DEGREE PROJECT IN CHEMICAL ENGINEERING,
SECOND CYCLE, 30 CREDITS
STOCKHOLM, SWEDEN 2020­2021
Scrubbing of sulfur
dioxide from secondary
process gases in a
copper smelter
KTH Thesis Report
Obiora Okolo
KTH ROYAL INSTITUTE OF TECHNOLOGY
SCHOOL OF ENGINEERING SCIENCES IN CHEMISTRY, BIOTECHNOLOGY AND HEALTH
Author
Obiora Okolo, okolo@kth.se
School of Engineering Sciences in Chemistry, Biotechnology and Health
KTH Royal Institute of Technology
Host companies
Boliden AB
Metso Outotec Sweden AB
Place for Project
Skellefteå, Sweden
Examiner
Kerstin Forsberg, KTH Royal Institute of Technology
Supervisors
Jörgen Johansson, Boliden AB
Olle Gunnarsson, Metso Outotec Sweden AB
Abstract
The processing industry is the largest source of sulfur dioxide (SO2 ) emissions in
Sweden, which includes the non­ferrous metals industry. The copper smelter Boliden
Rönnskär has an environmental permit to emit a maximum of 3500 tonnes of
SO2 /year, a limit that the smelter has been close to in recent years. To reduce the SO2
emissions at Rönnskär, wet scrubbing (together with a bag filter for dust cleaning) has
been proposed as a method for cleaning the SO2 ­bearing, intermittent tapping gases
from the flash furnace. To find the optimal wet scrubbing technique for the purpose,
wet scrubbing techniques based on the following reagents were investigated, evaluated
and compared in this report: caustic soda, soda ash, peroxide, lime and zinc oxide.
Measurements were also done on the secondary hood gases from the converters,
which could make use of the remaining capacity in the scrubber. Further, tests were
conducted on various process waters from other processes at Rönnskär, waters that
could be reused in the scrubber. The scrubber techniques were then evaluated based
on the input data and system requirements using simulation and design software as
well as theoretical calculations.
The results suggested that it is reasonable to clean secondary hood gases in the
scrubber, as they contained approximately 280 tonnes of SO2 /year. This could be
compared to the flash tapping gases that contained approximately 445 tonnes of
SO2 /year. Among the scrubbers, the peroxide scrubber evolved as the most attractive
technique due to its relatively low life cycle cost and due to its suitability with the
leaching plant and the flash cooling tower process water. The other packed tower
techniques, caustic soda and soda ash had the highest life cycle costs, mainly due to
their high reagent costs. The soda ash scrubber, which was the cheaper of the two
sodium­based scrubbers, could still be a suitable alternative due to its simplicity. The
open spray towers had lower life cycle costs than the packed towers. However, the lime
scrubber had several disadvantages that makes it an unsuitable alternative. In turn,
the zinc oxide scrubber is a relatively under­researched and unproven technique, but
should still be studied further as it could be integrated with the zinc smelting process
at Rönnskär. The use of process waters in the scrubber would lead to a net reduction of
process water to the process water treatment plant and would lead to reduced reagent
costs if a stripper is installed to remove the SO2 from the process waters before entering
the scrubber.
Sammanfattning
Processindustrin är den största källan för utsläpp av svaveldioxid (SO2 ) i Sverige, vilket
även inkluderar icke­järnmetallindustrin. Kopparsmältverket Boliden Rönnskär har
ett miljötillstånd som tillåter SO2 ­utsläpp på maximalt 3500 ton/år, en gräns som
Rönnskär har legat nära under de senaste åren. För att minska svaveldioxidutsläppen
på Rönnskär så har våtskrubbning (tillsammans med ett bagfilter för stoftrening)
föreslagits som en metod för att rena de SO2 ­bärande och intermittenta tappgaserna
från flashugnen. För att göra ett optimalt val av våtskrubberteknik för ändamålet,
undersöktes, utvärderas och jämfördes våtskrubbningstekniker baserade på de
följande kemikalierna i denna rapport: lut, soda, peroxid, kalk and zinkoxid.
Mätningar gjordes också på sekundärhuvsgaserna från konvertrarna, som också kan
användas i skrubbern då flashugnens tappgaser har ett fluktuerande flöde. Vidare
gjordes tester på olika processvatten på Rönnskär, som skulle kunna återanvändas i
skrubbern. Skrubbrarna utvärderades sedan baserat på indata och olika systemkrav
med hjälp av simulerings­ och designprogram samt teoretiska beräkningar.
Resultaten visade att det är rimligt att rena sekundärhuvsgaser i skrubbern, då de
innehöll ca 280 ton SO2 /år. Detta kan jämföras med flashugnens tappgaser som
innehöll ca 445 ton SO2 /år. Bland skrubbrarna, så utvecklade peroxidskrubbern sig till
att vara den mest attraktiva tekniken på grund av dess relativt låga livscykelkostnader
och dess lämplighet med både lakverket och flashkyltornsprocessvattnet. De andra
teknikerna med packade torn, lut­ och sodaskrubbern, hade högst livscykelkostnader,
huvudsakligen på grund av deras höga kemikaliekostnader. Sodaskrubbern, som var
den billigare av de två natriumbaserade skrubbarna, skulle kunna vara lämplig ändå på
grund av teknikens enkelhet. De öppna tornen hade lägre livscykelkostnader jämfört
med de packade tornen. Dock så hade kalkskrubbern flera nackdelar som gör att
den inte är ett lämpligt alternativ. Zinkoxidskrubbning har i sin tur inte studerats
tillräckligt och är en relativt oprövad teknik, men den bör studeras vidare då den
skulle kunna integreras med zinksmältningsprocessen på Rönnskär. Användningen
av processvatten i skrubbern skulle leda till en nettominskning av processvatten till
reningsverket. Det skulle även leda till minskade reagentkostnader ifall en stripper
installeras för att ta bort SO2 från processvattnen innan de går in i skrubbern.
Acknowledgements
First and foremost, I would like to thank my supervisors Jörgen Johansson (Boliden)
and Olle Gunnarsson (Metso Outotec) for providing me with guidance, support and
expertise throughout my work, which has been invaluable for me.
Further I would like to thank Robert Johansson, Peter Olsson and Kristoffer Renström
at Boliden as well as Leif Skilling at Metso Outotec, for sharing their expertise
throughout my work, but also for giving me the opportunity to write my master thesis
at both companies in the first place.
There were many other people who helped me at Rönnskär with various tasks and
questions, who I would like to thank as well. I have really enjoyed my time at both
companies and I believe that I have developed and learnt a lot. Thanks to all of
you!
Contents
1 Introduction
1
1.1 Aim and objectives . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .
2
1.2 Boliden Rönnskär & Metso Outotec . . . . . . . . . . . . . . . . . . . .
2
1.3 Delimitations . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .
3
2 Background
5
2.1 Overview of processes . . . . . . . . . . . . . . . . . . . . . . . . . . .
5
2.1.1 Flash smelting and tapping . . . . . . . . . . . . . . . . . . . . .
5
2.1.2 Converting . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .
6
2.2 Processes at Boliden Rönnskär . . . . . . . . . . . . . . . . . . . . . .
7
. . . . . . . . . . . . . . . . . . . . . . . .
7
2.2.2 Process water treatment plant (RV1) and stripper . . . . . . . .
8
2.2.3 Leaching plant . . . . . . . . . . . . . . . . . . . . . . . . . . . .
10
2.2.4 Process water sources . . . . . . . . . . . . . . . . . . . . . . .
10
2.3 Scrubber techniques . . . . . . . . . . . . . . . . . . . . . . . . . . . . .
10
2.3.1 SO2 absorption theory . . . . . . . . . . . . . . . . . . . . . . . .
10
2.3.2 Tower types . . . . . . . . . . . . . . . . . . . . . . . . . . . . .
11
2.3.3 Sodium­based scrubbing . . . . . . . . . . . . . . . . . . . . . .
13
2.3.4 Peroxide scrubbing . . . . . . . . . . . . . . . . . . . . . . . . .
14
2.3.5 Lime scrubbing . . . . . . . . . . . . . . . . . . . . . . . . . . . .
15
. . . . . . . . . . . . . . . . . . . . . . . .
16
2.2.1 Copper and gas flow
2.3.6 Zinc oxide scrubbing
3 Method
17
3.1 Input data . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .
17
3.1.1 Flash tapping gases and secondary hood gases . . . . . . . . .
17
3.1.2 Process waters . . . . . . . . . . . . . . . . . . . . . . . . . . . .
20
3.1.3 Costs . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .
21
3.2 System requirements . . . . . . . . . . . . . . . . . . . . . . . . . . . .
22
3.3 Scenarios . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .
22
3.4 Simulation and tower design . . . . . . . . . . . . . . . . . . . . . . . .
23
3.4.1 Simulation . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .
23
3.4.2 Tower design . . . . . . . . . . . . . . . . . . . . . . . . . . . . .
24
3.5 Cost analysis . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .
26
3.5.1 Operational costs . . . . . . . . . . . . . . . . . . . . . . . . . .
26
3.5.2 Capital investment cost estimation . . . . . . . . . . . . . . . . .
27
3.5.3 Life cycle costs and costs per unit reduction of SO2 emissions .
29
4 Results and analysis
30
4.1 SO2 from secondary hood gases . . . . . . . . . . . . . . . . . . . . . .
30
4.2 Base case scenario results . . . . . . . . . . . . . . . . . . . . . . . . .
30
4.2.1 Equilibrium temperatures . . . . . . . . . . . . . . . . . . . . . .
31
4.2.2 Water balances . . . . . . . . . . . . . . . . . . . . . . . . . . .
31
4.2.3 Bleed compositions . . . . . . . . . . . . . . . . . . . . . . . . .
32
4.3 Tower design . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .
34
4.4 Costs . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .
35
4.4.1 Operational costs . . . . . . . . . . . . . . . . . . . . . . . . . .
35
4.4.2 Capital investment costs . . . . . . . . . . . . . . . . . . . . . .
36
4.4.3 Life cycle costs and costs per unit reduction of SO2 emissions .
37
4.5 Further analysis . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .
38
4.5.1 Process water . . . . . . . . . . . . . . . . . . . . . . . . . . . .
38
4.5.2 Minimum and maximum SO2 scenarios . . . . . . . . . . . . . .
39
4.5.3 Costs to obtain different removal efficiencies . . . . . . . . . . .
40
5 Discussion
42
6 Conclusions and future work
46
Bibliography
48
A Tower design figures
53
B Life cycle costs and costs per unit reduction of SO2 emissions
(10 % WACC)
55
Acronyms
FGD ­ Flue gas desulfurization
RV1 ­ The process water treatment plant at Rönnskär
NaOH ­ Caustic soda
Na2 CO3 ­ Soda ash
BAT ­ Best available technology
BAT­AEL ­ BAT­associated emission levels
L/G ratio ­ Liquid­to­gas ratio
WACC ­ Weighted average cost of capital
Chapter 1
Introduction
Sulfur dioxide (SO2 ), a colorless gas with a sharp odor, has an acidifying effect on the
environment which harms both plants and animals. For humans, it may affect the
respiratory tract and lung function, as well as cause eye irritation. The processing
industry is the largest source of SO2 emissions in Sweden. This is due to the use of
sulfur­containing fuels and raw materials. [42]. In the non­ferrous metals industry,
the emissions originate from the processing of metal sulfides such as chalcopyrite
(CuFeS2 ) sphalerite ((Zn,Fe)S) and galena (PbS) [40]. Emissions of SO2 are regulated
on UN, EU and Swedish national level ([42]; [41]), which affects the non­ferrous metals
industry in different ways.
