Faculty of Bioscience Engineering
Academic year 2015-2016
In-situ electrochemical generation of iron and
alkalinity from sewage for sulfide control in sewer
rising main reactors
Grégory Baekelandt
Promotors
Prof. Dr. Ir. Korneel Rabaey
Prof. Dr. Ir. Ilje Pikaar
Tutor
Dr. Eleni Vaiopoulou
A thesis submitted for the Degree of Master in Bioscience Engineering: Environmental Technologies
University of Ghent, June 2016
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Copyright
“The author and promoters give the permission to use this thesis for consultation and to copy parts
of it for personal use. Every other use is subject to the copyright laws, more specifically the source
must be extensively specified when using results from this thesis.”
“De auteur en promotors geven de toelating deze scriptie voor consultatie beschikbaar te stellen en
delen ervan te kopiëren voor persoonlijk gebruik. Elk ander gebruik valt onder de beperkingen van
het auteursrecht, in het bijzonder met betrekking tot de verplichting uitdrukkelijk de bron te
vermelden bij het aanhalen van resultaten uit deze scriptie.”
Ghent, June 2016
The promotors,
The tutor,
The author,
Prof. Dr. Ir. Korneel Rabaey
Dr. Eleni Vaiopoulou
Grégory Baekelandt
Dr. Ilje Pikaar
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Preface
Starting with the education of Bio-engineering was a very big challenge for me. After four years of
studying I was finally at the point of starting my own master thesis. Since the start of the study I
knew that I wanted to have a foreign experience and I thought that this was the perfect moment for
it. After many doubts I decided to combine my love for science and my admiration for Australia.
During my second year of studying I met the most fascinating and interesting professor: Professor
Korneel Rabaey. I remembered that he had professional contacts in Australia and I decided to grab
my chance and ask him if he had an interesting thesis subject in Australia. He offered me a perfect
topic that had my sincere interest and that could take me six months to Brisbane, Australia. It’s
because of this that I want to thank Professor Korneel Rabaey in the first place and I will never forget
the energy and encouragement that you gave your students and the opportunity to make my thesis
an incredible and educational experience.
While preparing my stay in Brisbane I came in contact with a lot of people who helped me through
heaps of paperwork and I want to thank Vivienne Clayton especially for the assistance during this
period. Finally, I arrived in the big city and a little bit nervous I went for the first time to the
University of Queensland. This is the place where I first met Ilje Pikaar and Hui-Wen Lin (Winni).
From the first moment you helped me with all of the paperwork, organising all of my lab-work and
teaching me a huge amount of lab-experience. Winni, thank you for all your patience to teach me
everything you knew about the topic and especially the reactors, for all the calculations which
sometimes led to discussions, for the little snacks you brought, the stories during the countless
hours of work in the lab, etc. You made me feel very welcome and thought me more than I could
ever imagine. Ilje, a thousand times thank you for your help with the calculations, for the discussion
of the results and possible improvements. And at last, thank you for the very nice farewell lunch, I
will always remember how nicely you guys welcomed me and encouraged me to keep on going. Back
in Belgium, most of the time went to writing my thesis. At this point, Eleni Vaiopoulou gave me a
huge amount of help by giving me great advice to make my thesis a pleasant and complete work.
Also the advice during presentations really helped me, so thank you for all the work and time you
invested in me.
I kept the last and biggest acknowledgement for five people who are really important to me and who
I admire a lot. My sister Alexandra and her boyfriend Sander, thank you for the unbelievable journey
through Australia and keeping me up-to-date about what happened in Belgium. Sister, I know how
hard it was for you to see your brother leave for such a long period, but thank you for letting me
experience the adventure of my life. My girlfriend Marie, thank you for joining us on the trip through
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Australia. Being back in Belgium was like I never left, so thank you for everything and always
encouraging me to go on this wild adventure. And my biggest acknowledgement is for two very
special people that gave me the chance to experience all this. My parents, since I was a little boy you
always encouraged me to give every opportunity my best shot and you kept doing this until now. It is
impossible to express the immense gratitude I have for both of you. Thanks for the support during
this adventure, it was a huge decision that led to the most amazing experience.
This thesis is the result of countless hours of hard work and a huge amount of new experiences and
memories. All the difference was made by having the most amazing tutors and promotors, which
eventually led to this work and I can say that I’m proud.
“Dreams don’t work unless you do.”
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Abstract
Corrosion of the sewer infrastructure due to the formation of hydrogen sulfide is a major issue. The
costs associated with these problems are in the order of $100 million dollar annually in Australia
alone and for Flanders (Belgium) approximately €5 million per year, representing about 10% of total
cost for wastewater collection and treatment. Methane is another important sewer gas generated
by methanogens and forms a greenhouse gas (GHG).
Over the years a few abatement strategies were developed, where chemical dosing strategies with
iron salts, nitrate and nitrite, and oxygen are economic and popular choices. These methods all have
their own practical limitations and therefore there is the need to develop new technologies. In this
thesis there is the further development of an electrochemical technique to address corrosion in the
sewer systems, with the major advantage of in-situ production of iron and alkalinity to get rid of
transport and storage of hazardous chemicals. Previous research addressed several factors which
include demonstration of the removal potential of the electrochemical cell and further optimisation
of the cell design and operation from a practical point of view. This thesis starts with a summary of
relevant literature and background knowledge, followed by obtained results related to an innovative
and promising in-situ technology to address the issues (related to transport, safety, etc.) which are
the result of abatement strategies used nowadays.
The first part focusses on the activity of sulfate reducing bacteria (SRB) and methanogens in the
reactors. The proof of concept is to clearly indicate stability and reliability of the system to
precipitate sulfides. A Fe3+/Fe2+ mixture was produced with a molar dosing ratio of 0.68/0.32 and
was dosed to the reactor during the second part of the research. In addition, factors were also
addressed related to the impact on wastewater parameters like pH, which increases due to
production of alkalinity at the cathode, and phosphate concentrations, which decreases first due to
complex formation with iron ions and increase afterwards due to release from the complex.
The dosing of the produced iron and alkalinity to the reactors was the main topic in the second part
of the research and the main goal was to demonstrate a reliable and efficient removal of sulfides.
Next to this, it was important to show the differences between conventional abatement with iron
salts and the novel electrochemical technique. A higher inhibition of the SRB was achieved (-57% in
comparison to -45% for dosage with iron salts) in combination with a higher inhibition of the
methanogenic activity (-40% in comparison to -10% for dosage with iron salts) at a molar dosing
ratio of 0.5. In addition to the achievement of satisfying removal efficiencies (> 95% in order to
achieve sulfide concentrations lower than 0.5 mg S L-1) there is also the importance of inhibition of
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activity of SRB and methanogens to incorporate the impact of iron dosing on sewer biofilm activities.
This forms a clear advantage over other common abatement strategies that are used nowadays.
Results showed that molar dosing ratios of Fe/S lower than 1.0 form no problem to achieve
satisfying sulfide removal, which is interesting because of the lower operational cost resulting from
the lower cost for energy and iron plates.
The final part refers to the economic analysis to clearly indicate the position of the technology in the
current economical market and possible future perspectives which could lead to a widespread
practical implementation. A total operational cost of US$8.51 or AU$11.91 ML-1, including merely
the cost for iron plates and energy, makes the technology competitive. The practical implementation
will be done in the near future with the construction of a pilot plant at the Gold Coast, QLD, Australia
for the treatment of 1 ML wastewater d-1. The construction of the electrochemical dosing unit on
large scale is an exceptional opportunity because only minor large scale applications exist in the
electrochemical science. Overall, the lab scale experiments showed promising results indicating
potential on larger scale for a long term application.
Keywords
Sewer corrosion, ferrous ions, ferric ions, sulfide, sulfate reducing bacteria, methanogens, sewer,
biofilm, recovery, electrochemical systems, hydrogen sulfide
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Samenvatting
Er werd reeds veel onderzoek verricht naar corrosie van de rioolinfrastructuur vanwege de vorming
van waterstof sulfide (H2S) en verdere oxidatie hiervan tot zwavelzuur (H2SO4). Dit komt grotendeels
door de kosten die geassocieerd zijn met deze problematiek die in Australië kunnen oplopen tot een
grootte orde van $100 miljoen per jaar en voor Vlaanderen (België) ongeveer €5 miljoen per jaar
bedraagt (komt overeen met ongeveer 10% van de totale kost voor de collectie en behandeling van
afvalwater). Naast waterstofsulfide is ook methaan een belangrijk broeikasgas die een belangrijke
bron van onderzoek vormt.
Verschillende technologiën werden reeds ontwikkeld doorheen de jaren. Het doseren van
chemicaliën zoals nitraat, nitriet, zuurstof en ijzer zouten zijn enkele van de meest toegepaste
oplossingen om corrosie te bestrijden. Praktische problemen gerelateerd aan transport, behandeling
en opslag zorgen voor de noodzaak om een nieuwe in-situ technologie te ontwikkelen. In deze thesis
staat de verdere ontwikkeling van een electrochemische techniek voor de productie van ijzer en
hydroxide ionen centraal om dit probleem van corrosie van de rioleringssystemen aan te pakken.
Bij vorig onderzoek ging de aandacht reeds uitvoerig naar de verwijderingsefficiëntie van de
electrochemische cel en de verdere optimalisatie van design en uitvoering. Deze thesis start met de
vermelding van relevante literatuur en achtergrondkennis, gevolgd door de verkregen resultaten in
verband met deze innovatieve en veelbelovende in-situ technologie.
Het eerste gedeelte focust op de activiteit van sulfaat-reducerende bacteriën (SRB) en methanogene
bacteriën in de reactoren. Vervolgens komt de ‘proof of concept’ aan bod om de stabilteit en
betrouwbaarheid van het systeem om sulfiden te precipiteren te duiden. Een combinatie van 2- en
3-waardig ijzer werd hierbij geproduceerd en werd tijdens het tweede deel van het onderzoek
gedoseerd naar de reactoren. Verder werden eveneens parameters relevant voor afvalwater, zoals
pH, dat zal toenemen vanwege de productie van hydroxide ionen aan de cathode, en fosfaat
concentraties, dat eerst zal afnemen vanwege complexatie met ijzer ionen en vervolgens na
dosering naar de reactoren weer zal toenemen vanwege vrijstelling uit het complex, bestudeerd.
In het tweede gedeelte van het onderzoek staat het doseren in de reactoren centraal en was het
hoofddoel om een betrouwbare en efficiënte verwijdering van de sulfiden te bekomen. Daarnaast
stond het aantonen van de verschillen met de conventionele techniek, waarbij ijzer zouten worden
toegevoegd, centraal. Een hogere inhibitie van de SRB werd hierbij bereikt (-57% ten opzichte van 45% voor dosering met ijzer zouten) in combinatie met een hogere inhibitie van de methanogene
activiteit (-40% ten opzichte van -10% voor dosering met ijzer zouten) bij een molaire doseringsratio
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van 0.5. Naast het begrijpen van de inhibitie van sulfaat reducerende bacteriën en methanogene
archaea gerelateerd aan ijzer dosering, stond eveneens het bereiken van een voldoende hoge
verwijderingsefficiëntie centraal (> 95% om sulfide concentraties lager dan 0.5 mg S L-1 te bereiken).
Verder onderzoek toonde aan dat molaire doseringsratios van ijzer ten opzichte van sulfiden lager
dan 1.0 geen probleem vormen en nog steeds resulteerden in voldoende lage sulfide concentraties,
wat belangrijk is vermits de operationele kost met betrekking tot de energie en ijzer platen in dit
opzicht dus lager kan zijn.
Het laatste gedeelte behelst een economische analyse om de positie van de technologie op de
huidige economische markt te beduiden en finaal dus kan leiden tot een praktische implementatie
op grote schaal. Een eerste stap hierbij is de constructie van een ‘pilot plant’ aan de Gold Coast, QLD,
Australië voor de behandeling van 1 ML afvalwater d-1. Het onderzoek in de thesis werd dus gedaan
met het oog op de constructie van de electrochemische installatie op grote schaal, welke een
uitzonderlijke opportuniteit vormt vermits maar een klein aantal applicaties bekend zijn op
dergelijke schaal. De experimenten leiden grotendeels tot veelbelovende resultaten en duiden op
het belang van deze technologie op grote schaal en op lange termijn.
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Table of Contents
1
Introduction .................................................................................................................................... 1
2
Literature review............................................................................................................................. 3
2.1
Sewer systems......................................................................................................................... 3
2.2
Sulfur transformation and methane formation in sewer systems ......................................... 4
2.2.1
Electron acceptor availability under different redox conditions .................................... 4
2.2.2
Sulfur species and major sulfide processes .................................................................... 5
2.2.3
Microbial processes in sewer systems ............................................................................ 7
2.2.4
Methane production in sewer systems ........................................................................ 10
2.2.5
Competition and coexistence between sulfate reducing bacteria (SRB) and
methanogens ................................................................................................................................ 11
2.2.6
Problems related to H2S ................................................................................................ 12
2.2.7
Abatement strategies.................................................................................................... 15
2.3
Electrochemical introduction ................................................................................................ 20
2.3.1
Thermodynamics........................................................................................................... 22
2.3.2
Potential losses ............................................................................................................. 23
2.3.3
Iron electrocoagulation ................................................................................................. 24
2.4
Research objectives .............................................................................................................. 27
2.4.1
Objective 1: Demonstration of the sulfide removal within the rising main reactor using
in-situ electrochemical generation of iron and alkalinity from sewage ....................................... 27
2.4.2
Objective 2: Determination of the impact of iron ions on biofilm response in terms of
sulfide and methane production rate as well as the accumulated concentrations (sulfide and
methane) within the reactor......................................................................................................... 28
2.4.3
Objective 3: Determination of the change in wastewater characterisation such as pH
and phosphate concentrations of the sewage ............................................................................. 29
2.4.4
Objective 4: Determination of the impact of iron ions on biofilm response in terms of
inhibition effects of ferric addition on SRB and methanogenic archaea after the observation of
the decrease in sulfate reduction and methane production activity ........................................... 29
3
Materials and methods ................................................................................................................. 31
3.1
Experimental set-up .............................................................................................................. 31
3.2
Chemical analyses and measurements ................................................................................. 31
3.2.1
Ion Chromatography (IC) for the analysis of sulfur species .......................................... 32
3.2.2
Inductive coupled plasma - optical emission spectrophotometer (ICP-OES) ............... 34
3.2.3
Total suspended solids (TSS)/ Volatile suspended solids (VSS) .................................... 34
3.2.4
Volatile fatty acids (VFA) ............................................................................................... 34
3.2.5
Methane ........................................................................................................................ 34
3.3
Experimental procedure ....................................................................................................... 35
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4
3.3.1
Baseline phase .............................................................................................................. 35
3.3.2
Experimental phase....................................................................................................... 35
Results and discussion .................................................................................................................. 42
4.1
Electrochemical iron and alkalinity production .................................................................... 42
4.2
Baseline sewer reactor.......................................................................................................... 43
4.2.1
Stable and comparable activity ..................................................................................... 43
4.2.2
Batch test ...................................................................................................................... 46
4.3
Experimental phase .............................................................................................................. 46
4.3.1
Dosing ratio 0.5 ............................................................................................................. 46
4.3.2
Dosing ratio 1.0 ............................................................................................................. 54
4.4
Economic analysis ................................................................................................................. 62
5
Conclusion ..................................................................................................................................... 65
6
Future perspectives ...................................................................................................................... 68
7
Bibliography .................................................................................................................................. 70
8
Appendix ....................................................................................................................................... 75
Appendix I: experimental parameters for dosing ratio 0.5 .............................................................. 75
Appendix II: experimental parameters for dosing ratio 1.0 ............................................................. 77
Appendix III: Calibration curve for spectrophotometric iron measurement .................................... 79
Appendix IV: Basic economic calculations for the electrochemical treatment of in-situ iron
production......................................................................................................................................... 80
Appendix V: Methane production rate after dosing ratio 1.0 .......................................................... 82
Compared to the methane production rates during the baseline................................................ 82
Compared to the methane production rates during dosing ratio 0.5 .......................................... 82
Appendix VI: Removal efficiency as a function of dosing ratio......................................................... 83
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List of abbreviations
A/V
Area to Volume ratio
APS
Adenosine-5’-phosphosulfate
ATP
Adenosine triphosphate
AWMC
Advanced Water Management Centre
BNR
Biological Nutrient Removal
CEM
Cation Exchange Membrane
COD
Chemical Oxygen Demand
DO
Dissolved Oxygen
EC
Electrochemical
FISH
Fluorescence in situ hybridization
GHG
Greenhouse Gas
HRT
Hydraulic Retention Time
LAS
Linear Alkylbenzene Sulfonate
MICC
Microbially Induced Concrete Corrosion
MPR
Methane Production Rate
ORP
Oxidation-Reduction Potential
SAOB
Sulfide Anti-Oxidant Buffer
SOB
Sulfur Oxidizing Bacteria
SRB
Sulfate Reducing Bacteria
SRR
Sulfate reduction Rate
TDS
Total Dissolved Sulfide
TSS
Total Suspended Solids
VFA
Volatile Fatty Acids
VSS
Volatile Suspended Solids
WWTP
Wastewater Treatment Plant
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List of tables
Table 1: Oxidation-reduction potential (ORP) and cellular activity (modified from [11]) ...................... 5
Table 2: Oxidation states of sulfur (modified from [27]) ........................................................................ 9
Table 3: Human health effects at various hydrogen sulfide concentrations (modified from [17]) ...... 13
Table 4: Advantages and disadvantages of current methods for odour and corrosion control in
collection systems and pumping stations (modified from [54]) ........................................................... 17
Table 5: Sulfur species occurring in aqueous media. The most common species are printed in bold.
(modified from [74]) ............................................................................................................................. 33
Table 6: Pumping timeframe ................................................................................................................ 36
Table 7: Separate volumes of the different parts of the electrochemical (EC) system ........................ 37
Table 8: Measured iron species concentrations and Fe/S dosing ratio (n=3) ...................................... 38
Table 9: Measured iron species concentrations and Fe/S dosing ratio (n=4) ...................................... 38
Table 10: Separate volumes of different parts of the EC system ......................................................... 39
Table 11: Measured iron species concentrations and Fe/S dosing ratio (n=2) .................................... 39
Table 12: Measured iron species concentrations and Fe/S dosing ratio (n=2) .................................... 40
Table 13: Iron concentrations, coulombic efficiencies and speciation in 3h experiments (n=3) ......... 42
Table 14: Sulfate reduction rate (SRR) (mg S L-1/h) (n=4 for reactor 1; n=3 for reactor 1 and 3)......... 45
Table 15: Sulfide production rate (SPR) (mg S L-1 h-1) (n=3) .................................................................. 46
Table 16: Methane production rate (MPR) of different weeks and the average MPR (mg CH4 L-1 h-1)
(n=4) ...................................................................................................................................................... 46
Table 17: Average cell voltage (V), cathode potential (vs. Ag/AgCl) (V) and anode potential (vs.
Ag/AgCl) (V) during stable operation of the electrochemical system. ................................................. 48
Table 18: Summary of the results sulfur species and chloride concentration (n=3) of the control
system, conventional system, EC system and storage vessel. .............................................................. 51
Table 19: Summary of N and P results of the three systems (n=3) ...................................................... 52
Table 20: Average cell voltage (V), cathode potential (vs. Ag/AgCl) (V) and anode potential (vs.
