the electrolyte increases resulting

in higher concentration gradients which facilitates the foreign gas diffusion through the membrane [17],

[19], [21]. According to the report [19], while gas impurities with separated electrolytes were always

below 0.7 vol.% for all current densities and pressures, with mixed electrolyte cycles the gas impurities

increase significantly reaching critical values even at relatively high current densities during pressurized

operation. While at atmospheric pressure the gas impurity at 50 mA/cm2 is just slightly above the safety

limit of 2 vol.% H2 in O2, at 10 bar this limit is reached at 500 mA/cm2. For the tested cells, at 20 bar,

even at 700 mA/cm2, the gas purity was never below the safety limit.

Forcing depressurization after a certain stand-by time also reduces drastically the chances of producing

a hazardous mixture inside the cell but, when there is enough power to restart the cell, the restart time

will be significantly increased. The long restart time implies wasting the energy produced during this

period decreasing the overall efficiency.

23

2.3.3.4. Impacts on Reliability and Durability and Solution Approaches

Water electrolysis systems can suffer from degradation of different types over time and all its

components (stacks, control systems, BoP, etc.) are prone to this degradation. In most cases, the

systems degradation leads first to a decrease in efficiency and then to the failure of the system.

Electrolysis system materials, configurations, costs and efficiencies have been significantly improved

over the last decades however, failure mechanisms and degradation modes induced by intermittency

still need to be extensively studied. A few studies quantifying performance degradation with time are

available but generally the aggravating factors related to intermittency are not clearly identified.

The main degradation mechanism in alkaline electrolysers is corrosion due to the nature of the

electrolyte. The choice of the electrolyte and its concentration always requires a compromise between

performance and risk of material corrosion. Furthermore, the relatively high operating temperature of

AEC also greatly increases the effects of corrosion. These factors contribute to the degradation of the

sealing components of the electrolyser increasing the risk of electrolyte leakage that may lead to

premature stack degradation. The loss of integrity of the separator material is also one of the main

causes of stack failure, especially at high pressure. Diaphragm degradation leads to gas cross

contamination increasing the risk of formation of hazardous mixtures. The caustic environment inside

the electrolysis cell also causes wearing of the electrodes leading to reduced electrochemical

performances.

Despite the lack of root cause studies regarding stack degradation, two deterioration factors can be

intuitively pointed out as possibly strengthened by intermittent operation: mechanical stress induced by

frequent starts and stops, fluctuating product gas flow rates and frequent system depressurizations; and

high frequency of current and voltage setpoint changes across the stacks. To avoid stack failure during

operation, manufacturers usually establish a maximum number for starts/stops.

Electrolysis stacks are rather fragile mechanical assemblies that require perfect sealing and leak

tightness both inside the stacks and to the outside environment. Frequent temperature and pressure

variations induces mechanical stress which favor the degradation of sealing and tightening components,

potentially leading to containment and system failure.

Intermittent power availability and RES characteristics induce high frequency current and voltage

variations which generate stress on electrolysis stacks and power electronics potentially leading to

degradation and failure of these components. Nickel is usually the material of choice for AEC cathodes

since it presents good corrosion resistance at a fairly low price however, its corrosion is significantly

increased during intermittent operation as they start to present significant degradation after 5000 to

10000 start/stop cycles [17]. To prevent premature electrode degradation, efforts should be made to

reduce the number of cycles.

In fact, cathode degradation occurs when the electrolyser cell voltage drops below a minimum value

(~0.25V for nickel) after the hydrogen production has stopped. When the electrolysis process is stopped,

24

the current in the stack drops to zero almost immediately while the voltage decreases at a far slower

rate. This is due to the fact that the electrode’s electrochemical double layer acts as a capacitor and

delays the voltage breakdown after a power loss [18]. One of the solutions requires a strategy to allow

the electrolysis cell to operate below the lower operating limit for a certain short period of time before

shutting down the system. During this period, the voltage should remain above the minimum value at

which electrode degradation starts (~0.25 V for nickel) [2], [17], [18]. Also, during this period, the gas

contamination must also be below the safety limit of 2 vol.%.

Another approach to decreasing the number of start/stop cycles is the implementation of a back-up

battery system. The purpose of this storage system is to store the energy that the electrolyser cannot

use in a certain period (below the minimum load or above the maximum load) in order to provide it to

the electrolyser when the power production is below the lower operating limit in order to increase the

electrolyser operating time and reduce the number of cycles the system is subjected to [2], [17], [18].

2.4. Alkaline Water Electrolysis Fundamentals

An Alkaline Electrolysis Cell consists of two electrodes: the anode and the cathode; immersed in an

alkaline solution, usually aqueous potassium hydroxide (KOH) or sodium hydroxide (NaOH) at

concentrations ranging from 20 wt.% to 30 wt.%, separated by a diaphragm [3].

The water-splitting reaction requires the supply of electrical energy through a potential difference

between the two electrodes. A direct current is applied between the two electrodes to trigger the water

splitting reaction leading to an electron flow from the negative (cathode) to the positive (anode) terminal

of the DC source. In practical operating conditions, the cell voltage usually ranges from 1.3V to 2.2V

[22]. When the applied potential difference is within the practical range of operation, reduction and

oxidation of water take place simultaneously on the cathode and the anode, respectively. At the cathode,

which is polarized at a negative potential (relative to the RHE), electrons provided by the DC source are

consumed by hydrogen ions (protons) to form hydrogen. This is the hydrogen evolution reaction (HER)

(Eq. (6)). By keeping the electrical charge-balanced, hydroxide ions transfer through the electrolyte to

the anode is assured. At the anode, which is polarized at a potential close to 1.8V-2.0V RHE (reversible

hydrogen electrode), hydroxide ions give away electrons to the positive terminal of the DC source. This

is the oxygen evolution reaction (OER) (Eq. (7)). Both electrochemical reactions occur at the electrode

surface at the contact point between the electrolyte (H2O and OH-) and the metallic electrodes that

conduct electrons [3], [22].

Cathode 2H O + 2e → H + 2 OH (6)

Anode 2 OH → 1

2O + H O + 2e (7)

Overall Reaction H O (l) → H (g) +1

2O (g) (8)

25

For this electrochemical reaction process to take place, several barriers must be overcome, requiring a

sufficient electrical energy supply. These barriers, usually called overpotentials, include electrical

resistance of the electrical circuit, activation energies of the electrochemical reactions occurring at the

electrode surface, availability of electrode reaction sites due to partial coverage by gas bubbles formed,

and the resistances to the ionic transfer within the electrolyte solution. These barriers must be analysed

in the contexts of thermodynamics and kinetics as well as transport process principles [22]. Due to these

factors, the actual cell voltage, 𝑉 , is distinct from the reversible cell voltage, 𝑉 , and can be split

into its contributing factors (Eq. (9)): the reversible voltage, 𝑉 ; the activation overpotential, 𝜂 ; the

concentration overpotential, 𝜂 , and the ohmic overpotential, 𝜂 [5], [6], [23], [24].

𝑉 = 𝑉 + 𝜂 + 𝜂 + 𝜂 (9)

The activation overpotential, 𝜂 , results from the kinetics of the electronic charge transfer reactions

that take place at the electrodes’ reaction sites [24]. The nature and pre-treatment of the electrode

surface as well as the composition of the electrolytic solution adjacent to the electrode determine the

electrode reaction rate.

At high current densities, the reactions are not only controlled by electronic transfer but also by mass

transfer. Due to the electrochemical reactions taking place at the electrodes, a gradient in the reactant

concentrations arises in the vicinity of the electrode surface giving rise to a concentration or diffusion

overpotential, 𝜂 .

The ohmic overpotential, 𝜂 , is caused by the finite conductivity of electrolysis cells. The electrodes,

electrolyte and separator contribute to a further overpotential in an alkaline electrolyser cell due to the

resistances of these elements.

The model developed on this project will model each one of these contributions in order to obtain the

polarization curves for the electrolyser depending on its physical characteristics and operating

conditions. Polarization Curves are plots of current or current density versus operating voltage which

allow the estimation of the power consumption of the electrolyser.

2.4.1. Conventional vs Zero-Gap

In traditional cell designs (Figure 4), the cell assembly had a defined distance between the electrodes

that could range from approximately 1 mm to several centimetres. The bigger the distance between the

electrodes, the bigger the ohmic losses however, if the gap is too small, the reactant concentrations

change rapidly during operation which is not very convenient [2]. Also, as the specific conductivity of

the electrolyte gap is affected by the presence of gas bubbles, if the electrode distance is too small, the

gas bubbles will accumulate in the gap and lower the conductivity. With increasing distance, the bubble

detachment is enhanced and the specific conductivity increases [17]. Therefore, the size gap between

electrodes is a trade-off between maintaining the ohmic losses at its lowest possible value while not

compromising the conductivity of the electrolyte [2], [17].

26

Zero-gap electrolysis cell designs are becoming increasingly popular. In this type of configurations, two

porous electrodes are compressed into both sides of a hydroxide-ion conducting membrane. This leads

to an electrode gap equal to the thickness of the membrane (<0.5mm) which is much smaller than the

ones of traditional setups [25]. The very small electrode gap significantly reduces the ohmic resistance

contribution from the electrolyte between the electrodes. A gas diffusion layer provides an electrical

connection between the porous electrode and the bipolar plate while allowing a feed of electrolytic

solution to the reaction sites and the removal of product gas.

In the conventional setup, the bubbles migrate to the top of the cell from both sides of the electrodes

including the side facing the separator. On the electrode-diaphragm gap, the volume of solution

displaced by the bubble will not be available for the transport of hydroxide ions through the membrane

increasing the ohmic resistance. The zero-gap configuration forces the bubbles to be released from the

back side of the electrodes (the one not facing the separator) reducing significantly the void fraction and

hence reducing the resistance induced by the migration of hydroxide ions [25].

The membrane-electrode gap is a very relevant parameter for the modelling of the electrolyser since it

greatly affects the ohmic overpotential, 𝜂 .

Figure 4 - Conventional Cell Design vs Zero-Gap Cell Design [17]

2.4.2. Monopolar vs Bipolar

There are two alkaline electrolysis stack configurations: monopolar and bipolar (Figure 5).

The monopolar arrangement consists in alternate electrodes directly connected to the opposite

terminals of the DC power supply. This type of stack configuration results in several individual

electrolysis cells in parallel with one another. The total voltage applied to the electrolysis stack is the

same as that applied to each individual pair of electrodes in the cell.

In the bipolar arrangement, the only two electrodes are connected to the DC power supply. Every two

adjacent electrodes form a unit electrolysis cell, and these unit cells are electrically linked in series with

one another through the electrolytic solution that acts as conducting media. The total cell voltage is the

sum of the individual cell voltages.

In the unipolar configuration, the same electrochemical reaction takes place on both sides of the same

electrode, that is, each electrode acts solely as anode or cathode. In the bipolar configuration two

27

different reactions take place at each side of the electrodes that are not connected to the power source,

that is, one side of the electrode acts as anode and the other one acts as cathode [22].

From the manufacturing perspective, the monopolar configuration is easier to fabricate and maintain

since it requires fewer and simpler parts however, this type of configuration suffers from high electrical

currents at low voltages leading to large ohmic losses and so it must be operated at lower current

densities and lower temperatures [22], [26]. The bipolar configuration is much more compact which

allows shorter current paths in electrical wires and electrodes leading to lower ohmic losses, and

therefore increasing cell efficiency. Nevertheless, bipolar stacks require greater precision in design and

manufacturing to prevent electrolyte and gas leakage between cells and are harder to repair since they

cannot be repaired without shutting down the entire cell and servicing the entire stack [26].

Figure 5 - Monopolar (a) and Bipolar (b) electrolyser configurations [27]

2.4.3. Electrolyte

For the electrolysis reaction to occur, the use on an electrolytic solution is required since it consists of

ions with high mobility.

M. Sellami et al. [28] report that the amount of hydrogen produced by a cell with KOH electrolyte is

higher than that produced by the NaOH cell. The advantage of the KOH cell may start even before the

preparation of the electrolyte solution since the dry purity of KOH is higher than that of dry NaOH. The

higher the purity of the dry hydroxides the fewer the parasitic reactions on the resulting electrolyte and

the lower the energy losses caused by them which leads to higher electrolyser efficiency. Although the

main advantage of the KOH electrolyte is believed to be the lower cell resistance of the KOH cell and

the higher ionic mobility of K+ ions compared to the one of Na+ ions resulting in a higher electrical quantity

crossing the KOH cell implying a higher volume of produced hydrogen. The maximum conductivity of

KOH aqueous solution if obtained at around 30 wt.% and so it is the most commonly used concentration

in modern alkaline electrolysers [2], [23], [29] (Figure 6).The conductivity of NaOH is in general much

28

lower than the one of KOH and the highest conductivity is usually achieved at concentrations around 20

wt.% (Figure 6).

Figure 6 - KOH and NaOH Specific Conductivity at 60°C and 80°C (based on [23])

The amount of electrolyte present in an electrolytic cell must be adjusted continuously to make up for

the losses through product gases. Typically, a cell loses 1mg of KOH per Nm3 of H2 produced. Also, the

electrolyte must be changed periodically due to the accumulation of impurities coming from water and

from the degradation of electrolyser components that poison the electrolyte [2].

The type of electrolyte used and its concentration are important parameters that need to be taken into

account when modelling an AEC. The use of NaOH instead of KOH, for example, will significantly

increase the ohmic overpotential of the cell.

2.4.4. Electrodes and Electrocatalysts

Electrode materials for alkaline water electrolysis should have good corrosion resistance, high electronic

conductivity, and high catalytic activity regarding the two reactions of interest. Usually, the kinetics of

the hydrogen evolution is rapid and so, the reaction takes place with only a small overpotential at many

different electrode surfaces. In contrast, the kinetics of the oxygen evolution is much slower requiring a

substantial overpotential at the anode which is generally the biggest inefficiency problem in water

electrolysers [12]. Stainless steel and lead oxide were first identified as cheap electrode materials with

low overpotential for the OER however, they were not chemically stable at sufficiently high voltage in

highly concentrated alkaline solutions. Currently, nickel is recognized as one of the best materials for

OER offering rather good corrosion resistance in highly alkaline solutions, good electrical activity, and a

reasonable cost when compared to other suitable materials [2] with an average cost over the 2015-2020

period of 11.97 US$/kg [30] (≈10.05 €/kg [31]) while the average cost of platinum over the same period

is around 31.78 US$/g [32] (≈26,69 €/g [31]). Plain nickel electrodes present good durability in real

operation conditions yet accelerated corrosion phenomena occur during power-off periods [2].

29

Three main factors must be taken into consideration when developing a commercially practical electrode

material for alkaline water electrolysis namely, electrochemical efficiency, stability, and cost [33], [34].

