Final Report

on

Ethanol Dehydration to Green Ethylene

Catalysts, processes, difficulties and a plant design with an economic

evaluation

Presented to

COWI and Borealis

Principal Investigators

Josefina Jernberg, Øyvind Nørregård, Marianne Olofsson,

Oliver Persson, Maria Thulin

Tutors

Christian Hulteberg, Hans Karlsson

27th

May 2015

I

Abstract

Ethylene is an important chemical in the petrochemical industry. The main manufacturing

method is steam cracking of naphtha and shorter hydrocarbons. Naphtha is a product from the

petroleum industry which may have a negative impact on the environment. Therefore

processes to produce ethylene in a more environmentally friendly way are developed. In one

of these new processes bioethanol is dehydrated to ethylene over a catalyst. Bioethanol can be

of either 1st or 2

nd generation. Plants where 1

st generation bioethanol is used already exists but

few or no plant using 2nd

generation bioethanol is built in commercial scale. The potential

problems with 2nd

generation bioethanol is that it is not yet produced in large quantities and

that it contains impurities that possibly could cause problems in the process. Despite this, a

process where 2nd

generation bioethanol could be used would be preferred since the 2nd

generation bioethanol production does not compete with food industry. Borealis in

Stenungsund wants to complement their naphtha process with a process where 2nd

generation

bioethanol can be dehydrated to ethylene.

This study is divided into two parts, a literature review and a feasibility study. The aim of the

literature review was to give Borealis a background comprising catalysts, available processes

and contaminants that possibly can have an impact on the process. In the second part, the

feasibility study, a design of a green ethylene plant is presented together with an economical

evaluation.

Contaminants found in 2nd

generation bioethanol can possibly have an impact on catalysts and

further investigations are needed in this area. The types of catalysts that seemed to be the most

promising ones are alumina oxides, zeolites and heteropolyacids. The main problem using any

of these catalysts is coke formation. Modifications can be made in order to decrease coke

formation. There are companies claiming that they can sell processes which dehydrate ethanol

to ethylene. Some of these companies are British Petroleum, Chematur and Axens together

with Total and IFPEN.

The plant designed in the feasibility study consisted of one reactor with associated

downstream purification and recycle units. The reactor was designed to be isothermal and

contained a tube bundle filled with heteropolyacid catalyst. The process was simulated

successfully and 102,000 tonnes ethylene with a purity of 99.9 % was produced annually. The

yield of ethylene over the reactor, with recirculation of ethanol, was 96.8 %. The investment

cost of the plant was calculated to 28,105,200 USD and a total production cost of 1,136

USD/tonne ethylene was obtained. The ethanol price and the ethylene price was assumed to

be 537 USD/tonne and 1,280 USD/tonne respectively. This resulted in a pay-back time of

1.75 years in the economical evaluation.

II

Contents

1 Introduction ........................................................................................................................ 1

1.1 Purpose ....................................................................................................................... 1

1.2 Disposition ................................................................................................................. 1

2 Literature review ................................................................................................................ 2

2.1 Ethylene production from bioethanol ......................................................................... 2

2.2 Bioethanol based on lignocellulosic biomass ............................................................. 2

2.3 Potential contaminants formed in ethylene process ................................................... 5

2.4 Catalysts ..................................................................................................................... 5

2.5 Commercial processes available .............................................................................. 11

2.6 Ethylene purification ................................................................................................ 12

3 Discussion regarding literature study ............................................................................... 13

4 Feasibility study ............................................................................................................... 14

4.1 Process steps ............................................................................................................. 16

4.2 Design ....................................................................................................................... 17

4.3 Cost estimates ........................................................................................................... 20

5 Discussion ........................................................................................................................ 28

6 Conclusions ...................................................................................................................... 30

7 Appendix .......................................................................................................................... I-i

Appendix I – Ethanol calculations ....................................................................................... I-i

Appendix II – Heat and mass balances ............................................................................... II-i

Appendix III – Reaction specifications .............................................................................. III-i

Appendix IV – Design equations and assumptions ............................................................IV-i

Appendix V – Calculation of unit operation design ............................................................ V-i

Appendix VI – Economy equations and method ...............................................................VI-i

Appendix VII – Grass Root Capital ................................................................................. VII-i

Appendix VIII – Operating costs .................................................................................... VIII-i

Appendix IX – Investment calculations .............................................................................IX-i

1

1 Introduction

One of the most important polymer precursors in today’s society is ethylene. The traditional

production method, steam cracking, is a thermal cracking process with no catalysts involved.

A common feedstock is naphtha but lighter hydrocarbons such as ethane can be used as well.

The produced ethylene is an important monomer for a variety of plastics [1]. An ethylene

producing company is Borealis, which currently utilizes the steam cracking technique. Since

fossil fuels eventually will run low and with a desire to contribute to a more sustainable

society, an increasing interest in converting 2nd

generation bioethanol, derived from

lignocellulosic material, to ethylene has occurred.

1.1 Purpose

The aim of this study was to investigate the possibility of using second generation ethanol as a

feedstock for production of ethylene and to design a plant where this is possible. The report

identifies catalysts that can be used as well as potential contaminants and their effect on the

catalysts. A review of how a plant could be designed and economic aspects of the plant is also

presented.

1.2 Disposition

The study is divided into two parts, a literature review and a feasibility study. The literature

review shows an overall process design and which reactor types that can be used. A more

detailed description of bioethanol from lignocellulosic biomass and its contaminants are

presented as well. Possible catalysts are considered and some information about commercial

available processes is given. The first part is followed by a discussion regarding the literature

review. In the feasibility study a design of a process is presented together with an economic

evaluation. The report is finalized with a discussion and a conclusion. At the end, the

references are listed.

2

2 Literature review

2.1 Ethylene production from bioethanol

To produce ethylene from ethanol, acidic catalysts are often used. Homogeneous catalysts that

traditionally have been used are sulfuric acid or phosphoric acid. More modern catalysts are

acidic heterogeneous catalysts such as zeolites treated with e.g. alumina or manganese. These

catalysts provide a high selectivity to ethylene, which results in a relatively simple process. A

general process flow diagram is depicted in Figure 1 below, where the main parts needed for

ethylene production are presented. Ethanol with a certain composition is vaporized and

delivered to the reactor. The specification of the ethanol depends on the robustness of the

catalyst. After dehydration reaction in the reactor the gaseous product is separated from the

non-reacted ethanol. By-products are removed and transported to wastewater treatment. The

crude ethylene is separated from heavier products in a distillation tower and is then further

purified in a separate section [1].

Figure 1. A general process flow diagram of ethanol dehydration [1].

Heat supplement to the reactor can be either isothermally or adiabatically dependent of what

reactor conditions that are desired. An adiabatic reactor operates without heat exchange with

the environment. To provide this reactor with heat when producing ethylene from bioethanol,

steam is used as a heat carrier leading to a lowered temperature gradient. This results in a

reduced amount of catalyst needed and will decrease the by-product formation. When having

a catalyst that requires higher temperature (400-500 °C), the adiabatic reactor can be

advantageous since the isothermal reactor gives higher investment costs at these temperatures.

This is because the isothermal reactor operates with a circulating heating fluid and requires a

tube bundle to be able to operate at higher temperatures. The tube bundle consists of multiple

smaller tubes, which improves the heat transfer. The bundle gives a larger area, but results in

a higher amount of material, more welding and speciality materials, leading to higher

investment costs. When having a catalyst that can perform at lower temperatures (~300 °C)

an isothermal reactor is a good candidate. Because of the fact that the temperature in the

reactor is constant, it is easier to control the reaction. The reactor can be fixed bed or if

isotherm heat supplement is utilized, fluidized bed [1, 2].

2.2 Bioethanol based on lignocellulosic biomass

The bioethanol production is based on fermentation of sugars by using microorganisms like

yeast. In 1st generation bioethanol the ethanol is produced from sources where the sugars are

3

more easily obtainable. It can be from sugarcanes and sugar beets, where the sugars just need

to be extracted. Or from starchy sources like corn and grains, where hydrolysis of the starch is

used to convert it into sugars. The 2nd

generation bioethanol uses lignocellulosic biomass like

woody crops and agricultural residues which are much more complex to acquire sugars from.

The three main components in lignocellulose is cellulose, hemicellulose and lignin. Cellulose

may be hydrolyzed to glucose, but is more resistant to hydrolysis than starch. Another

difficulty is that lignin forms a protective barrier for the cellulose microfibrils [3]. The

hemicellulose may be depolymerized to form pentose and hexose sugars, but linkage between

hemicellulose and lignin makes it more difficult to degrade [4].

2.2.1 Production of 2nd generation bioethanol and its potential contaminants Production of ethanol from lignocellulosic biomass can be summarized in four main steps [5],

Figure 2. Pretreatment is first performed to break down the lignocellulosic structure. The

second step is hydrolysis which cleaves the cellulose and hemicellulose molecules into

monosaccharides. These sugars are then fermented to ethanol in the third step. In the final step

the fermentation broth is distilled in order to separate and purify the ethanol. In addition to the

core bioethanol process wastewater treatment is of high importance [6] but falls out of the

scope for this literature study.

Figure 2. A process flowchart describing the conversion of lignocellulose to ethanol.

2.2.1.1 Pretreatment

The pretreatment methods can be divided into physical, chemical, physicochemical and

biological methods [7]. Physical methods, such as milling, are generally too energy intense

which makes them unfeasible [8]. One method suitable for commercial use is dilute acid

pretreatment, which can be used on woody biomass like spruce and willow. This chemical

method aims to separate hemicellulose and lignin as well as open up the lignocellulosic

structure [9]. The major disadvantage with the dilute acid pretreatment, and other acid

pretreatment methods, is that a part of the sugars that could have been used for producing

ethanol is lost in formation of lignocellulose degradation by-products [10]. These by-products

can be divided into three main groups: furans, weak acids and phenolic compounds [9]. The

furans that exist after pretreatment are 5-hydroxymethyl-2-furaldehyde (HMF) and 2-

furaldehyde (furfural). These are formed by dehydration of hexoses respectively pentoses.

The concentration of furans is strongly dependent on how the dilute acid pretreatment is

performed, for example weather it is operated as a one-step or a two-step process. Three weak

acids that are generated during the lignocellulose treatment are acetic acid, levulinic acid and

formic acid. Acetic acid is formed by deacetylation of hemicellulose. Both formic and

levulinic acid are products from breakdown of HMF, but formic acid may also be created

from reaction of furfural in the acidic environment. The phenolic compounds that are formed

consist of a wide range of compounds. Lignin interacts in different ways in different

lignocellulosic materials, consequently the biomass source have big influence on the amount

Lignocellulosic

biomass

Pre-

treatment Hydrolysis Fermentation Distillation

Ethanol

4

and type of phenolic compounds which will form. Two of the more common phenolic

compounds that are formed from dilute acid pretreatment of spruce are vanillin and

syringaldehyde. The same groups of by-products as mentioned above have also been reported

for wet oxidation of wheat straw and steam pretreatment of sugar cane bagasse and corn

stover [9]. However, the amounts of by-products vary and not all of the above mentioned

compounds exist in these types of pretreatment.

A promising alkaline pretreatment method is AFEX (ammonia fiber explosion) in which the

biomass is treated with ammonia under pressure [6]. AFEX works better for biomass with low

lignin content and is therefore particularly suitable for agricultural residues rather than

softwoods. Unlike during acidic pretreatment no fermentation inhibiting by-products are

created during AFEX [7].

2.2.1.1.1 Inorganic contents

Perennial crops generally have lower inorganic content than annual crops [11]. Most woods in

Europe contain 0.1 to 1.0 % mineral components [12]. The most common ions of mineral

components in wood are shown in Table 1. These are mainly present in compounds of

carbonates or glucuronates. Also oxalate, phosphate and silicate anions can be found. Usually

sulphur and chlorine amounts for 0.1 % respectively of the biomass weight [13]. It is

important to point out that mineral contents vary between different species and locations and

that the mineral compounds and inorganics will exist in the hydrolysate solution. An

additional source of salts comes from after the pretreatment when the pH must be adjusted

[14].

Table 1. Ions present in mineral components of wood [12].

Ion Share in mineral components

Calcium 40 - 70 %

Potassium 10 - 30 %

Magnesium 5 - 10 %

Iron up to 10 %

Sodium low quantities

Manganese lower quantities

Aluminum lower quantities

2.2.1.2 Hydrolysis

After the lignocellulosic structure has been opened up in the pretreatment, hydrolysis can be

performed. The hydrolysis aims at converting the polysaccharides into fermentable sugars and

can be performed both with enzymes or acid [5]. The polysaccharides are mainly cellulose,

but depending on pretreatment method also parts of the hemicellulose remains. The present

commercialization of enzymatic hydrolysis has preceded by a possibility to decrease enzyme

loadings and lowered enzyme production costs [6]. Novel enzymes have also shown to be

very efficient in the conversion of cellulose [15]. The enzymatic hydrolysis uses a mixture of

enzymes, mainly cellulases, to release the remainder of the sugars [6].

5

2.2.1.3 Fermentation

The sugars that are to be fermented are both hexoses (glucose, galactose, mannose, rhamnose)

and pentoses (xylose, arabinose) [9]. There are many microorganisms with potential of being

used in the fermentation process, and normal baker’s yeast (Saccharomyces cerevisiae) is one

of them [6]. It is well established in large scale production and has a high ethanol tolerance.

Another advantage is that S. cerevisiae have shown relatively good resistance against the by-

products that are formed in the pretreatment, which inhibit fermentation. A large disadvantage

with S. cerevisiae is that it does not naturally ferment pentoses. Since xylose (a pentose) is the

second most occurring sugar in many plants it is highly desirable to ferment it. Different

methods have been proposed for solving this problem. One is to separate the hemicellulose

from the cellulose and fermenting it with other microorganisms while another is to modify S.

cerevisiae [5, 14]. During the metabolism of S. cerevisiae a variety of by-products is formed.

These are acetaldehyde, acetic acid/acetate, glycerol and carbon dioxide [14]. Also methanol

[3], fusel oils (n-propanol, amyl alcohol, isoamyl alcohol, isobutanol, phenethyl alcohol etc.)

and ethyl acetate are created during ethanol fermentation [16]. Volatile sulphur compounds

that can be created are diethyl sulfide and dimethyl sulfide [16]. It is further known that

recombinant S. cerevisiae slowly may reduce the furaldehydes created in the dilute acid

pretreatment to their corresponding alcohols (2-furan methanol, furan 2,5-dimethanol) [14].

2.2.1.4 Purification

The finalization of the ethanol production is to separate and purify the ethanol from the

fermentation broth. The main technique used in this process step is distillation. At 95.6 wt%

ethanol the distillation reaches the azeotropic mixture of ethanol and water. When the

azeotrope is boiled the vapor has the same composition as the unboiled mixture, resulting in

no separation in a distillation column. To purify the ethanol beyond the azeotrope the ethanol

must be dehydrated. However, for economic reasons dehydration above azeotrope is not

always implemented [3]. Most of the by-products formed and minerals solved during

bioethanol production are removed in the distillation, leaving only traces behind.

2.3 Potential contaminants formed in ethylene process

During the production of ethylene by catalytic dehydration of bioethanol, a range of by-

products is formed. The bioethanol itself may also already contain impurities, e.g. methanol

and fusel oil, as mentioned above.

Two ethanol molecules may react to form diethyl ether [17]. However, the diethyl ether is to

be considered more as an intermediate rather than a by-product [18]. Its formation is favored

mainly between 150 °C and 300 °C. The most important by-products are acetaldehyde and

hydrogen which are formed by decomposition of ethanol [17]. Other by-products that may be

formed in small quantities are alkanes and alkenes with low number of carbons (methane,

ethane, propane, butane, propylene, butylenes) and carbon monoxide [17, 18]. Also

hydrocarbons with five carbons or more may exist. Of all the hydrocarbons butylenes and

ethane are the most occurring in the ethylene product [19]. At too high temperatures the

amount of by-products like coke are increasing [17].