Ag/AgCl) (V) during stable operation of the electrochemical system. ................................................. 56
Table 21: Summary of the results for sulfur species and chloride concentrations (n=2) of the control
system, FeCl3 system, EC system and storage vessel............................................................................ 59
Table 22: Summary of N and P concentrations of the three systems (n=2) ......................................... 60
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List of figures
Figure 1: Illustration of the difference between gravity sewers and pressure sewers (modified
from [1]) 4
Figure 2: The dissolution of H2S at different pH, generated with dissolution constants pKa1 and pKa2
(modified from [17]) ............................................................................................................................... 5
Figure 3: Transfer of H2S across the air-water interface (from [22]) ...................................................... 6
Figure 4: Major sulfide processes and their effect factors associated with the sulfur cycle in the
sewer system. SOB: sulfur oxidizing bacteria; SRB: sulfate reducing bacteria; ORP: oxidationreduction potential; DO: dissolved oxygen (from [3], [15]) .................................................................... 7
Figure 5: Role of sulfur oxidizing bacteria (SOB) and sulfate reducing bacteria (SRB) (from [22])......... 8
Figure 6: Schematic representation of a sewer biofilm with sulfate partial penetration (adapted from
[38]) ....................................................................................................................................................... 12
Figure 7: Concrete biodeterioration influenced by SRB and SOB; (a) Initial state, (b) Swelling of
concrete and (c) Concrete cracking (modified from [46]) .................................................................... 15
Figure 8: Chemical and biological abatement strategies for H2S emission control in sewer systems
(from [50]) ............................................................................................................................................. 16
Figure 9: Contribution of each chemical to the overall sulfide control (modified from [51]) .............. 17
Figure 10: Mechanism of metal sulfide inhibition (from [59]).............................................................. 20
Figure 11: Simplified scheme of an electrochemical cell with CEM ..................................................... 21
Figure 12: Optimisation of the spacing. Blue path: Sewage is provided with iron ions; Red path: A
rereduction of iron ions can occur at the cathode (from [7])............................................................... 24
Figure 13: Simplified diagram of real application of in-situ electrochemical generation of iron and
alkalinity from sewage for sulfide control. (from research-plan made by Hui-Wen Lin (tutor in
Australia)) .............................................................................................................................................. 27
Figure 14: Simplified schematic overview of the in-situ electrochemical iron and hydroxide
production from sewage (from research-plan made by Hui-Wen Lin (tutor in Australia)) .................. 28
Figure 15: Schematic overview of experimental setup and process (from research-plan made by HuiWen Lin (tutor in Australia)) ................................................................................................................. 31
Figure 16: Picture of the experimental set-up under the bench (under the reactors)......................... 37
Figure 17: Experimental bench set-up (short-term) ............................................................................. 42
Figure 18: Total Dissolved Sulfide (TDS) concentration (mg S L-1) and pH profile for each reactor
during week 8........................................................................................................................................ 44
Figure 19: Total Dissolved Sulfide (TDS) concentration (mg S L-1) and pH profile for each reactor
during week 10. .................................................................................................................................... 45
Figure 20: Course of cell voltage, cathode potential and anode potential before and after cleaning.
Cleaning with 1M HCl was done after 6 days of operation .................................................................. 48
Figure 21: SRR of the control reactor, FeCl3 system and EC system (n=3) ........................................... 49
Figure 22: pH profile of the three systems on day 19........................................................................... 53
Figure 23: Comparison of the methane production rate (MPR) during baseline and experimental
phase (n=3) ........................................................................................................................................... 54
Figure 25: Course of cell voltage, cathode potential and anode potential before and after cleaning.
Cleaning with 1M HCl was done after 6 days of operation .................................................................. 56
Figure 26: Build-up of dirt in-between/on the electrodes after 6 days of operation........................... 56
Figure 27: SRR of the control reactor, FeCl3 system and EC system (n=2) ........................................... 57
Figure 28: Effect of pH on corrosion rate, related to SRB activity (from [80]) ..................................... 58
xviii
Figure 29: Effluent tubes with black deposit layer on the inside due to FeS formation. Left: EC
system; middle: FeCl3 system; right: control reactor............................................................................ 59
Figure 30: pH profile of the three systems on day 4............................................................................. 61
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1 Introduction
Sewer networks are related to the collected flows of wastewater originating from households,
industries and runoff from precipitation in urban areas. After collection, this wastewater stream is
transported for further treatment and disposal [1]. The origin of the sewer network date back to the
days of several ancient civilizations like those located in the Middle East and from the golden age of
the ancient Rome. A major reason for the start of collecting wastewater in underground systems was
the enormous problem of the unpleasant smell from the open sewers and the requirements for
space in the streets of densely populated cities [2]. The sewer network we know today is however a
relatively newly invented infrastructure of cities and makes it possible to live in densely populated
cities [1].
The wastewater that is transported includes substances with a pronounced chemical and biological
reactivity and the sewer network is therefore in addition to being a collection and conveyance
system also a reactor for transformation of wastewater [1]. One of the most important
transformations is the reduction of sulfate to generate hydrogen sulfide (H2S), which eventually
leads to sewer corrosion due to the formation of sulfuric acid (H2SO4) [3]. This acid attacks
cementitious and metallic materials of the concrete leading to rapid corrosion, sewer deterioration,
and eventual collapse. This leads to major economic issues, next to health concerns due to the
toxicity and odour nuisance related to H2S.
The economic effects of sewer corrosion for Flanders (Belgium) are approximated at €5 million per
year, representing about 10% of total cost for wastewater collection and treatment [4]. In Australia,
the costs can go up to $100 million per year [5]. This cost is related to several things: new sewer
materials, traffic costs due to blockage of the roads, etc. Another important process going on in the
sewer system is the production of methane (CH4). This can be related to modern climate change and
the ‘green wave’ which leads to increasing awareness about the environment. Despite the global
warming potential of methane, this methanogenic activity was not taken into account for a long time
[6].
Abatement strategies for sulfide control were developed with methods that involve the continuous
addition of chemicals (24/7) [3]. Next to significant operational costs there is also the issue of
transport, storage and handling the chemicals. Thus, it’s clear that an in-situ strategy becomes
inevitable to address sewer corrosion due to sulfide formation, since with this strategy issues related
to transport, storage and handling of chemicals are evaded. This strategy is based on the
electrochemical production of iron ions to precipitate the formed sulfides, next to the production of
alkalinity to increase the pH [7].
1
Preliminary studies done by the Advanced Water Management Centre (AWMC) during short term
lab experiments reported high removal efficiencies of electrochemical methods for sulfide [5, 7]. The
further optimisation of the electrochemical technique leads to the overcome of some drawbacks
such as electrode scaling and passivation.
In this thesis, a novel single chamber membrane-free reactor configuration is tested for the in-situ
production of iron and hydroxide. Subsequently, dosage of the iron and alkalinity is done to a
mimicking wet-well and reactor to indicate efficient sulfide removal in combination with a thorough
comparison with conventional dosing using iron salts. This novel approach is simple in operation, it
completely eliminates the need for chemical storage and is expected to be almost complete
maintenance-free, in comparison to other electrochemical approaches that require sensor
calibration. The attractiveness for practical applications is mostly determined by the ease of
operation, low degree of complexity, limited maintenance requirements and lower expected cost of
in-situ iron and caustic generation from sewage.
2
2 Literature review
2.1 Sewer systems
The three major ways of classification of sewers refer to the type of sewage collected, the type of
transport mode and the size and function of the sewer. A classification based on the type of sewage
collected leads us to the following three main types of sewer networks: sanitary sewers, storm
sewers and combined sewers [1]. Wastewater of domestic, commercial and industrial origin is
conveyed in both sanitary sewers and combined sewers, whereas the storm sewers in principle only
transport runoff water from urban surfaces and roads [8]. Sewer systems in Belgium and the rest of
Europe are typically combined sewer systems which have the drawback of potential overflows
during periods of heavy rainfall. Regulations in Australia prohibit the connection between
stormwater and sewage systems, which lead to a separation between sanitary sewers and storm
sewers [9].
According to a division that refers to the transport mode, one can distinguish gravity sewers and
pressure sewers. A gravity sewer is designed with a sloping bottom and the flow occurs by
gravitation (Figure 1). In contrast, the driving force for flow in a pressure sewer is pumping, for which
pumping stations are used, also called lift stations. Sewage is fed into and stored in an underground
pit, commonly known as a wet-well. The well is equipped with electrical instrumentation to detect
the level of sewage present. When the sewage level rises to a predetermined point, a pump will be
started to lift the sewage upward through a pressurized pipe system from where the sewage is
discharged into a gravity manhole (Figure 1). An important difference between these two types is
related to the gas phase. The water surface in a gravity sewer is most of the time exposed to a gas
phase (sewer atmosphere), making this by the presence of oxygen (dissolved oxygen concentrations
are typically around 0-5 mg L-1) a crucial factor for the microbial processes. In contrast, in
wastewater of pressure sewers anaerobic conditions are typically present as the pressure sewers are
completely filled with water [10-12]. More than 98% of the sulfide production occurs in biofilms,
which can be related to the surface area to volume ratios (A/V values) of sewers, since smaller
diameter pipes generally result in more sulfide production [13]. In addition, the sewage hydraulic
retention time is also important because stationary or slowly moving streams might result in
anaerobic conditions near the pipe wall, even though oxygen might still be present in the bulk liquid
phase [14].
3
Figure 1: Illustration of the difference between gravity sewers and pressure sewers (modified from [1])
2.2 Sulfur transformation and methane formation in sewer systems
2.2.1 Electron acceptor availability under different redox conditions
Anaerobes survive and degrade substrate most efficiently when the oxidation-reduction potential
(ORP) of their environment is between -200 and -400 millivolts (mV). Anaerobic activity (including
hydrolysis, acetogenesis and methanogenesis) is discouraged by the presence of dissolved oxygen
since dissolved oxygen raises the ORP. The ORP is a measurement of the relative amounts of
oxidized materials, such as nitrate ions (NO3-) and sulfate ions (SO42-), and reduced materials, such as
ammonium ions (NH4+) (Table 1). Free molecular oxygen is available at ORP values greater than +50
mV in the wastewater or sludge and may be used for the degradation of organic compounds by
aerobes and facultative anaerobes. This degradation occurs under aerobic conditions. Free
molecular oxygen is not available at ORP values between +50 and -50 mV, but nitrate or nitrite ions
(NO2-) are present and used for the degradation of organic compounds, which leads to anoxic
conditions. ORP values lower than -50 mV lead to absence of nitrate ions, nitrite ions and oxygen
and thus degradation occurs under anaerobic conditions. Although, sulfate ions are available for the
degradation of organic compounds. This sulfate is reduced and hydrogen sulfide is formed along
with a variety of acids and alcohols due to degradation of organic matter.
At ORP values less than -100 mV, the degradation of organic compounds proceeds as one portion of
the compound is reduced while another portion of the compound is oxidized. This form of anaerobic
degradation of organic compounds is commonly known as mixed-acid fermentation because a
mixture of acids are produced, such as acetate, butyrate, formate and propionate. Anaerobic
degradation of organic compounds and methane production occurs at ORP values lower than -300
mV. During methane production, simple organic compounds such as acetate are converted to
methane, and carbon dioxide (CO2) and hydrogen (H2) are combined to form methane, as can be
seen in the following chemical reactions [11]:
CH3COOH → CH4 + CO2
(1)
CO2 + 4H2 → CH4 + 2H20
(2)
4
Table 1: Oxidation-reduction potential (ORP) and cellular activity (modified from [11])
APPROXIMATE ORP (10-³ V)
>+50
+50 TO -50
<-50
<-100
<-300
CARRIER MOLECULE FOR
DEGRADATION OF ORGANIC
COMPOUNDS
O2
NO3 or NO2SO42Organic compounds
CO2
CONDITION
Aerobic/oxic
Anoxic
Anaerobic
Anaerobic
Anaerobic
2.2.2 Sulfur species and major sulfide processes
Hydrogen sulfide in sewage is primarily present as two sulfide species, depending on the pH [15]. At
the typical pH values of municipal sewage (6.5 < pH < 8.5) these two species are H2S (aq) and HS(Reaction (3)). At pH 7.0, approximately 50% of the hydrogen sulfide remains dissociated. However,
at pH 8.5, less than 4% is present as hydrogen sulfide. In other words: the higher the pH, the lower
the H2S to HS- ratio [16]. The S2- species is not taken into account in sewers because of its
insignificant presence at pH levels higher than 12.0 (Reaction (4)) (Figure 2).
H2S (aq) ↔ H+ + HS-
pKa1=7.04
(3)
HS- ↔ H+ + S2-,
pKa2=11.96
(4)
H2S (aq) ↔ H2S (g)
kc = 468 atm (mole fraction)-1 (5)
Figure 2: The dissolution of H2S at different pH, generated with dissolution constants pKa1 and pKa2 (modified from [17])
5
Only H2S (aq) can be transferred across the air-water interface, giving rise to the emission of
hydrogen sulfide from wastewater to the sewer atmosphere (Reaction (5)) (Figure 3). The process is
dependent on temperature, pH, hydraulic conditions of the water phase and ventilation of the air
phase. Field investigations in the USA performed by Pomeroy et al. (1946) [18], in Australia by
Thistlewayte et al. (1972) [19] and in Portugal by Matos et al. (1995) [20] showed clear evidence of
hydrogen sulfide being present in concentrations of up to about 300 mg/m3 in the atmosphere of
gravity sewers and sewer structures. Hydrogen sulfide can be removed due to the presence of
metals such as iron, zinc, lead and copper in the wastewater leading to precipitation. Oxidation of
sulfides is possible due to the volatilization of H2S and subsequently the condensate dissolves in the
sewer crown. Biological and chemical oxidation of sulfides in sewer network can occur, making the
total oxidation processes complex [3]. Biological oxidation of hydrogen sulfide can take place at the
sewer surfaces exposed to the sewer atmosphere. The sulfur oxidizing bacteria (SOB) oxidize the
dissolved H2S and other sulfur compounds (e.g., S2O32- and S0 ) to sulfuric acid (H2SO4) (Figure 3) [21].
The sulfuric acid generated will then attack the sewer infrastructure by reacting with corrodible
compounds, such as cement and metals (will be discussed further).
Figure 3: Transfer of H2S across the air-water interface (from [22])
Figure 4 outlines the major processes and associated pathways related to the bulk wastewater,
biofilms, sediments, sewer atmosphere and sewer surfaces exposed to the sewer atmosphere. These
processes are the formation of sulfide, volatilization of hydrogen sulfide to the sewer atmosphere,
chemical and biological oxidation of sulfide and precipitation of sulfide in the sewage [15]. Sulfide
formation is typically recognized as a problem in pressure mains due to the long residence time of
wastewater being more than 1-2 hours, insufficient reaeration potential and relatively high
6
temperatures [23]. This explains the difference between measured sulfide concentrations in
Australia and Belgium because countries like Australia have long pressure mains and high
temperatures, leading to higher observed sulfide concentrations. This results in an order of sewer
corrosion related to sulfide concentrations: dissolved sulfide concentrations of 0.5, 3 and 10 mg S L -1
are considered as low, moderate and high respectively, in terms of problems related to sewer
corrosion [18, 19].
Figure 4: Major sulfide processes and their effect factors associated with the sulfur cycle in the sewer system. SOB: sulfur
oxidizing bacteria; SRB: sulfate reducing bacteria; ORP: oxidation-reduction potential; DO: dissolved oxygen (from [3], [15])
2.2.3 Microbial processes in sewer systems
Wastewater characteristics play an important role for the nature and the course of the sewer
processes. Microbial transformations, and thereby biochemical transformations, characterize the
sewer environment in terms of wastewater quality. More specifically for sulfate reduction,
fundamental parameters influencing this process are the presence of sulfate, quantity and quality of
biodegradable organic matter, temperature and pH [24]. On the other hand, the physicochemical
characteristics, e.g. transfer of substances across the water-air interface (oxygen transferred into the
water phase and emission of volatile compounds from the water phase) and diffusion in the biofilm,
play an important role and are therefore strongly related to microbial transformations. As previously
7
mentioned, the hydraulics (e.g. wastewater flow velocity, area-to-volume ratio, etc.) and the sewer
solids transport processes also have a pronounced impact on the sewer performance [1, 24].
Figure 5: Role of sulfur oxidizing bacteria (SOB) and sulfate reducing bacteria (SRB) (from [22])
Bacteria can be subdivided into several groups depending on requirements for oxygen, source of
energy and type of environment in which they survive.
When using a classification based on oxygen, one can distinguish organisms that require oxygen for
survival and growth, termed obligate aerobes, while bacteria that require low levels of oxygen are
termed microaerophilic [25]. Furthermore, bacteria that can survive in an anaerobic environment
but prefer aerobic conditions are facultative anaerobes, while bacteria that cannot tolerate oxygen
are called obligate anaerobes. Even though obligate anaerobic bacteria do not grow in the presence
of oxygen, they are isolated routinely from oxygenated environments associated with particles in
association with other bacteria that effectively remove oxygen in the vicinity of the anaerobes.
The fundamental process in energy-conserving metabolism in all respiratory processes is the transfer
of hydrogen from a state more electronegative than that of H+/H2O to that of water [26].
Heterotrophic bacteria can assimilate almost any available carbon molecule, from simple alcohols
and sugars to complex polymers by dehydrogenation of organic compounds. Autotrophic bacteria
obtain the energy required for syntheses from absorbed light (photolithotrophic) or from chemical
sources (chemolithotrophic) and use inorganic nutrients for all metabolic requirements. When
autotrophic and heterotrophic mechanisms operate simultaneously, the metabolism is called
mixotrophic. Autotrophs can use elements or ions (e.g., ammonia (NH3), nitrite (NO2–), methane
(CH4), hydrogen gas (H2), sulfate (SO42-), ferrous iron (Fe2+) and manganese ion (Mn2+)) as sources of
8
energy. So eventually, the biological sulfur cycle is controlled by heterotrophic or mixotrophic SRB
and chemolithotrophic organisms. Acidophilic autotrophic microorganisms oxidise inorganic
compounds, such as iron and sulfur at low pH, to derive their metabolic energy [22].
Table 2: Oxidation states of sulfur (modified from [27])
OXIDATION STATE
+6
+5
+4
+4
+3
+2
0
-2
NAME
Sulfate
Dithionate
Sulfite
Disulfite
Dithionate
Thiosulfate
Elemental sulfur
Sulfide
FORMULA
SO42S2O62SO32S2O52S2O42S2O32S0
S2-
Sulfate reducing bacteria (SRB) are a metabolically versatile group of micro-organisms and include
Bacteria and Archaea. SRB are obligate anaerobic microorganisms that use sulfate or other oxidised
sulfur compounds as terminal electron acceptor (Table 2) [28]. Most of the SRB described to date
belong to one of five phylogenetic lineages [29]: (1) the mesophilic δ-Proteobacteria with genera
Desulfovibrio, Desulfobacterium, Desulfobacter, Desulfobulbus, Desulfomicrobium, Desulfomonas,
Desulfococcus, Desulfomonile, Desulfonema and Desulfosarcina; (2) the thermophilic gram-negative
bacteria with the genera Thermodesulfovibrio, Thermodesulfobacterium and Thermodesulfobium; (3)
the
gram-positive
bacteria with the
genera Desulfotomaculum, Desulfosporosinus and
Desulfosporomusa; (4) the Euryarchaeota with the genus Archaeoglobus; and (5) the Crenarchaeota
with the genera Thermocladium and Caldirvirga. The last two phylogenetic lineages (SRB belonging
to the Euryarchaeota and Crenarchaeota) have not been reported to be found in wastewater related
sources and are thus considered irrelevant for our further research [30]. The largest of the SRB
groups is affiliated with the δ-class of Proteobacteria, which currently contains over 30 gramnegative sulfate reducing genera. Two well identified genera of SRB are Desulfotomaculum (sporeforming straight or curved rods) and non-sporing genus Desulfovibrio (curved, motile vibrios or rods)
[25].
