Techno-economical comparison of evaporation

and/or membrane filtration

Study for the concentration of sulphite bleaching plant effluent

Marta de Malhão Lemos Ferreira

Thesis to obtain the Master of Science Degree in

Chemical Engineering

Supervisors: Professor Maria Norberta Neves Correia de Pinho (IST)

Professor Frank Lipnizki (LTH)

Examination Committee

Chairperson: Professor Henrique Matos (IST)

Supervisor: Professor Frank Lipnizki (LTH)

Member of the Committee: Professor Vítor Geraldes

November 2018

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i

Acknowledgments

I would like to express my enormous gratitude to Professora Maria Norberta de Pinho and to Professor

Frank Lipnizki for the opportunity of studying with one of the best Membrane research group. All the

knowledge given allowed me to grow immensely from a professional and personal viewpoint.

I am extremely grateful for the significant role that Gregor Rudolph and Johan Thuvander had through

my entire Thesis’ journey. Their guidance and support were crucial to the development of the report

and to the development of myself as a future Engineer. Additionally, my routine in Sweden wouldn’t

have been as cheerful without Catarina Oliveira, who have brightened my daily life the past six months.

Other great people made my stay in Lund incredible and unforgettable, so a big thanks to Mikael and

Ximo and Maria for your role in my life and for always having such a positive energy, to Miguel and

Lidia for all your support, to Anton and Johan for your kindness and Simon for your friendship.

To my friends in Portugal, the ones I met since ever and the ones I met in Técnico, I own you a great

thank you for having been always an example of strength and brilliance as colleagues and friends.

Particularly, Margarida, Marta, Rita and Mada, I am truly thankful for having you in my life.

My sincere gratitude for all the unconditional love and support given by my family, specially my mum,

my dad and my brother Afonso. Unquestionably, I must also express my gratitude to my grandma avó

Didi for always believing in me and to my grandpa avô Beto for being the role model of my life and for

being the reason why I chose to become an Engineer.

Lastly, I am forever grateful for having met the most incredible person ever. Miguel, you have fulfilled

my life in so many aspects and you have been the fuel that allowed to reach this point in the most

happy and flawless way possible. So, thank you for yesterday, for today and for tomorrow.

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Abstract

Evaporation plants have already been largely implemented in the pulp and paper industry (PPI) as

concentration units. On the other hand, the implementation of membrane filtration systems in the PPI

has been one of the main reasons to the significant low footprint and relatively decrease of energy

consumption. Therefore, a feasibility study concerning the integration of a membrane filtration process

as part of a concentration unit with the purpose of treating a bleaching effluent is the main subject of this

Thesis.

A comparison analysis is performed, from which both processes are examined as stand-alone units

along with an investigation regarding the integration as a combined process of both technologies, i.e.,

membrane filtration system followed by an evaporation unit.

Succinctly, concerning a sulphite bleaching plant effluent, a membrane filtration system configuration of

a housing containing three spiral-wound elements in series distributed in parallel is the most

economically viable option while a five multiple-effect evaporator (MEE) system was found to be the

most feasible option regarding the evaporation unit. Based on those configurations, for a general case

study, the economic optimal conditions for the combined concentration process require a retentate

concentration of 5,6% (w/w), which in turn corresponds to a volume reduction of 60%, along with an

evaporation final product concentration of 20% (w/w) both measured in a total solid (TS) content unit.

The complete system leads to a global volume reduction of 88%. A particular industrial case study was

studied, regarding a pulp mill located in Caima, Portugal, that has already implemented a 5-MEE system

for the treatment of of sulphite bleaching plant effluent. For that case, the installation of a membrane

filtration unit prior to the evaporation would be only economically favorable when concentrating the

retentate until 3,1% (w/w) and a volume reduction of 10% and enabling a evaporation final product

concentration of 20% (w/w) and a global volume reduction of 87%.

Key-words: Sulphite bleach plant effluent, Concentration methods, Spiral wound nanofiltration,

Evaporation, CAPEX, OPEX, TAC.

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Resumo

A implementação de membranas como unidades de separação e concentração na indústria da pasta

de papel tem promovido significativamente tanto a redução de utilização de água fresca como a

diminuição do consumo de energia. Por outro lado, unidades de evaporação capazes de desempenhar

semelhantes funções, maioritariamente, caracterizam-se por apresentarem vantagens económicas

superiores. Assim sendo, esclarece-se a necessidade de realizar uma análise económica comparando

os dois processos de concentração, com o intuito de prever qual o sistema e respetivas condições

operatórias que conduzem ao processo de concentração mais viável a nível económico e sustentável.

Tendo por base o estudo de um efluente de um processo de branqueamento de pasta de papel

proveniente de uma fábrica em Caima, Portugal, analisa-se paralelamente a implementação de um

sistema de membranas e de uma instalação de uma unidade de evaporação em conjunto com um

estudo relativo à integração de ambos processos como uma só unidade de concentração.

Adicionalmente, executa-se uma análise de sensibilidade de forma a compreender quais os parâmetros

preponderantes nos resultados económicos finais, medidos através das seguintes métricas, CAPEX,

OPEX e TAC.

Relativamente ao sistema de filtração por membranas, conclui-se que se deveria optar por um módulo

em espiral, cuja implementação deveria ser feita num sistema em cascata, agrupando três módulos em

série distribuídos em paralelo. Quanto à unidade de evaporação, esta deve ser constituída por um

sistema de 5 evaporadores conectados entre si por um sistema de evaporação de multi-efeito. Como

tal, definida a melhor configuração para cada tecnologia, o processo integrado, no qual se implementa

um sistema de membranas seguido por um sistema de evaporação, apresenta-se como sendo o

sistema mais viável. Para um caso geral, as condições operatórias mais favoráveis exigem que o

sistema de membrana concentre a corrente de concentrado até 5,6% (w/w) em sólidos totais, o

equivalente a um volume de redução de 60%, seguido por uma evaporação na qual o efluente é

concentrado até 20% (w/w) em sólidos totais, permitindo um volume de redução total de 87%. Por fim,

analisou-se o caso concreto da fábrica situada em Caima, a qual tem já instalado um sistema de

evaporação, concluindo-se que idealmente o sistema de filtração, a ser implementado como parte

integrante da unidade de concentração, deveria concentrar até 3,1% (w/w) em sólidos totais, o

correspondente a um volume de redução de 10%, seguido por uma evaporação capaz de concentrar

até 20% (w/w), promovendo um volume de redução total de 87%.

Palavras-chave: Efluente de branqueamento, Processos de concentração, Nanofiltração em módulo

espiral, Evaporação, CAPEX, OPEX, TAC

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vii

Table of Content

Chapter 1. Introduction ......................................................................................................................... 1

1.1. Context and Motivation ............................................................................................................ 1

1.2. Objectives ................................................................................................................................ 2

1.3. Thesis Outline .......................................................................................................................... 2

Chapter 2. Literature Review ........................................................................................................... 3

Pulp and paper industry ........................................................................................................... 3

2.1.1 Pulp and paper market ........................................................................................................ 3

Concentration Methods ........................................................................................................... 8

2.2.1 Membrane Filtration ............................................................................................................. 8

2.2.2 Evaporation ........................................................................................................................ 19

The method of least squares .................................................... Erro! Marcador não definido.

Concentration Methods in the Pulp and Paper Industry ........................................................ 27

Chapter 3. Techno-economical results ........................................................................................ 29

Techno-economical analysis outline ..................................................................................... 29

Membrane filtration system configuration analysis................................................................ 31

Evaporation system configuration analysis ........................................................................... 32

Membrane filtration and/or evaporation analysis................................................................... 34

3.3.1 Membrane filtration or evaporation .................................................................................... 34

3.3.2 Combined process or evaporation..................................................................................... 37

3.3.3 Process system price summary ......................................................................................... 40

Techno-economical models ................................................................................................... 41

Sensitivity Analysis ................................................................................................................ 42

Pulp and paper industry case study: Caima, Indústria da Celulose ...................................... 44

Techno-economic results discussion .................................................................................... 46

Chapter 4. Conclusion and future work ....................................................................................... 49

Chapter 5. References ................................................................................................................... 51

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ix

List of Acronyms and Nomenclature

Acronyms

BPE Bleach plant effluent

BPR Boiling point rise

CAPEX Capital expenditure

CC Convergence criteria

CIP Cleaning in place

CFV Cross-flow velocity

E-stage Alkaline extraction stage

EVAP Evaporation

MF Membrane filtration

NF Nanofiltration

OPEX Operational expenditure

PEC Purchase of delivered-equipment cost

PPI Pulp and paper industry

RO Reverse osmosis

SC Steam consumption

SE Steam economy

SWM Spiral-wound membranes

TAC Total annual costs

TCF Total chlorine free

TMP Transmembrane pressure

TS Total solids

UF Ultrafiltration

VR Volume reduction

VRF Volume reduction factor

x

Nomenclature

𝑨𝒊 Heat transfer area of an evaporator

𝑨𝒎 Membrane area

𝑪 Cost

𝑪𝑨𝒏𝒏𝒖𝒂𝒍 𝒓 Annual membrane replacement cost

𝑪𝒄𝒂𝒑 Capital cost

𝑪𝒃 Concentration of the solute in the bulk solution

𝑪𝑰𝑪 Instruments and controls cost

𝑪𝒎 Concentration of the solute in the membrane surface

𝑪𝑴𝑰 Miscellaneous equipment cost

𝑪𝒎𝒆𝒎 Membrane equipment cost

𝑪𝒑 Concentration of the solute in the permeate

𝑪𝑷𝑽 Pipes and valves cost

𝑪𝒓 Membrane replacement cost

𝑪𝒓𝒆𝒇 Reference cost

𝑪𝑻𝑭 Tanks and frames cost

𝒇𝒊 Multiplying factors

𝒉𝒍𝒊 Effluent enthalpy

𝒉𝒗𝒊 Water enthalpy

𝒉𝒗𝒊𝒗 Saturated steam enthalpy

∆𝒉𝒗𝒂𝒑𝒊 Saturated steam heat of vaporization

𝒌 Capacity

𝒌𝒓𝒆𝒇 Reference capacity

𝒊 Interest rate

𝑱 Permeate flux

𝑳𝒎 Lifespan of the membrane

𝑳𝒑 Lifespan of the plant

𝑳𝑷 Membrane permeability coefficient

𝑴𝒊 Mass flow in each evaporator

xi

𝑴𝒓 Sets of replacement pf membrane

𝑵 Number of effects in an evaporation unit

ℕ Investment period

𝒏 Cost capacity factor

𝑷𝒊𝒏𝒍𝒆𝒕 Inlet pressure to the plant

𝒑𝒔 Life steam pressure

∆𝑷 Frictional pressure drop

𝑸𝒉𝒐𝒖𝒔𝒊𝒏𝒈 Feed flow

𝒒𝒊 Heat transfer in each evaporator

𝑹𝒎 Intrinsic membrane resistance

𝑹𝒐𝒃𝒔 Observed retention coefficient

𝑹𝒕𝒓𝒖𝒆 True retention coefficient

𝑻𝒊 Evaporator temperature

𝑻𝒍𝒊 Liquor temperature

𝑻𝒗𝒊 Vapor temperature

𝑻𝒔 Life steam temperature

𝑼𝒊 Heat transfer coefficient

𝑾𝒇𝒆𝒆𝒅 Feed pump power

𝑾𝒍𝒊 Liquid mass flow in each evaporator

𝑾𝒓𝒆𝒄𝒊𝒓 Recirculation pump power

𝑾𝒗𝒊 Vapor mass flow in each evaporator

𝜼 Pump efficiency

𝝁𝒑 Permeate viscosity

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List of Figures

Figure 2-1. Pulp production by Worldwide region and by European country in 2016. Adapted from (CEPI

(Confederation of European Paper Industries), 2017). ........................................................................... 3

Figure 2-2. Three-dimension arrangement of cellulose, hemicellulose and lignin in a cell wall. (Energy

& Ise, 2017) ............................................................................................................................................. 5

Figure 2-3. Schematic illustration of the bleaching process of the pulp mill from Caima - Indústria da

Celulose, S.A., high lightening the effluent from which the study is going to be conducted. .................. 7

Figure 2-4. Schematic diagrams of the principal types of membranes: a) symmetrical microporous

membrane; b) asymmetric microporous membrane; c) thin film composite membrane. Adapted from

(Baker, 2004). .......................................................................................................................................... 9

Figure 2-5. Membrane filtration processes according to the average pore size of the particles retained

by the membrane (Epa, 2005). ................................................................................................................ 9

Figure 2-6. Membrane systems according to the flow process: a) dead-end filtration; b) cross-flow

filtration. ................................................................................................................................................. 10

Figure 2-7. Plate and frame module design. Adapted from (W. Baker, 2012) ...................................... 11

Figure 2-8. Spiral wound module design. Adapted from (W. Baker, 2012). .......................................... 11

Figure 2-9. Tubular module design. Adapted from (W. Baker, 2012) ................................................... 12

Figure 2-10. Hollow fiber membrane module design. Adapted from (W. Baker, 2012) ........................ 12

Figure 2-11. Example of a cascade plant design with retentate recycle composed by three stages in

which the first two the housings are in parallel. Addapted from (Nilsson, Lipnizki, Trägårdh, & Östergren,

2008). ..................................................................................................................................................... 15

Figure 2-12. Evaporators models: a) rising film tubular, b) falling film tubular and c) mechanical vapor

recompression evaporator. Adapted from (SPX Corporation, 2008). ................................................... 20

Figure 2-13. Block diagram for the 𝑖th effect. Adapted from (Kumar, Kumar, & Singh, 2013).............. 22

Figure 2-14. Quadruple-effect evaporator system parallel flow case. Adapted from (Kaya & Ibrahim

Sarac, 2007). ......................................................................................................................................... 22

Figure 2-15. Iterative method performed to calculate the steam economy (SE) and the steam

consumption (SC). ................................................................................................................................. 24

Figure 2-16. Installed cost of evaporators according to heat transfer area. Adapted from (S. M. Peters

& Timmerhaus, 1991) ............................................................................................................................ 25

Figure 3-1. Recirculation pump cost according to each case study. ..................................................... 32

Figure 3-2. Evaporation concentration process OPEX for different MEE systems according to each case

study. ..................................................................................................................................................... 33

Figure 3-3. Evaporation concentration process CAPEX for different MEE systems according to each

case study. ............................................................................................................................................ 33

Figure 3-4. Evaporation concentration process TAC for different MEE systems according to each case

study. ..................................................................................................................................................... 34

Figure 3-5. MF and EVAP CAPEX according to each case study. ....................................................... 35

Figure 3-6. MF and EVAP OPEX according to each case study. ......................................................... 35

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Figure 3-7. MF and EVAP TAC according to each case study. ............................................................ 36

Figure 3-8. Energy cost distribution percentage for MF and EVAP according to each case study. ..... 36

Figure 3-9. MF → EVAP and EVAP CAPEX according to each case study. ........................................ 38

Figure 3-10. MF → EVAP and EVAP OPEX according to each case study. ........................................ 38

Figure 3-11. MF → EVAP and EVAP TAC according to each case study. ........................................... 39

Figure 3-12. Energy cost distribution percentage for MF → EVAP and EVAP according to each case

study. ..................................................................................................................................................... 39

Figure 3-13. Process design parameters distribution of CAPEX (right) and OPEX (left) parameters for

solely MF, case study VI- α. .................................................................................................................. 42

Figure 3-14. Process design parameters distribution of CAPEX (right) and OPEX (left) parameters for

solely EVAP, case study VI- α. .............................................................................................................. 42

Figure 3-15. Process design influence regarding MF in TAC for MF →EVAP, case study VI- α.......... 43

Figure 3-16. Process design parameters influence regarding EVAP in TAC for MF → EVAP, case study

VI - α ...................................................................................................................................................... 43

Figure 3-17. MF → EVAP and EVAP TAC according to each case study for the PPI case study example.

............................................................................................................................................................... 44

Figure 3-19. Process design parameters influence regarding MF in TAC for MF → EVAP, case study I

– α for PPI case study example ............................................................................................................. 45

xiv

List of tables Table 2-1. Comparison between several membrane modules. Hollow fine fiber stands for capillary

membrane module. Ceramic membranes are grouped separately from polymeric membranes because

their preparation methods are significantly different. Adapted from (Wagner, 2001). .......................... 13

Table 2-2. Assumptions used based on literature data for the cost estimation .................................... 18

Table 2-3. Percentage of the capital cost parameters related to the PEC, adapted from (S. M. Peters &

Timmerhaus, 1991). .............................................................................................................................. 26

Table 3-1. Membrane filtration module characteristics ......................................................................... 29

Table 3-2. Experimental data used in the techno-economic analysis .................................................. 29

Table 3-3. Nomenclature related to each case study concerning only one concentration method,

membrane filtration or evaporation ........................................................................................................ 30

Table 3-4. Nomenclature related to each case study concerning both concentration methods,

membrane filtration followed by evaporation ......................................................................................... 31

Table 3-5. TAC per cubic meter of removed water ............................................................................... 40

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1

Chapter 1. Introduction

1.1. Context and Motivation

Nowadays the industry is focusing its attention on shifting towards more sustainable production methods

capable of producing new value-added products, one example of that approach is the pulp and paper

industry sector. Despite being considered as one of the largest consumer of water in a global scale

(Bajpai, 2017b), it is also designated as being an industry where large improvements have been

accomplished regarding the treatment of industrial effluent (Bai, Xiong, & Wang, 2017).

Nonetheless, achieving an even more sustainable and economic viable treatment of sulphite bleaching

plant effluent according to the production vision aforementioned is feasible. Furthermore, it can lead to

the valorization of by-products such as lignin and hemicelluloses. In fact, concentration methods as the

ones analyzed in this report are capable of separating those compounds from the effluent. This, in turn,

contributes to a possible transition to a bio-based economy (Obydenkova, Kouris, Hensen, Heeres, &

Boot, 2017).