2.4 Catalysts

A catalyst is used to change the reaction kinetics of a reaction. It cannot change the

equilibrium of a reaction but it can change the rate of reaction toward the equilibrium. It

lowers the activation energy, which is needed for the reaction to occur. The catalyst is not

6

consumed during the reaction but it can be deactivated and loose its ability to catalyze the

wanted reaction. There are some concepts that are important when talking about catalysts. The

main three concepts in focus within this report are activity, selectivity and deactivation. The

activity is a measurement of how fast the reaction reaches the equilibrium. Selectivity

describes the capability to produce a desired product. Deactivation is when a catalyst loses its

ability to catalyze a reaction and becomes less active [20, 21]. Deactivation of the catalysts

will be discussed later in the report.

Heterogeneous catalysts can be used for the dehydration of ethanol to ethylene. A

heterogeneous catalyst consist of three parts, these are carrier, support and active site. The

carrier provides structure to the catalyst and the reactor bed; it determines the heat and mass

transfer properties and also governs the pressure drop over the reactor. The support provides

surface area, on which the reaction can occur, and can be of the same material as the carrier.

The surface area of the support is important since this is where the active sites are placed. The

active site, or active phase, is where the reaction actually occurs and can be the material of the

support. Some common materials are alumina, silica and mixed oxides [20].

In the catalytic dehydration of ethanol to ethylene an acid catalyst, with weak, relatively

strong or strong acid sites, can promote the reaction [22]. Research has been done on different

types of catalysts and mainly three different groups of acid catalysts have been the focus in

this report. These are oxide catalysts (γ-Al2O3 based), molecular sieve catalysts (HZSM-5

zeolites) and heteropolyacids [22]. All of these can be modified and doped which results in

varying lifetime, conversion of ethanol and selectivity towards ethylene. See Table 2 for a

comparison of the different catalysts that are studied in this review.

2.4.1 Alumina oxides γ-Al2O3 catalysts are one of the earlier catalysts used for the production of ethylene from

ethanol [23]. γ-Al2O3 is a crystalline form of alumina oxide which have a porous structure

with a surface area of approximately 180 m2/g [20]. The yield of ethylene is quite low,

approximately 80 %, and requires relatively high temperatures, 450 °C [24]. The active

alumina-based catalyst is a straight forward and commercialized catalyst with good stability.

An Al2O3 catalyst doped with 10 % TiO2 has in microchannel reactor experiments shown a

relatively good stability during 400 h at temperatures between 410 and 430 °C [25].

To increase the ethanol conversion γ-Al2O3 can be modified in different ways, for example by

adding oxides like MgO/SiO2, Cr2O3, FeOx and TiO2 [22]. There are numerous of other ways

to improve the oxide catalysts and thus increase the ethanol conversion and selectivity. One of

the problems with this type of catalyst is the high reaction temperature required [24]. Another

problem is that water can deactivate active sites on γ-Al2O3 and inhibit the formation rate of

ethylene and diethyl ether [26].

2.4.1.1 Different phases

Al2O3 catalysts also exist in other phases than γ-Al2O3. In a study made by Phung et al. [27]

the catalytic activity of the following five transition alumina catalyst was investigated; P90

(90 ± 5 m2/g), P200 (190 ± 10 m

2/g), V200 (202 ± 5 m

2/g), D100 (100 ± 10 m

2/g) and SA330

(330 ± 10 m2/g). Manufacturers of the different catalysts are presented in Table 2. The X-ray

diffraction evidence the phase θ-Al2O3 for P90, γ, δ-Al2O3 for D100 and γ-Al2O3 for P200

and V200. SA330 has an amorphous structure. In the experiments 0.5 g of catalyst was used

and the temperature was varied between 150 °C and 450 °C at atmospheric pressure. The

reactor used was a tubular flow reactor. The feed consisted of 7.9 % v/v ethanol in nitrogen.

7

The result from the experiment showed that P200 was the most active at low temperatures.

The selectivity to diethyl ether, ethylene, ethane and butene was measured. For all of the

catalysts the conversion of ethanol increased with temperature. The selectivity to diethyl ether

was high at temperatures below 250 °C. At temperatures above 250 °C the selectivity to

ethylene increased and reached approximately 100 % for all the catalysts. At low temperature

and thus low conversions SA330 had the highest selectivity to ethylene [27].

2.4.2 Zeolites Zeolites are crystalline aluminosilicates with a fine pore structure. The surface area of a

zeolite is approximately 900 m2/g. The pores can contain exchangeable cations, which can be

exchanged to alter the performance of the zeolite. By changing the size of the cations, the

cross-section of the pores is changed and thereby allowing different types of molecules to be

adsorbed and pass through. The catalytic behavior of the zeolite can also be changed by

substituting the cations [28].

ZSM-5 zeolites have a high Si/Al ratio and have been used to catalyze the dehydration of

ethanol to ethylene [22, 28]. As well as the oxide catalysts, zeolites can be doped to enhance

the activity or to make it more stable. The stability is mainly affected by the acidity. High

acidity increases coke formation, which deactivates the catalyst. With this said one of the

difficulties is to find a way to reduce the acidity but still be able to keep a good conversion

and selectivity of ethylene [24]. Studies regarding these difficulties have been made on

different types of catalysts, which will be discussed below.

2.4.2.1 HZSM-5

HZSM-5 is a commercial zeolite that is in focus. It is able to catalyse the dehydration of

ethanol at lower temperatures, around 300 °C [24]. According to Zhang et al. [23] a

conversion of 98 % ethanol and ethylene selectivity of 95 % was achieved with HZSM-5 at

300 °C. The disadvantage with HZSM-5 is the high acidity, which promotes coking and

lowers the stability and lifetime of the catalyst. Desorption of NH3 is one method that can be

used to estimate the acidity of a catalyst. NH3 is adsorbed and corresponds to the amounts of

acid present in the catalyst; it is then desorbed by changing the temperature. This can be

made with a temperature-program which gradually increases the temperature. A higher

desorption temperature relates to stronger acid strength. HZSM-5 shows both strong and weak

sites and has a high desorption temperature. One conclusion from this is that the acid property

of HZSM-5 is high and also the activity, since the activity of the catalyst is dependent on the

acid strength and the amount of acid sites [23]. The stability on the other hand is dependent on

the deactivation of the catalyst. A strong acid site can polymerize the formed ethylene to

higher olefins and aromatics. Because of the fine microporous structure and the fact that these

substances can form structures that are not gaseous they will not be able to pass through.

Consequently coke can be formed and cover the active sites in the catalyst, which leads to

deactivation of the catalyst [29].

2.4.2.1.1 Lanthanum-phosphorous HZSM-5

Modifications of HZSM-5 catalyst have been made and lanthanum–phosphorous modified

HZSM-5 is one of them. To reduce the acidity of the HZSM-5 zeolite, phosphorous has been

used. This showed a lower amount of strong acidic sites and thereby a lower total acidic

strength. This will reduce the ability of coke formation and enhance the stability [30].

Lanthanum is added to minimize the required reaction temperature, which rises with the

addition of phosphorous. According to Zhan et al. the 0.5% La-2%P-HZSM-5 combination

8

had the best catalytic ability at relatively low temperatures and also better stability than the

original HZSM-5 [31].

2.4.2.1.2 Nano-scale HZSM-5

HZSM-5 can also be modified by making them nano-scale HZSM-5. Compared to micro-

scale HZSM-5 zeolites the nanoscale has a higher percentage of strong acid sites on the

surface of the catalyst, which enables for conversion without passing through the channels.

The diffusion path in the nanoscale HZSM-5 is also shorter than in the micro-scale HZSM-5

and thereby more resistant against coke formation which will be discussed further later on in

the report. The nanoscale catalyst showed a stable behaviour and the conversion of bioethanol

and selectivity of ethylene was almost constant during 630 h reaction. After 630 h of reaction

the conversion was 98.4 % and the selectivity of ethylene was 98.43 %, the reaction

temperature was 240 °C [32].

2.4.2.1.3 Alkali-treated HZSM-5

Alkali-treated HZSM-5 zeolites are zeolites with a changed pore structure. The structure of

traditional HZSM-5 is microporous but can be changed by treating the zeolite with NaOH.

Sheng et al. [29] performed experiments where treatment with NaOH resulted in more

mesopores and fewer strong acid sites. As mentioned above, the strong acid sites are more

likely to contribute to coke formation. Fewer strong acid sites will result in the catalyst not

being deactivated as fast. Mesopores allow higher flow through the catalyst and can also hold

the coke that has been formed. The reactant and the rest of the feed can still diffuse through

the catalyst to the active site. The modification that showed the best result was the HZSM-5

catalyst with 0.4 mol/L NaOH [29].

2.4.3 Heteropolyacids Heteropolyacids (HPAs) has become of great interest in the category of future catalyst. The

most known form is the Keggin type, with the general formula 𝑋𝑀12𝑂40, where X is the

heteroatom and M the addendum atom. 𝐻3𝑃𝑊12𝑂40 and 𝐻4𝑆𝑖𝑊12𝑂40 are two examples that

are commercially available, where 𝐻3𝑃𝑊12𝑂40 is the most acidic, regarded as a very strong

acid [33].

The main topic of interest in the research of new HPA catalysts is the economic benefit [34].

The economic benefits provided for the conversion of ethanol to ethylene is primarily the low

operating temperature, approximately 190 °C [24, 35]. Though the low operating

temperatures achieve large benefits, high temperatures are a problem because of the low

thermal stability of HPA catalysts.

2.4.3.1 Thermal stability and structure

Heteropolyacids’ acidity is beneficial for the ethanol conversion to ethylene. These catalysts

unfortunately have similar problems as zeolites regarding coking. Because of the low thermal

stability it is difficult to regenerate the catalyst from coke-deposits, which usually is

regenerated by burning the coke off with air [34]. A proposed way of decoking HPAs has

been tried using ozone in oxygen gas mixture at lower temperatures, approximately 125 °C,

instead of combustion of coke in air at high temperatures. This way was found promising in

lab scale and may be a way to regenerate the HPA catalyst without destroying the structure of

the catalyst [36]. Another way to enable regenerating of the HPA is by doping the material

with platinum group metals (PGM) to decrease the temperature of the combustion [34].

9

Another disadvantage with HPAs of the Keggin structure is the low surface area provided by

the crystal structure. This has been overcome by adding HPAs on supports with larger surface

area [37]. Though any better thermal stability of the catalyst is not always achieved, some

research has proven an increased thermal stability of the Keggin structure when added to

support materials [38]. Promising heteropolyacids added on support materials are mainly the

TPA-MCM-41 and STA-MCM-41, seen in Table 2, where MCM-41, which stands for

“Mobile Composition of Matter”, is the support. MCM-41 is a microporous silicate support

with high surface area, approximately 1,000 m2/g [39].

Some of the most promising heteropolyacids can be seen in Table 2; where especially

𝐴𝑔3𝑃𝑊12𝑂40 has proven high selectivity to ethylene production in the dehydration of ethanol.

One of the reasons for the exchange of hydrogen to silver as cation is to construct a water-

insoluble salt, which then is not affected by the water being produced in the process [40].

2.4.4 Coking Coking is a severe deactivation issue for most heterogeneous acid catalysts. Coking occur

mainly due to the oligomerisation and polymerization, linked to the dehydrogenation of

hydrocarbons, as seen in reaction 1.

𝐶𝑛𝐻𝑚 → 𝑝𝑜𝑙𝑦𝑚𝑒𝑟𝑎𝑠𝑎𝑡𝑖𝑜𝑛 → 𝑐𝑜𝑘𝑒 + 𝑥𝐻2 (reac. 1)

The higher ratio of carbon to hydrogen n/m in the hydrocarbon, and thus the more double

bonds, the more it tends to oligomerise. Thus the main issue of coking, in the conversion of

ethanol to ethylene, is the ethylene itself, which is very reactive [41].

To hinder the conversion of ethylene into coke the residence time, for the ethylene product,

has to be kept as low as possible [42]. This is one of the main problems considering zeolites,

as the micro-pores of the zeolite adds diffusion resistance and such the residence time

increases and coking evolves. As for most solid acid catalysts the strong acidity is a

contributing factor on coke formation as the reactivity increases with acidity and thus also the

chance of ethylene to react further. This can be suppressed by either decreasing the residence

time or by pretreating the catalyst to lower the reactivity [29]. Coke formation can also be

suppressed by adding water, or other substances, in the feed to compete with the precursors

[17].

10

Table 2. Comparison of the studied catalysts. Ethanol is either being inserted with nitrogen gas or air. The

reaction temperature is the lowest required to achieve the given selectivity and conversion.

Catalyst Max

ethylene

selectivity

Ethanol

conversion

Reaction

temperature

GHSVa/

WHSVb

Comments Ref

HZSM-5

96.4 % 100 % 260 °C 2 h-1 b

50 wt% ethanol

Lab scale

0.5 g catalyst

[31]

0.5% La 2%

P HZSM-5

99.9 % 100 % 240 °C 2 h-1 b

50 wt% ethanol

Lab scale

0.5 g catalyst

[31]

Nano HZSM-

5

98.4 % 98.4 % 240 °C 1 h-1 b

95 wt% ethanol

Lab scale

1 g catalyst

Run time: 620 h

[32]

0.4 mol/l

NaOH -

HZSM-5

99.6 % 99.7 % 265 °C 2.37 h-1 b

20 w% ethanol

Lab scale

1 g catalyst

Run time: 350 h

[29]

Ag3PW12O40 99.2 % 100 % 220 °C 6000 h-1 a

15 g/m3 ethanol

Lab scale

0.5 cm3 catalyst

Run time: 1 h

[43]

TPA-MCM-

41

99.9 % 98 % 300 °C 2.9 h-1 b

99.98 % ethanol

Lab scale

0.2 g catalyst

Run time: 2-3 h

[39]

STA-MCM-

41

99.9 % 99 % 250 °C 2.9 h-1 b

48 % ethanol

Lab scale

0.2 g catalyst

Run time: 8 h

[44]

P90 (Sasol)

θ-Al2O3

99.6 % 100.0 % 400 °C 1.43 h-1 b

7.9 wt % ethanol

Lab scale

0.5 g catalyst

[27]

P200 (Sasol)

γ, δ-Al2O3

100.0 % 99.7 % 350 °C 1.43 h-1 b

7.9 wt % ethanol

Lab scale

0.5 g catalyst

[27]

a Gas Hourly Space Velocity

b Weight Hourly Space Velocity

11

V200 (UOP)

γ-Al2O3

99.4 % 99.8 % 400 °C 1.43 h-1 b

7.9 wt % ethanol

Lab scale

0.5 g catalyst

[27]

D100

(Degussa/

Evonik)

γ-Al2O3

99.9 % 99.3 % 350 °C 1.43 h-1 b

7.9 wt % ethanol

Lab scale

0.5 g catalyst

[27]

SA330

(Strem)

amorphous

100.0 % 99.8 % 350 °C 1.43 h-1 b

7.9 wt % ethanol

Lab scale

0.5 g catalyst

[27]

2.5 Commercial processes available

Three available commercial processes have been investigated. These are developed by

Chematur, British Petroleum (BP) and Axens together with Total and IFPEN. The process by

BP is called Hummingbird, while the process by Axens is called Atol. Chematur and Axens

use adiabatic reactors, although the process of Chematur operates with four tubular reactors

while Atol utilizes two fixed beds [19, 45, 46]. The presented capacity from Chematur is

5,000 – 200,000 tonnes per year [19] while Axens states that Atol produce 50,000 – 400,000

tonnes per year [45]. A draft of Chematur’s process is presented in Figure 3. Syndol catalysts,

using Al2O3− MgO/SiO2, are utilized in the process by Chematur. This catalyst was

developed by American Halcon Scientific Design Inc. in the 1980’s [22]. The reactor in the

process from Atol runs at 400 – 500 °C [45]. A heteropolyacid is used as catalyst in the

process from British Petroleum and the reactor runs at temperatures between 160 and 270 °C.

