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KET050 Feasibility Studies on Industrial Plants
Dept. of Chemical Engineering
Lund University
A Feasibility Study on In-refinery Lignin
Hydrogenation
Presented to Preem
MAY 23, 2016
Principal investigators:
Gustav Améen, Tobias Essunger, Lukas Olsson, Luiz
Fellipe Ribeiro, Diego Alejandro Sánchez.
Tutors:
Christian Hulteberg, Per Tunå, Omar Abdelaziz
Department of Chemical Engineering
i
Abstract As renewable fuels have gained a great interest in recent year the use of lignin as fuel feedstock was
investigated. An existing refinery in Gothenburg belonging to Preem was examined for possible
integration of lignin reforming. A model of the process was built and different scenarios with lignin
feedstock were formulated and tested, the production pools sizes and compositions were examined.
Three different scenarios where lignin was added to the current oil refinery were simulated in Aspen
Plus 8.6v and compared to the present scenario with no lignin addition.
The three scenarios were based on the addition of lignin to process in different ways, the addition of
lignin into the crude distillation unit directly, the addition of lignin directly to the dehydrotreater, and
the pretreatment of lignin before its addition to the crude distillation unit. Material and energy
needed to drive the process for each unit were calculated by the simulation in Aspen and a cost
estimation for operational costs, equipment and investment was made. The yield for all the different
fuel pools and their quality were compared and a best scenario was chosen.
The addition of lignin directly into the dehydrotreater proved to be the best scenario based on both
operational costs, fuel yield and quality.
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Contents Abstract .................................................................................................................................................... i
1. Introduction ......................................................................................................................................... 1
2. Background .......................................................................................................................................... 1
3. Literature Review ................................................................................................................................ 2
3.1 Lignin ............................................................................................................................................. 2
3.2 Hydrogen Production .................................................................................................................... 3
3.3 Refinery ......................................................................................................................................... 5
3.3.1 API & Sulphur Content ............................................................................................................ 5
3.3.2 Overall Process ....................................................................................................................... 5
3.3.3 Desalting ................................................................................................................................. 5
3.3.4 Crude Distillation Unit ............................................................................................................ 6
3.3.5 Hydroprocessing ..................................................................................................................... 6
3.3.6 Thermal Process ..................................................................................................................... 7
3.3.7 Propane Deasphalting ............................................................................................................ 7
3.3.8 Catalytic Reforming ................................................................................................................ 8
3.3.9 Catalytic Isomerization ........................................................................................................... 9
3.3.10 Alkylation ............................................................................................................................ 10
3.3.11 Green Hydrotreater ............................................................................................................ 11
3.3.12 Octane and Cetane Number ............................................................................................... 11
4. Simulation .......................................................................................................................................... 12
4.1 Oil Composition ........................................................................................................................... 12
4.2 Crude Distillation Unit ................................................................................................................. 13
4.3 Dehydrotreater ............................................................................................................................ 14
4.4 Alkylation, Isomerization and Reforming .................................................................................... 15
4.5 Green Hydrotreater ..................................................................................................................... 15
4.6 Top Separations & Alkylation ...................................................................................................... 16
4.7 Estimating Octane and Cetane Numbers .................................................................................... 17
4.8 Adding Scenarios ......................................................................................................................... 19
4.9 Assumptions and Considerations for the Simulation .................................................................. 21
5. Cost Estimation .................................................................................................................................. 22
5.1 Lignin Cost ................................................................................................................................... 22
6. Results and Discussion ...................................................................................................................... 23
6.1 Pumping Costs ............................................................................................................................. 28
6.2 Raw Material Costs ...................................................................................................................... 28
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6.3 Investment Costs ......................................................................................................................... 28
6.4 Results on Octane and Cetane Numbers ..................................................................................... 29
6.5 Best Scenario ............................................................................................................................... 30
6.6 Crude Distillation Unit ................................................................................................................. 30
6.7 Pretreatment ............................................................................................................................... 30
6.8 Dehydrotreater ............................................................................................................................ 30
6.9 Other Alternatives ....................................................................................................................... 30
6.10 Further Findings......................................................................................................................... 30
6.11 Cost Discussion .......................................................................................................................... 31
6.12 Gasoline limitations ................................................................................................................... 32
7. Conclusion ......................................................................................................................................... 32
References ............................................................................................................................................. 33
Appendix A. Mass fractions in gasoline for different scenarios and raw data from Aspen. ................. 35
Appendix A1 ...................................................................................................................................... 35
Appendix A2 ...................................................................................................................................... 38
Appendix A3 ...................................................................................................................................... 41
Appendix A4 ...................................................................................................................................... 45
Appendix B. Detailed Illustration of Aspen Simulation for Different Scenarios .................................... 48
B1. No Vanillin Added ........................................................................................................................ 48
B2. Vanillin Added Directly to the Crude Distillation Unit ................................................................. 49
B3. Vanillin Added Directly to the Dehydrotreater .......................................................................... 50
B3. Vanillin Pretreated Before Being Added to the Crude Distillation Unit ...................................... 51
Appendix C. Aspen Parameters and Data. ............................................................................................ 52
C1. Table of Aspen Data .................................................................................................................... 52
C2. Temperature Profile without the Addition of Vanillin ................................................................ 53
C3. Temperature Profile with the Addition of Vanillin Directly to the Crude Distillation Unit ......... 54
C4. Temperature Profile with the Addition of Vanillin Directly to the Dehydrotreater .................... 55
C5. Temperature Profile with Preatment of Vanillin Before Addition to the Crude Distillation Unit 56
1
1. Introduction In recent year’s there´s been a great increase in demand for renewable fuels. To meet this demand,
the possibilities of using lignin as a substitute for fossil fuel is investigated. The lignin will be
pretreated and introduced in the already existing refinery in Gothenburg. It will be investigated in
which fuel pools the treated lignin will end up, and also where in the process it should be added.
The aim of this paper is to simulate the insertion of a lignin feedstock in an existing refining
processes, in order to estimate the behavior of major processing units and to elaborate an
economical evaluation for the investment needed. The simulation will be modeled based on the
current refining processes in Preems refinery in Gothenburg. Preem is the largest refiner in Sweden,
which accounts for 80% of the Swedish refinery capacity and 30% percent of the Nordic refinery
capacity, with a refining capacity of around 345,000 barrels per day.
2. Background The modern society has been built on the use of fossil fuels for energy supply, where slightly more
than 80% of global energy supply has its origin in fossil fuels. Of that amount, roughly 40% derives
from oil consumption, where approximately 90 million barrels per day are produced. According to
official estimates, the energy consumption is expected to increase as it has been doing in the recent
past, leading to increasing exigencies for satisfying energy demand, however thanks to current low
oil prices the fossil fuels does not have a supply problem. (Stenström, 2015) A gigantic industry was
built for oil refining and subsequent products, which have had a big impact on all levels from the
construction of plastics to transportation fuel. The main reserves of oil are located in the Middle East,
Russia and the U.S. (Konnov, 2016)
However, fossil fuels reserves are finite and depletion is far greater than formation of new ones (a
process which takes millions of years to occur), thus fossil fuels are considered a non-renewable
source of energy. Since the reserves of fossil fuels are limited, various types of global problems are
expected to arise in the future when the reserves are depleted. More easily available reserves of
better quality will be utilized first, thus the prices can be expected to rise in the future because of
higher costs for extraction and refining. The quality of fuel is also expected to decrease over time
(decrease in the net amount of energy obtained per unit of energy input) which leads to major
environmental problems and more refining efforts. Another issue is that existing reserves are
concentrated to a limited number of countries, thus economic and political tensions are expected.
(Stenström, 2015; Konnov, 2016)
An additional problem in using fossil fuels is its contribution to global warming and pollution (such as
acid rain and smog). A critical global question concerns which fuels should be used from now on and
in the future. The trend is to decrease the use of fossil fuels and instead develop and burn biofuels in
order to try to hinder global warming of the earth.
Nowadays fossil fuel feedstocks for oil refineries is the crude oil which consists of a large number of
hydrocarbons. The low-molecular fraction (propane in major composition) can be directly used as a
gaseous fuel (LPG). The liquid raw oil can be distilled in different fractions. Petrol (gasoline) is taken
from the fraction of crude oil in a boiling point range of 50-170°C consisting of hydrocarbons with a
carbon chain length of C4-C12 in a mixture of paraffins, cycloalkanes and oleofins which ratio depends
on the refinery processing units, the crude oil used and the grade of gasoline (octane rating).
Commercial jet aircraft fuel, so-called kerosene, has a slightly lower volatility fraction, and diesel fuel
has an even higher boiling point fraction of 280-480°C. Kerosene consists mostly of hydrocarbons
with a carbon chain length of C10-C16, and diesel C11-C24. (Konnov, 2016)
2
In order to remain competitive in this changing scenario, future refineries are nowadays expected to
shift fossil fuel feedstocks for renewable feedstocks, and develop new and more efficient processes
for the desired fuel pool obtained through refining.
Bio-fuels are liquid or gaseous fuels predominantly produced from biomass, which has been
recognized as an important energy source that could replace a fraction of fossil fuels. There are two
main incentives that make biomass attractive as a global energy source: firstly, it is renewable,
secondly it has beneficial environmental properties since there is no net release of carbon dioxide.
Biomass is a mixture of constituents of different structures: hemicellulose, cellulose, lignin and minor
amounts of extractives. After processing of the biomass to smaller pieces, the conversion to fuels can
take place through thermochemical and/or biological processes.