The copper smelter Boliden Rönnskär has an environmental permit from the Swedish
environmental court to emit 3500 tonnes of SO2 /year from 2019 and onwards, which
is a reduction from a previous permit to emit 4500 tonnes of SO2 /year. Since the
new permit was introduced, Rönnskär has been close to the emissions limit. There
are also plans for an increased copper production in the coming years, which would
be associated with increased SO2 emissions. To create a margin to the new emissions
limit, the company has identified the possibility of reducing the SO2 emissions from the
flash furnace’s tapping gases. This is as they are easily accessible and as they contain
a large share of the SO2 emissions at Rönnskär. The SO2 reduction is to be conducted
through the use of a wet scrubber, in which the SO2 is absorbed from the flue gases
through various techniques.
As the flash tapping gases are released in an intermittent pattern, periodically there will
be capacity in the scrubber to take in other process gases as well. An example of gases
1
CHAPTER 1. INTRODUCTION
that could be added to the scrubber are the gases from the secondary hoods, which
are the process gases that the primary hood of the converter does not have capacity
for. Today, the flash tapping gas and the secondary hood gas streams go through a bag
filter and a stack and reach the atmosphere without any reduction of SO2 . Installing a
scrubber would free up capacity in the bag filter, leading to reduced levels of SO2 in the
tapping area and in the converter aisle, which is positive from a working environment
perspective.
1.1 Aim and objectives
The aim of the project is to investigate, evaluate and compare different scrubbing
techniques to reduce the SO2 emissions from the flash furnace’s tapping gases. To
reach the aim, the following objectives will have to be fulfilled:
• Compile input data about the flash tapping gases.
• Investigate if it is reasonable to take in secondary hood gases into the scrubber.
• Perform simulations on and design of the different scrubber techniques.
• Perform cost analysis on the different scrubber techniques.
• Account for how the scrubbers are integrated into and affect Rönnskär’s other
processes.
1.2 Boliden Rönnskär & Metso Outotec
This thesis report was conducted together with Boliden AB and Metso Outotec Sweden
AB. Boliden were mainly involved in input data collection and process integration,
while Metso Outotec were mainly involved in the simulation and design of the scrubber
techniques. Practical work was performed at Boliden’s copper smelter site, Rönnskär.
The Rönnskär site is located in Skelleftehamn in the municipality of Skellefteå while
the local Metso Outotec office is located in Skellefteå in the same municipality.
Boliden is a metal company with its own mines and smelters that supplies base and
precious metals through the mining of ore and the production and delivery of metals
to the industry. Boliden has operations in Sweden, Finland, Norway and Ireland
[5]. The Rönnskär site is mainly a copper smelter, but other non­ferrous metals are
2
CHAPTER 1. INTRODUCTION
also produced. The plant receives copper concentrates from its own mines (in the
Boliden area and from Aitik) and from external suppliers. Rönnskär is Boliden’s
biggest production unit [4].
Metso Outotec provides technologies, end­to­end solutions and services for the
minerals processing, aggregates, metals refining and recycling industries globally [24].
Metso Outotec is a merger between Metso Minerals and Outotec, which was established
on the 1/7/2020 [23].
1.3 Delimitations
As Boliden Rönnskär is a site with many processes that are integrated in a complex
system (see Figure 2.2.1), many of the scrubber techniques affect other parts of the
site. The calculations however only focused on the scrubbers themselves as well as
on the incoming process waters, the stripper and the process water treatment plant.
Implications on other relevant operations at Rönnskär were only part of the analysis
and discussion and were not included in any calculations.
Further, there are many methods to abate SO2 from gas streams including using
a cleaner fuel, using a sulfur recovery unit and the use of end­of­pipe treatment
methods like flue gas desulfurization (FGD), biological technologies and electron beam
irradiation [13]. However, in this study only FGD systems were studied. This decision
was made based on a pre­study that was conducted by Boliden. Among the FGD
systems, only wet scrubbers were evaluated. This is as wet scrubbers have a higher
removal efficiency than dry and semi­dry scrubbers [43] and as dry and semi­dry
scrubbers have mainly been applied for low sulfur content process gases at high
temperatures in the coal industry [25]. Other system requirements were set based on
Boliden’s pre­study and based on legislation, which are presented in Section 3.2.
In a screening process, five wet scrubbers were selected for the evaluation in this report.
These are scrubbers based on the following reagents: caustic soda, soda ash, peroxide,
lime and zinc oxide. These scrubbers were selected due to the following reasons:
• The caustic soda and soda ash scrubbers were selected due to their simplicity
and technical maturity. They also produce a liquid effluent, which can be treated
further at Boliden’s process water treatment plant.
3
CHAPTER 1. INTRODUCTION
• The peroxide scrubber was selected as it is a proven technique. It was also chosen
as it had potential synergies with the leaching plant.
• The zinc oxide scrubber was selected due to its possible suitability with the zinc
smelting at Rönnskär.
• The lime scrubber was selected due to its relatively low reagent cost.
Other types of scrubbers were not considered due to their costs, removal efficiency,
complexity or the fact that the technique was not used on a commercial scale.
More information about the chosen scrubber techniques will be presented in Section
2.3.
In the cost analysis, the operational cost only included the reagent costs and the costs
for running the pumps and the fan. The costs included in the capital investment cost
are presented in Section 3.5.2.
4
Chapter 2
Background
In this chapter a general overview is given of flash smelting, flash tapping and
converting together with a presentation of the relevant processes at Boliden Rönnskär.
Further, the scrubber techniques are presented by describing their operating
principles, their advantages and disadvantages and by giving examples of industrial
plants using the scrubber in question. SO2 absorption theory and different tower types
used for wet scrubbing are also presented.
2.1 Overview of processes
In this section a general overview of flash smelting and tapping as well as converting
is given.
2.1.1 Flash smelting and tapping
Flash smelting is the most common process for copper (Cu) matte smelting. In the
flash furnace copper­iron­sulfur (Cu­Fe­S) concentrate reacts with oxygen gas (O2 )
continuously. This leads to oxidation of the concentrate’s Fe and S, heat evolution
and melting of the solids. Flash smelting results in [38]:
• Molten Cu­Fe­S matte with ∼60 % Cu, which is fed to the converters.
• Molten iron­silicate slag with 1­2 % Cu, which after treatment for Cu recovery
only contains a minor fraction of the Cu input.
5
CHAPTER 2. BACKGROUND
• Hot dust­laden off­gas containing 30­50 vol.­% SO2 , which can be further treated
for sulfuric acid (H2 SO4 ) production.
The matte and slag are tapped through water­cooled holes embedded in the furnace
walls. The holes typically have a 60­80 mm diameter. The tapping is only partial as
0.5 m deep reservoirs for matte and slag respectively are maintained in the furnace.
During steady­state operation the tapping of matte is done on a scheduled basis or
when it is needed in the converter. Tapping of slag is also done on a scheduled basis
or when it reaches a specific level in the furnace [38]. Tapping leads to SO2 emissions
around the furnace, especially on the matte side that contains more sulfur [33].
2.1.2 Converting
In the converter the Cu from the matte, containing approximately 60 % Cu, is oxidized
into molten blister copper that contains approximately 99 % Cu. Similarly to the flash
smelting, the converting also produces molten iron­silicate slag and SO2 ­containing
off­gas. The converting occurs through the oxidation of Fe and S from the matte with
oxygen­enriched air. [38].
Converting occurs in two stages, where the first one is the slag­forming stage (slag
blow) in which primarily Fe and S are oxidized to ferrous oxide (FeO), magnetite
(Fe3 O4 ) and SO2 . Silica flux is added to form a liquid slag with the Fe oxides. The
slag­forming stage is done when the Fe in the matte has been lowered to about 1 %.
The slag is separated from the remaining matte, which now mainly consists of impure
molten white metal (Cu2 S) [38].
The second stage is copper making (copper blow), when the S in the white metal is
oxidized to SO2 . Cu is not noticeably oxidized until almost all S is oxidized, which
leaves the produced blister copper with low levels of both S and O (0.001­0.03 % S
and 0.1­0.8 % O). The copper blow is terminated when copper oxide begins to appear
[38].
The matte is charged to the converter in several steps. Each step is followed by the
slag blow and slag being poured from the converter before new matte is charged. This
leads to accumulation of Cu in the converter until there is sufficient molten white metal
(250­300 tonnes) for the copper blow. The blister copper is poured from the converter
into ladles and is transported to the anode casting for further refining [38].
6
CHAPTER 2. BACKGROUND
The converter off­gas is collected by a primary hood that should fit as closely as possible
over the converter mouth. The off­gas entering the primary hood contains between 8­
10 vol.­% SO2 [38]. Outotec primary hoods capture over 95 % of the SO2 from the
converting process [6].
The secondary hood collects SO2 ­bearing gas that is not captured by the primary hood
[38]. Too much SO2 in the secondary hood during normal operation is an indication
of the primary system being insufficient and in need of modification. Secondary hoods
handle secondary gases such as emissions from [6]:
• rolling in and out the converter.
• charging the converter in both filling and blowing position.
• tapping matte, slag and blister copper.
• parked ladles.
Secondary hoods are important for a safer and cleaner working environment. They
also extend equipment lifetime due to less dust exposure [6].
2.2 Processes at Boliden Rönnskär
In this section the copper production process at Rönnskär and a few other processes
of interest for this project will be presented.
2.2.1 Copper and gas flow
In Figure 2.2.1 the various processes at Boliden Rönnskär are included. The copper
flow (in orange) represents the set of processes that are involved in the processing of
Cu. The Cu concentrate goes through either the flash furnace or the electric smelting
furnace while the Cu­containing electronic waste goes through the E­kaldo plant.
Secondary raw material, which contains copper, enters the electric smelting furnace
as well.
The next step from both furnaces and the E­kaldo plant is the converter, where the
blister copper is produced. The plant has three converters, which are more or less
identical, of which two are in operation simultaneously. During normal operation, one
7
CHAPTER 2. BACKGROUND
Figure 2.2.1: Operations at Boliden Rönnskär.
cycle is completed in one of the converters before the other one begins its cycle. Each
cycle takes around 8 hours.
Figure 2.2.1 also shows the gas flow (in black) which represents the SO2 ­dense (7­40
vol.­% SO2 ) primary gas flows from the furnaces, the Kaldo plant, the converter aisle
(the primary hood gases) and a fluidized bed roaster. But there are also other low SO2 ­
bearing gases that are not included in the figure such as the flash tapping gases and the
gases entering the secondary hood, which are of main interest in this report. The gases
that do not enter the primary hood of the converter (less than 1 %) either rise towards
the converter aisle ceiling or enter the secondary hood where they are transported to
a bag filter. The flash tapping gases are transported to the same bag filter before they
both exit through a stack.
