Conversion of Heat and Electricity into Chemical Components · Conversion of Heat and Electricity...
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Conversion of Heat and Electricity into Chemical Components Ivarsson, C.F.P.; Nolin, M.P.D.; Petersson, F.; Rudenius, O.N.E. 2013 Supervisors: Karlsson, H.T.; Hulteberg, P.C.; Dept. of Chemical Engineering, Faculty of Engineering, Lund University Børresen, B.T.; Mejdell, A.; Statoil
Transcript of Conversion of Heat and Electricity into Chemical Components · Conversion of Heat and Electricity...
Conversion of Heat and Electricity into Chemical Components
Ivarsson, C.F.P.; Nolin, M.P.D.; Petersson, F.; Rudenius, O.N.E. 2013
Supervisors:
Karlsson, H.T.; Hulteberg, P.C.;
Dept. of Chemical Engineering, Faculty of Engineering, Lund University
Børresen, B.T.; Mejdell, A.;
Statoil
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Disclaimer
This report was produced on the request of the Norwegian company Statoil, as a project in the course “Feasibility Studies on Industrial Plants”
(KET050), at the Department of Chemical Engineering, Faculty of Engineering, Lund University, Sweden. Neither the authors, Lund University
or Statoil can be held responsible for any of the effects or consequences following from using the information of this study. The authors, Lund
University or Statoil do not make any warranty, expressed or implied, or assumes any legal liability or responsibility for the accuracy or
completeness of this information. No reproduction is allowed without the written permission from the authors or supervisors.
Copyright © 2013 Pontus Ivarsson, Mikael Nolin, Filip Petersson, Oskar Rudenius.
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Abstract The aim of this report is to investigate the possibility of producing chemical components from carbon
dioxide, electricity and heat at locations where direct usage of the produced electricity and heat is
not possible due to lack of infrastructure. Carbon dioxide is an interesting raw material since it is
likely to be available as a feedstock in the future, due to raised emission taxes which eventually will
motivate usage of carbon capture technology. In the review part of this report, the utilization of heat
and electricity to produce chemical products such as diesel, algal oil and methanol is investigated.
Diesel is produced in the Fischer-Tropsch process from synthesis gas, a mixture of carbon monoxide
and hydrogen. The carbon monoxide is formed in a Reversed Water Gas Shift reactor from carbon
dioxide and the hydrogen is produced by electrolysis of water. To achieve high selectivity to diesel a
low temperature Fischer-Tropsch process should be used and hydrocracking of wax which is formed
in the reactor should be applied. This can lead to conversions of 90 % with a diesel yield of 80 %.
Algal oil is produced by cultivation of microalgae that accumulate oil in a bioreactor with a following
extraction step. Carbon dioxide can be used as the carbon source, and electricity to produce artificial
lighting if needed. The algal oil can be used for e.g. biofuels and food supplements, but is in need of
additional process steps before it can be used as such. Depending on the algae strain used for
production, the productivity can reach 142 mg/l day.
The classic way of producing methanol is by feeding synthesis gas to a reactor loaded with a
Cu/Zn/Al2O3 catalyst. However, new ways of producing methanol directly from carbon dioxide have
surfaced during recent years, using much the same catalyst with almost identical operating
parameters as the classical ways. This eliminates the need of using Reversed Water Gas Shift to
produce synthesis gas, making the process simpler and cheaper but still with a high conversion (95 %)
and selectivity (99.8 %).
The conclusion of the literature review is that the carbon dioxide to methanol process should be
investigated further. This is due to the high conversion and selectivity of the process, as well as it
being a rather simple process without the need of producing syngas in an extra process step,
compared to the diesel and the classic methanol processes. For similar reasons, the carbon dioxide to
methanol process is recommended instead of the algal oil process because the latter has a rather low
productivity and is more complex with need of more steps of purification, e.g. dewatering and oil
extraction. Therefore, the carbon dioxide to methanol process is chosen to be further examined.
The designed process has three pressure levels, at 1 bar, 30 and 100 bar. The hydrogen produced in
an electrolyser enters the process at 30 bar whereas the carbon dioxide enter the process at 1 bar
and is compressed to 30 bar before mixing with hydrogen and a recirculation stream. The mixture is
then compressed in a multi-stage compressor to 100 bar. The feed is pre-heated, using the product
gas stream leaving the reactor, and is then fed into the exothermic reactor which is cooled using
Dowtherm oil G, in turn producing steam at 1 bar. The product gas stream is cooled in two additional
heat exchangers, where additional steam is produced. After the heat exchangers, the product gas
stream enters a flash vessel at 30 bar where gas and liquid is separated and the gas stream is
recirculated. The liquid product stream is then fed into another flash vessel at 1 bar before being
distilled. The distillation process is divided into two columns. In the first column, light-end by-
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products (e.g. gases) are separated from the crude methanol. In the second distillation column,
methanol is separated from water.
Assuming 14 MW of electricity and 126 MW of heat (which is being converted via an Organic Rankine
Cycle to 9 MW of electricity), 3,393 tonnes of hydrogen can be produced and consumed annually
(based on 8,000 operating hours), with an assumed unlimited supply of carbon dioxide. The
simulations yield that 15,240 tonnes of methanol is produced annually with a conversion of the
hydrogen at 85.5 %; the remaining 14.5 % leaves the process with the purge.
The economic analysis is based on Ulrich’s method to find the grass-roots capital of the plant. The
production cost is calculated using the annuity method with a depreciation time of 15 years and an
interest rate of 10 %. The economic analysis yields a price of 1,009 USD/ton methanol which is well
above the market price as of the first quarter 2013 of 480 USD/ton. The sensitivity analysis indicates
that the production cost is sensitive to deviations in the interest rate used for calculations, and the
price of the Organic Rankine Cycle and electrolyser machines, leading to the conclusion that both of
these should be kept as low as possible.
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Table of Contents 1. Introduction.......................................................................................................................... 1
2. Processes .............................................................................................................................. 2
2.1. Synthesis Gas Production ........................................................................................................ 2
2.1.1. Generation of Hydrogen .................................................................................................. 3
2.1.2. Organic Rankine Cycle ..................................................................................................... 5
2.1.3. Reversed Water Gas Shift ................................................................................................ 7
2.2. Diesel Production - The Fischer-Tropsch synthesis ................................................................. 9
2.2.1. Fischer-Tropsch product distribution .............................................................................. 9
2.2.2. Catalysts......................................................................................................................... 10
2.2.3. Reactors ......................................................................................................................... 11
2.2.4. Product upgrading ......................................................................................................... 14
2.3. Algal-oil Production ............................................................................................................... 15
2.3.1. Microalgae Cultivation .................................................................................................. 15
2.3.2. Separation of microalgae .............................................................................................. 16
2.3.3. Species ........................................................................................................................... 19
2.4. Methanol Production ............................................................................................................ 21
2.4.1. Production ..................................................................................................................... 21
2.4.2. Catalysts......................................................................................................................... 22
2.4.3. Processes ....................................................................................................................... 23
2.5. Choice of Process ................................................................................................................... 26
3. Process design .................................................................................................................... 28
3.1. ORC and Electrolyser ............................................................................................................. 28
3.2. Compressors .......................................................................................................................... 29
3.3. Reactor and steam generation .............................................................................................. 30
3.4. Flash vessels .......................................................................................................................... 31
3.5. Distillation columns ............................................................................................................... 32
3.6. Heat exchanger network ....................................................................................................... 33
3.7. Dimensioning of unit operations ........................................................................................... 35
3.8. Process efficiencies................................................................................................................ 36
4. Economical evaluation ........................................................................................................ 36
4.1. Ulrich’s method ..................................................................................................................... 36
4.2. Results of economic evaluation ............................................................................................ 37
4.2.1. Capital costs ................................................................................................................... 37
4.2.2. Operating costs .............................................................................................................. 38
4.2.3. Total production cost .................................................................................................... 39
4.3. Sensitivity analysis ................................................................................................................. 39
5. Discussion and conclusions .................................................................................................. 43
6. Acknowledgments .............................................................................................................. 45
7. References .......................................................................................................................... 46
Appendix A – Flow sheet ............................................................................................................... I
Appendix B – Design and dimensioning ........................................................................................ II
Appendix C – Economy .............................................................................................................. VII
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1. Introduction At several places around the world, geothermal and other renewable energy plants are located so
remotely that the direct use of produced electricity and heat is hindered. An alternative to construct
new power lines is to convert the produced energy into a chemical component which easily can be
shipped or carried by trucks to the customer.
At the same time, the political developments that are reported in news media indicate that the taxes
on carbon dioxide emissions are to be raised in the future. Because of this, carbon capture might be
considered to eventually become worthwhile, which is why carbon dioxide is likely to become a
cheap feedstock available in large quantities in the future. It would therefore be of interest to
investigate the possibility and the profitability to convert carbon dioxide into chemical components
by using the remote energy.
This prospect has laid the foundation for this literature review, which, with eyes on the horizon, aims
to give an overview of how to produce diesel, algal oil or methanol by using carbon dioxide,
electricity, and heat. The study was assigned as a project from the Norwegian company Statoil, who
also recommended the aforementioned processes. In the literature review, the three different
processes will be described and compared with regards to the traditional selection rules, e.g. yield to
product and complexity of the process. One process will then be chosen to be designed, simulated
and analysed from an economic perspective.
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2. Processes Figure 2.1 is presented below to give an overview of the processes discussed in this review.
Fischer-Tropsch process
Algal oil production
DieselWax
Petrol
Reversed Water Gas Shift
Synthesis gas
Carbon dioxideNutrientsSun light
Algal slurry
Algal oil
Product upgrading
DieselPetrol
Hydrogen
Electrolyser
Carbon dioxide
Water
Methanol process
Crude methanol
Reversed Water Gas Shift
Synthesis gas
Product upgrading
Methanol
Hydrogen
Electrolyser
Carbon dioxide
Water
Product upgrading
Figure 2.1: A schematic overview of the review’s content.
Essentially, the three different products give three different “trees” of processes to be discussed. The
diesel is produced in the Fischer-Tropsch process from synthesis gas, whereas methanol is
conventionally produced from the same raw material but in a different process. Algal oil, on the
other hand, is produced by cultivating microalgae and harvesting the oil they accumulate during their
lifespan. In this chapter, the different processes presented in the figure above are described. Since
synthesis gas is a common intermediate for both the methanol and the Fischer-Tropsch processes,
the review will begin with a description of how synthesis gas can be produced from carbon dioxide
and water.
2.1. Synthesis Gas Production Synthesis gas, or syngas, is a gas mixture of hydrogen and carbon monoxide which is used for
industrial production of e.g. methanol and ammonia, but also as an important intermediate in
production of different synthetic fuels, such as Fischer-Tropsch diesel. Syngas is mostly produced via
catalytic steam reforming of methane, but can also be produced through gasification of coal or
biomass. If none of the traditional raw materials for syngas production is available, syngas could be
produced by first generating hydrogen gas through electrolysis of water which is then sent to a
reversed water gas shift (RWGS) reactor together with carbon dioxide.
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2.1.1. Generation of Hydrogen
The electrolysis process where water is decomposed to oxygen and hydrogen has been used for
hydrogen production for more than 100 years, and today circa 5 % of the total hydrogen production
origins from electrolysis of water (1). Two different kinds of electrolysers exist on the market today:
alkaline and polymer electrolyte membrane (PEM) electrolysers (2).
Alkaline electrolysis
Figure 2.1.1 shows the principal construction of an alkaline electrolysis cell. Anode and cathode
regions are separated from each other by a diaphragm and electrons are transported from the anode
to the cathode when a potential is applied to the electrodes. Oxygen is formed at the anode and
hydrogen is formed at the cathode according reactions 2.1.1-3:
Cathode: (2.1.1)
Anode:
(2.1.2)
Total cell reaction:
(2.1.3)
Figure 2.1.1: Shows the principal construction of an alkaline electrolysis cell (3).
The electrolyte in conventional alkaline electrolysers typically consists of a water solution of KOH (25-
35 wt %) but also water solutions of NaOH and HCl can be used. Pure water has a relatively low
conductivity and can therefore not be used as an electrolyte. The conductivity of conventional
electrolytes (and thereby also the productivity of the electrolyser) increases with increased
temperature, which is why the electrolysers are often run at temperatures just below the boiling
point of the electrolyte. If the electrolyser is operated under higher pressures, the boiling point of
the electrolyte increases, which enables higher temperatures and production rates. The electricity
consumption of pressurized electrolysers is not significantly higher than the electricity consumption
of those designed for operating at ambient pressure, and pressurized electrolysers also save
equipment and energy for hydrogen production as a whole, since the hydrogen often has to be
compressed in later process steps anyway (1).
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Because of the corrosive environment in an alkaline electrolyser, specific construction materials are
required. Carbon steel with coatings of plastics, ceramics and nickel in especially exposed areas are
commonly used in conventional electrolysers. The diaphragm was previously made from asbestos,
but the manufactures have replaced, or are planning to replace the asbestos with synthetic polymers
due to the health hazards associated with the asbestos (1).
Depending on the arrangement of the electrodes, the electrolysis cell can be divided into two
subcategories: unipolar or bipolar. In a unipolar cell, the anode and cathode have their own separate
cell region. In a bipolar cell however, the cells are arranged in series and each electrode serves as
anode on one side and as cathode on the other side. Most of the electrolysers today are made to run
in a bipolar mode (4).
An electrolyser typically has a water consumption of 0.8 l/Nm3 H2 and the efficiencies are typically in
the range of 70 – 80 % based on the higher heating value of hydrogen. The power consumption is
often in the range of 4-6 kWh/Nm3 H2. The productivity of an alkaline electrolyser varies largely
between different manufacturers. In Table 2.1.1 some industrially available electrolysers are listed
(4).