Increasingly strict environment polices have massively triggered the capability of effluent treatment

technologies capable to meeting environmental regulations while being cost effective and process

reliable. One solution is the zero-liquid discharge (ZLD) concept. It is a process where technologies that

can concentrate a liquid discharge up to near saturation level. With regard to the above mentioned,

evaporation process and membrane filtration process are within the most competitive existing markets

(Schwantes, Chavan, Winter, Felsmann, & Pfafferott, 2018).

One additional advantage of treating an effluent consists in being an imperative component in a

sustainable water management (Ebrahimi et al., 2015) which has been a worldwide major concern.

(O’Connell, 2017) The extension of water management challenges relies mostly on the global demand

for water as function of population growth, especially in emerging economies along with the

intensification of the water cycle due to climate change (WWAP, 2018).

Therefore, the present study explores the most suitable method of concentrating a sulphite pulp

bleaching effluent from a techno-economical viewpoint, ensuring a solution to the constant struggle of

having a process that provides simultaneously a sustainable and economically viable option.

2

1.2. Objectives

The focus of the present project is to compare two concentration methods, evaporation and membrane

filtration, applied to concentrate a sulphite pulp mill bleaching effluent (BPE) from an alkaline extraction

stage. The primary comparison is sustained by results that follow theoretic established laws along with

published experimental data related to each separation method. The secondary comparison relies on

the real case in which an evaporation plant is already integrated in the treatment of an effluent from a

pulp mill in Portugal operated by a company named Caima – Indústria da Celulose, S.A.

The performance of evaporation and membrane filtration as stand-alone systems is analyzed, as well

as the performance of both processes integrated. Furthermore, the best method and operation

conditions for concentrating the BPE is evaluated. The influence of certain operational parameters is

evaluated from a techno-economical perspective and represents the main points of this Thesis.

1.3. Thesis Outline

The Thesis is organized in the following sections, Chapter 2 contains a literature review about pulp and

paper industry along with the theoretical background concerning membrane filtration and evaporation.

Chapter 3 presents the methodology of the techno-economic results according to each system under

study, followed by a results discussion. Finally, in Chapter 4 the conclusion is expressed regarding the

selection of the most advantageous system according to the results from the previous chapter as well

as suggestions for future research directions.

3

Chapter 2. Literature Review

Pulp and paper industry

2.1.1 Pulp and paper market

Pulp and paper industry is facing innumerous and often digitalization-driven challenges. However, the

opportunities that certain steady-growing sectors are providing nowadays reveal the importance of

continuously investing in this sector (Carminati, 2017). Although a clear decline of the graphic-paper

market (newsprint, printing and writing paper) is visible, the growth of the pulp and paper industry has

recently been strongly driven by a rising demand for packing. That fact is boosted by the increase of e-

commerce and retail activities, along with the growth of tissue papers and hygienic products segments.

Despite being a comparatively small market today, pulp for textile application is growing as well (Berg,

2017) (Schaefer, 2016). Nonetheless, weaknesses of this segment have to be pointed out, related to

feedstock prices, i.e., the pulp prices, and to the ability to support significant investment costs to ensure

future growth (Livinec, 2018).

While China is the world’s leading country when it comes to paper and paper production, the pulp and

paper market is dominated by companies from North America, northern Europe and Japan. Concretely

International Paper and Kimberly-Clark from the United States, Stora Enso and UPM-Kymmene from

Finland, and Oji Paper and Nippon Paper Group from Japan (Statista, 2017). A global picture of the

market location is illustrated through Figure 2-1.

Figure 2-1. Pulp production by Worldwide region and by European country in 2016. Adapted from (CEPI

(Confederation of European Paper Industries), 2017).

4

Growth prospects can vary significantly among segments and regions, a phenomenon that is highly

influenced by demand trends, which in turn, can be translated into industry profitability. Another critical

factor is the continuous pursuit of value creation, which is enabling the shift in final product demand

modifying the pulp and paper structure. (O’Connell, 2017).

Lignocellulosic biomass is considered to be the most abundant feedstock on earth. In fact, an important

source of fractionated lignocellulosic is the waste stream from wood pulp and paper industry (Alexandri

et al., 2016). Currently, the industry is witnessing an outbreak of research activities to develop new

biobased-products from lignocellulosic biomass. In fact, new processes are being designed and

developed to extract hemicelluloses and lignin which are afterwards used as feedstock for a large range

of applications. The challenges are as countless as the accomplishments from finding new ways to

enlarge the product range of wood-based products in a more sustainable global economy (Berg, 2017).

One major and significant example is the concept of a pulp mill integrated in a biorefinery to reduce the

dependence on fossil resources and at the same time improving the economic sustainability (Marques,

Evtuguin, Magina, Amado, & Prates, 2009).

The industry is extremely capital and energy intensive mainly due to the involved processes of producing

pulp along with the industrial techniques required for the waste water treatments. Pulp is in fact the most

important raw material of the majority of the final products associated to the pulp and paper market,

(Ebrahimi et al., 2015). As a starting point of the process, the forest resource wood contains the following

main components: cellulose, hemicellulose and lignin.

Cellulose

Cellulose is the most abundant renewable material in nature and is characterized as a natural high

molecular polymer composed of glucose monomers, with cellobiose as the basic coupling unit. It is the

component responsible for the cell wall structure whereas lignin is the element that increases the

hardness of the cell wall (Chen, 2014), (Bonnin, Ralet, Thibault, & Schols, 2009).

Hemicelluloses

Lignocellulosic biomass provides essential raw materials for innumerous sectors and one of its

constituent is hemicellulose, the second most abundant renewable biopolymer found in nature.

Hemicelluloses are heterogeneous polysaccharides composed by several linear and branched

heteropolymers. Those heterpolymeres are bound to the cellulose fibrils trough hydrogen bonding, and

by that providing flexibility to the lignocellulosic structure (Phitsuwan, Sakka, & Ratanakhanokchai,

2013).

Hemicelluloses can be converted into bioethanol and other value added products, such as xylitol and

lactic acid, both having important applications in pharmaceutical and food industries (Menon, Rao, &

Prakash, 2010). Hemicelluloses are also used as hydrogels and as barrier films together with paper

additives. Thus, in essence it is a component with a vast range of use, justifying the significant market

value as well as the development and improvement of technologies capable of extracting it (Jönsson,

Nordin, & Wallberg, 2008).

5

Lignin

The PPI has lignin has an established byproduct. Howsoever, even if pulp mills are currently the world’s

largest producers of it they have benefited little from value-added lignin products. Despite the enormous

advances in lignin valorization research, the heterogeneous nature of lignin increases the difficulty of

isolation and in standardization which in turn affects the reproducibility and consistency of a lignin-based

product (Cline & Smith, 2017).

From a chemical point of view, basic units of lignin are phenylpropanoid derivatives that are combined

into high molecular networks by ether or carbon bonds. In an alkaline medium, the cleavage of the main

ether bonds leads to fragmentation and partial dissolution of the macromolecule (Chen, 2014).

In fact, research has found several applications for lignin in which its heterogeneous structure does not

represent a restriction. Currently lignin products can be segment in the following categories: binding

agents, emulsion stabilizers and dispersant agents. (Jönsson et al., 2008) Furthermore, lignin is used

for the production of the synthetic flavoring agent vanillin, (Cline & Smith, 2017) and for the production

of carbon fibers (Bajpai, 2013).

Figure 2-2. Three-dimension arrangement of cellulose, hemicellulose and lignin in a cell wall. (Energy & Ise,

2017)

Pulping process

Pulping is the process step after handling and debarking the wood. It has the purpose of breaking down

the structure of the fiber feedstocks into its constituent fibers. Chemical pulping is the process that is of

interest in this project. According to that process, the fibers are released from the wood matrix with the

use of chemicals in the presence of heat and pressure until the reaction reaches a certain degree of

delignification. The degree of delignification is one of the most important parameter for defining pulp

quality. Usually, it is expressed as the kappa number, which is directly related to the amount of lignin

present (Johssen, 2017).

6

The sulphate, or Kraft process, is the most common chemical pulping process. The Kraft process is

based on an alkaline solution of sodium hydroxide and sodium sulfide. However, an alternative process

to the Kraft process is the sulphite process and it is studied within this Thesis. It employs pulping cooking

with sulfites or bisulfates combined with magnesium or sodium (Bajpai, 2017b). The sulphite process

differs from the Kraft process in several aspects howbeit the most relevant one is its easiness and high

flexibility in bleaching. It can also lead to a higher yield of brighter beached pulp in comparison to the

Kraft process (Oeveren, V, Congress, & Diego, 2004).

Pulp washing is an essential step in the removal and recovery of the high amounts of chemicals required

in the cooking process due to several reasons expressed as follows: the dissolved chemicals interfere

with the downstream processing of the pulp, the chemicals are costly to replace and, finally, they can

be extremely harmful to the environment.

The following process, pulp bleaching, is used to obtain pulp products in which high purity and brightness

is required. The process can be classified according to the use of bleaching agents. It usually starts with

oxygen delignification where residual lignin is removed. However, the final brightness is only achieved

through further bleaching steps that can vary between each pulp mill plant (U.S. Environmental

Protection Agency, 2001).

The Portuguese pulp company Caima – Indústria da Celulose, S.A. is one of the few pulp mills in Europe

that produces bleached sulphite pulp from hardwood of Eucalyptus globulus using a magnesium

bisulphite solution. The bleaching of the pulp is total chlorine free (TCF), and it is characterized

according to the sequence E-O-P, which starts with an alkaline extraction (E) followed by an oxygen

delignification (O) and ends with the addition of hydrogen peroxide (P). Usually each bleaching stage

consists of a reactor followed by a pulp wash, as it is illustrated in Figure 2-3.

7

Figure 2-3. Schematic illustration of the bleaching process of the pulp mill from Caima - Indústria da Celulose,

S.A., high lightening the effluent from which the study is going to be conducted.

Regarding the bleaching process, the effluent from the pulp bleaching cannot be recycled to the recovery

boilers of the pulp mill due to the addition of sodium hydroxide. Clogging of the equipment is a

consequence, which is one of the reasons why it is so important to improve the treatment of that effluent

(Nordin & Jönsson, 2008).

In fact, a large variety of external treatment technologies are being used to treat pulp mills effluents. The

preferred technology depends on many factors, such as the properties of the matter to be removed,

environmental constraints, along with techno-economical factors. Membrane filtration and evaporation

are two advanced concentration methods that are already implemented in some pulp mills (Bajpai,

2017a).

8

For example, Caima – Indústria da Celulose, S.A. has successfully been using a series of evaporators

to treat the E -stage effluent from the bleaching process. However, it is still possible to improve the

separation, either by upgrading the evaporation system already implemented, by replacing it for a

membrane filtration system or even by integrated both methods in one effluent treatment system.

Concentration Methods

2.2.1 Membrane Filtration

A membrane is defined as permselective barrier between two homogeneous phases. A molecule or

particle is transported across a membrane from one phase to another because a force acts on that

particle. The extent of this force is determined essentially by the gradient in potential across the

membrane and the potential difference can arise as a result of differences in either pressure,

concentration, temperature or electrical potential (Mulder, 1996).

Materials and structure of synthetic membranes

The membrane itself can be manufactured from more than one material, such as polymers, ceramics,

glass and metals. Each material properties affects differently and significantly the design and

performance of the membrane filtration operation. Polymeric membranes are less expensive in

comparison with the other membrane materials hence represent the vast majority of membrane currently

being used. In case an application requires the solvent to be resistant and thermally stable, then

ceramics membranes are a better choice when compared with polymeric membranes.

In addition to the membrane material, the trans-wall symmetry of the membrane influences the operation

performance. In a symmetric membrane, the density or pore structure is uniform, while in an asymmetric

membrane there is a change in the density of the membrane material throughout the cross-section area.

A breakthrough to industrial applications was the development of asymmetric membranes. Those are

essentially used in pressure driven membrane processes and, in comparison to symmetric membranes,

the advantages are high flux, high rejection coefficient and good mechanical stability. The membranes

are composed by a very dense top skin layer supported by a highly porous sublayer. The latter sub-

layer serves only as a support and it has little effect on the separation characteristics of the process.

That facts leads to a membrane capable of having high selectivity, as it is excepted for a dense

membrane, combined with a high permeation rate of a thin membrane (Strathmann et al, 2006).

Another type of membrane is the composite membrane that presents an extremely thin surface layer

supported on a much thicker porous structure. Both layers can be originated from different materials,

usually polymers, and thus optimized independently, improving greatly the separation required and the

permeation rate of the process, nowadays, from a commercial point of view, vastly needed (Bungay,

Lonsdale, & de Pinho, 1986) (Baker, 2004).

9

Figure 2-4. Schematic diagrams of the principal types of membranes: a) symmetrical microporous membrane; b)

asymmetric microporous membrane; c) thin film composite membrane. Adapted from (Baker, 2004).

Pressure driven membrane processes

Pressure difference as the driving force in the separation process represents the most common

membrane filtration technologies. The processes can be distinguished according to the pore size in the

membrane and consequently regarding to the ability of separate specific molecules. In fact, regarding

microfiltration (MF), followed by ultrafiltration (UF) and nanofiltration (NF) and, reverse osmosis (RO),

the molecular weight of the particles separated diminishes as the membrane pore sizes decreases

ranging from 100 to 1 × 10−4 microns (𝜇𝑚) (Epa, 2005). Figure 2-5 illustrates the average pore size of

the particles retained by the membrane concerning the several distinct membrane filtration processes.

0,0001 0,001 0,01 0,1 1 10 100Pore size (µm)

MF

UF

NF

RO

Membrane

filtration process

Figure 2-5. Membrane filtration processes according to the average pore size of the particles retained by the

membrane (Epa, 2005).

Membrane systems can operate either through dead-end filtration or cross-flow filtration. Several

membrane filtration processes use a dead-end technique, in which the feed stream is directed

perpendicular to the filter surface, as it is illustrated in Figure 2-6, a). Nonetheless, when it is required

handling high concentration of small particles and molecules, the most suitable technique is undoubtedly

the one that uses cross-flow membranes, whose process is represented in the Figure 2-6, b).

Cross-flow filtration is a filtration technique in which the constant turbulent feed flow passes along the

membrane surface. This prevents the accumulation of matter and leads to a better filtration performance.

In addition to that, it enables less frequent membrane cleanings when compared to dead-end flow

system. This type of system is composed by a flow named as permeate, which contains the substances

a) b) c)

10

smaller than the membrane pores, while the remaining substances from the feed flow are present in the

stream designated as concentrate or retentate or even residue (Strathmann et al., 2006).

Figure 2-6. Membrane systems according to the flow process: a) dead-end filtration; b) cross-flow filtration.

Membrane modules

In an industrial scale, large membrane areas are usually required thus it is important to establish how

to properly pack the smallest unit of a membrane. That unit is called a module. The module is the central

part of a membrane installation and it can be present assorted designs that will be discussed further on.

The selection of module configuration, as well as the arrangement of module in a system is based solely

on economic considerations with the correct engineering parameters being employed to achieve the

final product specification (Mulder, 1996). The simplest one, the plate and frame module, is composed

by a set of two membranes and spacers, forming the feed flow channel. The set is clamped and stacked

between two endplates and placed in an housing, as illustrated in Figure 2-7. The latter module is

suitable for batch operations and it is mainly used in small-scale applications, especially in the food and

pharmacy industry.

11

Figure 2-7. Plate and frame module design. Adapted from (W. Baker, 2012)

A more complex configuration is the spiral wound module. It can be described as a plate and frame

system wrapped around a central collection pipe. The spiral wound module has a greater packing

density, providing a relatively large membrane area per unit volume. However, the spiral wound module

is quite sensitive to fouling and the feed channels can easily be blocked due to that it is often requiring

a pretreatment procedure (Bulletin, 2017).

Figure 2-8. Spiral wound module design. Adapted from (W. Baker, 2012).

12

While the previous described modules required flat sheet membranes for their manufacturing, tubular

and hollow fiber modules demand other membrane configurations. The tubular membrane module

consists of membrane tubes placed inside a porous support or bundled together. Phenomena as

concentration polarization and membrane fouling, can be controlled by using this type of module. One

drawback of implementing tubular design is its low surface area which entails high capital costs

(Strathmann et al., 2006). Ultimately, the hollow fiber module has the highest packing density of all

modules types available in the market and with a very cost-effective production. In opposite to tubular

membrane modules, hollow fiber modules are not the most suitable ones when it comes to concentration

polarization and membrane fouling (Baker, 2004). A tubular module design is expressed in Figure 2-9

whereas a hollow fiber module design is illustrated in Figure 2-10.

Figure 2-9. Tubular module design. Adapted from (W. Baker, 2012)

Figure 2-10. Hollow fiber membrane module design. Adapted from (W. Baker, 2012)

13

Table 2-1 presents typical characteristics of the aforementioned membrane modules.

Table 2-1. Comparison between several membrane modules. Hollow fine fiber stands for capillary membrane

module. Ceramic membranes are grouped separately from polymeric membranes because their preparation

methods are significantly different. Adapted from (Wagner, 2001).

Spiral

wound

element

Tubular Plate and

frame

system

Hollow

wide fiber

system

Hollow

fine fiber Ceramic High

price

Low

price

Membrane

density high low average average

very

high low

Plant

investment low high low high very high medium very high

Tendency

to fouling average low average low

very

high medium

Clean

ability good good good low poor good

Variable

costs low high low average average low high

Flow

demand medium high medium medium high low very high

Membrane fouling

Fouling is one the biggest challenges associated with the operation of pressure driven membrane

processes. This phenomenon occurs due to the adhesion and deposition of particles and colloids onto

membrane surfaces and into membrane pores. That enables the plug of pores and it also leads to the

cake formation on the surface of the membrane. Fouling increases the pressure required to generate

the desired volume of product water and it requires expensive chemical cleaning for removal. Moreover,

irreversible fouling, which can not be removed with chemical cleaning, reduces the performance of the

membranes overtime and it can lead to a membrane replacement. Therefore, reversing, removing

and/or mitigating membrane fouling will increase economic efficiency for membrane applications

through a reduction in the required transmembrane pressure or an increased membrane lifespan

(Guerra & Pelligrino, 2012), (Johan Thuvander, 2018).

