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
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 2nd generation. Plants where 1st 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.
I
Contents
1
2
Introduction ........................................................................................................................ 1
1.1
Purpose ....................................................................................................................... 1
1.2
Disposition ................................................................................................................. 1
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
II
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.
1
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
2
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.
Pretreatment
Fermentation
Hydrolysis
Distillation
Ethanol
Lignocellulosic
biomass
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 2furaldehyde (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
3
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
Iron
Sodium
5 - 10 %
up to 10 %
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].
4
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 byproducts 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 byproducts 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
5
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 m2/g), V200 (202 ± 5 m2/g), D100 (100 ± 10 m2/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.
6
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
7
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 microscale 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 𝑋𝑀12 𝑂40 , where X is the
heteroatom and M the addendum atom. 𝐻3 𝑃𝑊12 𝑂40 and 𝐻4 𝑆𝑖𝑊12 𝑂40 are two examples that
are commercially available, where 𝐻3 𝑃𝑊12 𝑂40 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].
8
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
𝐴𝑔3 𝑃𝑊12 𝑂40 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 waterinsoluble 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.
𝐶𝑛 𝐻𝑚 → 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].
9
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.
Ethanol
conversion
Reaction
temperature
GHSVa/
WHSVb
Comments
Ref
100 %
260 °C
2 h-1 b
[31]
0.5% La 2% 99.9 %
P HZSM-5
100 %
240 °C
2 h-1 b
Nano HZSM- 98.4 %
5
98.4 %
240 °C
1 h-1 b
0.4
mol/l 99.6 %
NaOH
HZSM-5
99.7 %
265 °C
2.37 h-1 b
Ag3PW12O40
99.2 %
100 %
220 °C
6000 h-1 a
TPA-MCM41
99.9 %
98 %
300 °C
2.9 h-1 b
STA-MCM41
99.9 %
99 %
250 °C
2.9 h-1 b
P90 (Sasol)
θ-Al2O3
99.6 %
100.0 %
400 °C
1.43 h-1 b
P200 (Sasol)
γ, δ-Al2O3
100.0 %
99.7 %
350 °C
1.43 h-1 b
50 wt% ethanol
Lab scale
0.5 g catalyst
50 wt% ethanol
Lab scale
0.5 g catalyst
95 wt% ethanol
Lab scale
1 g catalyst
Run time: 620 h
20 w% ethanol
Lab scale
1 g catalyst
Run time: 350 h
15 g/m3 ethanol
Lab scale
0.5 cm3 catalyst
Run time: 1 h
99.98 % ethanol
Lab scale
0.2 g catalyst
Run time: 2-3 h
48 % ethanol
Lab scale
0.2 g catalyst
Run time: 8 h
7.9 wt % ethanol
Lab scale
0.5 g catalyst
7.9 wt % ethanol
Lab scale
0.5 g catalyst
Catalyst
HZSM-5
Max
ethylene
selectivity
96.4 %
a
Gas Hourly Space Velocity
b
Weight Hourly Space Velocity
10
[31]
[32]
[29]
[43]
[39]
[44]
[27]
[27]
V200 (UOP)
γ-Al2O3
99.4 %
99.8 %
400 °C
1.43 h-1 b
D100
(Degussa/
Evonik)
γ-Al2O3
99.9 %
99.3 %
350 °C
1.43 h-1 b
SA330
(Strem)
amorphous
100.0 %
99.8 %
350 °C
1.43 h-1 b
7.9 wt % ethanol [27]
Lab scale
0.5 g catalyst
7.9 wt % ethanol [27]
Lab scale
0.5 g catalyst
7.9 wt % ethanol [27]
Lab scale
0.5 g catalyst
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]
11
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].
12
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.
13
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.
14
Ethanol
storage tank
Heat
exchanger 2
1
Reaction
column
5
4
27
Flue
gas
Oil heater
2
6
NaOH (aq)
storage tank
3
Heat
exchanger 1
Oxygen
13
12
10
26
14
15
Pure water
7
17
8
Heat exchanger 3
Gas/liquid
separator
9
Distillation 1
Absorber
16
Dryer
Waste water
treatment
18
25
21
20
24
11
22
19
Waste water
treatment
Heat
exchanger 4
Ethylene
column
Heat
exchanger 5
Figure 6. Flowsheet of the process.
15
C2
stripper
23
Ethylene
storage
tank
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.
16
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.
17
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
Heat exchanger 2
2
Cold stream: 10-80 °C
Area: 284 m
Hot stream: Dowtherm A oil, 288
°C
Cold stream: 81-240 °C
Reactor
Total vessel volume: 8.1 m3
Temperature: 240 °C
Number of tubes: 1,052
GHSV: 6,000 h-1
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
Heat exchanger 3
Hot stream: 70 – 25 °C
Area: 409 m2
Cold stream: 15-30 °C
Inside diameter: 1.47 m
Temperature: 25 °C
Height: 2.52 m
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
Nr of ideal stages: 35
Height: 26.3
Top temperature: 33.2 °C
Gas/Liquid separator
Bottom temperature: 29.7 °C
Distillation column 1
Inside diameter: 0.51 m
Nr of ideal stages: 25
Height: 18.8 m
Top temperature: 78 °C
18
Condenser area: 100 m2
Bottom temperature: 100 °C
2
Reboiler area: 161 m
Dryer, 2 pieces
Inside diameter: 2.5 m
Residence time: 12 h
Height: 5.94 m
Adsorbent: Zeolite
Adsorbent mass: 16,970 kg
Packing height: 4.94 m
Heat exchanger 4
Hot stream: 32 – -104 °C
Area: 62 m2
Cold stream: Cooling media,
-150 °C
Ethylene column
Inside diameter: 1.49 m
Nr of ideal stages: 15
Height: 11.3 m
Top temperature: -104.2 °C
Condenser area: 216 m2
Bottom temperature: -98.3 °C
2
Reboiler area: 50 m
Heat exchanger 5
Hot stream: -104 – -130 °C
Area: 511 m2
Cold stream: Cooling media,
-150 °C
C2 stripper column
Inside diameter: 2.30 m
Nr of ideal stages: 10
Height: 7.5 m
Top temperature: -109.7 °C
2
Condenser area: 13 m
2
Reboiler area: 9.0 m
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
19
Bottom temperature: -103.9 °C
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 4042.