Lignin was until recently considered a waste product of the pulping processes (lignocellulosic fibrous
material prepared by chemically or mechanically separating cellulose fibers from wood or fibers
crops), which was only used as fuel for heating on the separation process. Lignin is now gaining
attention from industries for its immense economical potential. Mainly for being a natural source of
aromatic compounds as well as a wide variety of hydrocarbons, which can be refined into a huge
variety of products. (Calvo-Flores, Dobado, Isac-García, & Martín-Martínez, 2015)
3. Literature Review
3.1 Lignin Lignin is a cross-linked amorphous phenolic polymer which has a complex macromolecule. It is a
biopolymer which is derived from plants. Its function is to protect against attacks from other
microorganisms. It also provides internal transport of nutrients and water as well as being a basic
structure supporting plant tissues, providing rigidity to the plant. It maintains the integrity of the
cellulose/hemicellulose/pectin matrix. It is the second most abundant natural polymer, representing
approximately 25% of the total components of plants. Together with cellulose and hemicellulose it is
one of the major sources of non-fossil carbon that make a special contribution to the carbon cycle.
Lignin represents an enormous reservoir of bound organic carbon. (Calvo-Flores, Dobado, Isac-
García, & Martín-Martínez, 2015)
From an economical point-of-view, roughly 1.3 billion cubic meters of lignin is obtained annually
along with the harvesting of timber. The paper industry and biorefineries produce so large quantities
of lignin that it´s considered a by-product. Lignin has several technical applications in many fields,
acting as a middle-term alternative for the production of chemicals, polymers, carbon fibers, fuels
and new materials. However, despite its potential, lignin is a fairly unused renewable raw material,
just recently gaining attention from industries. (Calvo-Flores, Dobado, Isac-García, & Martín-
Martínez, 2015)
Lignin has the highest heating value among all natural carbon polymeric compound for hard wood it
is often (HHV = 23.50 MJ/kg) and softwood lignin (HHV = 21.45 MJ/kg) (HHV=higher heating value)
which can be used directly as a fuel, or be incorporated as additive to several types of biomass-
derived fuels for industrial and domestic uses. From an economic standpoint, lignin isolated either
from biomass or the paper industry is cheaper than many other manufactured materials, but it plays
the same role that many other chemicals do. This makes it very competitive considering any market
in which it might be introduced. Between 40 and 50 million cubic meters are produced per year
worldwide from wood-pulping processes, mostly as non-commercialized waste products. (Calvo-
Flores, Dobado, Isac-García, & Martín-Martínez, 2015; Blunk & Jenkins, 2000)
3
For lignin to be added into fuel components, for example in the gasoline range, lignin has to undergo
pretreatment to be compatible with both the refinery processes and the evolving regulations for
transportation fuels. Different treatment processes are currently being developed for the
depolymerization and the removal of sulfur and oxygen into products that are predominately
mixtures of aromatic and naphthenic hydrocarbons that can be blended with gasoline.
Depolymerization by catalyzed cracking or hydrogenolysis, hydrodeoxygenation and organosolvolysis
showed promising results. Further treatment is also necessary due to the high content of moisture
and oxygen in lignin. It also contains volatile alcohols, esters and oils that have a solvent action on
some protective coatings applied to metals objects, thereby making the metal more susceptible to
the effects of moisture or other chemicals. Many species of timber can produce small amounts of
volatile corrosive substances and this more pronounced under warm, damp conditions. These
substances include many molecular weight carboxylic acids particularly acetic and formic acid, which
can both react directly to the metal and also accelerate the corrosion rate by increasing the solubility
of the primary corrosion product. (Johnson, Chornet, Zmierczak, & Shabtai, 2002; Joffres, o.a., 2013;
Thringa, Katikaneni, & Bakhshib, 2000; Umney, 1992)
A problem in the future might be that some of the products produced from hydrotreating lignin has a
value as chemicals which may have an effect of the price of lignin. Several factors restrict the use of
lignin, a non-uniform structure, unique chemical reactivity, and the presence of various organic and
inorganic impurities. (Vishtal & Kraslawski, 2011)
The lignin used in this project had the following composition seen in table 1. The products introduced
in the system is treated so that the composition is assumed to be mainly vanillin. This is done in order
to simplify the calculations in Aspen. Vanillin is fairly complex but is still simple enough to be
simulated by Aspen. The reactions for Deoxygenation are defined, and being an aromatic it gives a
fair representation of lignin.
Table 1 shows the general composition of lignin (Hulteberg, Tunå, & Abdelaziz, 2016).
Component Fraction [%]
Oxygen 21.35 %
Sulfur 0.45 %
Hydrogen 4.6 %
Carbon 73.4 %
Nitrogen 0.2 %
The molecular weight was 400-500 g/mol. (Hulteberg, Tunå, & Abdelaziz, 2016)
3.2 Hydrogen Production Hydrogen is an important raw material for many industry branches. By 2010 the hydrogen world’s
production as a chemical and as an energy source was valued as US 120 billion and world’s
consumption of hydrogen was mainly used for ammonia synthesis (62.4 %), oil refining (24.3 %) and
methanol production (8.7 %). In the past, refineries produced net hydrogen as a by-product of
catalytic reforming steps, with some minor portion being used internally in the industry in a variety
of hydrothermal processes. Nowadays, however, the internal demand for hydrogen in refineries has
increased largely due to stricter regulations on automobile emissions coupled with increasing
demand in hydrogen for hydrotreatment processes, making current refineries major net consumers
of hydrogen, establishing independent hydrogen production plants on their perimeter. The demand
for hydrogen is forecasted to increase in the major sectors by 10-15% per year. (Grupta, 2009;
Pagliaro & Konstandopoulos, 2012)
4
Natural Gas 48%
Coal 18%
Petroleum 30%
Electrolysis 4%
Hydrogen can be produced through thermal, electrolytic, or photolytic processes using fossil fuels,
biomass, or water as feedstock. Thermal processes used to produce hydrogen from methane include
steam methane reforming, partial oxidation and autothermal reforming (which combines the other
two). When heavy oils or coal is used, the gasification process is commonly used. Another way to
produce hydrogen is by water electrolysis, which is an electricity intensive process.
Currently hydrocarbons, majorly natural gas and petroleum, account for 78 % of world’s hydrogen
production. Only up to 4 % of hydrogen is produced by water electrolysis. The share that each
feedstock has in hydrogen production can be seen in Figure 2. The production of hydrogen requires
the least amount of energy if light hydrocarbons are used as feedstock, whereas hydrogen produced
by water electrolysis requires the highest amounts of energy. The energy consumption for each
process can be seen in Figure 1. (Grupta, 2009)
Environmentally speaking, the current production of hydrogen based on hydrocarbon fuels is a major
and important releaser of CO2, emitting around 100 million tonnes of CO2 equivalent per year.
Therefore, to become an environmental friendly fuel the production of hydrogen should shift to
water electrolysis, which is both renewable and allows for the zero net emission of CO2 if electricity
can be produced in a renewable way; to perform this shift, however, the price of hydrogen obtained
through electrolysis must decrease considerably. Estimates are that the price of hydrogen from
steam reformation of natural gas produced using traditional methods is 3 to 10 times less than the
hydrogen produced via electrolysis. Thus major improvements in the electrolysis process must be
attained in order to hydrogen produced via this method to be competitive, which is linked directly to
the cost of producing low cost renewable electricity. (Grupta, 2009; European Chemistry for Growth,
2013; Pagliaro &
Konstandopoulos, 2012)
Figure 2. Percentage share in hydrogen production for several feedstocks. Adapted from (Grupta, 2009).
Figure 1. Theoretical energy consumption for hydrogen production from different feedstocks. Adapted from (Grupta, 2009).
5
3.3 Refinery
3.3.1 API & Sulphur Content API gravity is a measurement of the density of oil crudes. The unit is an inverse from specific gravity.
This means that the higher the API value is, the lower the density of the crude is. If the API value is
lower than 10 the crude will sink in water, if it is higher than 10 it will float. API is measured at a
standard temperature of 60°F (15.6°C). Oil crudes used in industry today generally have an API value
of around 30 to 40. (Gary & Handwerk, 2001)
The sulfur content in oil crudes is measured in weight percentage of sulfur. Sulfur content in crudes
differ greatly and generally is between 0.1 to 5 wt. %. If a crude has a sulfur content which is higher
than 0.5 wt. % it is considered a sour crude. (Moulijn, Makkee, & van Diepen, 2013)
3.3.2 Overall Process
The main processes for production and treatment of the crude oil in the refinery can be seen in
Figure 1. This includes the Crude Distillation Unit (CDU), the catalytic processes (dark colored blocks)
and the thermal processes (white colored blocks).
Figure 3 showing the overall process for the refinery (Moulijn, Makkee, & van Diepen, 2013).
3.3.3 Desalting The crude feed usually contains small quantities of salts, trace metals, sand and other contaminants.
Even at low levels these severely increase the chances for fouling of different types of equipment, as
well as poisoning of the catalysts used in later process operations. The process at Preem in
Gothenburg uses a desalting step in order to separate and reduce the concentration of these
contaminants. (Moulijn, Makkee, & van Diepen, 2013; Heinemann & Speight, 2011)
The crude is added to the desalter together with hot water containing some additives. The salts,
trace metals, sand and other contaminants dissolve in the water. The crude and water phases can
then be separated. This is repeated once to make sure the concentrations of the contaminants are
sufficiently low before entering the CDU. (Moulijn, Makkee, & van Diepen, 2013; Heinemann &
Speight, 2011)
6
3.3.4 Crude Distillation Unit Before entering the crude distillation unit, the crude is preheated in a furnace powered by natural
gas. Since there are thousands of different species in the crude oil, the different components in the
CDU are fractioned according to volatility. The CDU at Preems´ refinery in Gothenburg only includes
an atmospheric CDU, as it´s a hydroskimming plant. A hydroskimming plant is one of the simplest
refineries with only atmospheric distillation and treatment of the top flows. There are no further
steps for the residue from the atmospheric distillation tower. This means that thermal processes are
not included. The CDU at Preem’s refinery contain 29 trays with temperatures rising from 190°C at
the top stage to 345°C at the bottom stage. The maximum temperature is set in order to reduce
cracking of the oil and formation of coke. (Moulijn, Makkee, & van Diepen, 2013; Heinemann &
Speight, 2011)
Products from the CDU are taken out at 5 different stages. The CDU has three strippers attached to
the three middle draw stages. From the top to the bottom, the strippers have 10, 6 and 6 stages,
respectively. The bottom products from the strippers are later dried and further processed. Steam is
used for both the side-strippers and at the bottom of the CDU. The top product in the CDU consisting
mainly of lighter products (C1-C16) which is after some separation taken out as LPG, gasoline and
kerosene.