Table 2.1.1: Conventional alkaline electrolysis systems (4).
Lurgi system (ELT/GTec)
MTU Teledyne Hydrogenics Norsk Hydro ASA
ABB & Cie
Cell type Bipolar Bipolar Bipolar Bipolar Bipolar Bipolar Operating Pressure (bar)
30 30 7 10/25* Atmospheric Ambient
Operating Temperature (oC)
90 130 80 - 80 80
Electrolyte 25 % KOH 30 % KOH KOH 30 % KOH 25 % KOH 25 % KOH
Current density (A/m2)
2000 7000-10000 3000 - - 2000
Cell voltage (V) 1.86 1.65 – 1.8 - - - 2.04
Current efficiency (%)
98.75 >99.5 - - 99.9 99.9
Power consumption (kWh/Nm3)
4.3 – 4.65 4-4.4 5.6 4.2 4.3 4.9
Maximum production (Nm3/h)
760 - 140 60 100 -
*: Hydrogen output pressure
To avoid a build-up of substances in the electrolyte solution, which may destroy the electrodes or
cause corrosion, the feed water has to be carefully purified before being mixed with the lye. The
required water purity (conductivity ≤ 1µS/cm) is obtained by the use of ion exchangers (1).
Necessary purification steps are determined of the purity requirements of the produced hydrogen.
The gas leaving the electrolyser contains lye residuals from the electrolyte and also small amounts of
oxygen (about 1 %). The lye residuals can be removed with a gas scrubber and the oxygen can be
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removed with a catalytic deoxidizer, where hydrogen and oxygen are reacted to form water. The
reaction is strongly exothermic and the catalyst used is often platinum or palladium. After
deoxidation, the gas contains a large amount of water which can be removed by absorption or
condensation (4). However, if the purity requirements of the product gas are low, the purification
step may just consist of a simple condenser to remove present moisture (1).
Polymer Electrolyte Membrane Electrolysis
A PEM electrolysis cell consists of a proton conducting membrane placed between an anode and a
cathode (see Figure 2.1.2). The most commonly used membrane consists of a perfluorosulfonic acid
polymer, the anode is usually made of titan with coatings of a noble metal catalyst such as iridium or
ruthenium, and the cathode is made from a porous carbon element such as graphite with catalyst
coatings of platinum.
Figure 2.1.2: Shows the principal construction of a PEM electrolysis cell (5).
Purified water is circulated in flow channels connected to the electrodes, and as a potential is
applied, water dissociates to oxygen and protons at the anode. The protons are then transported
through the membrane to the cathode together with some water molecules where they are
converted to hydrogen. The formed oxygen and hydrogen is then carried away by the circulating
water to water-gas separators (6).
PEM electrolysis is often used in smaller scales and the largest PEM electrolysis unit at present is
FuelGen® Hydrogen Fuelling System with a capacity of 6.5-6.8 Nm3 H2/h. The electricity consumption
of a PEM electrolysis stack is typically in the range of 4.2-5.5 kWh/Nm3. One of the advantages with
the PEM electrolysis compared to the alkaline version is that lower maintenance is required due to
the fact that no refill of electrolyte have to be made. Since the electrolyte membrane has a very low
permeability to gases, it is also possible to generate hydrogen with a purity greater than 99.99 %
without additional purification steps. One of the major drawbacks with the PEM electrolysers is their
high costs due to the fact that noble metals are used as catalyst coatings on the electrodes (6).
2.1.2. Organic Rankine Cycle
The organic Rankine cycle (ORC) is a way to utilize low-temperature heat to generate electricity,
which could be used in hydrogen production in order to beget syngas. The basic principle is to use
the low-temperature heat to evaporate a low-boiling, organic working fluid that is passed through a
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turbine to generate electricity. Typical temperatures are below 100 oC, and typical evaporation and
condensation pressures are 5 and 1 bar respectively, although to a degree dependent on the choice
of working fluid. A basic flow-sheet for the ORC is shown in Figure 2.1.3 (7).
Figure 2.1.3: A flow-sheet for ORC (7).
As can be seen above, the working fluid is passed through a turbine to generate electricity and is
then condensed using cooling water before it is evaporated once again; much like any cyclic process.
For a medium-temperature ORC (100-220 oC) the total efficiency is approximately 10 % but for a low-
temperature ORC (< 100 oC) the efficiency can be as low as 5-9 % which means that only 5-9 % of the
heat can be converted to electricity. Another down-side with the low temperature process is that the
heat exchanger areas must be large due to the low temperature difference between the heat source
and the working fluid (7). This is why the choice of working fluid is of high interest. There are several
preferable features for the working fluid including a low boiling point, high heat of vaporization. The
fluid should be non-toxic, flammable, and not prone to fouling or have any negative effect on the
ozone layer (8).
Another interesting parameter to consider is the appearance of the fluid’s T-s diagram; should the
saturation curve have a positive slope the fluid is called wet, if the line is straight the fluid is called
isentropic, and if the slope is negative the fluid is a dry fluid. It is desirable to use an isentropic or dry
fluid due to the fact that there will be no condensation in the turbine, since this will keep the turbine
blades from corroding (8). In Figure 2.1.4, a T-s diagram is shown for a number of plausible working
fluids (7).
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Figure 2.1.4: A T-s diagram for a number of working fluids (7).
Figure 2.1.4 shows that ammonia is a wet fluid and therefore not appropriate unless a high pressure
is used, this is not common in low-temperature applications. The remaining fluids, PF5050, HCFC123
and n-Pentane have isentropic or dry behavior but HCFC123 has negative effect on the ozone layer
and n-Pentane is highly flammable which is why PF5050 seems a good choice even though better
efficiencies can be reached with HCFC123 and n-Pentane (7).
2.1.3. Reversed Water Gas Shift
In the water gas shift (WGS) reaction, carbon monoxide and water forms carbon dioxide and
hydrogen gas in an exothermic, equilibrium reaction as shown in reaction 2.1.4.
(2.1.4)
The reaction is mostly utilized when a higher hydrogen share in a syngas is desired. Since the reaction
is reversible it is also possible to generate carbon monoxide and water from carbon dioxide and
hydrogen. However, this is not done to any great extent in industrial applications. Since the WGS
reaction is exothermic, the reversed reaction is endothermic and therefore supported by higher
temperatures. An expression that describes the temperature dependence of the equilibrium
constant that is often found in literature is shown as equation 2.1.1 (9).
(Eq. 2.1.1)
The forward WGS reaction is often conducted in two or more catalytic reactors. The temperature is
higher in the first reactor (350-600 oC) than in the following, low temperature reactors (150-300 oC),
and different catalysts are used depending on temperature (10). In the shift reactor, the side
reactions 2.1.5-7, which governs solid carbon deposition, can occur:
(2.1.5)
(2.1.6)
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(2.1.7)
Carbon deposition is undesirable since it will cause deactivation of the catalyst through fouling.
According to the reactions above, carbon deposition can be suppressed by increasing steam,
hydrogen, and carbon dioxide concentrations (11).
Catalysts
The reversed water gas shift reaction is favoured by higher temperatures, which is why a high-
temperature WGS catalyst should be chosen in this case. Two possible candidates are described
below.
Iron-Chrome
In industrial high temperature WGS, a catalyst combination of Fe3O4, which is the stable iron phase
under reaction conditions, and chrome is commonly used. The composition is typically 55 wt-% Fe3O4
and 6 wt-% chrome. The catalyst is normally not exposed to temperatures above 600oC in the high
temperature WGS reaction, but due to its stability it is likely that it can be exposed to even higher
temperatures without any risk of sintering. The iron-based catalyst is relatively cheap, but the
activated catalyst is pyrophoric and must therefore be stabilized via surface oxidation using an inert
gas, in the event of the catalyst having been exposed to air (10).
Many studies on the kinetics of different catalysts active for the WGS-reaction have been performed,
but there is still no consensus on which mechanism the reaction follows. A power-law rate model is
although often proposed, and the rate expression presented as equation 2.1.2 has been suggested
for a commercial iron-chromium based catalyst at 450o C (12):
(
)
) (Eq. 2.1.2)
Where r is the reaction rate in moles/(g cat s), R the gas law constant, T is the temperature in Kelvin,
pi is the partial pressure of component i in kPa and K is the equilibrium constant. The reaction rate of
the reversed reaction, as seen in equation 2.1.3, can be approximated from the rate equation for the
forward reaction and the equilibrium constant.
(Eq. 2.1.3)
Shale Ash
The catalytic activity of shale ash to the forward and reversed water gas shift reaction has also been
investigated. The shale ash, with a composition of 85.5 wt-% ash, 10.9 wt-% carbon, 0.2 wt-%
hydrogen, 0.1 wt-% nitrogen and 4.3 wt-% sulphur, was obtained from an oxygen blown pilot plant
gasifier, and the kinetics was studied in a temperature range of 700-800oC. The reaction rate was
fitted to a power-law expression that is given as equation 2.1.4 (13).
(
)
(Eq. 2.1.4)
Where r is the reaction rate in moles/(kg cat s) and pi is the partial pressure of component i in Pa. A
comparison between the reaction rates of the Fe/Cr catalyst and the shale ash shows that the Fe/Cr
gives higher reaction rates. The shale ash is more likely to endure higher temperatures than the
Fe/Cr catalyst. The catalyst has been reported to sinter first when temperatures exceed 800oC (13).
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2.2. Diesel Production - The Fischer-Tropsch synthesis In the Fischer-Tropsch synthesis (FT synthesis), synthesis gas is converted to a wide range of
hydrocarbons (e.g. diesel) in a catalytic, exothermic reaction that often is simplified by reaction 2.2.1.
(2.2.1)
The FT process was discovered in 1923 by the two German researchers Franz Fischer and Hans
Tropsch, and it has been used from time to time, especially at places where the availability of
petroleum has been low or restricted. The interest of the FT process has risen during the past ten
years due to the increasing debate of future diminishing petroleum reserves, and due to the fact that
carbon dioxide-neutral fuels can be produced if the synthesis gas is generated from biomass. Fuels
produced in a FT process gives lower emissions of NOx, SOx and particles than petroleum based fuels
since they are essentially free from nitrogen, sulphur and aromatics. Diesel that is produced in a FT
process also has a large advantage compared to other second generation biofuels since it can be
blended with conventional diesel in any proportion (14).
2.2.1. Fischer-Tropsch product distribution
The FT synthesis is a polymerization process in which the building block, or monomer, is the CH2-unit.
The synthesis occurs as a stepwise chain growth reaction of CO and H2 on the surface of a catalyst.
Due to the stepwise reaction mechanism, the products formed in the synthesis will consist of
hydrocarbons of various lengths. The product distribution can be described by the Anderson-Schulz-
Flory-theory (equation 2.2.1), where Wn is the weight fraction of a product containing n carbon
atoms and α is the chain growth probability which is defined as equation 2.2.2 (15).
(Eq. 2.2.1)
(Eq. 2.2.2)
The chain growth probability depends on the rates of chain growth propagation, rp, and the chain
growth termination, rt, which in their turn depends on choice of catalyst, reactor temperature,
pressure and feed gas composition (14).
When diesel is the desired product of the FT synthesis, it is of interest to generate a greater fraction
of long chained hydrocarbons (in this context longer chains of hydrocarbons are often referred to as
C5+) compared to shorter hydrocarbons such as methane. It is well known that the highest yield to
diesel is achieved by first making a wax, which then is hydrocracked into a diesel fraction. Figure
2.2.1 displays the product distribution as a function of the chain growth probability. It can be seen
that a maximized production of wax and, thereby also diesel, will be achieved when α = 0.85 – 0.9.
Such an alpha value corresponds to a selectivity to C5+ of 78 %. High alpha values can be reached if
the pressure is high and the temperature and the inlet H2/CO ratio is low. The alpha value is also
affected by the choice of catalyst (14).
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Figure 2.2.1: Product distribution according to the ASF-theory (14).
2.2.2. Catalysts
Four different metals can be used to catalyse the FT synthesis; cobalt, iron, nickel and ruthenium.
Nickel has however a too high selectivity to methane and the availability of ruthenium is low, why
the only feasible metals for FT synthesis are cobalt and iron (14).
Iron catalysts are cheap and have high selectivity to light hydrocarbons, which makes it appropriate
to used when petrol is the desired product. The catalyst can lose activity through oxidation by water
vapour, which is present in the synthesis, and by sintering. The lifespan of iron catalysts is
approximately three years (15).
Cobalt catalysts are more expensive than iron catalysts but have high selectivity to heavier
hydrocarbons, making it a feasible catalyst when diesel is the desired product. The cobalt catalyst
also gives higher conversion rates than the iron catalyst, and it is not as easily oxidized by water
vapour as the iron catalyst which gives it a longer lifespan (approximately five years). However, the
cobalt catalyst is more sensitive to sulphur compounds than the iron catalyst (15). Both iron and
cobalt catalysts can be deactivated by carbon deposition (reaction 2.1.5) which will occur more
frequently if an iron catalyst is used (15).
In the FT synthesis, hydrogen and carbon monoxide is consumed in a molar ratio of H2/CO = 2. If this
molar ratio is not attained, it can be regulated with the water-gas shift reaction (reaction 2.1.4). If an
iron catalyst is used the WGS-reaction takes place simultaneously with reaction 2.2.1, making it
possible for the H2/CO ratio to be adjusted inside the reactor. This can be an advantage if the
synthesis gas fed to the reactor has a low H2/CO ratio, since no investments then have to be made in
external WGS equipment. If the syngas is not low in hydrogen, the cobalt catalyst is to be preferred
since the WGS activity of an iron catalyst otherwise would cause a to high H2/CO ratio (14).