2.2.2 Process water treatment plant (RV1) and stripper
The process water treatment plant (RV1) is designed for 200 m3 water/h and handles
160 m3 water/h on average, of which 110 m3 /h is process water from Rönnskär’s
different processes. The flow rate is expected to increase with different expansions
8
CHAPTER 2. BACKGROUND
planned at Rönnskär. Most of the current process waters come from gas treatment
processes such as scrubbers, cooling towers, WET­ESPs etc., but also from other
processes such as the electrolysis plant. These waters contain various concentrations
of SO2 and H2 SO4 making the incoming water to RV1 acidic (pH 1­2).
All of the process water that comes to RV1 goes through a stripper. The stripper is
a packed tower, where air is used to remove SO2 from the process water. The SO2
from the stripper is then mixed with the concentrated SO2 gases entering the sulfuric
acid plant. The average volumetric flow and SO2 concentration of the incoming process
water are both above the correspondent values that the stripper is designed for, leading
to an average removal efficiency of 90% instead of the 95 % that it should remove
[36].
Green liquor, a dissolved smelt of mainly sodium carbonate (Na2 CO3 ) and sodium
sulfide (Na2 S), is used to precipitate metals and arsenic at RV1, but can also react with
other compounds such as SO2 . How much SO2 that is left in the process water after
the stripping determines how much additional green liquor that is consumed when
the water reaches RV1. SO2 dissolved in the incoming process water reacts with the
active precipitation chemical in green liquor, sulfide (S2− ), to form elementary sulfur
(S0 ) which reacts further with SO2 and water to form thiosulfate (S2 O2−
3 ). The reactions
are the following [36]:
2 S2–(aq) + SO2(aq) + 4 H+(aq)
S0(s) + SO2(aq) + H2O(l)
3 S0(s) + 2 H2O(l)
(2.1)
S2O32–(aq) + 2 H+(aq)
(2.2)
With a higher removal efficiency in the stripper, less SO2 would end up at RV1
consuming green liquor. Around half of Boliden’s green liquor costs can be attributed
to the poor function of the stripper. Further, high SO2 levels in the incoming water
result in working environment issues for the staff at RV1 [36].
The scrubber bleeds consist of both sulfites and sulfates. The sulfites turn into SO2 at
low pH values and consume green liquor at RV1 while the sulfates do not need to be
treated before being discharged into the sea. Sulfates are therefore not a problem for
RV1. Any H2 SO4 that reaches RV1 is neutralized using caustic soda (NaOH).
9
CHAPTER 2. BACKGROUND
2.2.3 Leaching plant
A new leaching plant is being built at Rönnskär where hazardous waste from the own
production is being reprocessed to reduce the amount of waste ending up in a deep
repository while the yield of already mined ore increases [14].
The leaching plant process will consist of a multi­stage leaching process.
Input
chemicals will mainly consist of H2 SO4 (70­76 %) from the H2 SO4 production at
Rönnskär, green liquor and NaOH. The estimated use of H2 SO4 (70­76 %) is around 1.3
m3 /h (11 000 m3 /year) and the plant is expected to be in operation 340 days/year.
2.2.4 Process water sources
The two process water that were evaluated for reuse as make­up water in the scrubber
originate from the clinker scrubber and the flash cooling tower. The clinker scrubber
is a caustic soda scrubber used to remove halogens and SO2 from gases from the
zinc clinker process. The flash cooling tower, cools the flash smelting off­gases from
around 65 ◦ C to 25 ◦ C and is then taken to the remaining steps of the sulfuric acid
production.
The compositions and other properties of both the clinker scrubber process water and
the flash cooling tower process water are presented in Section 3.1.2.
2.3 Scrubber techniques
In this section, SO2 absorption theory and the different scrubber techniques that were
evaluated in this report are presented.
2.3.1 SO2 absorption theory
Removal of SO2 from flue gases occurs in three steps [17]:
1. SO2 is absorbed into water droplets.
2. The dissolved SO2 reacts with an alkaline species.
3. The reaction product is removed.
When the liquid reaches its saturation limit of SO2 , the mass transfer rates are in
equilibrium and no additional SO2 can be removed. Therefore it is important to design
10
CHAPTER 2. BACKGROUND
and operate the scrubbers so that saturation conditions are not reached, which is done
by providing sufficient liquid or by a chemical reaction that prevents the dissolved
SO2 from returning to the gas phase [17]. Many commercial gas absorption processes
involve chemical reactions in the liquid phase [25]. The reaction between SO2 and
the alkaline solutions is, based on decades of studies, assumed to be instantaneous.
Instead it is the mass transfer that is considered the limiting and most important step
[26, 43].
Absorption of SO2 in water is a pH­dependent process and the pH of the liquid
decreases as it approaches saturation. Therefore, the removal efficiency is a function
of the liquid’s pH. Scrubbers are operated at pH levels above 5.5 and preferably
between 6.0 and 6.5 to avoid mass transfer limitations. The pH levels are maintained
by continually adding alkali to react with SO2 and therefore there is no significant
equilibrium limit of the absorption when the liquid’s pH is above 5.5 [17]. The alkalinity
in wet scrubbers keeps the following two equilibrium reactions progressing to the right
[39]:
SO2(g) + H2O(l)
H2SO3(aq)
(2.3)
The sulfurous acid (H2 SO3 ) further disassociates to form bisulfites or sulfites
[39]:
H2SO3(aq)
HSO3–(aq) + H+(aq)
SO3–2(aq) + 2 H+(aq)
(2.4)
Although a higher gas temperature leads to a lower absorption rate and vice­versa [26],
for wet scrubbing systems, the inlet gas temperature is mainly important with respect
to the construction materials in the scrubber. Due to the long contact time between gas
and liquid in the scrubber, the temperature of the outlet gas is close to the adiabatic
saturation temperature which is reached through vaporization of the liquid [17].
2.3.2 Tower types
Packed towers and open spray towers are the two most common contactors used to
facilitate the absorption in wet scrubbing.
11
CHAPTER 2. BACKGROUND
Packed tower
A packed tower is a column filled with packing material that provides a large surface
area for the contact of liquid and gas [26]. A countercurrent flow is preferred, making
the most diluted gas to be in contact with the cleanest scrubbing liquid [43]. A packed
tower normally consists of a tower shell, mist eliminator, liquid distributor, packing
material and a packing support. Both structured and random packings are used for
scrubbers and are mainly made of metal or plastic. Metallic packings are not suitable
for corrosive flue gases, while plastic packings are not suitable for high temperatures.
Ideally, the packing’s life time will be as long as the tower’s lifetime, but in adverse
environments it can be shorter due to corrosion, fouling and/or breakage [26].
The size of a packed tower is mainly affected by the gas velocity. When the diameter of
the tower is decreased, the gas velocity increases. At some point the velocity becomes so
high that it restricts the downward flow of the liquid. The liquid starts to accumulate at
the so called loading point, in which the pressure drop starts to increase and the mixing
between phases decreases. Increasing the gas velocity further passed the loading point
leads to the liquid filling the void space in the packing completely which in turn leads
to substantial increase in pressure drop and minimal mixing between phases. This is
called flooding and has to be avoided [18].
Advantages with packed towers are that they can reach high removal efficiencies
and that they have relatively low water consumption requirements. Disadvantages
include risk of fouling and clogging and extensive maintenance of the packing material.
Installation, operational and wastewater disposal costs may also be higher for packed
towers compared to other types of contactors [26]. Soda­based scrubbers normally use
packed towers [9].
Open spray tower
In open spray towers, the liquid is injected into the tower through a spray distribution
system, where it comes in contact with the flue gas that flows in the opposite direction
[26]. The towers are completely open and have no internal components except for
the spray nozzles and connecting piping [17]. This makes them simple to operate
and maintain and they also have lower pressure drops than packed towers. However,
compared to other scrubbers they have lower mass transfer, leading to lower removal
efficiencies for gases of lower solubility. To reach higher removal efficiencies they
12
CHAPTER 2. BACKGROUND
require higher water circulation rates [26]. Lime scrubbers are typically open spray
towers, due to the risk of scaling and plugging in a packed column [9, 12].
2.3.3 Sodium­based scrubbing
Among sodium­based scrubbers, the most common systems use aqueous solutions of
either caustic soda (NaOH) or soda ash (Na2 CO3 ) as reagents to absorb SO2 . They both
produce a clear liquid solution with a high pH and can be used interchangeably in most
scrubbers. Caustic soda dissolves in water to create the scrubbing solution through the
following reaction [39]:
Na+(aq) + OH–(aq) + H2O(l)
NaOH(s) + H2O(l)
(2.5)
For soda ash, the scrubbing solution is created through the following reactions
[39]:
Na2CO3(s) + H2O(l)
2 Na+(aq) + CO3–2(aq) + H2O(l)
(2.6)
HCO3–(aq) + OH–(aq)
(2.7)
CO3–2(aq) + H2O(l)
Both soda ash and caustic soda scrubbers produce a mixture of sodium sulfite
(Na2 SO3 ), sodium sulfate (Na2 SO4 ) and sodium bisulfite (NaHSO3 ). The proportions
of the different compounds depends on the pH and the degree of oxidation. The overall
reactions with SO2 are [39]:
SO2(g) + 2 Na+(aq) + 2 OH–(aq)
SO2(g) + 2 Na+(aq) + 2 OH–(aq) +
1
O2(g)
2
SO2(g) + Na+(aq) + OH–(aq)
Na2SO3(aq) + H2O(l)
(2.8)
Na2SO4(aq) + H2O(l)
(2.9)
NaHSO3(aq)
(2.10)
13
CHAPTER 2. BACKGROUND
As the systems have a relatively simple design, they require relatively little space.
Further, scaling and plugging risks are small and the power consumption is low
due to a low liquid­to­gas flow ratio [21]. Another advantage is that the sodium­
based scrubbers produce an aqueous solution instead of a slurry, which simplifies
the operation of the scrubber [15]. Caustic soda scrubbers tend to be slightly more
expensive than soda ash scrubbers and the cost of caustic soda also fluctuates more
than other reagents due to a high market demand [39]. The costs for the sodium­
based reagents are in general higher than for the reagents in lime­based scrubbers [17,
39]. The liquid effluent from sodium­based scrubbing may create disposal problems
at some locations and has to be treated before it is discharged into any watercourse
[39].
At the Freeport McMoRan Miami smelter a large 2 300 000 Nm3 /h scrubber, using 20
% caustic soda as reagent, is used to treat converter aisle fugitives and anode off­gas
[16]. Caustic soda is otherwise typically sold as a 50 % solution [39].
2.3.4 Peroxide scrubbing
In peroxide scrubbing, SO2 is removed through the following reaction with hydrogen
peroxide (H2 O2 ):
SO2(g) + H2O2(aq)
H2SO4(aq)
(2.11)
The Peracidox ® scrubber is one of the patented scrubber techniques using H2 O2 ,
owned by Metso Outotec, to remove residual SO2 from dilute off gases. In this process,
scrubbing is achieved through direct contact in a packed tower, where the hydrogen
peroxide reacts with SO2 in the first chamber and then overflows to a second chamber.