Very little is done to the bottom residue leaving the CDU which is used as heavy firing oil. As
mentioned earlier Preems refinery is a hydroskimming plant which removes the need for vacuum
distillation. Moreover, most thermal processes used on the residue such as visbreaking and propane
deasphalting are not implemented. The residue is stored and sold as fuel oil. (Preem, 2016;
Heinemann & Speight, 2011)
3.3.5 Hydroprocessing In order to obtain a higher yield of the wanted products hydrocracking is used to convert heavier
components to light end products by reaction with hydrogen. This is done since the fuel products are
often more valuable and the environmental laws are getting stricter (diesel, kerosene and gasoline).
Heteroatoms are also removed with hydrotreating. (Moulijn, Makkee, & van Diepen, 2013)
3.3.5.1 Dehydrotreater
The DHT, or hydrotreater, is a process implemented in almost all modern refineries. The purpose of
the process step is to remove heteroatoms which can cause complications in later process steps. The
heteroatoms removed in the process are commonly sulfur, oxygen and nitrogen. In current
hydrotreaters, the main focus is in the removal of sulfur. This is because of sulfur components ability
to deactivate the catalysts used in the reforming, isomerization and alkylation, even at low
concentrations. When hydrotreating heavier oil components a higher hydrogen pressure and higher
temperature is needed. The DHT removes sulfur by reacting it with hydrogen to produce H2S which is
later separated from the rest of the molecules. Usually a cobalt-molybdenum catalyst is used.
Equation 1, 2 and 3 shows the major reactions taking place during desulfurization, reacting
mercaptanes, sulfides and disulfides. It is usually not necessary to further remove oxygen in the
hydrotreating step, but this will have to be implemented in this project since, as mentioned before,
lignin has a fairly high oxygen content, which makes it highly corrosive. The reaction to remove
oxygen is shown in equation 4. (Moulijn, Makkee, & van Diepen, 2013; Gary & Handwerk, 2001;
Heinemann & Speight, 2011)
The reactions are exothermic and equilibrium is favored for low temperatures. The reaction rates are
favored by high temperatures. The temperature needs to be balanced in order to achieve sufficient
conversion while still maintaining a reasonable reaction rate. The reactors generally operate at a
7
temperature of 300°C. The pressures in the reactor is kept above atmospheric. (Moulijn, Makkee, &
van Diepen, 2013; Gary & Handwerk, 2001)
𝑅𝑆𝐻 + 2𝐻2 → 𝑅𝐻 + 𝐻2𝑆 (1)
𝑅2𝑆 + 2𝐻2 → 2𝑅𝐻 + 𝐻2𝑆 (2)
(𝑅𝑆)2 + 3𝐻2 → 2𝑅𝐻 + 2𝐻2𝑆 (3)
𝐶8𝐻8𝑂3(𝑣𝑎𝑛𝑖𝑙𝑙𝑖𝑛) + 5𝐻2 → 𝐶7𝐻8(𝑡𝑜𝑙𝑢𝑒𝑛𝑒) + 𝐶𝐻4 + 3𝐻2𝑂 (4)
3.3.5.2 Hydrocracking
In Hydrocracking heavy gas oils and vacuum gas oils are catalytically cracked into lighter products
such as naphtha, kerosene and diesel fuels. The hydrocracker has had an increasing importance as
transportation fuels has had an increased demand (Moulijn, Makkee, & van Diepen, 2013).
Hydrocracker is not present at Preems plant in Gothenburg.
3.3.6 Thermal Process As hydrocarbons are heated to a high temperature thermal cracking occurs (pyrolysis). There are 3
major cracking processes which are visbreaking, delayed coking and flexicoking. (Moulijn, Makkee, &
van Diepen, 2013)
3.3.6.1 Visbreaking
In visbreaking the viscosity of the vacuum residue is reduced. Visbreaking is a fairly mild thermal
process. Typically the gasoline yield from visbreaking is lower than 10%. Cracked residue is the main
product. (Moulijn, Makkee, & van Diepen, 2013)
3.3.6.2 Delayed Coking
Delayed coking has a much longer residence time than visbreaking which makes it more severe. As
follows from the longer residence time solid residue is formed (petroleum coke and simply coke).
Generally two batch reactors are used in order to enable continuous operations. One reactor
operates while the other one undergoes cleaning. (Moulijn, Makkee, & van Diepen, 2013)
3.3.6.3 Flexicoking
Flexicoking was developed in order to minimize the production of coke during the thermal processes.
The residue oil is fed into a reactor with a hot fluidized bed of coke and it produces gas, liquids and
more coke. Coke accounts for 30 wt. % of the cracking products. (Moulijn, Makkee, & van Diepen,
2013)
3.3.7 Propane Deasphalting By removing “asphaltenic” materials from the distillation products the formation of coke is reduced.
This is called propane Deasphalting. Propane Deasphalting is based on how good the hydrocarbons
are solved in propane. Liquid propane is the most common solvent but ethane, butane and pentane
are also used in many cases. (Moulijn, Makkee, & van Diepen, 2013; Gary & Handwerk, 2001)
Since Preem’s plant in Gothenburg does not have a vacuum distillation unit there are no thermal
processes at Preem. Instead, the heavier fractions are used in the process or sold as heavy fuel.
8
3.3.8 Catalytic Reforming In the catalytic reforming the paraffins and naphthenes undergo two types of reactions during the
conversion to higher octane components, cyclization and isomerization. Reaction conditions have to
be chosen so that the desired reactions are favored and undesirable products are inhibited. Desirable
reactions in a catalytic reformer all lead to the formation of aromatics and isoparaffins. The four
major reactions taking place in the catalytic reformer are the dehydrogenation of naphthenes to
aromatics (Figure 2), dehydrocyclization of paraffins to aromatics (Figure 3), hydrocracking (Figure 4)
and isomerization, which will be discussed in a separate topic. Typical feedstocks to catalytic
reformers are heavy straight run gasolines and naphthenes (82-190°C) and heavy hydrocracker
naphthenes. (Moulijn, Makkee, & van Diepen, 2013; Gary & Handwerk, 2001)
The catalytic reforming reactions are highly endothermic and industry operation conditions utilize
pressures from 15 to 35 bar(a) in order to maximize catalyst life. To maintain high reaction rates the
temperatures are kept between 450 and 500°C. The effluent from the reactor is cooled by heat
exchangers with the feed, and the net hydrogen produced is separated and partially recycled to the
process (for avoiding condensation into coke) and to other parts of the refinery. (Moulijn, Makkee, &
van Diepen, 2013)
The major dehydrogenation reactions occurring are:
Figure 4. Dehydrogenation of alkylcyclohexanes to aromatics. (Gary & Handwerk, 2001)
Figure 5 Dehydroisomerization of alkylcyclopentanes to aromatics. (Gary & Handwerk, 2001)
9
Hydrocracking reactions are exothermic and results in the production of lighter liquid and gas
products. The major reaction involves:
3.3.9 Catalytic Isomerization Since the octane number is higher for branched molecules, the naphtha is taken through an
isomerization process. This process usually uses a platinum catalyst in order to produce isomerized
molecules from n-alkanes such as n-pentane, n-hexane and n-heptane. The temperatures are kept
low during the reaction in order to benefit conversion (150°C). It is also beneficiary to keep the
pressure low, because although the number of moles during the reactions are kept the same, the
more branched nature of the isomers increases the volume of the molecule. (Moulijn, Makkee, & van
Diepen, 2013; Gary & Handwerk, 2001; Heinemann & Speight, 2011)
Figure 8 shows the general reaction taking place during catalytic isomerization. (Moulijn, Makkee, & van Diepen, 2013)
Figure 6 Dehydrocyclization of paraffins to aromatics (Gary & Handwerk, 2001)
Figure 7 Cracking and saturation of paraffins. (Gary & Handwerk, 2001)
10
3.3.10 Alkylation The aim of the alkylation is to form higher branched alkanes. This is mainly done by reacting
isobutene with low molecular alkenes. The reaction takes place with carbenium ions since an acid
solid catalyst is needed. The reactions are favored by high pressures and low temperatures, but since
side reactions in the process has the same favored conditions as the desired reactions, the applied
pressure is often low. (Moulijn, Makkee, & van Diepen, 2013)
The most common reactions in the alkylation is seen in figure 9 and 10.