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A relatively new way of structuring the catalyst in a reactor is to use monoliths. The monolithic
catalyst structure has been widely used in gas-solid applications but has recently become interesting
even in gas-liquid-solid processes, like the FT synthesis. The monoliths are ceramic network
structures with parallel channels made of cordierite or catalyst supports as alumina or silica. The
diameter of the channels varies between less than one millimetre up to a few millimetres and the
wall thickness is smaller than 300 µm. The active catalyst is either prepared directly on the monolith
or as a wash coat (15).
In a monolithic reactor the reactant gas is fed to the top of the reactor and then passed through the
catalyst covered monolith. An external heat exchanger is used to remove excess heat from the
reactor since the monolith has poor radial heat transfer properties. The lighter gaseous hydrocarbons
produced in the reactor are separated from the product stream by a gas-liquid separator in the
bottom of the reactor. The external heat exchanger is also used to reduce the temperature gradient
over the reactor which should be held as low as possible (15). Typical properties for the monolithic
reactor are:
No attrition of the catalyst.
A temperature gradient of 10-15 oC over the reactor.
Plug flows inside the monolith channels provides high productivity.
High selectivity is provided by short diffusion lengths in the monolith.
The pressure drop over the reactor is low.
2.2.3. Reactors
Conventional reactors that are used for FT synthesis today can be derived into two main categories:
High Temperature Fischer Tropsch reactors (HTFT) and Low Temperature Fischer Tropsch reactors
(LTFT). Except for the conventional reactors it is also possible to use microstructure reactors for FT
synthesis.
Independent of what kind of reactor that is utilized for the FT synthesis, it can be said that the
control of the temperature inside the reactor is key to reaching high conversion rates. It is of interest
to run the reactor isothermally because of the fact that an isothermal reactor can be operated at
slightly higher temperatures than a non-isothermal reactor. A slightly higher temperature would give
a slightly higher reaction rate which in turn would give a higher productivity. High heat exchanger
rates are therefore desirable to control the temperature in an FT reactor. This is made possible in
different ways in different reactor designs (14).
High Temperature Fischer Tropsch reactors
The HTFT reactors operate at a temperature of 300-350 oC and in a pressure range of 20 to 45 bar.
The catalyst used for these reactors are Fe-based and they are mainly used for production of petrol.
For the HTFT process there are two possible reactor choices: the Circulating Fluidized Bed (CFB)
reactor or the Fixed Fluidized Bed (FFB) reactor (14).
The Circulating Fluidized Bed reactor
In a CFB reactor the reactant as well as recycled syngas enters the bottom of the reactor. The syngas
entrains the iron-catalyst and carries the catalyst through the reaction zone, where excess heat is
removed by cooling coils. The product gas is then separated from the catalyst which is then recycled
to the CFB reactor (15). In Figure 2.2.2 below a simple form of a CFB reactor is shown.
12
Figure 2.2.2: A simple view of the two types of HTFT reactors. The CFB reactor is shown to the left (16) and the FFB reactor is shown to the right (17).
The Fixed Fluidized Bed reactor
In the FFB, the reactant gas is bubbled through a fluidized bed consisting of an iron catalyst. The
reactant gas is catalytically converted to hydrocarbons in gaseous phase. Before leaving the reactor
the product and unconverted gases passes through internal cyclones that very effectively retains the
catalyst in the reactor (17). To remove excess heat in the reactor cooling coils are immersed in the
catalytic bed (15) (17). A schematic picture of the FFB reactor is shown in Figure 2.2.2 above.
The Low temperature Fischer-Tropsch reactors
The LTFT reactors operate at temperatures between 200 and 240 oC and in a pressure range of 20 to
45 bars (14). Conventional reactors usually have a conversion of 90 % with a yield of 80 % to diesel
(16). The catalyst used for these reactors are either based on cobalt or iron. In this reactor type a
high selectivity to high molecular linear waxes can be achieved which means that the LTFT reactors
can be used to reach a high selectivity to diesel fuels if the waxes are hydrocracked (14). There are
two commercial alternative LTFT reactors, the multi tubular fixed bed (MTFB) reactor and the Slurry
Phase Reactor (SPR).
The Multi Tubular Fixed Bed reactor
In this reactor, the FT synthesis takes place in a bundle of tubes packed with catalyst, consisting of
either iron or cobalt. The reactant gas feed enters at the top of the reactor and the products exits at
the bottom. Excess heat that is generated during the synthesis is used to produce steam by letting
the heat move through the tube walls of the reactor (17). A schematic picture of the MTFB reactor is
shown in Figure 2.2.3.
This kind of reactor is complex, expensive, and difficult to mechanically scale up. The pressure drop
over the reactor is rather large and the removal of used catalyst is difficult (17).
13
Figure 2.2.3: A simplified view of the two LTFT reactors. The MTFB-reactor is shown to the left and the SPR is shown to the right (18).
The Slurry Phase Reactor
The slurry used in the slurry phase reactor consists of a solid phase catalyst (cobalt) which is
dispersed in wax produced in the FT synthesis. The feed enters the bottom of the reactor and is
bubbled through the slurry where it also is converted to hydrocarbons. The heat generated from the
reaction is passed through the slurry to cooling coils inside the reactor and steam is generated. The
lighter hydrocarbons exit at the top of the reactor together with the unconverted reactant. The
heavier hydrocarbons (diesel and wax) dissolve in the slurry and have to be separated outside the
reactor (15).
A significant process characteristic for the SPR is that it is well mixed and can therefore be operated
virtually isothermally. This makes it possible for the SPR to operate at higher temperatures than the
MTFB reactor which means that higher reaction rates can be achieved with this reactor
configuration. Since the catalyst operates at more optimal process conditions, the yield per reactor
volume for the SPR is higher than in the MTFB reactor even though the catalyst concentration is
lower. Because the reactor is well mixed, a disadvantage in terms of conversion exists in a once
through system, but by operating a number of SPR in series much of this problem can be reduced.
The pressure drop over the SPR is lower than over the MTFB reactor and it is also easier to remove
used catalyst in a SPR than in a MTFB reactor. Because of the isothermal conditions and the smaller
pressure drop in the SPR, the control of the reactor is simple and the operating costs are lower than
for a MTFB reactor. The design of a SPR is also simpler which reduces the investment cost (15).
Microchannel Reactors
A rather new way of producing FT fuels is through so called microchannel reactors, which are
compact reactors with channel diameters in the millimetre range. The small dimensions of the
reactor system is minimizing heat and mass transfer limitations and thus increasing the reaction rate
(19).
14
Conventional FT plants are designed to produce a minimum of 5000 barrels/day which can be
compared to a single microchannel reactor block that produces 30 barrels/day (19). Microchannel
reactors are therefore feasible for production of FT fuels in smaller scale.
A microchannel FT reactor can consist of several reactor blocks, each having thousands of thin
process channels filled with FT catalyst, which are interleaved with coolant channels. The fact that
the reactor consists of several blocks or modules makes it easy to scale up, and the whole process
does not have to be shut down when the catalyst has to be regenerated. The flexible construction
and the smaller size of the reactor is also causing the operation and investment cost to be low (19).
In a microchannel FT reactor it is possible to reach conversion efficiencies of 70 % per pass which can
be compared to 50 % that can be achieved by conventional FT reactors. The productivity is also much
higher in a microchannel reactor compared to a conventional reactor. With the right catalyst it is
possible to reach productivities of 1500 kg/m3h which can be compared to 200 kg/m3h that can be
achieved with a slurry phase reactor (19).
2.2.4. Product upgrading
In the FT synthesis, as declared in previous chapters, a wide range of hydrocarbons are formed.
Lighter hydrocarbons, water, CO2 and unreacted syngas exit the top of the reactor as a gaseous
mixture which can be recycled, burned for heat generation, or fed to a gas turbine for electricity
production (20). The heavier hydrocarbons produced in the FT reactor can be divided into four
different groups: naphtha (C5-C9), diesel (C10-C19), wax (C20+) and kerosene. The heavier hydrocarbons
have to be separated from each other as presented in Figure 2.2.4 (15).
Figure 2.2.4: A schematic overview of the necessary upgrading steps after the FT reactor.
At first, the wax is separated from the other hydrocarbons in a distillation column. The wax, which is
found in the bottom of the column, is then sent to a hydrocracker. Hydrocracking is a catalytic oil
refinery process in which C-C bonds are broken in the presence of hydrogen, causing heavier oil
fractions to be converted into lighter products such as naphtha, kerosene and diesel (14).
Hydrocracking is an endothermic process in which almost 100 % of the feed can be converted to
lighter hydrocarbons. It is possible to reach selectivity to diesel of 60 wt-%. A hydrocracker normally
operates at a temperature of 400o C (20).
15
The naphtha, kerosene and diesel leaving the hydrocracker is mixed with the distillate from the wax
distiller and is then separated from each other in two further distillation column. In the first column
the naphtha fraction is found in the top of the column while the bottom product is sent to a second
distillation column where diesel is the bottom product and small amounts of kerosene is found in the
top of the column. A naphtha fraction is obtained as a by-product in the production of FT diesel. Due
to the low octane number of naphtha, it is not possible to use as petrol. It can however be upgraded
to petrol but this requires extensive isomerization and reforming (14).
2.3. Algal-oil Production Microalgae are organisms that, just like land based plants grow via photosynthesis, utilizing carbon
dioxide and sunlight. Due to their unicellular structure they can easily produce chemical bound
energy such as oil from solar energy. This algal oil can be used to produce a wide spectrum of
chemicals, e.g. biofuels and food supplements (21).
2.3.1. Microalgae Cultivation
The fundamental requirements for algae cultivation are solar light and carbon dioxide, without these
there will be no cultivation. Therefore, the most important parameters for the cultivation for algae
are the supply of light for energy and the food supply of carbon dioxide. When taking the cultivation
of algae to an industrial level other parameters are important, especially for closed system. These
parameters are the oxygen oversaturation as well as the temperature and pH-levels in the reactor.
To efficiently cultivate algae a fix understanding of these parameters is therefore key (22) (23) (24).
The use of light as an energy source for cultivation is the principal limiting step for photobioreactors
(PBR) (23). Therefore the utilization of light in the reactors is very important. Problems arise when
the light intensity is too high and inhibits the growth of algae. A way to deal with this photo-
inhibition or light saturation effect is therefore important (22) (23). Different ways of controlling the
light intensity consists of partially shading the reactor or using fibre optics and artificial light as
mentioned earlier (22).
Through the photosynthesis the algae produces oxygen. In the closed PBR the build-up of oxygen can
lead to oversaturation which in turn is harmful for the algae, lowering the overall productivity of the
reactor (22) (23). An efficient way of removing oxygen from the system is therefore paramount.
The control of pH is important to maintain productivity in the reactor. For the algae to continuously
grow, additional carbon dioxide must be introduced to the system, but this lowers the pH of the
system and could affect the productivity. Although the respiration and consumption of carbon
dioxide increases the pH, a way of controlling the pH is important (22) (23).
Lastly the temperature control of the system is dependent of the species of algae used. Different
algae have different optimal temperature intervals where they thrive. It is therefore important to
make sure that adequate level of heating or cooling is possible. Low temperatures may slow or even
stop the algae growth whereas high temperatures can cause the culture to collapse (22).
There are different types of systems to cultivate microalgae, whereas the two main categories are
the open and the closed systems respectively.
16
The Open System
For the open system the cultivation of algae is done in either natural water such as lakes, lagoons
and ponds or artificial ponds and containers. The main advantages of the open system are that they
are simple to construct and therefore the initial cost for this kind of process is quite low. Also, they
are simpler to operate than most closed systems (25). However the open system has some major
drawbacks that overshadow the relative ease of construction and process operation. These include
poor light utilization by the cells, the loss of carbon dioxide by diffusion to the atmosphere,
significant evaporation loses and the constant threat of contamination and pollution of the system
(25) (23). Beyond these drawbacks the land needed should also be considered as the open systems
tend to be large. Because these systems have poor mixing the mass-transfer rates are low, which
results in low algae productivity (25).
The Closed Systems
The closed systems are characterized by the ability to regulate and control nearly all biotechnical
parameters for algae cultivation as well as a very low risk of contamination (23). The closed system
consists of a number of reactors in series. For the closed PBRs there are a number of different
suggested reactor designs. These are the flat-plate and tubular PBRs (24; 25).
The Flat-plate Photobioreactor
The flat-plate PBR are designed to achieve large illumination areas (25), they are therefore built as
narrow panels. Because the panels are narrow the light path for these kinds of reactors are narrow,
which is advantageous since a narrower path will induce lesser shading problems. Mixing in this
reactor is done by compressed air which is supplied from a perforated plastic tube in the bottom of
the reactor (24). The mixing makes the build-up of oxygen easier to control, however this kind of
mixing may also lead to a possibility of hydrodynamic stress to some algae species (25).
The Tubular Photobioreactor
Like the flat-plate reactor, the tubular reactor has a high illumination area, and because of this they
are suitable for outdoor production of algae. The reactor is usually made out of glass or transparent
plastic tubes. The form of the tubular reactor can be horizontal, conical, vertical or inclined. Mixing
and aeration of the reactor is usually done by either an air-pump or airlift system (24; 25). A
disadvantage with the tubular reactor is poor mass transfer. The poor mass transfer or oxygen build-
up can become a problem when the reactor is scaled up. Another problem for the tubular reactor
during scaling is the decrease of its illumination area which could lead to lower productivity. Beyond
the decrease of illumination area, gradients of oxygen and carbon dioxide transfer along the tubes
may occur during scale-up (25).