In addition to its obvious importance in the sulfur cycle, SRB are also important regulators in organic
matter turnover and biodegradation of recalcitrant pollutants. Sulfate has first to be activated at the
expense of two ATP-equivalents per sulfate molecule since it is a fairly stable molecule. This
activation is catalysed by ATP sulfurylase and yields adenosine 5’-phosphosulfate (APS) [28]. Typical
substrates, electron donors and energy (and carbon) sources, for SRB are lactate, ethanol,
propionate and H2. In the dissimilatory sulfate reduction, sulfide is released into the environments
9
whereas in the assimilatory sulfate reduction, sulfide is converted into organic sulfur compounds,
e.g. amino acids [22]. Sulfur compounds such as sulfite and thiosulfate are also reduced to sulfide if
present in wastewater. Nielsen et al. (1991) [31] shows that these compounds are reduced instead
of sulfate by the sulfate reducing bacteria and often result in increased sulfide production rates. On
the other hand, organic sulfur compounds usually seem to be insignificant sources for sulfide
production [32].
The results of a phylogenetic analysis by means of fluorescence in-situ hybridization (FISH) suggested
that at least six phylotypes of sulfur oxidizing bacteria (SOB) were involved in the microbial induced
concrete corrosion (MICC) process, being Thiothrix, Thiomonas, Thiobacillus, Halothiobacillus,
Acidiphilium and Acidithiobacillus-species [21]. Factors like the pH of the concrete surface as well as
trophic properties (e.g. autotrophic or mixotrophic), and the ability of the SOB to utilize different
sulfur compounds result in a certain sequence of SOB species present during the corrosion process
[21]. Furthermore, the vertical distribution of microbial community members revealed that A.
thiooxidans, a hyperacidophilic SOB and member of the genus Thiobacillus, was the most dominant
species present in the heavily corroded concrete layer, accounting for 70% of the population after
one year [33]. Two major pathways are important for sulfur oxidation in sewers, being the oxidation
of thiosulfate (S2O32-) and the oxidation of elemental sulfur (S0) (Figure 5). Since there are reactions
that cross both pathways, the pathways may occur at the same time, but the second pathway
dominates because of the dominance of A. thiooxidans (related to its adaptation to rapid sulfur
oxidation at low pH) [33]. It is most abundant at the highest surface layer and its presence decreases
with depth because of oxygen and H2S transport limitations. Thus, the production of sulfuric acid by
A. thiooxidans occurs mainly on the concrete surface and subsequently the sulfuric acid penetrates
through the corroded concrete layer and reacts with the concrete below, leading to progressive
corrosion [21, 34].
2.2.4 Methane production in sewer systems
Under anaerobic and reduced conditions, methanogens produce CH4 from either the reduction of
CO2 with H2 (hydrogenotrophic) or from the fermentation of acetate to CH4 and CO2 (acetoclastic). In
nature, the latter mechanism accounts for about two-thirds of the CH4 emitted [35].
Understanding methane production in sewer systems is important for several reasons:

Uncontrolled methane release is potentially unsafe since it forms an explosive mixture in air
at low concentrations and therefore poses occupational health and safety risks;
10

Methane contributes significantly to the greenhouse effect with a lifespan of about 12 years
and a global warming potential of roughly 21 times higher than carbon dioxide [6];

Methanogenesis causes loss of soluble COD which may have detrimental effects on the
operation of wastewater treatment plants (WWTP) with biological nutrient removal.
A significant contribution to the greenhouse effect is made by the produced methane since it will
remain dissolved until there is a point of release to the atmosphere [36]. In a rising main pipeline
that is full and pressurised the methane will become supersaturated. A vessel that is open to the
normal atmosphere will lead to a sudden discharge because of the large driving force that exists for
methane to escape from the liquid phase sewage to reach a new equilibrium due to the very low
methane concentrations in the atmosphere (1800 ppb) [37, 38]. Furthermore, the stripping rate will
be high because of a high mass transfer rate coefficient due to turbulence when the sewage is
discharged to a wastewater treatment plant (WWTP). Hence, the mass of methane released to the
atmosphere will be in the range approximately 40-250 tonnes CH4/year, based on a dissolved
methane content of approximately 20-120 mg L-1 as COD. At a global warming potential of roughly
21 times (relative to CO2), the released methane will contribute approximately 900-5300 tonnes CO2eq/year. This means an additional GHG contribution of roughly 12-72% from sewage methane over
and above that from the WWTP itself (designed for 100.000 person equivalents). This clearly shows
that methane production in sewer systems could be a very significant source of GHG in wastewater
systems, and should be managed [38].
2.2.5 Competition and coexistence between sulfate reducing bacteria (SRB) and
methanogens
It was found that methanogens were outcompeted by anaerobic bacteria using either SO42- or Fe3+ as
terminal electron acceptor for H2, a common methanogenic substrate. This is due to the fact that
SRB have a lower half-saturation constant and threshold for H2 uptake in comparison to
methanogens. Thus, SRB can lower the hydrogen partial pressure below levels that methanogens
could effectively utilize and eventually prevent the activity of methanogens [39, 40]. A similar
competitive mechanism for acetate between SRB and methanogens has also been reported [41].
Although, the supply of methanogenesis precursors (volatile fatty acids (VFA) or hydrogen) is
unlikely to be limiting within the biofilm, so the lower affinity of methanogens for these reactants
does not limit their growth when they grow in deeper layers of the biofilm. Since sulfate only
partially penetrates the biofilm, two different zones may appear in the biofilm: a sulfate reducing
anaerobic zone nearer the surface, dominated by SRB and a deeper anaerobic zone, dominated by
methanogens. Thus, the extent of methanogenesis in a sewer system is inversely proportional to the
11
sulfate penetration length into the biofilm [38] (Figure 6). Methane and hydrogen sulfide were
reported to be produced simultaneously in anaerobic sewer biofilms. This implies the coexistence
and function of SRB and methanogens in sewer even though they compete for VFA as discussed
above. The small thermodynamic difference between standard reduction potential E'0 of
methanogenesis (CO2/CH4, -240 mV) and sulfate reduction (SO42-/HS-, -217 mV) explains this
coexistence.
Figure 6: Schematic representation of a sewer biofilm with sulfate partial penetration (adapted from [38])
2.2.6 Problems related to H2S
2.2.6.1
Odour and health risks
Generation and emission of hydrogen sulfide is a universal sewer maintenance problem because of
its noxious odour, health hazard and severe corrosive attack on concrete sewers and related
materials. Although the fact that this is a universal problem, it is particularly widespread in countries
with warm climate [42]. H2S is potentially very dangerous because its unpleasant and strong smell is
quickly lost as the concentration increases over 10-50 ppm. It causes eye and respiratory injury when
concentration in the atmosphere goes to 50-300 ppm. When its concentration goes above 300 ppm,
hydrogen sulfide becomes life threatening (see Table 3) [17].
12
Table 3: Human health effects at various hydrogen sulfide concentrations (modified from [17])
EXPOSURE (PPM)
0.03
4
10
20
30
100
200
250
300
500
700
SYMPTOMS
Can smell.
May Cause eye irritation. Respiratory protection equipment must be used as it
damages metabolism.
Maximum exposure 10 minutes. Impairs sense of smell in three to 15 minutes. Causes
'gas eye' and throat injury.
Exposure for more than one minute causes severe injury to eye nerves.
Loss of smell, injury to blood barrier through olfactory nerves.
Respiratory paralysis in 30 to 45 minutes. Needs prompt artificial resuscitation. Will
become unconscious quickly (15 minutes maximum).
Serious eye injury and permanent damage to eye nerves. Stings eye and throat.
Prolonged exposure may cause the lung tissue to swell and fill up with water.
Loses sense of reasoning and balance. Respiratory paralysis in 30 to 45 minutes.
Respiratory distress. Will become unconscious in 3 to 5 minutes. Immediate artificial
resuscitation is required.
Breathing will stop and death will result if not rescued promptly. Permanent brain
damage may result unless rescued promptly.
Because hydrogen sulfide is a gas, inhalation is the major route of exposure to hydrogen sulfide.
Respiratory, neurological, and ocular effects are the most sensitive end-points in humans following
inhalation exposures. Health effects that have been observed in humans following exposure to
hydrogen sulfide include death and respiratory, ocular, neurological, cardiovascular, metabolic, and
reproductive effects. There have been numerous case reports of human deaths after single
exposures to high concentrations (≥ 300 ppm) of hydrogen sulfide gas, and most fatal cases
associated with hydrogen sulfide exposure have occurred in relatively confined spaces where the
victims lost consciousness quickly after inhalation of hydrogen sulfide [43].
2.2.6.2
Concrete and metal corrosion
Hydrogen sulfide emission in sewer systems is associated with several problems, including biogenic
corrosion of concrete, besides the release of obnoxious odours to the urban atmosphere and toxicity
of hydrogen sulfide gas to sewer workers. Minor problems of concrete corrosion have been reported
when the concentration of total sulfide in the wastewater is within the range of 0.1-0.5 mg S L−1 [18,
19]. Severe concrete corrosion may occur at sulfide concentrations from 2.0 mg S L−1. In Flanders
(Belgium), biogenic sulfuric acid corrosion of sewers is approximated at €5 million per year,
representing about 10% of total cost for wastewater collection and treatment [4]. In Australia, the
costs can go up to $100 million per year [5]. This cost is related to new sewer-materials, traffic costs
etc. Several methods have been investigated to solve the biogenic corrosion problem, , i.e.
optimizing the sewer hydraulic design to minimize sulfide generation, sulfate source control
technologies such as urine separation or pretreatment, improving the resistance of sewer pipes to
13
biogenic corrosion (via application of protective coatings such as bituminous and coal tar products,
vinyl and epoxy resins, cement and polyethylene linings …) and decreasing hydrogen sulfide emission
from sewage [3].
The corrosion process has several stages. The pH of newly manufactured ranges from 11.0 to 13.0
due to formation of calcium hydroxide (Ca(OH)2), which inhibits the growth of bacteria. Weathering
leads to conversion which converts calcium hydroxide into calcium carbonate (CaCO3), resulting in a
decrease of the surface pH due to dissociation to bicarbonate. In combination with carbon dioxide
(CO2) dissolving into the water a reduction of the surface pH to around 7.4 is established. Further
reduction of the pH (due to biological oxidation of H2S) leads to a surface pH below 5.0, which leads
to dominance of A. thiooxidans. The formed sulfuric acid (H2SO4) reacts with calcium hydroxide or
metallic components of concrete sewer walls. First, sulfuric acid reacts with calcium hydroxide
forming calcium sulfate (CaSO4) according to the following equation [44]:
Ca(OH)2 + H2SO4 → CaSO4 + 2H2O
(6)
Calcium sulfate is subsequently hydrated to form gypsum (CaSO4·2H2O), which would further react
with
the
tricalcium
aluminate
hydrate
(3CaO.Al2O3.6H2O)
to
form
ettringite
(3CaO.Al2O3.3CaSO4.32H20), an expansive product:
3CaSO4.2H2O + 4 CaO.Al2O3.13H2O → 3CaO.Al2O3.3CaSO4.32H20 (ettringite) (7)
The expansion leads to a decrease of the concrete strength and loss of bond between the cement
paste and aggregate. Wastewater can easily wash away gypsum and ettringite, which leads to
further exposure of fresh material to sulfuric acid attack and corrosion [34, 45].
Metal structures are also corroded by the corrosive effect of sulfuric acid. In addition, the bacterial
activity results in a decrease of the pH in the immediate area of the bacteria. The pH change results
in a lower electrical potential and thus, this area may act as a cathode. The adjacent metal becomes
an anode, and electrochemical corrosion may occur [34].
14
Figure 7: Concrete biodeterioration influenced by SRB and SOB; (a) Initial state, (b) Swelling of concrete and (c) Concrete
cracking (modified from [46])
2.2.7 Abatement strategies
As previously mentioned, sulfide production results in noxious odour, corrosive effects on sewer
infrastructure, and negative health aspects. In addition, sulfide in the wastewater may also affect the
biological processes in the wastewater treatment and it may be toxic for fish in streams affected by
overflow events [32]. Therefore, there are a few abatement strategies for sulfide control, e.g.
injection of oxygen, nitrate, hydrogen peroxide and chlorine or iron salts [3, 47-49].
15
Figure 8: Chemical and biological abatement strategies for H2S emission control in sewer systems (from [50])
Iron salts and oxygen are dosed at fewer sites, but these two chemicals have a much higher
contribution than the others in terms of sewage flow treated (Figure 9) [51]. Around 16% of the total
sewage is for example treated with oxygen, while the sewage receiving iron salts dosing accounts for
about 66% of the total sewage with chemical dosing [51]. Sodium hydroxide (NaOH) is only used in
systems with low flows (average dry weather flow lower than 0.5 ML d-1) and small pipe diameters
(more than 95% of the dosing is conducted in pipes with diameters smaller than 0.3 m). Typically, a
shock treatment is used to produce a pH of 12.5-13.0 in wastewater for a period of 20-30 minutes.
The inactivation effect on the SRB is only local because of the high buffer capacity of sewage, which
results in a high caustic requirement [3]. Thus, pipes with high surface area to volume ratios (A/V)
are more favourable for this treatment, since less amount of sodium hydroxide is required per
volume of sewage. Therefore, sodium hydroxide is a cost effective solution for sulfide control in
small systems with low flow rates and high A/V ratios.
Similarly, to sodium hydroxide, magnesium hydroxide is mainly applied in sewers with low flows and
small diameters. About 80% of the sites that are dosed with magnesium hydroxide are sewers with
average dry weather flows below 1 ML d-1 and pipe diameters between 0.15 and 0.3 m. Iron salts are
preferentially used in medium and large systems (around 80% of the sites with flows larger than 1
ML d-1). Accordingly, these systems have big pipe diameters. The dosing of iron salts is a simple and
cost effective method for the control of sulfides via precipitation reactions occurring in the bulk
liquid phase. Although iron dosing is appropriate for both small and large systems, it has gained a
wider application in large systems because other chemicals are less suitable for sulfide control in
such pipes [51].
16
Figure 9: Contribution of each chemical to the overall sulfide control (modified from [51])
Table 4 shows the advantages and disadvantages of the current methods that are nowadays most
applied for the control of odour and corrosion in the sewer system. As mentioned before, the
dosage of iron salts is the most common implemented technique in the world [52, 53]. Developing
an in-situ technique based on the addition of iron and alkalinity would address some of the
disadvantages related to iron salts. Storage, handling and transport of the chemicals would no longer
be necessary. Next to this, anions (e.g. chloride, nitrate and sulfate) are no longer dosed into the
sewer system. Also, the advantages of pH elevation thanks to alkalinity production at the cathode
are now obtained without the need of transport of caustic chemicals.
Table 4: Advantages and disadvantages of current methods for odour and corrosion control in collection systems and
pumping stations (modified from [54])
Method
Advantage
Disadvantage
Increasing DO concentration
by air (or pure oxygen) injection
to prevent anaerobic conditions
and expedite oxidation of the
sulfide in the bulk sewage
- air is readily available, so no
need for transportation
- no addition of chemicals
- no negative byproduct
formation
Dosage of nitrate salts to elevate
redox potential (encouraging
anoxic activity rather than
anaerobic)
- nitrate salts are highly soluble
thus high concentration may be
attained in the water, allowing for
full penetration into sewer
biofilm, with effective prevention
- low solubility of oxygen in water
results in a local effect and thus
multiple injection points are
required
- high energy and maintenance
requirements
- a preventative measure rather
than a sulfide removal method
- sulfide generated before point
of injection will not be treated
- unwanted addition of counter
17
of anaerobic conditions
- relatively inexpensive chemical
Addition of iron salts (either ferric
or ferrous or combination of
both)
- specific and effective oxidation
of sulfide by Fe3+, followed by FeS
precipitation
- not toxic
- no harmful byproducts
- relatively inexpensive
Elevation of pH to above 8.5 by
addition of strong base which
shifts the equilibrium of the
dissolved sulfide towards the nonvolatile species (S2−, HS−)
- may be effective in cases where
local odour abatement is required
cations to sewage (e.g. Na+, K+,
Ca2+)
- need for frequent transport of
chemicals to injection point
- possible negative effects on
wastewater treatment plant
because of nitrate load
- need for a control system in the
sewer to optimize dosages
- does not oxidize or precipitate
any other odorous compounds
apart from sulfides
- may add undesirable anions to
the water
- transport, handling and storage
of chemicals
- may cause unsolicited
flocculation and settling in the
sewer
- may precipitate with P
compounds thus demand will
increase beyond stoichiometry
- effective only locally (since the
pH is bound to decrease
downstream)
- high requirement due to high
buffering capacity of the sewage
The location is also important for the effectiveness of treatment and the operational cost. The
preferred dosing location for the majority of the chemicals is before the wastewater enters the rising
main, either at the wet well or the pumping station [51]. The effectiveness of the dosage of iron salts
is not affected by the location, but the hydraulic retention time (HRT) has to be high enough to allow
sufficient time for sulfide precipitation [55]. Thus, around 70% of iron dosing is done at upstream
locations (mainly at the wet well). However, recent lab-scale study demonstrated that the addition
of Fe3+ significantly inhibits the SRB activity of anaerobic sewer biofilms [56]. This inhibitory effect
needs to be verified for real sewers, but this would result in a better control of sulfide along the
entire pipe. In the absence of sulfide, the iron salts initially react with some other anions (e.g.
phosphate and hydroxide). Once sulfides are formed, iron ions will be made available for sulfide
precipitation due to the lower solubility of iron sulfide (FeS) in comparison to iron phosphate (FePO4)
and iron hydroxide (Fe(OH)3) precipitates.
18
Chemicals are dosed continuously or just during pumping events or periodically for several days or
weeks. It’s not a surprise that the dosing rate has a critical impact on the effectiveness of sulfide
mitigation and chemical consumption. Nitrate and oxygen oxidise sulfide, but they do not have longlasting inhibitory or toxic effects on SRB [13, 49]. The principle behind dosing iron salts is the
precipitation of sulfides. Thus, sulfide levels in the pipe are extremely important when dosing with
these three chemicals. Multiple techniques are nowadays available to be able to choose the most
(cost) effective approach. For instance, mathematical models can be used for the selection of the
most suitable chemical, dosing location and dosing strategy. On-line control can then further
improve the effectiveness of dosing and eventually reduce chemical costs [56].
Zhang et al. (2009) [56] showed that the long term addition of Fe3+ considerably reduces the sulfate
reducing and methanogenic activities of sewer biofilms. A first hypothesis for the mechanism likely
responsible for the inhibition is suggested by Lovley et al. (1983) [41] and Van Bodegom et al. (2004)
[57]. The idea is that Fe3+ in sediments leads to the inhibition of sulfate reduction and methane
production, caused by the competition of Fe3+-reducing bacteria with SRB and methanogens. The
hydrogen and acetate uptake by SRB and methanogens for metabolism is prevented because the
concentrations of hydrogen and acetate are maintained at very low level by Fe3+-reducing
organisms. Although, this mechanism is unlikely responsible for the inhibitory effect in sewers
because unlike in sediments, organic compounds in sewers are abundant and therefore not
expected to be a limiting substrate for the growth of Fe3+-reducing organisms, SRB and
methanogens. This is also confirmed by experimental data that showed excessive VFA
concentrations [58].