The efficiency is the initial evaluation criterion on any electrode development process. The most practical

methods to increase electrode efficiency involve increasing the electrolyser’s operating temperature,

increasing the electrochemically active surface area of the conventional nickel electrodes, or the

introduction of electrocatalysts. Increasing the cell’s operating temperature is usually advantageous

since it reduces the voltage required to maintain a given cell current density however, in some cases,

the increase in operating temperature may increase electrode corrosion. To increase the

electrochemically active surface of the conventional nickel electrodes, nickel polished sheets can be

replaced by other types of nickel structures such as corrugated plates, perforated plates, meshes, wire

cloths, etc [34]. In addition, a porous layer of nickel or nickel-iron alloys is usually coated over the base

nickel electrodes to catalyze the reactions. The electrocatalyst performance is surely dependent on the

catalyst composition and its surface area which is dependent on the catalyst microstructure and the

electrode support material in which the catalyst is applied [2], [12]. Electrocatalyst performance can be

translated by the Tafel Slope: the lower the Tafel Slope, the better the performance. This parameter and

its variation with temperature is of extreme importance for the physics model since the charge transfer

coefficient, 𝛼, can be computed from it.

After a satisfying efficiency is achieved, the electrocatalyst should be tested for chemical and physical

stability. A good electrocatalyst should remain chemically and physically stable for the entire lifespan of

the electrolysis system [33], [35]. Electrochemical corrosion at the surface of a metallic electrode

immersed in an electrolyte is the main stability issue of electrocatalysts. In addition, structural

transformations of electrocatalysts can lead to severe degradation [35].

The maximum practical electrode cost depends on the capital constraints of the whole electrolysis

system which, in turn, depends not only on the cost of competing systems but also on the price of

hydrogen produced through different methods.

2.4.4.1. Cathode

Traditionally, alkaline water electrolysers used to have cathodes made of iron-based materials however,

in highly alkali conditions and high temperatures, its resistance to corrosion decreases significantly

making iron and steel cathodes not adequate for alkaline electrolysers. Despite being more expensive

and a less active catalyst than iron and steel, nickel is now widely used as a cathode substrate material

in alkaline water electrolysers being also a common component of the cathode material [12]. The Tafel

slope for smooth nickel in alkali solutions of KOH or NaOH at room temperature is around ~115-120

mV/dec [36], [37] meaning that the rate-determining step in the hydrogen evolution is the initial electron

transfer from the cathode to water molecules. The overpotential for the HER in smooth nickel electrodes

is typically ~350-450mV, being far too large for modern electrolysers [12], [37]. One of the biggest

problems of nickel-based cathodes is the fact that, under certain conditions, hydrogen can be absorbed

30

into the nickel lattice leading to its deactivation which increases the cathodic overpotential. Fortunately,

the cathodic overpotential increase is much smaller at the electrolyser’s operating temperature range

343-373K. In fact, studies report that this overpotential decreases around 100 – 200mV if we shift the

operating temperature from 293 K to 363 K [12], [37] (Table 4). The sandblasting of the nickel-based

electrodes increases the electrochemically active surface area and, only by itself, can provide an

overpotential reduction in the HER of around 50 mV for current densities between 0.1 and 1.0 A/cm2

[12], [38].

If, after the sandblasting process, a Raney Nickel coating is applied, the overpotential for the hydrogen

evolution can be much further improved [2], [12], [39]. Compared to smooth nickel electrodes, Raney Ni

coated electrodes can achieve overpotentials 200 mV lower and Tafel slopes around 80 mV/dec lower

resulting in overpotentials for the HER of around 100mV at 0.3 A/cm2 and Tafel Slopes of around 40

mV/dec [12], [38]. These electrodes are usually prepared by sintering or electroplating but the most

successful one involves the encapsulation of NiAl or NiZn nanopowder in an electroplated Ni layer after

which the Al or Zn dissolves in KOH/NaOH electrolyte resulting in a very large surface area [12]. Tanaka

et al. [39] studied the influence of the Ni/Al ratio in the Raney Ni and reported that the lowest overpotential

for the HER is obtained when the Al content is high. Oxides of Ti, Zr, and Nb have also been added to

the layer to improve its properties. Despite providing significant improvements in the HER overpotential,

Raney Ni electrodes are not stable to current reversals when taking the cell off load [2], [12].

NiMo alloys electrocatalyst layers were proven to have several advantages over porous Raney Ni

structures and several authors [40]–[42] report the outstanding performance of these alloys. NiMo alloys

have very low overpotentials, around 60 mV at 0.5 A/cm2 in 30 wt% NaOH at 343 K and are able to

maintain this performance for many thousands of hours even when subjected to variations of the current

load and switching off the cell.

It is widely known that precious metals are excellent catalysts for the HER at all pH however, in

electrolysis cells, the main challenge is to obtain high surface area coating of precious metal catalysts

at an acceptable cost. As for other coating procedures, the Ni substrate electrode is pre-treated usually

using a degreasing solution followed by sandblasting and acid etching. Most commonly, the coating can

be deposited in by electroplating, by immersion plating where the nickel substrate itself is used as the

reducing agent or by thermal decomposition of a sprayed precious metal solution. A PtRu alloy with a

loading of 0.30-0.35 mg/cm2 of Pt and 0.10-0.15 mg/cm2 of Ru prepared by immersion plating after the

usual pre-treatment is recommended in the literature [12], [43], [44]. This coating provides an H2

evolution overpotential of around 50 mV at 0.3 A/cm2 at 363 K in 35 wt% NaOH electrolyte yet the

overpotential quickly rises to ~100mV [12]. This rise is commonly attributed to poisoning by iron and/or

organic compounds and it can be avoided by subjecting the coating to a post-treatment. Ru on Ni

substrate and Ir on Ni substrate cathodes were also studied and reported in the literature. In a 1M NaOH

electrolyte solution, these catalysts can provide overpotentials as low as 57 mV at 0.1 A/cm2 and Tafel

Slopes as low as 52 mV/dec [45] (Table 4).

31

In general, precious metal catalysts are the ones with the best performance and the best stability in the

AEC conditions however, their price very high which significantly influences the price of the hydrogen

produced. Currently, if precious metal catalysts are to be avoided, NiMo alloys seem to be the best

choice for cathode electrocatalyst for AEC.

Table 4 – Summary of cathode electrocatalyst performance

Substrate Catalyst Electrolyte i (A/cm2) T (K) A (mV/dec) 𝜂 (mV) Ref

Smooth Ni none 30wt% KOH 0.250

303 115 443

[37] 323 111 416

343 104 367

363 103 335

Steelblasted Ni none

50wt% KOH 0.100 363

270 [38]

[38]

[38]

Ni Raney Ni 240

Ni Raney Ni/Co 150

Ni NiMo alloys 30wt% NaOH 0.500 343 40-60 60 [41]

Ni Ru 1M NaOH 0.100

80 87 [45]

Ni Ir 52 57 [45]

Pt Pt-black 50wt% KOH 0.100 363 110 [38]

2.4.4.2. Anode

The overpotential for the oxygen evolution at the anode is always much larger than the one for the

hydrogen evolution at the cathode due to the fact that the kinetics of the oxygen evolution is much slower

since it is a complex multistep reaction involving at least four electron transfers and four proton transfers.

The oxygen evolution overpotential is usually around 400mV representing about 20% of the total cell

voltage [33]. Furthermore, the overpotential for the oxygen evolution reaction depends strongly on the

anode material, its surface state, and temperature.

None of the existent electrocatalysts have a satisfactory performance so continuous efforts are much

needed to develop lower overpotential and more stable electrocatalysts for the oxygen evolution. The

oxygen evolution reaction and the oxygen reduction reaction both require significant overpotentials and

so, even at the best electrocatalysts, the potential for the oxygen evolution reaction (OER) is substantially

positive compared to that for the oxygen reduction reaction (ORR) leading to the formation of an oxide

layer in the surface of the catalyst and the corrosion of both the electrocatalyst and the catalyst support

[12]. Furthermore, due to the difference in potential and the complex mechanisms involved in ORR and

OER, the best electrocatalysts for each reaction may be very different.

One of the biggest advantages of alkaline water electrolysis is that non-precious metals including nickel,

cobalt, and steels can be stable. D.E. Hall [33] and Miles et al. [46] reported that precious metals have

little or no advantage over nickel for oxygen evolution in alkaline solution since Ir, Pt, Ph, and Rh have

32

similar activities while Ni is slightly better. In later work, Miles et al. [47] prepared several precious metal

oxide electrocatalyst by a thermal decomposition method and tested them in a 30% KOH electrolyte at

353K reporting Ru to be the best performing of the tested electrocatalysts followed by Ir, Pt, Rh, Pd, Ni,

and Os which have similar performances to one another and are much more effective than Co and Fe.

For these electrocatalysts, Tafel Slopes were found to be between 46 and 191 mV/dec.

Traditionally, alkaline electrolyser anodes are made of nickel since it’s relatively inexpensive, highly

corrosion-resistant at positive potentials in alkaline electrolytes and the oxygen evolution efficiency for

nickel is among the highest for elemental metals making nickel a good benchmark against which

improved anodes should be compared [33].

In general, Ni and/or Co electrocatalysts (high active area metals, oxide/hydroxide layers, spinel oxides,

perovskites) outperform precious metal materials in terms of performance and stability while also being

much cheaper [12].

In early work, Hall [34], [48] corroborates the advantage of using high surface area Ni electrocatalysts.

The high surface area Ni coating fabricated by sintering Ni powder onto a steel electrode substrate

presented and OER overpotential of 250 mV at 0.4 A/cm2 in 30 wt% KOH at 353 K and a Tafel slope of

around 35 mV/dec (Table 5). Anodes resulting from this experiment were also proven to have good long-

term stability. In addition, the active area of the electrocatalyst can be further increased by impregnation

of the porous structure with nickel hydroxide Ni(OH)2 via electrochemical precipitation methods allowing

a reduction in the OER overpotential of 45-60 mV [49].

A different successful approach to obtaining low anode overpotential comprises the use of mixed

transition metal oxides/hydroxides notably iron-doped nickel hydroxide composites [50], [51]. The

precipitation of iron into nickel hydroxide thin films significantly improved the catalysis of the OER

reducing the Tafel slope from around 70 mV/dec in films without iron to about 25 mV/dec in films with

10-50 wt% iron. The oxygen overpotential on thin electrocatalytic films of Ni-Fe(OH)2 was more than 200

mV lower than that for Ni(OH)2 electrocatalytic films at 0.08 A/cm2 in 25 wt% KOH at 296 K [50]. Li et al.

[52] corroborated the good performance of iron-doped nickel hydroxide composites by comparing

several different Ni-based electrocatalyst coatings (Table 5).

Spinel oxides have been known for several decades to be effective anode electrocatalysts for OER with

promising long-term stability and performance. Tseung et al. [53]–[56] developed the first studies on the

use of NiCo2O4 and Li-doped Co3O4 as anode electrocatalysts. Teflon-bonded NiCo2O4 electrodes

allowed to reach current densities around 1.3 A/cm2 at 358 K in 5M KOH electrolyte at an overpotential

of 400 mV. These electrodes were proven to be stable for the OER at 358 K in 45 wt% KOH at 1 A/cm2

for more than 3000h with an increase in overpotential of less than 50 mV. Tseung et al. [56] also reported

Li-doped Co3O4 to have a better catalytic performance than NiCo2O4 giving 1 A/cm2 at an overpotential

of 300 mV in 5 M KOH at 343 K. Singh et al. [57] studied the influence of the oxide loading with different

microstructures in the electrocatalytic activity and reported an improvement in the electrocatalytic

activity with increasing oxide loading, particularly at relatively low loading levels. In an intermediate range

33

of loadings, the activity of the films was reported to be practically constant. Singh et al. reported Tafel

slopes as low as 50 mV/dec and current densities as high as 3 A/cm2 at an overpotential of 450 mV in

1M KOH at 298 K.

Table 5 - Summary of anode electrocatalyst performance

Substrate Catalyst Electrolyte T (K) i (A/cm2) A (mV/dec) 𝜂 (mV) Ref

Ni Steelblasted none 50wt% KOH 363 0.100 59 300 [38]

Ni Ni

50wt% KOH 363 0.100

53 280

[38] Ni Raney Ni 60 240

Ni Raney Ni/Co 55 240

Ni Raney Co 50 230

Mild Steel Ni 30wt% KOH 353 0.100

35 230

[34] 0.200 240

Mild Steel Ni flake 30wt% KOH 353 0.100

33 220

[34] 0.200 230

Mild Steel Fe-37Ni alloy 30wt% KOH 353 0.100

35 240

[34] 0.200 230

Mild Steel Ni(OH)2 30wt% KOH 353 0.200 23-27 180 [49]

99.97% Ni Ni (OH)2

25wt% KOH 296 0.080 70 371

[50] NiFe(OH)2 21 231

Stainless Steel none

1M NaOH 353 0.250

406

[52] Smooth Ni none 406

Ni Ni(OH)2 386

Ni FeNi(OH)2 311

Fe Co3O4 1M KOH 303 0.010 45 350 [57]

Ni foam Co3O4 1M KOH 0.010 328 [58]

Ni Ni Co2O4 1 M NaOH 353 0.250 347 [52]

Ni Ni Co2O4

5M KOH 358 1.3 400

[56] Li-Co3O4 343 1 300

Ni Li-Co3O4 1M KOH 298 0.100 60±5 400 [59]

2.4.5. Diaphragms

The diaphragm is a solid porous material placed between the anode and the cathode.

The diaphragm material must be electrically insulating to prevent any short circuit between the

electrodes. Besides, it must have a very low permeability to oxygen and hydrogen in order to avoid the

mixing of the produced gases which may lead to their recombination into water penalizing efficiency,

34

and to safety hazards. Furthermore, the diaphragm must also be a very good ionic conductor allowing

the transport of hydroxyl ions from the cathode to the anode [2]. While serving all these purposes, the

diaphragm should be stable in the standard conditions of operation of the electrolyser namely highly

alkali solution (30wt% KOH or 20% NaOH) at around 80°C, highly oxidizing conditions at the anode,

reducing conditions at the cathode, gas bubbling, etc.

When the electrolyser is operating, the ionic current passes through the liquid electrolyte that fills the

diaphragm’s pores hence, the ionic conductivity depends not only on the conductivity of the liquid

electrolyte but also on the porosity and tortuosity of the diaphragm material. To improve the ionic

conductivity of the separator and lower product gas crossover, the use of hydrophilic materials is

required. Despite the complexity of the calculation of the ionic conductivity of such porous media, some

simplistic models are available in the literature [60]–[62].

A well-performing diaphragm for alkaline water electrolysis should be temperature and corrosion

resistant, should not have any electron conduction capability of its own, should be sufficiently

mechanically stable and consequently, suitable for the manufacture of a very thin and porous layer that

can guarantee good ionic conductivity of the enclosed electrolyte. Moreover, these diaphragms should

have a relatively simple production method and a low cost of materials [63].

The ionic resistance of the membrane is another crucial parameter for the modelling of the electrolysis

cell since it is the most important contributor to the ohmic overpotential along with the electrolyte

resistance.