The exiting 50 % H2 SO4 is either recycled to a sulfuric acid plant or sold as a by­product
if a market exists [9]. The concentration of the outgoing H2 SO4 will be approximately
the same as the concentration of the incoming H2 O2 in a peroxide scrubber (Personal
communication with Leif Skilling, 15/2/2021).
14
CHAPTER 2. BACKGROUND
Metso Outotec lists the following benefits with the Peracidox ® process [2]:
• No additional chemicals (H2 SO4 as product, no wastewater or residues).
• Low opacity.
• High removal efficiency.
• High reliability.
• Low investment cost.
• High flexibility with respect to variations in SO2 content of the gas.
Disadvantages with the Peracidox ® process include a relatively high reagent cost and
the relatively low strength of the resulting H2 SO4 solution [9].
2.3.5 Lime scrubbing
Wet scrubbing systems based on lime and limestone are the most popular commercial
FGD systems [25]. In fact, they are used in over 90 % of all FGD systems in the
United States [34]. In lime scrubbers, the flue gas is scrubbed with calcium hydroxide
(Ca(OH2 )) which is formed when quicklime (CaO) reacts with water. The slurry is
pumped into an open spray tower where the SO2 is absorbed into the droplets of the
slurry, which leads to several chemical reactions that can be simplified as [25]:
CaSO3 ·
1
3
H2O(s) + H2O(l)
2
2
(2.12)
1
3
1
H2O(s) + H2O(l) + O2(g)
2
2
2
CaSO4 · 2 H2O(s)
(2.13)
SO2(g) + Ca(OH)2(s) + H2O(l)
CaSO3 ·
The product in Reaction 2.13. is gypsum (CaSO4 * 2 H2 O) that can be sold if it is clean
of impurities and if a market exists. Advantages of lime­based scrubbers include that
they are simple, use a widely available and inexpensive reagent and that they produce a
potentially usable product (gypsum). Disadvantages include high capital costs [25] and
the risk of calcium compounds accumulating in recirculating loops causing scaling [12].
Also, lime­based scrubbers can be adversely affected by rapid changes in concentration
[9] and operate poorly below 60 ◦ C (Personal communication with Olle Gunnarsson,
14/1/2021).
15
CHAPTER 2. BACKGROUND
2.3.6 Zinc oxide scrubbing
Zinc oxide (ZnO) reacts with water to create zinc hydroxide (Zn(OH)2 , which reacts
with the SO2 through several reactions to form zinc sulfite (ZnSO3 ) and zinc sulfate
(ZnSO4 ). These reactions can be simplified as [19]:
SO2(g) + Zn(OH)2(s)
ZnSO3(s) +
1
O2(g)
2
ZnSO3(s) + H2O(l)
(2.14)
ZnSO4(aq)
(2.15)
Zinc oxide scrubbing is suitable at a zinc smelter as the ZnO reagent can be obtained
from the fuming plant and the clinker furnace. The resulting ZnSO3 and ZnSO4 can
be processed further before they are reintroduced as raw material in the zinc smelting
process. As the zinc scrubber takes reagent from and later reintroduces the products to
the zinc smelting process, the scrubber has no reagent cost. ZnO has limited solubility
in water, making it the rate limiting step for the overall reaction with SO2 . This leads
to several problems [19]:
• Scaling and blockage of the equipment occurs easily.
• To ensure a high removal efficiency the liquid­to­gas ratio needs to be increased,
leading to high energy consumption.
• Poor buffering ability of pH in ZnO solution and low absorptive capacity.
Organic additives such as citric acid can be used to increase the absorption rate and SO2
removal efficiency, but also to prevent scaling and corrosion [19]. DynaWave, who have
installed zinc oxide scrubbers at zinc smelters in Belgium and Australia, have tackled
the low dissolution rate and achieved higher removal efficiencies than other zinc oxide
scrubbers by using a reverse jet nozzle. The nozzle is used to create an intense mass
and energy transfer region, where the flue gas is quenched and the SO2 is transferred
to the liquid [10, 11].
16
Chapter 3
Method
In this chapter a description is given of how the different parameters were acquired,
chosen or calculated. First the input data regarding the gas streams, process waters and
costs are presented. Then the system requirements and scenarios are presented which
the calculations, simulations and design were based on. Finally the set of methods used
for the cost analysis are presented.
3.1 Input data
Most of the input data were collected from existing process data at Boliden. For some
processes and parameters, additional measurements or tests were conducted.
3.1.1 Flash tapping gases and secondary hood gases
Data regarding the flash tapping gases from the flash furnace were collected from
Boliden’s existing process data regarding volumetric flows, temperatures and SO2
concentrations. The data used to calculate the average slag tapping and average matte
tapping values as well as the minimum and maximum SO2 scenario values for the flash
tapping gases were based on process data from the 24/9/2020 as well as from the
period 13/11/2020 ­ 17/11/2020.
There were no data regarding the secondary hood gases, which is why continuous
measurements were conducted between 16/12/2020 ­ 21/12/2020 to monitor the
volumetric flow, SO2 concentration and temperature. The results during the period
17
CHAPTER 3. METHOD
Figure 3.1.1: Volumetric flow and SO2 concentration of the flash tapping gases during
the 24/9/2020.
16/12/2020 ­ 17/12/2020 were used to calculate the values in Table 3.1.2. This is, as it
was only during these two days when the operations followed a normal pattern.
The average values of the different flow parameters of the flash tapping gases are listed
for matte tapping and slag tapping periods in Table 3.1.1. The fan is constantly running
at the matte tapping area while a damper is opened during slag tapping periods leading
to a higher volumetric flow. It should be noted that the matte tapping values in Table
3.1.1 represent the whole period with lower volumetric flows, not only the period when
matte tapping actually occurs. In Figure 3.1.1 it can be seen how the SO2 concentration
peaks during matte tapping. The figure also shows the fluctuating volumetric flow rate.
The average values for slag tapping were used in the base case scenario, which will be
presented in Section 3.3.
Table 3.1.1: Average values of different flow parameters of the flash tapping gases.
Slag tapping
(base case scenario)
3
Volumetric flow (Nm /h) 138 800
Temperature (◦ C)
38
SO2 (ppm)
175
Parameter
Matte tapping
49 100
37
245
The average values of the different flow parameters for the secondary hood gases
are listed during forced ventilation (for example during charging of material into
converter) in Table 3.1.2. Figure 3.1.2 shows that the majority of the highest SO2
concentration values coincide with the forced ventilation into the secondary hood. This
is logical, as one of the reasons for pulling gas into the secondary hood is to reduce the
18
CHAPTER 3. METHOD
Table 3.1.2: Average values of different flow parameters of the secondary hood gases.
Forced
ventilation
3
Volumetric flow (Nm /h) 109 800
Temperature (◦ C)
28
SO2 (ppm)
270
Parameter
Remaining
operation time
31 000
45
90
Base case
scenario
111 200
38
125
Figure 3.1.2: Volumetric flow and SO2 concentration of the secondary hood gases in
converter 3 during the period 16/12/2020­17/12/2020.
SO2 level in the converter aisle. Figure 3.1.2 also shows how often the forced ventilation
occurs. As the measurements on the secondary hood gases were only conducted on one
secondary hood (on converter 3), it was assumed that similar flow properties would be
present in the other converter in operation. Average SO2 and temperature values were
calculated for the secondary hood gases in the base case scenarios (see Table 3.1.2) in
which secondary hood gas was pulled to fill up the remaining scrubber capacity (250
000 Nm3 /h ­ 138 800 Nm3 /h = 111 200 Nm3 /h). These averages include both the
forced ventilation and the remaining operation time periods.
The values in the minimum and maximum SO2 scenarios were based on ten minute
averages. This is as the gas streams are likely to be mixed in the transport towards
the scrubber. The minimum SO2 scenario consists of the minimum volumetric flow
rate from two secondary hoods (29 000 Nm3 /h from each secondary hood) and the
volumetric flow during matte tapping. The minimum SO2 levels approach zero at
times, but were set to 5 ppm for some reagent to be consumed. The values of the
flow parameters in the minimum and maximum SO2 scenarios are presented in Table
3.1.3.
19
CHAPTER 3. METHOD
Table 3.1.3: Values of different flow parameters in the minimum and maximum SO2
scenarios for the flash tapping gases and secondary hood gases.
Minimum SO2 scenario
Volumetric flow (Nm3 /h)
Temperature (◦ C)
SO2 (ppm)
Maximum SO2 scenario
Volumetric flow (Nm3 /h)
Temperature (C)
SO2 (ppm)
Flash tapping
gas
Secondary hood
gas
49 300
19
5
58 000
27
5
140 000
49
340
110 000
74
3480
3.1.2 Process waters
Two process waters were evaluated for reuse as make­up water in the scrubber. This is
as there was an ambition to use process water from existing processes at Rönnskär
to avoid an additional water load to RV1 that already works close to maximum
capacity.
Table 3.1.4: Properties of the clinker scrubber process water and the flash cooling tower
process water.
Parameter
Clinker scrubber
water
Flash cooling
tower water
Fluorides (mg/l)
Chlorides (mg/l)
Dry matter (mg/l)
SO2 (g/l)
H2 SO4 (g/l)
pH
Temperature (◦ C)
Volumetric flow (m3 /h)
760
1.8
540
4.4
­
8.9
30
19
155
<0.1
10
18.9
8.1
1.1
25
3
The process water data is from ten measurements on the clinker scrubber process
water (between September 2019 and January 2021) and four measurements on the
flash cooling tower process water (between November 2020 and January 2021), which
are presented in Table 3.1.4. The H2 SO4 value in the flash cooling tower process water
is based on nine separate measurements between March 2020 and March 2021. The
temperature and flow rate for the clinker scrubber process water were obtained from
zinc smelting staff at Rönnskär in February 2021 while the same values were obtained
from existing process data for the flash cooling tower process water during the same
20
CHAPTER 3. METHOD
month. In Table 3.1.4 it can be seen that the clinker scrubber process water is alkaline
and has high values of fluorides, chlorides and dry matter while the flash cooling tower
process water is acidic and has high SO2 and H2 SO4 concentrations. The latter also has
a smaller volumetric flow and temperature compared to the clinker scrubber process
water.
3.1.3 Costs
Reagent costs
Typical reagent costs are listed in Table 3.1.5. The zinc oxide scrubber has no reagent
cost as it is obtained from the zinc smelting process at Rönnskär.
Table 3.1.5: Reagent costs.
Reagent
Cost (SEK/tonne)
Caustic soda (50 weight­%) 2450
Soda ash
2450*
Lime (quicklime)
1150**
Peroxide (49 weight­%)
2900
Zinc oxide
0
Green liquor
800
* Including 125 SEK/tonne for emissions rights.
** Including landfill cost of 235 SEK/tonne.
Equipment costs
The costs for the tower and the tower internals were rough estimates with some input
from Olle Gunnarsson from Metso Outotec.
The pump capital costs were based on cost indexes in a report by the National Energy
Technology Laboratory (NETL) of the U.S. Department of Energy [20] while the fan
and mixing tank capital costs were based on Peters et al. [35]. The bag filter and the
belt filter (for the open spray towers) were not included in the cost calculations.