The Internally Illuminated Photobioreactor
Some Photobioreactors can also be internally illuminated by a fluorescent lamp. These reactors can
be built to use both solar and artificial light, which could be used to supply the reactor with light
during both day and night cycles. The use of optical fibres has also been suggested to distribute solar
light into reactors to better control the light intensity (25).
2.3.2. Separation of microalgae
After the microalgae are passed through the reactor there is a need of separating the algae from
water to make further processing possible. After dewatering, the oil content must still be extracted
from the microalgae and be separated from the remaining biomass.
17
Dewatering
After the reactor, the algae concentration is often low; typically 0.5 g/l in an open pond and 5 g/l in a
photo-bioreactor and this is due to shading problems when the concentration is higher. The low
concentration means that there are large volumes of water in need of separation (26). The
separation, or harvesting, can be done in a number of ways all of which have different advantages
and disadvantages. The separation cost of microalgae can account for up to 25 % of the total
production cost which is why cost-effective dewatering and extraction methods are essential to find
(27).
Flocculation
Flocculation is a well-known technology based on the algae cells’ ability to form larger, coagulated
aggregates when mixed with a compound called flocculent. The larger algae aggregates are then
easily separated using gravity sedimentation as a second separation step. There are a number of
different techniques for flocculation including chemical flocculation, physical flocculation or bio
flocculation; the latter two are both relatively novel ideas under research.
Chemical flocculation is the most common flocculation method based on the naturally occurring
negative charge of the microalgae cell wall (21). If the negative charge is removed or reduced, the
algae will form larger aggregates. The additive (flocculent) is often a metal salt such as aluminium or
ferric chloride, both of which have been successfully used in different industries including water
treatment, for separating un-desirable particles from a desired product (26). The flocculent can also
be some kind of cationic polymer such as chitosan which is a very effective flocculent although only
at low pH-values (26). Another possibility to achieve good separation is to change the pH of the
solution which will change the ionization of the cell wall. Sometimes flocculation occurs
spontaneously due to the photosynthetic CO2-depletion that leads to an increased pH-value. This is
called auto flocculation (26). Auto flocculation can actually be somewhat controlled by interrupting
the CO2-supply (28).
Several new methods to physically flocculate the algae are under research. The large interest in
physical flocculation is due to the fact that there would be no contamination of the algae if no
substances are used. One way of flocculating the algae is to use ultrasound although this is not
feasible for large scale operations (26). Another way would be to use electrolytic release of metal
ions, a method called electro-coagulation flocculation (CEF) although recent research has shown that
an even better way would be to use polarity exchange (PE) to induce flocculation (29), the down-side
of these methods is that there may be some traces of metals after the operation (26). The last way of
physically inducing flocculation is to use magnetic nanoparticles, e.g. Fe2O3, which absorbs on the
surface of the algae cells. The algae cells may then be separated from the water when applying a
magnetic field. This method has several benefits; flocculation and separation can be completed in
one step and the nanoparticles can be separated after harvesting and used once again (26).
A final method to perform flocculation is to use biological substances, i.e. bio flocculation, for
example polymers which occur naturally in lakes and rivers. These polymers are mainly used in plants
focused on waste-water treatment. There are also examples of bacteria and fungi that induce
flocculation which can grow in the same medium as the algae. The down-side of this is that the algae
are biologically contaminated, which results in problems if the algae should be used as feed (26).
18
After flocculation several methods can be used to separate the coagulated algae from water before
further drying processes; centrifugation, filtration or more simple methods such as flotation or
gravity sedimentation (27).
Centrifugation
Another way to achieve separation of algae cells and water is to centrifuge the mixture, a very simple
but effective method since there are no additives in the product. The centrifugation can be used
both as an initial and as a second separation step to achieve higher concentration of the algae if the
first step is some kind of flocculation (26). The down-side of centrifugation is the high cost as well as
the fact that the cells might disrupt due to the high centripetal forces which slows down the speed of
centrifugation (21). The only feasible application of centrifugation is when the end product is on a
high price market and when the quantities are small. When there are large volumes of algae in need
of concentrating the centrifugation is way too expensive (26).
Filtration
The last, but relatively common and competitive method for separation is filtration. There are many
filtration techniques that are usable, including dead end filtration, tangential flow filtration, micro-
and ultra-filtration, and pressure- and vacuum filtration. For relatively large microalgae species the
pressure- and vacuum filtration is often enough but when the size of the microalgae is small other
filtration techniques has to be used. Studies have shown that tangential flow filtration is the most
effective method. The down-side of filtration as separation step is the high operational costs (21).
Extraction of Oil
The final step of the process is the extraction of algae oil which can be done either mechanically with
some kind of aid-chemical to aid the extraction process or via a purely chemical process. After
extraction, oil and biomass are typically separated by gravitation and the biomass is considered a
valuable by-product.
Mechanical extraction
The mechanical extraction processes are complicated, mainly because the internal pressure of the
algae can be up to 20 atmospheres (30) and due to the fact that the algae are in need of drying,
which is very energy intense, to even make the pressing possible. The most common mechanical
extractions methods include screw, expeller and piston presses. There is also a modern method
utilizing an ultrasonic reactor, that creates ultrasonic waves that disrupts the cell walls by creating
bubbles which collapses near the cell in a solvent (31).
Chemical extraction
Another way to achieve oil extraction is to add a solvent chemical. The most common chemicals
include hexane, benzene and ether but hexane is the most popular due to the availability and price
as well as its ability to work well together with the mechanical pressing. The downside using
chemicals is that they might interfere with the product and their toxicity. After extraction the oil and
hexane is separated by distillation. Algae oil can also be extracted using supercritical fluid (CO2).
Carbon dioxide can act both as a liquid and as a gas which extracts oil from the cell after liquefaction
under pressure (31).
19
Outlook
The most promising process to achieve dewatering and extraction of algae and oil in one single step
is currently being developed by the company OriginOil although a patent is still pending. The
OriginOil Single-Step Extraction is based on electromagnetic fields and modification of pH. In one
single step the microalgae cells are disrupted after a CO2-injection which lowers pH-value, a
“Quantum Fracturing” that creates a fluid effect to mechanically stress the microalgae cells before
electromagnetic pulses delivers the force that disrupts the cells. After this treatment the oil, water
and remaining biomass can be separated purely by gravitation sedimentation. The water can be
recirculated to the cultivation and the biomass can be transformed into valuable by-products (32).
This process has proven to be capable of separating five gallons of water per minute which is
equivalent to 94-97 % efficiency (27).
2.3.3. Species
When producing algal oil, the algae in question should be able to withstand variances in the growth
media (temperature, pH etc.) and have high oil content as well as a high growth rate, to grant a high
production rate (22). Starting with high oil content as a selection criterion, the following algae have
been examined closer (33).
Botryococcus Braunii
B. Braunii is a very common colony-forming microalga. It is able to produce a wide range of
hydrocarbons/oils (34) with an oil content of 25-75 % by dry weight (33), mostly comprised of C25-C31
(35) and an oil productivity of about 29-64 mg/l day (36). The colonies’ size is dependent on the light
intensity, and ranges from 0.07-0.15 mm (37). B. Braunii is a freshwater microalga, which means that
it needs desalinated water to thrive (38).
The production of oil is strongly coupled with the growth of B. Braunii. During the exponential and
early linear growth phase the production of oil is maximized, suggesting that the process should be
run at optimal conditions for growth. Compared to other microalgae as Chlorella sp. which produces
hydrocarbons during starvation, this seems an unusual but useful trait for microalgae (39). Optimal
conditions for growth have been determined to be at 30 oC, 0.2-5 % carbon dioxide in air (at 0.33 m3
gas/(m3 liquid minute)) and 850 µmol photons/m2s and a pH of about 6. These conditions have
rendered a specific growth rate of 0.5 /day (a doubling time of 1.4 days), which is the highest ever
reported for this species. The commonly used medium Chu 13 should be used, with the addition of
selenium and vitamin B12 (40).
Nannochloropsis Salina
N. Salina is a halotolerant microalga, meaning it thrives in salt water, which is able to produce oil at a
rate of 61 mg/l day (41). It has been reported to contain between 31-70 % oil of dry weight when
exposed to nitrogen starvation during the mid-logarithmic growth face (33; 38). It has also been
shown that the species can be grown continuously, and extracting cells in the stationary growth
phase shows an insignificant result to the oil content. This is due to the nitrogen deficiency that will
occur naturally during the stationary growth phase (42). This could imply a higher productivity since
the growth does not need to be quenched, but remains unconfirmed.
N. Salina is a unicellular microalga with ellipsoidal form, measuring on average 3.3 µm and 1.9 µm in
length and width respectively (22). During optimal conditions it has shown about the same specific
growth rate as B. Braunii of 0.599 /day. To grow at its maximum rate, N. Salina needs continuous
20
illumination of 120 µmole photons/m2s and 5 % carbon dioxide in air at a flow rate of 4 m3 gas/(m3
liquid hour) and at a temperature of 28 oC (42). It will of course also need salt water.
Chlorella Vulgaris
As its name suggests, C. Vulgaris is a quite common green microalga. It is a unicellular freshwater
alga with a diameter of about 3 µm, and an oil content of 28-32 % of dry weight (33; 43). As opposed
to e.g. B. Braunii, C. Vulgaris is in need of nitrogen starvation in order to accumulate oil and can
produce oil at a rate of 11.2-40 mg/l day (44). The starvation should be performed during 17 days,
since a shorter time could actually result in a decrease of oil. However, a compromise between
productivity and oil content could be reached by increasing the carbon dioxide concentration during
the nitrogen starvation, which could result in continued growth and an increased growth rate
compared to the end of the exponential growth phase (45). Research seems to be focused on the oil
accumulation from nitrogen starvation, and research on optimal conditions for growth is scarce.
Aurantiochytrium Mangrovei
Formerly Schizochytrium Mangrovei, this halotolerant microalga has been reported to contain
between 50-77 % oil (counting by dry weight) and to have a large proportion of docosahexaenoic
acid (DHA), an omega-3 polyunsaturated fatty acid (33; 46). This makes A. Mangrovei of particular
interest when producing algal oil for food supplements. The oil accumulation seems to happen
during the stationary growth phase when the carbon-containing nutrient in the growth media is at
the verge of being exhausted. To maximize the growth, reaching a specific growth rate of 0.15 /h, A.
Mangrovei should be put in salt water containing glucose of about 40 g/l at 30 oC (46). Research on
optimal lighting and about using carbon dioxide as a carbon source is scarce and these points remain
unclear.
Nannochloropsis Oculata
A microalga that might be overlooked when making a selection based on oil content is N. Oculata,
with its relatively low oil content of 22.7-29.7 % of dry weight. However, with a – in comparison –
staggering rate of oil production at 142 mg/l day (44), it could very well be considered a viable
option. This spherical microalga, of about 1-2 µm in size (47), accumulates oil in its stationary growth
phase. For a production in parity with the one specified above, the microalga should be grown with
an aeration with 2 % carbon dioxide at 0.25 m3 gas/(m3 liquid minute), 26 oC and a light intensity of
300 µmol/m2 s. The growth medium should be saline and contain several vitamins as well, such as
the f/2 medium. To optimize the oil production, it has been suggested to grow the alga semi
continuously, replacing half the growth medium once every day (48).
Comparison
To make data more easily overviewed and thereby making it easier to compare the discussed algae,
Table 2.3.1 has been set up.
21
Table 2.3.1: A comparison table for the discussed algae
Alga Halo-tolerant
Oil content Oil productivity (mg/l day)
Growth rate (/day)
B. Braunii No 25-75 % 29-64 0.5 N. Salina Yes 31-70 % 61 0.599 C. Vulgaris No 28-32 % 11.2-40 - A. Mangrovei Yes 50-77 % - 0.15 (/h) N. Oculata - 22.7-29.7 % 142 -
2.4. Methanol Production Methanol is a colourless, polar liquid that is miscible with water and is considered as one of the
world’s most important chemicals. With a production of more than 40 million tonnes (49) and a
growth rate of 4 % each year (50), the need for methanol seems to know no bounds. As can be seen
in Figure 2.4.1 below, methanol is used as a raw material for a wide range of products, including
acetic acid and terephtalic acid (49). In recent years, developments in production of fuels have given
methanol an even wider scope of use (e.g. production of di-methyl-ether (50) and petrol in the
Methanol-to-Gasoline process (18)).
Figure 2.4.1: A selection of reaction pathways with methanol as raw material (49).
2.4.1. Production
Methanol is most commonly produced using syngas. The process is divided into three categories:
low, medium and high pressure. Today, only the low pressure process is used due to economic
reasons (18). The general idea of the methanol synthesis is portrayed in Figure 2.4.2, and it starts
with the feed gas being compressed to 50-100 bar in multiple stages, before being led into the
reactor. The reactor runs at a temperature of about 200-300 oC. Here, reactions 2.4.1-3 take place
(49).
(2.4.1)
(2.4.2)
(2.4.3)
22
When using syngas as reactant, the methanol is mainly formed via equation 2.4.1 which has an
equilibrium conversion about 70-80 % using modern process conditions (18). The product stream
leaving the reactor is at first heat-exchanged with the compressed feed, and then cooled further
before it enters a gas/liquid separator, thus separating the crude methanol from the gases. The latter
can then be recycled, re-joining the feed gas stream at an intermediary pressure (49).