14

Process performance parameters

The membrane performance is characterized by a range of parameters, which in turn, are crucial to

design a membrane system. Firstly, the flux (𝐽) is defined as the flow of permeate per unit of area and it

is described by the resistance model shown in Equation (2.1).

𝐽 =𝑇𝑀𝑃. 𝐿𝑝

𝜇𝑝

(2.1)

Where 𝑇𝑀𝑃 is the pressure difference across the membrane named as the transmembrane pressure,

𝐿𝑝 is the membrane permeability coefficient. if the solution is pure water, then 𝐿𝑝 can be also called as

the hydraulic permeability. Finally, 𝜇𝑝 is the viscosity of the permeate. The intrinsic membrane

resistance, 𝑅𝑚, can be written as 1 𝐿𝑝⁄ (Cheryan, 1998).

The extent to which a solute is retained by a membrane is given by the retention coefficient expressed

by Equation (2.2), whereas the true retention coefficient is defined according to Equation (2.3). Note that

𝐶𝑝, 𝐶𝑏 and 𝐶𝑚 are the concentration of the solute in the permeate, in the bulk solution and at the

membrane surface, respectively (Mulder, 1996). If the flux increases, increasing also the concentration

at the membrane surface, the observed coefficient value decreases being always lower than the true

retention coefficient value (J. Thuvander, 2018).

𝑅𝑜𝑏𝑠 = 1 − 𝐶𝑝

𝐶𝑏

(2.2)

𝑅𝑡𝑟 = 1 − 𝐶𝑝

𝐶𝑚

(2.3)

Additionally, another the parameter that represents the concentration process is the volume reduction

(VR) which value is the ratio between the permeate volume, 𝑉𝑝, and the initial feed volume, 𝑉𝐹. The

volume reduction value increases with the increasing of the final product concentration required.

Consequently, the viscosity of the retentate will increase as the concentration at the surface of the

membrane, leading to a decrease in the flux until a point in which the process is no longer feasible

(Mulder, 1996). Another way of expressing the volume reduction VR parameter is through the volume

reduction factor VRF, as it is described in Equation (2.5).

𝑉𝑅 = 𝑉𝑝

𝑉𝐹

(2.4)

𝑉𝑅𝐹 =1

1 − 𝑉𝑅 (2.5)

15

Process design

The design of a membrane filtration system can differ immensely due to the amount of module

configurations possible to apply. The module is the central part of a membrane installation while a stage

is a certain number of modules connected together in series or in parallel. Moreover, both the selection

of a module configuration and the arrangement of the modules in a system are based solely on economic

considerations with the correct engineering parameters being employed to achieve the final product

specification.

Generally, the ideal configuration is in a parallel array which it will result in the lowest pressure drop

along with the highest flux value. However, the power consumption and the housing cost will be

considerably high (Cheryan, 1998). On the other hand, implementing more than one element in series

in an housing will result in a required lower power pump capacity hence a lower pump cost

comparatively. Despite that, it will also lead to a decrease of the TMP across the membrane housing,

decreasing the driven-pressure membrane performance. Thus one of the economic trade-offs is

definitely between the lower pump cost and the higher cost of more modules (Cheryan, 1998).

Usually to obtain the desired product specifications a cascade operation is needed, where the retentate

of each stage is the feed stream for next successive stage. The system design applied in this study is a

system with retentate recycle in which the retentate stream is fed back to the feed of the first stage. As

so, allowing an increase of the system recovery as well as the cross-flow within the pressure vessels.

In this type of system, a feed pump is required to assure a feed flow, combined with a recycle pump for

each stage, as presented in the Figure 2-11.

Figure 2-11. Example of a cascade plant design with retentate recycle composed by three stages in which the first

two the housings are in parallel. Addapted from (Nilsson, Lipnizki, Trägårdh, & Östergren, 2008).

16

Membrane system cost estimation

An economic evaluation of a membrane filtration process comprises the estimation of capital

expenditure (CAPEX) along with operating expenditure (OPEX). The first one can include direct and

indirect capital costs, whereas OPEX represents the sum of operating and maintenance costs. Lastly,

the total annual costs (TAC) express the sum of the CAPEX and OPEX.

Capital costs

The capital costs, 𝐶𝐶𝐴𝑃, are defined as fixed, onetime expenses that can include purchase of the

equipment, construction, engineering services, project development, membrane and pressure vessels,

automation and control, pipping, valves and fittings, cleaning in place (CIP) equipment and

miscellaneous costs. The latter parameter includes building, electrical supply and distribution. The total

capital cost is determined as the amortized capital expenditure (CAPEX) assuming a practical interest

rate, 𝑖, and considering an investment period, ℕ, of 10 years. The correlation that is being multiplied by

the capital cost is named as capital recovery factor, CRF and it presents for this case a value of 0,13.

𝐴𝑚𝑜𝑟𝑡𝑖𝑧𝑒𝑑 𝐶𝐴𝑃𝐸𝑋 = 𝐶𝐶𝐴𝑃 . [𝑖(1 + 𝑖)ℕ

(1 + 𝑖)ℕ − 1] (2.6)

Current purchasing data is difficult to obtain therefore assumptions based on literature data are

necessary. Hence the capacity method was used to estimate costs when specific data as the purchase

cost of the equipment, 𝐶𝑚𝑒𝑚, was lacking. Equation (2.7) expresses the capacity method law, in which

𝐶 and 𝐾 are the cost and capacity of a certain equipment, respectively, 𝐶𝑟𝑒𝑓 and 𝐾𝑟𝑒𝑓 are the reference

cost and capacity values from a valid reference data. Finally 𝑛 is the cost capacity factor related to the

economy of scale (Sethi, 1997). Consequently, it is required to correct the cost prices in order to

represent values of the year 2017. A solution to that is through the chemical engineering plant cost index

(CEPI) according to Equation (2.8) (M. S. Peters & Timmerhaus, 1991).

𝐶 = 𝐶𝑟𝑒𝑓 . (𝐾

𝐾𝑟𝑒𝑓

)

𝑛

(2.7)

𝐶 = 𝐶𝑟𝑒𝑓 . (𝑐𝑜𝑠𝑡 𝑖𝑛𝑑𝑒𝑥 𝑣𝑎𝑙𝑢𝑒 𝑎𝑡 𝑡ℎ𝑒 𝑝𝑟𝑒𝑠𝑒𝑛𝑡 𝑡𝑖𝑚𝑒

𝑐𝑜𝑠𝑡 𝑖𝑛𝑑𝑒𝑥 𝑣𝑎𝑙𝑢𝑒 𝑤ℎ𝑒𝑛 𝑒𝑠𝑡𝑖𝑚𝑎𝑡𝑒𝑑) (2.8)

The model developed by (Sethi, 1997) estimates the capital costs of a membrane system according to

each component category; pipes and valves, 𝐶𝑃𝑉, instruments and controls, 𝐶𝐼𝐶, tanks and frames, 𝐶𝑇𝐹,

and miscellaneous equipment, 𝐶𝑀𝐼 . The estimation requires a conversion that relates a certain constant

and coefficient to the membrane area value. In accordance with (Wagner et al., 2001) the membrane

area value required for a system is calculated by dividing the permeate volume by the flux. The following

equations are based on the previous mentioned model.

𝐶𝑃𝑉 = 5926,13 . (𝐴𝑚𝑒𝑚)0,42 (2.9)

17

𝐶𝐼𝐶 = 1445,5 . (𝐴𝑚𝑒𝑚)0,66 (2.10)

𝐶𝑇𝐹 = 3047,21 . (𝐴𝑚𝑒𝑚)0,53 (2.11)

𝐶𝑀𝐼 = 7865,02 . (𝐴𝑚𝑒𝑚)0,57 (2.12)

The feed pump is designed to being capable of delivering the plant feed flow to the recirculation pump

at the same pressure as the returning recirculation flow whereas the recirculation pump is sized to

deliver the plant feed and the recirculated volume at a pressure equal to the pressure drop across the

modules. Based on common practice a carbon steel centrifugal pump type was selected and the

purchasing and installation costs were estimated according to a study performed by (Symister, 2016).

The study aimed to evaluate distinct module costing technique outlined by Richard Turton et al. and by

Gavin Towler and Ray Sinnot (Symister, 2016) . The cost of each pump, feed pump, 𝐶𝑝,𝑓 and

recirculation pump, 𝐶𝑝,𝑟, is related to the power required, which is determined through the energy

required per m3 of permeate, as it is expressed by Equations (2.13) and (2.14), respectively,

𝑊𝑓𝑒𝑒𝑑 =𝑃𝑖𝑛𝑙𝑒𝑡

𝜂. 𝑉𝑅 (2.13)

𝑊𝑟𝑒𝑐𝑖𝑟𝑐 =∆𝑃𝑓 . 𝑄ℎ𝑜𝑢𝑠𝑖𝑛𝑔

𝜂. 𝐽. 𝐴𝑚𝑒𝑚

(2.14)

where 𝑃𝑖𝑛𝑙𝑒𝑡 is the inlet pressure to the plant, 𝜂 is the pump efficiency, ∆𝑃𝑓 and 𝐽 are the frictional

pressure drop and the flux in the housing and 𝑄ℎ𝑜𝑢𝑠𝑖𝑛𝑔 is the feed flow in the housing.

Finally, the total capital cost 𝐶𝐶𝐴𝑃 is the sum of the individual costs of all the capital equipment as it is

described by Equation (2.15).

𝐶𝐶𝐴𝑃 = 𝐶𝑚𝑒𝑚 + 𝐶𝑃𝑉 + 𝐶𝐼𝐶 + 𝐶𝑇𝐹 + 𝐶𝑀𝐼 + 𝐶𝑝,𝑓 + 𝐶𝑝,𝑟 (2.15)

Operating and Maintenance Costs

The annual operating and maintenance cost (OPEX) can cover the sum of the following parameters,

electrical energy, labor, maintenance, chemicals (cleaning and disposal) and membrane replacement.

The cost of the replacement of the membranes is calculated based on the number of sets of replacement

of membranes, 𝑀𝑟, multiplied by the membrane replacement cost, 𝐶𝑟. Furthermore, the annual cost of

membranes replacement, 𝐶𝐴𝑛𝑛𝑢𝑎𝑙 𝑟𝑒𝑝, is calculated from the total cost of membrane replacement

amortized aver the lifespan of the plant, 𝐿𝑃 (Guerra & Pelligrino, 2012). The latter parameters are

described in Equations (2.16) and (2.17).

18

Further, the annual energy cost is estimated as being equal to the unit energy cost times the annual

pump work from the membrane filtration system. Additionally, the annual cleaning cost is established

from literature data. Lastly, the sum of the annual labor and the maintenance costs 𝐶𝐿𝑎𝑏𝑜𝑟 and 𝐶𝑚𝑎𝑖𝑛,

are considered as being equal to 2% of the CAPEX (Arkell, Krawczyk, Thuvander, & Jönsson, 2013).

𝑀𝑟 =𝐿𝑃

𝐿𝑚

− 1 (2.16)

𝐶𝐴𝑛𝑛𝑢𝑎𝑙 𝑟𝑒𝑝 =. (𝑖(1 + 𝑖)𝐿𝑃

(1 + 𝑖)𝐿𝑃 − 1) . 𝑀𝑟 . 𝐶𝑚.𝐴𝑚

(2.17)

A summary of all the parameters values and ranges used in the membrane filtration cost model is listed

in Table 2-2.

Table 2-2. Overview about the parameters and assumed values for the cost calculation.

Parameter Value Reference

Annual operating hours (ℎ/𝑦) 8000 (Jönsson et al., 2008)

Plant lifespan (𝑦) 10 Typical practice

Membrane lifespan (𝑦) 1,5 (Suárez, Fernández, Iglesias, & Igesias,

2015)

Interest rate, 𝑖 (%) 5 Typical practice

Electricity price (€ 𝑀𝑊 ℎ⁄ ) 30 (Jönsson et al., 2008)

Pump efficiency (%) 70 (Suárez et al., 2015)

Membrane cost (𝑈𝑆 $ 𝑚2⁄ ) [490 - 2200] (Samhaber & Nguyen, 2014)

Cleaning cost (€ 𝑚2/𝑦⁄ ) 50 (Jönsson et al., 2008)

19

2.2.2 Evaporation

The design of an evaporator system is supported by the selection of the most suitable equipment and

its arrangement for the purpose of the process accomplishing efficiency and economic criteria.

Evaporation unit plants have been present in the industry for many years and evaporation equipment

has been improved and renovated massively by manufactures in order to address the challenging

product specifications. Thus, a large number of evaporators designs can be found in the market. In the

pulp and paper industry the most common evaporators types are the rising film tubular, the falling film

tubular and the mechanical vapor recompression (MVR). Usually, the arrangement is designed as a

multiple-effect evaporator system in which several evaporators are connected in series as a single unit.

Called as the first “modern” evaporated implemented in the industry, the rising film tubular unit was

developed commercially by using a vertical tube with steam condensing on its outside surface. Liquid

on the inside of the tube is boiling and generating vapor. This vapor forms a core in the center of the

tube. As the fluid moves up to the tube, more vapor is formed forcing the liquid to move to the tube wall

(Figure 2-12, a).) (SPX Corporation, 2008). In a falling film design, the liquid to be concentrated is

supplied to the top of the heating tubes and distributed in such a way as to flow down the inside of the

tube walls as a thin film. The liquid film starts to boil due to the external heating and is partially

evaporated as a result. The downward flow, caused initially by gravity, is enhanced by the parallel,

downward flow of the generated vapor. The residual film liquid and vapor is separated in the lower part

of the equipment. The main advantage of this equipment is the driving force non-limitation, allowing a

greater number of evaporators effects to be used within the same overall operating conditions

(Engineering, 2018), (SPX Corporation, 2008) (Figure 2-12, b) Lastly, mechanical vapor recompression

evaporators compress the vapor mechanically and then the vapor is used as a high pressure vapor in

the steam chest. (Figure 2-12, c). Those evaporators require significantly lower amounts of energy

because all the vapor generated is recycled.

20

Figure 2-12. Evaporators models: a) rising film tubular, b) falling film tubular and c) mechanical vapor

recompression evaporator. Adapted from (SPX Corporation, 2008).

a) b)

c)

STEAM

CONDENSATE

FEED

CONCENTRATE

21

Theoretical background

The evaporation process aims to separate a certain non-volatile component of a stream by boiling

volatile components, mainly water. That fact allows the concentration of the desired component. The

industry presents a large variety of evaporators models that differ from each other. The main difference

can be according to product specifications and due to economic constraints. Regardless the evaporator

model, the evaporation theory for a multiple-effect evaporator (MEE) system is valid for all the cases. In

such configuration, live steam is condensed in the first evaporator, generating vapor. The latter is sent

to condense in a second effect, where additional evaporation takes place. The process can be repeated

until reaching the last effect evaporator, where generated vapors are condensed in a condenser using

cooling water. The effects are numbered by convention in which the first effect is the one that has the

heating vapor at the highest pressure, thus to ensure the driving force, the relations expressed by

Equations (2.18) and (2.19) must be verified.

𝑝𝑁 < 𝑝𝑁−1 < ⋯ < 𝑝1 < 𝑝𝑆 (2.18)

𝑇𝑁 < 𝑇𝑁−1 < ⋯ < 𝑇1 < 𝑇𝑆 (2.19)

where 𝑁 is the number of the last effect and 𝑝𝑆 and 𝑇𝑆 are the pressure and temperature of the saturated

live steam, respectively. Furthermore, it is assumed that the heat transfer coefficients as well as the

heat transfer areas are the same for all effects. Therefore, the heat transfered in each effect, 𝑖, can be

described according to the following equation, in which the subscript 𝑖 = 0 corresponds to the saturated

live steam.

𝑞𝑖 = 𝑈𝑖 . 𝐴𝑖 . (𝑇𝑖−1 − 𝑇𝑖)

(2.20)

𝑞𝑖 = 𝑈𝑖 . 𝐴𝑖 . ∆𝑇𝑖 , 𝑖 = [1, 𝑁]

The evaporator temperature, 𝑇𝑖, is assumed to be equal to the liquor temperature, 𝑇𝑙𝑖, which is

calculated according to Equation (2.20), given by the fact that the vapor and liquid in 𝑖th effect are in

equilibrium and defined in terms of boiling point rise (𝐵𝑃𝑅) as it follows,

𝑇𝑙𝑖 = 𝑇𝑣𝑖 + 𝐵𝑃𝑅𝑖 (2.21)

Thereafter it is possible to determine the heat transfer area of each evaporator, 𝐴𝑖, required for the

evaporation process. Regarding the assumptions made, it must present an equal value for each effect.

Moreover, due to the relation given by Equation (2.20) the drop of temperature in each evaporator is

inversely proportional to the value of the respective heat transfer coefficient. The latter is a consequence

of considering the heat transfer and the heat transfer area equal for each evaporator. Thus, since the

temperature of the live steam, 𝑇𝑆, and the temperature of the last effect are known values, the driving

force ∆𝑇𝑖 is estimated according to Equation (2.22).

∆𝑇𝑖 = (𝑇𝑆 − 𝑇𝑁).1 𝑈𝑖⁄

∑ 𝑈𝑖𝑁𝑖=1

(2.22)

22

Since modeling and simulation are tools that will be used further on, mass and energy balances are

essential. Regarding the terms shown in Figure 2-13, a mass balance around the ith effect is given by

Equation (2.23).

Figure 2-13. Block diagram for the 𝑖th effect. Adapted from (Kumar, Kumar, & Singh, 2013).