The pressure in the reactor is between 1 bar and 45 bar. The process contains recirculation of

unreacted ethanol. [47]

12

Figure 3. Flow sheet of ethylene production – Chematur [19].

2.6 Ethylene purification

The amount of energy needed for dehydration depends on how pure the ethylene has to be [1].

Different reaction conditions give different amounts of by-products and have an impact on

what purification units are needed [48]. If the ethylene should be used for polymerization it

has to reach a certain level of purity [1]. According to an economic analysis made by Zhang

and Yu [22] the cost of purification after the dehydration is low compared to purification after

a process with petroleum.

In the ethylene production process from Chematur the purification of ethylene mainly consists

of a quench column, a caustic wash column, dryers, an ethylene column and a stripper [19].

The only information that was found about the purification section in the Atol process from

Axens was that it is simplified and contains no caustic tower or C2 splitter, which separates

compounds of two carbons [49]. The purification in the Hummingbird process from BP is

simplified [46].

13

3 Discussion regarding literature study

Most of the by-products formed and minerals solved during bioethanol production are

removed in the distillation, leaving only traces behind. Small amounts of a contaminant can

however still be a problem. It is desirable to have a continuous process operating during long

time intervals without downtimes. Having traces that deactivate the catalyst may hinder the

continuous operation, leading to a lower productivity. For most of the potential contaminants

mentioned in the literature study no information was found about how or if they affect the

catalysts. Thus it would be advantageous to further investigate how these trace compounds

influences the catalyst. This can be done in lab scale experiments.

In the zeolites there are exchangeable cations. When the ethanol contains inorganic ion

impurities from minerals, these ions can replace the zeolite ions. Due to this there are less

active sites, which results in lower activity of the catalyst. If the newly attached cation is too

large it is possible that it will cover larger parts of the zeolite. Consequently the cross

sectional area of the pores will be lowered and reduce the transportation in the particle. Trees

contain less inorganic material than agricultural crop residues. Therefore it could be an

advantage to use bioethanol from wood derived ethanol over bioethanol from agricultural

crops.

Coke is formed in the dehydration process, mainly due to polymerization of ethylene. In the

studied literature, coke is often mentioned as an important factor connected to the stability of

the catalysts. The coking is a central issue for all catalyst groups investigated. When building

a process it is relevant to take the coking into account and the regeneration step should be

considered carefully.

All catalysts that were studied seemed to have a high conversion of ethanol and high ethylene

selectivity. Alumina oxide catalysts are stable but require a high reaction temperature. The

zeolite catalysts can run at lower temperatures but without modifications they seem less

stable. The modifications that have been studied were only tested in lab scale; what happens

with phenomena like pressure drop and changes in mass transfer when the process is scaled

up is not known. Heteropolyacids has the possibility to provide good process conditions in the

conversion of ethanol to ethylene. Their advantage is the same as for the zeolites; high

activity because of the high acidity results in the possibility of low operating temperatures.

One disadvantage is the low thermal stability which will require gentle operating when

decoking. The main focus on the listed HPAs has been on ethanol conversion and the ethylene

selectivity. Further investigations should focus more on the long-term stability of the catalyst

to know for certain if these types of catalysts are appropriate for the process or not.

14

4 Feasibility study

The process mainly consists of one reactor and subsequent purification, see the block diagram

in Figure 4. Unreacted ethanol that is recycled in order to obtain a higher conversion of

ethanol is also presented in Figure 4.

Figure 4. Block diagram illustrating the two main parts of the process, reactor and subsequent purification.

Besides ethylene, the raw ethylene stream shown in Figure 4 contains water, water soluble

compounds, carbon dioxide, hydrocarbons heavier than ethylene, carbon monoxide and

hydrogen. These impurities are separated in five units in order to obtain polymer grade

ethylene. The units needed for purification can be seen in Figure 5. A flowsheet over the

process is presented in Figure 6.

Figure 5. Block diagram showing the purification units used in the process.

15

Heat exchanger 3

NaOH (aq) storage tank

Oil heater

C2 stripper

Ethylene column

Absorber Gas/liquid separator

Reaction column

Heat exchanger 1

Distillation 1

Heat exchanger 2

Ethylene storage

tank

Dryer

Waste water treatment

Heat exchanger 4

Flue gas1

3

2

4

5

6

10

9

8

7

12

13

11

14

16

1517

18

19

20

21

24

22

Oxygen

23

Waste water treatment

Heat exchanger 5

25

Ethanol storage tank

26

27

Pure water

Figure 6. Flowsheet of the process.

16

4.1 Process steps

The used process units in the plant and why they are used are described here.

4.1.1 Storage tanks Storage of feedstock, ethanol, and product, ethylene, is necessary to keep the process running

without being too dependent on transportation to and from the plant. A storage tank for

sodium hydroxide-water solution is also required.

4.1.2 Pumps Pumps were placed in the system to create a flow through the process. The pumps assure a

pressure of 1 bar throughout the system. The pumps are also designed to be able to raise the

pressure in the system to 10 bar, if this would be desirable.

4.1.3 Heating and vaporization The ethanol is delivered to a heat exchanger (Heat exchanger 1) where the ethanol is heated

and partly vaporized. The heat exchanger utilizes the product stream from the reactor as a hot

stream. Another heat exchanger (Heat exchanger 2) placed after the first one is then used

to vaporize the remaining liquid ethanol and to heat the now gaseous ethanol to the reaction

temperature which is 240 °C.

4.1.4 Reactor The reactor used in this process design is a fixed bed reactor with isotherm heat supplement.

This requires a tube bundle but enables to control the reaction temperature closely. The tube

bundle of the reactor is filled with heteropolyacid catalyst that favors the reaction from

ethanol to ethylene. It is important to keep the reaction temperature constant and not too high.

The reactor is kept at constant temperature by having a jacket of circulated heating media, see

section 4.1.5.

4.1.5 Circulating heating media Condensing Dowtherm A oil at 1 bar and 288 °C is chosen as circulating heating media. The

heating media is used in Heat exchanger 2, the reactor, the reboiler of distillation tower 1, the

reboiler of the ethylene column and the reboiler of the C2 stripper. The external energy

needed to heat the oil is lowered by recovering heat from the waste water and the flue gases.

The heating stream to the reboilers is not illustrated in Figure 6.

4.1.6 Cooling and phase separation The product stream from the reactor is cooled in two steps, first in Heat exchanger 1 where it

heats the feed and then in Heat exchanger 3. After cooling there will be both a gaseous and

a liquid phase. The liquid phase will mainly contain water and unreacted ethanol. In the

gaseous phase the main component will be ethylene, but other components such as diethyl

ether and acetaldehyde will also be present. The two phases are separated in a gas/liquid

separator in which water and water-soluble compounds hence are separated from the ethylene.

4.1.7 Recirculation and distillation of unreacted ethanol The stream containing water and water-soluble compounds is delivered to a distillation

column (Distillation 1) where unreacted ethanol is separated from the water. The water is sent

to a wastewater treatment step and the ethanol is recirculated back to Heat exchanger 2 before

entering the reactor together with the feed stream.

17

4.1.8 Caustic soda wash A caustic soda wash, an absorption tower using sodium hydroxide, is used. Carbon dioxide is

separated from the ethylene stream by letting the carbon dioxide react with sodium hydroxide.

A 50 wt % NaOH-water solution is used. [50]

4.1.9 Dryer system Two adsorption towers packed with zeolite is used to remove remaining water from the

gaseous stream coming from the caustic soda wash. Water is adsorbed to zeolites which

thereafter can be regenerated by heating. The two towers work alternately, one of the towers

is operating while the other is regenerated.

4.1.10 Cooling and ethylene column The ethylene gas, free from water, still contains impurities that need to be separated. The gas

stream is cooled in Heat exchanger 4 before entering a distillation tower (Ethylene column)

where remaining hydrocarbons are separated. In the bottom of the column, hydrocarbons with

three or more carbon atoms are taken out. The top stream, containing ethylene and lighter

gases, is led to a C2 stripper column.

4.1.11 C2 stripper column The ethylene stream coming from the ethylene column is cooled in Heat exchanger 5 and then

enters a final distillation tower (C2 stripper). In this last purification step carbon monoxide

and hydrogen is removed. Polymer grade ethylene can be taken out from the bottom of the

stripper; this ethylene stream contains at least 99.9 wt % ethylene.

4.1.12 Thermal fluid heater The by-products separated in the last two distillation columns have favorable heating values

and is therefore burnt with air in a thermal fluid heater. The thermal fluid heater heats the

Dowtherm A oil that is used as a circulating heating media throughout the process. The heat

from the by-products is not enough to supply the oil with heat and therefore also some natural

gas is added to the thermal fluid heater.

4.1.13 Wastewater treatment Wastewater is obtained when removing water from the process in the gas/liquid separator and

from the Caustic soda wash. The wastewater requires removal of a variety of compounds. The

compounds that mainly need to be removed are ethylene and other shorter hydrocarbons,

ethanol, acetaldehyde and ions such as hydroxide, sodium and carbonate.

4.1.14 Heat pump The heat pump is used for cooling of Heat exchanger 4, the condenser of the ethylene column,

Heat exchanger 5 and the condenser of the C2 stripper. It is needed since the temperatures of

the previously mentioned unit operations operate at very low temperatures and a cooling

media that can manage these temperatures is required. The heat pump is not illustrated in

Figure 6.

4.2 Design

To calculate the designs of each operating unit in the process seen in Figure 6, heat and mass

balances were obtained by simulating the system in Aspen Plus V8.2 – AspenONE.

To produce 100,000 tonnes of ethylene annually a certain amount of ethanol is needed. The

amount of ethanol depends on the estimated conversion and selectivity.

18

In calculations the conversion of ethanol was set to 95 mole-% and the selectivity to ethylene

was set to 98 mole-%. The calculations resulted in that 192,000 tonnes of 95 wt-% ethanol

was needed. For calculations see Appendix I. To be sure that the right amount of ethylene is

produced, a higher amount of ethanol was used when modelling the process. The yield of

ethylene over the reactor, with recirculation of ethanol, was 96.8 %. For amounts given from

Aspen see Appendix II, Table 10.

4.2.1 Design of process units To calculate the dimensions of each operation the results from Aspen were used, see

Appendix II, Table 10. The results of the dimensioning can be seen in Table 3Table 3.

Equations and assumptions can be seen in Appendix IV and Appendix V, Tables 13-39.

Table 3. Design results and operating conditions for the unit operations.

Unit Design Operating conditions

Ethanol storage tank Total vessel volume: 6,500 m3

Pump 1 Power: 12 kW Pressure increase: 10 bar

Heat exchanger 1 Area: 167 m2 Hot stream: 240-70 °C

Cold stream: 10-80 °C

Heat exchanger 2 Area: 284 m2 Hot stream: Dowtherm A oil, 288

°C

Cold stream: 81-240 °C

Reactor Total vessel volume: 8.1 m3

Number of tubes: 1,052

Tube length: 3.0 m

Tube diameter: 0.03 m

Catalyst: Heteropolyacid on

silica alumina

Catalyst mass: 1,870 kg

Heat exchanger area: 297 m2

Temperature: 240 °C

GHSV: 6,000 h-1

Heat exchanger 3 Area: 409 m2 Hot stream: 70 – 25 °C

Cold stream: 15-30 °C

Gas/Liquid separator Inside diameter: 1.47 m

Height: 2.52 m

Temperature: 25 °C

Residence time: 600 s

Pump 2 Power: 3 kW Pressure increase: 10 bar

NaOH storage tank Total vessel volume: 378 m3

Absorption column Inside diameter: 1.51 m

Height: 26.3

Nr of ideal stages: 35

Top temperature: 33.2 °C

Bottom temperature: 29.7 °C

Distillation column 1 Inside diameter: 0.51 m

Height: 18.8 m

Nr of ideal stages: 25

Top temperature: 78 °C

19

Condenser area: 100 m2

Reboiler area: 161 m2

Bottom temperature: 100 °C

Dryer, 2 pieces Inside diameter: 2.5 m

Height: 5.94 m

Adsorbent: Zeolite

Adsorbent mass: 16,970 kg

Packing height: 4.94 m

Residence time: 12 h

Heat exchanger 4 Area: 62 m2 Hot stream: 32 – -104 °C

Cold stream: Cooling media,

-150 °C

Ethylene column Inside diameter: 1.49 m

Height: 11.3 m

Condenser area: 216 m2

Reboiler area: 50 m2

Nr of ideal stages: 15

Top temperature: -104.2 °C

Bottom temperature: -98.3 °C

Heat exchanger 5 Area: 511 m2 Hot stream: -104 – -130 °C

Cold stream: Cooling media,

-150 °C

C2 stripper column Inside diameter: 2.30 m

Height: 7.5 m

Condenser area: 13 m2

Reboiler area: 9.0 m2

Nr of ideal stages: 10

Top temperature: -109.7 °C

Bottom temperature: -103.9 °C

Ethylene storage tank Total vessel volume: 4,920 m3

Heat pump Heat absorption rate: 3,710 kW

Wastewater treatment

plant

Capacity: 0.0033 m3/s

Thermal fluid heater Heating duty: 10,160 kW

20

4.3 Cost estimates

In order to determine the feasibility of the project, an economical evaluation was made. The

evaluation covered the capital cost needed to build the plant, annual operating costs and a

sensitivity analysis of how the pay-back time varies on feedstock cost, investment cost and the

price of ethylene. The initial case description was based on the following: an economic life

time of 15 years, a weighted average cost of capital of 10 % and an ethanol price of 537

USD/tonne [51]. The ethylene price was assumed to be 1,280 USD/tonne. The initial case

description can be seen in Table 6 and a summary of the total production cost in Table 7.

4.3.1 Capital costs A capital cost estimation method called Ulrich method was used for economical calculations.

The final result in capital cost for the plant is called the grass root capital. For this project, the

online database EconExpert was used to calculate the costs. All units and their Bare-Module

Cost can be seen in Table 4. Each cost is based on different properties, for example area and

volume. A more detailed description of the method, with equations, and all assumptions

regarding the unit operations can be found in Appendix VI and in Appendix VII, Tables 40-

42.

Table 4. Bare-Module Cost for all unit operations.

Unit Bare-Module Cost, USD

Ethanol storage tank 542,150

Pump 1 45,110

Heat exchanger 1 96,750

Heat exchanger 2 137,710

Reactor 332,710

Heat exchanger 3 178,370

Gas/Liquid separator 114,670

Pump 2 26,430

NaOH storage tank 85,190

Absorption column 878,860

Distillation column 1 514,840

Dryer, 2 pieces 868,820

Heat exchanger 4 52,700

Ethylene column 533,540

Heat exchanger 5 210,020

C2 stripper column 497,000

Ethylene storage tank 787,070

Heat pump 6,896,130

Wastewater treatment plant 991,410

Thermal fluid heater 1,175,270

21

The grass root capital for the plant was calculated to 28,105,200 USD, which includes the

fees, contingencies and auxiliary facilities. How the costs are distributed can be seen more

clearly in Figure 7 and 8. The largest cost in the grass root capital was the heat pump, which

accounted for 46 % of the total Bare-Module Cost. The heat pump is necessary to reduce to

electrical costs and creates a more profitable process.

Figure 7. Distribution of the grass root capital costs.

Figure 8. Detailed description of how the unit operation investment costs are distributed.

4.3.2 Operating costs The annual operating costs for the plant have been divided into three categories, fixed costs,

direct costs and indirect costs. The costs have been calculated by using rules of thumb

available in Project Handbook [52]. See Appendix VIII for calculations. The Appendix also

contains the annuity factor, fa which is used to calculate the annual capital cost [52]. The

annuity factor is calculated to 0.1315. To see how the operating costs are distributed on

different areas, see Figure 9 and 10.