2.4.5.1. Inorganic Materials

Asbestos

The first commercial diaphragms were made of microfibrous chrysotile asbestos despite not being very

resistant to corrosion caused by the strong alkaline medium and high temperatures. Asbestos was also

proven to have poor mechanical stability at the usual operating conditions. To improve its mechanical

stability, the thickness of the membrane would have to be increased to at least 2mm resulting in a high

voltage drop leading to further energy losses. In fact, under usual operation conditions (90ºC and 30

wt.% KOH), its surface specific electrical resistance is around 0.5 – 1 Ω/cm2 [64]. Asbestos dissolution

in the electrolyte was found to increase with time, temperature, and KOH concentration with weight

losses reaching up to 40% of the initial weight. Moreover, asbestos was proven to be carcinogenic and

its commercial use was forbidden in the European Union in 1999. The use of asbestos for the production

of separators is still authorized until 2025 for existing electrolysis installations until they reach the end of

their service life or until suitable asbestos-free substitutes become available.

Sintered Metals

The use of porous metallic diaphragms made from sintered metals might seem appealing due to the

high resistance to corrosion, good ionic conductivity, and the good gas sealing properties possessed by

35

some metals, however, there are two main reasons why diaphragms made of sintered metals can’t be

easily applied. Firstly, metals have good electrical conductivity and so they may not easily be

incorporated in an electrolytic cell and they are not compatible with zero-gap geometries. To overcome

this issue, the diaphragm must be completely insulated from the electrodes and other electrolyser

structures. Failure to achieve complete insulation will lead to the diaphragm taking the place of one of

the electrodes leading to the production of an explosive gas mixture. Secondly, only Pt-group metals

are corrosion resistant and neither nickel nor zirconia passivates sufficiently to guarantee the long-term

stability of these diaphragms under electrolysis conditions [65].

Divisek et al. [66] managed to reduce the electronic conductivity of the nickel structure by converting

the nickel surface into an oxide using thermal oxidation in air at temperatures above 1000°C. Sintering

of oxides of less noble metals which are, in general, hydrophilic would allow the production of highly

porous hydrophilic structures however they are not suitable for the production of very thin diaphragms

that would be corrosion resistant and have far lower ionic resistance than asbestos since these materials

are very brittle.

Metal-net supported oxide ceramic diaphragms

Materials for the supporting metal net need to be corrosion resistant in high-temperature caustic potash

solution in oxygen atmospheres. Only titanium, zirconium, and low carbon/low sulfur nickel are

theoretically stable in these conditions due to passivation. However, accelerated corrosion tests

conducted in 50 wt% KOH at 220 ºC showed that titanium and zirconium were completely dissolved

after approximately 100h and only low-C/low-S nickel was stable against caustic corrosion even in the

presence of oxygen [67].

Only a limited number of inorganic compounds especially oxides are expected to be stable in alkaline

electrolyser conditions (simultaneous contact with concentrated alkali solution and oxygen at the anode

and concentrated alkali solution and hydrogen at the cathode). Simultaneous stability against hydrogen

and will always be guaranteed for any oxide of the form MenOm if its free enthalpy per atom of oxygen is

more negative than the free enthalpy of formation of water [67]–[69]. In fact, only some zirconates and

titanates are expected to be stable in these conditions. In addition, as these oxides are intended to be

processed in a sintering procedure, their melting point must be low in order to allow them to be easily

sintered well below the metal-support melting point. The estimated sintering point of any substance is

estimated to be around 2/3 of its absolute melting temperature. Being nickel the most suitable metal to

be used as metal-net support and considering that its melting point is around 1453 °C (1726 K), the

chosen metal oxides should have melting temperatures below 2300°C [67], [68]. Given these limitations,

only alkaline earth zirconates and titanates seem to be promising materials. Moreover, the titanates

exhibit lower melting points than alkaline-earth zirconates and so, they are predicted to be processed

and sintered more easily [68].

The production of metal-net supported nickel oxide separators was initially considered to be a good

alternative to asbestos as separator material since the use of a metal net would eliminate the excessive

36

brittleness of sintered nickel oxide separators. However, the NiO free enthalpy per atom of oxygen

(Δ𝐺 = −203 𝑘𝐽/𝑚𝑜𝑙 ) is less negative than the free enthalpy of water (Δ𝐺 = −225 𝑘𝐽/𝑚𝑜𝑙 )

meaning that NiO may be reduced during alkaline electrolysis. Furthermore, oxygen adsorption and its

subsequent mounting into the NiO lattice may also occur. This reaction produces, in the case of

semiconductors like NiO, electron vacancies which lead to an increase in the electrical conductivity

which may cause short circuits in compact cell configurations [69]. Divisek et al. [69] synthetized a nickel

cloth supported NiO diaphragm by compressing carbonyl nickel powder with 2-3 µm particle size onto

a nickel carrier screen yielding a double-sided Ni “green body” which was then fired for 15 minutes in

air atmosphere at 1000 ºC oxidizing the nickel into NiO. The properties of the resulting diaphragm can

be found in Table 6. Despite the probable chemical instability predicted by the theory, the diaphragm

was proven to be stable in 10 M KOH at 120 ºC with a total weight loss of only 2% after 20000h. In a

different experiment, the diaphragm was tested under electrolysis conditions at 120 ºC in 10 M KOH at

a current density of 4000 A/m2 for 15000h. In this test, the total weight loss of the pure NiO-diaphragm

was negligible.

Table 6 - Physical Properties of pure NiO-diaphragm [69]

Property Value Unit

Thickness 0.5 mm Maximum temperature for continuous operation 120 ºC Porosity 50 % Pore radius 2-3 µm Gas Permeability (0.1 bar pressure difference) 5-10 l/min∙m2 Gas Purities (10M KOH, 100ºC, 4000 A/m2, 5 bar)

- Oxygen purity > 99.5 % - Hydrogen purity > 99.8 %

Ionic Resistance at 30ºC in 30 wt% KOH 0.2 Ω ∙ cm2 Tensile Strength 50 N/cm Bending Diameter 20 cm

Divisek et al. [69] also studied the influence of additives on the chemical stability of NiO in a nickel-cloth

supported nickel oxide diaphragm. The additives can be added to the original Ni powder either as a

metal power before oxidizing Ni into NiO or as an oxide. The Gibbs energy of the reaction between the

additive and NiO at the preparation temperature (≈1000 ºC) should be negative so that a new chemical

compound is created instead of forming a mixture of substrates. The additive of choice is usually TiO2

which, when processed with NiO at around 1000 ºC, gives rise to the formation of NiTiO3. Experiments

have shown that, with the addition of 5 wt% of TiO2, the corrosion rate is minimum (400 mg/m2∙1000h)

and inferior to the one of pure NiO. For TiO2 contents above 5 wt%, the corrosion rate increases being

superior to the one of pure NiO for TiO2 contents above 10 wt%. The use of ZrO2 as an additive is not

advantageous since its processing with NiO does not give rise to the formation of a new chemical

compound but a mixture of substrates instead and hence, the chemical stability of the separator remains

unchanged. In fact, experimental results show that the addition of ZrO2 leads to corrosion rates higher

than the one of pure NiO.

37

Fischer et al. [70] have suggested the coating of nickel meshes with an aqueous slurry of corrosion-

resistant ceramic materials, followed by drying and sintering. The ceramic coating so obtained insulates

the nickel mesh on both sides. Aqueous slurries of BaTiO3 and BaTiO3 (46%) + ZrO2 (46%) + K2Ti6O13

(4%) +Na2TiO3 (4%) were tested with other mixtures. Ceramic oxides have extraordinary chemical

resistance however, they cannot be used as diaphragms in the form of thin sintered sheets due to its

brittleness and sensitivity to thermal shocks. The thin coating applied on the nickel mesh would be

flexible enough to be bent without cracking around a radius of up to 3 cm. Tests were performed in 30%

KOH at 20°C and the ohmic resistances obtained were as low as 0.027 Ω/cm2 to 0.054 Ω/cm2. These

types of separators are expected to be less expensive than nickel sinters.

2.4.5.2. Organic Materials

Organic polymers are attractive alternatives to asbestos since they can be prepared as microporous

films or be spun into fibers which can be used to make woven cloths, felts, and other fabrics [68].

Although several manufacturing processes can be applied for the preparation of electrolysis separators,

very few organic polymers can withstand the environmental conditions existing in alkaline electrolysis

cells being the polysulfones and polytetrafluorethylene (PTFE) the most suitable ones for this type of

application [68].

The polysulfones (PSU) are a family of polymers composed of sulfone groups that provide thermal

stability and high-temperature rigidity, and ether groups that provide toughness. Hydrolytic stability of

all structure bonds is required for chemical stability on alkaline or acid environments even at high

temperatures. Although polysulfones present great thermal, structural, oxidative, and hydrolytic stability

when prepared as fibers, their maximum service temperature is rather low when compared to the usual

temperatures for alkaline electrolysis and the material was proven to be hydrophobic [68], hence the

PSU by themselves aren’t suitable for this type of applications. There are two main approaches to

improve the polysulfone’s wettability: grafting of hydrophilic functional groups directly to polymer chains

or impregnation of hydrophilic fillers into the polymer matrix directly in the course of formation of the

porous structure. Unfortunately, the improvements provided by the first approach are kept only for a

short time under the usual conditions of alkaline water electrolysis and so, it is not suitable for this type

of application. The second approach was proven to be a good alternative by several authors [64], [71]–

[74].

The most famous PSU-based separators for AEC are those where zirconia (ZrO2) is used as filler [64],

[71], [72], [75]. Zirfon® is a commercially available porous composite separator composed of a

polysulfone matrix and ZrO2 as hydrophilic filler manufactured based on the film-casting technique

(Table 7). Zirfon® contains 85 wt% of hydrophilic ZrO2 powder with a high specific surface area of about

22 m2/g as a hydrophobic agent and 15 wt% polysulfone, which gives the material its mechanical

strength [2]. The separator is highly stable in concentrated KOH solutions, up to 6 M KOH, at

temperatures up to 110ºC [76], [77]. With increasing loadings of zirconia, the structure of the diaphragm

becomes denser meaning that the total porosity and the average pore diameter decrease however, even

38

with a high loading of ZrO2 the separator remains very flexible with attractive mechanical properties [64].

The addition of the ZrO2 filler to the polysulfone matrix significantly increases the wettability of the

separator. Vermeiren et al. [64] studied the influence of the amount of zirconia in the separator wettability

and concluded that the higher the content in ZrO2, the greater the amount of water absorbed by the

separator and the faster the separator reaches its wettability limit. Experimental data regarding the

variation of the separator resistance with temperature of a 0.5 mm thick Zirfon® separator containing 85

wt% of ZrO2 in 30 wt% KOH electrolyte is also presented in their work. The separator resistance sharply

decreases with the increase in working temperature with a value of around 0.13 Ω ∙ cm2 at 80 ºC. In

addition, it is reported that no visible degradation was exhibited after the thick Zirfon® separator was

soaked in 25 wt% KOH at 90ºC for several weeks. The tested membranes also presented very low gas

crossover not permeating hydrogen at pressure differences up to 5.5 bar allowing gas purities higher

than 99.9% under usual operating conditions [64].

Table 7 - Zirfon® PERL properties [75], [76], [78]

Property Value Unit

Weight Density 1±0.2 g/cm3

Thickness 500±50 µm

Maximum temperature for continuous operation 110 ºC

Thermal Stability at 100 ºC (shrinkage) <1.5 %

Porosity 50±10 %

Tortuosity 2.04±0.14 —

Bubble Point 2±1 bar

Pore size 0.15±0.05 µm

Gas Permeability at 5 bar 4±1 l/min∙cm2

Ionic Resistance at 30 ºC in 30 wt% KOH <0.3 Ω ∙ cm2

Maximum Current 20 000 A/m2

Maximum Electrolyte Concentration (NaOH or KOH) 30 wt%

Life expectancy under normal operating conditions >5 years

Lee at al. [72] synthetized a porous composite diaphragm composed of a polysulfone network and ZrO2

(80 wt%) as the inorganic filler and reported it to have increased performance when compared to thick

Zirfon®. Polyvinylpyrrolidone (PVP) was also added to the composition of the separator in order to

eliminate macrovoids and better distribute the porosity as well as increase the mechanical properties of

the separator. A separator with the same composition but without PVP was also tested and results

showed that the separator exhibits a gradual decrease in the bubble point pressure and area resistance

as well as increases in hydrogen permeability, indicating gradual degradation. The difference in ohmic

resistance between the synthetized separator (≈0.2 Ω ∙ cm2) and the Zirfon® separator (≈0.3 Ω ∙ cm2)

with the same 500 µm thickness was not significant. A thinner separator with a thickness of 300 µm was

also tested. The bubble point pressure decreased to 1.3 bar, the area resistance also decreased to 0.1

Ω ∙ cm2 and the hydrogen permeability remained approximately the same of that of the 500 µm thick

separator (3×10-12 mol bar-1 s-1 cm-1 and 2×10-12 mol bar-1 s-1 cm-1, respectively). The authors attribute the

39

increased separator performance to the elimination of a PSU-rich surface layer, the minimization of the

number of large pores with diameters of around 1 µm due to the introduction of PVP and the reduction

of the separator thickness to 300 µm.

The Moscow Power Engineering Institute patented a separator based on a polysulfone matrix with TiO2

as hydrophilic filler [73], [74]. TiO2 is reported to be a cheaper and more available filler alternative when

compared to zirconia. Kuleshov et al. [74] compared the porosity and specific conductivity of several

types of separators namely an asbestos cardboard diaphragm, a 100 wt% polysulfone separator, a

separator based on a polysulfone matrix with TiO2 as a hydrophilic filler, and a separator based on a

polysulfone matrix with TiO2 as hydrophilic filler and PVP as pore-forming agent. All samples were tested

in 6 M KOH electrolytes at 18 ºC and 80 ºC. The conductivity of the polysulfone matrix with TiO2 (0.065

S/cm at 18 ºC; 0.256 S/cm at 80 ºC) was reported to be very close to the one of the asbestos cardboard

diaphragm (0.069 S/cm at 18 ºC; 0.270 S/cm at 80 ºC) and the specific conductivity of the PSU + TiO2

+ PVP (0.075 S/cm at 18 ºC; 0.287 S/cm at 80 ºC) was reported to be higher. The diaphragm constituted

by 27.15 wt% PSU, 63.35 TiO2 and 9.5 wt% PVP was the best performer of the tested samples with a

specific conductivity of 0.287 S/cm at 80 ºC in 6M KOH electrolyte ensuring lower cell voltage than

asbestos with excellent gas-separation properties.

PTFE is a polymer that can be found as fibers, porous films, or sheets. It has a maximum working

temperature of around 288 ºC and excellent chemical stability being highly resistant to alkaline media.

According to some authors, PTFE is stable up to 260 ºC in 50 wt% KOH [70] however reports show that

Teflon fibers can be dissolved under certain conditions at 200 ºC in 60 wt% KOH. Furthermore, it was

reported that large surface area fibers might favor the dissolution.