Figure 9 shows some of the reaction occurring in the alkylation. (Moulijn, Makkee, & van Diepen, 2013)
Figure 10 shows two of the reaction occurring in the alkylation. (Gary & Handwerk, 2001)
11
3.3.11 Green Hydrotreater GHT is a process developed by Preem in order to produce diesel effectively from tall oil into the
process, to produce their evolution diesel which introduces renewable diesel fuels to the market. The
process starts with premixing light gas oil and tall oil. This is done as late as possible before the
reactor in order to reduce strain on the process equipment, as the tall oil is highly corrosive. The
reaction takes place at high pressure (60 bar(a)) and high temperature (360°C). After the reactor the
product goes through some purification steps. First it enters a flash vessel which also removes the
water and separates the different components. As the reactions do not have 100% conversion a
majority of the flow is recirculated. The recirculation stream is later cleaned from sulfur, and is
purged to reduce the buildup of gases in the process. The stream not recirculated is distilled and
dried before being taken to treating and blending. The reactions can be seen in Figure 11. (Preem,
2016)
Figure 11 showing the reaction taking place during the GHT. (Arend, Nonnen, Hoelderich, Fischer, & Groos, 2011)
3.3.12 Octane and Cetane Number The octane number is a measurement of the quality of gasoline. This measurement is the quality of
the flammability and the ignition of the fuel, as well as the fuels ability to not self-ignite under high
pressures and temperatures. This is important for gasoline as the Otto engine depends on the spark
plug of the vehicle to ignite the fuel. If gasoline self-ignites in the engine, which is known as engine
knocking. This reduces the efficiency of the engine and causes damage. The octane number for most
hydrocarbons are usually between 0 and 100 octane. N-heptane has an octane value of 0, while
isooctane has a value of 100. N-heptane has an extremely low octane number because of its linearity
which increases its capability to self-ignite. Isooctane or 2,2,4-trimethylpentane is a highly branched
molecule, which have a lower tendency to self-ignite. This makes it a lot more stable during high
pressures and temperatures, and thus decreasing the chance of self-ignition. A simplified description
of the octane number is that the more branched the molecule is the higher its octane number will
be. (Moulijn, Makkee, & van Diepen, 2013; Demirbas, Balubaid, Basahel, Ahmad, & Sheikh, 2015)
On the other end of the spectrum the cetane number is a measurement of diesel fuel. The
measurements are the quality of the ignition and the ability for the fuel to auto ignite. This defines
the main difference between gasoline and diesel. While gasoline is ignited by a timed spark, diesel
uses self-ignition during compression to function. The cetane number also has a range between 0
and 100. Linear n-alkanes produce the highest numbers as their ability to self-ignite is very high. The
more branched the molecules are the lower their cetane number will be. (Moulijn, Makkee, & van
Diepen, 2013)
12
4. Simulation
HGO
H2
KEROSENEVAPOUT
CDU
DHT
GHT
SEPARATOR
H2
H2
HEAVY RESIDUE
DHT
Figure 12 A block flow diagram showing a simplification of the process simulated in Aspen.
The simulation of the process is carried out in Aspen plus 8.6v, the overall process can be seen in
figure 12. The assays for crude oil was based on Stat-fjord, Oseberg and ESCRCSR , which were
defined by specifying the distillation curve of each. The API characterization corresponding to each
cut from the distillation curve was also specified. The thermodynamic models used for the units are
BK-10 and Peng-Robinson. BK10 was used since it was based on hydrocarbons, and for its ability to
use pseudocomponents. Peng-Robinson was used since it is a very broad method, also because of its
ability to model high temperature process. It was also suggested for simulations of refineries in
aspen.
4.1 Oil Composition The oil-feed composition used in the simulation was based on a massfrac 0.3 of Oseberg (Norway), a
massfrac 0.3 of Statfjord (Norway) and a massfrac of 0.3 of Escravos (Nigeria). The sulfur content in
the combined oil was calculated to 0.225 wt-%. Since Aspen did not have parameters for heavier
components an estimation of two components was made. The sulfur components used in the
simulation was methyl-mercaptane and thiophene-2-hexadecyl. By using the sulfur content in the
different fractions of the oil, the amount of the different sulfur components was calculated to a
massfrac of methyl mercaptane: 0.00021 and a massfrac of thiophene-2-hexadecyl: 0.09569. The
reason why the second massfrac was much higher was because the thiophene-2-hexadecyl had a
much higher molar weight. The molar amount of sulfur did not differ very much.
13
4.2 Crude Distillation Unit The atmospheric column as well as its side strippers were simulated using BK-10 thermodynamic
method. With steam properties calculated through ASME tables from 1967, an equilibrium vapor-
liquid-free water algorithm was used for the simulation.
A rigorous model of PetroFrac was used for the simulation of the tower. Due to its complexity, the
CDU was initially simulated just as a single distillation tower with an integrated partial condenser, to
later have the side strippers, the external top feed which came from the cooler, flash section and the
pump-arounds added. The simulation was kept simple and at a low tolerance in order to ensure
convergence for the different run parameters. The main purpose being a deeper understanding of
the component distribution in the different outlets during different conditions. Just one of the two
physical CDUs was simulated as the other operates parallel and does not interact with the other.
Figure 13. A process flow showing the CDU unit simulated in Aspen.
CDU-1 at PREEM’s refinery in Gothenburg has a total of 29 trays. The top reflux is provided by the
liquid outlet of a flash unit (40°C and 1.4 bar(a)) fed with the cooled top flow from the CDU. In the
flash a tri-phase separation occurs, leading to the reflux stream consisting of liquid heavy-light end
compounds from the crude oil. It should be noted that not all the organic liquid phase is returned as
reflux. The reflux rate is about 3,400 m3/day according to provided data. As it will be explained later
in this section, the main concern is that the temperature profile is kept consistent with the obtained
data. To increase the stability of the simulation the reflux was adjusted to be 55% of the total liquid
obtained from the top flash. This causes an accumulation of liquid in the top of the CDU as well as
increased cooling. The increase in flowrates should however not affect the separation of
FRNICE
FLASH1
STR1
STR2
STR3
MIXER
CONSPLIT
MIXERVAP
MULT
SCALER
SPLIT SID
CDUSTREA FEED
VAPOR
SIDE1
SIDE21
SIDE3
BOT T OM
CDUSTEAM
VAPPROD
FREEW AT E
RECCOND
STESTR1
STESTR2
STESTR3
LLGO
LHGO
HGO
LLGOVAP
LHGOVAP
HGOVAP
SIDE22
LHGOIN
FREEW TCD
STRWT R1
STR2W AT
STR3W AT
CONPRE
MIXVAP
VAPOUT
CRUDEFLO
SIDE212
SIDE211
CDU
14
components, nor the flowrates of the different outlets from the CDU. The additional outlet of a free
water was made in the top of the tower led into a flash vessel in order to be able to achieve a mass
balance over the tower. In reality this water is in the same pipe as the main vapor outlet.
The CDU operates between 1.4 and 2 bar(a). Its temperature profile varies throughout the column
from about 180°C to 350°C. The temperature profile for the physical tower was used as the main
guideline for the simulation.
As the main concern was the temperature profile some simulation tools were used to get this
accurate. Some pump-arounds with desired return temperatures, were introduced through stages 1,
4, 8, 14 and 18. The function of these were to ensure a roughly equal temperature profile within the
CDU during the different operating conditions. A major drawback of this design is the loss of thermal
duty derived from the CDU, as well any changes of steam consumption.
The sizes of side outlets SIDE1 and SIDE22 (see figure 13) was given by PREEM to be 2000 m3/day and
1100 m3/day respectively. These values were used for the simulation for the current CDU, and
slightly decreased for the 3 modified runs. This was done to compensate for the decrease in the
heavy components, as a fraction of the crude mixture is exchanged for the lighter lignin. The side
outlets SIDE21 and SIDE3 were fitted to each simulation since no data for these flowrates were
provided. This led to variations during the different scenarios.
The flowrate of steam fed to the CDU could also be modified, but it was found to be difficult, since
the temperature profile varied with this parameter. This flow value is pretty important for the tower
as too low or too high values would lead to dry stages. The difference in steam consumption has a
direct effect on the cost estimations.
4.3 Dehydrotreater The reactions for the hydrotreaters were defined on a REquil model reactor which implies the
equilibriums are determined by the thermodynamic limitations. The method used was Peng-
Robinson.
Prior to each DHT reactor a separator was introduced as a simulation tool to ensure no thiophene
would get to the reactor, since Aspen does not have parameters for this component. The thiophene
was removed and methyl-mercaptane and 1-eicosyne was added in its place in order to keep the
mass balance correct. The removal of thiophene was done just before the DHTs in order for the
sulfur components to be distilled to the right streams in the CDU.
The parameters in both DHT reactors were set to 300°C and 5 bar(a). The reactions defined were the
following:
Hydrogenation of methyl-mercaptane:
𝐶𝐻4𝑆 + 𝐻2 → 𝐶𝐻4 + 𝐻2𝑆 (5)
The hydrogenation of di-benzothiophene to H2S:
𝐶12𝐻8𝑆(di−benzothiophene) + 5𝐻2 → 𝐶12𝐻16(1−phenyl,1−methylcyclopentane)+ 𝐻2𝑆 (6)
The hydrogenation of vanillin to water, toluene and methane:
𝐶8𝐻8𝑂3(𝑣𝑎𝑛𝑖𝑙𝑙𝑖𝑛) + 5𝐻2 → 𝐶7𝐻8(𝑡𝑜𝑙𝑢𝑒𝑛𝑒) + 𝐶𝐻4 + 3𝐻2𝑂 (7)
15
4.4 Alkylation, Isomerization and Reforming In the simulations the reactions occurring in the different catalytic processes were not calculated,
excluding DHTs and GHT. The assumption was based one main factor. This was the fact that lack of
information about all the reactions taking place in the different reactors were missing. Although the
most common reactions were defined, with petroleum being a very complex compound, adding only
these reactions wouldn’t correspond well to all the actual reactions taking place. Additionally, the oil
was defined as pseudocomponents which aspen could not compute.
4.5 Green Hydrotreater For the GHT simulation a REquil reactor was used with Peng-Robinson as the property method. The
operating temperature and pressure was set to 360°C and 60 bar(a).
The reactions modeled were:
For the hydrogenation of oleic acid:
𝐶18𝐻34𝑂2(𝑂𝑙𝑒𝑖𝑐 𝑎𝑐𝑖𝑑) → 𝐶17𝐻34(9−𝐻𝑒𝑝𝑡𝑎𝑑𝑒𝑐𝑒𝑛𝑒)+ 𝐶𝑂2 (8)
𝐶17𝐻34(9−𝐻𝑒𝑝𝑡𝑎𝑑𝑒𝑐𝑒𝑛𝑒)+ 𝐻2 → 𝐶17𝐻36(𝐻𝑒𝑝𝑡𝑎𝑑𝑒𝑐𝑎𝑛𝑒)
(9)
For the hydrogenation of vanillin:
𝐶8𝐻8𝑂3(𝑣𝑎𝑛𝑖𝑙𝑙𝑖𝑛) + 5𝐻2 → 𝐶7𝐻8(𝑡𝑜𝑙𝑢𝑒𝑛𝑒) + 𝐶𝐻4 + 3𝐻2𝑂 (10)
For the dearomatization of toluene produced during hydrogenation of vanillin:
𝐶7𝐻8(𝑡𝑜𝑙𝑢𝑒𝑛𝑒) + 4𝐻2 → 𝐶7𝐻16(𝑛−ℎ𝑒𝑝𝑡𝑎𝑛𝑒) (11)
Hydrogenation of methyl-mercaptane:
𝐶𝐻4𝑆 + 𝐻2 → 𝐶𝐻4 + 𝐻2𝑆 (12)
The hydrogenation of di-benzothiophene to H2S:
𝐶12𝐻8𝑆(di−benzothiophene) + 5𝐻2 → 𝐶12𝐻16(1−phenyl,1−methylcyclopentane)+ 𝐻2𝑆 (13)
The temperature approach for all reactions was set to 380 °C.