1%

99%

Grass Root Capital Cost

Material & supply

Unit operations

10% 4%

10%

6%

6%

8%

2%

46%

7%

1%

Unit operation investment cost

Storage Tanks & Pumps

Heat Exchangers

Distillation columns

Absorber

Dryers

Thermal fluid heater

Reactor

Heatpump

Wastewater treatment

22

4.3.2.1 Fixed capital

The fixed costs include the cost of storage for both feedstock and products and also the cost

for spare parts. Feedstock and products is assumed to be stored for 10 days and the cost for

spare parts have been set to 15 % of the cost for maintenance and repair. The total amount of

fixed costs sums up to 884,360 USD. The calculations can be seen in Appendix VIII, Table

43.

4.3.2.2 Direct costs

Direct costs include costs for consumables such as feedstock, solvents, catalyst, electricity and

cooling water. The price for some consumables can be seen in Table 5. It also includes costs

for maintenance and repair, operators, supervisors and laboratory work. The price for the

consumables was set according to market prices today, see Appendix VIII Table 44 and 45.

The cost for maintenance and repair has been set to 6 % of the grass root capital cost.

The plant will budget for five operators working in shifts. Their monthly salary is 3,300

USD/operator. There will also be one operator working day, with the same salary. Supervisor

costs and laboratory work are assumed to be 15 % of the operator cost respectively.

Table 5. Price of consumables.

Consumable Price Source/Comment

Feedstock - Ethanol 357 USD/tonne [51]

Solvents - NaOH 520 USD/tonne [53]

Natural gas 9 USD/GJ Assumptions by authors

Electricity 0.13 USD/kWh Assumptions by authors

Cooling water 0.013 USD/m3 Assumptions by authors

4.3.2.3 Indirect costs

Indirect costs are overhead costs for staff and personnel, administration and distribution as

well as sales. To calculate the overhead for staff 70 % was added on the cost for shift

personnel and 50 % was added on the cost for day personnel. Administration costs were

estimated to be 25 % of the overhead costs for staff. The calculations can be seen in Appendix

VIII, Table 46.

Figure 9. Distribution between operating costs for the process, the raw material is ethanol.

89%

11%

Operating costs

Raw material

Other operatingcosts

23

Figure 10. Detailed description of how the other operating costs than raw material is distributed.

4.3.3 Total production cost A summary of the plant’s production cost can be seen in Table 7. Based on the initial case

description, see Table 6, the production cost per tonne produced ethylene, can also be seen in

the table and, is calculated to 1,136 USD/tonne. See Appendix VII and VIII for all data and

Appendix VIII for calculations of the annuity factor.

Table 6. Initial case description.

Source/Comment

Economic lifetime 15 years Assumptions by authors

Weighted average cost of capital 10 % Assumptions by authors

Ethanol price 537 USD/tonne [51]

Ethylene price 1,280 USD/tonne Assumptions by authors

Table 7. Summary of the total production cost.

Cost

Fixed capital 884,360 USD/year

Direct costs 110,420,320 USD/year

Indirect costs 900,000 USD/year

Annual operating costs 112,204,680 USD/year

Grass root capital 28,150,200 USD

Annual capital cost 3,695,100 USD/year

Total annual production cost 115,899,800 USD/year

Production cost per tonne ethylene 1,136 USD/tonne

5% 7% 7%

14%

17% 20%

30%

Other operating costs

Natural gas

Electricity

Storage & spare parts

Maintenance & repair

Personnel

Other materials

Annual capital cost

24

4.3.4 Investment calculations An investment calculation was made to determine the pay-back time for the plant. Pay-back

time is a relative method which compounds the costs and revenues. The price of ethylene has

been set according to market prices and is 1,280 USD/tonne. This price is slightly higher than

the calculated production cost. With the given price of ethylene, the pay-back time is

calculated to 1.75 years. The interest rate has been chosen to 10 %. Table 8 displays a

summary of the calculations and the equations that were used can be seen in Appendix IX.

Table 8. Summary of investment calculations and pay-back time.

Cost

Grass root capital 28,150,200 USD

Annual operating costs 112,204,680 USD/year

Revenue 130,542,800 USD/year

Pay-back time 1.75 years

4.3.5 Sensitivity analysis A sensitivity analysis has been made to examine how the pay-back time and profit for the

plant varies with changes in ethanol price, ethylene price and grass root capital cost. The

initial case description, see Table 6, was used as a base to compare the results of the changes

with. The cost of ethanol is the largest operating cost for the plant. It is therefore important to

see how the ethanol price affects the plant economically. For green ethylene, a premium can

be added to the price. The premium can vary between 20-30 %. With a higher price on

ethylene the income will increase.

The sensitivity analysis of how the pay-back time changes with different ethanol price is

shown in Figure 11. With a 15 % increase in ethanol price the annual profit will be 3,700,000

USD. This relates to a pay-back time of 15 years. If the price instead decreases with 15 % the

annual profit will be 34,700,000 USD and the pay-back time will be 0.89 years.

Figure 11. Sensitivity analysis of how the pay-back time changes with different ethanol prices.

0

2

4

6

8

10

12

14

16

-50% -40% -30% -20% -10% 0% 10% 20%

Pay

-bac

k ti

me

(ye

ars)

Ethanol price

Sensitivity analysis of changes in ethanol price

25

There is a large difference between the two scenarios and compared to the regular price of

ethanol, an increase of ethanol price has a very large impact on the pay-back time and the

annual profit. As can be seen in Figure 11, an increase in ethanol price makes a larger impact

than a decrease. If the ethanol price increases with more than 18 % there will be no annual

profit. The operating costs will be higher than the revenues and the plant will not be

profitable. How the annual profit varies with ethanol price can be seen in Figure 12.

Figure 12. Sensitivity analysis of how the profit changes with different ethanol prices.

When adding 25 % to the price of ethylene the pay-back time decreases to 0.6 years compared

to 1.75 years for the original price. With the premium, the ethanol price could increase up to

50 % and the process would still generate an annual profit. The results of how pay-back time

changes with ethylene price can be seen in Figure 13. The price has been increased with 20-30

% and also decreased with 10 %, the results are compared to the original price of ethylene.

-20.000.000

-10.000.000

0

10.000.000

20.000.000

30.000.000

40.000.000

50.000.000

60.000.000

70.000.000

-50% -40% -30% -20% -10% 0 5% 10% 15% 20% 30%

Pro

fit

(USD

/ye

ar)

Ethanol price

Sensitivity analysis of the profit with changes in ethanol price

26

Figure 13. Sensitivity analysis of how the pay-back time changes with different changes of ethylene price.

Finally, a sensitivity analysis of the grass root capital was made. Ulrich's method does not

offer an exact accuracy, for example when calculating the installation cost. The installation

cost for the process is most likely higher than the one Ulrich's method has provided. It is

therefore interesting to see how the grass root capital will change in size and how this can

affect the process economy and pay-back time. See results in Figure 14 and 15.

Figure 14. Sensitivity analysis of how the grass root capital, capital cost and annual net income varies when the

grass root capital increases and decreases.

As can be seen in Figure 14 when the grass root capital increases it will result in a higher

capital cost each year. The capital cost is calculated with the grass root capital and annuity

factor, see Appendix IX. Its size will reflect on the annual net income, which is the annual

profit minus the capital cost, and lower it. Figure 15 displays the pay-back time for the

8,0

1,8

0,7

0,6

0,5

0,0 1,0 2,0 3,0 4,0 5,0 6,0 7,0 8,0 9,0

90%

100%

120%

125%

130%

Pay-back time (years)

Eth

yle

ne

pri

ce

Sensitivity analysis of changes in ethylene price

- 5.000.000

10.000.000 15.000.000 20.000.000 25.000.000 30.000.000 35.000.000 40.000.000 45.000.000 50.000.000

-30% 0 30% 50% 75%

USD

Grass root capital change

Sensitivity analysis of capital costs and income with changes in grass root capital

Grassrootcapital

Cost ofcapital

Annualnetincome

27

different grass root capital costs. A higher cost consequently means a longer pay-back time

for the plant.

Figure 15. Sensitivity analysis of how the pay-back time changes with changes of the grass root capital costs.

-

0,50

1,00

1,50

2,00

2,50

3,00

3,50

-30% 0 30% 50% 75%

Pay

-bac

k ti

me

(ye

ars)

Grass root capital change

Sensitivity analysis of pay-back time with changes in grass root capital

Pay-backtime

28

5 Discussion

The reaction is preferably run at as low a pressure as possible to push the equilibrium as far as

possible towards ethylene production. This has been overcome by running the simulations in

Aspen at the pressure of one bar. Other parts of the process such as the quench, the absorber

and the cryogenic distillation towers will on the other hand run optimal at higher pressures, to

minimize the cooling duty required in these systems. To improve the system a higher pressure

would preferably be set directly after the ethanol storage tank with the advantage of pumping

liquid ethanol compared to compressing gases in the downstream processing. This would

cause the reaction to not reach the optimal equilibrium, but also leads to a lesser amount of

catalyst needed because of the increase in concentration and thus an increase in reaction rate.

Since the lowered conversion and selectivity is included when calculating the mass balances

the higher pressure is considered to have a positive effect on the energy consumption, in the

downstream processing and the storage of ethylene. The size of each process equipment could

also become smaller. The usage of a higher pressure has been rejected in this study because of

the problem to make the simulations converge. This results in that the operating costs and

catalyst cost probably are estimated higher than the actual cost, which is considered better

than the opposite case.

In the process butylenes are produced as by-products. These are, due to their favorable

heating value, burnt in a thermal fluid heater to supply the process with heat. However,

butylene is a valuable chemical and it would therefore be more desirable to separate and

purify it rather than burning it.

The quench heat exchanger and the gas/liquid separator used in the process operates with

large amounts of water. To make the separation more energy efficient the separation could

have been performed in many steps connected in series. Removing more water increases the

carbon dioxide concentration which should make the absorber more efficient.

The sodium hydroxide scrubber, operating to remove the CO2 in the process, produces waste

water that is sent to wastewater treatment. In reality the sodium hydroxide would need to be

regenerated to make the process more sustainable. As Zeman and Lackner [50] explains the

produced sodium carbonate (Na2CO3) can be mixed with calcium hydroxide (Ca(OH)2) to

recover sodium hydroxide, see reaction 2. This procedure also creates a precipitate of calcium

carbonate (CaCO3). This precipitate can be filtered away, dried, washed and thermally

decomposed, see reaction 3. The calcium carbonate is decomposed to lime (CaO), which then

can be hydrated to produce calcium hydroxide to close the circle, see reaction 4. The

recycling of sodium hydroxide is well known in the pulp and paper industry (Kraft process).

𝑁𝑎2𝐶𝑂3(𝑎𝑞) + 𝐶𝑎(𝑂𝐻)2(𝑠) → 2 𝑁𝑎𝑂𝐻(𝑎𝑞) + 𝐶𝑎𝐶𝑂3(𝑠) (reac. 2)

𝐶𝑎𝐶𝑂3(𝑠) + 𝑒𝑛𝑒𝑟𝑔𝑦 → 𝐶𝑎𝑂(𝑠) + 𝐶𝑂2(𝑔) (reac. 3)

𝐶𝑎𝑂(𝑠) + 𝐻2𝑂(𝑙) → 𝐶𝑎(𝑂𝐻)2(𝑠) (reac. 4)

Two dryers are used to remove the final water in the process. While one dryer is operating the

other one is regenerated. Regeneration can be performed thermally, removing the water from

the zeolite by vaporization. This requires heat, but the cost of this heat is not included in the

economical evaluation.

The simulation implemented in Aspen is not optimized due to time constraints. By tweaking

the reflux ratios, distillation rates etc. of the distillation towers the energy requirements of

these could be lowered. The changes could have been done by using Aspen design

specifications. Implementing these improvements would probably lead to lower operational

29

costs due to lower energy requirements. Another thing that may be exaggerated is the number

of ideal trays in the absorption column. Lowering the number of trays would end up in a

smaller absorption tower and therefore lower investment cost.

The last two distillation columns and the last two heat exchangers in the process operate at

very low temperatures. To achieve this low temperatures a heat pump with cooling media

have to be used, for example pure ethylene. In the economical evaluation the cost of the

cooling media is not included. As mentioned above, there would be many advantages having

the process pressurized. Another advantage is that the last two distillation columns and the

last two heat exchangers could be operated at higher temperatures, making it easier to

construct the heat pump. This would also make the cost of the heat pump implemented in the

process more reliable. Ulrich method can’t handle heat pumps that works at temperatures

under -55 °C, but is still used even though the Aspen simulation displays cooling

requirements at under -100 °C.

The ethylene is stored in a tank at low temperature. The storage tank is not completely

adiabatic and will therefore receive heat from the surrounding. To keep the ethylene at low

temperature the tank needs to be cooled. The cost of this cooling is not included in the

economical evaluation.

The installation cost and maintenance and repair cost is dependent on how difficult it is to fill

and remove the catalyst from the reactor. Handling the catalyst will be more difficult if the

reactor has a tube bundle filled with catalyst instead of just using a packed reactor without

tube bundles. This is not something that Ulrich method takes into account. The reactor used in

the constructed ethylene process uses tube bundles and the cost for the reactor may therefore

be slightly underestimated.

The economic evaluation is rather optimistic and favorable given the current assumptions.

The database method that was used for calculating the capital costs is not completely reliable.

For example, as mentioned before, the installation cost will most likely be higher. Some

simplifications were made to run the simulation in Aspen. Therefore a couple of unit

operations have not been accounted for. Some understanding on how much the capital cost

could increase and affect the economy was given in the sensitivity analysis but the cost still

seems optimistic. An additional sensitivity analysis is recommended and possibly another

method to perform the cost estimations.

The sensitivity analysis was mainly made on pay-back time and it could be seen that the

process is sensitive for changes in market prices. If ethanol prices increase with more than 18

%, there will be no annual profit. It is not unlikely that the ethanol price increases and also, if

the ethanol is bought from outside the European Union there will be additional custom duty,

which consequently would increase the price. If it would be possible for the process, and

reactor, to run on ethanol with lower quality the ethanol price could be lower. This would

have a positive effect on the economy for the process.

As mentioned earlier there is a premium on green ethylene. The premium price is based on

what the customer is willing to pay for a green product and what the market looks like. There

is a demand for green ethylene and the price can therefore be kept high. If, on the other hand,

the production increases significantly the prices and possibly the premium will decrease.

30

6 Conclusions

Second generation bioethanol can contain contaminants that may have an impact on the

process. How these contaminants affect the catalysts’ performance in the long term has not

been found in the literature and therefore further investigations is needed in this area. Coke

formation is a problem in the dehydration process. Since none of the studied catalysts have

been used in industrial scale for manufacturing of green ethylene from 2nd

generation

bioethanol the choice of catalyst is complex.

The properties of heteropolyacids enable the process to operate at relatively low temperatures.

This together with good performance was the reasons why this catalyst was chosen in the

process design. In broad terms the process contained one reactor and several purification

units. The suggested process was simulated successfully at one bar. An ethanol flow of

192,000 tonnes/year resulted in 102,000 tonnes ethylene/year with a polymer grade purity of

99.9 %.

The economical evaluation displayed a total production cost of 1,136 USD/tonne ethylene and

a pay-back time of 1.75 years. Another method is needed for the economic calculations in

order to determine if the numbers are reliable. According to the sensitivity analysis market

prices had a large impact on the profitability of the plant.

31

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I-i

7 Appendix

Appendixes to the report are presented below.

Appendix I – Ethanol calculations

Calculation of mass ethanol needed to produce 100,000 tonnes ethylene annually.

Designations and values used in calculations are presented in Table 9.

Table 9. Designation and values used for variables used in calculations.