The exchange rates used for the equipment cost calculations were 8.37 SEK/$ and
10.03 SEK/€.
The remaining capital investment costs were estimated using the
method in Section 3.5.2.
21
CHAPTER 3. METHOD
3.2 System requirements
A few requirements were set for the scrubbing systems. As Boliden is regulated based
on total SO2 emissions (3500 tonnes/year), it is of interest to scrub as much SO2 as
possible for a reasonable cost. However, there is also a removal efficiency regulation
set by the EU commission when a unit for SO2 removal is installed. Wet scrubbers are
one of three techniques that are seen as the best available technique (BAT) for reducing
SO2 emissions from primary and secondary copper production. The BAT­associated
emissions levels (BAT­AEL) for SO2 emissions to air from a wet scrubber following
primary copper production is 350 mg/Nm3 [8] on a daily average. Boliden targets to
have a margin to the BAT­AEL.
Additionally, there were a few parameters that were fixed and decided based on
Boliden’s pre­study. One of these was the capacity of the scrubber which was set to 250
000 Nm3 /h. This is based on expected future flash furnace production that includes
both increased tapping and simultaneous tapping of slag and matte. It was also decided
that dust will be abated with a bag filter before the scrubber. The bag filter will be
designed to emit 1 mg/Nm3 , but should guarantee that the emissions do not exceed 2
mg/Nm3 .
3.3 Scenarios
The scrubber techniques were evaluated in different scenarios. First, five base case
scenarios were created based on the different reagents: caustic soda, soda ash,
peroxide, lime and zinc oxide. In these base case scenarios, the removal efficiency was
set to 90 % and clean industrial water of 10 ◦ C was used. The average values during
slag tapping were used for the flash tapping gases (see Table 3.1.1). The remaining
gas to fill up the capacity (250 000 Nm3 /h) was taken from the secondary hood gases
whose values for the different parameters in the base case scenarios are listed in Table
3.1.2. From these scenarios, results regarding the water balances, energy balances as
well as the sizes and compositions of the bleeds were obtained. It should be noted that
the base case scenarios are not averages of the volumetric flows, temperatures or SO2
concentrations for the flash tapping gases and the secondary hood gases. Instead they
represent one of the most probable operating modes when using both gas streams in
the proposed scrubber.
22
CHAPTER 3. METHOD
Further the use of the two process waters was evaluated. This was done by analyzing
how they affect both the scrubber and processes downstream of the scrubber. The
process water analysis was based on the make­up water need in the base case scenarios.
For the scrubbers that had the most promising results in the base case scenarios and
in operation with the process waters, a sensitivity analysis was conducted in which the
minimum and maximum SO2 scenarios (see Table 3.1.3) as well as the costs to obtain
different removal efficiencies were studied. The minimum and maximum scenarios
had the same removal efficiency (90 %) and used the same clean industrial water of
10 ◦ C as in the base case scenarios.
3.4 Simulation and tower design
The scenarios were studied using the simulation software METSIM, the design
software Rapsody and theoretical calculations. METSIM is a general­purpose process
simulation software designed to perform mass and energy balances in complex
processes [22] while Rapsody is a software for the design of packed columns [37].
3.4.1 Simulation
For the packed towers, the equilibrium temperature was first obtained in METSIM
for the maximum SO2 scenario, where the liquid flow rate was based on a liquid­to­
gas (L/G) ratio value from the literature (see Table 3.4.2). The obtained equilibrium
temperature was then used to design the packed tower in Rapsody, in which a more
suitable L/G ratio was chosen.
This new L/G ratio was then plugged back into
METSIM, in which it was used for the simulation of all scenarios. For the open spray
towers, the L/G ratio was based both on literature references and Metso Outotec’s
experience.
The sizes of the make­up water flow rates were decided by adding together the
water evaporated in the scrubber and the bleeds. While the evaporated water was
obtained directly from METSIM, the bleeds were calculated using the solubilities of
the scrubbing products that are listed in Table 3.4.1. The bleed sizes were calculated
in this way to avoid precipitation of the scrubbing products. The tabulated solubility
values are at various temperatures, but were the values encountered in the literature
that are the closest to the operating temperatures of the scrubbers.
23
CHAPTER 3. METHOD
Table 3.4.1: Solubilities of the products in the different scrubbers.
Solubility
(g compound/g water)
Temperature (◦ C) Reference
Caustic soda & soda ash
Na2 SO3
0.31
25
Na2 SO4
0.28
25
Peroxide
H2 SO4
*
­
Lime
CaSO3
0.000043
18
CaSO4 * 2 H2 O
0.0020
20
Zinc oxide
ZnSO3
0.0015
25
ZnSO4
0.56
25
* Soluble at all concentrations and temperatures.
[30]
[29]
[31]
[7]
[28]
[27]
[32]
As the reaction products of the lime and zinc oxide scrubbers are practically insoluble
(except ZnSO4 ), belt filters were used to dewater parts of the slurries and to obtain
bleeds that are 30 % solid. The peroxide scrubber bleed contains a 49 % H2 SO4 (based
on the input 49 % H2 O2 ). Bisulfites, such as sodium bisulfite, were not included in the
calculations and analyses.
3.4.2 Tower design
A packed tower and an open spray tower were designed based on the maximum SO2
scenario, to ensure that the scrubber is always able to clean 90 % of the SO2 . Only
one design for each type of tower was designed due to similar flow properties of the
packed tower scrubbers (caustic soda, soda ash and peroxide) and the open spray tower
scrubbers (lime and zinc oxide).
The towers consist of several parts. The top of the towers consists of a liquid/slurry
distributor, a gas outlet and a liquid/slurry inlet. The packing section consists of the
packed bed and the support grid for the packed towers while the reactive zone in the
open spray towers does not have any internal components. The bottom of the towers
includes the gas inlet and the liquid/slurry outlet.
The tower diameter for the packed towers was obtained from Rapsody and was decided
based on many design parameters which include packing type and size, flooding factor,
gas velocity, liquid load and pressure drop. The parameters mentioned above need to
24
CHAPTER 3. METHOD
have specific values or fall within specific ranges that were based both on the literature
and personal communication with Olle Gunnarsson. These common values/ranges are
listed for both packed and open spray towers in Table 3.4.2.
Table 3.4.2: Common values/ranges for various packed tower and open spray tower
parameters.
Common values/ranges
Packed tower
L/G ratio (L/m3 ) [18]
Gas velocity (m/s)
Operating pressure drop (kPa/m packing)
Total operating pressure drop (kPa)
Flooding factor (%) [18]
Liquid load (m3 /m2 /h)
Tank residence time (min)
Open spray tower
L/G ratio (L/m3 ) [18]
Gas velocity (m/s)
Total operating pressure drop (kPa)
Residence time (reactive zone) (s)
Tank residence time (min)
0.2­5.4
2­2.5
0.1­0.2
1.5
50­75
20
5
2.2­12.7
3­4
1
3­4
2
The height of the packing determines the removal efficiency and is decided with the
following formula [18]:
Z = HT U × N T U
(3.1)
where Z is the height of the packing, NTU is the number of transfer units and HTU
is the height of a transfer unit. The HTU and the NTU were both obtained from
Rapsody.
The tank height was calculated based on the tank residence time and the calculated
tower diameter. The gas inlet size was based on a 15 m/s velocity for the incoming gas
flow and the 250 000 Nm3 /h capacity. The section above the gas inlet has a height that
is 50 % larger than the gas inlet height which is to allow the gas to be well distributed
before reaching the packing/reactive section. 0.5 m height is added to the bottom of the
tower to compensate for the pump(s) being unable of pumping water from the absolute
bottom part of the tower. Heights for the remaining parts of the tower were decided
when all the heights mentioned above were decided.
25
CHAPTER 3. METHOD
3.5 Cost analysis
The costs included the operational costs and the capital investment costs. These costs
were later used to calculate the life cycle costs and the costs per unit reduction of SO2
emissions.
3.5.1 Operational costs
The operational costs include the reagent costs as well as the pump and fan power
costs.
The reagent costs were calculated based on the yearly reagent consumption which in
turn were based on the SO2 emissions during October 2019 ­ September 2020 for the
flash tapping gases and 16/12/2020 ­ 21/12/2020 for the secondary hood gases. The
size of these emissions is presented in Table 4.1.1 in the Results section. A 5 % surplus
of reagent was used to ensure that the desired removal efficiency was reached and that
an alkaline environment was maintained in the alkaline scrubbers.
The pump power costs were based on the shaft pump power which is calculated using
the following formula:
P =
qρgh
(3.6 × 106 )η
(3.2)
where P is the shaft pump power (kW), q is the flow (m3 /h), ρ is the density of the fluid
(kg/m3 ), g is the acceleration of gravity (m/s2 ), h is the differential head (m) and η is
the pump’s efficiency. The packed towers uses one pump that transports the liquid
from the bottom of the tower to the liquid distributor. The open spray towers uses
three pumps of equal size to transport the liquid to three different heights 2 meters
apart from each other. The utilization rate of the pumps in the open spray towers was
assessed to be 2/3 of the maximum capacity due to the fluctuating volumetric flows
and SO2 concentrations of the incoming gas streams. Further, the liquid was assumed
to drain down in the packed towers while it is sprayed into the open spray towers which
requires an additional 10 m head.
The fan power costs were based on the fan power which is calculated using the following
formula:
P =
dp × q
η
(3.3)
26
CHAPTER 3. METHOD
where P is the power used by the fan (kW), dp is the total pressure increase by the
fan (kPa), q is the gas volume flow delivered by the fan (m3 /s) and η is the fan’s
efficiency.
As previously shown in Table 3.4.2 the pressure drop for the packed towers and the
open spray towers were estimated to be 1.5 kPa and 1.0 kPa respectively. The electricity
cost was set to 0.45 SEK/kWh while the efficiency was set to 0.7 for both the pumps
and the fan.
3.5.2 Capital investment cost estimation
In this section, the set of methods used for estimating the capital investment costs are
presented. The capital cost estimates were study estimates, which are based on major
items of equipment which have a probable accuracy of estimate up to ± 30 %. A study
estimate is important for determining whether a proposed project should be given
further consideration or for comparison with alternative designs [35]. In Table 3.5.1
the items that are needed for a new facility and that were included in the calculations
are listed.
Table 3.5.1: Fixed­capital investment items for a chemical process [35].
Directs costs
Indirect costs
Delivered equipment
Delivered­equipment installation
Instrumentation and controls
Piping
Electrical systems
Buildings (including services)
Yard improvements
Service facilities
Land
Engineering and supervision
Legal expenses
Construction expenses
Contractor’s fee
Contingency
The delivered equipment costs were calculated using cost indexes, scaling and the data
and references presented in Section 3.1.3. The costs of the remaining fixed­capital
investment items were estimated using ratio factors based on the delivered equipment
cost. These methods are presented below.