𝑑

𝑑𝑡𝑀𝑖(𝑡) = 𝑊𝑙𝑖+1 − 𝑊𝑙𝑖 − 𝑊𝑣𝑖 (2.23)

Assuming a non-volatile component situation and a steady case scenario with a parallel flow system, as

it is illustrated in Figure 2-14 as an example, it is possible to simplify the problem solution through the

development of a simple yet efficient model. The model is based on the global mass balance along with

the component mass balance and is expressed by Equations (2.24) and (2.25)., respectively.

Additionally, a mass balance addressed to each effect is required, as it is given by Equation (2.26).

Finally, an energy balance is shown by Equation (2.27), used in accordance with the conventional

standard state.

Figure 2-14. Quadruple-effect evaporator system parallel flow case. Adapted from (Kaya & Ibrahim Sarac, 2007).

23

𝑊𝑙,𝑓𝑒𝑒𝑑 = 𝑊𝑙𝑁 + ∑ 𝑊𝑣𝑖

𝑁

𝑖=1 (2.24)

𝑥𝑖−1. 𝑊𝑙𝑖−1 = 𝑥𝑖𝑊𝑙𝑖 (2.25)

𝑊𝑙𝑖−1 = 𝑊𝑙𝑖 + 𝑊𝑣𝑖 (2.26)

𝑊𝑙𝑖−1. ℎ𝑙𝑖−1 + 𝑊𝑣𝑖−1. ∆ℎ𝑣𝑎𝑝𝑖−1 = 𝑊𝑙𝑖 . ℎ𝑙𝑖 + 𝑊𝑣𝑖 . ℎ𝑣𝑖 (2.27)

Process performance parameters

The efficiency of a MEE system is measured according to two major parameters, namely steam

economy (SE) and steam consumption (SC), also expressed as 𝑊𝑣𝑠. The steam economy is defined as

the ratio of vapor mass produced to the steam mass consumed. While the steam consumption value is

approximately the number of effects times lower than the steam consumption with only one effect.

However, the equipment cost can also be 𝑁 times higher when compared with a single evaporator cost,

which is a significant aspect that will be addressed in the Chapter 3.

Process model design

For an accurate estimation of SE and SC mathematical modeling couple with simulation is required. In

modeling of multi-effect evaporators, the pressure and temperature values of the live steam are set. In

addition to that, the necessary enthalpies and specific heat capacities are estimated from

thermodynamic equations from literature. While the heat transfer coefficients value is given by literature

data. Furthermore, the feed flow and initial concentration value must be known as well as the pressure

or temperature of the last effect and the desired final product concentration. The mass, component and

energy balances are provided for each evaporator by independent linear equations expressed in

Equations (2.24) – (2.27). The heat transfer area and the evaporator temperature are calculated through

Equations (2.20) and (2.22), respectively. All the parameters mentioned above represent the overhead

(OH) of the model applied, giving the input values required.

The iterative process is developed with the assumption that the heat transfer area must be equal for

each evaporator. Predominantly, the difference between the first iteration and the following ones relies

on the fact that firstly it is assumed that the mass flow of vapor generated in each evaporator is the

same, which it is not necessarily true. Thus, for the next iteration, new values of mass flow, temperatures

and heat transfer areas will be obtained applying once again Equations (2.24) - (2.27). Although for the

evaporator temperature and heat transfer area, Equations (2.28) and (2.29) are the ones implemented.

Lastly, a MATLAB function called fsolver is executed until fulfilling the convergence criteria of 10%

described in Equation (2.21) and enabling the determination of steam economy and steam consumption.

24

∆𝑇′𝑖 = ∆𝑇𝑖.𝐴𝑖

𝐴𝑚

(2.28)

𝐴𝑖 =𝑊𝑣. ∆ℎ𝑣𝑎𝑝

𝑖

𝑈𝑖. ∆𝑇𝑖

(2.29)

𝐴𝑚 =∑ (∆𝑇𝑖. 𝐴𝑖)

𝑁𝑖=1

∑ ∆𝑇𝑖𝑁𝑖=1

(2.30)

|𝐴𝑖 − 𝐴𝑚

𝐴𝑚

| ≤ 0,1 (2.31)

Figure 2-15. Iterative method performed to calculate the steam economy (SE) and the steam consumption (SC).

However, to model the complex process of evaporation temperature, dependent physico-thermal

properties of the effluent are crucial, therefore correlations from (Kumar et al., 2013) and (Khanam &

Mohanty, 2011) were taken into account and are expressed in the following equations.

Boiling point rise, ℃

𝐵𝑃𝑅 = (6,173. 𝑥𝑖 − 7,48. 𝑥𝑖1,5 + 32,747𝑥𝑖

2). [1 + 0,006. (𝑇𝑠 − 3,7316)] (2.32)

Effluent specific heat capacity, 𝑘𝐽. 𝑘𝑔−1. 𝐾−1

𝐶𝑝𝑙𝑖 = 4,187. (1 − 0,54. 𝑥𝑖) (2.33)

Effluent enthalpy, 𝑘𝐽. 𝑘𝑔−1

ℎ𝑙𝑖 = 𝐶𝑝𝑙𝑖 . 𝑇𝑙𝑖 (2.34)

Water enthalpy, 𝑘𝐽. 𝑘𝑔−1

ℎ𝑣𝑖 = 4,1832. 𝑇𝑙𝑖 + 0,127011 (2.35)

Saturated steam enthalpy, 𝑘𝐽. 𝑘𝑔−1

ℎ𝑣𝑖𝑣 = 1,75228. 𝑇𝑣𝑖 + 2503,35 (2.36)

Overhead (OH)MATLAB function

Fsolve

Xinput Model

Multi-effect evaporationSE and SC

X input

Convergence criteria

(CC)

CC 0,1

25

Saturated steam heat of vaporization, 𝑘𝐽. 𝑘𝑔−1

∆ℎ𝑣𝑎𝑝𝑖 = 2519,5 − 2,653. 𝑇𝑣𝑖 3.37

Evaporation system cost estimation

Analogously, an evaporation system economic analysis is performed according to the membrane

filtration cost estimation methodology. Hence the definition of CAPEX and OPEX as well as the prices

correction remains valid. Howsoever, the model to estimate the capital investment cost differs and it is

named as percentage of delivered-equipment cost method described by (S. M. Peters & Timmerhaus,

1991). According to the economic method, it is necessary to calculate the purchase of the equipment

cost, 𝑃𝐸𝐶. The parameter can be estimated through the relation between the installed cost of vertical

tube evaporators and the heat transfer area illustrated in Figure 2-16 based on prices of 1990. Then it

is necessary to use a CEPI value of 392,2 when taking into account the Equation (2.8).

Figure 2-16. Installed cost of evaporators according to heat transfer area. Adapted from (S. M. Peters &

Timmerhaus, 1991)

26

The other parameters included in the capital costs are estimated as percentages of the purchase of the

equipment cost value. Therefore, the capital cost, 𝐶𝐶𝐴𝑃, is calculated according to Equation (2.37), where

𝑓1,𝑓2,…, are multiplying factors. Those are related to instrumentation and control, piping, electrical

services, buildings and services, service facilities, land and yard improvements, and spare parts as

direct costs parameters, and as indirect costs parameters, engineering and supervision, construction

expenses, contractor´s fee and contingency. The percentages of each parameter are summarized in

Table 2-3. Additionally, the economic model considers another parameter, the working investment which

if 12% of capital costs. Therefore, for an evaporation plant, the CAPEX is the sum of capital costs and

the cost of working investment multiplied by the capital recovery factor.

𝐶𝐶𝐴𝑃 = 𝑃𝐸𝐶 + ∑(𝑓1. 𝑃𝐸𝐶 + 𝑓2. 𝑃𝐸𝐶 + ⋯ ) (2.37)

Table 2-3. Percentage of the capital cost parameters related to the PEC, adapted from (S. M. Peters &

Timmerhaus, 1991).

Capital cost parameters (% of 𝑷𝑬𝑪)

Dir

ect

Co

sts

Equipment 100

Instrumentation and control 10

Piping 10

Electrical services 5

Building and building services 15

Land and yard improvements 5

Service facilities 20

Spare parts 4

Ind

irect

Co

sts

Engineering and supervision 12

Construction expenses 10

Contractor’s fee 0,5

Contingency 8

The operating and maintenance costs for an evaporation system are deeply dependent on the steam

price and, consequently, on the steam consumption (SC). As stated by (Rajendran, Rajoli, Teichert, &

Taherzadeh, 2015), a typical practice is to consider 0,0085€/kg as a steam price value. The remaining

parameters, namely labor and maintenance can be estimated as being 1% and 2% of the total capital

investment costs, 𝐶𝐶𝐴𝑃, respectively (Pereira et al., 2018).

27

Concentration Methods in the Pulp and Paper Industry

The pulp and paper industry is focused on developing innovative methods aimed to increase the dry

solids concentration. Those methods are mainly to work as advanced wastewater treatment

technologies capable of meeting the environmental and economic constraints that the industry is facing.

In the present context evaporation systems along with membrane filtration systems are within the most

viable options.

Regarding the evaporation systems, they can consist in a certain number of evaporators disposal in a

large number of combinations. The evaporator type differs significantly according to the product and to

the purpose of the process.

The pulp mill Eldorado Cellulose e Papel, in Brazil has implemented a six-effect evaporators plant with

a tubular concentrator evaporator model capable of reaching a dry solids content TS (w/w %) of 80%.

Using the same evaporation process design, the pulp mill in Chile, Arouca Nueva Aldea, concentrates

the black liquor also until 80% in dry solids content. The swedish Mondi Packaging Dynäs mill applies a

configured six-effect tubular concentrator evaporation train capable of producing a 75% dry solids

content. In Portugal, Celbi built a complete seven-effect evaporation plant using the latter evaporator

model which has a capacity of concentrating until 74 (w/w %) dry solids content in the final product

(Gallagher, 2018).

A different evaporation system, a falling-film evaporation system, is successfully integrated in Stora

Enso Skoghall pulp mill, in Sweden, with nine effects. That configuration allows a given process stream

to reach a 80 (w/w %) dry solids product (Gallagher, 2018).

The reclamation of pulp mill effluents is being subject of great interest hence a wide range of distinct

processes have been implemented. Li & Watkinson show an example of a typical pulp mill wastewater

treatment in which an evaporation system made up of a mechanical vapor recompression evaporators

system is able to raise the total solids concentration from 2% to 20 (w/w %). The process is then followed

by a multiple effect evaporation system which increases the solids content further to roughly 35 (w/w

%). Another wastewater treatment example is situated in Austria, the Lenzing pulp mill installed in 1990

a multiple effect evaporation plant upstream the bleaching extraction stage filtrate capable of

concentrating the effluent from 2% up to 52 (w/w %) in dry solids content (Gleadow & Stratton, 2003).

Since 1981 an early use of ultrafiltration in a tubular configuration is implemented at Borregaard sulphite

pulp mill in Norway. The concentration plant has been applied to process spent sulphite liquor, with a

membrane area of 1120 m2 capable of having a dry solids content of 22 (w/w %) in the concentrate flow

(Judd & Jefferson, 2003).

Alfa Laval plate-and-frame modules are used for the production of lignosulphate at Biocel Paskov,

Czech Republic, from spent sulphite liquor. The operating pressure is 5 bar, the temperature is 60°C

and the volume reduction factor (VRF), the ratio between the volume of permeate withdrawn and the

initial feed volume, of the four-stage plant is 80 (Figoli, Cassano, & Basile, 2016).

28

Stora Enso’s Nymölla magnesium bisulphilte pulp mill, situated in the south of Sweden, it has installed

a tubular ultrafiltration plant used to concentrate an alkaline effluent from the bleaching process. The

technology is capable of reaching a high VRF, around 50 for the softwood line and 60 for hardwood.

(Nordin, 2008) The pulp mill needed to obtain a dry substance in the concentrate of about 18% to meet

the environmental discharge demands (Pabby K., Rizvi S.H., & Sastre, 2009).

29

Chapter 3. Techno-economical results

Techno-economical analysis outline

Within the scope of the Thesis, a comparison between membrane filtration and evaporation as

concentration processes is performed. A great number of case studies is evaluated and, several design

specifications and operational variables are defined upon an overall economic evaluation, based on

certain economic metrics (CAPEX, OPEX and TAC). Additionally, a sustainable metric is investigated.

the energy consumption and its cost.

Some variables such as the most viable membrane filtration module in the treatment of the bleaching

plant effluent along with certain operational conditions as transmembrane pressure (TMP) and cross-

flow velocity were studied and established by (São Pedro, 2016). Those parameters were considered

to be the most suitable thus no further investigation was required. However, instead of using the plate

and frame configuration, a corresponding spiral wound module was selected. The process design and

operational parameters are listed in Table 3-1 and Table 3-2.

Table 3-1. Membrane filtration module characteristics

Membrane filtration module NF270-400/34i

Membrane configuration Spiral wound element

Membrane material Polyamide thin-film composite

Active area (𝒎𝟐) 37

Maximum element pressure drop (𝒃𝒂𝒓) 1

pH Range, continuous operation 3 - 10

Maximum operating temperature (℃) 45

Table 3-2. Experimental data used in the techno-economic analysis

Feed volume (𝒎𝟑 𝒉⁄ ) 70

Feed concentration, TS (% 𝒘 𝒘⁄ ) 2,8

Transmembrane pressure (𝒃𝒂𝒓) 13

Operating and feed temperature (℃) 45

30

It is possible to proceed with both a comparison to analyze the most suitable process design parameters

for each process and a study balancing one and the other concentration method already having selected

the process configuration previously. Particularly, the major condition in order to perform a comparison

is comparing concentration system with the same feed conditions along with the same final product

specifications. In fact, the total solid (TS) content is the concentration parameter evaluated through the

comparison.

For the purpose of supporting a better understanding towards the analysis of the different case studies,

a nomenclature is provided. Firstly, it is performed an economical evaluation concerning the

configuration of a membrane filtration system, from which a parallel system (situation 1) and a parallel

and in series configuration (situation 2) is studied. Similarly, an examination regarding the economical

effect of the number of evaporators (3 – 7 evaporators) present in each system was also conducted.

The nomenclature listed in Table 3-3 is related to the standing-alone concentration processes,

membrane filtration and evaporation whereas the nomenclature concerning the combined process,

membrane filtration followed by evaporation is expressed in Table 3-4.

Table 3-3. Nomenclature related to each case study concerning only one concentration method, membrane

filtration or evaporation.

TS feed (%) = 2,83

TS final (%) Membrane Filtration (MF) Evaporation (EVAP)

3,1 I A

3,3 II B

3,7 III C

4,1 IV D

4,7 V E

5,6 VI F

6,8 VII G

31

Table 3-4. Nomenclature related to each case study concerning both concentration methods, membrane filtration

followed by evaporation (MF→EVAP).

Membrane Filtration and afterwards Evaporation

(MF→EVAP)

EVAP TS final (%)

MF TS final (%)

EVAP TS feed (%) 20 30 40

3,1 I – α I - β I - µ

3,3 II - α II – β II – µ

3,7 III – α III – β III – µ

4,1 IV – α IV – β IV – µ

4,7 V – α V – β V – µ

5,6 VI – α VI – β VI – µ

6,8 VII - α VII - β VII - µ

Membrane filtration system configuration analysis

Regarding the membrane filtration unit, two distinct scenarios are analysed. The parameters that differ

from each other are the number of loops, the number of housing per loop, 𝑎, and the number of

membrane elements per housing, 𝑏. Therefore, for the first membrane system (situation 1) it is assumed

that a membrane housing would carry only one spiral-wound element in a parallel configuration.

Whereas for the other membrane system (situation 2), a membrane housing contains three spiral-wound

elements distributed in series, while the housings are arranged in a parallel configuration. The

configuration chosen for each case study is described in Appendix A, Table A-2.

Effectively, as a consequence of the lower frictional pressure drop compared to situation 2, less energy

is required for situation 1. When the membranes are arranged only in parallel (situation 1), a larger

number of loops is needed thus a higher number of recirculation pumps is required. Consequently, the

capital investment for a parallel configuration exceeds the one for a parallel and in series arrangement

(situation 2).

Thereby, observing only the major parameter, the cost of the recirculation pump, a membrane system

configuration in which a housing contains three spiral-wound elements in series distributed in parallel is

the most economically viable option, as it is illustrated in Figure 3-1.

32

Figure 3-1. Recirculation pump cost according to each case study.

Evaporation system configuration analysis

A multiple-effect evaporator system (MEE) can be implemented and examined to study which

configuration results in the best CAPEX, OPEX and total annual costs (TAC) values as being a

concentration process placed in the treatment of a bleaching effluent. The number of effects, i.e.,

evaporators, is the parameter that will vary. Thereby, it enables different required live steam values

along with distinct purchasing equipment costs, which it will affect also the OPEX and the CAPEX

values. Therefore, a comparison between systems designs containing from three to seven evaporators

was performed. As expected, the CAPEX values for each system increase with the number of

evaporators while the OPEX values, in turn, do not follow the same relation. The lowest number of

evaporators system OPEX increases with the increasing of the desired final concentration, reaching a

point (C) from which the system is no longer an interesting option from an economic point of view.

Aiming for a high level of concentration for the final stream, 5-MEE, 6-MEE and 7-MEE systems

represent the lowest OPEX. Live steam values are the preponderant parameter in the determination of

the OPEX. However, for being estimated from an iterative process without being specifically restricted

by any convergence criteria, live steam values are almost identical between the first two latter mentioned

systems. Furthermore, concerning the 4-MEE system, an iterative error is associated since the live

steam value required is higher when compared with a 3-MEE system. The reason why is related to a

wrong first estimation when performing the evaporation simulation which can lead to an inaccurate

solution. The TAC is the following parameter under analysis, allowing to conclude that the 7-MEE is not

a suitable choice. Ultimately, the CAPEX parameter is based on the convergence criteria parameter.