One of the biggest setbacks associated with the use of PTFE as separator material is its lack of wettability

which leads to the formation of gas bubbles within the separator and on its surface, thus increasing the

ohmic resistance of the separator and reducing the gas purity. To increase the wettability of Teflon,

surface treatments like the grafting of stable anionic and cationic groups into the Teflon are usually

applied [68]. However, only anionic groups were proven to be stable in alkaline media at high

temperatures typical of AEC. A different technique for improving the wettability of Teflon is the in-situ

formation of potassium titanates in porous Teflon sheets. A Teflon sheet with potassium titanate crystals

inside its pores was proven to have good hydrophilic properties, average pore sizes of 1µm, and better

ionic conductivity that the original Teflon sheet with ionic resistance values of 0.28 Ω/cm2 for 0.4 mm

thick separators and 0.7 Ω/cm2 for 0.7 mm thick separators, in 30 wt% KOH at 25 ºC [79]. The potassium

titanate impregnation technique not only provides better results in increasing the PTFE wettability but it

is also easier to apply than grafting of stable anionic groups [68].

40

3 Model Development

In this chapter, a simulation tool for alkaline electrolysis cells is proposed. This chapter is divided into

four main subchapters. The first subchapter intends to model gas production at each electrode by

characterizing the partial pressure of the produced gases and its production rate. The second

subchapter models the reversible potential and overpotentials. The third one presents a set of auxiliary

modules that model some crucial parameters to be used on the main model. The fourth subchapter

presents several different ways of computing cell efficiencies along with their different meanings.

The model is based on a physics model with the intent of simulating the physical and electrical behaviour

of the alkaline electrolyser. This model evaluates separately the main physical phenomena taking place

in the electrolysis process inside the cell.

The electric response of an electrolysis cell is relatively fast (~50ms) when compared to the response

time of the global system and so, the static approach is the most suitable one to model the

electrochemical response of an electrolysis cell. In this type of model, the cell’s electrical equilibrium is

assumed to be established instantaneously for each time step meaning that the cell is considered as a

unique non-linear resistive element. Therefore, this model provides the mathematical description of the

current-voltage cell characteristics in the form of polarization curves [24].

3.1. Electrodes module

This module models the gas production at each electrode by characterizing the oxygen and the

hydrogen partial pressure and production rate.

3.1.1. Cathode Module

At the cathode, water is reduced to produce hydrogen (H2) and hydroxide ions (OH-). Plain water

The mole balance equations for the cathode can be written as

𝑑 𝑁

𝑑𝑡= − − (10)

𝑑 𝑁

𝑑𝑡= − + (11)

where , , and are the cathode inlet and outlet molar flow rates of water and hydrogen

respectively, is the molar flow rate of water consumed at the cathode and is the molar flow

rate of hydrogen produced at the cathode.

41

According to Faraday’s law, the molar flow rate of hydrogen produced at the cathode, , and the

molar flow rate of water consumed at the cathode, , can be expressed as

=𝐼

2𝐹 (12)

=𝐼

2𝐹 (13)

The molar fluxes of hydrogen and water at the cathode can be obtained from equations (14) and (15).

=

𝐴=

𝑖

2𝐹 (14)

=

𝐴=

𝑖

2𝐹 (15)

Being 𝐴 the electrode geometric area and 𝐹 the Faraday constant. Taking into account that the sum of

the molar fractions of hydrogen and water at the cathode is one (Eq. (16)) and assuming that the gases

present at the cathode are ideal and uniformly distributed and that the pressure is uniform throughout

the electrode gas flow channels, the effective partial pressure of hydrogen at the cathode is given by

equation (17).

𝜒 + 𝜒 = 1 (16)

𝑝 =𝑝

𝜒 (1 − 𝜒 ) (17)

Being 𝑝 the partial pressure and 𝜒 the molar fraction of the components. Adding the assumption of one-

dimensional geometry to the aforementioned assumptions, it is reasonable to use a simplification of the

Stefan-Maxwell equation [80] that describes diffusion in multicomponent systems where the molar flux

of water vapor normal to the surface of the cathode is set to zero [81]. The 1D gradient of the molar

fraction of water at the cathode can then be expressed by equation (18) that can be integrated from the

anode channel to the catalyst surface yielding equation (19).

𝑑 𝜒

𝑑𝑡 =

𝜖

𝜏∙

𝑅 𝑇 𝑖

2 𝐹 𝑃 𝐷 (18)

𝜒 =𝜖

𝜏∙ exp

𝑅 𝑇 𝑖 𝑙

2 𝐹 𝑃 𝐷 (19)

Being 𝜖 the cathode porosity, 𝜏 the cathode tortuosity, 𝑙 the cathode channel-to-catalyst

distance, 𝑅 the ideal gas constant, 𝑇 the working temperature, 𝑖 the current density, 𝑃 the total

pressure and 𝐷 the effective binary diffusion coefficient (H2O and H2) which is looked into with detail

in Section 3.3.1.

The partial pressure of hydrogen [5] can then be written as

𝑝 =𝜖

𝜏∙ exp

𝑅 𝑇 𝑖 𝑙

2 𝐹 𝑃 𝐷− 1 ∙ 𝑃 , (𝑇) (20)

where 𝑃 , (𝑇) is the temperature-dependent electrolyte saturated vapour pressure that can be

obtained from correlation equations derived from experimental data by Balej [82].

42

3.1.2. Anode Module

After being produced at the cathode, hydroxide ions (OH-) migrate to the anode where they are oxidized

to produces oxygen, electrons and water. The mole balance equations for the anode can be written as

𝑑 𝑁

𝑑𝑡= − − (21)

𝑑 𝑁

𝑑𝑡= − + (22)

where , , and are the anode inlet and outlet molar flow rates of water and oxygen

respectively, is the molar flow rate of water consumed at the anode and is the molar flow

rate of oxygen produced at the anode. According to Faraday’s law, the molar flow rate of oxygen

produced at the anode, , and the molar flow rate of water consumed at the anode, , can be

expressed as

=𝐼

4𝐹 (23)

=𝐼

2𝐹 (24)

The molar fluxes of hydrogen and water at the cathode can be obtained from equations (25) and (26).

=

𝐴=

𝑖

4𝐹 (25)

=

𝐴=

𝑖

2𝐹 (26)

Following the same steps and assumptions used for the anode module, the molar fraction of water at

the cathode can then be expressed by equation (27) [80], [81].

𝜒 =𝜖

𝜏∙ exp

𝑅 𝑇 𝑖 𝑙

2 𝐹 𝑃 𝐷 (27)

Being 𝜖 the anode porosity, 𝜏 the anode tortuosity, 𝑙 the anode channel-to-catalyst distance and

𝐷 the effective binary diffusion coefficient (H2O and O2) which is looked into with detail in Section

3.3.1.

The partial pressure of oxygen [5] can then be written as

𝑝 =𝜖

𝜏∙ exp

𝑅 𝑇 𝑖 𝑙

2 𝐹 𝑃 𝐷− 1 ∙ 𝑃 , (𝑇) (28)

3.2. Voltage module

In an alkaline electrolysis cell, the cell voltage can be described by equation (9). The reversible potential,

the activation overpotential, the concentration overpotential and the ohmic overpotential will be modelled

separately.

43

3.2.1. Reversible Potential

The reversible potential (Eq. (29)) is the minimum required potential to apply between the electrodes of

the electrolytic cell for the water dissociation reaction to occur. The first term, 𝑉 , accounts for the

reversible potential in reference conditions and the second term accounts for the deviation caused by

the operation in different conditions from the standard ones [24].

𝑉 = 𝑉 + 𝑅 𝑇

2 𝐹ln

𝑝 ∙ 𝑝

𝑎 (29)

With 𝑅 being the ideal gas constant, 𝐹 being the Faraday’s constant, 𝑝 being the partial pressure of

species 𝑖 and 𝑎 being the thermodynamic activity water. The thermodynamic water activity in KOH

and NaOH can be calculated from correlation equations derived from experimental data by Balej [82].

The reversible potential in standard conditions, 𝑉 , can be obtained from physical or empirical models

[5], [23], [24], [83]–[89]. The physical model (Eq. (30)) which is used in this work [5], [83]–[86] can itself

be split into two terms where the first one, 𝑉 , accounts for the reversible cell voltage at 25 ºC and 1

atm (1.229 V) and the second one accounts for the change in the reversible voltage at temperatures

different from the standard reference temperature [5], [24].

𝑉 = 1.229 + 𝑇 − 𝑇 × Δ𝑆

2 𝐹 (30)

Where 𝑇 is the operating temperature, 𝑇 is 25 ºC and 𝛥𝑆 /2𝐹 (≈ −8.5 × 10 − 3 𝐽 𝑚𝑜𝑙 𝐾 ) is the

standard state entropy change.

The reversible potential of the cell, 𝑉 , falls as the temperature increases due to the entropy

contribution, and rises as the pressure increases due to the free energy of the produced gases.

3.2.2. Activation overpotential

The activation overpotential is here modelled by the Butler-Volmer equation (Eq.(31)) which applies the

transition state theory assuming a single rate-limiting activated step in the complex reaction mechanism

at each electrode [5].

𝜂 = 𝑅 𝑇

𝛼 𝐹ln

𝑖

𝑖 (31)

Where 𝑘 represents the anode or the cathode, 𝛼 is the charge transfer coefficient of electrode 𝑘 and

𝑖 is the exchange current density of electrode 𝑘. The term that multiplies the logarithm is the Tafel

Slope. Several values for Tafel Slopes of different electrodes and electrocatalysts can be found in Table

4 and Table 5.

In this model, it is assumed that both electrodes have the same area and so, conserving the cell current

density implies equal current densities at both electrodes.

44

The values of the kinetic electrode parameters 𝛼 , 𝛼 , 𝑖 and 𝑖 , are dependent on the catalyst

material, operating temperature and the electrolyte type and weight-weight percentage (wt.%) [6].

The reference exchange current densities, 𝑖 , and 𝑖 ,

, are the values of the exchange current

density at a reference pressure and temperature for each one of the electrodes which is also dependent

on the available catalyst surface area. A physics model for the exchange current density is presented in

equation (32). This model is derived from the Arrhenius plot that relates the exchange current density

and the temperature.

𝑖 = exp −Δ𝐺

𝑅

1

𝑇−

1

𝑇 𝑖 , (32)

Where 𝑘 represents the anode or the cathode, Δ𝐺 is the free activation energy for the reaction taking

place at electrode 𝑘 and 𝑖 , is the exchange current density of the electrode 𝑘 at temperature 𝑇 .

The free activation energy for each reaction is the slope of the Arrhenius plot.

The charge transfer coefficient variation with temperature within the usual operating temperature range

is usually low. Its temperature-dependent variation is usually modelled through equations obtained from

the fitting of experimental results found in the literature [6], [23].

During the operation of the electrolysis cell, bubbles form in the electrodes surface. The detachment of

the bubbles occurs only after they reach a certain size leading to the accumulation of bubbles in the

electrode surface isolating the covered fraction of the surface from the reacting species and rendering

it inactive. The bubble coverage, Θ , reduces the active area of the electrodes which reduces the

exchange current density contributing to an increase in the activation voltage.

𝜂 = 𝑅 𝑇

𝛼 𝐹ln

𝑖

𝑖 (1 − Θ ) (33)

Calculating the bubble coverage of an electrode is a rather complex process since the bubble effect

depends on the surface characteristics of the electrode, the surface tension of the electrolyte, and the

circulation of the electrolyte, all of which influence the size at which the bubbles detach from the surface.

Vogt and Balzer [90] developed a mathematical model to predict the bubble coverage however, as

regarded by its authors, it is merely indicative. Therefore, an empirical model has been employed to

model the bubble coverage as a function of the current density, temperature, and pressure (Eq. (34)).

Θ = −97.25 + 182𝑇

𝑇− 84

𝑇

𝑇×

𝑖

𝑖

.

× 𝑃

𝑃 − 𝑃 , (𝑇) (34)

In equation (34), 𝑘 represents the electrode (anode or cathode), Θ is the bubble coverage, 𝑇 is the

working temperature at which we are calculation the bubble coverage, 𝑖 is the current density, 𝑖 is the

current density at which the electrode is 100% covered in bubbles and 𝑇 is the temperature at which

𝑖 was evaluated. The last term of equation (34) is a corrective term that allows this expression to be

valid at pressures different from atmospheric pressure. In this term, 𝑃 is the pressure at electrode 𝑘

and 𝑃 , is the temperature dependent electrolyte saturated vapor pressure.

45

According to Vogt and Balzer [90], the limiting current density, 𝑖 has a theoretical maximum value of

30 A/cm2 however this value may be much smaller since it depends on the type of electrolyte used, the

electrolyte concentration, electrode materials and surface characteristics and working pressure.

The total activation voltage is given by the sum of the anodic activation voltage and the cathodic

activation voltage translated by equation (35).

𝜂 = 𝜂 + 𝜂 (35)

3.2.3. Concentration overpotential

The mass flow through porous electrodes can be described as a diffusion phenomenon that can be

modelled by Fick’s Law. Applying Fick’s Law at the interface between the electrode and the electrolyte,

the molar concentration of the produced gases can be described by equations (36) and (37).

= 𝐷 𝐶 , − 𝐶 ,

𝛿 ⇒ 𝐶 , = 𝐶 , +

𝛿

𝐷 (36)

= 𝐷 𝐶 , − 𝐶 ,

𝛿 ⇒ 𝐶 , = 𝐶 , +

𝛿

𝐷 (37)

Where 𝛿 and 𝛿 are the anode and cathode thickness, respectively, 𝐶 , is the oxygen

concentration in the electrolyte on the anode side, 𝐶 , is the hydrogen concentration in the electrolyte

on the cathode side, 𝐶 , is the oxygen concentration in the anode channels and 𝐶 , is the hydrogen

concentration in the cathode channels. At the typical alkaline electrolyser operation pressures, the molar

concentration of the gaseous species is related to is partial pressure through the ideal gas law yielding

𝐶 , = 𝑃 𝜒

𝑅 𝑇+

𝛿

𝐷 (38)

𝐶 , =𝑃 𝜒

𝑅 𝑇+

𝛿

𝐷 (39)

The concentration overpotential can now be calculated from equation (40).

𝜂 =𝑅 𝑇

2𝐹ln

𝐶 ,

𝐶 ,

+ 𝑅 𝑇

4𝐹ln

𝐶 ,

𝐶 ,

(40)

𝐶 , and 𝐶 , represent the hydrogen and oxygen concentrations calculated at low current densities

because at low current densities the concentration gradient in the vicinity of the electrodes is negligible.

3.2.4. Ohmic overpotential

To model the ohmic overpotential, the electrodes, the electrolyte and the separator are modelled as

electrical resistances which are considered to be in series. The effective ohmic resistance of the cell

𝑅 can be obtained from the sum of the electrodes resistance 𝑅 , the electrolyte resistance

𝑅 , and the separator resistance 𝑅 (Eq.(41)). The ohmic overpotential can be then

obtained from the product of the effective ohmic cell resistance and the current (Eq. (42)).

𝑅 = 𝑅 + 𝑅 + 𝑅 (41)

46

𝜂 = 𝐼 ∙ 𝑅 (42)

Assuming that each of the main components of the cell behaves as an ohmic resistance implies that the

resistance of each element depends only on its material’s properties and on its geometry through the

relation 𝑅 = 𝜌 𝑙/𝐴 where 𝜌 is the material’s resistivity, 𝑙 is the path length through which the current

flows and 𝐴 is the cross-sectional area of the component. This relation can only be used as it is for simple

geometries. For more complex geometries such as porous materials, an equivalent 𝑙/𝐴 must be found.