Designation Description Value

methylene Mass ethylene 100,000 tonnes

Methylene Molar mass ethylene 28.05 g/mole [54]

Methanol Molar mass ethanol 46.07 g/mole [55]

nethylene Amount of substance

ethylene

Is calculated

xethanol Conversion of ethanol 0.95 mole/mole

Sethylene Selectivity to ethylene 0.98 mole/mole

nethanol Amount of substance

ethanol

Is calculated

methanol Mass ethanol Is calculated

Amount ethylene, expressed in mole, which should be produced, is calculated in Equation 1.

𝑛𝑒𝑡ℎ𝑦𝑙𝑒𝑛𝑒 =𝑚𝑒𝑡ℎ𝑦𝑙𝑒𝑛𝑒

𝑀𝑒𝑡ℎ𝑦𝑙𝑒𝑛𝑒↔ 𝑛𝑒𝑡ℎ𝑦𝑙𝑒𝑛𝑒 = 3.565 ∙ 109 𝑚𝑜𝑙𝑒 Equation 1

The molar ration between ethylene and ethanol is 1:1. Taking conversion of ethanol and

selectivity to ethylene into account the amount ethanol needed, expressed in mole, is

calculated in Equation 2.

𝑛𝑒𝑡ℎ𝑎𝑛𝑜𝑙 =𝑛𝑒𝑡ℎ𝑦𝑙𝑒𝑛𝑒

𝑋𝑒𝑡ℎ𝑎𝑛𝑜𝑙∙𝑆𝑒𝑡ℎ𝑦𝑙𝑒𝑛𝑒↔ 𝑛𝑒𝑡ℎ𝑎𝑛𝑜𝑙 = 3.829 ∙ 109 𝑚𝑜𝑙𝑒 Equation 2

Amount ethanol, expressed in mass, which is needed, is calculated in Equation 3.

𝑚𝑒𝑡ℎ𝑎𝑛𝑜𝑙 = 𝑛𝑒𝑡ℎ𝑎𝑛𝑜𝑙 ∙ 𝑀𝑒𝑡ℎ𝑎𝑛𝑜𝑙 ↔ 𝑚𝑒𝑡ℎ𝑎𝑛𝑜𝑙 = 176,000 𝑡𝑜𝑛𝑛𝑒𝑠 Equation 3

To be sure that at least the desired amount of ethylene will be produced an extra amount of

ethanol was added. In the simulation of the entire process 192,000 tonnes of 95 wt-% ethanol

was used.

II-i

Appendix II – Heat and mass balances

Summary of the process streams generated by Aspen Plus V8.2 – AspenONE.

Table 10. Specifications from simulation in Aspen Plus V8.2 – AspenONE.

Stream ID 1 3 4 5 6 7 8 9 10 11 12

From Storage Heat Ex 1 3 & 10 Heat ex 2 Reaction

column

Heat Ex 1 Heat Ex 3 G-L separator Distillation 1 Distillation 1 G-L separator

To 4 Heat ex 2 Reaction

column

Heat Ex 1 Heat Ex 3 G-L separator Distillation 1 4 Absorber

Temperature C 10 81,00146 80,81845 240 240 70 25 25,00015 75,8087 99,35311 25,00015

Pressure BAR 1 1 1 1 1 1 1 1 1 1 1

Vapor Fraction 0 0,786987 0,794862 1 1 0,677533 0,467026 0 1 0 1

Liquid Fraction 1 0,213013 0,205138 0 0 0,322467 0,532974 1 0 1 0

Mass Flow KG/HR 22000 22000 22931,65 22931,65 22931,65 22931,65 22931,65 9978,459 931,3098 9047,149 12953,19

Component Mass

Flow

ETHYLENE KG/HR 0 0 1,96998 1,969984 12323,1 12323,1 12323,1 1,969911 1,969911 5,79E-23 12321,13

ETHANOL KG/HR 20900 20900 21732,96 21732,96 1085,719 1085,719 1085,719 1018,694 832,624 186,0701 67,02491

WATER KG/HR 1100 1100 1100,335 1100,335 9108,672 9108,672 9108,672 8861,41 0,3310225 8861,079 247,2627

HYDROGEN KG/HR 0 0 0,000174 0,000174 8,695498 8,695498 8,695498 0,000174241 0,000174241 0,00E+00 8,695324

PROPENE KG/HR 0 0 0,348436 0,348436 3,884474 3,884474 3,884474 0,3484269 0,3484269 1,22E-16 3,536047

ISOBUTYLENE KG/HR 0 0 70,63443 70,63443 259,2231 259,2231 259,2231 70,63458 70,63458 3,29E-09 188,5886

ETHANE KG/HR 0 0 0,573134 0,573134 21,62941 21,62941 21,62941 0,5731143 0,5731143 1,58E-21 21,0563

DIETHYL-ETHER KG/HR 0 0 1,329672 1,329672 2,160144 2,160144 2,160144 1,329756 1,32975 5,44E-06 0,830388

ACETALDEHYDE KG/HR 0 0 23,42929 23,42929 48,10788 48,10788 48,10788 23,43021 23,43021 4,79E-07 24,67768

CO KG/HR 0 0 3,81E-06 3,81E-06 0,00362 0,00362 0,00362 3,81E-06 3,81E-06 5,35E-25 0,003616

CO2 KG/HR 0 0 0,060798 0,060798 67,80514 67,80514 67,80514 0,0607944 0,0607944 1,39E-19 67,74435

METHANE KG/HR 0 0 0,007819 0,007819 2,650044 2,650044 2,650044 0,00781874 0,00781874 1,40E-23 2,642225

Component Mass

Fraction

ETHYLENE 0 0 8,59E-05 8,59E-05 0,537384 0,537384 0,537384 0,000197416 0,00211521 6,40E-27 0,951204

ETHANOL 0,95 0,95 9,48E-01 9,48E-01 0,047346 0,047346 0,047346 0,1020893 0,8940355 2,06E-02 0,005174

WATER 0,05 0,05 0,047983 0,047983 0,39721 0,39721 0,39721 0,888054 0,000355438 0,979433 0,019089

HYDROGEN 0 0 7,60E-09 7,60E-09 0,000379 0,000379 0,000379 1,75E-08 1,87E-07 0 0,000671

PROPENE 0 0 1,52E-05 1,52E-05 0,000169 0,000169 0,000169 3,49E-05 3,74E-04 1,35E-20 0,000273

ISOBUTYLENE 0 0 3,08E-03 3,08E-03 0,011304 0,011304 0,011304 7,08E-03 0,0758443 3,64E-13 0,014559

ETHANE 0 0 2,50E-05 2,50E-05 0,000943 0,000943 0,000943 5,74E-05 0,000615385 1,75E-25 0,001626

DIETHYL-ETHER 0 0 5,80E-05 5,80E-05 9,42E-05 9,42E-05 9,42E-05 1,33E-04 0,00142783 6,01E-10 6,41E-05

ACETALDEHYDE 0 0 1,02E-03 1,02E-03 2,10E-03 2,10E-03 2,10E-03 0,00234808 0,0251583 5,30E-11 1,91E-03

CO 0 0 1,66E-10 1,66E-10 1,58E-07 1,58E-07 1,58E-07 3,82E-10 4,09E-09 5,91E-29 2,79E-07

CO2 0 0 2,65E-06 2,65E-06 2,96E-03 2,96E-03 2,96E-03 6,09E-06 6,53E-05 1,54E-23 5,23E-03

METHANE 0 0 3,41E-07 3,41E-07 0,000116 0,000116 0,000116 7,84E-07 8,40E-06 1,55E-27 0,000204

Mole Flow KMOL/HR 514,7262 514,7262 534,7337 534,7337 981,0151 981,0151 981,0151 515,9036 20 495,9036 465,1115

Component Mole

Fraction

ETHYLENE 0 0 0,000131 0,000131 0,447768 0,447768 0,447768 0,000136109 0,00351096 4,16E-27 0,944283

ETHANOL 0,881375 0,881375 0,88221 0,88221 0,024023 0,024023 0,024023 0,0428613 0,9036697 0,008145 0,003128

WATER 0,118625 0,118625 0,114221 0,114221 0,515393 0,515393 0,515393 0,9534398 0,000918727 0,991855 0,029509

II-ii

HYDROGEN 0 0 1,62E-07 1,62E-07 0,004397 0,004397 0,004397 1,68E-07 4,32E-06 0,00E+00 0,009274

PROPENE 0 0 1,55E-05 1,55E-05 9,41E-05 9,41E-05 9,41E-05 1,60E-05 0,000413999 5,87E-21 0,000181

ISOBUTYLENE 0 0 0,002354 0,002354 0,00471 0,00471 0,00471 0,00244021 0,0629457 1,18E-13 0,007227

ETHANE 0 0 3,56E-05 3,56E-05 0,000733 0,000733 0,000733 3,69E-05 9,53E-04 1,06E-25 0,001506

DIETHYL-ETHER 0 0 3,35E-05 3,35E-05 2,97E-05 2,97E-05 2,97E-05 3,48E-05 0,000896991 1,48E-10 2,41E-05

ACETALDEHYDE 0 0 0,000995 0,000995 0,001113 0,001113 0,001113 0,00103093 0,0265931 2,19E-11 0,001204

CO 0 0 2,54E-10 2,54E-10 1,32E-07 1,32E-07 1,32E-07 2,63E-10 6,80E-09 3,85E-29 2,78E-07

CO2 0 0 2,58E-06 2,58E-06 1,57E-03 1,57E-03 1,57E-03 2,68E-06 6,91E-05 6,38E-24 3,31E-03

METHANE 0 0 9,11E-07 9,11E-07 0,000168 0,000168 0,000168 9,45E-07 2,44E-05 1,76E-27 0,000354

Volume Flow m^3/HR 27,09045 11933,82 12515 22814,41 41854,99 18969,93 11368,06 10,30996 580,2716 9,906551 11529,74

Stream ID 14 15 16 17 18 19 21 20 22 23 24 25 26 27

From Absorber Absorber Dryer Dryer Heat ex 4 Ethylene

column

Ethylene

column

Heat ex 5 C2 Stripper C2 Stripper 21 & 24 Oil heater

To Absorber Dryer Heat ex 4 Ethylene

column

25 Heat ex 5 C2 Stripper Ethylene

storage tank

25 Oil heater Oil heater

Temperature C 25 33,24234 29,67447 32,31291 32,31291 -104 -98,3353 -104,149 -130 -103,9743 -109,7252 -104,3031 20 1252,258

Pressure BAR 1 1 1 1 1 1 1 1 1 1 1 1 1 1

Vapor Fraction 0 1 0 0 1 0,457764 0 1 0,0118454 0 1 0,463977 1 1

Liquid Fraction 1 0 1 1 0 0,542236 1 0 0,9881545 1 0 0,536023 0 0

Mass Flow KG/HR 2000 12823,71 2129,477 247,1647 12576,55 12576,55 627,7206 11948,83 11948,83 11642,29 306,5347 934,2553 5759,784 6694,039

Component Mass

Flow

ETHYLENE KG/HR 0 12299,55 21,57973 0 12299,55 12299,55 362,0619 11937,49 11937,49 11642,27 295,2183 657,2802 0 2,60E-35

ETHANOL KG/HR 0 31,08229 35,94262 0 31,08229 31,08229 3,11E+01 2,49E-97 2,49E-97 0 0 31,08229 0 2,71E-34

WATER KG/HR 1500 247,1647 1527,776 247,1647 0 0 0 0 0 0 0 0 0 1264,989

HYDROGEN KG/HR 0 8,695317 6,77E-06 0 8,695317 8,70E+00 5,94E-14 8,695317 8,70E+00 7,21E-07 8,695316 8,695316 0,00E+00 5,95E-04

PROPENE KG/HR 0 3,517747 0,018299 0 3,517747 3,517747 3,52E+00 4,96E-21 4,96E-21 0 0 3,517747 0 4,37E-36

ISOBUTYLENE KG/HR 0 185,4287 3,159874 0 185,4287 185,4287 1,85E+02 6,73E-37 6,73E-37 0 0 185,4287 0 4,23E-35

ETHANE KG/HR 0 21,012 0,044293 0 21,012 21,012 21,00915 0,002858 0,0028576 2,86E-03 8,58E-09 21,00915 0 1,04E-33

DIETHYL-ETHER KG/HR 0 0,776636 0,053753 0 0,776636 0,776636 7,77E-01 7,18E-59 7,18E-59 0 0 0,776636 0 5,54E-42

ACETALDEHYDE KG/HR 0 23,71615 0,961529 0 23,71615 23,71615 2,37E+01 2,44E-51 2,44E-51 0 0 23,71615 0 1,40E-33

CO KG/HR 0 0,003616 3,47E-07 0 0,003616 3,62E-03 1,13E-11 0,003616 3,62E-03 3,35E-06 0,00361278 0,003613 0,00E+00 2,14E-02

CO2 KG/HR 0 0,127987 0,000899 0 0,127987 0,127987 1,28E-01 5,79E-06 5,79E-06 5,79E-06 7,89E-12 0,127982 0 2832,458

METHANE KG/HR 0 2,639804 0,002421 0 2,639804 2,64E+00 1,56E-06 2,639802 2,639802 0,0223378 2,617464 2,617466 0 3,14E-25

H3O+ KG/HR 7,24E-15 0 2,57E-14 0 0 0 0 0 0 0 0 0 0 0

OH- KG/HR 212,614 0 160,3531 0 0 0 0 0 0 0 0 0 0 0

HCO3- KG/HR 0 0 0,00028 0 0 0 0 0 0 0 0 0 0 0

CO3-2 KG/HR 0 0 92,19792 0 0 0 0 0 0 0 0 0 0 0

NA+ KG/HR 287,386 0 287,386 0 0 0 0 0 0 0 0 0 0 0

OXYGE-01 KG/HR 0 0 0 0 0 0 0 0 0 0 0 0 5759,784 2596,57

Component Mass

Fraction

ETHYLENE 0 0,959125 0,010134 0 0,977975 0,977975 0,576788 0,999051 0,9990508 0,9999978 0,9630828 0,703534 0 3,88E-39

ETHANOL 0 0,002424 0,016879 0 0,002471 0,002471 0,049516 2,08E-101 2,08E-101 0 0 0,03327 0 4,04E-38

WATER 0,75 0,019274 0,717442 1 0 0 0,00E+00 0,00E+00 0 0 0 0 0 0,1889725

HYDROGEN 0 0,000678 3,18E-09 0 0,000691 0,000691 9,46E-17 0,000728 0,0007277 6,20E-11 0,0283665 0,009307 0 8,89E-08

PROPENE 0 0,000274 8,59E-06 0 0,00028 2,80E-04 0,005604 4,16E-25 4,16E-25 0 0 0,003765 0,00E+00 6,54E-40

ISOBUTYLENE 0 0,01446 1,48E-03 0 0,014744 0,014744 2,95E-01 5,63E-41 5,63E-41 0 0 0,198478 0 6,32E-39

ETHANE 0 0,001639 2,08E-05 0 0,001671 0,001671 3,35E-02 2,39E-07 2,39E-07 2,45E-07 2,80E-11 0,022488 0 1,55E-37

II-iii

DIETHYL-ETHER 0 6,06E-05 2,52E-05 0 6,18E-05 6,18E-05 1,24E-03 6,01E-63 6,01E-63 0,00E+00 0 0,000831 0 8,27E-46

ACETALDEHYDE 0 1,85E-03 4,52E-04 0 1,89E-03 0,001886 3,78E-02 2,04E-55 2,04E-55 0 0 0,025385 0 2,09E-37

CO 0 2,82E-07 1,63E-10 0 2,88E-07 2,88E-07 1,79E-14 3,03E-07 3,03E-07 2,88E-10 1,18E-05 3,87E-06 0 3,20E-06

CO2 0 9,98E-06 4,22E-07 0 1,02E-05 1,02E-05 2,04E-04 4,84E-10 4,84E-10 4,97E-10 2,57E-14 0,000137 0,00E+00 4,23E-01

METHANE 0 2,06E-04 1,14E-06 0 2,10E-04 0,00021 2,49E-09 2,21E-04 2,21E-04 1,92E-06 0,00853888 0,002802 0 4,69E-29

H3O+ 3,62E-18 0 1,21E-17 0 0 0 0 0 0 0 0 0 0 0

NA2CO3 0,00E+00 0 0,00E+00 0 0 0 0 0 0 0 0 0 0 0

OH- 0,106307 0 0,075302 0 0 0 0 0 0 0 0 0 0 0

HCO3- 0 0 1,31E-07 0 0 0 0 0 0 0 0 0 0 0

CO3-2 0 0 4,33E-02 0 0 0 0 0 0 0 0 0 0 0

NA+ 0,143693 0 0,134956 0 0 0 0 0 0 0 0 0 0 0

OXYGE-01 0 0 0 0 0 0 0 0 0 0 0 0 1 0,3878928

Mole Flow KMOL/HR 108,2645 461,9393 109,9003 13,71972 448,2196 448,2196 18,21959 430 430 415 15 33,21959 180 215,7242