27
CHAPTER 3. METHOD
Cost index
As prices could change with time due to changes in economic conditions, cost
indexes are used to update cost data applicable at a previous date to costs that are
representative of conditions today [35]. The present cost is calculated through the
following formula:
Present cost = Original cost ×
Index value at present
Index value when original cost was obtained
(3.4)
The Chemical Engineering Plant Cost Index (CEPCI) has been used since 1963, by
professionals in chemical process industry to adjust process plant construction costs
from one period to another [1]. The values used in this report are presented in Table
3.5.2.
Table 3.5.2: Chemical Engineering Plant Cost Index (CEPCI) values [3].
Year
CEPCI
1998 390
2002 396
2019 608
Scaling
If the two pieces of equipment are not of the same size, the capacity has to be adjusted.
The following formula can be used to predict the cost of a piece of equipment by using
the cost of a similar piece of equipment [35]:
Cost of equipment A = Cost of equipment B × X 0.6
(3.5)
where X is the capacity ratio between the two pieces of equipment [35].
Cost components
When the delivered equipment cost is determined, the other components of the capital
investment cost can be estimated by using ratio factors between the capital investment
items and the delivered equipment cost. This method is commonly used for study
28
CHAPTER 3. METHOD
Table 3.5.3: Percentages of the delivered equipment cost for fluid processing plants
[35].
Percentage of the delivered equipment cost (%)
Direct costs
Delivery of purchased equipment
10*
Installation of purchased equipment
60*
Instrumentation and controls (installed) 36
Piping (installed)
68
Electrical systems (installed)
15
Buildings (including services)
18
Yard improvements
10
Service facilities (installed)
6*
Indirect costs
Engineering and supervision
33
Construction expenses
41
Legal expenses
4
Contractor’s fee
22
Contingency
44
Working capital (15 % of TCI**)
­
* Percentage of purchased equipment.
** TCI ­ Total capital investment.
estimates with an expected accuracy of ± 20 to 30 % [35]. In Table 3.5.3 the ratio
factors are presented as percentages of the delivered equipment cost.
3.5.3 Life cycle costs and costs per unit reduction of SO2
emissions
The life cycle costs includes the previously presented capital investment costs and
operational costs, but also the cost of capital. A WACC (weighted average cost of
capital) rate of 6 % was used in the calculations, which has previously been found
reasonable for environmental investments at Rönnskär by Swedish authorities. The
same calculations were also done for a 10 % rate which is a normal WACC for an
investment in the production environment at Rönnskär (Personal communication with
Peter Olsson, 4/2/2021).
The economic lifetime was assumed to be 10 years and the capital investment costs
were assumed to be paid at equal installments over the economic lifetime. Further,
it was assumed that the economic lifetime is the same for all components of the
scrubbers.
29
Chapter 4
Results and analysis
4.1 SO2 from secondary hood gases
In Table 4.1.1 it can be seen that the potential SO2 from the secondary hood gases is
significant and amounts for around 39 % of the SO2 that could be removed. As the flash
tapping gases are expected to emit 780 tonnes of SO2 /year in the future, the total SO2
emissions from the flash tapping gases and the secondary hood gases would amount
to 1060 tonnes/year, of which 26 % would be secondary hood gases.
Table 4.1.1: SO2 emissions from the flash tapping gases and the secondary hood gases.
The emissions are based on the period October 2019 ­ September 2020 for the flash
tapping gases and for the period 16/12/2020 ­ 21/12/2020 for the secondary hood
gases.
Gas
SO2 emissions (tonnes/year)
Flash
445
Secondary hood 280
Total
725
4.2 Base case scenario results
In this section, all the tabulated values are in the cases where only sulfates are produced
in the reaction between SO2 and the reagent. The base case scenarios were based on
process data from the 24/9/2020 as well as from the period 13/11/2020 ­ 17/11/2020
for the flash tapping gases and from the period 16/12/2020 ­ 17/12/2020 for the
30
CHAPTER 4. RESULTS AND ANALYSIS
Table 4.2.1: Reagent consumption in the base case scenarios.
Reagent
(kg/h)
NaOH (50 weight­%)
Na2 CO3
H2 O2 (49 weight­%)
CaO
ZnO
260
170
115
90
130
secondary hood gases. The reagent consumption in the base case scenarios is listed
in Table 4.2.1.
4.2.1 Equilibrium temperatures
The equilibrium temperatures in the base case scenarios (see Table 4.2.2) do not vary
much. The variations that are observed are associated with the differences in reaction
enthalpies for the different reactions as well as the different water balances in the
scrubbers. It can be noted that the lime scrubber’s equilibrium temperature is far below
60 ◦ C, which is needed for the lime scrubber to operate properly.
Table 4.2.2: Equilibrium temperatures in the base case scenarios.
Scrubber
Equilibrium temperature (◦ C)
Caustic soda & Soda ash
Peroxide
Lime
Zinc oxide
20
23
19
20
4.2.2 Water balances
The water balances are presented in Table 4.2.3. The bleeds were based on the
solubilities in water for the caustic soda, soda ash and zinc oxide scrubbers meaning
that a bleed and make­up smaller than the listed minimum values would lead to
precipitation. For the peroxide scrubber the bleed was adjusted so that a 49 % H2 SO4
was produced and for the lime scrubber it was based on the belt filter that produces a
30 % solid.
The reaction between Na2 CO3 and water in the soda ash scrubber is endothermic
(31 400 kcal/h in the base case scenario) which makes the resulting NaOH/CO2 ­
solution cold (smaller than 0 ◦ C). To heat the mixture to at least 4 ◦ C, 38 kW heating of
31
CHAPTER 4. RESULTS AND ANALYSIS
Table 4.2.3: Water balances in the base case scenarios if only sulfates were produced.
Scrubber
Evaporation Minimum bleed
(m3 /h)
(m3 /h)
Minimum water added
(m3 /h)*
Caustic soda & Soda ash 2.76
0.78
3.54
Peroxide
2.32
0.23
2.55
Lime
2.90
0.57
3.47
Zinc oxide
2.68
0.45
3.13
* Includes both the make­up water and water from the reagent solutions.
the tank is needed. The heating costs are included in the operational cost for the soda
ash scrubber presented in Section 4.4.1. Another way to keep the temperature above 4
◦
C would be to add the make­up water and an additional 2 m3 water/h to the reagent
mixing tank. The latter option would also result in a larger bleed.
The caustic soda and soda ash scrubbers require a minimum bleed of 0.63 m3 /h if all
of the SO2 would be converted into Na2 SO3 .
A belt filter is used for the lime scrubber as the bleed based solely on the solubility of
CaSO4 would have been at least 160 m3 /h, which is unreasonably large. If only CaSO3
was formed, the bleed based on the solubility would be as large as 4660 m3 /h, while it
is 0.47 m3 /h using the belt filter.
As the zinc oxide scrubber produces ZnSO3 , it would require a belt filter as in the lime
scrubber. This is because if only ZnSO3 was produced in the reaction between Zn(OH)2
and SO2 , the bleed based on the solubility would be 150 m3 /h. Using a belt filter that
produces the same 30% solid as in the lime scrubber to separate the ZnSO3 , results in
a bleed of 0.53 m3 /h.
The biggest risk of accumulating impurities is in the peroxide scrubber as it has the
smallest bleed.
4.2.3 Bleed compositions
In Table 4.2.4 the bleed compositions for all base case scenarios are presented. The
bleed for the caustic soda and soda ash scrubbers end up in RV1, the lime scrubber
bleed ends up in a landfill (as the produced gypsum is not pure enough) and the zinc
oxide scrubber bleed is reintroduced into the zinc smelting process. The peroxide bleed
primarily ends up in the leaching plant, but as it is only in operation 340 days per year
there has to be alternative solutions.
32
CHAPTER 4. RESULTS AND ANALYSIS
Table 4.2.4: Bleed compositions in the base case scenarios. The rows with
sulfites/sulfates represent the extreme cases were only sulfites or only sulfates were
produced.
Mass flow rate (kg/h)
Caustic soda
H2 O (l)
NaOH (aq)
Na2 SO3 (aq)/Na2 SO4 (aq)
Soda ash
H2 O (l)
NaOH (aq)
Na2 SO3 (aq)/Na2 SO4 (aq)
CO2 (aq)
Peroxide
H2 O (l)
H2 O2 (aq)
H2 SO4 (aq)
Lime
H2 O (l)
Ca(OH)2 (s)
CaSO3 *0.5 H2 O (s)/CaSO4 *2 H2 O (s)
Zinc oxide
H2 O (l)
Zn(OH)2 (s)
ZnSO3 (s)/ZnSO4 (aq)
740
5
190/220
700
5
190/220
0.1
150
0.02
150
470
5
200/250
420
5
200/230
The alternative solutions are (in order):
• Mixing the peroxide bleed into the sulfuric acid plant.
• Not using the scrubber (the gases go through the converter filter where no SO2 is
removed).
• The peroxide bleed goes to RV1 where the H2 SO4 consumes NaOH.
The bleed of the peroxide scrubber consists of a 49 % H2 SO4 , which is directly linked
to the reagent H2 O2 with the same concentration. This can be compared to the 70­76
% used at the leaching plant. However, as the required H2 SO4 at the leaching plant
(2300 kg/h) is many times larger than the scrubber’s bleed (300 kg/h), the bleed can
be introduced to the leaching plant even though it has a lower concentration.
33
CHAPTER 4. RESULTS AND ANALYSIS
4.3 Tower design
The design parameters for the packed towers are listed in Table 4.3.1 and the design
is presented in Figure A.0.1 in Appendix A. The packing used is a 50 mm polypropene
Hiflow ® ring with a void fraction of 94 % and a surface area of 90 m2 /m3 . The flooding
factor is below the values found in the literature, but was chosen to have a reasonable
liquid load and operating pressure drop.
The design parameters for the open spray towers are also listed in Table 4.3.1 while the
design is presented in Figure A.0.2 in Appendix A. An L/G ratio at the upper end of the
values found in the literature was chosen for the open spray towers, as this was assessed
to be needed to reach similar removal efficiencies as in the packed towers.
Table 4.3.1: Tower design parameters for the packed towers and the open spray towers.
Packed tower
Packing
Liquid flow rate (m3 /h)
Volumetric L/G ratio (L/m3 )
Gas velocity (m/s)
Flooding factor (%)
Operational pressure drop (mbar/m)
Packed height (m)
Tank residence time (min)
Tank volume (m3 )
Tank height (m)
Open spray tower
Liquid flow rate (m3 /h)
Volumetric L/G ratio (L/m3 )
Gas velocity (m/s)
Residence time (s)
Height reactive section (m)
Tank residence time (min)
Tank volume (m3 )
Tank height (m)
Hiflow plastic 50­6
770
2.5
2.2
40
1.9
4
5
65
2.2
4000
12.7
3
3
9
2
130
5.1
The heights of the other tower parts are listed in Table 4.3.2 while the tower dimensions
are listed in Table 4.3.3. The latter table shows that even though the packed towers
have a slightly larger diameter, the open spray towers have a larger volume due to their
height.
34
CHAPTER 4. RESULTS AND ANALYSIS
Table 4.3.2: Heights of other tower parts (the same for both packed and open spray
towers).