Table 4. Bare-Module Cost for all unit operations.
Unit
Bare-Module Cost, USD
542,150
Ethanol storage tank
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
6,896,130
Heat pump
991,410
Wastewater treatment plant
1,175,270
Thermal fluid heater
20
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.
Grass Root Capital Cost
1%
Material & supply
Unit operations
99%
Figure 7. Distribution of the grass root capital costs.
Unit operation investment cost
Storage Tanks & Pumps
7%
10%
1%
Heat Exchangers
4%
Distillation columns
10%
6%
46%
6%
2%
8%
Absorber
Dryers
Thermal fluid heater
Reactor
Heatpump
Wastewater treatment
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.
21
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.
Operating costs
11%
89%
Raw material
Other operating
costs
Figure 9. Distribution between operating costs for the process, the raw material is ethanol.
22
Other operating costs
Natural gas
30%
5%7%
7%
14%
20%
17%
Electricity
Storage & spare parts
Maintenance & repair
Personnel
Other materials
Annual capital cost
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
884,360 USD/year
Fixed capital
110,420,320 USD/year
Direct costs
900,000 USD/year
Indirect costs
112,204,680 USD/year
Annual operating costs
28,150,200 USD
Grass root capital
3,695,100 USD/year
Annual capital cost
115,899,800 USD/year
Total annual production cost
1,136 USD/tonne
Production cost per tonne ethylene
23
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
28,150,200 USD
Grass root capital
Annual operating costs
112,204,680 USD/year
Revenue
130,542,800 USD/year
1.75 years
Pay-back time
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.
Sensitivity analysis of changes in ethanol price
16
14
Pay-back time (years)
12
10
8
6
4
2
0
-50%
-40%
-30%
-20%
-10%
0%
10%
Ethanol price
Figure 11. Sensitivity analysis of how the pay-back time changes with different ethanol prices.
24
20%
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.
Sensitivity analysis of the profit with changes in ethanol price
70.000.000
60.000.000
Profit (USD/year)
50.000.000
40.000.000
30.000.000
20.000.000
10.000.000
0
-50%
-40%
-30%
-20%
-10%
0
5%
10%
15%
20%
30%
-10.000.000
-20.000.000
Ethanol price
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.
25
Ethylene price
Sensitivity analysis of changes in ethylene price
130%
0,5
125%
0,6
120%
0,7
100%
1,8
90%
8,0
0,0
1,0
2,0
3,0
4,0
5,0
6,0
7,0
8,0
9,0
Pay-back time (years)
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.
USD
Sensitivity analysis of capital costs and income
with changes in grass root capital
50.000.000
45.000.000
40.000.000
35.000.000
30.000.000
25.000.000
20.000.000
15.000.000
10.000.000
5.000.000
-
Grass
root
capital
Cost of
capital
-30%
0
30%
50%
75%
Annual
net
income
Grass root capital change
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
26
different grass root capital costs. A higher cost consequently means a longer pay-back time
for the plant.
Sensitivity analysis of pay-back time with changes
in grass root capital
Pay-back time (years)
3,50
3,00
2,50
Payback
time
2,00
1,50
1,00
0,50
-30%
0
30%
50%
75%
Grass root capital change
Figure 15. Sensitivity analysis of how the pay-back time changes with changes of the grass root capital costs.
27
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).
𝑁𝑎2 𝐶𝑂3 (𝑎𝑞) + 𝐶𝑎(𝑂𝐻)2 (𝑠) → 2 𝑁𝑎𝑂𝐻(𝑎𝑞) + 𝐶𝑎𝐶𝑂3 (𝑠)
(reac. 2)
𝐶𝑎𝐶𝑂3 (𝑠) + 𝑒𝑛𝑒𝑟𝑔𝑦 → 𝐶𝑎𝑂(𝑠) + 𝐶𝑂2 (𝑔)
(reac. 3)
𝐶𝑎𝑂(𝑠) + 𝐻2 𝑂(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
28
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.
29
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.
30
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36
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
ethanol
methanol
Mass ethanol
of
substance Is calculated
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.
I-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
From
1
Storage
To
3
Heat Ex 1
4
Temperature
Pressure
Vapor Fraction
Liquid Fraction
C
BAR
Mass Flow