Hence it is strongly considered to be accurate and it suggests that the most viable alternative is indeed

the 5-MEE system. The latter conclusions are based in the results illustrated in Figure 3-2 - Figure 3-4.

I

II

III

IV

V

VI

VII

0

1000

2000

3000

4000

5000

6000

7000

8000

9000

10000

2,8 3,3 3,8 4,3 4,8 5,3 5,8 6,3 6,8

Recircula

tio

n P

um

p C

ost (€

)

TS final (%)

MF Configuration in Parallel (1) MF Configuration in Parallel and in Series (2)

33

Figure 3-2. Evaporation concentration process OPEX for different MEE systems according to each case study.

Figure 3-3. Evaporation concentration process CAPEX for different MEE systems according to each case study.

A

B

C

D

E F G

0

100

200

300

400

500

600

700

800

900

1000

2,8 3,3 3,8 4,3 4,8 5,3 5,8 6,3 6,8

CA

PE

X (

k€

/y)

TS final (%)

3 Evaporators 4 Evaporators 5 Evaporators 6 Evaporators 7 Evaporators

A

B

C

D

E

F

G

0

200

400

600

800

1000

1200

1400

2,8 3,3 3,8 4,3 4,8 5,3 5,8 6,3 6,8

OP

EX

(k€/y

)

TS final (%)

3 Evaporators 4 Evaporators 5 Evaporators 6 Evaporators 7 Evaporators

34

Figure 3-4. Evaporation concentration process TAC for different MEE systems according to each case study.

Membrane filtration and/or evaporation analysis

3.3.1 Membrane filtration or evaporation

In light of the purpose of contrasting concentration technologies it is suggested a comparison analyzing

an evaporation process with 5 evaporators (5-MEE system) and a membrane filtration process with a

configuration as described for situation 2 (Appendix, Table A-2.). Considering the evidences from

Chapters 3.1 and 3.2 those processes are considered to be the most techno-economically suitable.

Fundamentally, to conduct a fair comparison between the two concentration technologies evaporation

and membrane filtration, the procedure requires an analysis of two comparable points. Hence the

starting point, for instance, the feed conditions, as well as the ending point, such as the final product

concentration, must be identical for each case study. Moreover, due to the limiting factor of data

availability it is only feasible to compare membrane filtration and evaporation as stand- alone processes

and for low values of final concentration.

Both CAPEX and OPEX results prove that the membrane filtration is the option economically favorable

for a low bleaching effluent concentration in comparison with an evaporation process. An analysis

focused on TAC results support the conclusion of which concentration technology and process

conditions represent the lowest total annual cost.

A

B

C

D

E

F

G

0

200

400

600

800

1000

1200

1400

1600

1800

2000

2,5 3,0 3,5 4,0 4,5 5,0 5,5 6,0 6,5 7,0

TA

C (

k€/y

)

TS final (%)

3 Evaporators 4 Evaporators 5 Evaporators 6 Evaporators 7 Evaporators

35

Moreover, environmental effluent treatment constraints are a driving force to look for a better process

integration, capable of not only achieving higher concentration values but also leading to more

sustainable concentration methods. According to Figure 3-8, it is clear the minimal energy requirement

for the membrane filtration process. Consequently, the energy cost influence is significantly lower when

compared to the evaporation process. Therefore, the numerous advantages as the smaller foot print

make membrane filtration technology a potential candidate for the treatment of pulp and paper

wastewater. As so, it is crucial to understand how a membrane filtration system can be a viable option

when integrated in the treatment of the bleaching effluent aiming a higher final product concentration.

Figure 3-5. MF and EVAP CAPEX according to each case study.

Figure 3-6. MF and EVAP OPEX according to each case study.

I

II

III

IV

V

VI

VII

A

B

C

D

E F G

0

100

200

300

400

500

600

700

2,8 3,3 3,8 4,3 4,8 5,3 5,8 6,3 6,8

CA

PE

X (

k€/y

)

TS final (%)

MF EVAP

I

II

III

IV

V

VI

VII

A

B

C

D

E

F

G

0

100

200

300

400

500

600

700

800

900

2,8 3,3 3,8 4,3 4,8 5,3 5,8 6,3 6,8

OP

EX

(k€

/y)

TS final (%)

MF EVAP

36

Figure 3-7. MF and EVAP TAC according to each case study.

Figure 3-8. Energy cost distribution percentage for MF and EVAP according to each case study.

I

II

III

IV

V

VI

VII

A

B

C

D

E

F

G

0

200

400

600

800

1000

1200

1400

1600

2,8 3,3 3,8 4,3 4,8 5,3 5,8 6,3 6,8

TA

C (

k€/y

)

TS final (%)

MF EVAP

6

3

2

2

2

2

2

87

76

79

89

87

89

87

I / A

II / B

III / C

IV / D

V / E

VI / F

VII/ G

Energy cost influence (%) in TAC

Case s

tudy (

MF

/ E

VA

P)

MF EVAP

37

3.3.2 Combined process or evaporation

A comparison between a combined process and a stand-alone evaporation process is performed in light

of finding the techno-economical optimal point of the integrated system. This combined process consists

of a nanofiltration system followed by an evaporation unit. The alternative case enables a reduction of

the required duty of the evaporators. It was analyzed at multiple levels of implementation, i.e., several

levels of preconcentration were performed by the membrane unit prior to the concentration by

evaporators. From the CAPEX interpretation, an interesting phenomenon occurs. When increasing the

desired final concentration, the membrane filtration system presents a higher capital investment, as it is

expected. Whereas the evaporation system tends to show a lower capital investment cost in virtue of

having a more concentrated feed stream from the nanofiltration process. This leads to a lower live steam

requirement since a less amount of evaporated water is needed to achieve a certain final concentration.

Ultimately, from the CAPEX viewpoint an optimal case study resulting from the combined process is

more suitable than the stand-alone evaporation process, as identified in III – 𝛼 as it illustrated in Figure

3-9. Nevertheless, an examination of the OPEX (Figure 3-10) reveals a different conclusion regarding

the optimal point. When setting as a main goal achieving final product concentrations of 30 or 40%

(w/w), only two case studies with the lowest membrane pre-concentration values represent an

economically feasible option when compared with stand-alone evaporation. The remaining case studies

are characterized for having a TAC substantially lower in contrast with the other concentration process.

On the other hand, aiming for a final product concentration of 20% (w/w), the combined process OPEX

is remarkably lower when compared with the evaporation unit operational cost. Therefore, analyzing

also the TAC is crucial, by observing Figure 3-11. The case study represented as being an optimal point,

i.e., showing the lowest TAC when evaluated from a comparison perspective, is the one named as VI –

𝛼. Furthermore, the energy cost influence for the case study (MF TSfinal content of 5,6 % (w/w)) implies

once again that the integrated process has environmental advantages over the standing-alone

evaporation process, leading to a more reduced energy consumption, thus supporting the latter choice.

38

Figure 3-9. MF → EVAP and EVAP CAPEX according to each case study.

Figure 3-10. MF → EVAP and EVAP OPEX according to each case study.

I - α

II - α

III – α

IV – α

V – α

VI – α

VII - α

I - β

II – β

III – β

IV – β

V – β

VI – β

VII - β

I - µ

II – µ

III – µ

IV – µ

V – µ

VI – µ

VII - µ

750

770

790

810

830

850

870

890

910

930

950

2,8 3,3 3,8 4,3 4,8 5,3 5,8 6,3 6,8

CA

PE

X (

k€/y

)

MF final TS (%)

MF+EVAP 20% MF+EVAP 30% MF+EVAP 40% EVAP 20% EVAP 30% EVAP 40%

I - αII - α

III – αIV – α

V – α VI – α

VII - α

I - β

II – βIII – β

IV – β

V – β

VI – βVII - β

I - µ

II – µIII – µ

IV – µ

V – µ VI – µ VII - µ

700

800

900

1000

1100

1200

1300

1400

1500

1600

2,8 3,3 3,8 4,3 4,8 5,3 5,8 6,3 6,8

OP

EX

(k€/y

)

MF final TS (%)

MF+EVAP 20% MF+EVAP 30% MF+EVAP 40% EVAP 20% EVAP 30% EVAP 40%

39

Figure 3-11. MF → EVAP and EVAP TAC according to each case study.

Figure 3-12. Energy cost distribution percentage for MF → EVAP and EVAP according to each case study.

I - α

II - α

III – αIV – α

V – α

VI – α

VII - α

I - β

II – βIII – β

IV – βV – β

VI – β

VII - β

I - µ

II – µ III – µ

IV – µ

V – µ

VI – µ

VII - µ

1500

1600

1700

1800

1900

2000

2100

2200

2300

2400

2500

2,8 3,3 3,8 4,3 4,8 5,3 5,8 6,3 6,8

TA

C (

k€/y

)

MF final TS (%)

MF+EVAP 20% MF+EVAP 30% MF+EVAP 40% EVAP 20% EVAP 30% EVAP 40%

I - αII - α

III – αIV – α

V – α

VI – α

VII - α

I - βII – β

III – βIV – β

V – β

VI – β

VII - β

I - µ II – µ

III – µ IV – µ

V – µ

VI – µ

VII - µ

1000

1500

2000

2500

3000

3500

4000

2,7 3,2 3,7 4,2 4,7 5,2 5,7 6,2 6,7 7,2

TA

C (

k€/y

)

MF final TS (%)

MF+EVAP 20% MF+EVAP 30% MF+EVAP 40% EVAP 20% EVAP 30% EVAP 40%

43

37

36

30

25

21

17

50

48

43

41

34

30

23

51

49

43

41

36

28

24

50

50

50

50

50

50

50

51

51

51

51

51

51

51

52

52

52

52

52

52

52

3,1

3,3

3,7

4,1

4,7

5,6

6,8

Energy cost influence (%)

MF

fin

al

TS

(%

)

MF+EVAP 20% MF+EVAP 30% MF+EVAP 40% EVAP 20% EVAP 30% EVAP 40%

40

3.3.3 Process system price summary

The calculation of the TAC per volume of water removed is feasible and essentially, it provides an unique

analysis enabling a comparison capable of showing which system has the lowest TAC concerning only

the water removed as a concentration performance metric. The concentration process prices are

showed in Table 3-5. However, since the scope of this Thesis is to find an optimal system addressing

both economic and environmental constraints, parameters such as the final TS content and the energy

consumption are the decisive variables.

Table 3-5. TAC per cubic meter of removed water

Process system TAC per volume of removed water (€/m3 H2O)

Process system TAC per volume of removed water (€/m3 H2O)

V (MF) 2,09 V – µ (MF→EVAP)

3,59

VI (MF) 2,12 I - α (MF→EVAP)

3,60

IV (MF) 2,15 VII - µ (MF→EVAP)

3,60

I (MF) 2,26 B (EVAP) 3,63

III (MF) 2,28 VII – β (MF→EVAP)

3,63

II (MF) 2,36 V – β (MF→EVAP)

3,67

VI (MF) 2,69 II – α (MF→EVAP)

3,68

VII (MF) 2,56 IV – µ (MF→EVAP)

3,70

F (EVAP) 2,56 IV – β (MF→EVAP)

3,74

G (EVAP) 2,71 VI – µ (MF→EVAP)

3,75

E (EVAP) 2,74 III – µ (MF→EVAP)

4,17

D (EVAP) 3,19 EVAP – 40% 4,18

VI – α (MF→EVAP) 3,27 II – µ (MF→EVAP)

4,21

C (EVAP) 3,33 III – β (MF→EVAP)

4,22

III – α (MF→EVAP)

3,33 EVAP – 30% 4,24

IV – α (MF→EVAP)

3,35 II – β (MF→EVAP)

4,26

V – α (MF→EVAP)

3,41 EVAP – 20% 4,36

VI – β (MF→EVAP)

3,46 I - µ (MF→EVAP)

4,57

VII – α (MF→EVAP)

2,09 I – β (MF→EVAP)

4,64

A (EVAP) 4,97

41

Techno-economical models

Through the least-squares method, mathematical models were estimated in order to translate the

CAPEX, OPEX and TAC related to the most important configurations analyzed. For the membrane

filtration system, the total solids content concerns the retentate concentration whereas for the

evaporation unit, the variable TS content (𝑤 𝑤⁄ %) describes the final stream concentration of the

process. Finally, regarding the combined process, TS content (𝑤 𝑤⁄ %) represents the retentate

concentration from the membrane filtration system and also the feed concentration for the following

evaporation process.

Membrane filtration system

𝐶𝐴𝑃𝐸𝑋 (𝑘€ 𝑦⁄ ) = 9,56 × 101. 𝑇𝑆 − 1,82 × 102 (3.1)

𝑂𝑃𝐸𝑋 (𝑘€ 𝑦⁄ ) = 2,17 × 102. 𝑇𝑆 − 8,48. 𝑇𝑆2 − 5,11 × 102

(3.2)

𝑇𝐴𝐶 (𝑘€ 𝑦⁄ ) = 4,53 × 102. 𝑇𝑆 − 2,28 × 101. 𝑇𝑆2 − 1,01 × 103

(3.3)

5-MEE Evaporation system

𝐶𝐴𝑃𝐸𝑋 (𝑘€ 𝑦⁄ ) = 6,17 × 103. 𝑇𝑆 − 5,03 × 101. 𝑇𝑆2 + 1,21 × 104 (3.4)

𝑂𝑃𝐸𝑋 (𝑘€ 𝑦⁄ ) = 1,66 × 102. 𝑇𝑆 − 2,31 × 102

(3.5)

𝑇𝐴𝐶 (𝑘€ 𝑦⁄ ) = 1,07 × 103. 𝑇𝑆 − 8,10 × 101. 𝑇𝑆2 + 2,13 × 104 (3.6)

Combined process MF→EVAP TSfinal 20%

𝐶𝐴𝑃𝐸𝑋 (𝑘€ 𝑦⁄ ) = 1,12 × 105. 𝑇𝑆 − 5,12 × 104. 𝑇𝑆2 + 1,15 × 104. 𝑇𝑆3 − 1,25 × 103. 𝑇𝑆4

+ 5,33 × 101. 𝑇𝑆5 − 9,60 × 104 (3.7)

𝑂𝑃𝐸𝑋 (𝑘€ 𝑦⁄ ) = −5,65 × 102. 𝑇𝑆 + 4,17 × 101. 𝑇𝑆2 + 2,24 × 103

(3.8)

𝑇𝐴𝐶 (𝑘€ 𝑦⁄ ) = {

7,19 × 103. 𝑇𝑆 − 1,10 × 103. 𝑇𝑆2 − 9,94 × 103, 𝑇𝑆 ∈ [3,055; 3,657[

2,07 × 103. 𝑇𝑆 − 2,24 × 102. 𝑇𝑆2 + 2,99 × 101, 𝑇𝑆 ∈ [3,657; 4,662[

−1,01 × 103. 𝑇𝑆 + 9,01 × 101. 𝑇𝑆2 + 4,41 × 103, 𝑇𝑆 ∈ [4,662; 6,825]

(3.9)

42

Sensitivity Analysis

The capital and operating costs are based on numerous parameters as it is described in Chapter 2.

Therefore, it is essential the comprehension towards the influence of each variable. That understanding

leads to a preliminary estimation on which variable it will have the major impact on TAC. Importantly, to

fully understand how much a parameter can influence the economic metric a sensitivity analysis is

executed. Such an analysis it is measures the percentage of influence on TAC along with the percentage

related to its cost representation on the TAC of the chosen system, VI – 𝛼.

From a solely membrane filtration evaluation, miscellaneous equipment and spare parts costs, the latter

mostly known as membrane replacement costs, present the highest cost impact. Whereas an

evaporation system economy is mainly dependent on the purchasing equipment cost along with the

expected energy price, expectedly. The numeric parameter cost distribution for each process is

illustrated in Figure 3-13 and Figure 3-14.

Figure 3-13. Process design parameters distribution of CAPEX (right) and OPEX (left) parameters for solely MF, case

study VI - α

Equipment11%

Instrumentation and Control

18%

Piping11%

Miscellaneous Equipment

47%

Others13%

Spare parts 54%

Cleaning 34%

Labor and Maintenance

2%

Energy 10%

Labor and Maintenance

14%

Energy 86%

Equipment44%

Instrumentation and control4%Piping

4%

Miscellaneous equipment

18%

Spare parts2%

Others28%

Figure 3-14. Process design parameters distribution of CAPEX (right) and OPEX (left) parameters for solely EVAP,

case study VI - α

43

It is straightforward by now which parameters should affect the most the combined process, membrane

filtration followed by evaporation. Nevertheless, it is essential to measure the amount of influence on

the economy of the installation. In fact, a sensitivity analysis was carried out, where the parameter

original values experienced a single variation of ±5%, ±10% and ±20% while determining the new value

of case study VI – 𝛼 TAC for each scenario. That allows the calculation of how much in percentage the

TAC varies when a given parameter also varies, without changing the remaining process variables.

Moreover, it is also exhibited the numeric parameter distribution, which necessarily corresponds with

the previous one, coupled with the percentage values that support the aim of the sensitivity analysis.

Eminently, the steam price required for the evaporation plant, when experiencing a 10% variation, it has

the ability of varying approximately 3% the TAC of the system. That can have a tremendous impact on

the economy of the pulp mill as a whole. Secondly, the spare parts of a membrane filtration system

followed by the evaporation equipment are the secondary parameters to be taken into account. As a

result, TAC varies 1,2% and approximately 1%, respectively, when changed only 10% from the initial

values. The influence percentages of each process variable are represented in Figure 3-15 and Figure

3-16 and also listed in Table A-33.