Part of the tower
Height (m)
Distance top ­ liquid distributor
Distance liquid distributor ­ packing
Distance gas inlet ­ packing
Diameter gas inlet
Distance gas inlet ­ tank
2
0.5
3.75
2.5
0.3
Table 4.3.3: Tower dimensions of the packed tower and the open spray tower.
Packed tower
Diameter (m) 7.0
Height (m)
15.2
3
Volume (m ) 590
Open spray tower
6.1
23.1
680
4.4 Costs
The costs were based on the SO2 emissions during the period October 2019 ­ September
2020 for the flash tapping gases and the period 16/12/2020 ­ 21/12/2020 for the
secondary hood gases.
4.4.1 Operational costs
The reagent costs are listed in Table 4.4.1 and are as previously for the cases where only
sulfates are produced.
Table 4.4.1: Reagent costs for the different scrubber techniques.
Scrubber
Cost per year (SEK)
Caustic soda
Soda ash
Peroxide
Lime
Zinc oxide
4 190 000
2 790 000
2 150 000
690 000
0
For the sodium­based scrubbers, the actual reagent cost is higher as the sulfites that
are produced consume green liquor at RV1. If the whole bleed consists only of Na2 SO3
(190 kg/h) then 10 % of it would go to RV1 (90 % removed in the stripper) where it
consumes 50 kg/h green liquor to the cost of 40 SEK/h (350 000 SEK/year).
35
CHAPTER 4. RESULTS AND ANALYSIS
If the peroxide scrubber bleed would go to RV1 instead of going to the leaching plant
or the sulfuric acid plant, the costs would increase with 6 440 000 SEK/year due to the
use of caustic soda when H2 SO4 is neutralized at RV1.
In Table 4.4.2 it can be seen that the pump power is much higher for the open spray
towers due to a higher liquid flow rate, but also due to the extra power needed to spray
in the recirculating slurry (as opposed to the packed towers where the recirculating
liquid drains down from the liquid distributor). On the contrary the packed towers
have slightly higher fan power costs due to the higher pressure drop in the packed
towers. A cost of 150 000 SEK/year is added for the 39 kW heating in the soda ash
scrubber.
Table 4.4.2: Pump and fan power cost for the packed tower and open spray tower.
Pumps
Required shaft pump power (kW)
Electricity per year (kWh)
Cost per year (SEK)
Fan
Fan power (kW)
Electricity per year (kWh)
Cost per year (SEK)
Packed tower
Open spray tower
38
330 000
150 000
297
2 600 000
1 170 000
187
1 640 000
740 000
125
1 090 000
490 000
From Table 4.4.3 it can be concluded that the sodium­based scrubbers have the highest
operating costs, while the open spray towers (lime and zinc oxide) have the lowest
operating costs.
Table 4.4.3: Total operating costs for the different scrubber techniques.
Scrubber
Total operational costs per year (SEK)
Caustic soda
Soda ash
Peroxide
Lime
ZnO
5 080 000
3 830 000
3 040 000
2 350 000
1 660 000
4.4.2 Capital investment costs
In Table 4.4.4 it can be seen that the purchased equipment costs are higher for the open
spray towers due to a slightly larger tower volume and the larger pumps. This leads to
36
CHAPTER 4. RESULTS AND ANALYSIS
the open spray towers having higher total capital investment costs which are presented
in Table 4.4.5.
Table 4.4.4: Purchased equipment costs for the packed tower and the open spray tower.
Column shell
Internals
Pumps
Fan
Mixing/storage tank
Total purchased equipment
Packed tower costs (SEK)
Open spray tower costs (SEK)
9 630 000
1 240 000
330 000
770 000
120 000
12 090 000
10 180 000
570 000
1 190 000
770 000
120 000
12 830 000
Table 4.4.5: Total capital investment costs for the packed tower and the open spray
tower.
Packed tower
costs (SEK)
Purchased equipment
12 090 000
Delivery of purchased equipment
1 210 000
Installation of purchased equipment
7 250 000
Instrumentation and controls (installed) 4 790 000
Piping (installed)
9 040 000
Electrical systems (installed)
1 990 000
Buildings (including services)
2 390 000
Yard improvements
1 330 000
Service facilities (installed)
730 000
Engineering and supervision
4 390 000
Construction expenses
5 450 000
Legal expenses
530 000
Contractor’s fee
2 930 000
Contingency
5 850 000
Working capital
8 450 000
Total capital investment
68 420 000
Open spray tower
costs (SEK)
12 830 000
1 280 000
7 700 000
5 080 000
9 590 000
2 120 000
2 540 000
1 410 000
770 000
4 660 000
5 790 000
560 000
3 100 000
6 210 000
8 960 000
72 600 000
4.4.3 Life cycle costs and costs per unit reduction of SO2
emissions
The costs in Table 4.4.6 include the previously presented operating costs and capital
investment costs as well as the cost of capital. The results show that the sodium­
based scrubbers have the highest costs even if only sulfates were formed. The peroxide
scrubber is the second cheapest if the bleed goes to the leaching plant, while it is the
37
CHAPTER 4. RESULTS AND ANALYSIS
most expensive if the bleed goes to RV1. The same results for a 10 % WACC can be
found in Table B.0.1 in Appendix B.
Table 4.4.6: Life cycle costs and costs per unit reduction of SO2 emissions for a 6 %
WACC during the 10 year economic lifetime.
Caustic soda
Only sulfates formed
Only sulfites formed
Soda ash
Only sulfates formed
Only sulfites formed
Peroxide
To leaching plant
To RV1
Lime
Zinc oxide
Life cycle costs
(SEK)
Costs per unit reduction of SO2 emissions
(SEK/kg reduction of SO2 emissions)
138 000 000
141 000 000
21
22
125 000 000
129 000 000
19
20
117 000 000
182 000 000
118 000 000
109 000 000
18
28
18
17
4.5 Further analysis
In the further analysis, the use of process waters was evaluated for the different
scrubbers. A sensitivity analysis, including the minimum and maximum SO2 scenarios
as well as studying the costs to obtain different removal efficiencies, was conducted for
the soda ash and peroxide scrubbers which were considered to have the most promising
results.
4.5.1 Process water
If process waters are used as make­up water in the scrubbers, an increased reagent use
and cost is required to react with the SO2 in the process waters. As 90 % of the SO2
from the process waters would have been removed in the stripper, the green liquor use
at RV1 reduces with 10 % when the process waters are used in the scrubber instead.
This reduction does not occur if only sulfites are formed in the sodium­based scrubbers
as the sulfites also consume green liquor at RV1. Further, the flash cooling process
water contains H2 SO4 , which has to be neutralized in the alkaline scrubbers (all except
the peroxide scrubber in which H2 SO4 is the product). The costs associated with the
clinker scrubber process water and the flash cooling tower process water are presented
in Table 4.5.1 and 4.5.2 respectively. The results in the tables are based on the data in
38
CHAPTER 4. RESULTS AND ANALYSIS
Table 3.1.4 and represent the cases where only sulfates are formed in the scrubbers.
The results show that the costs for removing SO2 are high for many scrubbers, while
the costs are lower for neutralizing H2 SO4 in the scrubbers than at RV1. The latter is
as all the scrubbers have cheaper reagents than the caustic soda that is used at RV1 for
neutralizing H2 SO4 (expect the caustic soda scrubber).
Table 4.5.1: Costs of using clinker scrubber process water in the scrubber.
Scrubber
SO2 removal cost
(SEK/year)
Caustic soda 790 000
Soda ash
540 000
Peroxide
330 000
Lime
130 000
Zinc oxide
0
Green liquor cost
Total
at RV1 (SEK/year) (SEK/year)
­ 60 000
­ 60 000
­ 50 000
­ 60 000
­ 50 000
730 000
480 000
280 000
70 000
­ 50 000
Table 4.5.2: Cost of using flash cooling tower process water in the scrubber.
Scrubber
Caustic soda
Soda ash
Peroxide
Lime
Zinc oxide
SO2
removal
cost
(SEK/year)
3 210 000
2 210 000
1 340 000
530 000
0
Green liquor
cost at RV1
(SEK/year)
­ 240 000
­ 250 000
­ 200 000
­ 240 000
­ 220 000
H2 SO4
removal
cost
(SEK/year)
950 000
650 000
0
160 000
0
NaOH cost
at RV1
(SEK/year)
Total
(SEK/year)
­ 950 000
­ 980 000
­ 770 000
­ 960 000
­ 880 000
2 970 000
1 630 000
370 000
­ 510 000
­ 1 100 000
4.5.2 Minimum and maximum SO2 scenarios
The amount of Na2 CO3 added in the soda ash scrubber and 49 % H2 O2 added in
the peroxide scrubber in the minimum SO2 scenarios were 2.4 kg/h and 1.6 kg/h
respectively while they were 2370 kg/h and 1550 kg/h respectively in the maximum
SO2 scenarios. Tables 4.5.3 and 4.5.4 show the equilibrium temperatures, water
balances and bleed compositions for the soda ash scrubber and the peroxide scrubber.
In the maximum SO2 scenario, the soda ash scrubber has a bleed size that would imply
a significant load on RV1 while the H2 SO4 in the peroxide scrubber bleed has to be of
a higher concentration in the maximum SO2 scenario if it is to be used in the leaching
plant. The size of the peroxide bleed is also larger than what is used in the leaching
plant.
39
CHAPTER 4. RESULTS AND ANALYSIS
Table 4.5.3: The equilibrium temperatures and water balances in the minimum and
maximum SO2 scenarios for the soda ash scrubber and for the peroxide scrubber.
Soda ash
Peroxide
Minimum SO2 scenario
Equilibrium temperature (◦ C)
13
13
3
Evaporation (m /h)
0.6
0.6
Minimum bleed (m3 /h)
0.1
0.006
3
Minimum water added (m /h)* 0.7
0.6
Maximum SO2 scenario
Equilibrium temperature (◦ C)
31
39
3
Evaporation (m /h)
7.0
5.5
Minimum bleed (m3 /h)
8.9
3.1
3
Minimum water added (m /h)* 15.9
8.6
* Includes both the make­up water and water from the reagent solutions.
Table 4.5.4: Bleed composition in the minimum and maximum SO2 scenarios for
the soda ash scrubber and for the peroxide scrubber. The rows with sulfites/sulfates
represent the extreme cases were only sulfites or only sulfates were produced.
Minimum SO2 scenario Maximum SO2 scenario
mass flow rate (kg/h)
mass flow rate (kg/h)
Soda ash
H2 O (l)
NaOH (aq)
Na2 SO3 /Na2 SO4 (aq)
CO2 (aq)
Peroxide
H2 O (l)
H2 O2 (aq)
H2 SO4 (aq)
80
0.1
2.2/2.4
0.0001
8350
70
2210/2490
20
3
0.000006
0.2
2200
36
2080
4.5.3 Costs to obtain different removal efficiencies
Table 4.5.5: Cost differences between different removal efficiencies and the 90 %
removal efficiency for the soda ash scrubber and peroxide scrubber. The changes in
operational costs for the peroxide scrubber are in parenthesis in the last column.