Component Mole

Fraction

ETHYLENE 0 0,949103 0,006999 0 0,978154 0,978154 0,708359 0,989586 0,9895856 0,9999964 0,7015537 0,705286 0 4,29E-39

ETHANOL 0 0,001461 0,007099 0 0,001505 0,001505 0,037031 1,26E-101 1,26E-101 0 0 0,02031 0 2,72E-38

WATER 0,769067 0,0297 0,771649 1 0 0 0 0 0 0 0 0 0 0,325497

HYDROGEN 0 0,009338 3,06E-08 0 0,009623 0,009623 1,62E-15 0,010031 0,0100311 8,62E-10 0,2875606 0,129845 0 1,37E-06

PROPENE 0 0,000181 3,96E-06 0 0,000187 0,000187 4,59E-03 2,74E-25 2,74E-25 0 0 0,002516 0 4,82E-40

ISOBU-01 0 0,007154 0,000512 0 0,007373 0,007373 0,181392 2,79E-41 2,79E-41 0 0 0,099486 0 3,50E-39

ETHAN-01 0 0,001513 1,34E-05 0 0,001559 1,56E-03 0,038348 2,21E-07 2,21E-07 2,29E-07 1,90E-11 0,021032 0,00E+00 1,60E-37

DIETH-01 0 2,27E-05 6,60E-06 0 2,34E-05 2,34E-05 5,75E-04 2,25E-63 2,25E-63 0 0 0,000315 0 3,46E-46

ACETA-01 0 0,001165 0,000199 0 0,001201 0,001201 2,95E-02 1,29E-55 1,29E-55 0 0 0,016206 0 1,47E-37

CO 0 2,79E-07 1,13E-10 0 2,88E-07 2,88E-07 2,21E-14 3,00E-07 3,00E-07 2,88E-10 8,60E-06 3,88E-06 0 3,55E-06

CO2 0 6,30E-06 1,86E-07 0 6,49E-06 6,49E-06 1,60E-04 3,06E-10 3,06E-10 3,17E-10 1,19E-14 8,75E-05 0 0,2983426

METHA-01 0 0,000356 1,37E-06 0 0,000367 0,000367 5,35E-09 3,83E-04 0,0003827 3,36E-06 0,010877 0,004911 0 9,06E-29

DOWTH-01 0 0,00E+00 0,00E+00 0 0,00E+00 0,00E+00 0,00E+00 0,00E+00 0,00E+00 0,00E+00 0,00E+00 0 0,00E+00 0,00E+00

H3O+ 3,52E-18 0,00E+00 1,23E-17 0 0,00E+00 0 0,00E+00 0,00E+00 0,00E+00 0,00E+00 0,00E+00 0 0 0

NA2CO3 0 0 0,00E+00 0 0 0,00E+00 0 0 0,00E+00 0 0 0 0 0

OH- 0,115466 0 0,085788 0 0 0 0 0 0 0 0 0 0 0

HCO3- 0,00E+00 0 4,17E-08 0 0 0 0 0 0 0 0 0 0 0

CO3-2 0 0 0,01398 0 0 0 0 0 0 0 0 0 0 0

NA+ 0,115466 0 0,113748 0 0 0 0 0 0 0 0 0 0 0

NAOH 0 0 0,00E+00 0 0 0 0 0 0 0 0 0 0 0

OXYGE-01 0 0 0 0 0 0 0 0 0 0 0 0 1 0,3761555

Volume Flow m^3/HR 1,573762 11702,71 1,709154 0,250685 11383,52 2897,645 0,999688 6042,049 80,38169 20,48199 203,8151 217,3445 4387,223 27359,74

III-i

Appendix III – Reaction specifications

Reactor The composition of the stream after the reactor is based on data from Chematur [19]. To

obtain this composition reactions have been set up. Two Aspen reactors, RStoic and REquil,

are used to simulate the real process reactor. Reactions specified in the RStoic reactor and the

selectivities for these are displayed as reaction 1-7 in Table 11. A REquil reactor is used to

balance the carbon dioxide and the carbon monoxide, in order to make it coincide with

Chematur data. The reaction for this reactor is shown as reaction 8 in Table 11.

Table 11. Reactions specified in Aspen reactors.

# Reaction Selectivity (%)

1 𝐶2𝐻5𝑂𝐻 → 𝐶2𝐻4 + 𝐻2𝑂 98.0

2 2 𝐶2𝐻5𝑂𝐻 → (𝐶2𝐻5)2𝑂 + 𝐻2𝑂 0.005

3 2 𝐶2𝐻5𝑂𝐻 → 𝐶4𝐻8 + 2 𝐻2𝑂 1.5

4 𝐶2𝐻5𝑂𝐻 → 𝐶2𝐻4𝑂 + 𝐻2 0.125

5 2 𝐶2𝐻5𝑂𝐻 → 𝐶3𝐻6 + 𝐶𝑂2 + 3 𝐻2 0.0375

6 2 𝐶2𝐻5𝑂𝐻 → 𝐶2𝐻6 + 2 𝐶𝑂 + 3 𝐻2 0.3125

7 2 𝐶2𝐻5𝑂𝐻 → 3 𝐶𝐻4 + 𝐶𝑂2 0.02

8 𝐶𝑂 + 𝐻2𝑂 ↔ 𝐶𝑂2 + 𝐻2 equilibrium

Sodium hydroxide scrubber The absorption of CO2 to the sodium hydroxide liquid stream is according to the reactions in

Table 12 [50]. These are implemented in Aspen.

Table 12. Reactions specified in Aspen absorption tower.

# Reaction

1 2 𝐻2𝑂 ↔ 𝐻3𝑂+ + 𝑂𝐻−

2 𝐻𝐶𝑂3− + 𝐻2𝑂 ↔ 𝐻3𝑂+ + 𝐶𝑂3

2−

3 𝐶𝑂2 + 𝑂𝐻− → 𝐻𝐶𝑂3−

4 𝐻𝐶𝑂3 → 𝐶𝑂2 + 𝑂𝐻−

IV-i

Appendix IV – Design equations and assumptions

The equations and assumptions used in the dimensioning calculations for the different unit

operations are presented below.

Reactor The tubes in the reactor were set to be 3 m long and have a diameter of 3 cm. By assuming an

ethanol space velocity of 6000 GHSV [40], the minimum amount of catalyst required could

be calculated:

𝑉𝑐𝑎𝑡 =𝐹𝐸𝑡𝑂𝐻

𝐺𝐻𝑆𝑉 Equation 4

Where the volume flow could be calculated as:

𝐹𝐸𝑡𝑂𝐻 =�̇�𝐸𝑡𝑂𝐻

𝑀𝐸𝑡𝑂𝐻∙ 𝜌𝐸𝑡𝑂𝐻

1 𝑏𝑎𝑟,0 ℃ Equation 5

Where �̇�𝐸𝑡𝑂𝐻 is the mass flow of ethanol, 𝑀𝐸𝑡𝑂𝐻 is the molar mass of ethanol and 𝜌𝐸𝑡𝑂𝐻1 𝑏𝑎𝑟,0 ℃

is the molar density of ethanol at standard conditions. An extra 20 % of catalyst was added to

ensure that as much ethanol as possible will react. By diving the total catalyst volume, Vcat, by

the volume of one tube, Vtube, the number of tubes was calculated:

𝑛𝑡𝑢𝑏𝑒𝑠 =𝑉𝑐𝑎𝑡

𝑉𝑡𝑢𝑏𝑒 Equation 6

With the total number of tubes the cross sectional area of all the tubes was calculated with

Equation 7:

𝐴𝑡𝑢𝑏𝑒𝑠 = 𝑛𝑡𝑢𝑏𝑒𝑠 ∙ 𝜋𝑟𝑡𝑢𝑏𝑒2 Equation 7

Where rtube is the radius of one tube. To be able to design the reactor, the packing of the tubes

inside the reactor had to be assumed. This packing factor, 𝑝𝑓𝑡𝑢𝑏𝑒, of the tubes was set to 0.25.

The total reactor cross sectional area was calculated by using Equation 8 and the diameter of

the reactor could then be calculated with Equation 9.

𝐴𝑟𝑒𝑎𝑐𝑡𝑜𝑟 =𝑛𝑡𝑢𝑏𝑒𝑠∙𝐴𝑡𝑢𝑏𝑒𝑠

𝑝𝑓𝑡𝑢𝑏𝑒 Equation 8

𝑑𝑟𝑒𝑎𝑐𝑡𝑜𝑟 = 2 ∙ √𝐴𝑟𝑒𝑎𝑐𝑡𝑜𝑟

𝜋 Equation 9

When running at industrial scale, a guard bed is usually added to adsorb impurities and thus

the prolonging the reactor catalyst lifetime. Consequently, a guard bed was added and

assumed to be 20 % of the total reactor catalyst volume:

𝑉𝑔𝑢𝑎𝑟𝑑 𝑏𝑒𝑑 = 0.2 ∙ 𝑉𝑐𝑎𝑡 Equation 10

The height of the guard bed was calculated with Equation 11. A total height of the reactor was

then calculated with Equation 12. To the total height 0.5 m was added at the top and the

bottom of the reactor.

ℎ𝑔𝑢𝑎𝑟𝑑 𝑏𝑒𝑑 =𝑉𝑔𝑢𝑎𝑟𝑑 𝑏𝑒𝑑

𝐴𝑡𝑢𝑏𝑒𝑠 Equation 11

ℎ𝑟𝑒𝑎𝑐𝑡𝑜𝑟 = ℎ𝑡𝑢𝑏𝑒𝑠 + ℎ𝑔𝑢𝑎𝑟𝑑 𝑏𝑒𝑑 + 0.5 ∙ 2 Equation 12

The heat transfer from the heating fluid could be calculated with Equation 13 to ensure that

the supplied heat were enough compared to the reactor requirement.

𝑄 = 𝑘 ∙ 𝐴𝑠𝑢𝑟𝑓𝑎𝑐𝑒 𝑡𝑢𝑏𝑒𝑠 ∙ ∆𝑇 Equation 13

IV-ii

𝑘 is the overall heat transfer coefficient, 𝐴𝑠𝑢𝑟𝑓𝑎𝑐𝑒 𝑡𝑢𝑏𝑒𝑠 is the total surface area of the tubes

and ∆𝑇 is the temperature difference between inside and outside of the tubes. The heat

transfer coefficient was assumed to be 500 𝑊

𝑚2∙𝐾, [56]. When comparing the heat load required

for the reaction, given from Aspen simulations, with the calculated available heat from the

heating fluid the available heat load was sufficient.

Distillation columns The design of the distillation columns were obtained by calculating the diameter and the

height of the columns. A maximum allowable vapor velocity, uv, was estimated from the plate

spacing, lt, with Equation 14:

𝑢𝑣 = (−0.17𝑙𝑡2 + 0.27𝑙𝑡 − 0.047) ∙ (

𝜌𝑙−𝜌𝑣

𝜌𝑣)

0.5 Equation 14

Where ρl is the liquid density and ρv is the vapor density. The plate spacing was assumed to

be 0.6 m. An average vapor density was estimated both at the top and the bottom of the

column and hence two maximal velocities were obtained. These two velocities were used in

Equation 15 with its vapor density respectively estimating the two column diameters:

𝐷𝑐 = √4𝑉𝑤

𝜋𝜌𝑣𝑢𝑣 Equation 15

Where Dc is the column diameter and Vw is the vapor mass flow in the column. The largest

diameter calculated with Equation 15 was then used for the column design. By using Equation

16 and 17 the column height could be calculated.

ℎ𝑐 = 𝑁𝑅 ∙ 𝑙𝑡 Equation 16

𝐸0 =𝑁𝐼

𝑁𝑅 Equation 17

Where hc is the column height, NR is the number of real stages, NI is the number of ideal

stages and E0 is the column efficiency. The column efficiency was set to 0.8 for all the

distillation towers. Both the condenser and the reboiler were seen as heat exchangers and

designed thereafter [57].

Heat exchangers The required heat exchanger area, A, was calculated with Equation 18, where Q is the effect,

k is the overall heat transfer coefficient and TL is the logarithmic average temperature,

calculated with Equation 19. The numbers 1 and 2 in Equation 19 denotes different sides of

the heat exchanger. Equation 20 was used to calculate the amount of needed cooling media of

the condenser [56].

𝐴 =𝑄

𝑘∙𝑇𝐿 Equation 18

𝑇𝐿 =(𝑇2,𝑖𝑛−𝑇1,𝑜𝑢𝑡)−(𝑇2,𝑜𝑢𝑡−𝑇1,𝑖𝑛)

ln((𝑇2,𝑖𝑛−𝑇1,𝑜𝑢𝑡)

(𝑇2,𝑜𝑢𝑡−𝑇1,𝑖𝑛))

Equation 19

�̇�𝑐𝑜𝑜𝑙𝑖𝑛𝑔 =𝑄

𝐶𝑝∙∆𝑇 Equation 20

Gas/liquid separator The height and diameter of the vessel used for separating gas from liquid were calculated. The

gas velocity (uG) was determined by using the vapor and liquid densities (ρG and ρL):

IV-iii

𝑢𝐺 = 0.064√𝜌𝐿

𝜌𝐺− 1 Equation 21

From the gas velocity and the gas flow rate (QG) the diameter (D) of the vessel could be

calculated:

𝐷 = 2√𝑄𝐺

𝑢𝐺𝜋 Equation 22

The liquid height (hL) was calculated based on an assumed residence time of the liquid

fraction (θ) of 600 seconds. QL denotes the liquid flow from the vessel.

ℎ𝐿 =𝑢𝐺𝑄𝐿𝜃

𝑄𝐺 Equation 23

The total height of the vessel (h) was calculated as the sum of the gas height (hg) and the

liquid height. The vessel was designed so that hg was equal to the column diameter D or, if the

column diameter was less than 1 m, was set to 1 m [58].

ℎ = ℎ𝐺 + ℎ𝐿 , 𝑤ℎ𝑒𝑟𝑒 ℎ𝐺 = max {1 𝑚, 𝐷} Equation 24

Absorption column The absorption column was designed in the same manner as the distillations columns, but

without condenser and reboiler at top and bottom. [57]

Storage tanks The storage tanks were designed to store ethanol, ethylene and sodium hydroxide in 10 days

respectively. The volume of the tanks (V) was calculated from the flows (F) and the amount

of storage days (θ):

𝑉 = 𝐹𝜃 Equation 25

Dryer The breakthrough time could be calculated with Equation 26 [59], in which V is total mass

flow, c0 is concentration of water, S is adsorbent mass and X is an adsorption factor.

𝑡𝑏𝑟𝑒𝑎𝑘𝑡ℎ𝑟𝑜𝑢𝑔ℎ =𝑆∙𝑋

𝑉∙𝑐0 Equation 26

Assumptions that were made was that 21 g water can be adsorbed by 100 g zeolite, 3A

Molecular Sieve. A zeolite from Interra Global was used. This zeolite has a bulk density of ca

0.70 g/ml [60].