Representing solid materials with dell-defined charge carriers such as the electrodes and the separator

and liquid electrolytes in which charge build-up is negligible by a temperature dependent resistivity is a

very reasonable approximation.

The intrinsic resistivities of the porous electrodes and separator, 𝜌0𝑒𝑙𝑒𝑐𝑡𝑟𝑜𝑑𝑒 and 𝜌

0𝑠𝑒𝑝𝑎𝑟𝑎𝑡𝑜𝑟 , are not affected

by its porous nature or by the presence of bubbles that may interrupt the electrolyte conduction path. In

fact, these factors only affect the detailed geometry of the elements.

The electrodes, the separator and the bubble zone are inhomogeneous conductors and therefore, there

is not an exact theory to describe their properties. These components’ properties must be described by

approximate models. To obtain conductivity of porous materials, many approximate models have been

developed to describe the conduction properties of a porous conductive matrix with inclusions of other

material in its pores in terms of an effective conductivity 𝜌𝑒𝑓𝑓

. This approach modifies the intrinsic

resistivity 𝜌0 into an effective resistivity 𝜌 that accounts for the effects of the porosity so that the

geometric factors used to calculate the materials’ resistance can be taken as the gross external

dimensions of the element. At low concentrations, most models provide approximately the same results

regardless of the detailed arrangement and morphology of the inclusions[60]. Here it is assumed that

the electrolyte does not penetrate the electrodes pores and that the pores are insulating relative to the

electrode material and therefore, its conductivity is zero [60]–[62]. A similar approach is taken regarding

the bubble zone where the bubbles are considered to be insulating relatively to the electrolyte itself.

Solving the Bruggeman’s equation for the conductivity of a two-phase conductor [5], [60],

1 − 𝜑 = 𝜎 − 𝜎

𝜎 − 𝜎

𝜎

𝜎

/

(43)

where 𝜑𝑖𝑛𝑐𝑙𝑢𝑠𝑖𝑜𝑛

is the volume fraction of the inclusion phase, 𝜎𝑖𝑛𝑐𝑙𝑢𝑠𝑖𝑜𝑛 is the conductivity of the inclusion

phase, 𝜎0 is the conductivity of the matrix and 𝜎𝑒𝑓𝑓 is the effective conductivity of the porous material.

Setting the conductivity of the inclusion phase 𝜎𝑖𝑛𝑐𝑙𝑢𝑠𝑖𝑜𝑛 to zero, the relation between the effective

conductivity of the porous material and the matrix conductivity is obtained (Eq. (44)).

𝜎

𝜎= (1 − 𝜑 ) / (44)

The separator cannot be modelled as a conducting matrix with insulating pores but instead it must be

modelled as an insulating matrix with pores filled with conducting material. The pore filling material of

the separator cannot be considered to be at low concentrations as it was the case for the electrodes

47

and so, a different approach to model the separator must be taken. In this model, the separator will be

modelled based on an assumed microstructure [5].

3.2.4.1. Electrodes

The resistance of both electrodes can be modelled according to equation (45).

𝑅 = 𝜌𝑒𝑓𝑓𝑘

𝛿𝑘

𝐴𝑒

(45)

Being 𝜌 the effective conductivity, 𝛿 the thickness of the electrode, 𝐴 the geometric area of the

electrode and 𝑘 the anode or the cathode.

The effective conductivity of the electrode 𝜌 consists in modifying the intrinsic resistivity of the

electrode material 𝜌 in order to account for the effect of the porosity in this property. According to

equation (44), the effective resistivity of the electrode can be written as

𝜌𝑒𝑓𝑓𝑘 =

𝜌0𝑘

(1 − 𝜖𝑘)3

2

[1 + 𝜅𝑘(𝑇 − 𝑇𝑟𝑒𝑓)] (46)

Where 𝑘 the anode or the cathode, 𝜌 is the resistivity of the 100%-dense electrode material at

reference temperature 𝑇 , 𝜖 is the electrode porosity and 𝜅 is the temperature coefficient of

resistivity. The term 𝜌 [1 + 𝜅 (𝑇 − 𝑇 )] provides a good approximation for the resistivity of most of

the used electrode materials at the usual working temperatures however, in some cases, a better

approximation of the electrode’s resistivity can be obtained from empirical correlations that should

replace this term. The resulting electrode resistance can be written as

𝑅𝑒𝑙𝑒𝑐𝑡𝑟𝑜𝑑𝑒𝑠 =𝜌

0𝑎𝑛

(1 − 𝜖𝑎𝑛)3

2

𝛿𝑎𝑛

𝐴𝑒

1 + 𝜅𝑎𝑛 𝑇 − 𝑇𝑟𝑒𝑓 +𝜌

0𝑐𝑎𝑡

(1 − 𝜖𝑐𝑎𝑡)3

2

𝛿𝑐𝑎𝑡

𝐴𝑒

1 + 𝜅𝑐𝑎𝑡 𝑇 − 𝑇𝑟𝑒𝑓 (47)

3.2.4.2. Electrolyte

The resistance of the electrolyte varies in space inside the electrolyser. In the region close to the

electrodes, the bubble formation is significant so the electrode resistivity is higher. In the bubble free

zone, the conductivity of the electrolyte is the expectable conductivity of the electrolyte solution at that

working temperature. Therefore, the electrolyte resistance will be modelled as the sum of the electrolyte

resistivity in the bubble free zone and the electrolyte resistivity in the bubble zone (Eq. (48)) [5].

𝑅𝑒𝑙𝑒𝑐𝑡𝑟𝑜𝑙𝑦𝑡𝑒 = 𝑅𝑒𝑙𝑒𝑐𝑡𝑟𝑜𝑙𝑦𝑡𝑒𝑎𝑛 + 𝑅𝑒𝑙𝑒𝑐𝑡𝑟𝑜𝑙𝑦𝑡𝑒

𝑐𝑎𝑡 = 𝑅𝑏𝑓𝑎𝑛 + 𝑅𝑏𝑧

𝑎𝑛 + 𝑅𝑏𝑓𝑐𝑎𝑡 + 𝑅𝑏𝑧

𝑐𝑎𝑡 (48)

Being 𝑅 and 𝑅 the electrolyte resistance on the anode side and on the cathode side,

𝑅 and 𝑅 the resistance of the electrolyte in the bubble-free zone of the anode side and the cathode

side, and 𝑅 and 𝑅 the resistance of the electrolyte in the bubble zone of the anode side and the

cathode side, respectively.

48

In the bubble zone, the current flows in a path length 𝛽 which corresponds to the bubble zone thickness

(Eq. (49)). Experimental studies found that the bubble zone thickness in an electrolysis cell ranges from

0.4 mm to 0.6 mm [5], [91]. Consequently, in the bubble-free zone, the current flows in a path length

𝑙𝑒−𝑠 − 𝛽 which corresponds to the distance between the bubble zone and the separator (Eq. (50)).

𝑅𝑏𝑧 = 𝜌𝑏𝑧

𝛽

𝐴𝑒

(49)

𝑅𝑏𝑓 = 𝜌𝑒𝑙

𝑙𝑒−𝑠 − 𝛽

𝐴𝑒

(50)

Being 𝜌 the resistivity of the electrolyte in the bubble zone and 𝜌 the resistivity of the bubble-free

electrolyte.

To obtain the effective resistivity of the electrolyte on the bubble zone, the Bruggeman’s model (Eq. (44))

is applied once again yielding

𝜌𝑏𝑧

= 𝜌

𝑒𝑙

1 − 𝜑𝑏𝑢𝑏𝑏𝑙𝑒𝑠

3/2 (51)

The effective resistivity of the bubble-free electrolyte 𝜌 can be obtained from empirical relations such

as the ones summarized by Damien Le Bideau et al. [29].

The total electrolyte resistance is then

𝑅𝑒𝑙𝑒𝑐𝑡𝑟𝑜𝑙𝑦𝑡𝑒 = 𝜌𝑒𝑙

𝑙𝑎𝑛−𝑠 − 𝛽

𝑎𝑛

𝐴𝑒

+ 1

1 − 𝜑𝑎𝑛𝑏𝑢𝑏𝑏𝑙𝑒𝑠

32

𝛽𝑎𝑛

𝐴𝑒

+ 𝑙𝑐𝑎𝑡−𝑠 − 𝛽

𝑐𝑎𝑡

𝐴𝑒

+ 1

1 − 𝜑𝑐𝑎𝑡𝑏𝑢𝑏𝑏𝑙𝑒𝑠

32

𝛽𝑐𝑎𝑡

𝐴𝑒

(52)

The relationship between 𝜑𝑏𝑢𝑏𝑏𝑙𝑒𝑠

and Θ under specific conditions applicable to electrolysers operating

conditions was obtained by Vogt [92] and can be written as

𝜑𝑏𝑢𝑏𝑏𝑙𝑒𝑠

= 3/2 Θ (53)

3.2.4.3. Separator

The role of the separator in an electrolysis cell is to prevent gas mixing while providing a continuous

ionic conduction path. In order to comply with these requirements, electrolyser separators are usually

made of porous structures with pore volume fraction 𝜖 and effective pore length 𝑙 . The effective pore

length can be obtained from the product of the separator tortuosity 𝜏 by the separator thickness 𝛿 (Eq.

(54)).

𝑙 = 𝜏 𝛿 (54)

Following one of the main global assumptions of this model, uniform distribution of current across the

geometric area of the separator 𝐴 , the equivalent conductor area of the porous structure is given by

𝐴 𝜔 𝜖 /𝜏 and the equivalent conductor length is given by the effective length (Eq. (54)).

The ratio between the volume of electrolyte absorbed inside the pores of the separator to the total

volume of the separator is the wettability factor 𝜔 . The effective porosity of the separator is then given

by the product 𝜔 𝜖 . The separator resistance then yields

49

𝑅 = 𝜌 𝜏 𝛿

𝐴 𝜔 𝜖 (55)

In some cases, the results provided by equation (55) are an overestimation of the real values of the

separator resistance hence, when available, the use of empirical correlations to estimate the separator

resistance is highly advised.

3.3. Auxiliary Modules

In this section, models to compute diffusion, electrolyte saturated vapor pressure, water activity and

electrolyte electrical conductivity are presented.

3.3.1. Diffusion

The diffusion of a molecular species with mean free path 𝜆 through a porous medium with average pore

radius can be modelled by two main mechanisms depending on the ratio between the mean free path

and the mean pore radius [5]. Molecular diffusion mechanisms are used for 𝜆 ⁄ ≫ 1 and Knudsen

diffusion mechanisms are used for 𝜆 ⁄ ≪ 1. Molecular diffusion explains the phenomena related to the

flux of molecules from a region of higher concentration to a region of lower concentrations in order to

try to cancel the concentration gradient. Knudsen diffusion [93] is used to explain the diffusion

mechanisms when the molecules collisions with the pore walls occur much more frequently than

collisions with other molecules. In most porous media, both mechanisms are relevant and so, the

effective diffusion coefficients at the cathode and at the anode can be expressed by equations (56) and

(57) [93], [94].

𝐷 =𝜖

𝜏

1

𝐷+

1

𝐷, (56)

𝐷 =𝜖

𝜏

1

𝐷+

1

𝐷, (57)

Where 𝐷 is the effective molecular diffusion coefficient for the H2–H2O binary system, 𝐷 is

the effective molecular diffusion coefficient for the O2–H2O binary system and 𝐷 , is the effective

Knudsen diffusion coefficient for water. The absence of the Knudsen diffusivities for H2 and O2 in water

in equations (56) and (57) is a consequence of the assumption of uniform pressure that implies equi-

molar counter diffusion yielding zero total molar flux. The effective Knudsen diffusion can be determined

𝐷 , =4

3

8 𝑅 𝑇

𝜋𝑀 (𝑚 /𝑠) (58)

Where is the mean pore radius (m) and 𝑀 is the molecular weight (kg/kmol).

To obtain the molecular binary diffusion coefficients, 𝐷 and 𝐷 , the Chapman-Enskog theory

of ideal gas can be applied (Eq. (59) and (60)) [95].

50

𝐷 = 0.00266 𝑇 /

𝑃 𝑀 𝜎 Ω (𝑐𝑚 /𝑠) (59)

𝐷 = 0.00266 𝑇 /

𝑃 𝑀 𝜎 Ω (𝑐𝑚 /𝑠) (60)

𝑀 = 2 1

𝑀+

1

𝑀 (61)

𝑀 = 2 1

𝑀+

1

𝑀 (62)

Being 𝜎 and 𝜎 the mean molecular radii of species H2-H2O and O2-H2O (Å), respectively, Ω

the dimensionless diffusion collision integral and 𝑀 , 𝑀 and 𝑀 the molecular weights of water,

hydrogen and oxygen, respectively.

According to Poling et al. [95], the dimensionless diffusion collision integral Ω can be expressed as

Ω =1.06036

𝑡 .+

0.19300

exp(0.47635 t)+

1.03587

exp(1.52996 t)+

1.76474

exp(3.89411 t) (63)

𝑡 =𝑘𝑇

𝜀 (64)

𝑡 =𝑘𝑇

𝜀 (65)

The Lennard-Jones energies 𝜀 and 𝜀 can be calculated according to [5], [95]

𝜀 = (𝜀 𝜀 ) . (66) 𝜀 = (𝜀 𝜀 ) . (67)

The Lennard-Jones potentials, 𝜀 /𝑘, for oxygen, hydrogen and water are, respectively, 106.7 K, 59.7 K

and 809.1K [95].

The mean molecular radii 𝜎 and 𝜎 can be obtained from equations (68) and (69) where

𝜎

= 2.827 Å, 𝜎

= 3.467 Å and 𝜎 = 2.641 Å [5], [95].

𝜎 =𝜎

+ 𝜎

2 (68)

𝜎 =𝜎

+ 𝜎

2 (69)

The values of the effective diffusion coefficients can be obtained by solving equations (56) and (57).

3.3.2. Electrolyte Saturated Vapor Pressure

Balej [82] condensed several series of experimental data and converted them to the same concentration

scale, namely molarity, in order to obtain water vapour partial pressures in potassium and sodium

hydroxide solutions over wide concentration and temperature ranges. The obtained correlations are

dependent on temperature and molarity (Eq. (70)) which, in turn, is dependent on the temperature and

electrolyte density. The electrolyte density is also a temperature dependent property that can be

modelled using correlations obtained from experimental data [29], [96], [97].

𝑚 = 𝑌𝑑

𝑀 (70)

51

3.3.2.1. Potassium Hydroxide

An experimental correlation for the density of the potassium hydroxide solution is presented by Gilliam

et al. [96] (Eq. (71)). Gilliam at al. correlation is applicable for temperatures ranging from 0 ºC to 200 ºC

and molar fractions ranging from 0 to 0.5 providing results with deviations inferior to 1% [29] compared

to experimental data from Zaytcev [97] and Mashovets et al. [98].