16
4.6 Top Separations & Alkylation The outlet streams of the REquil reactor from the top DHT (DHT1) is connected to a separator which
removes the gaseous H2S. This separator aims to represent the sweetening process for the gaseous
effluent made with amine currents. It is also intended to ensure sulfur free feeding to the
subsequent flash separators.
A sequence of Flash separators were used in order to separate the components into four fractions
with increasing boiling point. These four fractions was later named according to which type of
reaction they would undergo in an actual process. These were, LPG, isomerization, reforming, and
kerosene, which doesn’t undergo a reaction after separation in the simulation. A detailed view of the
schematics of these separators can be seen in figure 14.
For the simulation in Aspen, flash vessels were used with Peng-Robinson as the method for the
calculations. The simulation of the separators used the following temperature conditions: 150 °C for
FLASH11, 25 °C for FLASH12 and 205 °C for FLASH13. All pressures were kept at 1 bar(a). These
conditions were based on the boiling points of the different fuel pools. The isomerization and
reforming fraction is lastly mixed in order to simulate the mixing of the products from these to
reaction steps in the refinery in Gothenburg.
FLASH11
FLASH13
FLASH12
MIXBENS
LPGIN
FLASH11L
LPG
ISO
REF
KEROSINE
GASOLINEDHT1
DHT1SSEP
NULLDHT1
DHT1PROD H2SUT
DHT1DESU
Figure 14. Simulation of Flash Separators for separating LPG, gasoline and kerosene.
17
4.7 Estimating Octane and Cetane Numbers Estimation of the octane number for gasoline was performed based on the methodology presented
in (T.A. Albahri, 2002) which can be seen in Eq. 14:
𝑅𝑂𝑁 = 𝑎 + 𝑏 𝑇 + 𝑐 𝑇2 + 𝑑 𝑇3 + 𝑒 𝑇4 (14)
Where RON is the predicted Research Octane Number. This can also be measured experimentally
under low speed condition by ASTM D 908, and T = Tb/100 in which Tb is the normal boiling point in
“°C” and coefficients a through e are given in Table 2 below:
Hydrocarbon family a b c d e
n-Paraffins 92.809 -70.97 -53 20 10 iso-Paraffins 2-Methylpentanes 95.927 -157.53 561 -600 200 3-Methyl-pentanes 92.069 57.63 -65 0 0 2,2-Dimethyl-pentanes 109.38 -38.83 -26 0 0 2,3-Dimethyl-pentanes 97.652 -20.8 58 -200 100 Naphthenes 77.536 471.59 -418 100 0 Aromatics 119 144.8 -12 0 0
The RON is calculated assuming that the fuel is a mixture of four model compounds from n-paraffins,
iso-paraffins, naphthenes and aromatic families. The RON of the mixture is afterwards calculated
according to eq. 15:
𝑅𝑂𝑁 = 𝑥𝑁𝑃 (𝑅𝑂𝑁)𝑁𝑃 + 𝑥𝐼𝑃 (𝑅𝑂𝑁)𝐼𝑃 + 𝑥𝑁 (𝑅𝑂𝑁)𝑁 + 𝑥𝐴 (𝑅𝑂𝑁)𝐴 (15)
Where 𝑥𝑁𝑃,𝑥𝐼𝑃, 𝑥𝑁 and 𝑥𝐴 are volume fractions of n-paraffins, iso-paraffins, naphthenes and
aromatic groups, respectively.
Estimation of the cetane number was performed using the methodology presented in (Stournas,
1992), which is based on the aromaticity, the distillation curve and the density of the fuel, as seen in
Eq. 16:
𝐶𝑒𝑡𝑎𝑛𝑒 𝑁𝑢𝑚𝑏𝑒𝑟 = 𝑎 𝐼𝑃 + 𝑏 𝐷10 + 𝑐 𝐷50 + 𝑑 𝐷90 + (𝑒
𝐸𝑃) + 𝑓 (
1
𝐷𝐸𝑁𝑆2) + 𝑔 (1
𝐴𝑅𝑂𝑀4) + ℎ (16)
Where IP is the initial boiling point in “°C”, Dn is “n% volume recovered in °C”, EP is the end boiling
point in “°C”, DENS is the specific gravity at 15°C and AROM is the percentage weight of aromatics.
The coefficients a through h are seen in Table 3 below:
Table 3. Coefficients for calculation the cetane number using Equation 16.
Const. Value
a -0.011
b 0.092554
c 0.119366
d 0.130821
e 7083.031
f 110.258
g 16096.35
h -219.705
Table 2. Coefficients for RON estimation.
18
The fuel aromaticity (AROM) can be estimated using Eq. 17:
𝐴𝑅𝑂𝑀 = 𝑎 𝐼𝑃 + 𝑏 𝐷10 + 𝑐 (𝐷50 )0.5 + 𝑑 𝐷90 + 𝑒 𝐸𝑃 + 𝑓 (1
𝐷𝐸𝑁𝑆2) + 𝑔 (17)
Where the coefficients from a to g are given in Table 4:
Table 4. Coefficients for estimating aromaticity of fuel based on Eq. 17.
Const. Value
a -0.25931
b 0.514474
c -13.157
d -0.03947
e 0.059787
f -166.654
g 400,452
Estimation on the cetane number using the presented methodology can have an error up to 6% in
comparison to the experimental ASTM D 976 methodology. (Stournas, 1992)
19
4.8 Adding Scenarios In order to decide where the lignin should be added a three different scenarios were investigated. In
these scenarios the distribution of the lignin was investigated in order to determine in which fuel
pools they ended up. The lignin, represented by vanillin is mixed so that it at most is 10 vol. % of the
total inflow in the CDU. This is done in order to minimize corrosion.
In scenario one the vanillin is mixed with the crude oil before entering the CDU, without any
pretreatment. In this scenario an additional dearomatization step of the vanillin products
must be added in order to get the right hydrogen usage. This is crucial since diesel should
contain straight chains and not aromatics. The crude flow in this scenario is corresponding to
95% of the original crude flow.
HGO
H2
KEROSENEVAPOUT
CDU
DHT
DHT
GHT
SEPARATOR
H2
H2
HEAVY RESIDUE
Figure 15. A block flow diagram showing a simplification of the process simulated in Aspen for Scenario 1.
20
The second scenario is when Lignin is mixed with the top outlet flow before it enters the
DHT. The crude flow in this scenario is corresponding to 95% of the original crude flow.
HGO
H2GASOLINE
KEROSENEVAPOUT
SEPARATOR
H2
H2
HEAVY RESIDUE
LIGNIN
GHT
DHT
DHT
CDU
Figure 16. A block flow diagram showing a simplification of the process simulated in Aspen for Scenario 2.
In scenario 3 the vanillin is mixed with hydrotreated lignin (recirculated) and then entered to
the CDU. The crude flow in this scenario is corresponding to 95% of the original crude flow.
HGO
H2
KEROSENEVAPOUT
CDU
DHT
GHT
SEPARATOR
H2
H2
HEAVY RESIDUE
DHT
RECIRKULATION
PRETREATMENT
Figure 17. A block flow diagram showing a simplification of the process simulated in Aspen for Scenario 3.
21
4.9 Assumptions and Considerations for the Simulation The simulation is based on the assumption that vanillin was the only main lignin component. The
validity of this assumption is directly depending on the quality of the pretreatment. The outcome of
the different scenarios would vary widely depending on the composition of the entering lignin. With
vanillin being a fairly simple aromatic molecule it will break down into light components, with
toluene as the heaviest. There is a possibility that during pretreatment or hydrogenation of the lignin
in the DHT, the resulting compounds would consist of a mixture of lighter compounds such as,
methane, toluene and water, as well as more complex molecules consisting of multiple
interconnected aromatics connected by carbon-carbon bonds. These complex molecules would in all
probability alter the properties of the resulting products in the process if implemented. This would
lead to a more complex pretreatment step being required than the one simulated in Aspen.
Another issue with the simulation is that some data concerning the CDU was missing, leading to
some approximation giving a different result when it comes to the different scenarios. For example,
the flow in the different parts of the CDU was defined but it was not changed when the different
scenarios were played out. This could have a great effect on the spread of the different fuel pools.
The same problem occurs in the scenario when the vanillin is added in the DHT, this is because the
flows to the flash vessels are increased. This may be one explanation why a lower amount of toluene
ends up in the gasoline pools.
Since aspen did not have parameters for the heavier sulfur component the approximation of only
using two different sulfur components made the simulation quite simplified giving us a much simpler
Another problem with the simulation is that no reactions of the crude oil concerning the
isomerization and reforming is performed. This will have a great effect of the octane number as well
as the cetane number. It will also give a quite different production and usage of hydrogen since large
amounts of hydrogen is produced in reforming.
As can be seen in table 9 the production of gasoline increased in the DHT step only. In both the other
scenarios the gasoline yield decreased. We can also see that in all cases but the CDU the diesel pool
was reduced. This was partly because the inflow of crude oil was reduced to 95% of the original flow.
The second reason why the flows may change is because of the CDU (mentioned above).
Depending on the quality of the pretreatment of the lignin, the true composition (that will differ
from assumed 99.9% vanillin) might in a hydrotreater react into heavier products than toluene. If
these aromatic-carbon chains are long enough they might end up in the kerosene pool. Depending
on the size of these products it may be more desirable to add a flash vessel in which the vapor would
be of a volatility desired for kerosene and where the liquid is directed into the dearomatization step
for the diesel production.