Thus, a new hypothesis was suggested by Utgikar et al. (2002) [59], where the deposit of metal
sulfides on the sewer biofilm could cause the inhibition of the activity of the cells present. During the
Fe3+ treatment, the wastewater contained high concentrations of Fe2+ sulfide precipitates because of
high sulfide production in rising main sewers. These insoluble sulfides of heavy metals may reduce
access of reactants (sulfate, VFA and organic matter) in vicinity of bacterial cell to the necessary
enzymes, thus reducing the further metabolism of bacteria. Fe3+, like many heavy metals (e.g.
copper, zinc and nickel), could also deactivate enzymes of microorganisms by reacting with their
functional groups, denature proteins of microorganisms and compete with essential cations utilized
by microorganisms. These processes cause adverse effects on the activities of microorganisms [60].
The metal sulfides are although not overtly toxic to the SRB, as the SRB cultures are found to be still
viable in presence of metal sulfides. Jong et al. (2003) [61] stated that metal toxicity and inhibition in
SRB systems are strongly influenced by the chemical and physicochemical properties of the
surrounding SRB environment, where the resistance against the inhibitory effect is bigger in mixed
19
cultures (as can be found in sewer systems). Further expansion of this hypothesis would result in the
idea of less inhibition for the methanogens since these organisms are situated deeper in the biofilm
layer, which results in less interaction with the iron salts (thus less deactivation of enzymes, etc.).
Although, the access of reactants is still reduced because of the formation of the crust on the biofilm
layer, which results in (lower) inhibition for the methanogens.
Figure 10: Mechanism of metal sulfide inhibition (from [59])
2.3 Electrochemical introduction
An electrochemical cell contains two electrodes that allow transport of electrons. These electrodes
are separated by an electrolyte that allows movement of ions but blocks movement of electrons.
Electrons travel from one electrode to another through an external conducting circuit, doing work or
requiring work in the process [62]. Figure 11 shows an example of an electrochemical cell where
sulfide (HS-) is oxidized to elemental sulfur (S0) on the anode, which results in a layer of elemental
sulfur (S0) on the anode that causes passivation of the electrode [63]. Oxygen is reduced at the
cathode to form water. Both reactions are separated by a cation exchange membrane (CEM).
20
e-
H2O
S0
C+
O2
HSAnode
Cathode
Reduction
Oxidation
Figure 11: Simplified scheme of an electrochemical cell with CEM
Reduction occurs at one electrode (the cathode) and oxidation occurs at the other electrode (the
anode), so the two processes are in fact separated. Thus, the complete redox reaction is broken into
two half-cells. The rate of these reactions can be controlled by externally applying a potential
difference between the electrodes, for example with an external power supply. Even though the
half-cell reactions occur at different electrodes, the rates of reaction are coupled by the principles of
conservation of charge and electro neutrality. The flow of current is continuous in this case. At the
interface between the electrode and the electrolyte, the flow of current is still continuous, but the
identity of the charge-carrying species changes from being an electron to being an ion. In the
electrolyte, electro neutrality requires that there should be the same number of equivalents of
cations as anions [62]:
∑𝑖 𝑧𝑖 𝑐𝑖 = 0
(8)
Where ∑i is the sum over all species i in solution, and 𝑐𝑖 and 𝑧𝑖 are the concentration and the charge
number of species i, respectively.
The rate of reaction is related to the current through Faraday’s law. It states that the rate of
production of a species is proportional to the current, and the total mass produced is proportional to
the amount of charge passed multiplied by the equivalent weight of the species:
21
𝑚𝑖 =
𝐼 𝑡 𝑀𝑖
𝐹 𝑧𝑖
(9)
Where 𝑚 is the mass of the substance liberated at an electrode in grams, 𝑄 is the total electric
charge passed through the substance and was replaced by I.t (current multiplied by time), 𝐹 =
96485.3 C.mol−1 is the Faraday constant, 𝑀 is the molar mass of the substance and 𝑧 is the valency
number of ions of the substance (electrons transferred per ion).
2.3.1 Thermodynamics
As previously mentioned, the reaction in an electrochemical cell can either occur spontaneously or
may need an external input of energy in order to take place. This can be expressed by the energy
change in the reaction described by a change in Gibbs free energy (∆𝐺) for each half cell:
∆𝐺 = (∑𝑖 𝑠𝑖 µ𝑖 )𝑎𝑛𝑜𝑑𝑒 − (∑𝑖 𝑠𝑖 µ𝑖 )𝑐𝑎𝑡ℎ𝑜𝑑𝑒
(10)
With µ𝑖 = chemical potential of species i, 𝑠𝑖 = stoichiometric coefficient of species i
The change in free energy (∆𝐺) is a measure of the maximum amount of work that can be
performed during a chemical process (∆𝐺 = 𝑤𝑚𝑎𝑥 ). For ∆𝐺 < 0 the reaction occurs spontaneously
and the electrochemical cell produces electricity (galvanic or fuel cell). If ∆𝐺 is positive, the reaction
will require energy and the electrochemical cell becomes an electrolysis cell. Thus, there must be a
relationship between the potential of an electrochemical cell and ∆𝐺:
ΔG = −𝑛 𝐹 𝐸𝑐𝑒𝑙𝑙
(11)
𝐸𝑐𝑒𝑙𝑙 is the electromotive force (also called cell voltage) between two half-cells and is also given by
𝐸𝑐𝑎𝑡ℎ𝑜𝑑𝑒 − 𝐸𝑎𝑛𝑜𝑑𝑒 . The greater the 𝐸𝑐𝑒𝑙𝑙 of a reaction, the greater the driving force of electrons
through the system and the more likely the reaction will proceed (more spontaneous). The Gibbs
free energy at defined pressure and temperature can be derived using:
ΔG = ΔG° + R T ln(Q)
(12)
With 𝑄 the reaction quotient given as the activities of the products divided by the activities of the
reactants for the reaction 𝑠𝐴𝐴 + 𝑠𝐵𝐵 → 𝑠𝐶𝐶 + 𝑠𝐷𝐷. Activities are used to express the effective
concentration of a component in a mixture (depending on temperature, pressure and other present
chemicals).
𝑎 𝑠𝐶 + 𝑎 𝑠𝐷
𝑄 = 𝑎𝐶 𝑠𝐴 + 𝑎𝐷 𝑠𝐵
𝐴
𝐵
(13)
22
For dilute systems activities can be given as concentrations and equation 13 becomes
[𝑟𝑒𝑑]
𝑄=
(14)
[𝑜𝑥]
with [𝑟𝑒𝑑] the concentration of the products at the reduced side and [𝑜𝑥] the concentrations of the
reactants or oxidized species.
0
We also know that ΔG = −𝑛 𝐹 𝐸𝑐𝑒𝑙𝑙 and ΔG° = −𝑛 𝐹 𝐸𝑐𝑒𝑙𝑙
, so via substitution we obtain:
0
− 𝑛 𝐹 𝐸𝑐𝑒𝑙𝑙 = − 𝑛 𝐹 𝐸𝑐𝑒𝑙𝑙
+ 𝑅 𝑇 ln(𝑄)
(15)
Dividing both sides of this equation by −𝑛 𝐹,
0
𝐸𝑐𝑒𝑙𝑙 = 𝐸𝑐𝑒𝑙𝑙
−
𝑅𝑇
𝑛𝐹
ln(𝑄)
(16)
which is called the Nernst equation and which is used to calculate theoretical potentials of half-cell
reactions under specific pH, temperature and concentrations.
2.3.2
Potential losses
It is expected that there should be an agreement between the calculated amount of substances
dissolved as a result of passing a definite quantity of electricity (derived from Faraday’s law) and the
experimental amount determined. Significant error may be introduced if insufficient attention is
given to the geometry of the electrode and the optimum conditions of operation of the EC cell. One
area of uncertainty is in the measurement of potential of the EC cell. The measured potential is the
sum of three components:
ηAP = ηκ + ηMt + ηIR
(17)
where ηAP is the applied overpotential (V), ηκ the kinetic overpotential (V), ηMt the concentration
overpotential (V), ηIR the overpotential caused by solution resistance or ohmic drop (IR-drop, V). The
IR-drop is related to the distance (d in cm) between the electrodes, surface area (A in m2) of the
cathode, specific conductivity of the solution (κ in mS.m−1) and current (I in A) by the equation
shown below
ηIR =
Id
Aκ
(18)
An important choice of parameters that can be controlled in this set-up is the choice of spacing. The
IR-drop, which is the result of the resistance of the electrons through the electrodes and of ions
through the electrolyte (equation 18). Engineered systems are designed to maximize the process
23
efficiency and minimize the energy input. In this case this means a limitation of the ohmic resistance,
which means a lower current density and smaller spacing [7]. Because of the reduction of the initially
formed iron ions to Fe0 along the flow path in case when the spacing is too low, there is the
possibility to select an optimal spacing which is a compromise between energy input losses due to
ohmic drop and avoiding a decrease in efficiency due to rereduction of the previously formed iron
ions (Figure 12).
Figure 12: Optimisation of the spacing. Blue path: Sewage is provided with iron ions; Red path: A rereduction of iron ions
can occur at the cathode (from [7])
Concentration overpotential (ηMt, V), also known as mass transfer or diffusion overpotential, is
caused by the change in anolyte concentration occurring in the proximity of the electrode surface
due to electrode reaction. This overpotential can be reduced by increasing the masses of the metal
ions transported from the anode surface to the bulk of the solution and this can be achieved by
enhancing the turbulence of the solution.
Kinetic overpotential (ηκ, also called activation potential) has its origin in the activation energy
barrier to electron transfer reactions. The kinetic overpotential is especially high for evolution of
gases on certain electrodes. Both kinetic and concentration overpotential increase as the current
increases [64]. However, the effects of these changes need to be investigated for specific types of
physical and chemical species in aqueous solution.
2.3.3 Iron electrocoagulation
Theoretically, the electrolytic oxidation of iron results in ferrous (Fe2+) or ferric (Fe3+) generation (eq
19-21) at the anode (vs. SHE) [65]:
24
Fe(s) → Fe2+ + 2e-
E0 = +0.44 V
(19)
Fe2+ → Fe3+ + e-
E0 = -0.77 V
(20)
Fe(s) → Fe3+ + 3e-
E0 = +0.04 V
(21)
In aqueous environments Fe2+ ions react with sulfide as shown below:
Fe2+ + HS- → FeS(s) + H+
(22)
Fe3+ ions oxidize sulfide to elemental sulfur while being reduced into Fe2+ ions, which then
precipitates with sulfide to form ferrous sulfide precipitants:
2Fe3+ + HS- → 2Fe2+ + S0 + H+
(23)
The overall reaction of Fe3+ with sulfide can be expressed as:
2Fe3+ + 3HS- → 2FeS(s) + S0 + 3H+
(24)
When using Fe2+ ions a higher Fe2+ to S2− ratio is often attributed to the reactions of Fe2+ with other
competing anions in sewage which reduce the availability of Fe2+ for sulfide precipitation, such as
carbonate (CO32−), nitrilotriacetate (C6H6NO6), ethylenediaminetetraacetate (EDTA, C10H13N2O8) and
linear alkylbenzene sulfonate (LAS) [54, 66-70]. pH also has a significant impact on the Fe2+ to S2−
ratio with a lower pH increasing the Fe2+ demand. Furthermore, a higher Fe2+ to S2− ratio is required
when a lower sulfide concentration is to be achieved because of the relation with the solubility
product of FeS [54]. Firer et al. (2008) [54] recommended a Fe2+ to S2− ratio of 1.3/1 mole mole-1 in
order to maintain a dissolved sulfide concentration below 0.1 mg S L−1.
Stoichiometrically, it can be hypothesized from equations 22 to 24 that Fe3+ is more effective than
Fe2+. In the complete absence of dissolved oxygen, Fe3+ is demonstrably better at reducing sulfide
levels than Fe2+. Tomar et al. (1994) [71] reported that the dosage requirement of a ferric salt
solution was 20% lower than the ferrous salt solution for complete sulfide control. However,
experiments have shown that even a very low concentration of dissolved oxygen (approximately 0.2
mg L-1) greatly improves the effectiveness of Fe2+ [56]. Since the cost of ferrous and ferric salts
(normalized per g Fe) is almost identical this suggests that the use of ferric salts should be preferred
[54]. Further on in this thesis it will become clear that dosing a mixture of ferrous and ferric ions is
even better for the efficiency of the process. Padival et al. (1995) [12] observed that the dosage of a
25
mixture of ferric and ferrous chloride with a molar ratio of 1.9/1 at 16 mg Fe L-1 was able to reduce
sulfide levels in the 40 km downstream sewer pipe line to below 0.5 mg S L-1. The average sulfide
level prior to the dosage was 6.4 mg S L-1, so the ratio between iron injected (Fe2+ + Fe3+) and sulfide
removed was approximately 1.5/1 mole mole-1. The reason why a mixture of Fe3+ and Fe2+ is actually
more efficient than just dosing a single form of iron ions is because of the following equation:
Fe2+ + Fe3+ + 4HS- → Fe3S4(s) + S0 + 4H+
(25)
Due to the formation of Fe3S4 as an intermediate and subsequently transformation to FeS2, a mineral
more stable and less soluble than FeS is formed [72, 73].
26
2.4 Research objectives
In brief, in-situ electrochemical generation of iron and alkalinity from sewage is dosed to a lab-scale
system. The aim is to design and test this system as a representation of the real application (pilot
plant at Gold Coast, QLD, Australia). Four research objectives are proposed in order to fill up cavities
in previous research. A mimicking wet-well and sewer rising main reactors will be used to determine
the feasibility of this technology.
e-
eSewage
out
OHFe3+
reactor
Sewage
in
Pumping station
Rising main
Figure 13: Simplified diagram of real application of in-situ electrochemical generation of iron and alkalinity from sewage for
sulfide control. (from research-plan made by Hui-Wen Lin (tutor in Australia))
2.4.1 Objective 1: Demonstration of the sulfide removal within the rising main reactor using
in-situ electrochemical generation of iron and alkalinity from sewage
The proof of concept about in-situ electrochemical generation of iron and alkalinity for sulfide
control has been demonstrated [7]. Sulfide removal efficiency was above 98% resulting in very low
sulfide concentrations (i.e. < 0.5 mg L-1) at the Fe/S molar ratio of 1.0. The stability of the process
during the course of the previous experiment with the average anode potential and cell voltage of 0.04±0.23 V (vs. NHE) and 2.90±0.54 V, respectively [7].
The main goal of this research is to use sacrificial iron electrodes both as anode and as cathode to
produce concentrated iron solution from sewage (without sulfide addition). The produced iron
solution (sewage based) is then fed to the rising main reactor in order to precipitate sulfides present
in the rising main reactor as FeS. Relevant reaction mechanisms for the anode and cathode are
described below.
27
Anodic reaction
Fe → Fe3+ + 3e−
E 0 = +0.04V
(25)
E 0 = −0.83V
(26)
Cathodic reaction
2H2 O + 2e− → 2OH − + H2
SEWAGE EFFLUENT
pH
e-
Iron
Fe
Cathode
3OH- + 1.5H2
Anode
Iron
Fe3+ + 3e-
3H2O
CONTINUOUS SEWAGE FLOW
(Sulfide addition)
Figure 14: Simplified schematic overview of the in-situ electrochemical iron and hydroxide production from sewage (from
research-plan made by Hui-Wen Lin (tutor in Australia))
2.4.2 Objective 2: Determination of the impact of iron ions on biofilm response in terms of
sulfide and methane production rate as well as the accumulated concentrations
(sulfide and methane) within the reactor
Understanding the effect of Fe3+/Fe2+ addition on sulfate reduction and methane production in the
anaerobic sewer system is important because it would allow for the optimisation of chemical dosing
for sulfide removal and control of methane gas emissions. A decrease of the sulfate reduction and
methane production rate is expected. Despite the very low solubility of iron sulfide, complete
control of dissolved sulfide is difficult and iron salts must be added in excess to obtain adequate
control. Getting dissolved sulfide concentrations lower than 0.2 mg S L-1 in practice is quite difficult,
but possible accumulation of sulfides will possibly lead to better precipitation and eventually a
stable/controlled sulfide concentration.
28
2.4.3 Objective 3: Determination of the change in wastewater characterisation such as pH
and phosphate concentrations of the sewage
Because of the production of alkalinity an increase of the pH is expected. This will result in a change
in S-fractions leading to less H2S in the gas phase and an improved precipitation process. When
sulfides are not present yet in the wastewater the iron ions form a complex with the phosphates and
hydroxides present in the wastewater and thus a decrease of the phosphate concentration is
expected. Relevant reaction mechanisms in the mimicking wet-well are described below.
Fe3+ + PO43- → FePO4
(27)
Fe3+ + 3OH- → Fe(OH)3
(28)
Because of the reduction of sulfate to sulfide and subsequently the formation of the iron sulfide
precipitate an increase of the phosphate concentration is expected after dosage to the reactors.
Relevant reaction mechanisms in the rising main reactor are described below.
FePO4 + HS- → FeS(S) + PO43-
(29)
2Fe3+ + HS- → 2Fe2+ + H+ + S0
(30)
Fe2+ + HS- → FeS(S) + H+
(31)
2.4.4 Objective 4: Determination of the impact of iron ions on biofilm response in terms of
inhibition effects of ferric addition on SRB and methanogenic archaea after the
observation of the decrease in sulfate reduction and methane production activity
The deposition of iron sulfide on the sewer biofilm leads to a reduced access to reactants (sulfate,
VFA and organic matter) in the vicinity of the cells. In addition a deactivation of enzymes of the
microorganisms by reaction with their functional groups, denaturation of the proteins and
competition with essential cations is possible. Deposition on the biofilm layer should result in
inhibition effects for SRB and methanogenic archaea. It is assumed that the inhibition of the
methanogenic activity will be lower than the inhibition of the sulfate reducing activity because of
their position in the biofilm layer. It is assumed that inhibition of the SRB will proceed for two to
three weeks after we stop dosing. A lack of recovery of methanogenic activity is expected because of
the slow growth rate and their inability to compete for substrates with SRB in the outer layer of
biofilms.
29
30
3 Materials and methods
3.1 Experimental set-up
Three laboratory scale rising main sewer reactors (A, B and C) with an internal volume of 0.75 L are
operated at the Advanced Water Management Centre (AWMC) as dosing and control reactors
(Figure 15). To provide additional biofilm growth area, each reactor is equipped with 10 plastic
biofilm carriers and 8 pieces of carbon cloth (size: 4.5 cm x 1.5 cm).
PLC control
Sewage + Fe3+
Raw sewage
(3)
pH
A
Fe3+
Fe3+
pH
Fe3+
Fe3+
Electrochemical
ferric dosing
Electrochemical cell
Mimicking
(Ferric production)
wet-well A
PLC control
Sewage + Fe3+
Raw sewage
(3)
(1)
Heating Unit
pH
PLC control
PLC control
Fe3+
(2)
Fe3+
Fe3+
3+
Sewage Tank
(4 oC)
Fe
3+
Fe
Fe3+
(4)
B
Fe3+
Fe3+
Mimicking
wet-well B
Concentrated ferric
chloride solution
(3)
Conventional ferric
chloride dosing
Raw sewage
PLC control
C
Control reactor
(No dosing)
Figure 15: Schematic overview of experimental setup and process (from research-plan made by Hui-Wen Lin (tutor in
Australia))
3.2 Chemical analyses and measurements
Chemical analysis and measurements are done related to the parameters that have to be measured
at experimental stage: sulfate reduction rate, sulfide production rate/sulfide accumulation, methane
production rate, phosphate concentration, iron concentration (both mimicking wet-well and the
dosing reactor) and sulfide removal efficiency.