V-i

Appendix V – Calculation of unit operation design

The calculations of the unit operations were made in excel, see specifications below.

Reactor

Table 13. Reactor parameters.

Parameters Unit Source/Comment

Residence time in reactor 6,000 h-1

[40]

Mass flow EtOH 201,016,000 kg/year Aspen Plus V8.2

Density EtOH 240 °C 135.5 kg/m3

Molar gas volume EtOH 0 °C 1

bar

22.4 l/mol Assumption ideal gas

Number of operating days 365 days/year Assumption by authors

Molar mass EtOH 0.0461 kg/mol [61]

ΔTL (condensing oil-catalyst) 48 °C Aspen Plus V8.2

Overall heat transfer coefficient, k 500 W/(m2·°C) Assumption by authors

Table 14. Flows and assumed dimensions of reactor.

Calculations flows Unit

Mass flow EtOH 22,931 kg/h

Volumetric flow

EtOH

169.23 m3/h

Normal volumetric

flow EtOH

11,150 Nm3/h

Assumed dimensions

High inner tubes 3 m

Diameter tubes 0.03 m

Table 15. Calculations of tube bundle inside reactor.

Calculations tubes Unit

Volume catalyst 1.858 m3

Extra catalyst (marginal) 0.3717 m3

Catalyst volume reactor 2.2230 m3

Guardbed (catalyst) volume 0.4460 m3

Volume/tube 0.0021 m3

Total needed number of tubes 1,052 tubes

Cross sectional area all tubes 0.7433 m2

Surface area all tubes 297.3 m2

Guard bed height 0.15 m

V-ii

Possible heat transfer tubes, Q 7134 kW

Table 16. Calculations of reactor vessel.

Distillation 1 Distillation tower 1. Water/Ethanol

Table 17. Calculations for distillation tower 1.

Parameters Unit Source/Comment

Pressure column 100,000 Pa Aspen Plus V8.2

Molar mass ethanol 0.046 kg/mol [61]

Molar mass water 0.018 kg/mol [61]

Gas law constant 8.315 J/mol∙K [61]

Liquid density 967.8 kg/m3

Aspen Plus V8.2

Tray distance 0.6 m Assumption by authors

Mass flow rate vapor 0.259 kg/s Aspen Plus V8.2

Ideal number of stages 25 Aspen Plus V8.2

Column efficiency. E0 0.8 Assumption by authors

Calculations Top Bottom

Volumetric fraction ethanol 0.9 0.01 m3/ m

3 Aspen Plus V8.2

Temperature column 351.2 373.2 K Aspen Plus V8.2

Density water vapor 0.617 0.581 kg/ m3 Assumption ideal gas law

Density ethanol vapor 1.578 1.485 kg/ m3 Assumption ideal gas law

Vapor density 1.482 0.590 kg/ m3 Choose average

Maximal vapor velocity 1.374 2.179 m/s Aspen Plus V8.2

Diameter column 0.402 0.506 m

Continues with the largest

diameter

Real number of stages 31.25

Calculations reactor vessel Unit

Total height 4,15 m

Cross sectional area whole

reactor

2.97 m2

Inside diameter whole reactor 1.95 m

Total volume whole reactor 12.3 m3

Catalyst density 700 kg/ m3

Total amount catalyst needed 1,873 kg

V-iii

Column height 18.75 m

Column surface area 29.82 m2

Table 18. Calculations for the condensor in distillation tower 1.

Condenser

Parameters Unit Source/Comment

Temp. Cooling

media in

20 °C Assumption by authors

Temp. Cooling

media out

40 °C Assumption by authors

Temp. Vapor (top) 76 °C Aspen Plus V8.2

Heat transfer coeff.

(k)

312.3 W/(m2·°C) Assumption by authors, made on chilled water [56]

Heat transferred, Q 1,409 kW Aspen Plus V8.2

log. mean temp 45.27 °C

Area 99.7 m2

Mass rate cooling

water

16.9 kg/s

Table 19. Calculations for the reboiler in distillation tower 1.

Reboiler

Parameters Unit Source/Comment

Temp. heat. Media in 280 °C Aspen Plus V8.2

Mass flow oil 38.89 kg/s Aspen Plus V8.2

Temp. heat. Media out 248.7 °C Aspen Plus V8.2

Temp. Liquid 100 °C Aspen Plus V8.2

Heat transfer coeff. (k) 85 W/(m2·°C) Assumption by authors, made on oil [56]

Heat transferred, Q 2,245 kW Aspen Plus V8.2

log. mean temp 163.9 °C

Area 161.2 m2

Ethylene column Distillation tower ethylene/higher hydrocarbons

Table 20. Calculations for the ethylene column.

Parameters Unit Source/Comment

V-iv

Pressure column 100,000 Pa

Molar mass ethylene 0.028 kg/mol Aspen Plus V8.2

Molar mass butene 0.056 kg/mol Aspen Plus V8.2

Gas law constant 8.315 J/mol∙K [61]

Liquid density 627.9 kg/m3

Aspen Plus V8.2

Tray distance 0.6 m Assumption by authors

Mass flow rate 3.319 kg/s Aspen Plus V8.2

Ideal number of stages 15 Aspen Plus V8.2

Column efficiency. E0 0.8 Assumption by authors

Calculations Top Bottom

Volumetric fraction ethylene 1 0.79 m3/m

3 Aspen Plus V8.2

Temperature column 169 174 K Aspen Plus V8.2

Density ethylene vapor 1.996 1.939 kg/m3 Assuming ideal gas

Density butene vapor 3.993 3.878 kg/m3 Assuming ideal gas

Vapor density 1.996 2.346 kg/m3 Assuming ideal gas

Maximal vapor velocity 0.953 0.878 m/s

Diameter column 1.491 1.432 m

Continues with the

largest diameter

Real number of stages 18.75

Column height 11.25 m

Column surface area 52.69 m2

Table 21. Calculations for the condensor in the ethylene column.

Condenser

Parameters Unit Source/Comment

Temp. Cooling media -150 °C Assumption by authors

Temp. Vapor (top) -104 °C Aspen Plus V8.2

Heat transfer coeff. (k) 114 W/(m2·°C) Assumption by authors, made on ethylene

liquid-ethylene vapor [56]

Heat transferred, Q 1,135 kW Aspen Plus V8.2

log. mean temp 46 °C

Area 216.4 m2

V-v

Table 22. Calculations for the reboiler in the ethylene column.

Reboiler

Parameters Unit Source/Comment

Temp. heat. Media in 245 °C Assumption by authors

Mass flow oil 38.9 kg/s Aspen Plus V8.2

Temp. heat. Media out 218.8 °C Aspen Plus V8.2

Temp. Liquid -98.33 °C Aspen Plus V8.2

Heat transfer coeff.

(k)

114 W/(m2·°C) Assumption by authors, made on ethylene

liquid-ethylene vapor [56]

Heat transferred. Q 1,884 kW Aspen Plus V8.2

log. mean temp 330.0 °C

Area 50.07 m2

C2 Stripper Distillation tower ethylene/light gases.

Table 23. Calculations for the C2 stripper column.

Parameters Unit Source/Comment

Pressure column 100,000 Pa

Molar mass ethylene 0.028 kg/mol Aspen Plus V8.2

Molar mass hydrogen 0.0020 kg/mol Aspen Plus V8.2

Gas law constant 8.315 J/mol∙K [61]

Liquid density 148.7 kg/m3 Aspen Plus V8.2

Tray distance 0.6 m Assumption by authors

Mass flow rate 3.319 kg/s Aspen Plus V8.2

Ideal number of stages 10 Aspen Plus V8.2

Column efficiency. E0 0.8 Assumption by authors

Calculations Top Bottom Source/Comment

Volumetric fraction ethylene 0.7 1 m3/m

3 Aspen Plus V8.2

Temperature column 163 169 K Aspen Plus V8.2

Density ethylene vapor 2.070 1.996 kg/m3 Assuming ideal gas

Density hydrogen vapor 0.149 0.143 kg/m3 Assuming ideal gas

Vapor density 1.493 1.996 kg/m3 Assuming ideal gas

V-vi

Maximal vapor velocity 0.534 0.461 m/s

Diameter column 2.302 2.142 m

Continues with the

largest diameter

Real number of stages 12.5

Column height 7.5 m

Column surface area 54.23 m2

Table 24. Calculations for the condensor in the C2 stripper column.

Condenser

Parameters Unit Source/Comment

Temp. Cooling media -150 °C Assumption by authors

Temp. Vapor (top) -110 °C Aspen Plus V8.2

Heat transfer coeff. (k) 114 W/(m2·°C) Assumption by authors, made on ethylene

liquid-ethylene vapor [56]

Heat transferred, Q 57 kW Aspen Plus V8.2

log. mean temp 40 °C

Area 12.5 m2

Table 25. Calculations for the reboiler in the ethylene column.

Reboiler

Parameters Unit Source/Comment

Temp. heat. Media in 215 °C Assumption by authors

Temp. heat. Media out 210.64 Aspen Plus V8.2

Mass flow oil 38.89 kg/s Aspen Plus V8.2

Temp. Liquid -104 °C Aspen Plus V8.2

Heat transfer coeff. (k) 114 W/(m2·°C) Assumption by authors, based on

ethylene liquid-ethylene vapor [56]

Heat transferred, Q 313 kW Aspen Plus V8.2

log. mean temp 316.8 °C

Area 8.67 m2

V-vii

Heat Exchanger 1 Ethanol feed and oil

Table 26. Calculations for heat exchanger 1.

Parameter Unit Source/Comment

Ethanol feed. in 10 °C Aspen Plus V8.2

Ethanol feed. out 80 °C Aspen Plus V8.2

Stream from reactor

(ethylene) in

240 °C Aspen Plus V8.2

Stream from reactor

(cooled ethylene) out

70 °C Aspen Plus V8.2

Temp. transfer efficiency 0.739

Calculations

Heat transfer. Q 5,808 kW Aspen Plus V8.2

Overall heat transfer

coefficient (k)

341 Assumption by authors, based on [56]

Log. mean temp. 102.0

Area 167.0 m2

Heat Exchanger 2 Ethanol preheating.

Table 27. Calculations for heat exchanger 2.

Parameter Unit Source/Comment

Cold stream in (ethanol) 81 °C Aspen Plus V8.2

Cold stream out (ethanol) 240 °C Aspen Plus V8.2

Hot stream in (dowtherm A) 288 °C Aspen Plus V8.2

Hot stream out (dowtherm A) 288 °C Aspen Plus V8.2

Calculations

Heat transfer, Q 3,088 kW Aspen Plus V8.2

Overall heat transfer

coefficinet (k)

100 Assumption by authors, based on [56]

Log. mean temp. 108.8

Area 283.8 m2

V-viii

Heat Exchanger 3 Heat exchanger connected to quench column.

Table 28. Calculations for heat exchanger 3.

Parameters Unit Source/Comment

Temp. Ethylene in 70 °C Aspen Plus V8.2

Temp. Ethylene out 25 °C Aspen Plus V8.2

Temp. Cooling media

in

15 °C Water, assumption by authors

Temp. Cooling media

out

30 °C Water, assumption by authors

Calculations

Heat transfer coeff (k) 350 W/(m2·°C) Assumption by authors, based on

ethylene vapor-chilled water [56]

Heat transferred, Q 3,097 kW Aspen Plus V8.2

log.Temp. 21.6 °C

Area 408.9 m2

Mass flow cooling

water

49.40 kg/s

Heat Exchanger 4 Cooling of stream before ethylene column.

Table 29. Calculations for heat exchanger 4.

Parameters Unit Source/Comment

Temp. Ethylene in 32 °C Aspen Plus V8.2

Temp. Ethylene out -104 °C Aspen Plus V8.2

Temp. Cooling media

in

-150 °C Assumption by authors

Temp. Cooling media

out

-150 °C Assumption by authors

Calculations Assumption by authors, based on

ethylene vapor-ethylene vapor [56]

Heat transfer coeff (k) 114 W/(m2·°C) Aspen Plus V8.2

Heat transferred, Q 697 kW

log.Temp. 98.88 °C

Area 61.83 m2

V-ix

Heat Exchanger 5 Cooling of stream before C2 Stripper column.

Table 30. Calculations for heat exchanger 5.

Parameters Unit Source/Comment

Temp. Ethylene in -104 °C Aspen Plus V8.2

Temp. Ethylene out -130 °C Aspen Plus V8.2

Temp. Coooling

media in

-150 °C Assumption by authors

Temp. Cooling

media out

-150 °C Assumption by authors

Calculations

Heat transfer coeff

(k)

114 W/(m2·°C) Assumption by authors, based on ethylene

vapor-ethylene vapor [56]

Heat transferred, Q 1,819 kW Aspen Plus V8.2

log.Temp. 31.22 °C

Area 511.2 m2

Gas/Liquid separator Separates gas from liquid after heat exchanger 3.

Table 31. Calculations for the gas/liquid separator.

Parameters Unit Source/Comment

Liquid density (bottom) 954.6 kg/

m3

Aspen Plus V8.2

Vapor density (top) 1.123 kg/

m3

Aspen Plus V8.2

Volumetric flow rate gas 3.155 m3/s Aspen Plus V8.2

Volumetric flow rate

liquid

0.0029 m3/s Aspen Plus V8.2

Residence time 600 s Assumption by authors.

Calculations

Gas velocity 1.865 m/s

Diameter flash column 1.468 m

Height gas 1.468 m

Choose either height=1

or if D>1, choose height of gas [58].

V-x

Height liquid 1.050 m

Height column 2.51749 m

Absorption column

Table 32. Calculations for the absorption column.

Parameters Unit Source/Comment

Pressure column 100,000 Pa

Molar mass ethylene 0.028 kg/mol Aspen Plus V8.2

Molar mass water 0.018 kg/mol Aspen Plus V8.2

Gas law constant 8.315 J/mol∙K [61]

Liquid density 1,246 kg/m3 Aspen Plus V8.2

Tray distance 0.6 m Assumption by authors

Mass flow rate vapor 3.562 kg/s Aspen Plus V8.2

Ideal number of stages 35 Aspen Plus V8.2

Column efficiency. E0 0.8 Assumption by authors

Calculations Top Bottom

Volumetric fraction

ethylene

0.97 0.98 m3/m

3 Aspen Plus V8.2

Temperature column 306.4 302.8 K Aspen Plus V8.2

Density ethylene vapor 1.101 1.114 kg/m3 Assumption ideal gas

Density water vapor 0.707 0.716 kg/m3 Assumption ideal gas

Vapor density 1.089 1.106 kg/m3 Assumption ideal gas

Maximal vapor velocity 1.819 1.805 m/s

Diameter column 1.513 1.507 m

Continues with the largest

diameter

Real number of stages 43.75

Column height 26.25 m

Column surface area 124.8 m2

Storage tanks Ethanol tank, liquid.

Table 33. Calculations for the ethanol storage tank.

Parameter Unit

V-xi

Flow ethanol 27.09 m3/h

Days of storage 10 days

Volume 6,502 m3

NaOH tank, liquid.

Table 34. Calculations for the soduim hydroxide storage tank

Parameter Unit

Flow NaOH 1.57 m3/h

Days of storage 10 days

Volume 377.8 m3

Ethylene tank, liquid/gas.

Table 35. Calculations for the ethylene storage tank

Parameter Unit

Flow ethylene 20.48 m3/h

Days of storage 10 days

Volume 4,916 m3

Dryer Two pieces with regeneration.

Table 36. Calculations for the two dryers.