Figure 2.4.2: The general methanol production process. a) Reactor b) Heat exchanger c) Cooler d) Separator e, f) Compressors (49)
In terms of energy conservation, the condensation energy of methanol in the cooler can be used as
well as the heat of the purge gases. However, the methanol leaving the gas/liquid separator contains
not only water, but also by-products such as higher alcohols, hydrocarbons and waxes, esters,
dimethyl ether and ketones, and should therefore be distilled to get the desired product. The by-
products can be divided into the low-boiling light ends, including dissolved gases, dimethyl ether,
methyl formiate and acetone, and the high-boiling heavy ends. The latter include higher alcohols and
ketones, lower esters and long hydrocarbons. The waxes are not categorised as such, but are still
easily extracted from the distillation bottoms, as they are not particularly polar or volatile (49).
2.4.2. Catalysts
Since the reactions taking place are exothermic equilibrium reactions, it is favourable to lower the
temperature of the reactor in order to push the equilibrium to the right hand side. A lower
temperature will also lower the risk of damaging the catalyst via sintering. However, lowering the
temperature too much will lower the reaction rate in an undesired fashion. As such, there is an
optimal process temperature. In a similar way, there is an optimal process pressure. According to Le
Chatelier’s principle, the equilibrium will be pushed to the product side of the reaction when
increasing the pressure. The negative effect of this is largely economical. A higher pressure means a
higher investment cost (e.g. higher-pressure vessels) as well as operating costs (e.g. compressors).
In the low pressure process, copper-zinc oxide-alumina catalysts are considered the most cost
effective and are widely used. Due to the catalyst’s high activity, the process can be run at low
pressure and temperature (50-100 bar, 200-300 oC) which is quite a bit lower compared to the older,
high pressure process’ conditions (250-350 bar, 300-350 oC). The high selectivity, which renders the
methanol at a minimum of 99.5 % purity, in combination with a long technical lifespan and low costs
of production make for very compelling arguments to use the low pressure process with the
23
Cu/ZnO/Al2O3 catalysts. Using this catalyst, the ignition temperature of the reactor would be about
180 oC (51). However, these catalysts are quite sensitive to catalyst poisons such as hydrogen
sulphide and chlorine compounds and require syngas of the purest quality, with the amount of
hydrogen sulphide of 0.1 ppm at the most (49), although this should no longer propose an actual
obstacle with modern technology (52).
As in the Fischer-Tropsch process, it is possible to apply the monolithic catalyst structure (as
described in section 2.2.2 above) in the methanol reactor.
2.4.3. Processes
ICI
The ICI low-pressure methanol process is the most common process for methanol production
worldwide and accounts for 60 % of the produced methanol. The first ICI low-pressure plant was
built in the 1960’s and since then, not much has happened. The reactor in this process is an adiabatic,
single bed, quench reactor. The fact that the reactor is quenched by 60 % of the fresh syngas
reassures that the reaction rate is high at all times. The remaining 40 % of the fresh syngas is sent
directly to the reactor after it is compressed, heated and mixed with a recirculating flow of unreacted
gas. The produced crude methanol is cooled and condensed by water and high pressure steam is
produced. This high pressure steam can be used in other parts of the process, i.e. the compressors or
reboilers in the distillation columns. The crude methanol is then separated from the gas in a
pressurized vessel before it is passed through two distillation columns to separate the methanol from
water and by-products. The separated gas is purged to keep the inert components at a minimum
before it is recirculated with the fresh syngas (18). A process flow-sheet for the classic ICI process is
shown in Figure 2.4.3 (49):
Figure 2.4.3: A process flow-sheet for the ICI process. a) Pure methanol column; b) Light ends column; c) Heat exchangers; d) Cooler; e) Separator; f) Reactor; g) Compressor; h) Compressor recycle stage (49).
24
Lurgi
The Lurgi process for methanol production is the second most common in the world and accounts for
30 % of the worldwide production. It is very similar to the ICI process described above but the main
difference is the reactor which in the Lurgi methanol process is a multi-tubular reactor, almost
isothermal reactor cooled by water which generates high pressure steam at around 40 bar. A process
flow-sheet for the Lurgi process is shown in Figure 2.4.4 (49):
Figure 2.4.4: A process flow-sheet for the ICI process. a) Pure methanol columns; b) Light ends column; c) Heat exchangers; d) Cooler; e) Separator; f) Reactor; g) Compressor recycle stage (49).
Lurgi has in recent years developed a new and advanced technology for methanol production from
syngas named Lurgi MegaMethanol. The most obvious advantage is the high capacity of the process,
5,000 mtpd compared to 2,500 mptd using classical processes of about the same size (53). The high
capacity of the MegaMethanol plant is due to the fact that two reactors are used; one which is water
cooled and one which is cooled by fresh syngas, see Figure 2.4.5 (54):
25
Figure 2.4.5: A process flow-sheet for the Lurgi MegaMethanol process (54).
As can be seen above, the water reactor is an isothermal and water cooled reactor which operates
like the reactor in the classic Lurgi process. The methanol containing gas leaves the isothermal water
reactor and enters a second, downstream reactor without being cooled. In this second reactor the
fresh syngas enters at a low temperature and flows upstream, inside the tubes while cooling the hot
methanol containing gas which flows in the catalyst bed outside the tubes. The temperature profile
in the gas-cooled reactor is shown in Figure 2.4.6 (54):
Figure 2.4.6: Temperature profile for the gas-cooled reactor (54)
A large advantage with this counter-current is that the methanol containing reaction gas is
continuously cooled, thus reassuring that there is always a driving force towards the methanol side in
the equilibrium reaction which in turn enables conversion rates above 80 %. The recycle ratio can
therefore be reduced by 50 % compared to the recycle ratio in a classic process (54). Another great
advantage is that the need for syngas pre-heating is largely decreased due to the fact that the gas-
cooled reactor works like a heat-exchanger.
26
Carbon dioxide and hydrogen as raw materials
Another interesting option would be to use carbon dioxide and hydrogen directly in a reactor and
thereby bypassing the need of producing syngas. This would be of particular interest when supplies
of pure hydrogen and carbon dioxide exist in abundance, and could very well be considered as a
“green” way of producing methanol, should a renewable energy source be used. Such a process is
currently being researched by Lurgi. Using carbon dioxide instead of carbon monoxide as a raw
material will lead to reaction 2.4.2 and 2.4.3 being the dominant reactions. Since reaction 2.4.2 is less
exothermal than reaction 2.4.3, this will lead to smaller temperature spikes in the reactor which in
turn will lead to a smaller formation of by-products. However, the productivity is slightly lower than
when using standard syngas and there will be a need of developing and testing catalysts designed for
the purpose of converting carbon dioxide to methanol (55).
In late 2011, the world’s first commercial carbon dioxide-to-methanol plant, called the George Olah
Plant, was established in Reykjanes, Iceland, commissioned by Carbon Recycling International. It has
since then been producing five million litres per year whilst recycling about 5,000 tonnes carbon
dioxide per year, using 5 MW of power and geothermal energy as heat supply (56). This is performed
via reaction 2.4.2, using catalysts much alike the commercial copper-zinc oxide-alumina catalysts that
have been widely used in the traditional low-pressure process for the last decades (vide supra). Run
at a temperature of 260 oC, with selectivity to methanol at 99.8 % has been achieved (57), with a
possible equilibrium conversion of 83 % (49).
Outlook - Haldor Topsoe’s Condensing Reactor
Haldor Topsoe, a Danish catalyst and chemical process company, have developed and patented a
type of fixed bed reactor in which the methanol is continuously condensed on cooled surfaces
embedded within the fixed bed. The temperature of the catalyst is kept above the dew point of
methanol, as to limit the condensing of methanol on the particles of catalyst, thereby limiting the
negative effect of mass transfer of gas through a liquid film (58).
2.5. Choice of Process The three processes described in this report have advantages and disadvantages in comparison to
each other. Concerning methanol processes, the carbon dioxide to methanol is chosen since there is
no need of producing syngas via RWGS. As can be seen from Table 2.1, the production of algal oil
does not demand as high temperatures and pressures, which makes for a lower need of heating and
pressurizing. However, conversions and yields are hard to come by, but what is known is that the
potential productivities are very low, at 142 mg/l day compared to FT and methanol processes that
have potential productivities of thousands of tonnes per day. Of course, this depends on the
dimensions of the plants in question but could be used as an indicator of scale-up possibilities. At the
same time, the algal oil is in need of further down-stream processing to be able to use it as food
supplements or bio fuel. This renders more process steps than what is presented in Table 2.2. Since
the algal oil process is still relatively new and under development, with problems concerning shading
and mass transport a recommendation cannot be at this time to further investigate algal oil.
27
Table 2.1: Comparison table for the operational conditions of the processes.
Process Phase of reaction
Temp. (o C)
Press. (bar)
One pass conversion
Total yield Selectivity By-products
LTFT Gas/liquid 200-240 20-45 90 % 80 % to diesel
89 % Petrol, kerosene.
Algal oil Liquid <30 Atm. - - - Biomass CO2 to Methanol
Gas 200-300 50-100 83 % 80 % 99.8 % Hydrocarbons, wax
Table 2.2: Comparison table for the equipment of the processes.
Process Reactors Catalyst Distillation columns
Other steps of purification
Other equipment
Recirculation
LTFT 1 Fe/Co 3 1 hydrocracking RWGS and ORC equipment
Yes
Algal oil 1 Alga, nutrients
- 1 dewatering, 1 oil extracting
Artificial lighting. Later process steps.
Yes
CO2 to Methanol
1 Cu-Zn-Al2O3
1-3 - ORC equipment Yes
When comparing the potential of FT and methanol synthesis, the processes are quite similar. The
processes are run almost alike with almost the same amount of process steps, but there is a small
advantage for carbon dioxide to methanol since the FT synthesis requires one additional purification
step and has a slightly lower selectivity. The methanol is also a much more versatile product that can
be used both as a fuel and as a raw material for a lot different products, even though there can be
issues regarding toxicity and corrosion e.g. when using it as a fuel. Another advantage is that there is
no need of converting carbon dioxide to carbon monoxide which excludes the RWGS investment and
operational costs. Also, the methanol process seems not to be as sensitive to political decisions. As a
speculative example, diesel might be phased out during the next 40 years to due developments in
the automobile sector but the aforementioned scope of use for methanol is key to keeping it on the
market. Even though the carbon dioxide to methanol process is quite new it builds upon the same
principle, reactor design and catalyst as the classic processes. Therefore the conclusion must be to
further investigate the carbon dioxide to methanol process.
28
3. Process design The process investigated consists of three major blocks and the overall principal is shown in figure
3.1. The first block is an Organic Rankine Cycle, where additional electricity is produced from low
grade heat, and the second block consists of an electrolyser where hydrogen and oxygen are
produced from water. Hydrogen and carbon dioxide are mixed, compressed and sent to a catalytic
reactor, where the gases are converted to methanol and water. The pressure of the gaseous mixture
leaving the reactor is first lowered in two flash vessels and the methanol is then separated from
dissolved gases and water in two distillation columns.
Organic Rankine Cycle Electrolyser Methanol processLow grade heat Electricity Hydrogen Methanol
OxygenCarbondioxide
Figure 3.1: Shows the overall principal of the investigated process.
To be able to estimate the investment and operating costs of the plant it is necessary to set up mass-
and energy balances over the different process steps. This is done in the simulation program Aspen
Plus, which is a flow sheeting tool with many predefined unit operations such as distillation columns,
heat exchangers and compressors, which makes simulating large systems simpler and much more
time-efficient compared to using tabula rasa simulation environments such as MatLAB.
The flow sheet produced and simulated in Aspen Plus is presented in appendix A. In the coming
chapters, the specifications used, and simulation results given, when the different unit operations
were simulated, are presented separately.
3.1. ORC and Electrolyser 14 MW electricity and 126 MW heat at 80oC is available for the production of methanol. Since there
is no need for the low degree heat, this is instead used for generation of additional electricity in an
Organic Rankine Cycle. The efficiency of the ORC is assumed to be 7 %, which corresponds to an
electricity generation of about 9 MW. The ORC is not incorporated in the Aspen Plus simulations,
since no pre-programmed ORC block exists. The electricity generated in this unit is just added to the
ingoing electricity to the electrolyser unit. Since 9 MW of electricity is generated from the ingoing
126 MW, 117 MW has to be removed in the condenser.
There is no pre-determined electrolyser block in Aspen Plus, which is why the electrolyser unit is not
incorporated in the simulations, and the hydrogen production from a given amount of electricity is
instead calculated by hand. The power consumption, operating temperature and pressure of the
electrolyser unit is assumed to be 4.5 kWh/Nm3 H2, 90oC and 30 bar respectively. This is equal to the
operating conditions of a Lurgi type electrolyser. Since electricity is needed for compression work
later in the process, all of the electricity available cannot be used for hydrogen production. If 21.5
MW of the available electricity (electricity from ORC included) is used in the electrolyser, 210.4 kmol
H2/h and 105.2 kmol O2/h can be generated (calculations can be seen in appendix A). If this amount
of hydrogen is processed, the simulations shows that the compression work will reach 1.38 MW, and
29
the total electricity consumption of the process will therefore be 22.88 MW. The water consumption
of the electrolyser will be 3.82 m3/h (calculations can be seen in appendix B).
Since there is no need for oxygen in the methanol process, it is suggested that the pressurized
oxygen formed in the electrolysis should be stored in gas tubes and sold to e.g. hospitals in the
vicinity of the plant.
3.2. Compressors In order to pressurize the system, two compressors are needed. In Aspen Plus, these are simulated
using the Pressure Changers Compr and MCompr respectively. The hydrogen in stream 1, as seen in
Figure 3.2.1, is produced in the electrolyser and enters the system at 90oC and 30 bar. The hydrogen
will then be mixed with carbon dioxide, which arrives by pipeline in stream 2 at 0oC and 1 bar but is
compressed in the single-stage compressor to 30 bar and 302oC before mixing. The hydrogen and the
carbon dioxide are also mixed together with a recirculation stream from the first flash vessel at 26oC
and 30 bar, which is first purged by 6 % (not seen in Figure 3.2.1).