𝑑 = (−3.25 × 10 T + 0.111 T + 1.00171 × 10 ) exp(0.86 × 𝑌 ) [𝑘𝑔/𝑚 ] (71)

Balej’s [82] experimental correlation for the water vapour partial pressures (in bar) in potassium

hydroxide solutions is applicable for temperatures between 0 ºC and 300 ºC and for molar

concentrations between 0 and 18 M (Eq. (72)).

𝑙𝑜𝑔 𝑃 , = −0.01508 𝑚 − 0.0016788 𝑚 + 2.25887 × 10 𝑚 + (1 − 0.0012062𝑚

+ 5.6024 × 10 𝑚 − 7.8228 × 10 𝑚 ) × (35.4462 −3343.93

𝑇

− 10.9 𝑙𝑜𝑔 (𝑇) + 0.0041645 𝑇)

(72)

3.3.2.2. Sodium Hydroxide

An experimental correlation for the density of the sodium hydroxide solution (kg/m3) is presented by

Zaytcev [29], [97] (Eq. (73)). This correlation is applicable for temperatures ranging from 0 ºC to 200 ºC

and molar fractions ranging from 0 to 0.5 providing results with deviations inferior to 1% compared to

experimental data [29].

𝑑 = (1000 + 6.2 × 10 T − 3.55 × 10 T ) × 10^(0.425 − 1.15 × 10 𝑇) × 𝑌 (73)

Balej’s [82] experimental correlation for the water vapour partial pressures (in bar) in sodium hydroxide

solutions is applicable for temperatures between 0 ºC and 300 ºC and for molar concentrations between

0 and 25 M (Eq. (74)).

𝑙𝑜𝑔 𝑃 , = −0.010986 𝑚 − 1.461 × 10 𝑚 + 2.03528 × 10 𝑚 + (1

− 1.34141 × 10 𝑚 + 7.07241 × 10 𝑚 − 9.5362 × 10 𝑚 ) × (35.4462

−3343.93

𝑇 − 10.9 𝑙𝑜𝑔 (𝑇) + 0.0041645 𝑇)

(74)

3.3.3. Thermodynamic Water Activity

In order to obtain the water activity in potassium and sodium hydroxide solutions over wide concentration

and temperature ranges, Balej [82] followed a similar procedure as for the saturated vapour pressure

which also led to temperature and molar concentration dependent correlations.

3.3.3.1. Potassium Hydroxide

Balej’s [82] experimental correlation for the water activity (in bar) in potassium hydroxide solutions is

applicable for temperatures between 0 ºC and 150 ºC and for molar concentrations between 0 and 18

M (Eq. (75)).

52

𝑙𝑜𝑔 𝑎 , = −0.02255 𝑚 + 0.001434 𝑚 +1.38 𝑚 − 0.9254 𝑚

𝑇 (75)

3.3.3.2. Sodium Hydroxide

Balej’s [82] experimental correlation for the water activity (in bar) in sodium hydroxide solutions is

applicable for temperatures between 0 ºC and 200 ºC and for molar concentrations between 0 and 25

M (Eq. (76)).

𝑙𝑜𝑔 𝑎 , = −0.01332 𝑚 + 0.002542 𝑚 − 3.06 × 10

+1.5827 𝑚 − 1.5669 𝑚 + 0.021296 𝑚

𝑇

(76)

3.3.4. Electrolyte Electrical Conductivity

3.3.4.1. Potassium Hydroxide

Based on experimental data from several authors, Gilliam et al. [96] derived a temperature and molarity

dependent correlation to model the KOH electrolyte specific electrical conductivity Κ in S/cm (Eq.

(77)). This correlation provides good results when compared to experimental data for temperatures

between 0 ºC and 100 ºC and concentrations between 0 and 12 M [29], [96].

Κ = −2.041 𝑚 − 0.0028 𝑚 + 0.005332 𝑚 𝑇 + 207.2 𝑚

𝑇+ 0.001043 𝑚 − 3 × 10 𝑚 𝑇 (77)

3.3.4.2. Sodium Hydroxide

An experimental correlation to model the specific electrical conductivity of the NaOH electrolyte, Κ

in S/cm, is suggested by Le Bideau et al. [29] (Eq. (78)) based on experimental data from Zaytsev [97]

and Maksimova [99]. This correlation is far more limited and less accurate than the one presented for

the KOH electrolyte due to the reduced amount of experimental data available in the literature and the

fact that there are significant differences in the data provided by different authors.

Κ = −0.457 + 1.02 × 10 ( 𝑇 − 273.15) + 32 𝑌 − 29.9 𝑌 + 7.84 𝑌 (78)

This correlation is applicable for temperatures between 0 ºC and 50 ºC and for mass fractions between

0 and 0.25. Within these ranges, the average deviation from Zaytsev’s experimental data [97] is around

1.5% having a maximum deviation of 11.7% at 50 ºC and 0.08 NaOH mass fraction[29]. However, given

the fact that Maksimova’s data [99] is not in agreement with Zaytsev’s data [97], the correlations presents

an average deviation of 20% from Maksimova’s data [29].

53

3.4. Cell Efficiencies

The proportion of effective voltage to split water in the total voltage applied to the electrolysis cell is

known as voltage efficiency (Eq. (79)) [22].

𝜂 = 𝑉 − 𝑉

𝑉 (79)

The percentage of the theoretical energy needed to break water molecules in the real cell is designated

Faradic efficiency (Eq. (80)) [22]. It is a measure of the cell efficiency from a purely cell voltage point of

view. At 25 ºC, 𝐸 is approximately 1.23 V.

𝜂 = ΔG

ΔG + Losses =

𝐸

𝐸 (80)

Conversely, the thermal voltage (Eq. (81)) translates the need of an additional voltage, above the

reversible voltage, to maintain the thermal balance [22]. Thus, the percentage of the actual energy input

in the cell voltage defines the thermal efficiency. If the cell operates in endothermic mode, this efficiency

can be above 100% as the system may absorb heat from the surroundings (Figure 7). At an operating

temperature of 25 ºC, 𝐸 is around 1.48 V.

𝜂 = ΔH

ΔG + Losses =

𝐸

𝐸 (81)

Figure 7 - Potential for hydrogen production as a function of temperature [22]

Another very useful way to evaluate the efficiency of the system is to compare the produced hydrogen

output with the total electric energy provided to the system in terms of the energy carried by the

hydrogen produced (Eq. (82)) [22].

𝜂 = 𝐻𝐻𝑉

𝑉 × 𝑖 × 𝑡 (82)

In this equation, HHV is the high heating value of one mole of hydrogen (≈283.8 kJ), 𝑉 is the cell

voltage, 𝑖 is the applied current and 𝑡 is the time needed to produce one mole of hydrogen.

An alternative approach to determine the energy efficiency is to subtract the energy losses, 𝐸 , from

the total energy input, 𝐸 (Eq. (83)). Energy losses comprise energy losses in the electric circuits and

54

due to the electrical resistances of the electrode materials, transport related resistances (gas bubbles in

the electrolyte and covering the electrodes, ionic transfer in the electrolyte and through the membrane,

etc.), and electrochemical reaction resistances.

𝜂 = 𝐸 − 𝐸

𝐸 = 1 −

𝐸

𝐸 (83)

55

4 Experimental Approach

In this chapter, description of the construction of the lab-scale electrolyser and the experimental setup

and procedure will be done. The first subsection consists of a detailed description of the construction of

the lab-scale electrolyser built with the purpose of later validating the model and the second subsection

is dedicated to the experimental setup and procedure.

4.1. Construction of the Lab-Scale Electrolyser

The lab-scale electrolyser (Figure 8) built for the validation of this model is a one cell alkaline atmospheric

electrolyser with conventional electrode-membrane spacing. This model uses sodium hydroxide, NaOH,

as electrolyte at concentrations around 20 wt.%.

The electrolyser is constituted by two electrodes made of pure nickel sheet recovered from the anode

grid of Ni-MH batteries in their end-of-life [100] and a Zirfon® PERL UTP 500 membrane as gas

separator. The electrodes were glued to the separator with a small glue drop on three of the electrode

corners in order to try to obtain a zero-gap configuration which, in practice, was not verified and this

construction resulted in a conventional configuration electrolyser with a membrane-electrode gap of

around 1mm. The membrane electrode assembly was placed on a transparent acrylic container.

a) Anodic Side

b) Cathodic Side

56

c) Exploded View

Figure 8 - Lab-scale Electrolyser 3D Model

For the electrical connection between the electrodes and the power source, two holes were made on

the upper face of the acrylic container in order to introduce two stainless steel bolts. This way, one end

of the bolt would be inside the electrolyser to be connected to the electrode and the other end would

be outside to be connected to the power source. The connection to the power source was made by

winding an electric wire around the bolt and then squeezing the wire between two washers by tightening

two nuts that are in contact with the washers (Figure 9).To connect the electrodes to the power source,

a hole was made in one section of the electrodes in order to make the tip of the bolt pass through it. The

section of the electrode would then be squeezed between two nuts to increase the surface contact area

and assure good electrical conductivity (Figure 9).

Figure 9 - 3D Model Detail of the Electrical Connection Mechanisms

Five additional holes were made on the acrylic container: two for gas extraction, two for electrolyte

recirculation and one for water feeding.

The two holes for gas extraction – one on the anodic side for oxygen extraction and one on the cathodic

side for hydrogen extraction – were made on the top face next to the ones for electrical connectors.

57

The two holes for the electrolyte recirculation (one on the anodic side and one on the cathodic side)

were made on one of the lateral faces of the acrylic container as close as possible to the bottom of the

electrolyser to avoid the recirculation of product gas along with the electrolyte. The recirculation channel

(Figure 10, Figure 11) is crucial to assure electrolyte concentration uniformity and pressure equilibrium.

The hole for the introduction of the water feeding channel was made on the opposite face of the holes

for the electrolyte recirculation. The water feeding channel (Figure 10, Figure 11) which works based on

the communicating vessels theory, allows a continuous feed of water to make up for the water losses

due to the electrolysis process and it is crucial to assure that the electrolyte concentration remains

constant over time.

Figure 10 - Connecters, Inlets and Outlets (Lab-scale Electrolyser Top View)

Figure 11 - Connections, Inlets and Outlets

4.1.1. Experimental Setup and Procedure

The main goal of this experiment was to collect several cell voltage data points at different current

densities and at different operating temperatures in order to compare them with the polarization curves

produced by the model.

The experimental setup (Figure 12, Figure 13) consisted of:

The Lab-scale Electrolyser;

58

One multimeter – Newport TrueRMS Supermeter;

One k-type thermocouple;

One DC power supply – TDK Lambda Power Supply (0 – 150 V; 0 – 5 A);

One Incubator – Memmert U40 2000 W (0 – 220 ºC) (Figure 14).

The lab-scale electrolyser was connected to the DC power supply through the stainless-steel bolts by

winding the electric wire around the bolt and then squeezing the wire between two washers by tightening

two nuts that are in contact with the washers. The alligator clips of the multimeter were then connected

to the stainless-steel bolts to measure the potential difference between the anode and the cathode. A k-

type thermocouple was immersed in the electrolyte through the oxygen outlet hole and connected to

the multimeter to measure the working temperature of the electrolyser (Figure 12, Figure 13).

Since this is a small cell and thus the power supplied to it is relatively low, the electrolyser was placed

inside an incubator (Figure 14) to increase the electrolyte temperature more efficiently in order to collect

data at different working temperatures (30 ºC, 40 ºC, 50 ºC and 60 ºC).

For safety reasons, since the incubator is a confined space, a gas extraction tube was inserted on the

hydrogen outlet hole in order to remove the produced hydrogen to the exterior of the incubator. The

end of the tube was connected to a hotte to assure a safe extraction of the gas.

Figure 12 - Schematic Representation of the Experimental Setup

Figure 13 - Experimental Setup

Figure 14 - Memmert U40 Incubator

59

4.1.1.1. Experimental Procedure

Step 1: The electrolyser is connected to the power supply and the multímeter and the thermocouple is

introduced in the anodic side.

Step 2: The electrolyte is introduced in the electrolyser thought the water feeding channel resulting in

an active area of 18 cm2. The water feeding channel is then connected to a reservoir containing destilled

water.

Step 3: The incubator is turned on and set to the lowest of the desired temperatures.

Step 4: The electrolyser is connected to the power supply and the multimeter and placed inside the

incubator.

Step 5: The power supply is turned on and the current is slowly increased until a potential difference of

around 2.2 V is reached (read on the multimeter).

Step 6: The current is continuously adjusted in order to maintain the voltage at 2.2 V as the increase in

temperature reduces the potential difference for the same current.

Step 7: When the desired working temperature is reached (read on the multimeter), the current is

decreased progessively in order to obtain the voltage values for each current set point until the current

reaches approximately zero. These (i,V) data points are the ones needed to validate the model (Appendix

A3).

Step 8: The incubator is set to the following lowest desired temperature and the current is again

increased in order to obtain a voltage of arround 2.2 V. Repeat all steps from Step 6 on untill all the

necessary data is collected.

The uncertainty associated with the experimental measurements is the uncertainty of the multimeter in

voltage measurements and the uncertainty of the power supply in current measurements. The

multimeter provides voltage measurements with 4 decimal places so the uncertainty is around 0.0001

V. The power supply displays the current values with three decimal places, so the uncertainty is around

0.001A. Any substantial power losses were evaluated from the wiring. Uncertainty associated to the type

K thermocouple is at most of ±1 ºC.

60

5 Results and Discussion

In this chapter, the discussion of the results will take place.

Model validation is an essential step in the model conception process to guarantee that the simulation

results provided by the model are representative of the reality of the events. The validation of the model

will be done by comparison with experimental data from the literature and by comparison with

experimental data obtained from the lab-scale electrolyser built for this purpose, as described in the

previous section.

5.1. Model Validation with Data from the Literature

In this section, the previously presented model will be validated by comparison with sets of experimental

data available in the literature [5], [6], [23] from two different alkaline electrolyser models: the HRI

electrolyser and the PHOEBUS electrolyser.

5.1.1. HRI Electrolyser

In the early 2000s, the Hydrogen Research Institute (HRI) developed a stand-alone renewable energy

system based that uses an alkaline electrolyser to produce hydrogen when the produced energy

surpasses the demand. The electrolyser in question is a 5-kW alkaline atmospheric electrolyser

manufactured by Stuart Energy Systems, Inc. built with conventional electrode-membrane spacing. The

electrolyser is constituted by two electrodes made of 99.99% pure nickel separated by a Zirfon®

membrane and immersed in a 30 wt.% KOH electrolyte. The HRI electrolyser summarized information

can be found in Table 8.

The distance between the electrode channels and the electrocatalyst layer (𝑙 , 𝑙 ) was assumed

to be half the electrode thickness (0.1 cm) since in AEC electrodes the electrocatalysts are usually

applied to the electrode surface.

The bubble zone thickness at each electrode (𝛽 , 𝛽 ) was assumed to be the reported maximum

thickness of a bubble layer at AEC electrodes (0.06 cm) [5], [91] since this electrolyser works at ambient

pressure and the experimental data obtained from Ref. [23] refers to relatively low temperatures and

none of this conditions favour bubble departure leading to bigger bubbles and thicker bubble layers.