31
3.2.1 Ion Chromatography (IC) for the analysis of sulfur species
3.2.1.1
Sulfide anti-oxidant buffer preparation (SAOB)
The sulfide anti-oxidant buffer was made up as follows: 0.8 g NaOH and 0.7 g ascorbic acid were
dissolved in a 250 mL volumetric flask with the sparged and filtered distilled, deionised water.
Sparging was done with helium gas and oxygen was excluded from the glass vials during filling of the
caps. The vials are wrapped into aluminium foil to be stored at 4°C for a maximum of 4 days. After
opening one of the vials to use SAOB solution, the rest of the content was discarded.
3.2.1.2
Sample preparation
Just before the sampling a glass bead and 0.5 mL of the SAOB is added to each IC vial. The samples
are drawn into a 5 mL plastic syringe avoiding any contact with air under any circumstances. A sterile
Millipore PES 0.22 µm express filter with a 200 µL plastic tip attached is then pushed onto the
syringe. The first two millilitres of sample going through the filter are discarded. The plastic tip is
immersed and the sample is slowly dispensed into the vial, making sure the vial is filled right to the
top. However care must be taken that the liquid does not overflow. Immediately after that the vial
has to be capped and vigorously shaken. The glass bead ensures that the liquid is sufficiently mixed.
3.2.1.3
Instrumentation of Dionex ICS-2000
A compact Dionex ICS-2000 ion chromatograph with an AD25 absorbance (230 nm) and a DS6
heated conductivity detector (35oC) were used in series. Preceding the conductivity detector a
Dionex ASRS-ULTRA II 4 mm suppressor (131 mA) was attached. The samples were injected with a
Dionex AS50 autosampler. The data processing was done with the Dionex Chromelon software. A
potassium hydroxide gradient was applied with the Dionex automatic eluent generator using an
EluGen cartridge (EGC II KOH). The gradient started at 12 mM KOH, was ramped up in 5 minutes to
34 mM where it was kept for 3 minutes, then in one minute it was ramped up from 34 to 52 mM and
kept at that concentration for another 5.5 minutes. The data acquisition time is 14.5 minutes and
the total analysis time 19.5 minutes. The eluent was degassed with a Dionex ICS-2000 degasser. The
injection volume was 25 μL and the flow rate 1mL/min. The separation was achieved with a Dionex
IonPac AG18 (4x50 mm) guard and an IonPac AS18 (4x250 mm) separating column. Both columns
were heated to 35oC.
32
Table 5: Sulfur species occurring in aqueous media. The most common species are printed in bold. (modified from [74])
SULFUR SPECIES (Y)
SXO62- X >= 3
S2O82S2O72SO42S2O62S2O52SO32S2O42S4O62MMX+(S2O3)Y(MX-2Y)
S2O32S0 & S8
CH3SXCH3
RSH
SCNSX2- X >= 2
HS-
NAME
Polythionates
Peroxodisulfate
Disulfate
Sulfate
Dithionate
Disulfite
Sulfite
Dithionite
Tetrathionate
Metal
thiosulfate
complexes
Thiosulfate
Elemental sulfur
Dimethylpoly-sulfide
(DMPS)
Sulfhydryl thiols
Thiocynate
Polysulfides
Sulfide,
hydrogen
sulfide
OXIDATION
STATE
0, V
VII
VI
VI
V
IV
IV
III
II1/2
II
II
0
0, I
0
0
0, III
PK
0, 0.9
1.98, -3
OXIDATION
PRODUCTS
SO42SO42Very stable
1.89, 7.21
0.35, 2.45
SO32-, SO42SO42S2O32-, SO32-, SO42SO42-
0.6, 1.72
SO42-
-1.8
6.99, 12.9
S0, S2O32S2O32-, SO32-, SO42-
As previously mentioned, sulfide is one of the species that is rapidly oxidised by oxygen/air,
especially in the presence of heavy metals or when exposed to light with the main product being
thiosulfate. The oxidation of sulfite to sulfate seems to be even faster, esp. in acidic conditions
and/or in the presence of transition metal ions (e.g. Fe3+ and Cu2+). Therefore any sample
preservation method has to either minimise oxidation by eliminating any air or oxygen input or
transform the unstable sulfur species to a stable compound by the addition of chemicals. Sulfide,
sulfate, thiosulfate and elemental sulfur seem to be the major components in aqueous systems, as
indicated in Table 5 [74]. Wastewater from storage vessel is analysed after each wastewater
collection. The dissolved sulfur species are measured using IC: sulfide (S2-), sulfite (SO32-), sulfate
(SO42-) and thiosulfate (S2O32-).
3.2.1.4
S::CAN spectrometer probe (sulfi::lyser)
The S::CAN sensor is used for visualisation of the activity in the reactors. The online measurement of
the HS- concentration gives the opportunity to monitor all changes. The spectrometer probes
measure optical spectra from 190 to 720 nm directly in liquid media. The substances contained in
the medium weaken a light beam emitted by a lamp that moves through the liquid. After contact
with the medium its intensity is measured by a detector over a range of wavelengths specific to the
33
application. It’s a big advantage that a detailed knowledge of the chemical and physical basics of
measuring is not required.
3.2.2 Inductive coupled plasma - optical emission spectrophotometer (ICP-OES)
ICP-OES is an analytical technique used for the detection of trace metals. It uses the inductively
coupled plasma to produce excited atoms and ions that emit electromagnetic radiation at
wavelengths characteristic for a particular element. The intensity of this emission is indicative of the
concentration of the element. Total metal ions (Na, K, Mg, Ca, Al, Fe, Ba, Cd, Cr, Co, Cu, Mn, Mo, Ni,
Zn, Pb, Si, P ,B ,S , Se and As) are measured via unfiltered samples. 10-100% HNO3 is added to the
tube before the sample is taken were the pH should be < 2. The aim is to have an optimal sample
volume of 20 mL.
3.2.3 Total suspended solids (TSS)/ Volatile suspended solids (VSS)
Gravimetric analysis is used for determination of the Total Suspended Solids (TSS) and Volatile
Suspended Solids (VSS) present in a wastewater sample. A standard glass-fibre filter is dried for at
least 2 hours at a temperature of 550°C in order to test the TSS. After weighing, a well-mixed sample
should be filtered through a weighed standard glass-fibre filter. The residue left on the filter is dried
to a constant weight at a temperature between 103°C and 105°C overnight. The increase in weight
of the filter represents the total suspended solids of the sample. Large floating particles or
submerged agglomerates of nonhomogeneous materials from the sample may be excluded in the
total suspended solids measurements if it is determined that their inclusion is not representative of
the entire sample.
After the total suspended solids value is determined a VSS test may be performed. The filter used for
TSS testing is ignited at 550°C for at least 1 hour. The weight lost on ignition of the solids represents
the volatile solids in the sample.
3.2.4 Volatile fatty acids (VFA)
Volatile fatty acids, ethanol, butanol, propanol are measured via the addition of 0.1 mL 10% formic
acid to 0.9 mL filtered sample (add the formic acid first and then the 0.9 mL of sample) into a 2 mL
GC/HPLC vial. The analysis is done with GC/FID (gas chromatograph with flame ionisation detector).
3.2.5 Methane
Vacuumed exetainers are weighed before and after sample addition. These exetainers are put
upside down in the fridge and are taken out of the fridge 24 hours before analysis. The analysis is
done with GC/FID/ECD (GHG gas chromatograph with flame ionisation detector and electron capture
detector).
34
3.3 Experimental procedure
The experimental period was divided into two phases: (1) the baseline phase and (2) the
experimental phase.
3.3.1 Baseline phase
The intention of the baseline phase is to get a stable and comparable sulfide (and potential
methane) production in all reactors before the start of the experimental phase, in order to have
three reactors that can be compared without problems after the experimental phase. All reactors
receive fresh sewage without chemical dosing of iron salts/electrochemically produced iron and
hydroxide ions. Online measurement of the sulfide concentration using one S::CAN sensor
(sulfi::lyser) is done during this phase. IC samples are taken from the reactors for calibration of the
S::CAN sensor (sulfi::lyser) during batch tests. Using these IC samples, the sulfate reduction rate is
determined. The methane production rate is determined based on methane samples (taken in
accordance to the procedure described in the chemical analyses and measurements) during batch
tests. Biofilm samples were taken at this stage for the measurement of biofilm structure (carbon
cloth) and DNA extraction for sequencing (carriers).
In total, there were thirteen weeks of the baseline phase. The reactors started running six weeks
before the start of the baseline phase, where they received fresh wastewater two times per day to
promote the development of anaerobic biofilms on the inner walls and carrier material of the
reactors.
Subsequently, batch tests were carried out a few times a week in order to get a blank result for the
SPR and SRR in all three reactors. Reactors are not disconnected from each other and fresh sewage
is pumped into the reactors to get the whole content of the reactor replaced (to get rid of residual
sulfate/sulfide). The pumping of fresh sewage lasts for 10 minutes to ensure a thorough
replacement of liquid in the reactors with fresh sewage. Wastewater samples from each reactor
were taken at 0, 15, 45 and 60 minutes after feeding because these were reasonable sampling
intervals to represent the biggest change in sulfide production and thus, also the activity.
3.3.2 Experimental phase
At experimental stage, continuous ferric dosing (i.e. electrochemical ferric production) exposes the
sewer biofilms in reactor A. Iron (hot rolled plate carbon steel) is used as anode and cathode
materials in a single-chamber electrochemical system for the production of ferric ions. Fresh sewage
is used as the electrolyte. Experiments are galvanostatically operated at a fixed current density to
generate iron and hydroxide. The sewer biofilms in reactor B are continuously exposed to ferric
35
dosing (i.e. purchased ferric chloride), while reactor C is operated as a control system without ferric
dosing.
During this stage there was sampling from:

The reactors themselves (IC, FIA, CH4 at time=2 min, 15 min, 30 min, 1 hour, 3 hours +
temperature measurements of each reactor)

The wet well (FIA, UV)

Storage vessel (FIA, CH4, IC, pH)
All reactors are exposed to 8 pumping events daily (during baseline and experimental phase), as can
be seen in Table 6 (pump numbers refer to Figure 15).

Pump (1) - each pumping event lasts 2 minutes with 0.75 L sewage delivered to the
mimicking wet-well A and B.

Pump (3) - each pumping event lasts 2 minutes with 0.75 L sewage delivered to the rising
main reactor A, B and C.

Pump (4) - each pumping event lasts 15 seconds with 5.14 mL ferric chloride solution
delivered to the mimicking wet-well B.
Table 6: Pumping timeframe
ST
1 PUMPING
2ND PUMPING
3RD PUMPING
4TH PUMPING
5TH PUMPING
6TH PUMPING
7TH PUMPING
8TH PUMPING
PUMP (1)
PUMP (4)
PUMP (3)
01:00
04:00
07:00
10:00
13:00
16:00
19:00
22:00
01:02
04:02
07:02
10:02
13:02
16:02
19:02
22:02
01:10
04:10
07:10
10:10
13:10
16:10
19:10
22:10
Reaction 22-23 reveals that 2 moles Fe3+ remove 3 moles HS-, so a minimal molar dosing ratio of 2/3
or 0.67 should be applied. Related to the objectives, there was the choice to first dose two weeks at
a Fe/S-ratio of 0.5 and afterwards a second period of two weeks in which there is dosage at a Fe/Sratio of 1.0. The goal is to see if lower concentrations of iron can be used to remove sulfides because
this would mean a lower input of energy, saving up on operational costs while still eliminating
sulfides to a satisfying level.
To make it all a little bit more understandable a small illustration of the set-up during this
experimental phase.
36

In front: The experimental set-up of the electrochemical dosing-technique. On the left there
is the cup with the electrodes which are connected to a potentiostat. On the right the
mimicking wet-well that receives fresh wastewater every three hours. Two pumps make
sure that there is continuous circulation between the electrochemical cell itself and the wetwell. A third pump was added after problems with flooding and has the function to pump all
the water that would overflow back to the wet-well.

In the back: The experimental set-up of the conventional dosing-technique. On the left the
stock-solution and on the right the mimicking wet-well which receives fresh wastewater and
stock-solution every three hours.
Figure 16: Picture of the experimental set-up under the bench (under the reactors)
3.3.2.1
3.3.2.1.1
Dosing ratio 0.5
Electrochemical (EC) iron generation and dosing calculations
The applied current was 4.7 mA in combination with an area of the anode that was under water of
2.24 cm², which leads to a current density of 2.09 mA/cm². To calculate the concentration of total
iron in the wet-well we have to know the volume of the system:
Table 7: Separate volumes of the different parts of the electrochemical (EC) system
EC cup
Wet-well basic rest volume
Pumped to wet-well
102.72 mL
150
mL
900
mL
37
The molar amount of total iron that we produce (taking in account the Fe3+/Fe2+-ratio of 0.69/0.31) is
0.202545 mmole. Calculation leads to a theoretical Fe/S molar dosing ratio of 0.89 for the EC system.
By use of UV, the concentration in the wet-well was measured in triplicate on different days, this led
to the following result:
Table 8: Measured iron species concentrations and Fe/S dosing ratio (n=3)
5.31±1.01 mg Fe2+ L-1
0.80±0.80 mg Fe3+ L-1
6.11±0.23 mg Fe L-1
Fe2+
Fe3+
0.87±0.14
0.13±0.14
6.11 mg total Fe L-1
0.109 mmole total Fe
0.48±0.02 dosing ratio
So, a theoretical dosing ratio of 0.89 eventually leads to a dosing ratio of 0.48±0.02 in practice. There
are several reasons that can lead to this lower dosing ratio. First, the coulombic efficiency of the
process is lower than 100%. Secondly, there was a lot of adsorbance to the tubes and the wet-well.
3.3.2.2
Conventional dosing system
For the conventional dosing system a stock-solution of 1.24 g Fe3+ L-1 was made and dosed during 15
seconds via 5.14 mL. This leads to a theoretical dosing ratio of 0.50.
Table 9: Measured iron species concentrations and Fe/S dosing ratio (n=4)
4.73±0.35 mg Fe2+ L-1
0.43±0.53 mg Fe3+ L-1
5.16±0.32 mg Fe L-1
Fe2+
Fe3+
0.92±0.10
0.08±0.10
5.16 mg total Fe L-1
0.092 mmole total Fe
0.41±0.02 dosing ratio
So, a theoretical dosing ratio of 0.50 eventually leads to a dosing ratio of 0.41±0.02 in practice.
Again, adsorbance to the tubes and the walls of the wet-well could cause this difference in
theoretical and practical dosing ratio.
38
3.3.2.3
3.3.2.3.1
Dosing ratio 1.0
Electrochemical system
For the EC system the applied current was 7.9 mA in combination with an area of the anode that was
under water of 3.92 cm², which leads to a current density of 2.02 mA/cm², which is very similar to
the current density during the first stage of the experimental phase. To calculate the concentration
of total iron in the wet-well the volumes of the system are given:
Table 10: Separate volumes of different parts of the EC system
EC little cup
Wet-well basic rest volume
Pumped to wet-well
126
126
900
mL
mL
mL
The molar amount of total iron that we produce (taking in account the Fe3+/Fe2+-ratio of 0.69/0.31) is
0.340448 mmole. After calculation, this leads to a theoretical Fe/S molar dosing ratio of 1.00 for the
EC system. By use of UV, the concentration in the wet-well was measured in triplicate on different
days. This leads to the following result:
Table 11: Measured iron species concentrations and Fe/S dosing ratio (n=2)
10.90±0.18 mg Fe2+ L-1
2.19±0.05 mg Fe3+ L-1
13.09±0.13 mg Fe L-1
Fe2+
Fe3+
0.83±0.01
0.17±0.01
13.09 mg total Fe L-1
0.234 mmole total Fe
0.69±0.01 dosing ratio
Thus, a theoretical dosing ratio of 1.00 eventually leads to a dosing ratio of 0.69±0.01 in practice.
Again, the coulombic efficiency of the process is lower than 100% and there was a lot of adsorbance
to the tubes and the wet-well, which explains the difference between the theoretical dosing ratio
and the one measured. Although, during the first stage there was a difference of 46% (from 0.89 in
theory to 0.48 in practice) between the theoretical and practical dosing ratio, which is now only 31%.
This could indicate that the influence of adsorption is smaller when dosing at higher ratios.
3.3.2.3.2
Conventional system
For the conventional dosing system a stock-solution of 3.70 g Fe3+ L-1 was made and dosed during 15
seconds via 5.14 mL. This leads to a theoretical dosing ratio of 1.00.
39
Table 12: Measured iron species concentrations and Fe/S dosing ratio (n=2)
8.38±0.09 mg Fe2+ L-1
3.94±0.31 mg Fe3+ L-1
12.32±0.40 mg Fe L-1
Fe2+
Fe3+
0.68±0.01
0.32±0.01
12.32 mg total Fe L-1
0.221 mmole total Fe
0.65±0.02 dosing ratio
So, a theoretical dosing ratio of 1.00 eventually leads to a dosing ratio of 0.65±0.02 in practice.
Again, adsorbance to the tubes and the wall of the wet-well could cause this difference in theoretical
and practical dosing ratio.
40
41
4 Results and discussion
4.1 Electrochemical iron and alkalinity production
The first step for the control of sulfides via the in-situ production of iron and hydroxide ions was to
run a short-term experiment to produce a stable molar ratio of Fe3+/Fe2+ and a set concentration,
using wastewater as electrolyte. This leads to the following coulombic efficiencies and molar ratios
(see Table 13).
Figure 17: Experimental bench set-up (short-term)
Table 13: Iron concentrations, coulombic efficiencies and speciation in 3h experiments (n=3)
Surface area electrodes (cm²)
Current applied (mA)
Current density (mA/cm²)
Time running experiment (min)
Volume cup (L)
Charge Fe2+ (C)
Concentration Fe2+ (mg Fe2+ L-1)
Charge Fe3+ (C)
Concentration Fe3+ (mg Fe3+ L-1)
Total charge (C)
Total concentration (mg Fe L-1)
2+
Molar ratio (Fe /Total Fe)
Molar ratio (Fe3+/Total Fe)
Coulombic efficiency (%)
Run 1
2.52
5.20
2.063
180
0.103
14.11
Run 2
2.52
5.20
2.063
180
0.103
11.57
Run 3
2.52
5.20
2.063
180
0.103
12.34
Average
2.52
5.20
2.063
180
0.103
12.67 (±1.30)
39.82
37.58
32.64
47.19
34.76
37.95
35.74(±3.69)
40.91(±5.44)
70.69
51.69
88.77
58.76
71.28
50.29
76.91(±10.27)
53.58(±4.54)
110.51
0.36
0.64
92.04
121.41
0.27
0.73
104.62
106.04
0.33
0.67
93.12
112.65(±7.91)
0.32(±0.05)
0.68(±0.05)
96.59(±6.97)
42
The results confirm previous work [7]. The molar ratio of Fe2+/Total Fe was 0.31±0.02, the molar
ratio of Fe3+/Total Fe was 0.69±0.02. So the ratio of Fe3+/Fe2+ in the samples from the experiments
without sulfide dosing show that 1/3 of the iron is present as ferrous and 2/3 is present under the
form of ferric ions. Based on these results one can see that, as expected, also Fe3+ is generated at
this anode potential.
4.2 Baseline sewer reactor
4.2.1 Stable and comparable activity
The aim of the baseline phase is to get a stable and comparable sulfide production in all three
reactors before the experimental phase can start to make sure that the obtained results after the
experimental phase are comparable (as previously mentioned in the experimental procedure). There
is also the possibility to compare the methanogenic activity of the three reactors but previous
knowledge and experience showed that this is much more difficult. This is because the acetoclastic
methanogenic activity accounts for about two-thirds of the CH4 emitted and these organisms are
more affected by changes in the pH (compared to other methanogens) that occur in the wastewater
feed [75]. At pH levels away from the optimum the methanogens must expend energy to maintain
homeostasis rather than perform anabolism [76]. Changes of methanogenic activity could have also
been caused by increased free ammonia (NH3) concentrations, which disrupt the proton motive
force and methanogenic homeostasis.