0,10

0,21

0,41

0,17

0,35

0,69

0,10

0,21

0,42

0,46

0,92

1,84

±5%

±10%

±20%

Process design parameter influence (%) in TAC

Variatio

n f

rom

base c

ase (

%)

Equipment (MF) Instrumentation and Control (MF)

Piping (MF) Miscellaneous Equipment (MF)

0,62

1,24

2,47

0,40

0,80

1,59

0,02

0,04

0,08

0,12

0,23

0,47

±5%

±10%

±20%

Process design parameter influence (%) in TAC

Variatio

n f

rom

base c

ase (

%)

Spare parts (MF) Cleaning (MF) Labor and Maintenance (MF) Electricity (MF)

0,48

0,96

1,92

0,05

0,10

0,19

0,05

0,10

0,19

0,19

0,38

0,77

0,02

0,04

0,08

±5%

±10%

±20%

Process design parameters influence (%) in TAC

Variatio

n f

rom

base c

ase (

%)

Equipment (EVAP) Instrumentation and Control (EVAP)

Piping (EVAP) Miscellaneous Equipment (EVAP)

Spare parts (EVAP)

0,25

0,51

1,02

1,52

3,04

6,09

±5%

±10%

±20%

Process design parameters influence (%) in TAC

Variatio

n f

rom

base c

ase (

%)

Labor and Maintenance EVAP) Steam Price (EVAP)

Figure 3-15. Process design parameters influence regarding MF in TAC for MF → EVAP, case study VI - α

Figure 3-13. Process design parameters influence regarding EVAP in TAC for MF → EVAP, case study VI - α

44

Pulp and paper industry case study: Caima, Indústria da

Celulose

The treatment of wastewater effluents using multi-effect evaporator systems as a concentration method

has been largely present in the pulp and paper industry. Accordingly, the introduction of a membrane

filtration unit should complement the already installed concentration plant instead of replacing it entirely.

That being the case, brings the necessity of performing a distinct economic evaluation and sensitivity

analysis, building up on the steps described in the previous Chapters.

When an evaporation plant is already installed, the number of evaporators is established, and the

associated CAPEX is no longer a determinant economic variable. According to available data regarding

a bleaching effluent from the pulp mill Caima, Indústria da Celulose a 5 multi-effect evaporator system

is implemented to concentrate the alkaline extraction stage effluent. The associated stream conditions

and specifications were also used to support the general case. Therefore, the definition of the most

suitable system from a techno-economic viewpoint for the present particular case is the analysis subject

of the observation of Figure 3.17. It is clear that the case study I – α represents the economically

preferential option, which implies a final total solid concentration of 3,1% (w/w) for the membrane

filtration as being the feed to the evaporation system, enabling a final process total solids content of

20% (w/w).

Figure 3-14. MF → EVAP and EVAP TAC according to each case study for the PPI case study example.

I - α II - α

III – α IV – α

V – α

VI – α

VII - α

I - βII – β

III – βIV – β

V – β

VI – β

VII - β

I - µ II – µIII – µ IV – µ

V – µ

VI – µ

VII - µ

0

500

1000

1500

2000

2500

3000

3500

4000

2,7 3,2 3,7 4,2 4,7 5,2 5,7 6,2 6,7 7,2

TA

C (

k€

/y)

MF final TS (%)

MF+EVAP 20% MF+EVAP 30% MF+EVAP 40% EVAP 20% EVAP 30% EVAP 40%

45

Once again, a sensitivity analysis is performed, and it is focused on the industrial case of the pulp mill

located in Caima. The stream price greatly influences the total annual costs of the concentration

technology, where a variation of 10% leads to an increase or decrease of TAC of approximately 4%.

Most impressively, a variation of 20% in steam price could vary by almost 7% the TAC. In fact, steam

price is a variable of great concern and its variation is described in the following Chapter. In addition,

the capital cost parameters related to the purchase of the membrane filtration system can also have an

equal impact on the TAC by varying it around 2%.

1,7

1,7

1,8

1,7

1,7

1,8

1,7

1,8

1,8

1,8

1,9

2,1

±5%

±10%

±20%

Process design parameter influence (%) in TAC

Variatio

n f

rom

base c

ase (

%)

Equipment (MF) Instrumentation and Control (MF)

Piping (MF) Miscellaneous Equipment (MF)

0,3

0,6

1,1

0,03

0,1

0,1

0,005

0,01

0,02

0,03

0,1

0,1

±5%

±10%

±20%

Process design parameter influence (%) in TAC

Variatio

n f

rom

base c

ase (

%)

Spare parts (MF) Cleaning (MF) Labor and Maintenance (MF) Electricity (MF)

Figure 3-18. Process design parameters influence regarding MF in TAC for MF → EVAP, case study I – α for PPI case study

example

0,5

1,0

2,0

1,8

3,6

7,1

±5%

±10%

±20%

Process design influence (%) in TAC

Variatio

n f

rom

base c

ase (

%)

Labor and Maintenance EVAP) Steam Price (EVAP)

Figure 3-19. Process design parameters influence regarding EVAP in TAC for MF → EVAP, case study I – α for

PPI case study example

46

Techno-economic results discussion

Research investigating the potential of the combined process, membrane filtration system followed by

an evaporation plant, with the purpose of concentrating a certain flow has been greatly conducted.

However, the novelty of this Thesis is the main focus on the techno-economic analysis based on results

from a previous study investigated by (São Pedro, 2016), concerning a bleach effluent from Caima,

Indústria da Celulose. Although those results were related to a plate and frame module, an analysis

using a spiral wound configuration is valid due to the similar flat-sheet structure of both modules. One

of the reasons why a spiral-wound module was chosen to support the present Thesis results is because

of its considerably lower cost in comparison to plate and frame costs (Samhaber & Nguyen, 2014). More

assumptions were made, particularly the one considering that the flow operational conditions remain the

same along the membrane filtration system which is not always correct. In addition to the latter, with the

fact that the pressure drop is negligible on the feed side.

A downside of using spiral-wound nanofiltration membrane is the necessity of installing a pre-treatment

technology to prevent membrane fouling, enabling also longer operation times without cleaning. Some

very efficient pre-treatment methods have been already installed in the pulp and paper industry, which

includes sand filter, back-washable screen filter and biologically treatment as an activated sludge plant

(Nystro, 2007). Another more conventional pre-treatment processes involve coagulation/flocculation,

sedimentation and media filtration. Furthermore, an innovative spiral-wound nanofiltration pre-treatment

study was carried by (Geraldes, Anil, de Pinho, & Duarte, 2008) in which the effectiveness of dissolved

air flotation (DAF) was investigated. The results indicated that, in the range of operating conditions

studied, another distinct treatment capable of removing residual suspended particles and colloidal

matter should be integrated. In general, according to (Gorenflo & Frimmel, 2002), nanofiltration plants

pre-treatment costs are estimated to be in the same range as costs for the NF process itself, however

specific cost data was not find in literature.

Therefore, a cost estimation was conducted regarding a sand filter system, in light of fully understanding

the economic impact of that technology. The capital cost calculated from an equation expressed in

(Rural & Cowi, 2006) is tantamount to the capital cost estimated from (Slezak & Sims, 1984), as being

29762€, considering the operational feed conditions. The latter amount does not insert within the range

of the NF plant costs, in fact it is significantly lower. In fact, it affects only in 6% and in 2% the CAPEX

of the case studies I and III, respectively. Therefore, the implementation of a pre-treatment as sand filter

would not compromise the economic advantages of installing the optimal chosen system.

47

Regarding the nanofiltration module used to obtain the results, a major concern must be underlined.

Although pH and temperature adjustments are generally not a problem when using a membrane filtration

system, since innumerous different membrane products with distinct operational specifications are

available in the market, the membrane module applied in this case can be a concern. The bleaching

effluent comes from an alkaline stage thus its pH is approximately 10 which is the maximum value

allowed for the membrane operate successfully. In addition to that, (São Pedro, 2016) concluded that

the optimal operating temperature is 50 ℃ whereas 45 ℃ is the temperature limit of the membrane in

this project considered, which can affect the membrane performance results when applying the

suggested membrane to the effluent treatment unit.

The economic model of (Sethi, 1997) has some advantages regarding to those divulged in literature for

the fact of being based on real published data and for considering an economy of scale. Although the

model addresses ultrafiltration installations, authors such as (Costa & de Pinho, 2006) used it to estimate

nanofiltration plant costs, thereby validating the model. The consequent techno-economic investigation

presented the impact of the membrane filtration parameters. That analysis revealed the challenge of

keeping membrane costs low, because of the frequent need for membrane replacement, expressed as

spare parts, that is associated with these applications. In addition, the miscellaneous equipment factor

cost is highly case dependent. However, when installing a membrane unit into an existing facility, the

cost of buildings, construction and electrical supply is plausibly lower in contrast with miscellaneous

equipment cost associated to a completely new plant. In fact, the industrial case study example was

analyzed considering the latter factor however for further investigations it is suggested to be reassessed.

Another critical factor is the influence of the energy consumption and, consequently, the effect that the

steam price has in the total annual cost of the evaporation system. Hence it is crucial analyzing briefly

how the steam is generated and its price dependency. A industrial plant can have multiples steam

sources and several fuels, as so determining the true cost of steam becomes far more complex (Kumana

& Associates, 2003). In most companies, the reported cost of steam is the average cost of generation

at a particular production rate and it is affected by innumerous parameters, the fuel cost, the operating

steam pressure and even the boiler efficiency (Swagelok Energy Advisors, 2011).

The steam cost can then vary immensely, although its effect differs according to each system. With

regard to a standing-alone evaporation plant, an additional sensitivity analysis revealed that a steam

price variation of 10% would lead to a TAC variation of 8%, more than twice as much as comparing to

the combined system (3%). Moreover, a fascinating research performed by (Wang, Tung, & Ward, 2017)

shows that, in case of a significant increase of the steam price, the membrane filtration unit followed by

a multi-effect evaporation process was found to have the lowest cost. Positively, this Thesis results are

in line with other studies, reinforcing one of the advantages of opting for choosing a membrane filtration

system followed by an evaporation installation over to an unique evaporation plant regardless the

influence of steam price.

48

Process design and operational parameters economic impact the influence expected to be found in the

total annual cost of each system. In fact, even though it might not be clear through Figure 3-9, the

evaporator capital cost slightly decreases as the retentate concentration increases. Whereas the

membrane filtration system capital cost substantially increases. Firstly, the evaporation unit cost

decreases because the equipment is designed according to the amount of water needed to be removed.

It means exactly the same as being sized according to the final concentration required. With higher

retentate concentration as feed to the evaporation system, a lower heat transfer area is required, which

is the design parameter that determines the equipment cost. On the other hand, higher retentate

concentration values increase membrane area requirements, leading to a higher membrane filtration

cost.

Undoubtedly, a membrane filtration as a concentrator unit upstream of evaporation enables the

reduction of the total energy process requirements, as it is illustrated in Figure 3-8 and Figure 3-12.

However, the large CAPEX of the initial membrane installation and operational and maintenance costs,

which includes membrane replacement, are both parameters capable of greatly reducing the process

economic advantages. So continuous efforts on improving membrane performance along with cost are

mandatory in the pursuing of having a membrane system as a cost-effective concentration process.

The sensitivity analysis that was performed showed quantitively the importance and impact of some

process parameters involved in the design and maintenance. Regarding membrane filtration unit the

system coupled with the membrane replacement cost represent the highest expense. While, concerning

the evaporation plant, steam price along with equipment purchase cost present the most significant

costs. Although numerous conclusions can be highlighted, another study focused on more process

design parameters would complement immensely the present results but was not feasible due to the

limited data available. Suárez et al., 2015 carried out a reverse osmosis sensitivity analysis addressing

different process parameters. The results from that study show the great influence of the plant capacity,

the permeate flux and the outlet temperature have on the CAPEX. Therefore, in case of a future analysis,

an evaluation integrating those parameters as well is suggested.

49

Chapter 4. Conclusion and future work

Compared to the literature, the results from the developed model showed good accuracy in predicting

the feasibility of a nanofiltration unit followed by an evaporation plant set up. For the innumerous case

studies analyzed both CAPEX and OPEX calculated values were in good agreement with those reported

in the field of pulp and paper industry. Regarding the performance and cost results, the use of membrane

concentrator upstream to the evaporation unit reduces the size and energy use of the subsequent

evaporator. That fact provides economic improvements for different case studies. The most suitable

case study is VI – α, where the retentate is concentrated until 5,6% (w/w) and the final product

concentration is 20% (w/w). For the particular case of the pulp mill in Caima, the only economic viable

scenario is the one in which the total solid content of the retentate is 3,1% (w/w) and the final product

concentration of the combined process is 20% (w/w).

Membrane filtration as a solo unit is indeed economically arduous to an industrial company to

considerer. However, when integrated with distinct technologies, as evaporation plants, it can bring

several advantages. It is a compact system, easy to install in industrial plants. It can also be used as a

process booster, capable of rapidly concentrate an effluent, increasing the process capacity while

significantly reducing the energy consumption requirement when compared to other concentration

methods.

One of the concerns of the wastewater management is definitely meeting the several environmental

constraints by applying innovative technologies capable of eliminating the organic matter of the effluents

in the most suitable way. Moreover, from an economic management viewpoint, there are several

benefits of membrane filtration processes in lignin and hemicellulose extraction, since it does not require

pH or temperature adjustment. Furthermore, it is possible to control both component fraction molecular

mass by membrane cut-off. In chemical pulp mills, lignin and hemicelluloses can be extracted and

afterwards used in several applications, as mentioned in Chapter 2. Hemicellulose fermentation or

hydrolysis leads to value-added products, such as ethanol and xylitol whose significance in food and

pharmaceutic industry is unquestionable. While lignin, in addition to several related value-added

products, can be used as an external fuel. However, the profitability of lignin and hemicelluloses

extraction depends to a great extent on the process design parameters, specially the steam/electricity

price and whether the quality of lignin and hemicelluloses fulfills the demands for the related end-

products.

Several issues remain to be further investigated, however, the results of this study offer a starting point

for further optimization of concentration processes in the pulp and paper industry. Prospectively, the

development of a data-model capable of analyzing simultaneously experimental and economic data

would be extremely interesting. The least squares method provided a first estimation of a model that

predicts the obtained techno-economical results for each concentration process. Howsoever, more data

should be included and tested to fully validate those model conclusions.

50

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51

Chapter 5. References

Alexandri, M., Papapostolou, H., Komaitis, M., Stragier, L., Verstraete, W., Danezis, G. P., … Koutinas,

A. A. (2016). Evaluation of an integrated biorefinery based on fractionation of spent sulphite liquor

for the production of an antioxidant-rich extract, lignosulphonates and succinic acid. Bioresource

Technology, 214, 504–513. https://doi.org/10.1016/j.biortech.2016.03.162

Arkell, A., Krawczyk, H., Thuvander, J., & Jönsson, A. S. (2013). Evaluation of membrane performance

and cost estimates during recovery of sodium hydroxide in a hemicellulose extraction process by

nanofiltration. Separation and Purification Technology, 118, 387–393.

https://doi.org/10.1016/j.seppur.2013.07.015

Bai, F., Xiong, L., & Wang, Q. (2017). Accepted Manuscript, 11–19.

https://doi.org/10.1016/j.neubiorev.2017.04.006

Bajpai, P. (2013). Chapter 4 - Products from Hemicelluloses*. Biorefinery in the Pulp and Paper Industry.

https://doi.org/http://dx.doi.org/10.1016/B978-0-12-409508-3.00004-3

Bajpai, P. (2017a). External Treatment Technologies Used for Pulp and Paper Mill Effluents. Pulp and

Paper Industry, 71–92. https://doi.org/10.1016/B978-0-12-811099-7.00006-X

Bajpai, P. (2017b). Introduction. Pulp and Paper Industry, 1–8. https://doi.org/10.1016/B978-0-12-

811099-7.00001-0

Baker, R. W. (2004). Membrane Technology and Applications. Membrane Technology and Applications.

https://doi.org/10.1016/S0376-7388(00)83139-7

Berg, P. L. (2017). Pulp, paper, and packing in the next decade: Transformation change. Paper & Forest

Products McKinsey&Company. Retrieved from https://www.mckinsey.com/industries/paper-and-

forest-products/our-insights/pulp-paper-and-packaging-in-the-next-decade-transformational-

change

Bonnin, E., Ralet, M. C., Thibault, J. F., & Schols, H. A. (2009). Enzymes for the valorisation of fruit- and

vegetable-based co-products. Handbook of Waste Management and Co-Product Recovery in

Food Processing (Vol. 2). Woodhead Publishing Limited.

https://doi.org/10.1533/9781845697051.3.257

Bulletin, T. (2017). Spiral-Wound Elements, 19–20.

Bungay, E. P. ., Lonsdale, H. K., & de Pinho, M. N. (1986). Synthetic Membranes: Science, Engineering

and Applications. (N. A. S. C. M. and C. Science, Ed.). D. Reidel Publishing Company.

Carminati, R. (2017). Market&trends: the pulp and paper sector is not dead yet. Paper Industry World.

Retrieved from http://www.paperindustryworld.com/markettrends-pulp-paper-sector-not-dead-yet/

CEPI (Confederation of European Paper Industries). (2017). Key Statistics 2016.

Http://Www.Cepi.Org/Keystatistics2016, 1–31. Retrieved from https://www.e-

control.at/documents/20903/443907/Statbro_englisch_FINAL.pdf/0feaeaa0-d46f-4496-a498-

52

39a81e5e67a2

Chen, H. (2014). Biotechnology of lignocellulose: Theory and practice. Biotechnology of Lignocellulose:

Theory and Practice. https://doi.org/10.1007/978-94-007-6898-7

Cheryan, M. (1998). Ultrafiltration and Microfiltration Handbook. Pennsylvania: Western Hemisphere.