Component Mass
Flow
ETHYLENE
ETHANOL
WATER
HYDROGEN
PROPENE
ISOBUTYLENE
ETHANE
DIETHYL-ETHER
ACETALDEHYDE
CO
CO2
METHANE
Component Mass
Fraction
ETHYLENE
ETHANOL
WATER
HYDROGEN
PROPENE
ISOBUTYLENE
ETHANE
DIETHYL-ETHER
ACETALDEHYDE
CO
CO2
METHANE
Mole Flow
Component Mole
Fraction
ETHYLENE
ETHANOL
WATER
4
5
3 & 10
Heat ex 2
Heat ex 2
Reaction
column
6
Reaction
column
Heat Ex 1
7
8
9
Heat Ex 1
Heat Ex 3
G-L separator
Heat Ex 3
G-L separator
Distillation 1
10
Distillation 1
11
Distillation 1
4
12
G-L separator
Absorber
10
1
0
1
81,00146
1
0,786987
0,213013
80,81845
1
0,794862
0,205138
240
1
1
0
240
1
1
0
70
1
0,677533
0,322467
25
1
0,467026
0,532974
25,00015
1
0
1
75,8087
1
1
0
99,35311
1
0
1
25,00015
1
1
0
KG/HR
22000
22000
22931,65
22931,65
22931,65
22931,65
22931,65
9978,459
931,3098
9047,149
12953,19
KG/HR
KG/HR
KG/HR
KG/HR
KG/HR
KG/HR
KG/HR
KG/HR
KG/HR
KG/HR
KG/HR
KG/HR
0
20900
1100
0
0
0
0
0
0
0
0
0
0
20900
1100
0
0
0
0
0
0
0
0
0
1,96998
21732,96
1100,335
0,000174
0,348436
70,63443
0,573134
1,329672
23,42929
3,81E-06
0,060798
0,007819
1,969984
21732,96
1100,335
0,000174
0,348436
70,63443
0,573134
1,329672
23,42929
3,81E-06
0,060798
0,007819
12323,1
1085,719
9108,672
8,695498
3,884474
259,2231
21,62941
2,160144
48,10788
0,00362
67,80514
2,650044
12323,1
1085,719
9108,672
8,695498
3,884474
259,2231
21,62941
2,160144
48,10788
0,00362
67,80514
2,650044
12323,1
1085,719
9108,672
8,695498
3,884474
259,2231
21,62941
2,160144
48,10788
0,00362
67,80514
2,650044
1,969911
1018,694
8861,41
0,000174241
0,3484269
70,63458
0,5731143
1,329756
23,43021
3,81E-06
0,0607944
0,00781874
1,969911
832,624
0,3310225
0,000174241
0,3484269
70,63458
0,5731143
1,32975
23,43021
3,81E-06
0,0607944
0,00781874
5,79E-23
186,0701
8861,079
0,00E+00
1,22E-16
3,29E-09
1,58E-21
5,44E-06
4,79E-07
5,35E-25
1,39E-19
1,40E-23
12321,13
67,02491
247,2627
8,695324
3,536047
188,5886
21,0563
0,830388
24,67768
0,003616
67,74435
2,642225
0
0,95
0,05
0
0
0
0
0
0
0
0
0
514,7262
0
0,95
0,05
0
0
0
0
0
0
0
0
0
514,7262
8,59E-05
9,48E-01
0,047983
7,60E-09
1,52E-05
3,08E-03
2,50E-05
5,80E-05
1,02E-03
1,66E-10
2,65E-06
3,41E-07
534,7337
8,59E-05
9,48E-01
0,047983
7,60E-09
1,52E-05
3,08E-03
2,50E-05
5,80E-05
1,02E-03
1,66E-10
2,65E-06
3,41E-07
534,7337
0,537384
0,047346
0,39721
0,000379
0,000169
0,011304
0,000943
9,42E-05
2,10E-03
1,58E-07
2,96E-03
0,000116
981,0151
0,537384
0,047346
0,39721
0,000379
0,000169
0,011304
0,000943
9,42E-05
2,10E-03
1,58E-07
2,96E-03
0,000116
981,0151
0,537384
0,047346
0,39721
0,000379
0,000169
0,011304
0,000943
9,42E-05
2,10E-03
1,58E-07
2,96E-03
0,000116
981,0151
0,000197416
0,1020893
0,888054
1,75E-08
3,49E-05
7,08E-03
5,74E-05
1,33E-04
0,00234808
3,82E-10
6,09E-06
7,84E-07
515,9036
0,00211521
0,8940355
0,000355438
1,87E-07
3,74E-04
0,0758443
0,000615385
0,00142783
0,0251583
4,09E-09
6,53E-05
8,40E-06
20
6,40E-27
2,06E-02
0,979433
0
1,35E-20
3,64E-13
1,75E-25
6,01E-10
5,30E-11
5,91E-29
1,54E-23
1,55E-27
495,9036
0,951204
0,005174
0,019089
0,000671
0,000273
0,014559
0,001626
6,41E-05
1,91E-03
2,79E-07
5,23E-03
0,000204
465,1115
0
0,881375
0,118625
0
0,881375
0,118625
0,000131
0,88221
0,114221
0,000131
0,88221
0,114221
0,447768
0,024023
0,515393
0,447768
0,024023
0,515393
0,447768
0,024023
0,515393
0,000136109
0,0428613
0,9534398
0,00351096
0,9036697
0,000918727
4,16E-27
0,008145
0,991855
0,944283
0,003128
0,029509
KMOL/HR
II-i
HYDROGEN
PROPENE
ISOBUTYLENE
ETHANE
DIETHYL-ETHER
ACETALDEHYDE
CO
CO2
METHANE
Volume Flow
m^3/HR
Stream ID
From
0
0
0
0
0
0
0
0
0
27,09045
14
To
Absorber
0
0
0
0
0
0
0
0
0
11933,82
15
Absorber
1,62E-07
1,55E-05
0,002354
3,56E-05
3,35E-05
0,000995
2,54E-10
2,58E-06
9,11E-07
12515
16
Absorber
1,62E-07
1,55E-05
0,002354
3,56E-05
3,35E-05
0,000995
2,54E-10
2,58E-06
9,11E-07
22814,41
17
Dryer
18
Dryer
21
0,004397
9,41E-05
0,00471
0,000733
2,97E-05
0,001113
1,32E-07
1,57E-03
0,000168
11368,06
25
20
Ethylene
column
Heat ex 5
-98,3353
1
0
1
-104,149
1
1
0
-130
1
0,0118454
0,9881545
Ethylene
storage tank
-103,9743
1
0
1
Ethylene