Single-stage compressor
Carbon dioxide (Stream 2)
Compressed CO2 (Stream 3)
Cooling water inlet Compressed mixed stream (Stream 5)
Mixer
Hydrogen (Stream 1)
Recirculated stream (Stream 14)
Mixed stream (Stream 4)
Cooling water outlet
Multi-stage compressor
Figure 3.2.1: A schematic view of the compressors of the system
The mixed stream exits the mixer at 75oC and 30 bar and is led to a multi-stage compressor where
the gas mixture is compressed to 100 bar, rendering it at a temperature of 106oC. The intercooler in
the multi-stage compressor has a total cooling duty of 0.937 MW and a required net-work of 1.16
MW whereas the net-work required in the single-stage compressor is 0.26 MW. Required flow rate of
cooling water in the multi-stage compressor is 6.24 kg/s (for calculations see Appendix B). Relevant
data for the compressor system is summarized in Table 3.2.1.
Table 3.2.1: Relevant data given from the simulations of the compressor system
Stream 1 2 3 4 5
Temperature (oC) 90 0 302 74.6 105.6 Pressure (bar) 30 1 30 29.9 100 Mole Flow (kmol/h) 215 71.7 71.7 829.6 829.6
Compressor Single stage Multi-stage
Net work required (MW) 0.26 1.16 Net intercooling required (MW) - 0.937 Required cooling water flow (kg/s) 6.24
30
3.3. Reactor and steam generation For the production of methanol from carbon dioxide, a fixed bed equilibrium reactor is used. In
Aspen Plus, the reactor module is a REquil-reactor. The catalyst, Cu/ZnO/Al2O3, has a remarkably high
selectivity at 99.8 %. The by-products are assumed not to form since the selectivity to methanol is so
high, which is why they are excluded from the simulations. The reactor is run at a temperature of
260oC, which is maintained by the exothermic reactions. The pressure in the reactor is assumed to be
100 bar, with a pressure drop of 5 bar. The feed is pre-heated by the product gases to 187oC, a
temperature above the ignition temperature of 180oC (51). The reactor is cooled by Dowtherm G oil
which is used to produce steam at one bar in an external heat-exchanger. The amount of heat
released during the reaction is simulated to be 0.5 MW which equals a steam mass flow of 0.22 kg/s.
The mass flow of Dowtherm oil used for the steam production is 10.56 kg/s which is calculated by
assuming a 10 degree cooling of the oil. The heat exchanger area is 4.04 m2 (for calculations se
Appendix B). A flow-sheet describing the reactor and the heat exchanger setup is shown in Figure
3.3.1 and relevant data retrieved from the simulations of the reactor block is shown in Table 3.3.1:
Equilibrium reactor
Feed+recirculated stream (Stream 6)
Liquid product stream (Stream 8)
External heat exchanger
Steam, 1 bar 105 C
Water, 1 bar 25 C
Dowtherm G oil
Dowtherm G oil
Vapor Product Stream (Steam 7)
Figure 3.3.1: The reactor and heat exchanger setup.
Table 3.3.1: Relevant data retrieved from the reactor simulations.
Stream 6 7
Temperature (oC) 187 260 Pressure (bar) 100 95,0 Mass Flow (kg/h) 7,225 7,225 Volume Flow (m3/h) 320 333 Mass Fraction of components
Carbon Monoxide 0.028 0.030 Carbon dioxide 0.771 0.396
Water 997 PPM 0.154 Methanol 0.008 0.279 Hydrogen 0.192 0.141
31
3.4. Flash vessels The Aspen Plus module for separation of gases and liquid is called Flash2. The flash vessel system is
presented in Figure 3.4.1.
Figure 3.4.1: A schematic overview of the flash vessel system.
The first flash vessel is used to lower the pressure from 100 bar to 30 bar and to separate the liquid
products from the gases that are recirculated to the multi-stage compressor. The second flash vessel
is used to further lower the pressure so that the distillation columns will run with a lesser energy
demand. Both flash vessels are assumed to be working adiabatically. A schematic view of the flash
vessels is shown in figure 3.4.1 and relevant data retrieved from the simulation of the flash vessels
are shown in Table 3.4.1.
Table 3.4.1 Relevant data retrieved from the simulations of flash vessels.
Stream 11 12 15 16 17
Temperature (oC) 30 25.9 25.9 22.8 22.8 Pressure (bar) 94.7 30 30 1.2 1.2 Vapor Fraction in stream 0.804 0 1 1 0 Mass Flow (kg/h) 7,225 3,952 3,305 167.7 3,138 Volume Flow (m3/h) 161.6 485.4 3.3834 80.78 3.660 Mass Fraction of components
Carbon Monoxide 0.030 0.054 0.001 0.022 157 PPM
Carbon dioxide 0.396 0.669 0.074 0.920 0.029 Water 0.154 0.002 0.335 0.006 0.353
Methanol 0.79 0.016 0.590 0.052 0.618 Hydrogen 0.141 0.259 8 PPB 162 PPB TRACE
32
3.5. Distillation columns To simulate the separation of gases, methanol and water, the Aspen Plus column model RadFrac is
used. The first distillation column is run at a slightly elevated pressure of 1.2 bar, and is used to
separate the dissolved gases from the crude methanol mixture. The column consists of ten stages
and the feed enters on stage two. The condenser is of type partial-vapour so that the evaporated
methanol is condensed back into the column with a reflux ratio of 1. The pressure drop over the
condenser is set to 100 mbar and it is assumed to be cooled by cooling water at a temperature of
15oC. The cooling duty is simulated to be 0.027 MW, which corresponds to a cooling water flow of
3.20 kg/s, the calculations can be found in appendix B. The reboiler is of type kettle, and the required
heat to run the column is simulated to be 0.2 MW which is found by iteration so that a methanol
mass recovery of 99.6 % is achieved. A steam mass flow of 0.073 kg/s (1 bar, 100oC) is required in the
reboiler. A flow sheet describing the distillation columns is presented in Figure 3.5.1 and relevant
data retrieved from the simulation of the distillation column 1 is shown in Table 3.5.1.
Distillation column 1
Gases, CO, CO2, H2 (Stream 18)
Water, Methanol(Stream 19)
Distillation column 2
Methanol (Stream 20)
Water (Stream 21)
Mixed Stream(Stream 17)
Figure 3.5.1: A schematic view of the distillation columns.
Table 3.5.1: Relevant data retrieved from the simulations of distillation column 1.
Stream 17 18 19
Temperature (oC) 22.8 21.8 75.5 Pressure (bar) 1.2 1.1 1.1 Vapor Fraction in stream 0 1 0 Mass Flow (kg/h) 3,137.6 99.34 3,038.3 Volume Flow (m3/h) 3.660 51.803 3.840 Mass Fraction of components
Carbon Monoxide 157 PPM 0.005 TRACE Carbon dioxide 0.029 0.914 TRACE
Water 0.353 0.003 0.364 Methanol 0.618 0.078 0.636 Hydrogen TRACE TRACE TRACE
Oxygen - - - Energy requirements (MW)
Reboiler 0.196 Condenser -0.027
33
The second column is run at a pressure of 1.1 bar and is used to separate the methanol from the
water. The feed enters on the tenth stage and the column consists of 20 stages. The simulated
column has a total condenser with a reflux ratio of 1.7 and the cooling duty is 1.6 MW, which
corresponds to a cooling water flow of 10.74 kg/s. A pressure drop of 100 mbar is assumed. The
column has a distillate to feed ratio of 0.4 which yields a methanol mass recovery of 99.9 %. The
reboiler is of type kettle and the reboiler duty is 1.60 MW, which is equal to a steam mass flow of
0.60 kg/s (1 bar, 100oC). It should be said that is not possible to use steam of 1 bar, 100oC to re boil
water at 99oC, since this would result in infinite heat exchanger area requirements. In reality steam
of higher pressure has to be produced to enable the reboiling, but in this case 1 bar and 100oC steam
has been used in the dimensioning of the reboiler. Relevant data retrieved from the simulations of
distillation column 2 is shown in Table 3.5.2.
Table 3.5.2: Relevant data retrieved from the simulations of distillation column 2.
Stream 20 21
Temperature (oC) 64.3 99.5 Pressure (bar) 1 1 Vapor Fraction in stream 0 0 Mass Flow (kg/h) 1,933.97 1,104.332 Volume Flow (m3/h) 2.597 1.203 Mass Fraction of components
Carbon Monoxide - - Carbon dioxide TRACE TRACE
Water 0.002 0.998 Methanol 0.998 0.002
Energy requirements (MW) Reboiler 1.611
Condeser -1.596
3.6. Heat exchanger network To achieve high total energy efficiency in the process, a well-planned heat exchanger-network had to
be designed. For example; the exothermic reactions can be used to produce steam, the hot product
gases can pre-heat the feed and the product stream can condense in an additional heat exchanger
and steam can once again be produced. In Aspen Plus, the heat exchanger modules used are the
Heater and HeatX. It is assumed that the pressure drop in each heat exchanger is 100 mbar. A flow
sheet describing the setup of the heat exchangers is shown in Figure 3.6.1and relevant data retrieved
from the simulations of the heat exchangers is shown in Table 3.6.1:
34
HE1 HE2 HE3
Warm product stream
(Stream 7)
Gaseous product stream(Stream 9)
Cooled product stream(Stream 11)
Cold feed (Stream 5)
Preheated feed (Stream 6) Water 1 bar, 100 C
Partially condensed
product stream (Stream 10)
Steam 1 bar, 105 C
Cooling water, 1 bar, 25 C
Cooling water, 1 bar, 60 C
Figure 3.6.1: The heat exchanger network.
Table 3.6.1: Data retrieved from the simulations of the heat exchangers.
Stream 5 6 7 9 10 11
Temperature (oC) 105.7 187 260 173.2 110 30 Pressure (bar) 100 95 100 94.9 94.8 94.7 Vapor Fraction in stream 1 1 1 1 0.85 0.804 Mass Flow (kg/h) 7,225 7,225 7,225 7,225 7,225 7,225 Volume Flow (m3/h) 269,8 327 333 277 211,02 161,6 Mass Fraction of components
Carbon Monoxide 0.028 0.028 0.03 0.03 0.03 0.03 Carbon dioxide 0.771 0.396 0.771 0.396 0.396 0.396
Water 997 PPM
997 PPM
0.154 0.154 0.154 0.154
Methanol 0.008 0.008 0.279 0.279 0.279 0.279 Hydrogen 0.192 0.192 0.141 0.141 0.141 0.141
Heat exchanger 1
Since the product gases leaves the reactor at a temperature of 260oC and the feed is in need of pre-
heating to at least the ignition temperature of 180oC the logical choice is to use the product gases for
this purpose. Since both the feed and the products are gaseous the overall heat transfer coefficient
will be low, approximately 10 W/m2K. Another issue that might occur is condensation of the product
gases, which is why the outlet temperature from the heat exchanger is set to 173oC, just above its
condensation temperature. The simulations showed that the heat duty of the heat exchanger is 0.59
MW and that the required heat transfer area is 845 m2.
Heat exchanger 2
The condensing heat of the product stream is used to produce steam. The condensing heat released
in heat exchanger 2 is 1.4 MW which corresponds to a steam production of 0.61 kg/s. The steam is
chosen to be of 1 bar and 100oC.
Heat exchanger 3
Since separation of methanol and the remaining gases is made easier if all the methanol and water is
condensed, an additional cooler is installed before the flash vessels. Cooling water with a
temperature of 25oC is assumed to be available which is why the product gas/liquid stream is cooled
to 30 OC. The overall heat flow in the cooler is 0.9 MW which leads to a cooling water demand of 6
kg/s.
35
3.7. Dimensioning of unit operations The simulation results were used to design and dimension the different process unit operations.
Stainless steel seems an appropriate choice of material for the process and was therefore chosen as
construction material for all unit operations (59). The calculations are shown in appendix B and the
results are summarized in Table 3.7.1.
Table 3.7.1: Results from dimensioning calculations.
Unit operation Design Operating conditions (oC / bar)
Mixer Diameter: 0.10 m H2 inlet: 90 / 30 CO2 inlet: 302 / 30 Recirculation: 26 / 30 Outlet: 74 / 30
Compressor 1 Fluid power: 254.5 kW 0 – 302 / 1 – 30 Compressor 2 Fluid power, 1
st: 599.1 kW
Fluid power, 2nd
: 530.0 kW Intercooler area: 576.38 m
2
74 – 159 / 30 – 55 30 – 106 / 55 – 100
Heat exchanger 1 Area: 818.76 m2 Hot stream: 260 – 173 / 95
Cold stream: 106 – 187 / 100 Heat exchanger 2 Area: 171.75 m
2 Hot stream: 173 – 110 / 95 Cold stream: 25 – 60 / 1
Heat exchanger 3 Area: 183.80 m2 Hot stream: 110 – 30 / 95
Cold stream: 25 – 60 / 1 Flash vessel 1 Inside diameter: 0.506 m
Height: 4.14 m 30 / 95 – 30
Flash vessel 2 Inside diameter: 0.48 m Height: 4.32 m
26 / 30 – 1.2
Distillation column 1 Inside diameter: 0.35 m Height: 6 m Reboiler area: 4.09 m
2
Condenser area: 4.0 m2
Feed: 23 / 1.2 Distillate: 22 / 1.1 Bottoms: 76 / 1.1
Distillation column 2 Inside diameter: 1.05 m Height: 12 m Reboiler area: 805.5 m
2
Condenser area: 51.9 m2
Feed: 76 / 1.1 Distillate: 64 / 1 Bottoms: 99 / 1
Reactor Total vessel volume: 33 m3
Heat exchanger area: 500 m2
Inlet: 187 / 100 Outlet: 260 / 95
External heat exchanger Area: 4.04 m2 Hot stream: 255 – 235 / 1 Cold stream: 100 / 1
Storage vessel Volume: 1600 m3 -
36
3.8. Process efficiencies The energy flows were summarized and an energy flow diagram was created which is presented in
figure 3.8.1.