4.2.1.1
Performance of the reactors at week 8
Figure 18 presents the pH and Total Dissolved Sulfide (TDS) during two pumping events for each
reactor in week 8. The pH in reactor 1 decreased from 7.16±0.02 to 7.02±0.05 during a pumping
event of three hours. Reaction 3 (pKa1 = 7.04) indicates that the TDS is present as H2S(aq) and HS- at
this pH. Longer retention of wastewater leads to a lower pH due to formation of protons, so
hydrogen sulfide (H2S) will go into the sewer atmosphere and result in corrosion. The lowering of the
pH corresponded at the same time with an increase of the TDS concentration (due to the reduction
of sulfate) from 4.32±3.86 to 11.26±0.41 mg S L-1. The initial sulfide concentration was relatively
high, which can probably be related to a bad plug-flow that results in a partial replacement of the
wastewater present in the reactor. Attempts to improve the plug-flow were done by increasing the
speed of the pump to 1500 mL in two minutes, but due to unsafe situations there was a reset to 750
mL in two minutes.
For reactor 2 there was a similar result at the start of the pumping event related to the pH
(7.16±0.04), as expected. For reactor 2 the pH decreased to a value of 7.09±0.03. The TDS
43
concentration increased for reactor 2 from 3.18±0.63 to 13.67±1.73 mg S L-1, which is also similar to
reactor 1.
For reactor 3 there was also a similar result at the start of the pumping event related to the pH
(7.23±0.07) and this value decreased to 7.12±0.04. The TDS concentration increased from 6.29±2.39
to 13.43±0.79 mg S L-1. The initial sulfide concentration is slightly higher than for reactor 1 and 2, but
Figure 18 shows that the profiles of the three reactors are similar. Although, reactor 1 has a slightly
lower activity than the other two reactors, but this is attributed to the scarce reproducibility of
biological processes.
Figure 18: Total Dissolved Sulfide (TDS) concentration (mg S L-1) and pH profile for each reactor during week 8.
4.2.1.2
Week 10
Figure 19 presents the pH and Total Dissolved Sulfide (TDS) during two pumping events for each
reactor in week 10. The pH in reactor 1 decreased from 7.48±0.09 to 7.31±0.05 during a pumping
event of three hours. The wastewater collection was done after a long period of dryness and a lot of
soap in the wet-well which result in higher pH values (in comparison to week 8), although, the pH of
the wastewater feed was actually lower (7.35 instead of 7.80 during week 8). Figure 19 shows an
increase of the TDS (due to the reduction of sulfate) from 5.66±0.32 to 11.51±1.30 mg S L-1, which is
similar to the results in week 8. The initial concentration was relatively high, but again, this could be
related to a bad plug-flow situation that results in a partial replacement of the wastewater present
in the reactor.
For reactor 2 (7.36±0.07) there was a similar result at the start of the pumping event related to the
pH, as expected. The pH decreased to a value of 7.25±0.02. The TDS concentration increased for
reactor 2 from 10.28±3.65 to 14.75±0.24 mg S L-1.
For reactor 3 there was a similar result at the start of the pumping event related to the pH
(7.44±0.05). The pH decreased to a value of 7.29±0.07. The TDS concentration increased from
6.33±0.30 to 14.30±1.34 mg S L-1. The initial average concentration is slightly higher than for reactor
1 and 2, but Figure 19 shows that the profiles of the three reactors are similar.
44
Again, reactor 1 had the lowest activity in comparison to the other two reactors. On Figure 19 it can
be seen that the sulfide production rate (SPR) is lower for reactor 1 and it looks like it has not
received its ‘plateau’ after three hours. Sulfidogenic activity can occur in a wide range of pH (5.09.0), but the optimal pH has previously been reported to be in the range of 7.0-8.0. Thus, during the
whole baseline phase there have been optimal pH-conditions which result in stable (and similar)
activity in the three reactors [77].
Figure 19: Total Dissolved Sulfide (TDS) concentration (mg S L-1) and pH profile for each reactor during week 10.
4.2.1.3
Sulfate reduction rate (SRR), sulfide production rate (SPR), methane production rate (MPR)
Table 14 presents the SRR of the three reactors. The SRR values are the average of week 7, 9, 10 and
12 for reactor 1. For reactor 1 and 3 the SRR is the average of week 7, 9 and 12. The SRR, SPR and
MPR were determined using the linear parts of the profiles (first hour data points: 0, 15, 45 and 60
minutes). The SRR of the reactors are similar and indicate that the baseline phase was successful.
Table 14: Sulfate reduction rate (SRR) (mg S L-1/h) (n=4 for reactor 1; n=3 for reactor 1 and 3)
Reactor 1
Reactor 2
Reactor 3
Average SRR (mg S L-1 h-1)
-5.37±1.14
-5.95±1.00
-5.47±0.44
Table 15 presents the SPR of the reactors. This is the average of week 9, 10 and 12. This table gives a
confirmation of the findings in Figure 18 and Figure 19 (where it was already mentioned that the
sulfidogenic activity in reactor 1 is slightly lower). Previous research showed SPR of 5.7±0.5 mg S L-1
h-1, which is in agreement with the average SPR of the three reactors during the baseline phase [13].
A possible explanation for the lower activity in reactor 1 could be a less favourable pH in the first
reactor (which lowers the sulfidogenic activity), but as previously mentioned the pH was similar for
the three reactors during the whole baseline phase. Another explanation could be a small difference
of the flow pattern in the reactor. Fresh feed is pumped into the reactor via a four-way dispenser at
the bottom of the reactor, effluent leaves the reactor via the top. Partial blockage of the dispenser in
the reactor can have a negative influence on the availability of the fresh reactants.
45
Table 15: Sulfide production rate (SPR) (mg S L-1 h-1) (n=3)
Reactor 1
Reactor 2
Reactor 3
Average SPR (mg S L-1 h-1)
4.87±1.92
6.35±2.80
6.61±2.38
Table 16 presents the MPR, which is the average of week 4, 6, 8 and 11. As previously stated it is
difficult to decide if the activity of the three reactors is similar based on the methanogenic activity
because the results vary every week. The average MPR gives a good representation of the separate
weeks during the baseline phase and indicates that the methanogenic activity of reactor 2 and 3 are
similar. The activity of the first reactor is again lower, which corresponds to the SRR and SPR.
Table 16: Methane production rate (MPR) of different weeks and the average MPR (mg CH4 L-1 h-1) (n=4)
Reactor 1
Reactor 2
Reactor 3
Week 4
1.39
2.90
2.33
Week 6
2.10
4.60
4.40
Week 8
1.91
3.89
3.18
Week 11
1.34
1.24
3.45
Average
1.69±0.38
3.16±0.85
3.34±0.85
4.2.2 Batch test
As previously stated, the main goal of the research is to demonstrate the effectiveness of the in-situ
production of iron and hydroxide ions in comparison to the conventional technique were iron salts
are dosed. This is important to remember in relation to the choice of which reactor is the control
reactor, which reactor receives conventional dosing, and which reactor is treated with the
electrochemical dosing technique. Since the first reactor always had the lowest activity and reactor 2
and 3 are very similar, it is logic to dose at reactor 2 and 3 and keep reactor 1 as the control reactor.
So eventually this leads to:

Reactor 1: control reactor

Reactor 2: conventional dosing (FeCl3.6H2O)

Reactor 3: electrochemical production of iron and hydroxide ions
4.3 Experimental phase
4.3.1 Dosing ratio 0.5
The choice of a dosing ratio that is actually lower than the minimal dosing ratio is because of the
interest in the difference between both techniques (related to sulfide removal, inhibition, etc.) and
certainly don’t overdose. Sulfate concentration of the wastewater from the storage vessel was at
15.35±1.22 mg S L-1 in the last three weeks. At the start of the experimental phase concentrations of
46
10.33 mg S L-1 were measured, so a sulfide concentration of 10 mg L-1 in the sewer system to be
treated was proposed.
4.3.1.1
4.3.1.1.1
Results
Long-term performance
For the removal of the proposed concentration of 10 mg S L-1 we need a molar dosing ratio of 0.67.
For the comparison of both systems it would be ideal to have two dosing ratios that are identical.
The experimental measurement of the concentrations in the wet-well revealed that the dosing ratio
of the EC system is 0.48±0.02 and the dosing ratio of the conventional system is 0.41±0.02. It’s clear
that there is a big difference between the theoretical and practical dosing ratio of the EC system.
This difference is smaller for the conventional system. This is mainly related to the set-up, since the
EC system has a lot of tubing in comparison to the conventional method because of the need of
continuous circulation (Figure 16). The difference in dosing ratio between the systems means that
there is also a different removal of sulfides. The EC system can remove 7.24 mg S L-1. The
conventional system can remove 6.11 mg S L-1. Despite the minor difference in dosing ratios
between the two lines it should be possible to compare these two systems.
Previous research stated that there is an increase of the cell voltage due to the formation of a
concretion layer on the electrodes [7]. The passivation layer on the electrodes can be described by a
slimy and sludge like structure filling up the entire spacing, which results in additional resistance and
an unstable cell voltage. I cleaned the electrodes once a week with 1M HCl solution to make sure
that the original cell voltage could be recovered and keep the fluctuations more or less to a
minimum. This can be seen in Figure 20 were the cell voltage increases after cleaning and remains
stable during the following week of the experiment. As previously stated, the current density is kept
low (2 mA cm-²) to avoid a substantial loss of energy input due to ohmic drop in the wastewater.
Another option to keep the cell voltage stable is the use of polarity switching, but the effectiveness
of this technique was already shown during previous research [7].
47
2.5
cell voltage before cleaning
cell voltage after cleaning
cathode potential before cleaning
cathode potential after cleaning
anode potential before cleaning
anode potential after cleaning
2
1.5
Potentia (V)
1
0.5
0
-0.5 0
2000
4000
6000
-1
-1.5
-2
Number of scans
Figure 20: Course of cell voltage, cathode potential and anode potential before and after cleaning. Cleaning with 1M HCl
was done after 6 days of operation
Table 17: Average cell voltage (V), cathode potential (vs. Ag/AgCl) (V) and anode potential (vs. Ag/AgCl) (V) during stable
operation of the electrochemical system.
Cell voltage (V)
Cathode potential (vs. Ag/AgCl) (V)
Anode potential (vs. Ag/AgCl) (V)
4.3.1.1.2
1.76±0.04
-1.49±0.02
-0.38±0.01
Sulfate reduction rate (SRR) and presence of phosphate
Understanding the effect of iron addition on sulfate reduction in the sewer environment is
important because it will allow for the optimisation of chemical dosing for the removal of sulfides.
Figure 21 presents the average SRR of the control reactor, conventional system reactor and EC
system reactor. The control reactor did not receive any dosage of chemicals/electrochemical
produced ions, so the SRR should stay the same. The SRR decreased from 5.37±1.14 to 4.90±0.51 mg
S L-1 h-1, which is only a difference of 0.46 mg S L-1 h-1 and this is perfectly explainable by the scarce
reproducibility of the biological process. One of the factors that is important is the sulfate
concentration present in the feed. During the baseline phase concentrations of 15.35±1.22 mg S L-1
were measured. During experimental phase feed concentrations of 10.33 mg S L-1 were measured,
which can cause a lower activity of the SRB and can explain the small difference. This difference can
also be related to a possible temperature difference in the room where the reactors stood because
there is a doubling of the activity with every 10°C [78]. To take into account this temperature
difference there was a regular measurement of the temperature in the reactor. The temperature
was controlled around 24-25°C in the reactors, via air conditioning, so the temperature was not
responsible for the difference of SRR for the control reactor. The most interesting difference is
present between the FeCl3 system and the EC system. The SRR of the FeCl3 system decreased from
48
5.95±1.00 to 3.30±0.22 mg S L-1 h-1, which is a decrease of 2.65 mg S L-1 h-1 or 45%. The SRR of the EC
system decreased from 5.47±0.44 to 2.38±0.29 mg S L-1 h-1, which is a decrease of 3.09 mg S L-1 h-1 or
57%. Taking into account the decrease of SRR of the control reactor, this still means a difference of
2.18 mg S L-1 h-1 for the FeCl3 system (after completion) and 2.63 mg S L-1 h-1 for the EC system. As
previously mentioned in literature, these inhibition effects on the SRB are related to the deposit of
metal sulfides on the surface of SRB, which cause the inhibition of the activity of these cells [59]. The
insoluble sulfides may reduce access of reactants (sulfate, VFA and organic matter) in vicinity of
bacterial cell to the necessary enzymes, thus reducing the further metabolism of bacteria. Fe3+ could
also deactivate enzymes of microorganisms by reacting with their functional groups, denature
proteins of microorganisms and compete with essential cations utilized by microorganisms. These
processes cause adverse effects on the activities of microorganisms. Next to this also the
surrounding SRB environment is of great importance, because the resistance against the inhibitory
effect is bigger in mixed cultures (as can be found in sewer systems) [61].
Sulphate reduction rate (mg S/L/h)
8.00
7.00
6.00
Control - baseline
Control - experiment
5.00
FeCl3 - baseline
4.00
FeCl3 - experiment
3.00
EC - baseline
EC - experiment
2.00
1.00
0.00
Figure 21: SRR of the control reactor, FeCl3 system and EC system (n=3)
Table 18 presents a summary of the results for sulfur species and chloride concentration (done in
triplicate) of the three systems. It is important to see if the samples are well taken. This may be seen
in the chloride column by making sure that the chloride concentrations of the samples taken at
different points of time for a certain system are similar. The higher average values of the samples
from the FeCl3 system are of course related to the chloride that is present in the dosing medium and
results in approximately 10% increase of the salinity.
Differences in sulfate concentration between the start of a pumping event (time=2 minutes) and
after 1 hour are important. The value after 1 hour is important and not after 3 hours because of the
49
linear regression, which was done using the data points of the first hour. Attention is drawn at
values given in italics. For the control reactor this means a decrease of 4.90 mg S L-1, for the FeCl3
system a decrease of 3.30 mg S L-1 and for the EC system a decrease of 2.38 mg S L-1 (lower decrease
because of the higher inhibition of the SRR). The sulfate will be converted to sulfide, so this should
be seen in the sulfide column. Although, the iron ions will first remove the sulfides which are present
in the wastewater in storage vessel. Thus, 5.93 mg S L-1 will be removed. As previously mentioned,
the EC system and FeCl3 system are able to remove respectively 7.24 mg S L-1 and 6.11 mg S L-1. Thus,
after removal of sulfides present in the storage vessel, the EC system and FeCl3 system can remove
respectively 1.31 mg S L-1 and 0.18 mg S L-1 of the sulfides which are originated from the reduction of
sulfate.
For the FeCl3 system this means that there should be an increase of 3.30 mg S L-1 of sulfides, minus
0.18 mg S L-1 which is removed by iron ions, so eventually an increase of 3.12 mg S L-1. The results
show an increase of 2.12 mg S L-1. For the EC system there should be an increase of 2.38 mg S L-1,
minus 1.31 mg S L-1 which is removed by iron ions, so eventually an increase of 1.07 mg S L-1. The IC
results show an increase of 1.84 mg S L-1. For the control reactor there should be an increase of 4.90
mg S L-1, but when looking at the results this is only 4.16 mg S L-1. Making the balance at each point
of time results in a difference of approximately 1 mg S L-1 for the three systems between the
measurement at 2 minutes and 1 hour. This could be due to the formation of H2S, which maybe
builds up in the reactor or the cavity of the S::CAN sensor (and which is not measured using the ICsamples).
It’s obvious that the estimated concentration of sulfides to treat is important here. The dosing of
iron ions removes all of the sulfides which are already present in the wastewater in the storage
vessel, but the removal of sulfides derived from sulfate is only a very small part for both systems.
This will be addressed by dosing at a dosing ratio of 1.0, where the removal of sulfides related to
sulfate will be higher. Although, the fact that the removal of the EC system is better than the
conventional system is confirmed (as previously stated) because a mixture of both forms of iron has
shown to reach higher removal efficiencies than one form used separately [12].
50
Table 18: Summary of the results sulfur species and chloride concentration (n=3) of the control system, conventional system,
EC system and storage vessel.
Chloride
(ppm)
Sulfide-S
(ppm)
Sulfite-S
(ppm)
Sulfate-S
(ppm)
Thiosulfate-S
(ppm)
CONTROL
2min
15min
30min
1 hour
3 hours
116.05
115.14
114.80
114.70
114.66
7.88
9.50
10.60
12.04
13.36
0.21
0.30
0.32
0.36
0.42
8.45
6.54
5.26
3.55
1.16
0.55
0.58
0.55
0.49
0.39
FeCl3
2 min
15 min
45 min
1 hour
3 hours
129.57
128.74
128.50
128.04
130.11
3.01
3.45
3.97
5.13
6.91
0.20
0.23
0.25
0.29
0.36
6.65
5.46
4.59
3.35
1.02
0.53
0.41
0.37
0.33
0.34
EC
2 min
15 min
45 min
1 hour
3 hours
116.33
115.79
115.20
115.42
116.58
3.23
3.60
4.12
5.07
7.41
0.25
0.30
0.35
0.36
0.40
9.27
8.76
7.95
6.89
4.20
1.02
0.94
0.84
0.67
0.38
-
115.13
5.93
0.12
10.33
0.49
STORAGE
VESSEL
The formation of FeS can be related to the phosphate concentrations present. During the first step
there is the formation of FePO4 (reaction 27) and Fe(OH)3 (reaction 28) in the wet-well. In a real
sewer system with upstream dosage this means that due to the low availability of sulfide, the Fe3+
ions will first precipitate with other anions such as phosphate and hydroxide. However, Fe3+ seems
to become available for sulfide precipitation when sulfide is produced downstream, accompanied by
phosphate release. This means that there is a mixture of FePO4, Fe(OH)3 and FeS present in the wetwell, although the presence of FeS will be limited because of the lower sulfide concentrations. After
dosage to the reactor the sulfides will replace the phosphate and hydroxides to form more FeS,
which liberates the phosphate and hydroxide ions and results in higher measured phosphate
concentrations and higher pH in the EC system. Overall this means that the removal rate of
phosphates is higher at increasing Fe/S ratios because sulfide will precipitate first, and then the left
over iron will react with present phosphates and form FePO4. The trapping/release of the phosphate
ions may be seen when comparing the concentration in the storage vessel and the wet-well, which
means a decrease from 7.51 mg PO4-P L-1 to 4.35 mg PO4-P L-1 (-42%). Measured concentrations after
51
dosage to the reactor show an increase of 4.35 mg PO4-P L-1 to 5.73 mg PO4-P L-1 (after 3 hours)
(+32%), which means an overall removal of phosphate of 24%. A similar trend may be seen for the
FeCl3 system, where a decrease is measured from 7.51 mg PO4-P L-1 to 6.29 mg PO4-P L-1 (-16%),
followed by an increase of 6.29 mg PO4-P L-1 to 7.21 mg PO4-P L-1 (after 3 hours) (+15%), and thus an
overall removal of 4%. In the EC system there is a clear removal of phosphate on top of the sulfide
removal, which is nice advantage but the principle behind this is unclear, although it’s probably
related to the kinetics of the precipitation reactions. The Fe-P-S interactions are quite complex and
for instance organic matter can complicate the chemistry even further.