column
22
1,68E-07
1,60E-05
0,00244021
3,69E-05
3,48E-05
0,00103093
2,63E-10
2,68E-06
9,45E-07
10,30996
Heat ex 5
0,00E+00
5,87E-21
1,18E-13
1,06E-25
1,48E-10
2,19E-11
3,85E-29
6,38E-24
1,76E-27
9,906551
25
0,009274
0,000181
0,007227
0,001506
2,41E-05
0,001204
2,78E-07
3,31E-03
0,000354
11529,74
23
C2 Stripper
24
C2 Stripper
26
25
-109,7252
1
1
0
-104,3031
1
0,463977
0,536023
20
1
1
0
1252,258
1
1
0
21 & 24
29,67447
1
0
1
32,31291
1
0
1
32,31291
1
1
0
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
KG/HR
KG/HR
KG/HR
KG/HR
KG/HR
KG/HR
KG/HR
KG/HR
KG/HR
KG/HR
KG/HR
KG/HR
KG/HR
KG/HR
KG/HR
KG/HR
KG/HR
KG/HR
0
0
1500
0
0
0
0
0
0
0
0
0
7,24E-15
212,614
0
0
287,386
0
12299,55
31,08229
247,1647
8,695317
3,517747
185,4287
21,012
0,776636
23,71615
0,003616
0,127987
2,639804
0
0
0
0
0
0
21,57973
35,94262
1527,776
6,77E-06
0,018299
3,159874
0,044293
0,053753
0,961529
3,47E-07
0,000899
0,002421
2,57E-14
160,3531
0,00028
92,19792
287,386
0
0
0
247,1647
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
12299,55
31,08229
0
8,695317
3,517747
185,4287
21,012
0,776636
23,71615
0,003616
0,127987
2,639804
0
0
0
0
0
0
12299,55
31,08229
0
8,70E+00
3,517747
185,4287
21,012
0,776636
23,71615
3,62E-03
0,127987
2,64E+00
0
0
0
0
0
0
362,0619
3,11E+01
0
5,94E-14
3,52E+00
1,85E+02
21,00915
7,77E-01
2,37E+01
1,13E-11
1,28E-01
1,56E-06
0
0
0
0
0
0
11937,49
2,49E-97
0
8,695317
4,96E-21
6,73E-37
0,002858
7,18E-59
2,44E-51
0,003616
5,79E-06
2,639802
0
0
0
0
0
0
11937,49
2,49E-97
0
8,70E+00
4,96E-21
6,73E-37
0,0028576
7,18E-59
2,44E-51
3,62E-03
5,79E-06
2,639802
0
0
0
0
0
0
11642,27
0
0
7,21E-07
0
0
2,86E-03
0
0
3,35E-06
5,79E-06
0,0223378
0
0
0
0
0
0
295,2183
0
0
8,695316
0
0
8,58E-09
0
0
0,00361278
7,89E-12
2,617464
0
0
0
0
0
0
657,2802
31,08229
0
8,695316
3,517747
185,4287
21,00915
0,776636
23,71615
0,003613
0,127982
2,617466
0
0
0
0
0
0
0
0
0
0,00E+00
0
0
0
0
0
0,00E+00
0
0
0
0
0
0
0
5759,784
2,60E-35
2,71E-34
1264,989
5,95E-04
4,37E-36
4,23E-35
1,04E-33
5,54E-42
1,40E-33
2,14E-02
2832,458
3,14E-25
0
0
0
0
0
2596,57
0
0
0,75
0
0
0
0
0,959125
0,002424
0,019274
0,000678
0,000274
0,01446
0,001639
0,010134
0,016879
0,717442
3,18E-09
8,59E-06
1,48E-03
2,08E-05
0
0
1
0
0
0
0
0,977975
0,002471
0
0,000691
0,00028
0,014744
0,001671
0,977975
0,002471
0
0,000691
2,80E-04
0,014744
0,001671
0,576788
0,049516
0,00E+00
9,46E-17
0,005604
2,95E-01
3,35E-02
0,999051
2,08E-101
0,00E+00
0,000728
4,16E-25
5,63E-41
2,39E-07
0,9990508
2,08E-101
0
0,0007277
4,16E-25
5,63E-41
2,39E-07
0,9999978
0
0
6,20E-11
0
0
2,45E-07
0,9630828
0
0
0,0283665
0
0
2,80E-11
0,703534
0,03327
0
0,009307
0,003765
0,198478
0,022488
0
0
0
0
0,00E+00
0
0
3,88E-39
4,04E-38
0,1889725
8,89E-08
6,54E-40
6,32E-39
1,55E-37
II-ii
Oil heater
27
Oil heater
33,24234
1
1
0
Mass Flow
Component Mass
Flow
ETHYLENE
ETHANOL
WATER
HYDROGEN
PROPENE
ISOBUTYLENE
ETHANE
DIETHYL-ETHER
ACETALDEHYDE
CO
CO2
METHANE
H3O+
OHHCO3CO3-2
NA+
OXYGE-01
Component Mass
Fraction
ETHYLENE
ETHANOL
WATER
HYDROGEN
PROPENE
ISOBUTYLENE
ETHANE
C2 Stripper
4,32E-06
0,000413999
0,0629457
9,53E-04
0,000896991
0,0265931
6,80E-09
6,91E-05
2,44E-05
580,2716
25
1
0
1
C
BAR
Heat ex 4
19
Heat ex 4
0,004397
9,41E-05
0,00471
0,000733
2,97E-05
0,001113
1,32E-07
1,57E-03
0,000168
18969,93
Ethylene
column
-104
1
0,457764
0,542236
Temperature
Pressure
Vapor Fraction
Liquid Fraction
Dryer
0,004397
9,41E-05
0,00471
0,000733
2,97E-05
0,001113
1,32E-07
1,57E-03
0,000168
41854,99
Oil heater
DIETHYL-ETHER
ACETALDEHYDE
CO
CO2
METHANE
H3O+
NA2CO3
OHHCO3CO3-2
NA+
OXYGE-01
Mole Flow
Component Mole
Fraction
ETHYLENE
ETHANOL
WATER
HYDROGEN
PROPENE
ISOBU-01
ETHAN-01
DIETH-01
ACETA-01
CO
CO2
METHA-01
DOWTH-01
H3O+
NA2CO3
OHHCO3CO3-2
NA+
NAOH
OXYGE-01
Volume Flow
KMOL/HR
0
0
0
0
0
3,62E-18
0,00E+00
0,106307
0
0
0,143693
0
108,2645
6,06E-05
1,85E-03
2,82E-07
9,98E-06
2,06E-04
0
0
0
0
0
0
0
461,9393
2,52E-05
4,52E-04
1,63E-10
4,22E-07
1,14E-06
1,21E-17
0,00E+00
0,075302
1,31E-07
4,33E-02
0,134956
0
109,9003
0
0
0
0
0
0
0
0
0
0
0
0
13,71972
6,18E-05
1,89E-03