Cline, S. P., & Smith, P. M. (2017). Opportunities for lignin valorization: an exploratory process. Energy,

Sustainability and Society, 7(1). https://doi.org/10.1186/s13705-017-0129-9

Costa, A. R., & de Pinho, M. N. (2006). Performance and cost estimation of nanofiltration for surface

water treatment in drinking water production. Desalination, 196(1–3), 55–65.

https://doi.org/10.1016/j.desal.2005.08.030

Ebrahimi, M., Busse, N., Kerker, S., Schmitz, O., Hilpert, M., & Czermak, P. (2015). Treatment of the

bleaching effluent from sulfite pulp production by ceramic membrane filtration. Membranes, 6(1),

1–15. https://doi.org/10.3390/membranes6010007

Energy, S., & Ise, S. (2017). Base-catalyzed Reformation of Kraft Lignin for Base-catalyzed Reformation

of Kraft Lignin for Carbon Fiber Precursors In partial fulfillment of a Master ´ s of Science Chair of

Forest Biomaterials, (March). https://doi.org/10.13140/RG.2.2.23616.12807

Engineering, G. P. (2018). Falling Film Evaporator. Retrieved July 22, 2018, from

https://www.gea.com/pt/products/falling-film-evaporator.jsp

Epa, U. S. (2005). Note on the Membrane Filtration Guidance Manual. Environmental Protection.

Figoli, A., Cassano, A., & Basile, A. (2016). Mmebrane Technologies for Biorefining.

Gallagher, B. (2018). EVAPS - High-Performance Evaporation Systems. Retrieved July 12, 2018, from

https://businessdocbox.com/Green_Solutions/78099852-Evaps-high-performance-evaporation-

systems.html

Geraldes, V., Anil, A., de Pinho, M. N., & Duarte, E. (2008). Dissolved air flotation of surface water for

spiral-wound module nanofiltration pre-treatment. Desalination, 228(1–3), 191–199.

https://doi.org/10.1016/j.desal.2007.10.008

Gleadow, P., & Stratton, S. (2003). NATIONAL COUNCIL FOR AIR AND STREAM IMPROVEMENT

PULP MILL PROCESS CLOSURE : A REVIEW OF GLOBAL TECHNOLOGY DEVELOPMENTS

AND MILL EXPERIENCES IN THE 1990s.

Gorenflo, A., & Frimmel, F. H. (2002). Nanofiltration of a German groundwater of high hardness and

NOM content : performance and costs, 1, 253–265.

Guerra, K., & Pelligrino, J. (2012). Investigation of Low-Pressure Membrane Performance, Cleaning,

and Economics Using a Techno-Economic Modeling Approach, (174).

Johssen, B. (2017). The Pulp and Paper Industry. Physical Review Letters, 1(3), 501–509.

https://doi.org/10.1021/ie50507a025

Jönsson, A. S., Nordin, A. K., & Wallberg, O. (2008). Concentration and purification of lignin in hardwood

53

kraft pulping liquor by ultrafiltration and nanofiltration. Chemical Engineering Research and Design,

86(11), 1271–1280. https://doi.org/10.1016/j.cherd.2008.06.003

Judd, S., & Jefferson, B. (2003). Membranes for Industrial Wastewater Recovery and Re-use.

Kaya, D., & Ibrahim Sarac, H. (2007). Mathematical modeling of multiple-effect evaporators and energy

economy. Energy, 32(8), 1536–1542. https://doi.org/10.1016/j.energy.2006.09.002

Khanam, S., & Mohanty, B. (2011). Development of a new model for multiple effect evaporator system.

Computers and Chemical Engineering, 35(10), 1983–1993.

https://doi.org/10.1016/j.compchemeng.2010.11.001

Kumana & Associates. (2003). How To Calculate The True Cost of Steam. Wshington.

Kumar, D., Kumar, V., & Singh, V. P. (2013). Modeling and dynamic simulation of mixed feed multi-

effect evaporators in paper industry. Applied Mathematical Modelling, 37(1–2), 384–397.

https://doi.org/10.1016/j.apm.2012.02.039

Li, Y., & Watkinson, A. P. (2011). Deposit formation in evaporation of a pulp mill effluent. Heat Transfer

Engineering, 32(3–4), 258–263. https://doi.org/10.1080/01457632.2010.495614

Livinec, M. (2018). What to Watch? Unpack growth from global trade and e-commerce.

Marques, A. P., Evtuguin, D. V., Magina, S., Amado, F. M. L., & Prates, A. (2009). Chemical Composition

of Spent Liquors from Acidic Magnesium–Based Sulphite Pulping of Eucalyptus globulus. Journal

of Wood Chemistry and Technology, 29(4), 322–336.

https://doi.org/10.1080/02773810903207754

Menon, V., Rao, M., & Prakash, G. (2010). Value added products from hemicellulose - Biotechnological

perspective. Global Journal of Biochemistry (Vol. 1).

https://doi.org/10.1017/CBO9781107415324.004

Mulder, M. (1996). Basic Principles of Membrane Technology (first). Dordrecht, the Netherlands: Kluwer

Academic Publishers.

Nilsson, M., Lipnizki, F., Trägårdh, G., & Östergren, K. (2008). Performance, energy and cost evaluation

of a nanofiltration plant operated at elevated temperatures. Separation and Purification

Technology, 60(1), 36–45. https://doi.org/10.1016/j.seppur.2007.07.051

Nordin, A. K. (2008). Process Design for Ultrafiltration of Complex Process Streams. Lund University.

Nordin, A. K., & Jönsson, A. S. (2008). Optimisation of membrane area and energy requirement in

tubular membrane modules. Chemical Engineering and Processing: Process Intensification, 47(7),

1090–1097. https://doi.org/10.1016/j.cep.2007.09.010

Nystro, M. (2007). Membrane filtration for tertiary treatment of biologically treated effluents from the pulp

and paper industry. https://doi.org/10.2166/wst.2007.217

O’Connell, E. (2017). Towards Adaptation of Water Resource Systems to Climatic and Socio-Economic

Change. Water Resources Management, 31(10), 2965–2984. https://doi.org/10.1007/s11269-017-

54

1734-2

Obydenkova, S. V., Kouris, P. D., Hensen, E. J. M., Heeres, H. J., & Boot, M. D. (2017). Environmental

economics of lignin derived transport fuels. Bioresource Technology, 243, 589–599.

https://doi.org/10.1016/j.biortech.2017.06.157

Oeveren, P. V, V, H. I. N., Congress, W. B., & Diego, S. (2004). Best Available Techniques, (July), 1–

18. https://doi.org/10.1016/B978-0-12-811099-7.00003-4

Pabby K., A., Rizvi S.H., S., & Sastre, A. M. (2009). Handbook of Membrane Separations.

Pereira, G. C. Q., Braz, D. S., Hamaguchi, M., Ezeji, T. C., Maciel Filho, R., & Mariano, A. P. (2018).

Process design and economics of a flexible ethanol-butanol plant annexed to a eucalyptus kraft

pulp mill. Bioresource Technology, 250(September 2017), 345–354.

https://doi.org/10.1016/j.biortech.2017.11.022

Peters, M. S., & Timmerhaus, K. D. (1991). Plant Desing and Economics for Chemical Engineers.

Peters, S. M., & Timmerhaus, D. K. (1991). Plant Design and Economics for Chemicals Engineers

(Fourth Edi).

Phitsuwan, P., Sakka, K., & Ratanakhanokchai, K. (2013). Improvement of lignocellulosic biomass in

planta: A review of feedstocks, biomass recalcitrance, and strategic manipulation of ideal plants

designed for ethanol production and processability. Biomass and Bioenergy, 58, 390–405.

https://doi.org/10.1016/j.biombioe.2013.08.027

Rajendran, K., Rajoli, S., Teichert, O., & Taherzadeh, M. J. (2015). Impacts of retrofitting analysis on

first generation ethanol production: Process design and techno-economics. Bioprocess and

Biosystems Engineering, 38(2), 389–397. https://doi.org/10.1007/s00449-014-1278-2

Rural, T., & Cowi. (2006). Appendix 2 : Documentation of Expenditure Functions - Water Supply. Water

Supply, 1–9. Retrieved from https://www.oecd.org/env/outreach/36228967.pdf

Samhaber, W. M., & Nguyen, M. T. (2014). Applicability and costs of nanofiltration in combination with

photocatalysis for the treatment of dye house effluents. Beilstein Journal of Nanotechnology, 5,

476–484. https://doi.org/10.3762/bjnano.5.55

São Pedro, T. (2016). Ultrafiltration and Nanofiltration of E-stage bleaching plant effluent of a sulphite

pulp mill, (December).

Schaefer, K. (2016). Outlook for the World Pulp Market. Brazil.

Schwantes, R., Chavan, K., Winter, D., Felsmann, C., & Pfafferott, J. (2018). Techno-economic

comparison of membrane distillation and MVC in a zero liquid discharge application. Desalination,

428(October 2017), 50–68. https://doi.org/10.1016/j.desal.2017.11.026

Sethi, S. (1997). Transient permeate flux analysis, cost estimation, and Design Optimization in

Crossflow Membrane Filtration.

Slezak, L. A., & Sims, R. C. (1984). Application and Effectiveness of Slow Sand Filtration in the United

55

States. Journal / American Water Works Association, 76(12), 38–43.

SPX Corporation. (2008). Evaporator Handbook, 70.

Statista. (n.d.). Paper Industry.

Strathmann, H., Giorno, L., & Drioli, E. (2006). An Introduction to Membrane Science and Tecnhnolgy.

Rend, Italy: Institute on Membrane Technology.

Suárez, A., Fernández, P., Iglesias, J., & Igesias, E. (2015). Cost assessment of membrane processes:

A practical example in the dairy waswater reclamation by reverse osmosis. Journal of Membrane

Science, 493, 389–402.

Swagelok Energy Advisors. (2011). Knowing the Cost of Steam, (31), 4. Retrieved from

http://chicago.swagelok.com/Services/Energy-Services/~/media/Distributor Media/C-

G/Chicago/Services/ES - Knowing Cost of Steam_BP_31.ashx

Symister, O. J. (2016). An Analysis of Capital Cost Estimation Techniques for Chemical Processing.

Thuvander, J. (2018). Recorvery of Hemicelluloses Extracted from Spruce and Wheat Bran. Lund

University\\.

Thuvander, J. (2018). Recovery of Hemicelluloses Extracted from Spruce and Wheat bRAN. Lund

University.

U.S. Environmental Protection Agency. (2001). Pulp and Paper Combustion Sources National Emisson

Standards for Hazardous Air Pollutants. Chicago.

W. Baker, R. (2012). Membrane Technology and Applications (Second edi). California: John Wiley &

Sons, Ltd.

Wagner, J. (2001). Membrane Filtration Handbook: Practical Tips and Hints, 127.

https://doi.org/10.1007/s13398-014-0173-7.2

Wang, H. Y., Tung, K. L., & Ward, J. D. (2017). Design and economic analysis of membrane-assisted

crystallization processes. Journal of the Taiwan Institute of Chemical Engineers, 81, 159–169.

https://doi.org/10.1016/j.jtice.2017.09.023

WWAP. (2018). The United Nations Worls Water Development Report 2018: Nature-Based solutions

for water. Paris.

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57

Appendix A

This appendix includes data related to techno-economical results. The analyzed nanofiltration system

concerning all the case studies can differ considerably based on some design parameters, such as

volume reduction, average flux and permeate flow. Those parameters lead to the required membrane

area as well as the safety membrane area, which in turn, enables the determination of the number of

membrane elements, as it is described in Table A-1. Additionally, a nomenclature is given and

expressed in Table A-2. in accordance with the configuration proposed for situation 1 and 2 and following

the example illustrated in Figure A-1.

Table A-1. Process and design parameters for each membrane filtration case study.

Case VR (%)

𝑱𝒂𝒗𝒆𝒓𝒂𝒈𝒆

(𝑳. 𝒉−𝟏. 𝒎−𝟐)

𝑸𝒑

(𝒎𝟑. 𝒉−𝟏)

𝑨𝒎

(𝒎𝟐)

𝟏, 𝟐. 𝑨𝒎

(𝒎𝟐)

Membrane elements

I 10 59 7 119 142 4

II 20 33 14 423 508 14

III 30 26 21 803 964 27

IV 40 25 28 1138 1365 37

V 50 22 35 1569 1883 51

VI 60 17 42 2482 2979 81

VII 68 13 48 3609 4331 118

Figure A-1. Example of a multistage plant design, with 3 loops where the first two are composed by two housings

(a), each one containing three membrane elements (b), (2×3), while the third one has only one housing with two

membrane elements, (1×2). Adapted from (Nilsson et al., 2008).

58

Table A-2. Membrane filtration system configuration according to the case study, expressing the number of loops

and the number of housings (a) and membrane elements (b) for each loop.

Situation 1 Situation 2

Case Loops (𝒂 × 𝒃) Loops (𝒂 × 𝒃)

I 1 (4 × 1) 1 (2 × 3)

II 2 (7 × 1) 1 (5 × 3)

III 3

1

(8 × 1)

(3 × 1)

1

1

(8 × 3)

(1 × 3)

IV 4

1

(8 × 1)

(5 × 1)

1

1

(8 × 3)

(5 × 3)

V 6

1

(8 × 1)

(3 × 1)

2

1

(8 × 3)

(1 × 3)

VI 11 (8 × 1) 3

1

(8 × 3)

(3 × 3)

VII 15 (8 × 1) 5 (8 × 3)

59

On the other contrary, data related to the multiple-effect evaporator system was obtained through a

simulation using a MATLAB code developed for the present Thesis and is expressed as it follows.

60

61

The economic evaluation concerning the case studies is listed in the following tables. As matter of

example regarding the evaporation costs, only evaporation unit systems aiming a final product

concentration of 20% are listed, since they represent the configuration of highest economic interest.

Table A-3. Economic variables used for cost evaluation.

Name Description Value

𝒊 Annual interest rate 0,05

𝑵 Investment period (y) 10

𝑪𝑹𝑭 Capital recovery factor 0,130

US $ to € Currency conversion factor 0,87

CEPCI

Chemical engineering plant

cost index

662 (2017)

576 (2014)

392 (1990)

Case study MF I

Table A-4. Capital costs for case study I.

Name Description Cost (€)

𝑪𝒎𝒆𝒎 Set of membranes 47953

𝑪𝑷𝑽 Pipes and valves 84036

𝑪𝑰𝑪 Instrumentation and controls 67392

𝑪𝑻𝑭 Tanks and frames 74560

𝑪𝑴𝑰 Miscellaneous equipment 234668

𝑪𝒑,𝒇 Feed pump 558

𝑪𝒑,𝒓 Recirculation pump 1673

𝑪𝑪𝑨𝑷 Total capital cost 509989

Annual CAPEX 66046

62

Table A-5. Operational parameters and costs for case study I.

Name Description Cost (€)

𝑪𝑨𝒏𝒏𝒖𝒂𝒍 𝒓𝒆𝒑 Membrane replacement cost 37261

𝑾𝒇𝒆𝒆𝒅 (𝒌𝑾. 𝒉) Feed pump power 36

𝑾𝒓𝒆𝒄𝒊𝒓 (𝒌𝑾. 𝒉) Recirculation pump power 17

𝑬 𝑪𝒑,𝒇 Energy cost of feed pump work 6066

𝑬 𝑪𝒑,𝒓 Energy cost of feed pump work 2806

𝑪𝒄𝒍𝒆𝒂𝒏𝒊𝒏𝒈 Membrane cleaning cost 7124

𝑪𝒍𝒂𝒃𝒐𝒓 Labor cost

1321

𝑪𝒎𝒂𝒊𝒏 Maintenance cost

Annual OPEX (€) 54577

Case study MF II

Table A-6. Capital costs for case study II.

Name Description Cost (€)

𝑪𝒎𝒆𝒎 Set of membranes 102818

𝑪𝑷𝑽 Pipes and valves 143332

𝑪𝑰𝑪 Instrumentation and controls 155950

𝑪𝑻𝑭 Tanks and frames 146256

𝑪𝑴𝑰 Miscellaneous equipment 484333

𝑪𝒑,𝒇 Feed pump 822

𝑪𝒑,𝒓 Recirculation pump 883

𝑪𝑪𝑨𝑷 Total capital cost 1034394

Annual CAPEX 133959

63

Table A-7. Operational parameters and costs for case study II.

Name Description Cost (€)

𝑪𝑨𝒏𝒏𝒖𝒂𝒍 𝒓𝒆𝒑 Membrane replacement cost 79892

𝑾𝒇𝒆𝒆𝒅 (𝒌𝑾. 𝒉) Feed pump power 36

𝑾𝒓𝒆𝒄𝒊𝒓 (𝒌𝑾. 𝒉) Recirculation pump power 42

𝑬 𝑪𝒑,𝒇 Energy cost of feed pump work 6066

𝑬 𝑪𝒑,𝒓 Energy cost of feed pump work 7006

𝑪𝒄𝒍𝒆𝒂𝒏𝒊𝒏𝒈 Membrane cleaning cost 25397

𝑪𝒍𝒂𝒃𝒐𝒓 Labor cost

2679

𝑪𝒎𝒂𝒊𝒏 Maintenance cost

Annual OPEX (€) 121040

Case study MF III

Table A-8. Capital costs for case study III.

Name Description Cost (€)

𝑪𝒎𝒆𝒎 Set of membranes 150986

𝑪𝑷𝑽 Pipes and valves 187565

𝑪𝑰𝑪 Instrumentation and controls 237981

𝑪𝑻𝑭 Tanks and frames 205360

𝑪𝑴𝑰 Miscellaneous equipment 697702

𝑪𝒑,𝒇 Feed pump 822

𝑪𝒑,𝒓 Recirculation pump 1673

𝑪𝑪𝑨𝑷 Total capital cost 1482089

Annual CAPEX 191937

64

Table A-9. Operational parameters and costs for case study III.