22
5. Cost Estimation For the economic analysis the consumption of energy and material was compared for each scenario.
Most flows are assumed to remain the same in the three different scenarios, and thus they are not
reported on the economic analysis. Since the plant is already constructed and the economic impact
of the addition of lignin to the refinery process, only changes the flows and energy consumptions,
only these will be taken into account.
In order to estimate the effect from the lignin addition to the process, and not focusing on the
expenses in terms of consumption, but also taking into account the profit from its addition, the
changes on product currents are taken into account. Moreover, an estimation over the change in
cetane number and octane number for the diesel and gasoline currents is proposed.
For the calculation of utility usage, it was assumed that water was used for refrigeration with a
thermal difference at lowest of 5°C. No heat integration was considered. For heating purposes high
and medium pressure steam was used to ensure at least 30°C as thermal gradient between utility
and hottest outlet current.
The price of oil was based on current prize at middle of May 2016, which was 44.08 USD per
American barrel. Four scenarios were considered where the prices for crude were considered to
remain the same as actual, double and four times as actual prize, as well as calculated lignin price
and twice and four times the calculated lignin cost. For the lignin-adding scenarios, lignin replaced 5%
of the mass flow rate for crude oil, thus for the non-lignin scenario the crude oil cost was estimated
as 886,850 kr/h, while for all the lignin-adding scenarios the cost of the crude was estimated as
853,950 kr/h.
5.1 Lignin Cost The cost for the lignin feedstock can be difficult to estimate, due to different types of lignin that can
be obtained depending on the process conditions in the paper industry. In this report an indirect,
more conservative approach was chosen to estimate the cost of lignin. Since most of the lignin is
combusted to produce heat used internally in the paper industry, the estimation was based on its
energy content (average on hard wood and soft wood lignin), specifically the higher heating value
(HHV). A mean between the hardwood lignin (HHV = 23.50 MJ/kg) and softwood lignin (HHV = 21.45
MJ/kg) was used to estimate the energy content for lignin. (Blunk & Jenkins, 2000)
For the different scenarios the lignin inflow was set at 17,122 kg/hr, thus the cost for the lignin
feedstock based on its energy content was estimated to 60,075 kr/hr.
23
6. Results and Discussion The results from the different scenarios are shown in the figures below.
The temperature profile in the CDU simulated is shown in figure 18 without lignin present in the
process. The temperature in the different stages is shown in °C. The temperature in the CDU was
supposed to replicate the temperature of the given CDU. The Profile below gave a good
approximation of the given data.
Figure 18. The temperature profile obtained from the simulation of the CDU without the addition of lignin.
Block CDU: Temperature Profile for Main Column
Stage
Tem
pera
ture
C
1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29180
200
220
240
260
280
300
320
340
360
Temperature C
24
Figure 19. shows the amount of oil and lignin in the different fuel pools for the different scenarios.
As can be seen in figure 19 the amounts of the different fuel pools varied depending on where the
lignin was added. The scenario giving the highest yield in the gasoline pool was the DHT adding
scenario. The diesel has a higher yield in the CDU which is not desired since most of the toluene
should be in the gasoline. All the scenarios gave a lower amount of Heavy Fireing oil. The pretreated
vanillin scenario gave a lower yield of the different fuels than original, although it gave a higher
kerosene yield.
The increased production of gasoline and diesel for the different scenarios can be seen in table 5.
Table 5 showing the change in percent in the gasoline pool and the diesel pool.
% gasoline % diesel
CDU -8.23 2.34
DHT 7.37 -2.76
PRE -5.61 -17.05
As can be seen the production of gasoline yield only increased in the DHT case while the diesel only
increased in the CDU adding case.
0
20000
40000
60000
80000
100000
120000
140000
LPG Gasoline Kerosene Diesel Heavy FireingOil
[kg/
h]
Production Pools
Original
Vanillin In CDU
Vanillin in DHT 1
Pretreated Vanillin
25
The usage of hydrogen for the different scenarios is plotted in figure 18. All the scenarios are
compared to the run with regular oil.
Figure 20. showing the consumption of hydrogen in the different blocks of the process for the different scenarios.
As the figure shows the Vanillin directly in the CDU caused the highest hydrogen use. Overall the
hydrogen consumption increase for all the scenarios, this since the treatment of lignin demanded
hydrogen. The reason why the CDU scenario needed the most hydrogen was because the vanillin
ending up in the Diesel pool needs further hydrotreatement in order to increase cetane number.
0
500
1000
1500
2000
2500
3000
Pretreatment DHT1 DHT2 GHT TotalConsumption
[kg/
h]
H2 consumption
RunWithoutVanillin
VanillaInCDU
VanillinDHT1
Pretreated Vanillin
26
The composition in the gasoline and diesel pool for the different scenarios is plotted in figure 19 and
20 respectively.
Figure 21. showing the composition of the gasoline pool for the different scenarios.
As can be seen in figure 21 the case where the higher amount of toluene ended up in the gasoline
was the pretreated scenario.
Figure 22. showing the mass fractions of the diesel for the different scenarios.
0
0,05
0,1
0,15
0,2
0,25
0,3
0,35
Mass Fraction of Diesel
Diesel Original
Diesel 5 mass% Vanillin into CDU
Diesel 5 mass% Vanillin into DHT1
Diesel 5 mass% Vanillin pretreated
0
0,05
0,1
0,15
0,2
0,25
0,3
0,35
0,4
0,45
0,5Mass Fraction of Gasoline
Gasoline Original
Gasoline 5 mass% Vanillin into CDU
Gasoline 5 mass% Vanillin into DHT1
Gasoline 5 mass% Vanillin pretreated
27
The amount of lignin as toluene that reached the gasoline fraction for the different scenarios is
depicted in figure 23.
Figure 23. Showing how much of the lignin as toluene in mole-% ended up in the gasoline product.
The mass-% of vanillin that is converted into toluene and located in the gasoline pool is depicted in
figure 24.
Figure 24. mass-% of vanillin as toluene in the gasoline pool.
Figure 23 indicates how close the process is to its maximum yield of toluene in the gasoline pool
whilst figure 24 shows the true amount of product per mass unit vanillin. The reason for the lower
numbers in figure 24 compared to figure 23 is the production of water and methane in the
hydrotreater.
0
0,1
0,2
0,3
0,4
0,5
0,6
0,7
0,8
0,9
1
CDU Fed DHT Fed Pretreated
Mole-% of the lignin as toluene in the gasoline
0
0,1
0,2
0,3
0,4
0,5
0,6
0,7
0,8
0,9
1
CDU Fed DHT Fed Pretreated
Mass-% of the Vanillin as toluene in the gasoline pool
28
6.1 Pumping Costs Pumping costs were estimated for each scenario based on the different pumping necessities of the
scenarios, which are: the inlet crude flow (CRUDEFLO), the outlets for the first and second DHT
(DHT1DESU, DHT1PROD, DHT2VAP, DHT2LIQ), the outlets for diesel (DIESEL) and tall oil (TALLOIL).
Summarized costs can be seen in Table 5 below:
Table 6. Pumping costs for the different scenarios given in kr.
Scenario CRUDEFLO DHT1DESU DHT1PROD DHT2LIQ DHT2VAP DIESEL TALLOIL TOTAL (kr/hr)
No Lignin 219 13.14 13.14 2.43 6.26 235 1.26 491
Lignin CDU 211 14.18 14.20 2.25 6.23 239 1.28 489
Lignin DHT1 211 15.31 15.33 2.37 6.11 229 1.23 481
Lignin Pretreated 211 16.25 16.26 1.51 4.82 191 1.04 443
6.2 Raw Material Costs Costs for the consumption of natural gas, water and hydrogen were estimated based on average
prices obtained from (Hydrogen Fuel Cost vs Gasoline, s.d.) (Natural Gas Price Statistics, s.d.) (Svenskt
Vatten, s.d.), and summarized in Table 7:
Table 7. Average prices considered for the estimation of material costs.
Item Unit Price (Cost) Reference Year
Natural Gas SEK/kWh 0.425 2015 (annual average)
Water SEK/ton 0.014 2013 (annual average)
Hydrogen SEK/Kg 38.385 2016 (april average)
The material costs were estimated for each scenario by considering the different mass flows. The
non-lignin scenario was used as comparison for the others scenarios where lignin is added to the
process. Costs for each scenario can be seen in Table 8:
Table 8 Summarized material costs. % increase is the increase in material usage considering No-lignin scenario as reference.
Utility Costs (SEK/h)
Scenario Natural Gas Water Hydrogen Total % increase
No-lignin 49,000 962 32,000 82,000 -
Lignin-CDU 52,000 946 75,000 127,000 55.6
Lignin-DHT 44,000 922 74, 000 119,000 45.5
Lignin-Pretreated 44,000 922 74, 000 119,000 45.5
6.3 Investment Costs The addition of lignin to the current refinery processes requires additional equipment, such as piping,
storage tanks and pumps. Estimation of investment costs was performed for the centrifugal pumps
need based on catalog prices. Summarized information used to estimate the equipment costs can be
seen in Table 9: DHT (DHT1DESU, DHT1PROD, DHT2VAP, DHT2LIQ) and the outlets for diesel (DIESEL)
and tall oil (TALLOIL) were categorized as one of the pump numbers as can be seen in Table 9:
Accounting for the different pumping necessities, such as volumetric flow rates, motor power and
pressure drop for each scenario, the inlet crude flow (CRUDEFLO), the outlets for the first and second
29
DHT (DHT1DESU, DHT1PROD, DHT2VAP, DHT2LIQ) and the outlets for diesel (DIESEL) and talloil
(TALLOIL) were categorized as one of the pump numbers as can be seen in Table 9:
Thus the overall investment cost on pumps is estimated as 1,241,000 kr.