2,88E-07
1,02E-05
2,10E-04
0
0
0
0
0
0
0
448,2196
6,18E-05
0,001886
2,88E-07
1,02E-05
0,00021
0
0
0
0
0
0
0
448,2196
1,24E-03
3,78E-02
1,79E-14
2,04E-04
2,49E-09
0
0
0
0
0
0
0
18,21959
6,01E-63
2,04E-55
3,03E-07
4,84E-10
2,21E-04
0
0
0
0
0
0
0
430
6,01E-63
2,04E-55
3,03E-07
4,84E-10
2,21E-04
0
0
0
0
0
0
0
430
0,00E+00
0
2,88E-10
4,97E-10
1,92E-06
0
0
0
0
0
0
0
415
0
0
1,18E-05
2,57E-14
0,00853888
0
0
0
0
0
0
0
15
0,000831
0,025385
3,87E-06
0,000137
0,002802
0
0
0
0
0
0
0
33,21959
0
0
0
0,00E+00
0
0
0
0
0
0
0
1
180
8,27E-46
2,09E-37
3,20E-06
4,23E-01
4,69E-29
0
0
0
0
0
0
0,3878928
215,7242
m^3/HR
0
0
0,769067
0
0
0
0
0
0
0
0
0
0
3,52E-18
0
0,115466
0,00E+00
0
0,115466
0
0
1,573762
0,949103
0,001461
0,0297
0,009338
0,000181
0,007154
0,001513
2,27E-05
0,001165
2,79E-07
6,30E-06
0,000356
0,00E+00
0,00E+00
0
0
0
0
0
0
0
11702,71
0,006999
0,007099
0,771649
3,06E-08
3,96E-06
0,000512
1,34E-05
6,60E-06
0,000199
1,13E-10
1,86E-07
1,37E-06
0,00E+00
1,23E-17
0,00E+00
0,085788
4,17E-08
0,01398
0,113748
0,00E+00
0
1,709154
0
0
1
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0,250685
0,978154
0,001505
0
0,009623
0,000187
0,007373
0,001559
2,34E-05
0,001201
2,88E-07
6,49E-06
0,000367
0,00E+00
0,00E+00
0
0
0
0
0
0
0
11383,52
0,978154
0,001505
0
0,009623
0,000187
0,007373
1,56E-03
2,34E-05
0,001201
2,88E-07
6,49E-06
0,000367
0,00E+00
0
0,00E+00
0
0
0
0
0
0
2897,645
0,708359
0,037031
0
1,62E-15
4,59E-03
0,181392
0,038348
5,75E-04
2,95E-02
2,21E-14
1,60E-04
5,35E-09
0,00E+00
0,00E+00
0
0
0
0
0
0
0
0,999688
0,989586
1,26E-101
0
0,010031
2,74E-25
2,79E-41
2,21E-07
2,25E-63
1,29E-55
3,00E-07
3,06E-10
3,83E-04
0,00E+00
0,00E+00
0
0
0
0
0
0
0
6042,049
0,9895856
1,26E-101
0
0,0100311
2,74E-25
2,79E-41
2,21E-07
2,25E-63
1,29E-55
3,00E-07
3,06E-10
0,0003827
0,00E+00
0,00E+00
0,00E+00
0
0
0
0
0
0
80,38169
0,9999964
0
0
8,62E-10
0
0
2,29E-07
0
0
2,88E-10
3,17E-10
3,36E-06
0,00E+00
0,00E+00
0
0
0
0
0
0
0
20,48199
0,7015537
0
0
0,2875606
0
0
1,90E-11
0
0
8,60E-06
1,19E-14
0,010877
0,00E+00
0,00E+00
0
0
0
0
0
0
0
203,8151
0,705286
0,02031
0
0,129845
0,002516
0,099486
0,021032
0,000315
0,016206
3,88E-06
8,75E-05
0,004911
0
0
0
0
0
0
0
0
0
217,3445
0
0
0
0
0
0
0,00E+00
0
0
0
0
0
0,00E+00
0
0
0
0
0
0
0
1
4387,223
4,29E-39
2,72E-38
0,325497
1,37E-06
4,82E-40
3,50E-39
1,60E-37
3,46E-46
1,47E-37
3,55E-06
0,2983426
9,06E-29
0,00E+00
0
0
0
0
0
0
0
0,3761555
27359,74
II-iii
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 𝐶2 𝐻5 𝑂𝐻 → 𝐶2 𝐻4 + 𝐻2 𝑂
98.0
2 2 𝐶2 𝐻5 𝑂𝐻 → (𝐶2 𝐻5 )2 𝑂 + 𝐻2 𝑂
0.005
3 2 𝐶2 𝐻5 𝑂𝐻 → 𝐶4 𝐻8 + 2 𝐻2 𝑂
1.5
4 𝐶2 𝐻5 𝑂𝐻 → 𝐶2 𝐻4 𝑂 + 𝐻2
0.125
5 2 𝐶2 𝐻5 𝑂𝐻 → 𝐶3 𝐻6 + 𝐶𝑂2 + 3 𝐻2
0.0375
6 2 𝐶2 𝐻5 𝑂𝐻 → 𝐶2 𝐻6 + 2 𝐶𝑂 + 3 𝐻2
0.3125
7 2 𝐶2 𝐻5 𝑂𝐻 → 3 𝐶𝐻4 + 𝐶𝑂2
0.02
8 𝐶𝑂 + 𝐻2 𝑂 ↔ 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 𝐻2 𝑂 ↔ 𝐻3 𝑂+ + 𝑂𝐻 −
2 𝐻𝐶𝑂3− + 𝐻2 𝑂 ↔ 𝐻3𝑂+ + 𝐶𝑂32−
3 𝐶𝑂2 + 𝑂𝐻 − → 𝐻𝐶𝑂3−
4 𝐻𝐶𝑂3 → 𝐶𝑂2 + 𝑂𝐻 −
III-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
Equation 4
𝐺𝐻𝑆𝑉
Where the volume flow could be calculated as:
𝐹𝐸𝑡𝑂𝐻 =
𝑚̇𝐸𝑡𝑂𝐻
𝑀𝐸𝑡𝑂𝐻
1 𝑏𝑎𝑟,0 ℃
∙ 𝜌𝐸𝑡𝑂𝐻
Equation 5
1 𝑏𝑎𝑟,0 ℃
Where 𝑚̇𝐸𝑡𝑂𝐻 is the mass flow of ethanol, 𝑀𝐸𝑡𝑂𝐻 is the molar mass of ethanol and 𝜌𝐸𝑡𝑂𝐻
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
Equation 6
𝑉𝑡𝑢𝑏𝑒
With the total number of tubes the cross sectional area of all the tubes was calculated with
Equation 7:
2
𝐴𝑡𝑢𝑏𝑒𝑠 = 𝑛𝑡𝑢𝑏𝑒𝑠 ∙ 𝜋𝑟𝑡𝑢𝑏𝑒
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
Equation 8
𝑝𝑓𝑡𝑢𝑏𝑒
𝐴𝑟𝑒𝑎𝑐𝑡𝑜𝑟
𝑑𝑟𝑒𝑎𝑐𝑡𝑜𝑟 = 2 ∙ √
Equation 9
𝜋
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
Equation 11
𝐴𝑡𝑢𝑏𝑒𝑠
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-i
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
𝑊
transfer coefficient was assumed to be 500 2 , [56]. When comparing the heat load required
𝑚 ∙𝐾
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:
𝜌𝑙 −𝜌𝑣 0.5
𝑢𝑣 = (−0.17𝑙𝑡2 + 0.27𝑙𝑡 − 0.047) ∙ (
𝜌𝑣
)
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:
4𝑉𝑤
𝐷𝑐 = √
Equation 15
𝜋𝜌𝑣 𝑢𝑣
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
𝐸0 =
Equation 16
𝑁𝐼
Equation 17