Parameters Unit Source/Comment

Adsorption factor 0.21 kg water/kg

adsorbent

[60]

Concentration water 0.019 kg water/kg mass

flow

Aspen Plus V8.2

Total mass flow 12,824 kg/h Aspen Plus V8.2

Drying time 12 h Assumption by authors

Density adsorbent 700 kg/m3 [60]

Calculations Unit

Adsorbent mass 14,143 kg adsorbent

Volume adsorbent 20.20 m3 adsorbent

Extra adsorbent 4.040 m3 adsorbent

Total volume

adsorbent

24.24 m3 adsorbent Assuming cylindrical geometry

Total mass adsorbent 16,971 kg adsorbent

V-xii

Diameter dryer 2.5 m

Height dryer 4.94 m

Total height 5.94 m

Centrifugal pumps Pump 1.

Table 37. Calculations for centrifugal pump 1.

Parameter Unit Source/Comment

Inlet pressure 1 bar

Outlet pressure 10 bar

Mass flow 22,000 kg/h Ethanol/water

Heat transfer, Q 12 kW Aspen Plus V8.2

Pump 2.

Table 38. Calculations for centrifugal pump 2.

Parameter Unit Source/Comment

Inlet pressure 1 bar

Outlet pressure 10 bar

Mass flow 2,000 kg/h NaOH/water

Heat transfer, Q 3 kW Aspen Plus V8.2

Wastewater treatment plant Table 39. Calculations for the wastewater treatment plant.

Wastewater from Unit

Absorber 1.71 m3/h

0.000475 m3/s

EtOH distillation 9.91 m3/h

0.002753 m3/s

Total wastewater capacity 0.003228 m3/s

VI-i

Appendix VI – Economy equations and method

Ulrich Method A capital cost estimation method called Ulrich method was used for economical calculations.

Ulrich method is based on a calculation where the Bare-Module Cost, CBM, for each utility is

calculated by multiplying the Purchased Equipment Cost, Cp, with a Module Factor, FBMα, see

Equation 27 [52].

𝐶𝐵𝑀 = 𝐶𝑝 ∙ 𝐹𝐵𝑀𝛼 Equation 27

The sum of all Bare-Module Costs are then calculated and multiplied with two factors, one

taking fees and contingencies, ffees/contingency, into account and one taking auxiliary facilities,

fauxiliary facilities, into account, Equation 28 [52]. The cost K is direct and indirect plant costs as

well as auxiliary facilities costs.

𝐾 = (∑ 𝐶𝐵𝑀𝑖

𝑛𝑖=1 ) ∙ 𝑓𝑓𝑒𝑒𝑠/𝑐𝑜𝑛𝑡𝑖𝑛𝑔𝑒𝑛𝑐𝑦 ∙ 𝑓𝑎𝑢𝑥𝑖𝑙𝑖𝑎𝑟𝑦 𝑓𝑎𝑐𝑖𝑙𝑖𝑡𝑖𝑒𝑠 Equation 28

Purchased Equipment Costs and Module Factors were taken from diagrams from January

2004 in which the costs were given in USD. A factor, see Equation 29, was used for updating

of the utilities cost, see Equation 30. In Equation 29 IAK stands for Equipment Construction

Cost Index and applies to US. X and Y in Equation 29 and 30 denotes which year the values

are related to [52].

𝑓 =(𝐼𝐴𝐾)𝑥

(𝐼𝐴𝐾)𝑌 Equation 29

𝐾$,𝑋 = 𝑓 ∙ 𝐾$,𝑌 Equation 30

Fixed costs The annual cost for storage of feedstock and product was calculated with Equation 31, in

which Q is annual consumption or production, P is cost or income per tonne and D is days of

storage. fA is the annuity factor, which can be calculated with Equation 32, in which X is life

time and X is the interest rate. The factor can also be used to depreciate the grass root capital

cost [62].

𝐴𝑛𝑛𝑢𝑎𝑙 𝑐𝑜𝑠𝑡 𝑓𝑜𝑟 𝑠𝑡𝑜𝑟𝑎𝑔𝑒 = 𝑄 ∙ 𝑓𝐴 ∙ 𝑃 ∙ 𝐷/(365 𝑑𝑎𝑦𝑠/𝑦𝑒𝑎𝑟) Equation 31

𝑓𝐴 =𝑋

1−(1+𝑋)−𝑁 Equation 32

Annual maintenance and repair cost were estimated to amount to 6 % of Grass Roots Capital.

The cost of spare parts was calculated to be 15 % of the cost for maintenance and repair times

the factor fA calculated with Equation 32.

Direct costs The annual cost for operators was calculated with Equation 33. S stands for direct monthly

salary and n is number of workers per shift [62].

𝐴𝑛𝑛𝑢𝑎𝑙 𝑐𝑜𝑠𝑡 𝑓𝑜𝑟 𝑜𝑝𝑒𝑟𝑎𝑡𝑜𝑟𝑠 = 12 𝑚𝑜𝑛𝑡ℎ𝑠/𝑦𝑒𝑎𝑟 ∙ 𝑆 ∙ 𝑛 ∙ 5 Equation 33

The cost for supervisors for operators and laboratory work was 15 % on operators each.

Consumables are also included as direct costs. The cost for land and license fees was

neglected.

VI-ii

Indirect costs To calculate the overhead for staff 70 % was added on the cost for shift personnel and 50 %

was added on the cost for day personnel. Administration costs were estimated to be 25 % of

the overhead costs for staff. Costs for research, development, distribution and sale were

neglected.

VII-i

Appendix VII – Grass Root Capital

Purchased cost for equipment and Bare-Module Cost have been retrieved from the online

database EconExpert, using Ulrich’s calculation method [63].

Table 40. All units and their pruchased cost for equipment and Bare-Module Cost.

Unit Dimensions/Operating

conditions

Material Cp

(Purchased

cost, USD)

CBM

(bare module

cost, USD) Ethanol storage

tank

6,502 m3 Atmosperic

pressure- Cone roof

Stainless steel 154,899 542,146

Sodium hydroxide

storage tank

378 m3 Atmosperic

pressure- Cone roof

Stainless steel 24,339 85,188

Ethylene storage

tank

4,916 m3 Atmospheric

press-Gas holder

Stainless steel 224,877 787,069

Centrifugal pump

1

Pressure rise 10 bar

12 kW, Centrifugal pump

Stainless steel 9,348 45,111

Heat exchanger 1 168 m2 Shell and Tube,

Fixed tube sheet and U-

tube

Stainless steel 16,731 96,749

Heat exchanger 2 284 m2 Shell and Tube,

Fixed tube sheet and U-

tube

Stainless steel 23,814 137,710

Reactor vessel

D=1.94 m, H=4.15 m,

Process vessel, vertically,

no packing

A=297.3 m2, Shell and

tube, fixed tube

Stainless steel

Stainless steel

20,209

24,585

190,545

142,168

Thermal fluid

heater

10,158 kW

Furnaces, Thermal fluid

heater, Mineral oil heaters

- 534,216 1,175,274

Heat exchanger 3 409 m2

Shell and Tube, Fixed

tube sheet and U-tube

Stainless steel 30,845 178,368

Gas/Liquid

separator

D=1.47 m, H=2.52 m,

Process vessel, Vertically

oriented, No packing or

trays

Stainless steel 12,161 114,666

Centrifugal pump

2

Pressure rise 10 bar, 3kW,

Centrifugal

Stainless steel 5,477 26,433

Absorber D=1.52 m, H=26.3 m

43 Trays, Process vessel,

Vertically oriented,

Sieve-trays

Stainless steel and

stainless steel trays

78,444 878,861

Dryer - 2 st

D=2.5 m, H=5.94 m,

Packed height 4.94 m

Process vessel, vertically

Stainless steel

Zeolite adsorbent

66,658 868,824

Heat exchanger 4 62 m2

Shell and Tube, Fixed tub

Stainless steel 9,113 52,700

VII-ii

Ethylene column

D=1.49 m, H=11.25 m

12 Trays, Process equip,

vertically, sieve trays

Condensor: Heat

exchanger, A= 216m2

Shell and Tube, Fixed

tube sheet and U-tube

Reboiler: Heat exchanger,

A= 50m2

Shell and Tube, Fixed

tube sheet and U-tube

Stainless steel and

stainless steel trays

Stainless steel

Stainless steel

34,666

19,755

8,080

372,576

114,235

46,727

Heat exchanger 5 511 m2 Shell and Tube,

Fixed tub

Stainless steel 36,318 210,016

C2 stripper

D=2.3 m, H=7.5 m

13 Trays, Process vessel,

vertically, sieve trays

Condensor: Heat

exchanger, A= 12.5m2

Shell and Tube, Fixed

tube sheet and U-tube

Reboiler: Heat exchanger,

A= 8.7m2

Shell and Tube, Fixed

tube sheet and U-tube

Stainless steel and

stainless steel trays

Stainless steel

Stainless steel

35,937

4,080

3,503

453,142

23,593

20,258

Ethanol

distillation tower

D=0.506 m, H=18.75 m

19 Trays, Process vessel,

vertically, sieve trays

Condensor: Heat

exchanger, A= 100 m2

Shell and Tube, Fixed

tube sheet and U-tube

Reboiler: Heat exchanger,

A= 161 m2

Shell and Tube, Fixed

tube sheet and U-tube

Stainless steel

Stainless steel

Stainless steel

36,470

12,071

16,275

350,919

69,803

94,113

Heatpump Heat absorption

rate=3,708 kW

Auxiliary, Mechanical ref

unit

Coolant temp:

- 55°C

- 6,896,134

Wastewater

treatment plant

Capacity: 0.0043 m3/sek

(minimum)

- - 991,409

Total CBM 14,964,737

VII-iii

Table 41. Calculation for the investment cost for the material needed upon start-up.

Material supply

Catalyst Catalyst volume 2.68 m3

Density 700 kg/m3

Catalyst mass 1,873 kg

Catalyst cost 60 USD/kg

Cost per "set" 112,390 USD

VKK 2012 7.5

Total cost

catalyst

112,390 USD

Dowtherm oil Oil flow 140,000 kg/h

circle time 0.5 h

Required oil 70,000 kg

Oil density 468.2 kg/m3

Oil volume 150 m3

Oil cost 4,400 USD/tonne

Total cost oil 308,000 USD

Table 42. Summary of the total capital cost for the plant.

Total cost of plant

Contingency and fees 1.15 15%

Auxiliary and facilities 1.25 25%

Grass Root Plant Cost 21,511,809 USD

Cost updating IAK 2004 115

IAK 2012 148

factor, f 9.39

Grass Root Plant Cost 27,684,763 USD

Total Grass Root Capital

Units and material

28,105,154 USD

VIII-i

Appendix VIII – Operating costs

Operating costs for the plant are calculated with excel and divided into three groups; fixed capital, direct costs and indirect costs.

Fixed capital

Table 43. Fixed capital cost including storage and spare parts.

Storage of feedstock Unit Source/Comment Storage of product Unit Source/Comment

Ethanol Ethylene

Days of storage days 10 Assumption by authors Days of storage days 10 Assumption by authors

Consumption kg/hr 22,000 Aspen Plus V8.2 Production kg/hr 11,642 Aspen Plus V8.2

Density kg/ m3 812 [61] Operating time hours/day 24

Consumption volume m3/hr 27.09 Year Days/year 365

Operating time hours/day 24 Annual production tonnes/year 101,987

Year Days/year 365

Annual consumption m3/year 237,340 Income per tonne dollar/tonne 1,280 Assumption by authors

Cost per m3 USD/ m3 436 [51] VKK 2012 7.5 [62]

Annual cost USD/year 372,679 Annual cost USD/year 470,218

NaOH Total cost of storage USD/year 851,101

Days of storage days 10 Assumption by authors

Consumption kg/hr 500 Aspen Plus V8.2

Operating time hours/day 24 Spare parts for plant

Year Days/year 365 Percent % 0.15

Annual consumption tonne/year 4,380

Cost per tonne dollar/tonne 520 [53] Total cost spare parts USD/year 33,256

VKK 2012 7.5 [62]

Annual cost USD/year 8,204 Total fixed capital USD/year 884,357

VIII-ii

Direct cost

Table 44. Direct costs for the plant, part one.

Feedstock Unit Source Dowtherm oil Unit Source

Ethanol 95 wt % Oil exchange % 5 Assumption by authors

Consumption kg/hr 22,000 Aspen Plus V8.2 Exchange of oil kg/year 3,500

Density kg/ m3 812 [61] Oil price USD/tonne 4,400 [64]

Consumption volume m3/hr 27.1 Total cost of oil USD/year 115,500

Operating time hours/day 24 Cooling water

Year Days/year 365 Required energy removal kW 6,992 Aspen Plus V8.2

Annual consumption m3/year 237,340 Required water flow kg/s 96

Cost per m3 USD/ m3 436 [51] Water density kg/ m3 1,000 [61]

Total cost of

feedstock

USD/year 103,463,894 Water flow m3/s 0.096

Solvents m3/year 3,027,456

NaOH Cost per m3 USD/m3 0.013 Assumption by authors

Consumption kg/hr 500 Aspen Plus V8.2 Total cost of cooling

water

USD/year 40,366

Operating time hours/day 24 Supervisors

Year Days/year 365 Percent 10% 0.1 [62]

Cost per tonne USD/tonne 520 [53] Total cost supervisors USD/year 104,000

Total cost NaOH USD/year 2,277,600 Laboratory work

Electricity Percent 10% 0.1 [62]

Required energy kW 757 Total cost laboratory USD/year 104,000

hours/year h/year 8,760 Maintenance & repair

Energy hours/year kWh/year 6,627,816 Percent 6% 0.06 Assumption by authors

Cost per kWh USD/kWh 0.13 Assumption by authors Total cost of maintenance USD/year 1,686,309

Total cost electricity USD/year 883,709

VIII-iii

Table 45. Direct costs for the plant, part two.

Catalyst Unit Source

Catalyst changes 2 Assumption by authors

catalyst cost per set dollar 112,390 Assumption by authors

total cost for changes dollar 224,780

Total cost of catalyst USD/year 29,553

Natural gas

Required energy kW 2,380 Aspen Plus V8.2

Energy every year GJ/year 75,043

Natural gas price USD/GJ 9 Assumption by authors

Total cost of natural gas USD/year 675,388

Operators

Shift workers N, s 5 Assumption by authors

Direct monthly salary USD/month 3,333 Assumption by authors

Shifts 5 Assumption by authors

Day workers N, d 1 Assumption by authors

Cost for shift workers USD/year 1,000,000

Cost for day workers USD/year 40,000

Total cost of operators USD/year 1,040,000

Total direct cost USD/year 110,420,318

VIII-iv

Indirect cost

Table 46. Indirect costs for the plant.

Overhead for staff Unit Source

Shift personnel 70 % 0.7 [62]

Day personnel 50 % 0.5 [62]

Total cost for overhead USD/year 720,000

Administration

Percent 25 % 0.25 [62]

Total cost of distribution USD/year 180,000

Total indirect cost USD/year 900,000

IX-i

Appendix IX – Investment calculations

Costs and revenues are calculated to determine the annual net income, ai.

ai =Ii – Ui

Grass root capital, G

Annual operating costs, Ui

Annual income, Ii

Annual net income, NI

Annuity factor, fa

Interest rate, X = 10 %

The pay-back time, n, can be estimated by using Equation 34 [52].

−𝐺 + ∑ 𝑎𝑖𝑛𝑖=1 ∙ (1 + 𝑋)−𝑖 > 0 Equation 34

ni is considered to be constant so the equation can be simplified to Equation 35 below.

𝑛 = − (ln(1−𝐺∙(

𝑋

𝑎𝑖))

ln(1+𝑋)) Equation 35

Annual net income is calculated with Equation 36 below and the capital cost with Equation 37

[52].

𝑁𝐼 = 𝑎𝑖 − 𝑓𝑎 ∙ 𝐺 Equation 36

𝐶𝑎𝑝𝑖𝑡𝑎𝑙 𝑐𝑜𝑠𝑡 = 𝑓𝑎 ∙ 𝐺 Equation 37

Annuity factor, fa = 0.1315, see Appendix VI for calculations.