Name Description Cost (€)

𝑪𝑨𝒏𝒏𝒖𝒂𝒍 𝒓𝒆𝒑 Membrane replacement cost 144595

𝑾𝒇𝒆𝒆𝒅 (𝒌𝑾. 𝒉) Feed pump power 36

𝑾𝒓𝒆𝒄𝒊𝒓 (𝒌𝑾. 𝒉) Recirculation pump power 75

𝑬 𝑪𝒑,𝒇 Energy cost of feed pump work 6066

𝑬 𝑪𝒑,𝒓 Energy cost of feed pump work 12600

𝑪𝒄𝒍𝒆𝒂𝒏𝒊𝒏𝒈 Membrane cleaning cost 48184

𝑪𝒍𝒂𝒃𝒐𝒓 Labor cost

3839

𝑪𝒎𝒂𝒊𝒏 Maintenance cost

Annual OPEX (€) 241702

Case study MF IV

Table A-10. Capital costs for case study IV.

Name Description Cost (€)

𝑪𝒎𝒆𝒎 Set of membranes 186088

𝑪𝑷𝑽 Pipes and valves 217119

𝑪𝑰𝑪 Instrumentation and controls 299502

𝑪𝑻𝑭 Tanks and frames 247004

𝑪𝑴𝑰 Miscellaneous equipment 850963

𝑪𝒑,𝒇 Feed pump 822

𝑪𝒑,𝒓 Recirculation pump 2000

𝑪𝑪𝑨𝑷 Total capital cost 1803498

Annual CAPEX 233561

65

Table A-11. Operational parameters and costs for case study IV.

Name Description Cost (€)

𝑪𝑨𝒏𝒏𝒖𝒂𝒍 𝒓𝒆𝒑 Membrane replacement cost 144595

𝑾𝒇𝒆𝒆𝒅 (𝒌𝑾. 𝒉) Feed pump power 36

𝑾𝒓𝒆𝒄𝒊𝒓 (𝒌𝑾. 𝒉) Recirculation pump power 108

𝑬 𝑪𝒑,𝒇 Energy cost of feed pump work 6066

𝑬 𝑪𝒑,𝒓 Energy cost of feed pump work 18194

𝑪𝒄𝒍𝒆𝒂𝒏𝒊𝒏𝒈 Membrane cleaning cost 68265

𝑪𝒍𝒂𝒃𝒐𝒓 Labor cost

4671

𝑪𝒎𝒂𝒊𝒏 Maintenance cost

Annual OPEX (€) 241792

Case study MF V

Table A-12. Capital costs for case study V.

Name Description Cost (€)

𝑪𝒎𝒆𝒎 Set of membranes 225678

𝑪𝑷𝑽 Pipes and valves 248507

𝑪𝑰𝑪 Instrumentation and controls 370296

𝑪𝑻𝑭 Tanks and frames 292889

𝑪𝑴𝑰 Miscellaneous equipment 1022102

𝑪𝒑,𝒇 Feed pump 822

𝑪𝒑,𝒓 Recirculation pump 2629

𝑪𝑪𝑨𝑷 Total capital cost 2162923

Annual CAPEX 280108

66

Table A-13. Operational parameters and costs for case study V.

Name Description Cost (€)

𝑪𝑨𝒏𝒏𝒖𝒂𝒍 𝒓𝒆𝒑 Membrane replacement cost 175358

𝑾𝒇𝒆𝒆𝒅 (𝒌𝑾. 𝒉) Feed pump power 36

𝑾𝒓𝒆𝒄𝒊𝒓 (𝒌𝑾. 𝒉) Recirculation pump power 142

𝑬 𝑪𝒑,𝒇 Energy cost of feed pump work 6066

𝑬 𝑪𝒑,𝒓 Energy cost of feed pump work 23856

𝑪𝒄𝒍𝒆𝒂𝒏𝒊𝒏𝒈 Membrane cleaning cost 94150

𝑪𝒍𝒂𝒃𝒐𝒓 Labor cost

5602

𝑪𝒎𝒂𝒊𝒏 Maintenance cost

Annual OPEX (€) 305032

Case study MF VI

Table A-14. Capital costs for case study VI.

Name Description Cost (€)

𝑪𝒎𝒆𝒎 Set of membranes 297166

𝑪𝑷𝑽 Pipes and valves 301297

𝑪𝑰𝑪 Instrumentation and controls 501197

𝑪𝑻𝑭 Tanks and frames 373482

𝑪𝑴𝑰 Miscellaneous equipment 1327479

𝑪𝒑,𝒇 Feed pump 822

𝑪𝒑,𝒓 Recirculation pump 4035

𝑪𝑪𝑨𝑷 Total capital cost 2805478

Annual CAPEX 363322

67

Table A-15. Operational parameters and costs for case study VI.

Name Description Cost (€)

𝑪𝑨𝒏𝒏𝒖𝒂𝒍 𝒓𝒆𝒑 Membrane replacement cost 230906

𝑾𝒇𝒆𝒆𝒅 (𝒌𝑾. 𝒉) Feed pump power 36

𝑾𝒓𝒆𝒄𝒊𝒓 (𝒌𝑾. 𝒉) Recirculation pump power 225

𝑬 𝑪𝒑,𝒇 Energy cost of feed pump work 6066

𝑬 𝑪𝒑,𝒓 Energy cost of feed pump work 37800

𝑪𝒄𝒍𝒆𝒂𝒏𝒊𝒏𝒈 Membrane cleaning cost 148936

𝑪𝒍𝒂𝒃𝒐𝒓 Labor cost

7266

𝑪𝒎𝒂𝒊𝒏 Maintenance cost

Annual OPEX (€) 430975

Case study MF VII

Table A-16. Capital costs for case study VII.

Name Description Cost (€)

𝑪𝒎𝒆𝒎 Set of membranes 371969

𝑪𝑷𝑽 Pipes and valves 352574

𝑪𝑰𝑪 Instrumentation and controls 641605

𝑪𝑻𝑭 Tanks and frames 455409

𝑪𝑴𝑰 Miscellaneous equipment 1643086

𝑪𝒑,𝒇 Feed pump 822

𝑪𝒑,𝒓 Recirculation pump 5585

𝑪𝑪𝑨𝑷 Total capital cost 3471049

Annual CAPEX 449517

68

Table A-17. Operational parameters and costs for case study VII.

Name Description Cost (€)

𝑪𝑨𝒏𝒏𝒖𝒂𝒍 𝒓𝒆𝒑 Membrane replacement cost 289030

𝑾𝒇𝒆𝒆𝒅 (𝒌𝑾. 𝒉) Feed pump power 36

𝑾𝒓𝒆𝒄𝒊𝒓 (𝒌𝑾. 𝒉) Recirculation pump power 333

𝑬 𝑪𝒑,𝒇 Energy cost of feed pump work 6066

𝑬 𝑪𝒑,𝒓 Energy cost of feed pump work 55944

𝑪𝒄𝒍𝒆𝒂𝒏𝒊𝒏𝒈 Membrane cleaning cost 216528

𝑪𝒍𝒂𝒃𝒐𝒓 Labor cost

8990

𝑪𝒎𝒂𝒊𝒏 Maintenance cost

Annual OPEX (€) 576558

69

Case study MF→EVAP I - α

Table A-18. Capital costs for case study I – α.

Capital cost parameters (% of 𝑷𝑬𝑪) Cost (€)

Dir

ect

Co

sts

Equipment 100 2468880

Instrumentation and control 10 246888

Piping 10 246888

Electrical services 5 123444

Building and building services 15 370332

Service facilities 20 493776

Land and yard improvements 5 123444

Spare parts 4 98755

Ind

irect

Co

sts

Engineering and supervision 12 296266

Construction expenses 10 246888

Contractor’s fee 0,5 123444

Contingency 8 197510

Working investment 12% of 𝐶𝐶𝐴𝑃 604382

Annual CAPEX 730522

Table A-19. Operational parameters and costs for case study I - α.

Name Description Value (€)

𝑪𝒔𝒕𝒆𝒂𝒎 Steam average cost 956680

𝑪𝒍𝒂𝒃𝒐𝒓 Labor cost 56409

𝑪𝒎𝒂𝒊𝒏 Maintenance cost 112818

Annual OPEX (€) 1125907

70

Case study MF→EVAP II - α

Table A-20. Capital costs for case study II - α

Capital cost parameters (% of 𝑷𝑬𝑪) Cost (€)

Dir

ect

Co

sts

Equipment 100 2468880

Instrumentation and control 10 246888

Piping 10 246888

Electrical services 5 123444

Building and building services 15 370332

Service facilities 20 493776

Land and yard improvements 5 123444

Spare parts 4 98755

Ind

irect

Co

sts

Engineering and supervision 12 296266

Construction expenses 10 246888

Contractor’s fee 0,5 123444

Contingency 8 197510

Working investment 12% of 𝐶𝐶𝐴𝑃 604382

Annual CAPEX 730522

Table A-21. Operational parameters and costs for case study II - α.

Name Description Value (€)

𝑪𝒔𝒕𝒆𝒂𝒎 Steam average cost 830904

𝑪𝒍𝒂𝒃𝒐𝒓 Labor cost 56409

𝑪𝒎𝒂𝒊𝒏 Maintenance cost 112818

Annual OPEX (€) 1000131

71

Case study MF→EVAP III - α

Table A-22. Capital costs for case study III - α

Capital cost parameters (% of 𝑷𝑬𝑪) Cost (€)

Dir

ect

Co

sts

Equipment 100 1925727

Instrumentation and control 10 192573

Piping 10 192573

Electrical services 5 96286

Building and building services 15 288859

Service facilities 20 385145

Land and yard improvements 5 96286

Spare parts 4 77029

Ind

irect

Co

sts

Engineering and supervision 12 231087

Construction expenses 10 192573

Contractor’s fee 0,5 96286

Contingency 8 154058

Working investment 12% of 𝐶𝐶𝐴𝑃 471418

Annual CAPEX 569807

Table A-23. Operational parameters and costs for case study III - α.

Name Description Value (€)

𝑪𝒔𝒕𝒆𝒂𝒎 Steam average cost 714491

𝑪𝒍𝒂𝒃𝒐𝒓 Labor cost 43999

𝑪𝒎𝒂𝒊𝒏 Maintenance cost 87998

Annual OPEX (€) 846488

72

Case study MF→EVAP IV - α

Table A-24. Capital costs for case study IV - α

Capital cost parameters (% of 𝑷𝑬𝑪) Cost (€)

Dir

ect

Co

sts

Equipment 100 1925727

Instrumentation and control 10 192573

Piping 10 192573

Electrical services 5 96286

Building and building services 15 288859

Service facilities 20 385145

Land and yard improvements 5 96286

Spare parts 4 77029

Ind

irect

Co

sts

Engineering and supervision 12 231087

Construction expenses 10 192573

Contractor’s fee 0,5 96286

Contingency 8 154058

Working investment 12% of 𝐶𝐶𝐴𝑃 471418

Annual CAPEX 569807

Table A-25. Operational parameters and costs for case study IV - α.

Name Description Value (€)

𝑪𝒔𝒕𝒆𝒂𝒎 Steam average cost 358670

𝑪𝒍𝒂𝒃𝒐𝒓 Labor cost 43999

𝑪𝒎𝒂𝒊𝒏 Maintenance cost 87998

Annual OPEX (€) 490667

73

Case study MF→EVAP V - α

Table A-26. Capital costs for case study V - α

Capital cost parameters (% of 𝑷𝑬𝑪) Cost (€)

Dir

ect

Co

sts

Equipment 100 1925727

Instrumentation and control 10 192573

Piping 10 192573

Electrical services 5 96286

Building and building services 15 288859

Service facilities 20 385145

Land and yard improvements 5 96286

Spare parts 4 77029

Ind

irect

Co

sts

Engineering and supervision 12 231087

Construction expenses 10 192573

Contractor’s fee 0,5 96286

Contingency 8 154058

Working investment 12% of 𝐶𝐶𝐴𝑃 471418

Annual CAPEX 569807

Table A-26. Operational parameters and costs for case study V - α.

Name Description Value (€)

𝑪𝒔𝒕𝒆𝒂𝒎 Steam average cost 476915

𝑪𝒍𝒂𝒃𝒐𝒓 Labor cost 43999

𝑪𝒎𝒂𝒊𝒏 Maintenance cost 87998

Annual OPEX (€) 608912

74

Case study MF→EVAP VI - α

Table A-27. Capital costs for case study VI - α

Capital cost parameters (% of 𝑷𝑬𝑪) Cost (€)

Dir

ect

Co

sts

Equipment 100 1388745

Instrumentation and control 10 138875

Piping 10 138875

Electrical services 5 69437

Building and building services 15 208312

Service facilities 20 277749

Land and yard improvements 5 69437

Spare parts 4 55550

Ind

irect

Co

sts

Engineering and supervision 12 166649

Construction expenses 10 138875

Contractor’s fee 0,5 69437

Contingency 8 111100

Working investment 12% of 𝐶𝐶𝐴𝑃 339965

Annual CAPEX 410919

Table A-28. Operational parameters and costs for case study VI - α.

Name Description Value (€)

𝑪𝒔𝒕𝒆𝒂𝒎 Steam average cost 358670

𝑪𝒍𝒂𝒃𝒐𝒓 Labor cost 31730

𝑪𝒎𝒂𝒊𝒏 Maintenance cost 63460

Annual OPEX (€) 453860

75

Case study MF→EVAP VII - α

Table A-29. Capital costs for case study VII - α

Capital cost parameters (% of 𝑷𝑬𝑪) Cost (€)

Dir

ect

Co

sts

Equipment 100 1277645

Instrumentation and control 10 127765

Piping 10 127765

Electrical services 5 63882

Building and building services 15 191647

Service facilities 20 255529

Land and yard improvements 5 63882

Spare parts 4 51106

Ind

irect

Co

sts

Engineering and supervision 12 153317

Construction expenses 10 127765

Contractor’s fee 0,5 63882

Contingency 8 102212

Working investment 12% of 𝐶𝐶𝐴𝑃 312768

Annual CAPEX 378045

Table A-30. Operational parameters and costs for case study VII - α.

Name Description Value (€)

𝑪𝒔𝒕𝒆𝒂𝒎 Steam average cost 262269

𝑪𝒍𝒂𝒃𝒐𝒓 Labor cost 29192

𝑪𝒎𝒂𝒊𝒏 Maintenance cost 58383

Annual OPEX (€) 349844

76

Case study EVAP 20%

Table A-31. Capital costs for case study EVAP 20%.

Capital cost parameters (% of 𝑷𝑬𝑪) Cost (€)

Dir

ect

Co

sts

Equipment 100 73992338

Instrumentation and control 10 7399234

Piping 10 7399234

Electrical services 5 3699617

Building and building services 15 11098851

Service facilities 20 14798468

Land and yard improvements 5 3699617

Spare parts 4 2959694

Ind

irect

Co

sts

Engineering and supervision 12 8879081

Construction expenses 10 7399234

Contractor’s fee 0,5 3699617

Contingency 8 5919387

Working investment 12% of 𝐶𝐶𝐴𝑃 18113324

Annual CAPEX 21893745

Table A-32. Operational parameters and costs for case study EVAP 20%.

Name Description Value (€)

𝑪𝒔𝒕𝒆𝒂𝒎 Steam average cost 1067802

𝑪𝒍𝒂𝒃𝒐𝒓 Labor cost 1690577

𝑪𝒎𝒂𝒊𝒏 Maintenance cost 3381154

Annual OPEX (€) 6139532

77

The sensitivity analysis results are listed in Table A-33 for the general case study and in Table A-34, for

the industry case study.

Table A-33. Process design parameter influence (%) in TAC for the general case study, VI-α..

General Case Study - Influence percentage in TAC (%)

Parameter variation ±20% ±10% ±5%

Equipment (MF) ±0,41 ±0,21 ±0,10

Instrumentation and Control (MF) ±0,69 ±0,35 ±0,17

Piping (MF) ±0,42 ±0,21 ±0,10

Miscellaneous Equipment (MF) ±1,84 ±0,92 ±0,46

Equipment (EVAP) ±1,92 ±0,96 ±0,48

Instrumentation and Control (EVAP)

±0,19 ±0,10 ±0,05

Piping (EVAP) ±0,19 ±0,10 ±0,05

Miscellaneous Equipment (EVAP) ±0,77 ±0,38 ±0,19

Spare parts (EVAP) ±1,24 ±0,62 ±0,02

Spare parts (MF) ±2,47 ±1,24 ±0,62

Cleaning (MF) ±1,59 ±0,80 ±0,40

Labor and Maintenance (MF) ±0,08 ±0,04 ±0,02

Electricity (MF) ±0,47 ±0,23 ±0,12

Labor and Maintenance (EVAP) ±1,02 ±0,51 ±0,25

Steam Price (EVAP) ±6,09 ±3,04 ±1,52

Table A-34. Process design parameter influence (%) in TAC for the PPI case study, I-α.

Caima, Indústria da Celulose - Influence percentage in TAC (%)

Parameter variation ±20% ±10% ±5%

Equipment (MF) ±1,77 ±1,73 ±1,70

Instrumentation and Control (MF) ±1,81 ±1,75 ±1,71

Piping (MF) ±1,84 ±1,76 ±1,72

Miscellaneous Equipment (MF) ±2,13 ±1,91 ±1,79

Spare parts (MF) ±1,15 ±0,57 ±0,29

Cleaning (MF) ±0,11 ±0,05 ±0,03

Labor and Maintenance (MF) ±0,02 ±0,01 ±0,005

Electricity (MF) ±0,13 ±0,07 ±0,03

Labor and Maintenance (EVAP) ±1,97 ±0,99 ±0,49

Steam Price (EVAP) ±7,13 ±3,56 ±1,78

A nanofiltration pre-treatment technology can be a (slow) sand filter, which capital cost is calculated as

expressed in equation A.1, according to (Rural & Cowi, 2006).

𝐶𝑎𝑝𝑖𝑡𝑎𝑙 𝑐𝑜𝑠𝑡 (€) = 9900 . 𝑄𝑓𝑒𝑒𝑑−0,634 (𝑚3 𝑑𝑎𝑦⁄ ) A -1