Removal
efficiency (%)
Change in
packing height (m)
Change in
capital cost (SEK)
99
95
90
85
80
1.6
0.5
0
­ 0.3
­ 0.5
5 730 000
1 790 000
0
­ 1 070 000
­ 1 790 000
Change in soda ash
scrubber operational cost
(peroxide scrubber) (SEK)
280 000 (220 000)
160 000 (120 000)
0
­ 160 000 (­ 120 000)
­310 000 (­ 240 000)
40
CHAPTER 4. RESULTS AND ANALYSIS
At 91.2 % removal efficiency, the SO2 emissions are below the BAT­AEL (350 mg
SO2 /m3 ) at all times, including the maximum SO2 scenario used in this report. Table
4.5.5 shows the cost differences for various removal efficiencies for the soda ash and
peroxide scrubber. For removal efficiencies of 95 % and 99 % the costs increase with
2 % and 6 % respectively compared to the 90 % removal efficiency, while for removal
efficiencies of 85 % and 80 % the costs decrease with 1 % and 2 % respectively compared
to the 90 % removal efficiency.
41
Chapter 5
Discussion
As the flash tapping occurs intermittently, part of the study was to investigate if the
secondary hood gases could be taken into the scrubber to make use of the remaining
capacity of the scrubber when flash tapping does not occur. The results show that the
potential is high as at least 280 tonnes of SO2 /year could be cleaned from the secondary
hood gases, which is also a significant amount in relation to the SO2 from the flash
tapping gases. As emissions regulations have been and keep getting more stringent,
Rönnskär has a possibility of creating an even bigger margin to the 3500 tonnes of SO2
emissions/year permit. And as previously stated, taking in secondary hood gases into
the scrubber improves the working environment in the converter aisle. However, the
use of secondary hood gases in the scrubber would imply an increased cost for piping
as the converters are further away from the location of the proposed scrubber than the
flash furnace. Controlling the two secondary hoods in operation in addition to the flash
tapping would also make process control more complex. The two latter issues have not
been investigated in this report.
The sodium­based scrubbers have the highest operational costs, which could be even
higher depending on the share of Na2 SO3 in the bleeds. The simplicity of the sodium­
based scrubbers could still make the soda ash, which is the cheaper of the two, worth
considering. As the incoming gas streams have approximately the same concentrations
of oxygen as ambient air, the maximum amount of Na2 SO3 is oxidized to Na2 SO4 in
the scrubbers. To get a more precise cost estimate of the sodium­based scrubbers the
equilibrium between sulfites and sulfates needs to be studied.
42
CHAPTER 5. DISCUSSION
The peroxide scrubber is compatible with the flash cooling tower process water as it
does not need to neutralize the H2 SO4 as in the other scrubbers. This is an advantage
as it removes the neutralization cost of the flash cooling tower process water at RV1
without any added reagent cost. The peroxide scrubber also works well with the
fluctuating flow of the incoming gas streams. The bleed from the peroxide scrubber was
a H2 SO4 solution between 0.006­3.1 m3 /h. In the minimum and base case scenarios
(0.23 m3 /h) the contribution was small compared to the total 70­76 % H2 SO4 used in
the leaching plant (1.3 m3 /h on average) making the exact concentration of the bleed
less important. However, in the maximum scenario the bleed size is larger (3.1 m3 /h)
than the average need at the leaching plant which would require it to have the same
concentration as the H2 SO4 used in the leaching plant. The excess bleed would also
require intermediate storage or to be taken to the sulfuric acid plant. Another issue
is that the leaching plant is only expected to operate 340 days/year during which the
bleed has to be taken somewhere else. Due to the high cost of neutralizing the H2 SO4 in
the peroxide scrubber in RV1, mixing the peroxide bleed into the sulfuric acid plant or
pausing the use of the scrubber should be the alternative solutions considered before
taking it to RV1. Yet another issue is that the peroxide bleed is the smallest among
the scrubbers, meaning that it has the biggest risk of accumulating impurities from the
various gas and water streams. If the peroxide bleed is to be used at the leaching plant,
this would require a new pipeline to transport the H2 SO4 from the scrubber. The two
latter issues have not been investigated in this report.
The scrubbers operate at relatively low temperatures (13­39 ◦ C) which is positive for
the solubility of SO2 in water. However, the temperatures are far from the optimal
temperature range for the lime scrubber. Other disadvantages for the lime scrubber
are the fluctuating flow rates and SO2 concentrations of the incoming gas streams and
that Rönnskär preferred not to produce non­regenerative solid scrubbing products.
Considering the disadvantages and the fact that the life cycle costs for the lime scrubber
are higher than for the peroxide scrubber and only slighter lower than for the soda ash
scrubber, it is unlikely that the lime scrubber is to be selected.
The zinc oxide scrubber process is interesting as it has no reagent costs, which also gives
the scrubber the lowest life cycle costs. The cost of treating the ZnSO3 and ZnSO4 before
reintroducing it to the zinc smelting process is unknown and should be estimated if the
zinc oxide technique is chosen for the proposed scrubber. It is also unknown which of
the insoluble ZnSO3 or the soluble ZnSO4 is preferred as a product from the scrubber,
43
CHAPTER 5. DISCUSSION
as this depends on the cost and feasibility of the treatment methods. There is limited
literature regarding the zinc oxide scrubber, but as previously mentioned, there are
a few zinc oxide FGD systems in operation from who Rönnskär could learn more.
Especially regarding the issue of ZnO having low solubility in water.
The bleed streams are all approximately 1 m3 /h in the base case scenarios which is a
minor contribution when comparing it to the 160 m3 /h that RV1 handles on average
and the 200 m3 /h that it is designed for. The maximum bleed values of, for instance,
the soda ash scrubber (8.9 m3 /h) would however have a considerable impact. The
use of process waters in the scrubbers has the advantage that the water load to RV1
decreases due to the evaporation in the scrubbers. The net reduction is even bigger if
the bleed does not go to RV1 at all which is the case for the peroxide scrubber.
The SO2 content in the process water leads to an increase in the reagent consumption
and cost. In the meanwhile, the stripper currently removes 90 % of the SO2 that enters
it. Therefore using process water in the scrubber leads to increased costs for Rönnskär
as most of the SO2 would have been removed in the stripper for a much lower cost.
Therefore, it is recommended to install a new stripper before the scrubber to remove
the SO2 before it enters the scrubber to reduce reagent costs. On the other hand,
neutralizing the H2 SO4 of the flash cooling tower process water in the scrubber instead
of in RV1 leads to cost reductions as all scrubber reagents (except NaOH) are cheaper
than the NaOH used for neutralizing H2 SO4 in RV1. This makes the flash cooling tower
process water suitable for the proposed scrubber. However, it should be noted that
the flash cooling tower only provides the scrubber with approximately 3 m3 process
water/h, which would not be enough, for example, in the maximum cases where a
larger amount of water is needed in the scrubber. In these cases, the process water
would have to be used together with other water. Another process water that was not
studied in this report is the flash scrubber process water, which contains less SO2 and
more H2 SO4 than the flash cooling tower process water. From what has been learnt
in this report, this process water could be interesting to study further to achieve even
more cost reductions, especially with the peroxide scrubber.
Based on the reasoning above, the clinker scrubber process water, that has a pH in
the alkaline range, should not be used in the scrubber. Instead, it should go straight
towards RV1 where it contributes in neutralizing the other process waters that are
mainly acidic. The clinker scrubber process water is also high in fluorides and has
44
CHAPTER 5. DISCUSSION
some chlorides, which means that there would be a risk of scaling and corrosion in the
scrubber. Another disadvantage that the clinker scrubber process water has, is that the
clinker furnace is located further away from the location of the proposed scrubber than
the flash cooling tower, which would lead to higher piping costs.
A removal efficiency of 91.2 % would be needed to have emissions below the BAT­AEL
at all times, including the maximum SO2 scenario. As the BAT­AEL is stated as a daily
average, Rönnskär has to evaluate how often the SO2 emissions will be close to the
values in the maximum SO2 scenario in the future. Based on the data used in this
report a 90 % removal efficiency would be enough to comply with the BAT­AEL, but it
has to be monitored if this will be the case in the future as well.
45
Chapter 6
Conclusions and future work
The aim of the project was to evaluate various scrubber techniques to remove SO2 from
the flash tapping gases at Boliden Rönnskär. First, it was concluded that including the
secondary hood gases in the scrubber was a suitable alternative to make use of the
scrubber’s capacity and to reduce SO2 emissions further.
Among the scrubbers, the peroxide scrubber evolved as an attractive alternative due
to its suitability with the leaching plant and the flash cooling tower process water. It
also had the lowest life cycle costs of the packed towers. The sodium­based scrubbers
were the most expensive mainly due to their high reagent costs. However, due to the
simplicity of the sodium­based scrubbers, the soda ash scrubber might still be worth
considering.
The open spray towers had the lowest operational costs, due to their low reagent costs,
but also the highest capital costs. The lime scrubber had several disadvantages that
made it unsuitable for the process gases and waters in question, while the zinc oxide
scrubber had the issue of zinc oxide’s low solubility in water. As the zinc oxide scrubber
had the lowest life cycle cost and could be integrated with the zinc smelting process, it
should be investigated further.
The use of process water in the scrubbers lead to a net reduction of process water to RV1
due to evaporation in the scrubbers. Further, it was cheaper (equally expensive for the
caustic soda scrubber) to neutralize H2 SO4 in the scrubbers than in RV1, while it was
more expensive to take care of SO2 in the scrubber compared to stripping. Therefore it
is recommended to use the flash cooling tower process water in the scrubber together
with a stripper unit before the scrubber to minimize the reagent costs.
46
CHAPTER 6. CONCLUSIONS AND FUTURE WORK
For future work it is proposed to:
• Perform a complete analysis of metals in the process gases and waters to evaluate
the risk of scaling and corrosion.
• Study the flash scrubber process water as it could reduce reagent costs even more
than the flash cooling tower process water.
• Study the process control of the different process gases and waters.
• Study the sulfite/sulfate equilibrium to obtain more precise operational costs, if
any of the sodium­based scrubbers or the zinc oxide scrubber is preferred.
47
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52
Appendix A
Tower design figures
Figure A.0.1: Packed tower design.
Figure A.0.1 and Figure A.0.2 show the tower designs of the packed towers and open
spray towers respectively. The figures of both towers are drawn to scale, while the
equipment outside the tower (pumps, fans etc.) are not.
53
APPENDIX A. TOWER DESIGN FIGURES
Figure A.0.2: Open spray tower design.
54
Appendix B
Life cycle costs and costs per unit
reduction of SO2 emissions (10 %
WACC)
Table B.0.1: Life cycle costs and costs per unit reduction of SO2 emissions for a 10 %
WACC during the 10 year economic lifetime.
Reagent
Caustic soda
Only sulfates formed
Only sulfites formed
Soda ash
Only sulfates formed
Only sulfites formed
Peroxide
To leaching plant
To RV1
Lime
Zinc oxide
Life cycle costs
(SEK)
Costs per unit reduction of SO2 emissions
(SEK/kg reduction of SO2 emissions)
150 000 000
154 000 000
23
24
138 000 000
141 000 000
21
22
130 000 000
194 000 000
131 000 000
122 000 000
20
30
20
19
55