Table 19: Summary of N and P results of the three systems (n=3)
CONTROL
2 min
1 hour
3 hours
FeCl3
43.07
42.97
43.97
6.84
7.12
7.21
3 hours
42.53
42.83
43.17
4.98
5.38
5.73
-
39.80
42.50
43.90
4.35
6.29
7.51
2 min
1 hour
3 hours
EC
2 min
1 hour
WET WELL EC
WET WELL FeCl3
STORAGE VESSEL
4.3.1.1.3
NH4-N (ppm) PO4-P (ppm)
44.40
7.51
45.23
7.62
45.83
7.57
pH
Figure 22 reveals a difference in pH between the EC and FeCl3 system. The pH of the control was
7.22±0.03, of the conventional dosing system 7.07±0.09 and of the EC system 7.39±0.04. The pH is
on average 0.32 units higher than for the EC system (in comparison to the conventional system). The
fractionation of S-species will go to the right, so less H2S and more HS- will be formed (Figure 2). This
may be linked to the IC-results (Table 18) when making the overall balance (sum of the different Sspecies on a certain moment in time): the difference between the starting point and the
measurement after one hour resulted in a difference of 1.30 mg S L-1 for the FeCl3 system and 0.78
mg L-1 for the EC system. This confirms the shift to less H2S, because the ‘loss’ of S-species is indeed
lower for the EC system. Thus, the higher pH has a double effect: one way hand it results in less H2S
52
(and thus less corrosion), and on the other hand it results in more HS-, which makes it possible to
trap more sulfides via iron ions and thus a higher efficiency.
When looking at the pH profile of the EC system there is also another interesting appearance: the pH
starts at a higher value and decreases during the pumping event. This means that dosage should be
done regularly, to ensure that we do not lose this advantage.
8
7.75
pH
7.5
EC
7.25
FeCl3
7
Control
6.75
6.5
13:09
15:47
18:25
21:04
Time on day 19 (h)
Figure 22: pH profile of the three systems on day 19
4.3.1.1.4
Methane production rate (MPR)
Understanding the effect of iron addition on methane production is important because it will allow
for the optimisation of chemical dosing for the control of methane gas emissions. The MPR during
baseline was similar for the FeCl3 reactor and EC reactor, but the MPR of the first reactor was lower.
After further development it can be seen that the MPR is considerably higher in the control system
(4.40 mg CH4 L-1 h-1) in comparison to the other two systems that received dosage of iron (2.86 and
2.00 mg CH4 L-1 h-1 for the FeCl3 and EC system, respectively). The big increase of methanogenic
activity in comparison to the small decrease in SRR (Figure 21) in the control reactor shows that the
methanogens thrive after a longer period (this can be explained on the basis of their kinetic
properties (Ks and μmax)). The MPR of the FeCl3 system decreased 10%, while the MPR of the EC
system decreased 40%. The SRR decreased respectively 45 and 57%. This is interesting because this
confirms the hypothesis of the inhibition effects on the SRB due to the deposition of metal sulfides
on the surface of SRB, which causes the inhibition of these cells. This can be seen on Figure 6,
because the methanogens typically dominate the deeper biofilm layers, and are thus less affected by
the inhibition effects of the metal sulfides because sulfide is produced by SRB in the inner layer and
will diffuse outward, reacting with Fe3+ in the outer layer of biofilms, thus forming a protection
mechanism for inner layer biofilms. Since sulfide precipitation is sensitive to pH (precipitation
decreases at lower pH), and since the pH value of the inner layers of sewer biofilms is expected to be
53
lower than those in the outer layers and the bulk liquid phase because of the accumulation of VFAs
and CO2 (because of inhibition of SRB and methanogens), this means that FeS may not form in deep
layers of biofilms even if iron ions managed to diffuse into deep layers [79]. Although, this protection
layer results in low supply of reactants, which causes the inhibition. The reduction of the MPR has
also an additional positive effect in relation to the production of GHG and consumption of COD. The
application of the in-situ technology is in this respect more beneficial because of the bigger
reduction of MPR. This will discussed further in the economic analysis of the technology.
Methane production rate (mg CH4/L/h)
6
5
Control baseline
4
3
2
Control experimental
FeCl3 baseline
FeCl3 experimental
EC baseline
EC experimental
1
0
Figure 23: Comparison of the methane production rate (MPR) during baseline and experimental phase (n=3)
4.3.2 Dosing ratio 1.0
During the second part of the experimental phase there was the choice for a dosing ratio of 1.0. This
means a dosage higher than the minimal dosing ratio to make sure that all sulfides are removed and
to see how efficient both technologies are. The increase of the pH is also of interest to see if
production of hydroxide ions is beneficial. Based on the sulfate and sulfide concentration in the
wastewater from the storage vessel (respectively 13.45 and 1.76 mg S L-1) and the removal of
sulfides which were related to the conversion of sulfate in the reactor I chose to estimate the sulfide
concentration to be treated at 15 mg S L-1 (instead of 10 mg S L-1 during the first stage of the
experimental phase with dosing ratio 0.5). The increase of the concentration is logic because of the
combination of higher temperatures and very low rainfall in Brisbane, which results in higher
concentrations in the collected wastewater. This was approximately confirmed by a measured
concentration of the collected wastewater of 13.45 mg S L-1.
54
4.3.2.1
4.3.2.1.1
Results
Long-term performance
For the removal of the proposed concentration of 15 mg S L-1 we need a molar dosing ratio of 0.67.
For the comparison of both systems it would be ideal to have two dosing ratios that are identical.
The experimental measurement of the concentrations in the wet-well revealed that the dosing ratio
of the EC system is 0.69±0.01 and the dosing ratio of the conventional system is 0.65±0.02. During
the first stage of the experimental phase there was a bigger difference in dosing ratios between both
systems (difference was 0.07 between both systems). As previously mentioned, the difference in
dosing ratios between the systems leads to a different removal of sulfides. The EC system can
remove 15.50 mg S L-1. The conventional system can remove 14.59 mg S L-1. The minor difference in
dosing ratios between the two lines is almost inevitable (maybe a perfectly equal set-up for the two
systems with continuous circulation in the conventional system would result in even better results
due to similar adsorption). During the first stage there was the choice to get the practical dosing
ratios of both systems as similar as possible. During this second stage the goal is to make the
theoretical dosing rate both 1.0, which is actually a more logic approach because of better
comparability afterwards.
Again, it was important to make sure that the operation of the system was stable to make sure that
the dosage was constant at all time. The stability of the system can be seen when looking at the
anode potential, cathode potential and cell voltage on Figure 24. Again, the electrodes were cleaned
with 1M HCl after 6 days of operation and not the use of polarity switching (despite the build-up of
dirt in between/on the electrodes (see Figure 25)) because of the stable cell voltage and anode
potential. After cleaning, there was a relatively small increase of the cell voltage but it remained
stable afterwards.
55
cell voltage before cleaning
cell voltage after cleaning
cathode potential before cleaning
cathode potential after cleaning
anode potential before cleaning
anode potential after cleaning
2.5
2
Potential (V)
1.5
1
0.5
0
-0.5
0
500
1000
1500
2000
2500
3000
3500
-1
-1.5
-2
Number of scans
Figure 24: Course of cell voltage, cathode potential and anode potential before and after cleaning. Cleaning with 1M HCl
was done after 6 days of operation
Table 20: Average cell voltage (V), cathode potential (vs. Ag/AgCl) (V) and anode potential (vs. Ag/AgCl) (V) during stable
operation of the electrochemical system.
Cell voltage (V)
Cathode potential vs. SHE (V)
Anode potential vs. SHE (V)
2.03±0.06
-1.46±0.03
-0.48±0.02
Figure 25: Build-up of dirt in-between/on the electrodes after 6 days of operation
4.3.2.1.2
Sulfate reduction rate (SRR) and presence of phosphate
Inhibition of the SRB is not an instantaneous process. It takes approximately three days for the
maximum inhibition on the sulfate reducing activity to be induced. This is why it was important to
wait long enough after the start of dosage to assure us of correct results. On the other hand, the
termination of Fe3+ dosage does not lead to immediate elimination of the inhibitory effects. The
56
recovery can be described by an exponential rise function with a time constant of 0.05 days. When
Fe3+ injection was stopped, the SRR started to recover and reached 90% of the previous level after
about three weeks [79]. There was a period of two to three weeks without dosage of
chemicals/electrochemically produced iron between the end of the first experimental stage and the
start of the second experimental stage. This period should be enough to let the SRB recover, taking
into account that [79] used a dosing ratio of about 1.0, which is higher in comparison to the ones
achieved during the first experimental stage (0.48 and 0.41 for the EC and conventional system,
respectively).
The control reactor did not receive any dosage, so the SRR should be almost the same as during the
baseline phase. The SRR increased from 5.37±1.14 to 8.17±0.89 mg S L-1 h-1. A possible reason is the
change in wastewater characteristics. Ismail et al. (2014) [80] showed that when the pH increases
from 5.5 to 7.0, corrosion rate starts to decrease and is followed by a peak. Afterwards the corrosion
rate starts to increase again with a maximum rate at pH 9.5 (Figure 27). During the baseline phase a
pH of 7.54±0.23 was measured of the wastewater in the storage vessel, while this was 7.21 during
the second part of the experimental phase. Thus, this would result in higher SRR values for the three
reactors. Since the control reactor does not receive any dosage, this confirms the higher SRR in
comparison to the SRR during the baseline phase. The SRR of the conventional system and the EC
system decreased. Again, the most important difference is the difference between the conventional
and EC system. The SRR of the FeCl3 system decreased from 5.95±1.00 to 5.64±0.12 mg S L-1 h-1,
which is a decrease of 0.31 mg S L-1 h-1 or 5%. The SRR of the EC system decreased from 5.47±0.44 to
3.97±0.33 mg S L-1 h-1, which is a decrease of 1.50 mg S L-1 h-1 or 27%.
10.00
Sulphate reduction rate (mg S/L/h)
9.00
8.00
7.00
6.00
5.00
4.00
3.00
Control - baseline
Control - experiment
FeCl3 - baseline
FeCl3 - experiment
EC - baseline
EC - experiment
2.00
1.00
0.00
Figure 26: SRR of the control reactor, FeCl3 system and EC system (n=2)
57
Figure 27: Effect of pH on corrosion rate, related to SRB activity (from [80])
Table 21 presents a summary of the results for sulfur species and chloride concentrations of the
three systems. Attention is drawn at values given in italics. Differences of the sulfate concentration
between the start of a pumping event (2 minutes) and after 1 hour are important. For the control
reactor this means a decrease of 8.17 mg S L-1, for the FeCl3 system a decrease of 5.64 mg S L-1 and
for the EC system a decrease of 3.98 mg S L-1 (lower decrease because of the higher inhibition of the
SRR). Iron ions will remove the sulfides which were already present in the wastewater in the storage
vessel first, so 1.76 mg S L-1 will be removed. The EC system and FeCl3 system were able to remove
respectively 15.50 mg S L-1 and 14.59 mg S L-1, thus the EC system and FeCl3 system can remove
respectively 13.74 mg S L-1 and 12.83 mg S L-1 of the sulfides which are originated from the reduction
of sulfate. For the FeCl3 system this means that there should be an increase of 5.64 mg S L-1 of
sulfides, minus 12.83 mg S L-1 which is removed by iron ions, so eventually there should be no
sulfides left in the reactor because of the over dosage. The results show an increase of 0.28 mg S L-1,
so the system removed almost all of the sulfides related to the reduction of sulfate. After 3 hours a
final concentration of 0.55 mg S L-1 was measured, so over dosage removed most of the sulfides in
the reactor. For the EC system there should be an increase of 3.98 mg S L-1, minus 13.74 mg S L-1
which is removed by iron ions, so eventually all of the sulfides should be removed. The results show
an increase of 0.27 mg S L-1, so most of the sulfides are removed. After 3 hours a final concentration
of 0.76 mg S L-1 was measured, which is only a small accumulation of sulfides after each pumping
event. The accumulation of the sulfide concentration in both systems is not really a problem since
higher concentrations should be removed more easily. Thus, the EC system functions as good as the
conventional system which is used in practice nowadays. The removal of sulfides was also visible
during operation of the systems because of the formation of FeS which results in a black deposit
layer on the transparent effluent tubes (replaced weekly).
58
When making the overall S-balance for each system there is again loss of S. This loss is actually
bigger in comparison to the loss during dosing ratio of 0.5, which may be explained by the lower pH
of the wastewater feed and thus more formation of H2S.
Figure 28: Effluent tubes with black deposit layer on the inside due to FeS formation. Left: EC system; middle: FeCl3 system;
right: control reactor.
Table 21: Summary of the results for sulfur species and chloride concentrations (n=2) of the control system, FeCl3 system, EC
system and storage vessel.
Chloride
(ppm)
Sulfide-S
(ppm)
Sulfite-S
(ppm)
Sulfate-S
(ppm)
Thiosulfate-S
(ppm)
CONTROL
2 min
15 min
30 min
1 hour
3 hours
102.60
100.77
100.62
101.07
102.02
4.90
7.26
8.32
10.46
11.62
0.12
0.15
0.21
0.27
0.41
11.39
8.56
6.65
3.22
0.14
0.36
0.37
0.38
0.38
0.30
FeCl3
2 min
15 min
30 min
1 hour
3 hours
145.17
144.20
145.10
146.20
146.42
0.40
0.53
0.75
0.68
0.55
0.06
0.07
0.09
0.08
0.07
7.22
5.31
3.92
1.58
0.12
0.10
0.03
0.04
0.06
0.02
EC
2 min
15 min
30 min
1 hour
3 hours
102.13
101.69
101.03
101.98
101.02
0.54
0.63
0.82
0.81
0.76
0.07
0.08
0.11
0.10
0.10
8.22
6.84
5.70
4.24
1.64
0.25
0.21
0.12
0.14
0.06
-
101.13
1.76
0.05
13.45
0.00
STORAGE
VESSEL
59
The formation of FeS can again be related to the phosphate concentrations present. As mentioned
earlier, a mixture of FePO4, Fe(OH)3 and FeS is present in the wet-well. After dosage to the reactor
the sulfides will replace the phosphate and hydroxides to form more FeS, which liberates the
phosphate and hydroxide ions and results in higher measured phosphate concentrations and higher
pH in the EC system. The overall trapping/release phenomena of the phosphate ions may be seen
when comparing the concentrations in the storage vessel and in the reactor after 3 hours. For the EC
system this means an overall decrease from 13.90 mg PO4-P L-1 to 10.20 mg PO4-P L-1 (-27%). A lower
phosphate removal for the FeCl3 system was detected with a decrease from 13.90 mg PO4-P L-1 to
12.70 mg PO4-P L-1 (-9%).
Table 22: Summary of N and P concentrations of the three systems (n=2)
CONTROL
NH4-N
50.60
53.75
54.30
PO4-P
13.50
13.65
13.80
49.80
53.10
53.70
8.09
11.10
12.70
3 hours
45.90
47.60
47.90
8.58
9.49
10.20
-
50
13.90
2 min
1 hour
3 hours
FeCl3
2 min
1 hour
3 hours
EC
2 min
1 hour
STORAGE VESSEL
4.3.2.1.3
pH
Figure 29 reveals a difference in pH between the EC and FeCl3 system. The pH of the control was
7.02±0.05, of the conventional dosing system 6.94±0.10 and of the EC system 7.09±0.07. The pH
values are lower for the three systems, which can be explained (as previously stated) by the lower
pH of the collected wastewater. The pH in the EC system is on average 0.15 units higher than for the
conventional system, which is actually a smaller difference than during the stage with a dosing ratio
of 0.5.
The average pH values can be related to the effects on the SRR on Figure 26 because the pH of the
control reactor is the closest to 7.0, which results in the highest SRB activity. Since the pH of the EC
system has the biggest difference with pH 7.0 this results in a bigger decrease of the SRB activity. A
60
higher pH in the entire sewer pipe results in a lower level of total dissolved sulfide and almost no
production of methane. Furthermore, there will be an increase in both total COD and VFA
concentrations. The higher COD concentration is expected because almost no COD was used for
methane production. In addition, a lower sulfide production rate requires a smaller amount of COD.
The combination of a lowered fermentation rate at elevated pH, the absence of methane production
and also reduced consumption of VFAs for sulfate reduction leads to an increased VFA
concentration. Next to the bigger decrease of SRB activity, there is also the advantage of
fractionation of S-species (Figure 2) which results in less H2S in the sewer atmosphere (less
pH
corrosion) and more efficient precipitation of sulfides in the sewage.
7.5
7.4
7.3
7.2
7.1
7
6.9
6.8
6.7
6.6
6.5
13:19:12
pH EC
pH FeCl3
pH control
15:57:36
18:36:00
21:14:24
Time on day 4 (h)
Figure 29: pH profile of the three systems on day 4
4.3.2.1.4
Methanogenic production rate
As mentioned earlier, the inhibition is not an instantaneous process. It takes approximately 7 days
for the maximum inhibition on the methanogenic activity to be induced. It’s clear that there is a
difference between SRB and methanogens with a view to start of inhibition. This can be explained by
a difference in time constants (0.46 for SRB, 0.31 for methanogens) in exponential decay functions,
which are related to the different spatial locations of SRB and methanogens in the sewer biofilm
layer [79]. As mentioned earlier, the second period of dosage was done after a period of 3 weeks
without dosage. Previous studies showed that methanogenic activity did not show any recovery in 5
weeks but these were done with a dosing ratio of about 1.0 (so higher in comparison to 0.48 and
0.41 of the EC and conventional system, respectively) [79]. Methane samples and methane
production rates were calculated and added to the appendix, although these results are probably
not reliable.
The previously discussed hypotheses explains the longer transient period for methanogens in
comparison to SRB because they are less abundant in the outer layer of sewer biofilms. When Fe3+
addition is terminated, the FeS containing biofilm layer will be broken off and replaced by new
61
biomass. Thus, slow growth rate and their inability to compete for substrates with SRB results in the
lack of recovery of methanogenic activity [49].
4.4 Economic analysis
To be able to use the novel electrochemical technology on large scale for the treatment of sewer
corrosion means that it has to be competitive on economic scale. The overall costs of the novel
technology are the cost of current, the cost of iron plates (sacrificial electrode material) and the cost
of the development and optimisation of the dosing unit Best scenario for the implementation of this
novel technology would be were the electrochemical treatment is able to compete or even improve
on a cost-effective level in comparison to conventional chemical dosing with FeCl3.
The sulfide concentrations present in sewer systems form one of the main factors that have an
influence on the operational cost. Padival et al. (1995) [12] mentioned concentrations of 6.4 mg S L-1
upstream, where a dosing ratio of 2.5/1 with FeCl3 led to 97% average elimination of sulfide, which
eventually led to a cost of US$8.40 or AU$10.8 kg S-1. The range in cost is US$4.32-8.40 or AU$5.5510.8 kg S-1, and these are dependent on sulfide removal, dosage of chemicals and current prices of
the chemicals. Despite the very low solubility of iron sulfide, complete control of dissolved sulfide is
difficult and iron salts must be added in excess to obtain adequate control. Getting dissolved sulfide
concentrations lower than 0.2 mg S L-1 in practice is quite difficult [12]. Average concentrations
measured in Australia in the collected wastewater were around 10-15 mg S L-1. During the
experimental phase an average sulfide concentration of 12.49±1.23 mg S L-1 was measured, were the
removal efficiency of both technologies were 95.60% for the FeCl3 technology and 93.92% for the EC
technology. These levels of sulfide to treat are rather high in practice and previous research
mentioned values of 2-6 mg S L-1 over a 40 km sewer network, hence the use of 6.4 mg S L-1 in the
economic analysis [12].