ORC Electrolyser Methanol process
Power9 MW H2
CO20 MW
Cooling water0 MW
Power1.4 MW
Heat126 MW
Methanol11.05 MW
Process water3.54 MW
Cooling water0 MW
Cooling water0 MW
Cooling water3.4 MW
Power12.5 MW
Cooling water117 MW
O20 MW
Losses from dist.1, flash 2 and purge
2.11 MW
Figure 3.8.1: : An energy flow diagram of the process. The reference temperature is set to 25oC.
By comparing the energy inflows and outflows it can be seen that the energy losses are 2.8 MW
which is equal to 2.0 % of the total energy inflow. Of the total energy input, 7.7 % is converted to
methanol; however 78.9 % of the electricity input is converted to the final product.
Another interesting parameter is the conversion of the raw materials. If the hydrogen and carbon
dioxide losses are summarized it can be calculated that the conversion of hydrogen and carbon
dioxide is 85.5 % and 86.9 % respectively.
4. Economical evaluation To be able to investigate the feasibility of the methanol plant an economical evaluation is made. The
evaluation takes capital costs and operating costs into account and from this, a production cost of
the methanol can be found. Detailed tables containing information of plant cost etc. is found in
appendix C.
4.1. Ulrich’s method The objective of Ulrich’s method for determining the capital cost of the plant is to find the module
cost, CBM, for each of the apparatus. The module cost can be found via an online database which uses
the equation 4.1 below:
(Eq. 4.1)
Here, Cp is the standard cost of the apparatus and FαBM is a factor which takes operating conditions
and the materials of construction into account, as well as installation, instrumentation, et cetera. The
sum of all the module costs will then account for the total capital cost of the plant, K, according to
equation 4.2.
37
(∑ ) (Eq. 4.2)
The factor fc takes contingency and contracting into account and faux makes up for auxiliary
equipment and buildings, such as offices, laboratories and workshops. As rule of thumb, these can be
set to 1.15 and 1.25 respectively.
The capital cost is updated from the 2004 CE plant cost index of 400 to the year average of 2012,
584.6 (60) and a currency conversion of 1 € per 1.3 USD and 7 SEK per 1 USD is used.
To determine the operating costs, these are divided into three major categories: fixed capital (e.g.
spare parts and storage), direct variable costs (e.g. raw material and catalyst) and indirect variable
costs (e.g. overhead costs and research and development costs).
4.2. Results of economic evaluation The details of the economic evaluation are found in appendix C. The procedure and assumptions
used to calculate the production cost per ton produced methanol follows in this chapter.
4.2.1. Capital costs
To calculate the cost of the unit operations of the process, the online database EconExpert (Ulrich)
was used. However, for the ORC and electrolyser, estimates have been used. For the ORC, Opcon
have submitted an estimate of roughly 1.71-2.14 million USD for 800 kW (61). For the electrolyser,
DOE estimates the cost to be about 800 USD per kilowatt (62). The ORC and electrolyser costs are
excluded from the contingency, construction and auxiliary add-ons since they are assumed to be
delivered as fully constructed blocks and that the repair and service is managed by the
manufacturers. The annuity factor, used to calculate the depreciation, is calculated to 0.1315
according to equation 4.3, using a depreciation time N 15 years and an interest rate i of 10 %.
(Eq. 4.3)
The grass roots capital amounts to 62,110,325 USD and in Figure 4.2.1 it can be seen how the capital
cost is distributed between the three major process blocks: ORC, electrolysis unit, and methanol
process. Detailed calculations can be seen in Appendix C.
Figure 4.2.1: Shows how the capital cost is distributed between the three major process blocks.
38
4.2.2. Operating costs
Operating costs are divided into three sub-categories – fixed capital (FC), direct variable costs (DVC)
and indirect variable costs (IVC) – and calculated using rule of thumb. In the first case, the cost for
service and repair on the ORC and the electrolyser is included in the post, whereas it is excluded in
the second case. The total cost for each category is presented in Table 4.2.1 below.
Fixed Capital
Spare parts are calculated as 15 % of the service/repair post in DVC. Storage of raw materials and
products are calculated using market prices (as of May, 2013) over a four week period, with the same
annuity factor fa of 0.1315 as when calculating the depreciation in chapter 4.2.1 above. The largest
post of the fixed capital is the spare parts, accounting for 229 008 USD/a.
Direct Variable Costs
The DVC include raw materials, catalyst, heating and cooling, service and repair, and labour and
management salaries. As for raw materials, the CO2 is considered to be free of charge in the standard
case, whereas the price of the water going to the electrolyser is set to be the same as for the cooling
water, 1.5 USD/m3. However, a large proportion of the cooling comes from the ORC. This is assumed
to be made by using sea water, free of charge, and is thus excluded from these calculations. The
catalyst cost is calculated from the technical lifespan of five years and the same interest rate as
before. The price of the catalyst is assumed to be 80 USD/kg.
Service/repair is assumed to be 7 % of the grass roots capital in the first case and in the second case
to be 7 % of the total CBM due to service/repair being managed free of charge by the manufacturers
of the ORC and electrolyser respectively, e.g. because of warranties.
The personnel cost is based on two sets of three people working on a five-shift schedule with a salary
of 3,570 USD per month. The management post is calculated as 15 % of the personnel post. The
largest post of the DVC is the service and repair, accounting for 1,526,723 USD/a, followed by cooling
at 1,136,592 USD/a, and personnel at 1,285,200 USD/a.
Indirect Variable Costs
The IVC include overhead costs, administration, distribution and sales, and R’n’D. The overhead cost
is 70 % of the personnel cost and 50 % of the service/repair cost. Administration is calculated as 25 %
of the overhead costs. Distribution and sales is 5 % of the total DVC and IVC, whereas R’n’D is 1 % of
the total DVC and IVC. The largest post of the IVC is the personnel, costing 899,640 USD/a.
From Figure 4.2.2it is obvious that the DVC account for the largest part of the operating costs.
39
Figure 4.2.2: A comparison of the operating costs
4.2.3. Total production cost
The production cost of methanol per ton is found by summarizing the annual operating costs and the
depreciation based on 15 years and an interest rate of 10 %. The summary of the economical
evaluation presented in appendix C is shown in Table 4.2.1:
Table 4.2.1: Results of the economical evaluation
Cost
Fixed Capital 303,004 USD/a Direct Variable Costs 4,440,127 USD/a Indirect Variable Costs 2,469,884 USD/a
Operating cost 7,213,015 USD/a Grass Roots Capital 62,110,325 USD Depreciation 8,167,508 USD/a Total annual production cost 15,380,523 USD/a Production cost per ton produced methanol 1,009 USD/ton
4.3. Sensitivity analysis A sensitivity analysis was made on three different cases; grass roots capital, cost per ton of carbon
dioxide and interest rate. The grass roots capital was chosen as an interest of study due to the fact
that Ulrich’s method offers an accuracy of +/- 30 %. The cost per ton of carbon dioxide is interesting
due to possible political decisions surrounding greenhouse gases. Future possible incentives might
result in negative costs for carbon dioxide, i.e. it might become cheaper to capture and convert
carbon rather than emitting it to the atmosphere. Finally, the interest rate is studied because of the
large impact it has on the depreciation cost. The minimum, standard and maximum parameter values
are shown Table 4.3.1. The standard case production cost has been presented in Table 4.2.1 and is
1,009 USD/ton.
40
Table 4.3.1: The different cases of the sensitivity analysis
Minimum Standard Maximum
Grass Roots Capital -30 % 62,110,325 +30 % Price of CO2 (USD/ton) -39 0 39 Interest rate (%) 5 10 20
The result of the grass roots capital sensitivity analysis is shown in Figure 4.3.1:
Figure 4.3.1: Sensitivity analysis of grass roots capital
In the cheapest case, the production cost will be 835 USD/ton. At the other end, the production cost
will be 1,082 USD if the grass roots capital is 30 % higher than calculated. The result of the carbon
dioxide price analysis is shown in Figure 4.3.2:
41
Figure 4.3.2: Sensitivity analysis of carbon dioxide price
In the case where taxes have been raised enough to make the price of the carbon dioxide to be
regarded as negative, i.e. the capturing and conversion of carbon dioxide will be much cheaper than
emitting it, the production cost will be 942 USD/ton. However, if not so, the price of the carbon
dioxide will make for a production cost of 1,076 USD/ton. Also, to get a production cost equal to that
of the market price of methanol at 370 €/ton (63) (equivalent of 480 USD/ton) the price of the
carbon dioxide would have to be -306,5 USD/ton. Finally, the interest rate sensitivity analysis is
shown in Figure 4.3.3:
Figure 4.3.3: Sensitivity analysis of interest rate
42
The interest might have a huge impact on the production cost. If the interest rate is kept low, the
production cost will amount to 866 USD/ton. However, if not so, the production cost will climb up to
1,345 USD/ton.
To easily overview the impact of the analysed cases, a summary of the sensitivity analysis is shown in
Figure 4.3.4. As can be seen, it will be important to keep the cost of the ORC/electrolyser and the
interest rate as low as possible.
Figure 4.3.4: The summary of the sensitivity analysis
43
5. Discussion and conclusions The scope of this project was to evaluate the possibility of converting heat, electricity and carbon
dioxide into chemical components. In the literature review the conclusion was to further investigate
the feasibility of converting carbon dioxide into methanol which is why this process was designed
and simulated. The specifications submitted by Statoil included 14 MW of electricity and 126 MW of
heat at 80 oC being available to convert an unlimited supply of carbon dioxide into the desired
product. The carbon dioxide was assumed to be free of charge. The 126 MW of heat was converted
into electricity using an Organic Rankine Cycle, although with a low efficiency of 7 %. This was since
heat at 80 oC is not very useful in the process and because the limiting factor for production is the
amount of hydrogen produced in the electrolyser. The usage of an electrolyser for hydrogen
production was another specification by Statoil but there might be other interesting options but this
is outside of the scope of this project.
The efficiency of the process, discussed in chapter 3.8, is reasonable with an electrical efficiency of
78.9 % and conversions of hydrogen and carbon dioxide to be close to 90 %. The total efficiency, i.e.
the total energy input over the energy found in the product, is quite low which is because of the very
low efficiency of the ORC. However, without the ORC the 126 MW of heat would have been utterly
useless. Therefore, it is believed that the electricity efficiency should be used as a key indicator and
not the total efficiency. An improvement of the ORC efficiency is still interesting, though, as it would
potentially have a large beneficial impact on the process. A greater efficiency would mean that a
larger proportion of the heat could be converted into electricity which in turn would produce more
hydrogen.
The economic analysis based on the simulations yielded a production cost of 1,009 USD/ton with an
annual production of 15,240 tonnes based on 8,000 hours of operational time. Hence, the production
cost is more than double of the market price of the first quarter in 2013 which was 370 €/ton (63)
which corresponds to 480 USD/ton. Thus, it is a process that might not seem to be economically
viable. However, further studies including, but not limited to, optimization and scale-up might
indicate that larger systems are more profitable e.g. due to larger production rates but not
necessarily as large scale-up factor on the grass-roots capital. The depreciation cost of the grass roots
capital accounts for more than half of the annual production cost, but the sensitivity analysis shows
that the interest rate and the grass roots capital have a large impact on the production cost which
should therefore be kept at a minimum. The insecurity regarding the ORC and electrolyser makes the
sensitivity analysis an interesting point of study, especially since it is suspected that machines with
higher production capacities might be available and therefore the cost per produced MW will be
lower.
For further studies, it is recommended to start by performing a risk-analysis of the plant, followed by
adding by-products to the simulations. The by-products would affect not only the reactor but also
the separation steps. It should also be of interest to make use of the purged hydrogen and carbon
dioxide rather than just emitting it to the atmosphere and thereby losing valued raw materials. The
purged stream could for instance be burned for generation of high quality heat that could be used on
the process site. Another interesting point of study would be to improve heat usage, e.g. by
improving the ORC or finding other ways to utilize the heat. Finally, since this is a small plant in the
44
pilot scale, a scale-up would be interesting to investigate in order to discover the full potential of
carbon capturing and utilisation.
45
6. Acknowledgments We, the authors, want to direct our deepest gratitude to Hans T. Karlsson and Christian P. Hulteberg
at the Department of Chemical Engineering, Lund University for the support and encouragement
they have offered throughout the entire project. Also from the department, we would like to thank
Laura Malek in particular for the help she provided during the simulations as well as the literature
study. We would also like to thank Ola Wallberg and Per Tunå for the assistance they provided.
We would also like to extend our gratitude to Børre T. Børresen and Astrid Mejdell at Statoil for
giving us the opportunity to work with this very interesting project.
Lastly, we would like to thank our fellow student Maria Sundqvist for supplying an important piece of
literature.
46
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I
Appendix A – Flow sheet
II
Appendix B – Design and dimensioning In this appendix the dimensioning calculations of the different process apparatus are described.
Necessary data used in the calculations is retrieved from the Aspen Plus simulations.
ORC and Electrolyser The efficiency of the electrolyser is assumed to be 7 %, which is why 9 MW of the ingoing 126 MW of
heat will be converted to electricity.