The limiting current density at which the electrodes are considered to be 100% covered by bubbles was

considered to be around 3 A/cm2 at 25 ºC (~10% of the theoretical maximum value [90]) as suggested

by Abdin et al. [5] in their simulation for this model.

61

Table 8 - HRI Electrolyser information [6], [23], [75], [83]

Operating information Value Unit

Operating Voltage Range 48 - 56 V

Operating Pressure 1 bar

Operating Temperature Range 0 – 80 ºC

Hydrogen Production Rate (at 80 ºC) 1 Nm3 h-1

Electrical Power Reference (at 60 ºC) 5 kW

Electrolyte Concentration — KOH 30 wt. %

Electrolyser’s Parameter

𝐴 Electrode area 300 cm2

𝑙 , 𝑙 Distance between electrode and separator 0.125 cm

𝛿 , 𝛿 Electrode thickness 0.2 cm

𝜖 , 𝜖 Electrode porosity 0.3 —

𝜏 , 𝜏 Electrode tortuosity 3.65 —

𝐴 Separator area 300 cm2

𝛿 Separator thickness 0.05 cm

𝑁 Number of electrolytic cells in the electrolyser 24 —

To model the activation overpotential using the Butler-Volmer equation (Eq.(31)), values for the charge

transfer coefficients, 𝛼, the exchange current densities, 𝑖 , are required. These values are dependent on

the catalyst material, operating temperature and the electrolyte type and composition [6] and are usually

obtained experimentally. To model the anodic and the cathodic charge transfer coefficients, empirical

correlations (Eq.(84) and (85)) were obtained from experimental data available in the literature [101]–

[103].

𝛼 = 0.0043 𝑇 − 0.5697 (84)

𝛼 = 0.0025 𝑇 + 1.9650 × 10 (85)

The exchange current densities for the anode and the cathode were also modelled from experimental

data. The parameters required to solve equation (32) were fitted from Arrhenius plots (Table 9) obtained

from experimental data available in the literature [38], [101]–[103].

Table 9 - Fitted parameters for the HRI electrolyser's anode and cathode.

Electrode 𝚫𝑮𝒄𝒌 (kJ/mol) 𝒊𝟎.𝒓𝒆𝒇

𝒌 (A/cm2)

Anode 41.50 1.3435 × 10-5 at 𝑇𝑟𝑒𝑓 = 296.15 K

Cathode 23.45 1.8456 × 10-3 at 𝑇𝑟𝑒𝑓 = 296.15 K

The reference conditions used to compute the concentration overpotential, 𝜂 , were ambient

temperature and pressure (20 ºC, 1 atm) and a working current density of 0.001 A/cm2 at which the gas

production is residual therefore not inducing concentration overpotentials and hence being a good

reference working point.

62

The calculation of the membrane ohmic resistance through the model suggested by equation (55)

requires the separator porosity, tortuosity and wettability. However, as already discussed, in most cases,

the results obtained from equation (55) are an overestimation and for that reason, the use of

experimental correlations is highly advised for better accuracy [6], [64]. Experimental values for the

porosity and tortuosity of different types of Zirfon® separators are reported by Stojadinovic et al. [75].

For the separator wettability, 𝜔 , Abdin et al. [5] suggest a fitted value of 0.85.

To compute the ohmic resistance of the Zirfon® separator, in this work, a correlation obtained from

experimental data provided by the Zirfon® manufacturer, Agfa [77], was used (Figure 15, Eq.(86)) since

the separator resistance values obtained from equation (55) could be up to twice the resistance values

provided by the manufacturer for a given temperature. This correlation can be applied at ambient

pressure for AEC typical temperatures presenting a maximum error of 2.9% when compared to the

experimental data.

Figure 15 - Zirfon® Separator areal ionic resistance in 30 wt.% KOH: experimental data by Agfa [77]

and fitting curve.

The experimental data points [23] at each temperature and the corresponding polarization curve

produced by the model can be found in Figure 16. The maximum and average deviations for each

temperature can be found in Table 10. The average maximum deviation is around 1.02% and the total

average deviation is 0.47% which proves that the model predicts the electrolyser polarization curves

accurately. As expected, the operating voltage decreases with temperature since it increases the

electrolyte conductivity, decreases membrane resistance (Figure 15) and increases electrode

electrochemical efficiency.

𝑅,

=54.01834 exp(−0.01727 𝑇)

𝐴 (86)

63

Table 10 - Maximum and average deviations between HRI electrolyser experimental data and the model at each temperature.

Temperature 𝚫𝑽𝒎𝒂𝒙 (%) 𝚫𝑽 (%)

23 ºC 0.7935 0.3092

35 ºC 0.7483 0.3816

40 ºC 1.3404 0.6387

45 ºC 1.2124 0.6002

53.5 ºC 0.9899 0.4066

Average 1.0169 0.4673

Figure 16 – HRI electrolyser polarization curves at different temperatures with corresponding experimental data [23].

5.1.2. PHOEBUS Electrolyser

The PHOEBUS electrolyser is a zero-gap 26 kW alkaline electrolyser that operates at 7 bar. This

electrolyser model has been widely used in the literature for the validation of several semi-empirical

models [104], [105]. Summarized information about the PHOEBUS electrolyser can be found in Table

11. The membrane is a nickel oxide membrane (NiO) supported by a nickel mesh. The anode is a nickel

plate with a Ni/Co3O4/Fe electrocatalyst layer applied to its surface and the cathode is a nickel plate with

a C-Pt electrocatalyst layer applied to its surface. This electrolyser uses KOH as electrolyte with

concentrations between 30 wt.% and 40 wt.% [23].

The distance between the electrode channels and the electrocatalyst layer (𝑙 , 𝑙 ) was assumed

to be half the electrode thickness (0.1 cm) since in AEC electrodes the electrocatalysts are applied to

the electrode surface.

The bubble zone thickness at each electrode (𝛽 , 𝛽 ) was assumed to be the reported minimum

thickness of a bubble layer at AEC electrodes (0.045 cm) [5], [91] since this electrolyser works at a

pressure above ambient pressure which favours bubble departure leading to smaller bubbles and

thinner bubble layers. The limiting current density at which the electrodes are considered to be 100%

covered by bubbles (𝑖 ) was considered to be the theoretical maximum value [90], around 30 A/cm2.

64

Table 11 - PHOEBUS electrolyser information [5], [6], [23], [104]

Operating information Value Unit

Operating Voltage Range 30 - 40 V

Operating Pressure 7 bar

Operating Temperature Range 0 – 80 ºC

Electrical Power Range 5 - 26 kW

Electrolyte Concentration — KOH 30 - 40 wt. %

Electrolyser’s Parameter

𝐴 Electrode area 2500 cm2

𝑙 , 𝑙 Distance between electrode and separator 0 cm

𝛿 , 𝛿 Electrode thickness 0.2 cm

𝜖 , 𝜖 Electrode porosity 0.41 —

𝜏 , 𝜏 Electrode tortuosity 4.25 —

𝐴 Separator area 2500 cm2

𝑁 Number of electrolytic cells in the electrolyser 21 —

The distance between the electrode channels and the electrocatalyst layer (𝑙 , 𝑙 ) was assumed

to be half the electrode thickness (0.1 cm) since in AEC electrodes the electrocatalysts are applied to

the electrode surface.

The bubble zone thickness at each electrode (𝛽 , 𝛽 ) was assumed to be the reported minimum

thickness of a bubble layer at AEC electrodes (0.045 cm) [5], [91] since this electrolyser works at a

pressure above ambient pressure which favours bubble departure leading to smaller bubbles and

thinner bubble layers. The limiting current density at which the electrodes are considered to be 100%

covered by bubbles (𝑖 ) was considered to be the theoretical maximum value [90], around 30 A/cm2.

The temperature-dependent separator resistance correlation (Eq.(87)) was adapted from the semi-

empirical correlation provided by Ulleberg [104] for the total cell resistance.

𝑅 = 2,2500 × 10 𝑇 − 1,3392

𝐴 (87)

As the PHOEBUS electrolyser is a zero-gap model, the electrolyte resistance, 𝑅 , is

approximately zero.

The reference conditions used in the calculation of the concentration overpotential, 𝜂 , are the same

as those used in the simulation of the HRI electrolyser at which the gas production is residual therefore

not inducing concentration overpotentials and hence being a good reference working point.

To model the anodic and the cathodic charge transfer coefficients, 𝛼 and 𝛼 , empirical correlations

(Eq.(88)and (89)) were adapted from experimental data available in Refs. [57] and [106].

𝛼 = 0.0020 𝑇 + 0.1612 (88)

𝛼 = 0.0013 𝑇 + 0.1348 (89)

65

The exchange current densities for the anode and the cathode, 𝑖 and 𝑖 ,were also modelled from

experimental data. The reference exchange current density (𝑖 , ) and the activation energy (Δ𝐺 ) for

the cathode are provided by Durst et al. [106]. The anodic reference exchange current density (𝑖 , )

and the corresponding activation energy (Δ𝐺 ) were fitted from Arrhenius plots (Table 12) obtained

from experimental data available in the literature [57], [104].

Table 12 – Fitted Parameters for the PHOEBUS electrolyser’s anode and cathode.

Electrode 𝚫𝑮𝒄𝒌 (kJ/mol) 𝒊𝟎.𝒓𝒆𝒇

𝒌 (A/cm2)

Anode 94.36 1.2622 × 10-7 at 𝑇𝑟𝑒𝑓 = 313.15 K

Cathode 18.00 1.6000 × 10-1 at 𝑇𝑟𝑒𝑓 = 313.15 K

The experimental data points from at each temperature [104] and the corresponding polarization curve

produced by the model can be found in Figure 17. The maximum and average deviations for each

temperature can be found in Table 13. The average maximum deviation is around 0.42% and the total

average deviation is 0.19% proving that the model predicts the electrolyser polarization curves

accurately.

Table 13 – Maximum and average deviations between the PHOEBUS electrolyser experimental data and the model at each temperature.

Temperature 𝚫𝑽𝒎𝒂𝒙 (%) 𝚫𝑽 (%)

30 ºC 0.1348 0.0950

40 ºC 0. 4957 0.2225

50 ºC 0.7995 0.2316

60 ºC 0.5045 0.1831

70 ºC 0.2396 0.1349

80 ºC 0.3616 0.2441

Average 0.4226 0.1852

Figure 17 – PHOEBUS electrolyser polarization curves at different temperatures with corresponding experimental data [104].

66

5.2. Model Validation with Data from a Lab-scale Electrolyser

In this subsection, the developed model will be validated with sets of experimental data obtained from a

lab-scale model built for this purpose. The experimental data was obtained according to the experimental

setup and procedure described in Chapter 4.

The lab-scale electrolyser built for the validation of this model is a one cell alkaline atmospheric

electrolyser with conventional electrode-membrane spacing. This model uses sodium hydroxide, NaOH,

as electrolyte at concentrations around 20 wt.%.

5.2.1. Physical Parameters and Model Validation

Summarized information about the lab-scale electrolyser can be found in Table 14.

Table 14 - Lab-scale Electrolyser summarized information.

Operating information Value Unit

Operating Pressure 1 bar

Operating Temperature Range 0 – 80 ºC

Electrical Power Range 0.4 - 14 W

Electrolyte Concentration — NaOH 20 wt. %

Electrolyser’s Parameter

𝐴 Electrode area 18 cm2

𝑙 , 𝑙 Distance between electrode and separator 0.1 cm

𝛿 , 𝛿 Electrode thickness 0.1 cm

𝜖 , 𝜖 Electrode porosity 0.2 —

𝜏 , 𝜏 Electrode tortuosity 4.25 —

𝐴 Separator area 18 cm2

𝑁 Number of electrolytic cells in the electrolyser 1 —

The initial intention was to build a zero-gap electrolyser in order to obtain a lower cell resistance and

higher efficiency however, after analysing the experimental data it was concluded that the actual cell

resistance was not compatible with a zero-gap cell since the experimental data points would always be

considerably above the voltage values computed by the model (Figure 18). Despite not being able to

assure a zero-gap configuration, the gap between the electrodes and the membrane was still very

reduced thus, through an iterative process, the average electrode-membrane gap (𝑙 , 𝑙 ) was

found to be around 1 mm.

The electrolyser is constituted by two electrodes made of pure nickel sheet recovered from the anode

grid of Ni-MH batteries in their end-of-life. The electrodes are made of four layers of perforated 100%

pure nickel sheet [100] which resulted in an approximate thickness of 1mm ( 𝛿 , 𝛿 ) and an

approximate porosity of 20% (𝜖 = 𝜖 = 0.2). The tortuosity (𝜏 , 𝜏 ) was assumed to be around 4.25

similarly to the one assumed in the HRI electrolyser [5] since it is difficult to estimate with reasonable

67

accuracy and it does not significantly affect the results. The distance between the electrode channels

and the electrocatalyst layer (𝑙 , 𝑙 ) was assumed to be half the electrode thickness (0.05 cm)

since in AEC electrodes the electrocatalysts are usually applied to the electrode surface.

Figure 18 - Difference between assuming a zero-gap configuration and assuming a 1 mm gap

The electrolyser is constituted by two electrodes made of pure nickel sheet recovered from the anode

grid of Ni-MH batteries in their end-of-life. The electrodes are made of four layers of perforated 100%

pure nickel sheet [100] which resulted in an approximate thickness of 1mm ( 𝛿 , 𝛿 ) and an

approximate porosity of 20% (𝜖 = 𝜖 = 0.2). The tortuosity (𝜏 , 𝜏 ) was assumed to be around 4.25

similarly to the one assumed in the HRI electrolyser [5] since it is difficult to estimate with reasonable

accuracy and it does not significantly affect the results. The distance between the electrode channels

and the electrocatalyst layer (𝑙 , 𝑙 ) was assumed to be half the electrode thickness (0.05 cm)

since in AEC electrodes the electrocatalysts are usually applied to the electrode surface.

Due to the lack of experimental data on the membrane resistance in NaOH electrolyte, the Zirfon®

membrane resistance in NaOH could not be modelled directly from experimental data. Since the

membrane resistance is proportional to the electrolyte ionic resistivity according to equation (55),

equation (86) multiplied by 𝜌 /𝜌 at each temperature was used to model the membrane

resistance in NaOH (𝑅 , , eq.(90)). As the NaOH resistivity is higher than the KOH resistivity

(Figure 6), the Zirfon® membrane resistance will also be higher when immersed in sodium hydroxide

(Figure 19). The electrolyte ionic resistivity at each temperature for KOH and NaOH can be obtained

from equations (77) and (78), respectively.

𝑅,

=𝜌

𝜌 𝑅

, (90)

Figure 19 shows the temperature dependence of the membrane ionic resistance for both KOH and

NaOH. The KOH curve is the one obtained from equation (55) represented in Figure 15 which presents

a low deviation from experimental data provided by Agfa [77]. The NaOH curve was obtained by applying

a correction factor to equation (55) which yielded equation (90).