A minimal dosing ratio of 0.67 should result in a total removal of the present sulfide. Previous
research [7] showed that a Fe/S ratio of 0.67 results in an efficiency of 97.20% (second order
polynomial: y=-92.589x² + 159.29x + 32.043, R²=0.9781). This was almost confirmed during the
experimental phase were a dosing ratio of 0.69 resulted in 93.92% removal efficiency (second order
polynomial: 97.87%). Taking into account an influent concentration of 6.4 mg S L-1, a dosing ratio of
0.67 would lead to an iron concentration of 7.45 mg F L-1 to effectively control the present sulfides.
In practical terms this makes up for 7.45 kg Fe ML-1 of sewage treated. The iron plates cannot be
dissolved completely, so it is assumed that only 60% of the iron plates can be used in practice before
replacement. This means that 12.41 kg Fe ML-1 is needed for control. Taking into account the world
price for hot rolled plate carbon steel (US$436 or AU$610.4 /tonne in January 2016) this means a
62
cost of US$5.41 or AU$7.58 ML-1. It has to be mentioned that these world prices can vary a lot, for
instance from US$617 or AU$863.8 tonne-1 in February 2015 to US$425 or AU$595 in December
2015. Overall a downward trend of the price is visible, which makes the EC technique even more
interesting.
A second important cost that has to be considered for this technology is the cost of energy. During
the second experimental stage a fixed current of 7.9 mA was applied, which resulted in a recorded
cell voltage of 2 V. In experiments 0.75 L of sewage was treated in 3 hours, but since the sulfide
concentration was approximately 3 times as high as the one proposed during this cost analysis, it
should be possible to treat 1 L of sewage in 1 hour. Use of the Faraday law results in an energy use
of 22.10 kWh ML-1. The energy price is on average US$0.14 or AU$0.20 kWh-1, thus an energy cost of
US$3.09 or AU$4.33 ML-1. This sums up to a total cost of US$8.51 or AU$11.91 ML-1 (includes the
cost for iron plates and energy).
Padival et al. (1995) [12] mentioned a cost of US$8.40 or AU$10.80 kg S-1 for the conventional dosing
system, so since these costs are in the same order it should be possible to compete with the
conventional technologies for application on large scale. Ganigue et al. (2011) [51] mentioned costs
of AU$10.90-170.6 ML-1 for dosing with iron salts, so comparing these prices with the cost of the
electrochemical technology clearly indicates the potential of this new technology. Since sulfide
concentration and flows have a dynamic behaviour over a day, a continuous on-line control is
necessary. Frequent sensor calibration and maintenance leads to the need for operators, which will
come with a huge economic impact. In addition, the current high price of dissolved sulfide sensors
makes it economically non-viable. An alternative could be dosage on historical data, but this is
obviously not ideal. This indicates the key benefit of the electrochemical system because the
maintenance is (in theory) extremely low.
A third important cost is the capital cost of the pilot plant. Due to confidentiality restrictions it is not
possible to incorporate plans of the unit in this thesis. Outer dimensions of the container are
approximately 2.5 x 2.5 x 6.06 m. Inner dimensions of the actual electrochemical system are given by
1.04 x 0.825 (water level limit) x 2.25 m. The basic capital cost per m² of the entire unit can be
estimated at US$5800 or AU$8120 m-2. Thus, with an area of 15.15 m² this result in a capital cost of
US$87870 or AU$123018.
The production and emission of methane as a GHG from many natural and anthropogenic anaerobic
systems (rice paddies, animals, wastewater and sludge treatment …) has been extensively studied.
However, methane from sewer systems appears to have been largely overlooked. Data presented by
De Haas et al. (2004) [81] showed that a typical biological nutrient removal (BNR) WWTP for 100.000
63
person equivalents may be expected to produce around 4300-7400 tCO2-e year-1 of GHG emissions.
When all dissolved methane from the sewer is discharged to the atmosphere, then the mass of
methane released to the atmosphere will be in the range of 40-250 tCH4 year-1. At a global warming
potential of roughly 21 times (relative to CO2), the released methane will contribute approximately
900-5300 tCO2-e year-1. This means an additional GHG contribution of 12-72% (in comparison to the
worst case scenario of 7400 tCO2-e year-1) from sewage methane over and above that from the
WWTP itself. Experimental results showed a decrease of the methane production rate of the FeCl3
system with 10%, while the MPR of the EC system decreased with 40% (using a dosing ratio lower
than the minimal dosing ratio). Taking into account the costs to reduce emissions during modern
climate change this could be a major advantage for this novel technology. In addition to this it was
also clear that methanogens do not recover immediately from dosage, so dosage will result in a
permanent decrease of methane production. Best scenario would be to dose upstream, to decrease
the overall methanogenic activity in the sewer system. Overall, this will result in a decrease of the
most important negative effects of methanogenic activity in the sewer system: the significant
greenhouse gas emissions, consumption of the valuable COD required for the downstream biological
nutrient removal facilities, and built-up of potentially explosive gas mixtures in sewer atmosphere
[13].
64
5 Conclusion
In a first step, the aim was to show the efficient removal of sulfides via precipitation due to iron ion
formation at the anode and hydroxide formation at the cathode in a single chamber membrane-free
cell. Overall removal efficiencies of > 93% resulted in effluent sulfide concentrations <1 mg L-1 thanks
to a combined production of Fe3+/Fe2+ leading to a more efficient removal of sulfides. Hydroxide
formation leads to an increase of the pH, which results in a double effect in practice: lower H2S
concentrations in the sewer atmosphere (so less corrosion) due to a shift to a more dissolved ionic
form (HS-), and a more effective FeS precipitation of the sulfides in sewage.
During the second part of this thesis, the focus was mainly on the inhibition effects related to sulfate
reducing activity and methanogenic activity. Deposit of metal sulfides on the surface of SRB and
methanogenic cells reduce access of reactants (sulfate, VFA and organic matter) in vicinity of
bacterial cells to the necessary enzymes, thus reducing the further metabolism of bacteria.
Furthermore, Fe3+ could also deactivate enzymes, denature proteins and compete with essential
cations used by these microorganisms. A decrease of the sulfate reduction rate (SRR) of 57% (molar
dosing ratio of 0.48±0.02) was measured and this clearly states the advantage of this novel
electrochemical technique in comparison to conventional dosing with FeCl3 (SRR decrease of 45%).
Continuous dosing, preferably upstream, will be necessary for a continuous removal of the formed
sulfides. Even during a short breakdown of the system, the continuous inhibition of the SRB and
methanogens will lead to less sulfides being formed during this period and no greenhouse gasses
(GHG) due to no recovery of methanogens in such a short period.
For the overall design and configuration of the cell, there are a few parameters that have to be taken
into account and optimised. A flow of sewage should run in parallel with the electrodes which are
placed in several rows. Furthermore, parameters such as spacing and choosing optimal current
density are definitely crucial for successful operation of the system. Polarity switching can easily be
implemented in practice and is necessary to dissolve both electrodes, not only the anode. The effect
of polarity switching on scaling is not clear yet, further research should create clearness.
Finally, an economic analysis showed that the electrochemical technology can compete on large
scale with the conventional dosing strategies. A total cost of US$8.51 or AU$11.91 ML-1, merely
including the cost for iron plates and energy, is in the same order as a cost of US$8.40 or AU$10.80
kg S-1 for the conventional dosing system. In addition, a decrease of the methanogenic activity (-40%
for a molar dosing ratio of 0.48±0.02) leads to a decrease of GHG emissions addressing the
contribution of approximately 900-5300 tCO2-e/year for a WWTP (designed for 100.000 people).
Next to that, this will also lead to a lower chemical oxygen demand (COD) removal, which increases
65
the availability of organic carbon for the downstream biological nutrient removal due to reduced
consumption of organic carbon for methane formation.
66
67
6 Future perspectives
Previously performed research proved the concept
of electrochemical production of iron ions and
alkalinity from sewage for the control of sulfides in
the sewer system. This thesis focussed on the use
of this technology for dosage to a reactor, showing
that stable operation of the system is possible.
Comparison between this novel technology and
conventional dosing via iron salts showed that the removal of sulfides via EC dosing is comparable or
even better, and offers numerous advantages related to increased safety and higher inhibition
effects for SRB and methanogens which eventually leads to less sewer corrosion, lower GHG
emissions and better downstream biological nutrient removal. Future research related to the
influence of a minimal dosing ratio on methanogenic activity should increase knowledge even
further.
Electrochemical production of iron ions and alkalinity based on iron plates in combination with
polarity switching will be used in a pilot plant located at the Gold Coast, QLD, Australia aiming to
demonstrate the effectiveness and usability in a long term operation. Since the flow leaves the pilot
upwards through the plates, the build-up of dirt in between the electrodes should not happen
because of a higher flow upwards between the electrodes. Although, the pilot should confirm these
predictions. If not, periodic cleaning by e.g. rinsing with chemicals should address this problem.
Further optimization through modelling will lead to the best operational parameters for the system.
Important parameters will be the spacing, geometric formation of the electrodes, related to the flow
pattern between the electrodes and in the reactor, and overall form of the unit. In addition, the
frequency of dosing should also be optimised to assure us of the advantage related to an increased
pH. Further research will evoke new questions and matters to investigate, but the opportunity of a
large scale electrochemical unit and possible implementation in the future as a common technology
certainly makes it all worth it. The operational cost and electricity cost will determine the success of
this novel technology. The electricity need for remote areas could be covered by solar power or wind
turbines. Next to this, new developments in for instance the field of bioleaching can result in
recovery of iron necessary for the iron plates via the use of bacteria for oxide ores, which will
become more important as high grade ores are running out in the future. This combining effect can
further increase and extend the already existing environmental value of this technology.
68
69
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74
8 Appendix
Appendix I: experimental parameters for dosing ratio 0.5
These are the experimental parameters which were used during the experimental phase with a
dosing ratio of 0.5. The most important parameters are indicated in red. These parameters include
the length of the electrode that was under water, applied fixed current, obtained current density,
volume that was pumped into the reactor and the obtained dosing ratios for both systems.
given
number
unit
width
length
1.4
1.6
cm
cm
depth
big surface
0.3
2.24
cm
cm²
side surface
bottom surface
Surface area electrode
Current applied
Current I
Current density
minutes running experiment
total coulombs
Faraday constant
z (# e- transferred for Fe3+)
z (# e- transferred for Fe2+)
MM Fe
0.48
0.42
2.24
4.7
0.0047
2.098214286
180
50.76
96485.3
3
2
55.845
cm²
cm²
cm²
mA
A
mA/cm²
min
C
C/mol
# e- transferred
# e- transferred
g/mol
Reactor volume
Fe3+ (based on current applied assuming 100% CE)
0.10272
0.000175364
9.793174712
95.33853886
0.000263045
14.68976207
143.0078083
L
mol
mg
mg/L
mol
mg
mg/L
Fe2+ (based on current applied assuming 100% CE)
Volume
HS- MM
0.75
33.0729
L
g/mol
Proposed HS- concentration in sewer reactor
HS- in reactor
10
7.5
mg/L
mg
0.0075
0.000226772
g
mol
75
EC system - calculations
Ferric MM
Ferric/sulfide ratio
Ferric necessary (theory)
55.845
0.5
0.000113386
0.006332035
6.332034687
1150
1.15
900
0.9
0.000151181
g/mol
0.008442713
8.442712916
8.442712916
0.008442713
0.00970912
0.000173858
0.000202545
0.000188367
0.893166048
9.835753732
g/L
mg/L
mg/L
g/L
g
mol
mol (100% CE)
mol (93% CE)
Fe3+ MM
55.845
g/mol
FeCl3.6H2O MM
270.3
g/mol
Fe3+ necessary (theory)
0.000113386
mol
Volume of FeCl3 we dose through PLC to wet-well
0.006332035
6.332034687
5.136
g
mg
mL
Ratio molar massas
Amount of FeCl3.6H2O for stock solution
0.005136
4.840182648
3.0045
L
g
Concentration stock-solution FeCl3.6H2O
6.009
g/L
Concentration stock-solution Fe3+
1.241482075
g/L
Volume of wastewater to wet-well
Concentration Fe3+ in the wet-well
900
0.9
0.007044524
mL
L
g/L
Mol Fe3+ that we dose
Dosing ratio (theoretical)
7.044523629
0.000114178
0.503491552
mg/L
mol Fe3+
Total volume electrochemical cell
Dosing-volume into wet-well
Concentration Fe3+ that reactor gets
Concentration in the wet-well
(same concentration as reactor gets)
Mol total Fe that we produce (theory)
Mol total Fe that we produce (really)
Dosing ratio (theoretical)
Concentration in the wet-well
mol
g
mg
mL
L
mL
L
mol/L
mg/L
FeCl3 system - calculations
76
Appendix II: experimental parameters for dosing ratio 1.0
These are the experimental parameters which were used during the experimental phase with a
dosing ratio of 1.0. The most important parameters are indicated in red. These parameters include
the length of the electrode that was under water, applied fixed current, obtained current density,
volume that was pumped into the reactor and the obtained dosing ratios for both systems.
given
number
unit
width
length
1.4
2.8
cm
cm
depth
big surface
0.3
3.92
cm
cm²
side surface
bottom surface
Surface area electrode
Current applied
Current I
Current density
minutes running experiment
total coulombs
Faraday constant
z (# e- transferred for Fe3+)
z (# e- transferred for Fe2+)
MM Fe
0.84
0.42
3.92
7.9
0.0079
2.02
180
85.32
96485.3
3
2
55.845
cm²
cm²
cm²
mA
A
mA/cm²
min
C
C/mol
# e- transferred
# e- transferred
g/mol
Reactor volume
Fe3+ (based on current applied assuming 100% CE)
0.10272
0.00029476
16.46086813
160.2498845
0.00044214
24.6913022
240.3748267
L
mol
mg
mg/L
mol
mg
mg/L
Fe2+ (based on current applied assuming 100% CE)
Volume
HS- MM
0.75
33.0729
L
g/mol
Proposed HS- concentration in sewer reactor
HS- in reactor
15
11.25
mg/L
mg
0.01125
0.000340158
g
mol
77
Electrochemical system - calculations
Ferric MM
Ferric/sulfide ratio
Ferric necessary (theory)
55.845
1
0.000340158
0.018996104
18.99610406
1150
1.15
900
0.9
0.000453544
g/mol
0.025328139
25.32813875
25.32813875
0.025328139
0.02912736
0.000521575
0.000340448
0.000316616
1.00
16.53243712
g/L
mg/L
mg/L
g/L
g
mol
mol (100% CE)
mol (93% CE)
Fe3+ MM
55.845
g/mol
FeCl3.6H2O MM
270.3
g/mol
Fe3+ necessary (theory)
0.000340158
mol
Volume of FeCl3 we dose through PLC to wet-well
0.018996104
18.99610406
5.136
g
mg
mL
Ratio molar massas
Amount of FeCl3.6H2O for stock solution
0.005136
4.840182648
8.951
L
g
Concentration stock-solution FeCl3.6H2O
17.902
g/L
Concentration stock-solution Fe3+
3.70
g/L
Volume of wastewater to wet-well
Concentration Fe3+ in the wet-well
900
0.9
0.02098703
mL
L
g/L
Mol Fe3+ that we dose
Dosing ratio (theoretical)
20.98702979
0.000340158
1.000000639
mg/L
mol Fe3+
Total volume electrochemical cell
Dosing-volume into wet-well
Concentration Fe3+ that reactor gets
Concentration in the wet-well
(same concentration as reactor gets)
Mol total Fe that we produce (theory)
Mol total Fe that we produce (really)
Dosing ratio (theoretical)
Concentration in the wet-well
mol
g
mg
mL
L
mL
L
mol/L
mg/L
FeCl3 system – calculations
78
Appendix III: Calibration curve for spectrophotometric iron measurement
Calibration standards are made with a stock solution of 2 g L-1 Fe2+ as FeSO4.7H2O + 10 g L-1 ascorbic
acid adjusted to a pH of 2.7 with 2M H2SO4. Dilutions are made with MilliQ water to obtain the Fe2+
concentrations of 0.1, 0.3, 0.5, 1, 1.5, 2.5, 5 and 10 mg Fe2+ L-1. The absorbance is then measured
UV510 absorbance (-)
using a quartz cuvette.
1.5
y = 0.139x + 0.0109
R² = 0.9997
1
0.5
0
0
2
4
6
Fe2+
Concentration (mg Fe2+ L-1)
0
0.1
0.3
0.5
1
1.5
2.5
5
10
8
10
(mg/L)
UV510
0.0022
0.0224
0.0521
0.0816
0.1479
0.22
0.3617
0.7243
1.3905
79
Appendix IV: Basic economic calculations for the electrochemical treatment of in-situ
iron production.
GENERAL
exchange rate AU dollar/US dollar
exchange rate US dollar/euro
energy cost
World price for hot rolled plate carbon steel
Estimated cost of installation
Daily flow
Minimal molar dosing ratio
Influent sulfide concentration
Influent sulfide load
Iron load
ELECTRICITY
Faraday constant
z (# e- transferred for Fe3+)
cell voltage
Total coulomb
Current (C/sec)
kW
kWh
Electricity cost
IRON FOR PLATES
Amount of iron needed daily
Plates can dissolve
Amount of iron needed daily
Cost of iron
MAINTENANCE
Area of the electrode installation (inside container)
Area of entire EC unit
1.4
1.16
0.14
0.20
436
610.4
5000
5800
8120
1
0.67
6.40
200
133
96485.3
3
2
38594120
446.69
460.51
0.92
22.10
3.09
4.33
US$ kWh-1
AU$ kWh-1
US$ tonne-1
AU$ tonne-1
€ m-2
US$ m-2
AU$ m-2
ML
mg S L-1
mol ML-1
mol ML-1
C/mol
# e- transferred
V
C
A
in case of CE 100%
A
in case of CE 97%
kW
kWh
US$
AU$
7.45 kg Fe daily
60 %
12.41 kg Fe daily
5.41 US$ ML-1
7.58 AU$ ML-1
2.34 m2
15.15 m2
80
Volume of tank (contains electrodes)
Volume of iron (25% of tank volume)
density of iron plates
amount of iron in the tank
Days before replacement
TOTAL COSTS
Total operational cost (energy + plates)
Total capital cost
1.95
0.487266
8000
3898.125
314.08
8.51
11.91
87870
123018
m3
m3
kg m-3
kg Fe
days
US$ ML-1
AU$ ML-1
US$
AU$
81
Appendix V: Methane production rate after dosing ratio 1.0
Compared to the methane production rates during the baseline
Methane production rate (mg
CH4/L)
9
8
7
Control baseline
6
Control experimental
5
FeCl3 baseline
4
FeCl3 experimental
3
EC baseline
2
EC experimental
1
0
Compared to the methane production rates during dosing ratio 0.5
Compared to the methane production rates after the dosing ratio 0.5, because methanogenic
activity does not recover after a period of three weeks without dosage. All methane production rates
are higher in comparison to the methane production rate during dosing ratio of 0.5.
Methane production rates (mg
CH4/L)
8
7
6
5
4
3
2
Control 0.5
Control 1.0
FeCl3 0.5
FeCl3 1.0
EC 0.5
EC 1.0
1
0
82
Appendix VI: Removal efficiency as a function of dosing ratio
120
removal efficiency (%)
100
y = -92.589x2 + 159.29x + 32.043
R² = 0.9781
80
60
40
20
0
0
0.2
0.4
0.6
0.8
1
1.2
Fe/S dosing ratio
Fe/S dosing ratio
0.28
0.52
0.81
1.04
removal efficiency
68.5
92.2
97.9
98.50
83
84
85