𝑁𝑅
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
Equation 18
𝑘∙𝑇𝐿
𝑇𝐿 =
(𝑇2,𝑖𝑛 −𝑇1,𝑜𝑢𝑡 )−(𝑇2,𝑜𝑢𝑡 −𝑇1,𝑖𝑛 )
Equation 19
(𝑇2,𝑖𝑛 −𝑇1,𝑜𝑢𝑡 )
ln(
)
(𝑇2,𝑜𝑢𝑡 −𝑇1,𝑖𝑛 )
𝑚̇𝑐𝑜𝑜𝑙𝑖𝑛𝑔 =
𝑄
Equation 20
𝐶𝑝 ∙∆𝑇
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-ii
𝜌𝐿
𝑢𝐺 = 0.064√
𝜌𝐺
−1
Equation 21
From the gas velocity and the gas flow rate (QG) the diameter (D) of the vessel could be
calculated:
𝑄𝐺
𝐷 = 2√
Equation 22
𝑢𝐺 𝜋
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 =
𝑢𝐺 𝑄𝐿 𝜃
Equation 23
𝑄𝐺
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
Equation 26
𝑉∙𝑐0
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].
IV-iii
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
-1
Residence time in reactor
6,000
Mass flow EtOH
201,016,000 kg/year
Density EtOH 240 °C
h
[40]
Aspen Plus V8.2
3
135.5
kg/m
Molar gas volume EtOH 0 °C 1 22.4
bar
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
2
500
W/(m ·°C)
Table 14. Flows and assumed dimensions of reactor.
Calculations flows
Mass flow EtOH
Volumetric
EtOH
Unit
22,931
kg/h
flow 169.23
m3/h
Normal
volumetric 11,150
flow EtOH
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-i
Assumption by authors
Possible heat transfer tubes, Q
7134
kW
Table 16. Calculations of reactor vessel.
Calculations reactor vessel
Unit
Total height
4,15
m
Cross sectional area whole 2.97
reactor
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
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/ m3
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
V-ii
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
Cooling 20
°C
Assumption by authors
Temp.
Cooling 40
media out
°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
log. mean temp
45.27
°C
Area
99.7
m2
Mass rate cooling
water
16.9
kg/s
Temp.
media in
Aspen Plus V8.2
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
log. mean temp
163.9
°C
Area
161.2
m2
Aspen Plus V8.2
Ethylene column
Distillation tower ethylene/higher hydrocarbons
Table 20. Calculations for the ethylene column.
Unit
Parameters
V-iii
Source/Comment
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/m3
Aspen Plus V8.2
Temperature column
169
174
K
Aspen Plus V8.2
3
Density ethylene vapor
1.996
1.939
kg/m
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
largest diameter
Real number of stages
18.75
Column height
11.25
m
Column surface area
52.69
m2
the
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
log. mean temp
46
°C
Area
216.4
m2
Aspen Plus V8.2
V-iv
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
°C
Aspen Plus V8.2
-98.33
Heat transfer coeff. 114
(k)
W/(m2·°C) Assumption by authors, made on ethylene
liquid-ethylene vapor [56]
Heat transferred. Q
1,884
kW
log. mean temp
330.0
°C
Area
50.07
m2
Aspen Plus V8.2
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
3
3
Volumetric fraction ethylene
0.7
1
m /m
Aspen Plus V8.2
Temperature column
163
169
K
Aspen Plus V8.2
Density ethylene vapor
Density hydrogen vapor
Vapor density
2.070
0.149
1.493
1.996
0.143
1.996
V-v
3
Assuming ideal gas
3
Assuming ideal gas
3
Assuming ideal gas
kg/m
kg/m
kg/m
Maximal vapor velocity
0.534
0.461
m/s
Diameter column
2.302
2.142
m
Continues with
largest diameter
Real number of stages
12.5
Column height
7.5
m
Column surface area
54.23
m2
the
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
log. mean temp
40
°C
Area
12.5
m2
Aspen Plus V8.2
Table 25. Calculations for the reboiler in the ethylene column.
Reboiler
Parameters
Unit
Source/Comment
°C
Assumption by authors
Temp. heat. Media in
215
Temp. heat. Media out
210.64
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
log. mean temp
316.8
°C
Area
8.67
m2
Aspen Plus V8.2
Aspen Plus V8.2
V-vi
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
°C
Aspen Plus V8.2
°C
Aspen Plus V8.2
kW
Aspen Plus V8.2
Stream from
(ethylene) in
reactor 240
Stream from reactor 70
(cooled ethylene) out
Temp. transfer efficiency
0.739
Calculations
Heat transfer. Q
Overall heat
coefficient (k)
5,808
transfer 341
Log. mean temp.