If 21.5 MW of electricity is used in the electrolyser and it is assumed that the power consumption is
4.5 kWh/Nm3 H2 the following quantities of hydrogen can be formed in the electrolysis unit:
Since one mole of water gives one mole of H2 and ½ mole O2, 105.2 moles or 2390 Nm3 O2/h will be
formed in the electrolyser.
The water consumption of the electrolyser is 0.8 l/Nm3 H2, which is why the total water requirements
of the electrolysis unit will be 3.82 m3/h.
Heat exchanger 2 The product gases from the reactor are cooled in heat exchanger 2 from 173oC to 110oC with boiling
water. The gas stream is partially condensed and the heat transfer coefficient can be estimated with
rules of thumb from (64) to 0.25 kW/m2 oC. The following equations are used for calculating the heat
transfer area and the steam production in the heat exchanger.
Heat exchanger area
Steam production
Heat exchanger 3 The gases leaving heat exchanger 2 are further cooled in heat exchanger 3 with water that is
assumed to enter the heat exchanger at 25oC and leave at 60oC. Some condensation also occurs in
heat exchanger 3, why the heat transfer coefficient is assumed to be equal to that in heat exchanger
2. The following equations are used for calculating the heat transfer area and the required cooling
water flow.
III
Heat exchanger area
Cooling water flow
Compressor 2 The feed gases are compressed in a multistage compressor with intercooling. The gases are leaving
the first compressor stage at 159.8oC and shall be cooled to 30oC before entering the second stage. It
is assumed that the cooling water enters at 25oC and leaves at 60oC and that the heat transfer
coefficient is 0.05 kW/m2 oC. The following equations are used for calculating the required cooling
water flow and the heat transfer area:
Mixer The diameter of the mixer is estimated from the economic optimum velocity. According to (65) gases
with a density in the range of 0.2 – 20 kg/m3 has an optimum velocity between 40 and 9 m/s. The
density of the gases entering the mixer is 9.13 kg/m3, why the optimum velocity in this case can be
estimated to 26 m/s by assuming a linear dependence between gas density and optimum velocity.
The volumetric gas flow through the mixer is 0.22 m3/s, which is why the required diameter of the
mixer can be estimated to 0.10 m.
Flash vessel 1 The volumetric flow rate of liquid through flash vessel 1 is 3.79 m3/h. If it is assumed that the
residence time in the flash vessel is 10 min, the required volume of the liquid phase can be calculated
to 0.63 m3. The maximum gas velocity, ug, can be calculated with the following equation (66):
√
Where ρl = 862.98 kg/m3 and ρg = 8.36 kg/m3
are the densities of the liquid and gas through the flash
vessel. The maximum gas velocity can be calculated to 0.65 m/s.
The volumetric flow rate of gas leaving flash vessel 1 is 469.10 m3/h. If the gas velocity is set to the
maximum value, the flash vessel area and diameter can be calculated to 0.2 m2 and 0.51 m
respectively.
IV
Since the area and the volume of the liquid filled part of the flash vessel are known, the height of the
liquid filled part can be calculated to 3.16 m. The height of the gas filled part of the vessel is set to 1
m, which is why the total height of the flash vessel can be estimated to 4.16 m.
Flash vessel 2 Flash vessel 2 is dimensioned in the same manner as flash vessel 1. The volumetric flow rate of liquid
through flash vessel 2 is 3.61 m3/h. If it is assumed that the residence time in the flash vessel is 10
min, the required volume of the liquid phase can be calculated to 0.60 m3.
The densities of the liquid and gas through the flash vessel are 857.87 and 2.08 kg/m3 respectively,
which is why the maximum gas velocity can be calculated to 1.30 m/s.
The volumetric flow rate of gas leaving flash vessel 1 is 84.67 m3/h. If the gas velocity is set to 10 % of
the maximum value, the flash vessel area and diameter can be calculated to 0.18 m2 and 0.48 m
respectively.
Since the area and the volume of the liquid filled part of the flash vessel are known, the height of the
liquid filled part can be calculated to 3.32 m. The height of the gas filled part of the vessel is set to 1
m, which is why the total height can be estimated to 4.32 m.
Reactor The volumetric feed flow rate to the reactor is 319.73 m3/h (187oC, 100 bar) which is equal to 18975
Nm3/h. If it is assumed that the weight hourly space velocity is 6.0 Nm3/kg kat s, the catalyst mass
required can be calculated to 3163 kg. The catalyst bulk density is assumed to be 1200 kg/m3 which is
why the catalyst volume can be calculated to 2.635 m3.
If it is assumed that the optimal gas velocity through the catalyst with respect to pressure drop and
mass transfer is 2 m/s at NTP, the required catalyst area can be calculated to 2.635 m2 and since the
catalyst volume is known, the height of the catalyst bed can be calculated to 1.00 m. It is although
assumed that the catalyst is layered with inert catalyst, which is why the catalyst height will be
doubled. It is also assumed that 1.5 m is required in the top and bottom of the reactor, why the total
reactor vessel height will be 5.0 m.
If it is assumed that the catalyst bed occupies 40 % of the total vessel area, the vessel area can be
calculated to 6.6 m2 which corresponds to a diameter of 2.9 m. Since the height is known, the total
vessel volume can be estimated to 33 m3.
It is assumed that the catalyst is packed in tubes with a diameter of 3.81 cm. The number of tubes
required can then be calculated to 2313. The heat transfer area can be calculated to 500 m2.
The reactor is cooled with Dowtherm oil which is assumed to enter the reactor at 235oC and leave at
255oC. The specific heat of the oil is 2.367 kJ/kg K.
The oil is then used for steam generation in an external heat exchanger. Boiling water enters at
100oC, 1 bar and the heat transfer coefficient is assumed to be 0.85 kW/m2 o C.
V
Distillation column 1 To be able to estimate the height and diameter of the distillation columns, the Sounders and Brown
equation is used (67):
(
)
Where:
is the maximal allowed vapor velocity based on total column area in m/s.
is the spacing between trays in m.
and is the liquid and vapor density in kg/m3.
The average liquid and vapor density in distillation column 1 is 785 and 1.13 kg/m3, and if the spacing
between the trays is assumed to be 0.6 m, the maximal allowed vapor velocity can be calculated to
1.41 m/s. The column diameter can then be calculated with the following equation:
√
Where is the maximum vapor mass flow velocity, which in this case is 0.16 kg/s. The column
diameter can be calculated to 0.36 m.
The number of trays in distillation column 1 is 10, and the spacing between the trays is assumed to
be 0.6 m, which is why the column height can be estimated to 6 m.
The reboiler shall operate at a temperature of 76oC and steam at 1 bar, 100oC is used as heating
medium. If the heat transfer coefficient is assumed to be 2 kW/m2 oC the heat exchanger area and
required steam mass flow can be calculated:
The condenser shall operate at 22oC and it is assumed that cooling water enters the condenser at
15oC and leaves at 17oC and that the heat transfer coefficient is 2 kW/m2 oC. The heat transfer area
and the necessary cooling water flow can then be calculated.
VI
Distillation column 2 The column diameter and height of distillation column 2 is estimated in the same manner as for
distillation column 1. The average liquid and vapor density in distillation column 2 is 745 and 1.16
kg/m3 respectively, and the spacing between the trays is assumed to be 0.6 m, the maximal allowed
vapor velocity can be calculated to 1.41 m/s and the column diameter can then be calculated to 1.08
m.
The column has 20 trays, which is why the column height can be estimated to 12 m.
The reboiler of distillation column 2 operates at 99oC and the heating media is steam at 1 bar, 100oC.
If the heat transfer coefficient is assumed to be 2 kW/m2 oC the required heat transfer area and
required steam mass flow can be calculated.
The condenser operates at 64oC and it is assumed that cooling water enters the condenser at 25oC
and leaves at 60oC and that the heat transfer coefficient is 2 kW/m2 oC. The heat transfer area and
necessary cooling water flow can then be calculated.
Storage vessel A storage vessel for the produced methanol shall be able to hold the production of 4 weeks. The
production is 2.56 m3/h and an operating time of 8000 h/year is assumed.
VII
Appendix C – Economy
Capital costs Unit Op. Operating Conditions Material Cp (USD) Fα
BM CBM (USD)
Mixer ∙ Flow rate: 0.219 m3/s
∙ Density: 9.1 kg/m3
∙ Linear flow rate: 26.0 m/s ∙ Diameter: 0.1 m
Stainless steel
2,008 2.9 5,823
Comp 1 ∙ Type: Centrifugal gas compressor ∙ Fluid power: 254.5 kW
Stainless steel
303,948 6.3 1,914,873
Comp 2 ∙ Type: Multistage centrifugal gas compressor, approx. as two single stage compressors. ∙ Fluid power, 1
st: 599.1 kW
∙ Fluid power, 2nd
: 530.0 kW ∙ Intercooler (IC) type: Fixed tube sheet and U-tube ∙ IC area: 576.38 m
2
∙ Pressure: 53.68 barg, tube side only
Stainless steel
1st: 672,889 2nd: 600,533 IC: 58,079
6.3 6.3 6.0
4,239,199
3,783,361
349,225
Heat exchanger 1
∙ Type: Fixed tube sheet and U-tube ∙ Area: 818.76 m
2
∙ Pressure: 99 barg, shell and tube side both
Stainless steel
76,032 6.0 512,047
Heat exchanger 2
∙ Type: Fixed tube sheet and U-tube ∙ Area: 171.75 m
2
∙ Pressure: 95 barg, tube side only
Stainless steel
24,806 6.2 152,800
Heat exchanger 3
∙ Type: Fixed tube sheet and U-tube ∙ Area: 183.80 m
2
∙ Pressure: 95 barg, tube side only
Stainless steel
25,933 6.2 159,743
Flash vessel 1 ∙ Type: Process vessel, vertically oriented, no packing or trays ∙ Inside diameter: 0.506 m ∙ Height: 4.14 m ∙ Pressure: 29 barg
Stainless steel
11,837 20.9 246,947
Flash vessel 2 ∙ Type: Process vessel, vertically oriented, no packing or trays ∙ Inside diameter: 0.48 m ∙ Height: 4.32 m ∙ Pressure: 0.2 barg
Stainless steel
12,054 9.4 113,654
VIII
Unit Op. Operating Conditions Material Cp (USD) FαBM CBM (USD)
Distillation column 1
∙ Column type: Process vessel, vertically oriented, sieve trays. ∙ Inside diameter: 0.35 m ∙ Height: 6 m ∙ Pressure: 0.2 barg ∙ No. of trays: 10 ∙ Reboiler type: Fixed tube sheet and U-tube ∙ Reboiler duty: 196.3 kW ∙ Reboiler area: 4.09 m
2
∙ Pressure: 0.1 barg ∙ Condenser type: Fixed tube sheet and U-tube ∙ Condenser duty: 27.0 kW ∙ Condenser area: 4.0 m
2*
∙ Pressure: 0.1 barg
Stainless steel
Col.: 14,872 Reb.: 3,860 Cond.: 3,830
9.7 5.8 5.8
143,529
22,320
22,150
Distillation column 2
∙ Column type: Process vessel, vertically oriented, sieve trays. ∙ Inside diameter: 1.05 m ∙ Height: 12 m ∙ Pressure: 0.1 barg ∙ No. of trays: 20 ∙ Reboiler type: Fixed tube sheet and U-tube ∙ Reboiler duty: 1611.0 kW ∙ Reboiler area: 805.5 m
2
∙ Pressure: 0 barg ∙ Condenser type: Fixed tube sheet and U-tube ∙ Condenser duty: 10595.7 kW ∙ Condenser area: 51.9 m
2
∙ Pressure: 0 barg
Stainless steel
Col.: 43,590 Reb.: 75,067 Cond.: 12,056
10.5 5.8 5.8
456,714
434,088
69716
Reactor ∙ Type: Jacketed vessel – with coil - Heat exchanger area: 500 m
2
∙ Total vessel volume: 33 m3
∙ Pressure: 99 barg
Stainless steel
36,989 38.5 1,423,323
External heat exchanger
∙ Type: Fixed tube sheet and U-tube ∙ Area: 4.04 m
2
∙ Pressure: 0 barg
Stainless steel
3,844 5.8 22,226
Storage vessel
∙ Type: Cone-roof ∙ Volume: 1600 m
3
Stainless steel
85,903 3.5 300,662
Total CBM 15,172,400 * The minimum value for usage in the database
IX
Unit Op. Operating Conditions CBM (USD)
ORC ∙ Generation: 9 MW ∙ Price of 800 kW OPCON ORC Powerbox: estimated to 1,925,000 USD.
23,100,000
Electro-lyser
∙ Electricity usage: 21.5 MW ∙ Price of electrolyser: estimated to 800 USD/kW
17,200,000
Cost (USD)
Total bare module cost 15,172,400 Contingency, construction and auxiliary (15% and 25% of total bare module cost)
6,637,925
ORC and electrolyser cost 40,300,000
Grass Roots Capital 62,110,325
Operating costs
Fixed Capital
Post Cost (USD/a)
Storage: Raw materials 0
Product 73,996 Spare parts 229,008
Total 303,004
Direct Variable Costs
Post Cost (USD/a)
Raw materials: CO2 0
Water 45,792 Catalyst 253,040 Heating 0 Cooling 1,136,592 Service/repair 1,526,773 Personnel 1,285,200 Management 192,780
Total 4,440,127
Indirect Variable Costs
Post Cost (USD/a)
Overhead: Personnel 899,640
Service/repair 763,361 Administration 415,750 Distribution/sales 325,944 Research and development 65,189
Total 2,469,884
X
Total
Post Cost
Production cost 7,213,015 USD/a Depreciation cost 8,167,508 USD/a Production cost per ton produced methanol
1009 USD/ton