Estimation on piping and storage tanks, plus any extra infrastructure needed was done by using
Ulrich´s factor methodology. Ulrich suggests using a 30% over the bare module cost to estimate the
cost for auxiliary facilities: 2 – 6% to auxiliary buildings, 17 – 25% to auxiliary processes and 43 – 63%
on equipment for installation costs. (Ulrich, 1984)
Estimating the storage tank as auxiliary buildings and pipes as auxiliary process equipment:
Storage Tanks (6%) = 74,500 kr, pipes, valves and control (25%) = 310,000 kr, installation costs (63%
on equipment) = 1,024,000 kr. Total investment cost is estimated as 2,300,000 kr.
6.4 Results on Octane and Cetane Numbers The octane and cetane numbers were calculated for each scenario in order to estimate the impact of
adding lignin into the process. Results are summarized in Table 11:
Table 11. Results for octane and cetane numbers estimation for different scenarios
Scenario Octane# Cetane#
No-lignin 97.7 57.8
Lignin-CDU 98.3 57.7
lignin-DHT 100.1 57.9
Lignin Pretreated 95.6 57.5
It can be seen that the diesel cetane number is practically the same, not being affected by the
differences in the scenarios. The gasoline octane number, increases for scenarios in which lignin is
added to the CDU and even more when lignin is added to the directly DHT.
Table 10. Cost estimation of pumping for each scenario based on classification.
Table 9. Soma catalogs used to estimate pump costs.
CRUDEFLO DHT1DESU DHT1PROD DHT2LIQ DHT2VAP DIESEL TALLOIL
Pump Number 1 2 2 2 2 4 2
Cost (kr) 34,000 23,000 23,000 23,000 23,000 1,094,000 23,000
Number Pump Type Capacity (m3/h) Head (m) Cost Cost (SEK) Catalog Brand
1 Centrifugal 7,500 500 4,200 34,000 BB1 PumpWorks
2 Centrifugal 1,200 320 2,800 23,000 OH2 PumpWorks
3 Centrifugal 450 700 6,100 49,500 RON Ruhrpumpen
4 Centrifugal 20,000 720 135,000 1,094,000 ZMII Ruhrpumpen
30
6.5 Best Scenario The most important criteria’s to look at when deciding which scenario is the best is in which fuel pool
the lignin ends up and the usage of hydrogen. Based on the prices today Gasoline and Diesel are the
most desired products. By combining the amounts produced with the amount of hydrogen used the
different scenarios are compared.
6.6 Crude Distillation Unit The pros of using the scenario when vanillin is mixed with the CDU stream untreated is that very little
changes in the process. The flow rates in the CDU will remain the same. A problem with the scenario
is that a lot of the toluene ends up in the diesel pool which requires more hydrotreatment and the C7
chain is in the smallest parts of the diesel.
6.7 Pretreatment The benefits of using the pretreatment scenario before adding the lignin to the CDU it that all
components entering the CDU is treated meaning a lower corrosion as well as a greater part of the
lignin ending up in the gasoline pool. Since lower amounts end up in the diesel pool lower amounts
of hydrogen is needed.
6.8 Dehydrotreater The DHT scenario has the benefits of minimizing the flows in the CDU which save a lot of energy and
it sends the lignin to the pools where it (based on products of hydrotreatment of vanillin) should end
up. This way all the toluene reacted will be in the top flow which is where it would have the best
effect. The downside with this scenario is that a high concentration of lignin in the flow before the
hydrotreated is obtained which may have a corrosive effect. To avoid that corrosive effect, the
products from the reactor could be recycled in order to dilute the lignin, and lignin should be added
as late as possible into the reactor.
In this case the DHT adding scenario and Pretreatment scenario gave the best outcome. This was
because the products attained from the hydrotreatment of lignin is mainly toluene, methane and
water. With toluene being used commercially as an octane booster the DHT adding scenario is the
one which seems most logic, as the presence of toluene in the diesel fraction would result in a lower
cetane number.
6.9 Other Alternatives Another alternative for adding the lignin is to introduce it in the GHT. This scenario was not simulated
since the products received from treating vanillin is not suitable in the diesel composition. Also in the
scenario where the vanillin was added directly to the CDU most of the vanillin ended up in the GHT.
Considering this the result (excluding heat and lower repair cost because of less exposure of
corrosive vanillin) would be quite identical to the CDU scenario. Although there would probably be
some savings in material cost and heating when adding the vanillin in GHT instead of the CDU.
6.10 Further Findings It can also be seen that the scenario adding vanillin directly to the CDU has the lowest amount of
toluene in the gasoline pool. By comparing massfrac of gasoline in figure 19 a fair estimation of how
much of the toluene that ends up in the gasoline pool can be made. This way the effectiveness of the
process can be estimated.
Even though vanillin is not a precise estimation of the composition of lignin most of the lignin
components are based on aromatic rings which gives the same indication as the vanillin gives.
31
Based on this the overall product (including all other products) will probably result in aromatic rings,
smaller methyl groups and water. With this said vanillin is probably a quite good guide to where the
lignin should be added. This is a theory that indicates that the adding of lignin in the DHT should be
the most efficient scenario.
Kerosene is the product that is least favorable, this since there is very low taxes on this fuel. When it
comes to diesel and gasoline the taxes are higher which makes renewable gasoline and diesel more
desirable since environmental taxes are much lower for renewable fuels.
In the third scenario when lignin is pretreated before entering the CDU a phenolic compound
hydrogenation could be carried out, by mixing lignin with water and methanol and optimizing the
pressure and temperature (YU, o.a., 2013).
6.11 Cost Discussion The operation cost estimation considered only the differences in each scenario when compared to
the No-lignin reference scenario. The total operating costs is presented in Table 12 below, this
includes the sum of the material (lignin, natural gas, water and hydrogen) and energy costs. It should
be emphasized that these are not the actual total costs for all the processes in the refinery, but
rather the estimated differences in costs for each scenario.
Table 12. Total costs for each scenario is shown below along with different scenarios where the oil price and lignin is changed. The original scenario is with the current oil price [392 kr/barrel] and the ordinarily calculated lignin cost.
The profitability is closely tied to the oil price as well as the lignin stock price. Economical profit is
only achieved when the oil price is 4 times higher than today. Most likely the pretreated lignin will
have a higher cost than the originally estimated cost based on the heating value. This since there is a
lot of cost concerning the separation of lignin from black liquor. The prices do not consider the
environmental taxes which will probably have a positive effect on the scenarios using lignin.
Since the amount of lignin added to each scenario is the same, analysis on the costs should be
focused on the energy consumption (pumping) and increase in raw materials.
The Lignin-CDU scenario had the highest cost increase. This was expected due to the larger amount
of energy needed for heating the added lignin stream into the CDU. The Lignin-DHT and Lignin-
Pretreated scenario had roughly the same increase in costs, mostly due to the increase in material
usage, mainly in hydrogen consumption, which was increased 2.37 times. The pumping costs can be
considered roughly the same for each scenario, it is below 0.5% of the total operating costs and can
therefore be neglected. Major differences in material usage were found for each scenario. The lignin
costs account for roughly one third of the total operating cost in the different adding scenarios. The
consumption of natural gas increased in the Lignin-CDU scenario. This is mostly due to the increase in
heat necessary to preheat the feed to the CDU. The consumption of natural gas decreased in the
Scenario Original
Double the Oil Price
Quadruple the Oil Price
Double the Lignin Price
Quadruple the Lignin Price
Total (kSEK)
% increase
Total (kSEK)
% increase
Total (kSEK)
% increase
Total (kSEK)
% increase
Total (kSEK)
% increase
No-lignin 969 - 1,856 - 3,630 - 969 - 969 -
Lignin-CDU 1,042 7.5 1,896 2.1 3,604 -0.72 1,102 13.70 1,222 26.09
Lignin-DHT 1,034 6.6 1,888 1.7 3,595 -0.95 1,094 12.84 1,214 25.28 Lignin-
Pretreated 1,035 6.8 1,889 1.7 3,597 -0.91 1,095 12.99 1,215 25.38
32
Lignin-DHT and Lignin-Pretreated scenarios. This since less heat is need to drive the CDU (only 95% of
the original crude mass flow now enters the CDU). Water consumption was kept roughly constant in
every scenario.
A spectrum analysis was also made to estimate how sensitive are the costs based on the lignin price.
Thus, the operational costs were estimated for half the estimated lignin cost, where the Lignin-CDU,
Lignin-DHT and Lignin-Pretreated scenarios had an increase of 4.4%, 3.5% and 3.7% respectively
when compared to the no-lignin scenario. For doubling the lignin price, the operational costs had an
increase of 10.6%, 9.7% and 9.9% respectively.
6.12 Gasoline limitations As it can be seen in the pie charts from Appendix A. The amount of toluene in the gasoline pool is
increased with the introduction of vanillin. The smallest mass fraction increase occurs when
introducing the lignin into the CDU tower (104% increase) and large increases are obtained when
adding pretreated lignin to the CDU and lignin on the DHT1 (371% and 374% respectively). The
increasing toluene fraction in gasoline when introducing the lignin as part of the vanillin is no surprise
because of the boiling range of vanillin and toluene. The increased content of toluene impacts the
pool Octane number as seen in table 11 due to octane number for toluene being above 100.
Advantage may be taken from this high aromatic content since gasoline-ethanol mixtures could be
produced in order to dilute the aromatics but keeping the octane number requirements. Considering
that the limits for aromatics (OK-Q8, 2009)in gasoline is about 40 vol-% aromatics and 1 vol-%
benzene, although the limitations have increased in the past years, the gasoline pool would probably
meet the environmental requirements. Also it should be considered the possibility of mixing the high
octane gasoline pool with a low quality gasoline to dilute the aromatic content until matching
environmental and quality requirements.