102.0
Area
167.0
Assumption by authors, based on [56]
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
3,088
kW
Aspen Plus V8.2
Calculations
Heat transfer, Q
Overall
heat
coefficinet (k)
transfer 100
Log. mean temp.
108.8
Area
283.8
Assumption by authors, based on [56]
m2
V-vii
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 15
in
°C
Water, assumption by authors
Temp. Cooling media 30
out
°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
water
cooling 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 -150
in
°C
Assumption by authors
Temp. Cooling media -150
out
°C
Assumption by authors
Assumption by authors, based on
ethylene vapor-ethylene vapor [56]
Calculations
Heat transfer coeff (k)
114
W/(m2·°C)
Heat transferred, Q
697
kW
log.Temp.
98.88
°C
Area
61.83
m2
V-viii
Aspen Plus V8.2
Heat Exchanger 5
Cooling of stream before C2 Stripper column.
Table 30. Calculations for heat exchanger 5.
Parameters
Unit
Source/Comment
-104
°C
Aspen Plus V8.2
Temp. Ethylene out -130
°C
Aspen Plus V8.2
Temp.
Coooling -150
media in
°C
Assumption by authors
Temp.
Cooling -150
media out
°C
Assumption by authors
W/(m2·°C)
Assumption by authors, based on ethylene
vapor-ethylene vapor [56]
Temp. Ethylene in
Calculations
Heat transfer coeff 114
(k)
Heat transferred, Q
1,819 kW
log.Temp.
31.22 °C
Area
511.2 m2
Aspen Plus V8.2
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
3
m /s
Aspen Plus V8.2
600
s
Assumption by authors.
Gas velocity
1.865
m/s
Diameter flash column
1.468
m
Height gas
1.468
m
Volumetric
liquid
flow
Residence time
rate 0.0029
Calculations
Choose either height=1
or if D>1, choose height of gas [58].
V-ix
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
Volumetric
ethylene
fraction 0.97
Bottom
0.98
m3/m3
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
3
Assumption ideal gas
Vapor density
1.089
1.106
kg/m
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-x
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 [60]
adsorbent
Concentration water
0.019
kg water/kg mass Aspen Plus V8.2
flow
Total mass flow
12,824
kg/h
Drying time
12
h
Density adsorbent
700
Calculations
Aspen Plus V8.2
Assumption by authors
3
kg/m
[60]
Unit
Adsorbent mass
14,143
kg adsorbent
Volume adsorbent
20.20
m3 adsorbent
Extra adsorbent
4.040
m3 adsorbent
volume 24.24
m3 adsorbent
Total
adsorbent
Total mass adsorbent
16,971
kg adsorbent
V-xi
Assuming cylindrical geometry
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
Aspen Plus V8.2
kW
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
Aspen Plus V8.2
kW
Wastewater treatment plant
Table 39. Calculations for the wastewater treatment plant.
Wastewater from
Absorber
EtOH distillation
Total wastewater capacity
1.71
0.000475
9.91
0.002753
0.003228
Unit
m3/h
m3/s
m3/h
m3/s
m3/s
V-xii
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].
𝛼
𝐶𝐵𝑀 = 𝐶𝑝 ∙ 𝐹𝐵𝑀
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 = (∑𝑛𝑖=1 𝐶𝐵𝑀𝑖 ) ∙ 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].
(𝐼
)
𝑓 = (𝐼𝐴𝐾)𝑥
Equation 29
𝐾$,𝑋 = 𝑓 ∙ 𝐾$,𝑌
Equation 30
𝐴𝐾 𝑌
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)
𝑓𝐴 =
𝑋
Equation 31
Equation 32
1−(1+𝑋)−𝑁
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-i
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.
VI-ii
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 Material
conditions
Ethanol storage
tank
Sodium hydroxide
storage tank
Ethylene storage
tank
Centrifugal pump
1
Heat exchanger 1
6,502 m3 Atmosperic
pressure- Cone roof
378 m3 Atmosperic
pressure- Cone roof
4,916 m3 Atmospheric
press-Gas holder
Pressure rise 10 bar
12 kW, Centrifugal pump
168 m2 Shell and Tube,
Fixed tube sheet and Utube
284 m2 Shell and Tube,
Fixed tube sheet and Utube
D=1.94 m, H=4.15 m,
Process vessel, vertically,
no packing
A=297.3 m2, Shell and
tube, fixed tube
Heat exchanger 2
Reactor vessel
Thermal fluid
heater
Heat exchanger 3
Gas/Liquid
separator
Centrifugal pump
2
Absorber
Dryer - 2 st
Heat exchanger 4
10,158 kW
Furnaces, Thermal fluid
heater, Mineral oil heaters
409 m2
Shell and Tube, Fixed
tube sheet and U-tube
D=1.47 m, H=2.52 m,
Process vessel, Vertically
oriented, No packing or
trays
Pressure rise 10 bar, 3kW,
Centrifugal
D=1.52 m, H=26.3 m
43 Trays, Process vessel,
Vertically oriented,
Sieve-trays
D=2.5 m, H=5.94 m,
Packed height 4.94 m
Process vessel, vertically
62 m2
Shell and Tube, Fixed tub
Cp
(Purchased
cost, USD)
CBM
(bare module
cost, USD)
Stainless steel
154,899
542,146
Stainless steel
24,339
85,188
Stainless steel
224,877
787,069
Stainless steel
9,348
45,111
Stainless steel
16,731
96,749
Stainless steel
23,814
137,710
Stainless steel
20,209
190,545
Stainless steel
24,585
142,168
534,216
1,175,274
Stainless steel
30,845
178,368
Stainless steel
12,161
114,666
Stainless steel
5,477
26,433
Stainless steel and
stainless steel trays
78,444
878,861
Stainless steel
Zeolite adsorbent
66,658
868,824
9,113
52,700
-
Stainless steel
VII-i
Ethylene column
Heat exchanger 5
C2 stripper
Ethanol
distillation tower
Heatpump
Wastewater
treatment plant
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
34,666
372,576
Stainless steel
19,755
114,235
Stainless steel
8,080
46,727
511 m2 Shell and Tube,
Fixed tub
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
36,318
210,016
Stainless steel and
stainless steel trays
35,937
453,142
Stainless steel
4,080
23,593
Stainless steel
3,503
20,258
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
36,470
350,919
Stainless steel
12,071
69,803
Stainless steel
16,275
94,113
Heat absorption
rate=3,708 kW
Auxiliary, Mechanical ref
unit
Capacity: 0.0043 m3/sek
(minimum)
Coolant temp:
- 55°C
-
-
6,896,134
-
991,409
Total CBM
VII-ii
14,964,737
Table 41. Calculation for the investment cost for the material needed upon start-up.