68

Figure 19 - Zirfon® Ionic Resistance in 20 wt.% NaOH vs Zirfon® Ionic Resistance in 30 wt.% KOH

The bubble zone thickness at each electrode (𝛽 , 𝛽 ) was assumed to be the reported maximum

thickness of a bubble layer at AEC electrodes (0.06 cm) [5], [91] since this electrolyser works at ambient

pressure and the experimental data was obtained at relatively low temperatures and none of this

conditions favour bubble departure leading to bigger bubbles and thicker bubble layers.

The limiting current density at which the electrodes are considered to be 100% covered by bubbles was

considered to be around 2 A/cm2 at 25 ºC (~6.7% of the theoretical maximum value [90]). Despite being

an estimation, this value proves to be a good estimation since it provides good results when compared

to the experimental data (Figure 20).

Figure 20 - Comparison between several limiting current densities at a working temperature of 60 ºC

The reference conditions used to compute the concentration overpotential, 𝜂 , were again ambient

temperature and pressure (20 ºC, 1 atm) and a working current density of 0.001 A/cm2 at which the gas

production is residual therefore not inducing concentration overpotentials and hence being a good

reference working point.

69

To model the anodic and the cathodic charge transfer coefficients, 𝛼 and 𝛼 , empirical correlations

(Eq. (91) and (92)) were adapted from experimental data available in Refs. [101]–[103].

𝛼 = 1.1530 × 10 𝑇 + 0.344890 (91)

𝛼 = 1.72100 × 10 𝑇 + 0.06112 (92)

The charge transfer coefficient is the only parameter that influences the Tafel Slope for a given

temperature so it must be well adapted to the experimental data so that the model predicts the

electrolyser operation accurately. From Figure 21 it is noticeable that the polarization curve is more

sensitive to variations of the anodic charge transfer coefficient, 𝛼 , since the cell overpotential is

strongly dominated by the anodic activation overpotential (Figure 22) due to the slower kinetics of the

oxygen reaction.

Figure 21 - Influence of the charge transfer coefficient on the Polarization Curve

Figure 22 - Anodic and Cathodic Contribution to the Activation Overpotential at 60 ºC

The exchange current densities for the anode and the cathode, 𝑖 and 𝑖 ,were also modelled from

experimental data. The reference exchange current density (𝑖 , ) and the activation energy (Δ𝐺 ) for

the cathode were fitted from an Arrhenius plot obtained from experimental data by Kibria et al [102]

(Table 15). The anodic exchange current density (𝑖 , ) and the corresponding activation energy (Δ𝐺 )

(Table 15) were fitted from to experimental data assuming the cathodic reference exchange current

70

density and activation energy. These values are in accordance with the average values obtained from

ref. [103].

Table 15 - Fitted Parameters for the Lab-scale electrolyser anode and cathode.

Electrode 𝚫𝑮𝒄𝒌 (kJ/mol) 𝒊𝟎.𝒓𝒆𝒇

𝒌 (A/cm2)

Anode 64.53 4.7250 × 10-8 at 𝑇𝑟𝑒𝑓 = 303.15 K

Cathode 26.58 2.2459 × 10-3 at 𝑇𝑟𝑒𝑓 = 303.15 K

The maximum and average deviations for each temperature can be found in Table 16. The average

maximum deviation is around 0.67% and the total average deviation is 0.32% proving that the model

predicts the electrolyser polarization curves accurately.

Table 16 - Maximum and average deviations between the lab-scale electrolyser experimental data and the model at each temperature.

Temperature 𝚫𝑽𝒎𝒂𝒙 (%) 𝚫𝑽 (%)

30 ºC 0.6335 0.4224

40 ºC 0.7361 0.3369

50 ºC 0.7940 0.3177

60 ºC 0.4996 0.2172

Average 0.6658 0.3236

The experimental data points from at each temperature and the corresponding polarization curve

produced by the model can be found in Figure 23. For the same temperature step of 10 ºC, it is

noticeable that the polarization curve grows at increasingly smaller rates. This is due to the better

electrode electrocatalytic performance at higher temperatures, the increase in electrolyte conductivity

which decreases the separator resistance and electrolyte resistance reducing the ohmic overpotential,

the fact that higher temperatures favour bubble departure leading to lower activation overpotentials.

Figure 23 – Lab-scale electrolyser polarization curves at different temperatures with corresponding experimental data.

71

Figure 24 explores the contributions of each source of overpotential to the total cell voltage of the lab-

scale model at 60 ºC.

Figure 24 - Overpotential Contribution to the total Cell Voltage

The concentration overpotential is negligible in this case (≈ 7 × 10 𝑉). In fact, this is the case for most

atmospheric pressure alkaline electrolysers such that most authors choose not to include the

concentration overpotentials in their models [6], [23]. For this reason, its contribution is not visible in

Figure 24.

The ohmic contribution is relatively small reaching a maximum of 0.2287 Ω at around 0.27 A/cm2 (Figure

24 and Figure 25). The biggest contribution to the ohmic overpotential is the resistance caused by the

membrane-electrode gap ( 𝑅 = 0.0352 Ω at 0.27 A/cm2) followed by the membrane ionic

resistance (𝑅 = 0.0132 Ω). The electrode resistance is negligible being 1.3469 × 10 Ω in this case

and for this reason is not visible in Figure 25.

Figure 25 - Contributions to the Ohmic Overpotential at 60 ºC

The biggest contributor to the cell overpotential is the activation overpotential due to reaction kinetics

and bubble accumulation on the electrodes (Figure 26). The dominant contribution is the activation

overpotential of the electrodes themselves due to the reaction kinetics specially on the anodic side due

to its relatively small exchange current density.

72

Figure 26 - Contributions to the Anodic and Cathodic Activation Overpotentials at 60 ºC

73

6 Conclusions

6.1. Achievements

The main goal of this thesis was to develop and validate a steady state model of an alkaline electrolysis

cell by modelling the physical processes involved in the electrolysis process. The main goal was

accomplished successfully since the model simulates with good accuracy the polarization curves of

alkaline electrolysers working with different electrolytes and at different temperatures and pressures.

The accuracy of the model, however, is strongly dependent on four main parameters: the membrane-

electrode gap, the limiting current density, the charge transfer coefficients, and the exchange current

densities. Apart from the membrane-electrode gap, all these parameters are difficult to estimate

accurately.

By displaying the total cell voltage as a sum of its several components, this model is useful to investigate

the sensitivity of the polarization curve to certain parameters.

The electrolyser data collected from the literature allowed the validation of the model resulting in very

low deviations between the data obtained from the model and the experimental data. Since the

PHOEBUS electrolyser and the HRI electrolyser are commercial models, their detailed construction

process and materials are not easily found in the literature thus, some of the parameters used to perform

the simulations rely on assumptions and fittings. Nonetheless, taking into account the also very good

accuracy of the simulation regarding the lab-scale electrolyser, it is not expected that the accuracy of

the simulation regarding the HRI and PHOEBUS electrolysers may vary greatly if the manufacturers

would provide the real simulation parameters.

Despite not being very polished and optimized, the lab-scale model built for the purpose of validating

the physics model succeeded to accomplish its main goal. The experimental data collected from the

electrolyser allowed the validation of the model, resulting in maximum deviations under 0.7% and

average deviations under 0.35%. The deviation values obtained from these simulations can be somewhat

considered more reliable and realistic since the characteristics of this electrolyser were widely known,

which allowed the use of simulation parameters that were surely true, without the need of establishing

assumptions.

The measurement of the flow rate of hydrogen produced by the lab-scale electrolyser would have been

of great interest since it would allow the estimation of the several cell efficiencies listed in this work.

However, none of the measurement techniques used during this project was reliable.

74

6.2. Future Work

In the future, since the activation overpotential is the heaviest contributor for the cell overpotential, for a

more accurate estimation of the activation overpotentials of the bubble covered electrodes, the model

of the bubble coverage should be improved since the one used in this work is based on a semi-empirical

correlation and, in some cases, may not provide good estimates for the real percentage of the covered

area under certain working conditions. The development of a physic-based model to compute the bubble

coverage of the electrode surfaces and the bubble void fractions in the bubble zones is a major challenge

that one should try to overcome in the future.

Additionally, in order to reduce the activation overpotential of the bubble free anode – the responsible

for the biggest share of the total activation voltage – research effort should continue to be directed

towards the development of new catalysts and electrode structural parameters that allow the reduction

of the overpotential caused by the OER.

Furthermore, the electrodes of the lab-scale model should be replaced by nickel electrodes with a

different structure in order to increase the cell efficiency. According to theory, replacing these electrodes

by a nickel cloth or mesh would lead to lower cell voltages since it would greatly increase the active

area.

As future work, it would also be interesting to successfully measure the product gas flow rate in order to

calculate the several cell efficiencies listed in this thesis.

75

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81

Appendix A A1. HRI Electrolyser Experimental Data

Data collected from plots available on ref [23]

T =23 ºC T =35 ºC T =40º C T =45 ºC T =53.5 ºC

i V i V i V i V i V

0.0026 1.4220 0.0026 1.3790 0.0035 1.3748 0.0027 1.3460 0.0044 1.3690

0.0049 1.4600 0.00745 1.4610 0.0057 1.41908 0.01017 1.4650 0.0066 1.4110

0.0084 1.5030 0.0115 1.5010 0.0094 1.4539 0.0138 1.5071 0.0089 1.4416

0.0145 1.5451 0.0260 1.5860 0.0128 1.5030 0.0207 1.5410 0.0169 1.4965

0.0194 1.5860 0.0370 1.6210 0.0185 1.5440 0.0295 1.5680 0.0238 1.5130

0.0291 1.6240 0.0520 1.6610 0.0282 1.5730 0.0450 1.6080 0.0339 1.5540

0.0436 1.6670 0.0732 1.700 0.0406 1.6040 0.0639 1.6420 0.0498 1.5940

0.0617 1.7090 0.1010 1.7440 0.0577 1.6490 0.0899 1.6910 0.0701 1.6360

0.0860 1.7540 0.1350 1.8020 0.0811 1.6930 0.1250 1.7350 0.1000 1.6760

0.1350 1.8310 0.1620 1.8300 0.1110 1.7310 0.1680 1.7960 0.1380 1.7250

0.1810 1.8490 0.1510 1.7820 0.1960 1.8270 0.1790 1.7890

0.2370 1.9250 0.1810 1.8270 0.2210 1.8590 0.1860 1.7980

0.2000 1.8490 0.2450 1.8690

82

A2. PHOEBUS Electrolyser Experimental Data

Data collected from plots available on ref [104]

T=30 ºC T=40 ºC T=50 ºC

i (A/cm2) V i (A/cm2) V i (A/cm2) V

0.1240 1.8514 0.0425 1.6692 0.0417 1.6158

0.1266 1.8564 0.0601 1.7087 0.0439 1.6236

0.0765 1.7377 0.0469 1.6315

0.0805 1.7372 0.0496 1.6389

0.0947 1.7668 0.0566 1.6502

0.0972 1.7687 0.0641 1.6705

0.1216 1.8116 0.0684 1.6754

0.1384 1.8322 0.0725 1.6852

0.1582 1.8543 0.0774 1.6921

0.1609 1.8543 0.0932 1.7157

0.0974 1.7216

0.0987 1.7335

0.1011 1.7246

0.1060 1.7409

0.1104 1.7443

0.1224 1.7586

0.1254 1.7546

0.1300 1.7670

0.1317 1.7763

0.1346 1.7713

0.1389 1.7851

0.1484 1.7964

0.1664 1.8072

0.1770 1.8234

0.1883 1.8282

0.2016 1.8564

0.2040 1.8539

83

T=60 ºC T=70 ºC T=80 ºC

i (A/cm2) V i (A/cm2) V i (A/cm2) V

0.0660 1.6268 0.0723 1.5901 0.1153 1.6154

0.0698 1.6342 0.0741 1.5985 0.1305 1.6346

0.0736 1.645 0.0801 1.6054 0.1341 1.6380

0.0790 1.6479 0.0833 1.6103 0.1434 1.6503

0.0826 1.6598 0.0910 1.6221 0.1475 1.6568

0.0859 1.6658 0.1063 1.6457 0.1507 1.6591

0.0960 1.6756 0.1189 1.6629 0.1525 1.6606

0.1010 1.6829 0.1241 1.6678 0.1554 1.6635

0.1088 1.6948 0.1283 1.6762 0.2902 1.7925

0.1174 1.7066 0.1362 1.6841 0.2981 1.7973

0.1282 1.7233 0.1401 1.6885

0.1320 1.7307 0.1438 1.6969

0.1424 1.7395 0.1482 1.7018

0.1509 1.7513 0.1584 1.7101

0.1551 1.7532 0.1659 1.7209

0.1669 1.7630 0.1702 1.7229

0.1719 1.7729 0.2657 1.8205

84

A3. Lab-Scale Electrolyser Experimental Data

T = 30 ºC T = 40 ºC T = 50 ºC T = 60 ºC

i (A/cm2) V i (A/cm2) V i (A/cm2) V i (A/cm2) V

0.1300 2.15 0.1692 2.1990 0.1944 2.1970 0.2222 2.1980

0.1226 2.1318 0.1614 2.1857 0.1886 2.1900 0.2107 2.1824

0.1168 2.1201 0.1499 2.1625 0.1770 2.1661 0.1991 2.1625

0.1072 2.1014 0.1383 2.1353 0.1654 2.1474 0.1895 2.1423

0.0994 2.089 0.1269 2.1102 0.1558 2.124 0.1779 2.1262

0.0898 2.0703 0.1152 2.088 0.1439 2.1076 0.1664 2.1030

0.0783 2.0411 0.1056 2.0685 0.1327 2.0806 0.1548 2.0852

0.0667 2.0123 0.0941 2.0451 0.1211 2.0586 0.1433 2.0645

0.0571 1.9801 0.0825 2.0210 0.1116 2.0394 0.1337 2.0495

0.0494 1.9623 0.0709 1.9890 0.0999 2.0185 0.1221 2.0261

0.0378 1.9326 0.0556 1.9491 0.0884 1.9925 0.1106 2.0005

0.0282 1.8933 0.0440 1.9125 0.0768 1.9700 0.1009 1.9831

0.0224 1.8658 0.0324 1.8728 0.0672 1.9438 0.0894 1.9598

0.0166 1.8344 0.0208 1.8255 0.0557 1.9130 0.0778 1.9349

0.0108 1.7909 0.0112 1.7676 0.0441 1.8805 0.0662 1.9087

0.0051 1.7198 0.0054 1.6975 0.0326 1.8412 0.0547 1.8779

0.0028 1.6643 0.0042 1.6728 0.0210 1.7939 0.0451 1.8484

0.0014 1.6024 0.0028 1.6329 0.0114 1.7343 0.0335 1.8113

0.0014 1.5719 0.0056 1.6649 0.0219 1.7613

0.0042 1.6400 0.0104 1.6785

0.0028 1.5962 0.0056 1.6134

0.0014 1.5300 0.0042 1.5842

0.0028 1.5506

0.0014 1.4976

85

Appendix B

B1. HRI Electrolyser Polarization Curves

a)

b)

c)

d)

e)

86

B2. PHOEBUS Electrolyser Polarization Curves

a)

b)

c)

d)

e)

f)

87

B3. Lab-scale Electrolyser Polarization Curves

a)

b)

c)

d)