Material supply
Catalyst
2.68 m3
700 kg/m3
1,873 kg
Catalyst volume
Density
Catalyst mass
Catalyst cost
Cost per "set"
VKK 2012
60 USD/kg
112,390 USD
7.5
112,390 USD
Oil flow
circle time
Required oil
Oil density
Oil volume
140,000
0.5
70,000
468.2
150
Total cost
catalyst
Dowtherm oil
Oil cost
kg/h
h
kg
kg/m3
m3
4,400 USD/tonne
308,000 USD
Total cost oil
Table 42. Summary of the total capital cost for the plant.
Total cost of plant
Contingency and fees
Auxiliary and facilities
1.15 15%
1.25 25%
Grass Root Plant Cost
Cost updating
21,511,809 USD
IAK 2004
IAK 2012
factor, f
115
148
9.39
27,684,763 USD
28,105,154 USD
Grass Root Plant Cost
Total Grass Root Capital
Units and material
VII-iii
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
Ethanol
Storage of product
Unit
Source/Comment
Ethylene
Days of storage
days
Consumption
kg/hr
10
22,000
3
Density
kg/ m
812
Consumption volume
m3/hr
27.09
Operating time
hours/day
24
Year
Days/year
365
Annual consumption
m3/year
3
Assumption by authors
Aspen Plus V8.2
[61]
237,340
3
Cost per m
USD/ m
Annual cost
USD/year
436
[51]
372,679
NaOH
Days of storage
days
10
Assumption by authors
Consumption
kg/hr
500
Aspen Plus V8.2
Operating time
hours/day
24
Year
Days/year
365
Annual consumption
tonne/year
4,380
Cost per tonne
dollar/tonne
VKK 2012
Annual cost
USD/year
days
10
Production
kg/hr
11,642
Operating time
hours/day
24
Year
Days/year
365
Annual production
tonnes/year
101,987
Income per tonne
dollar/tonne
1,280
VKK 2012
7.5
Annual cost
USD/year
470,218
Total cost of storage
USD/year
851,101
Spare parts for plant
520
[53]
7.5
[62]
8,204
Days of storage
Percent
%
Total cost spare parts
USD/year
Total fixed capital
USD/year 884,357
VIII-i
0.15
33,256
Assumption by authors
Aspen Plus V8.2
Assumption by authors
[62]
Direct cost
Table 44. Direct costs for the plant, part one.
Feedstock
Unit
Source
Ethanol 95 wt %
Consumption
kg/hr
3
Density
kg/ m
3
Dowtherm oil
Unit
Source
Oil exchange
%
5
22,000
Aspen Plus V8.2
Exchange of oil
kg/year
3,500
812
[61]
Oil price
USD/tonne
4,400
USD/year
115,500
Consumption volume
m /hr
27.1
Total cost of oil
Operating time
hours/day
24
Cooling water
Year
Days/year
365
Required energy removal
kW
6,992
Required water flow
kg/s
96
Water density
kg/ m3
1,000
3
Annual consumption
m /year
237,340
Cost per m3
USD/ m3
436
Total
cost
feedstock
of
USD/year
[51]
103,463,894
Water flow
m /s
0.096
m3/year
3,027,456
Cost per m3
USD/m3
0.013
Total
water
USD/year
40,366
Solvents
NaOH
Aspen Plus V8.2
3
Consumption
kg/hr
500
Operating time
hours/day
24
Supervisors
Year
Days/year
365
Percent
10%
0.1
Cost per tonne
USD/tonne
520
Total cost supervisors
USD/year
104,000
Total cost NaOH
USD/year
2,277,600
Percent
10%
0.1
USD/year
104,000
[53]
cost of cooling
Assumption by authors
[64]
Aspen Plus V8.2
[61]
Assumption by authors
[62]
Laboratory work
Electricity
Required energy
kW
757
Total cost laboratory
hours/year
h/year
8,760
Maintenance & repair
Energy hours/year
kWh/year
6,627,816
Percent
6%
0.06
Cost per kWh
USD/kWh
0.13
Total cost of maintenance
USD/year
1,686,309
Total cost electricity
USD/year
883,709
Assumption by authors
VIII-ii
[62]
Assumption by authors
Table 45. Direct costs for the plant, part two.
Catalyst
Unit
Catalyst changes
Source
2
Assumption by authors
Assumption by authors
catalyst cost per set
dollar
112,390
total cost for changes
dollar
224,780
Total cost of catalyst
USD/year
29,553
Required energy
kW
2,380
Energy every year
GJ/year
75,043
Natural gas price
USD/GJ
9
Total cost of natural gas
USD/year
675,388
Shift workers
N, s
5
Assumption by authors
Direct monthly salary
USD/month
3,333
Assumption by authors
5
Assumption by authors
Assumption by authors
Natural gas
Aspen Plus V8.2
Assumption by authors
Operators
Shifts
Day workers
N, d
1
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-iii
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
Percent
25 %
0.25
Total cost of distribution
USD/year
180,000
Administration
[62]
Total indirect cost USD/year 900,000
VIII-iv
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 + ∑𝑛𝑖=1 𝑎𝑖 ∙ (1 + 𝑋)−𝑖 > 0
Equation 34
ni is considered to be constant so the equation can be simplified to Equation 35 below.
𝑋
𝑛 = −(
ln(1−𝐺∙(𝑎 ))
𝑖
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.
IX-i