7. Conclusion Based on the study done in this paper the best scenarios were adding the lignin in either the DHT or
as pretreated lignin in the CDU. The DHT adding scenarios led to the smallest decrease on diesel pool
production and the gasoline pool was even increased by 7% roughly. Furthermore, those two
scenarios have the fewer increases in operational cost (6.6% DHT and 6.8% Pretreated-CDU), also the
DHT adding scenario had the highest increase in quality for the gasoline pool. As it can be seen in
Figure 21 the amount of lignin ending in desired fuel pools is close to 100% for the in-DHT adding
scenario and CDU-Pretreated. Nevertheless, it should be considered that any lignin adding scenario
would lead to an increase in hydrogen usage about 100%.
Although operational costs increase with the addition of lignin, the production of the desired product
pools is also increased, both in absolute values and in quality of the fuels, enabling the possibility to
further increase production by diluting the higher quality fuels with lower quality ones. Therefore,
there is an economic justification for investment in using lignin as feedstock.
Major drawbacks in using lignin as feedstock for the refinery is the increase in demand of hydrogen
for its treatment. Current production of hydrogen depends on hydrocarbon fuels and thus are not
free from greenhouse effect emission.
33
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35
Appendix A. Mass fractions in gasoline for different scenarios and raw data
from Aspen.
Appendix A1
Gasoline Original
N-BUTANE I-PENTAN N-PENTAN TOLUENE
PSEUDO COMPONENT 1 PSEUDO COMPONENT 2 PSEUDO COMPONENT 3 PSEUDO COMPONENT 4
Figure A1. A pie chart showing the mass fraction of different components in the outlet gasoline, without the addition of Lignin.
36
Table A1 Component mass flow, Standard vol. fraction and mass fraction data for different streams taken from Aspen. (Without the addition of vanillin).
Component Mass Flow LPG Gasoline Kerosene Diesel Heavy Fireing Oil
METHANE KG/HR 33,5749 15,72619 0,000113 0 1,03E-16
ETHANE KG/HR 6,639443 18,57475 0,000327 0 4,61E-16
PROPANE KG/HR 91,92039 928,9048 0,047911 3,08E-10 1,41E-12
I-BUTANE KG/HR 29,29861 761,4451 0,09517 3,21E-09 1,12E-11
N-BUTANE KG/HR 80,69543 2993,596 0,515291 2,04E-08 7,15E-11
I-PENTAN KG/HR 28,5007 2669,788 1,160108 2,19E-07 6,60E-10
N-PENTAN KG/HR 31,21662 3999,103 2,220047 3,83E-07 1,10E-09
CYCLO-01 KG/HR 3,548052 666,201 0,577622 2,13E-07 4,56E-10
TOLUENE KG/HR 1,284907 2627,62 19,70205 6,32E-04 1,90E-07
BENZENE KG/HR 1,901836 963,2187 2,100928 4,07E-06 7,69E-09
WATER KG/HR 4,841755 24,80712 0,000286 1,38E+01 0,853474
METHY-01 KG/HR 0,000323 0,011011 1,69E-06 7,686644 2,56E-12
H2 KG/HR 1,438697 0,075348 3,32E-07 0,00E+00 0
OLEIC-01 KG/HR 0 0 0 8638,431 0
THIOP-02 KG/HR 0 0 0 0 6,26E+03
CARBO-01 KG/HR 0 0 0 622,1274 0
N-HEP-02 KG/HR 0 0 0 1451,335 0
1-EIC-01 KG/HR 3,10E-09 3,001599 25,94549 23650,84 0,00E+00
1-HEP-01 KG/HR 0 0 0 1931,675 0
PS1 KG/HR 18,19302 45666,11 424,2284 0,000717 1,05E-06
PS2 KG/HR 0,463749 23302,84 1960,832 1,054372 4,73E-05
PS3 KG/HR 0,01948 20228,56 8600,857 890,9996 0,003849
PS4 KG/HR 9,92E-05 2866,668 5085,041 24500,69 0,569381
PS5 KG/HR 5,27E-09 5,661671 4,25E+01 30431,71 99,91866
PS6 KG/HR 1,75E-14 0,000694 0,022454 2,00E+04 2,13E+03
PS7 KG/HR 0 1,52E-11 3,16E-09 5961,586 7119,772
PS8 KG/HR 0 0 0 0 90065,58
Mass Flow KG/HR 333,538 107742 16165,89 118068 105676 Volume Flow M3/DAY 4908,209 149795 580,5542 4638,155 3376,563
Component Std. Vol. Fraction
METHANE METHANE 0,166079 3,66E-04 1,92E-08
ETHANE ETHANE 0,027714 3,65E-04 4,67E-08
PROPANE PROPANE 0,269154 0,01281 4,80E-06 4,40E-15
I-BUTANE I-BUTANE 0,077321 0,009464 8,59E-06 4,13E-14 1,78E-16
N-BUTANE N-BUTANE 0,205119 0,035837 4,48E-05 2,53E-13 1,10E-15
I-PENTAN I-PENTAN 0,067788 0,029906 9,44E-05 2,54E-12 9,47E-15
N-PENTAN N-PENTAN 0,073511 0,044351 0,000179 4,40E-12 1,57E-14
CYCLO-01 CYCLO-01 0,007023 0,00621 3,91E-05 2,06E-12 5,45E-15
TOLUENE TOLUENE 0,002188 0,021077 0,001148 5,25E-09 1,96E-12
BENZENE BENZENE 0,003193 0,007616 0,000121 3,33E-11 7,78E-14
WATER WATER 0,007188 1,73E-04 1,45E-08 9,96E-05 7,64E-06
METHY-01 5,49E-07 8,82E-08 9,83E-11 6,37E-05
37
H2
0,056635 1,40E-05 4,47E-10
OLEIC-01 OLEIC-ACID 0,069848
THIOP-02
0,05327
CARBO-01 CARBON-DIOXIDE 0,005469
N-HEP-02 N-HEPTADECANE 0,013445
1-EIC-01 1-EICOSYNE 5,77E-12 2,63E-05 0,001651 0,214574
PS1 PSEUDO COMPONENT 1 0,036175 0,427634 0,028853 6,95E-09 1,25E-11
PS2 PSEUDO COMPONENT 2 0,000876 0,207364 0,12673 9,71E-06 5,39E-10
PS3 PSEUDO COMPONENT 3 3,54E-05 0,173149 0,534702 0,007896 4,22E-08
PS4 PSEUDO COMPONENT 4 1,73E-07 0,023591 0,30393 0,208732 6,00E-06
PS5 PSEUDO COMPONENT 5 9,03E-12 4,57E-05 0,002494 0,254249 0,001033
PS6 PSEUDO COMPONENT 6 2,89E-17 5,40E-09 1,27E-06 0,160974 0,021236
PS7 PSEUDO COMPONENT 7 1,75E-13 0,046906 0,0693
PS8 PSEUDO COMPONENT 8 0,855147
Component Mass Fraction
METHANE 0,100663 1,46E-04 7,01E-09 0 9,75E-22
ETHANE
0,019906 1,72E-04 2,02E-08 0,00E+00 4,36E-21
PROPANE 0,275592 0,008622 2,96E-06 2,61E-15 1,33E-17
I-BUTANE
0,087842 0,007067 5,89E-06 2,72E-14 1,06E-16
N-BUTANE 0,241938 0,027785 3,19E-05 1,73E-13 6,77E-16
I-PENTAN 0,08545 0,024779 7,18E-05 1,85E-12 6,24E-15
N-PENTAN 0,093592 0,037117 0,000137 3,25E-12 1,04E-14
CYCLO-01
0,010638 0,006183 3,57E-05 1,81E-12 4,32E-15
TOLUENE
0,003852 0,024388 0,001219 5,35E-09 1,80E-12
BENZENE
0,005702 0,00894 0,00013 3,45E-11 7,27E-14
WATER
0,014516 2,30E-04 1,77E-08 1,17E-04 8,08E-06
CARBO-01 CARBON-DIOXIDE 0 0 0 0,005269 0
N-HEP-02 N-HEPTADECANE 0 0 0 0,012292 0
1-EIC-01 1-EICOSYNE 9,29E-12 2,79E-05 0,001605 0,200316 0
1-HEP-01
0 0 0 0,016361 0
PS1 PSEUDO COMPONENT 1 0,054546 0,423847 0,026242 6,07E-09 9,90E-12
PS2 PSEUDO COMPONENT 2 0,00139 0,216284 0,121294 8,93E-06 4,48E-10
PS3 PSEUDO COMPONENT 3 5,84E-05 0,18775 0,532037 0,007547 3,64E-08
PS4 PSEUDO COMPONENT 4 2,97E-07 0,026607 0,314554 0,207513 5,39E-06
PS5 PSEUDO COMPONENT 5 1,58E-11 5,25E-05 0,002632 0,257747 0,000946
PS6 PSEUDO COMPONENT 6 5,25E-17 6,44E-09 1,39E-06 0,169107 0,020148
PS7 PSEUDO COMPONENT 7 0 1,41E-16 1,96E-13 0,050493 0,067374
PS8 PSEUDO COMPONENT 8 0 0 0 0 0,85228
38
Appendix A2 Table X2 Component mass flow, Standard vol. fraction and mass fraction data for different streams taken from Aspen. (Addition of vanillin directly to the CDU).
Component Mass Flow LPG Gasoline Kerosene Diesel Heavy Fireing Oil
METHANE KG/HR 580,8838 12,39718 0,002908 0 4,70E-17
ETHANE KG/HR 21,54762 2,730724 0,000673 0,00E+00 2,10E-16
PROPANE KG/HR 674,8178 308,0805 0,098476 2,40E-10 6,41E-13
I-BUTANE KG/HR 351,6405 409,6616 0,195028 2,56E-09 5,09E-12
N-BUTANE KG/HR 1108,796 1850,874 1,054816 1,63E-08 3,25E-11
I-PENTAN KG/HR 499,4943 2097,439 2,360667 1,77E-07 3,00E-10
N-PENTAN KG/HR 572,8591 3305,556 4,510143 3,12E-07 5,02E-10
CYCLO-01 KG/HR 68,27974 576,0077 1,168795 1,79E-07 2,07E-10
TOLUENE KG/HR 60,13554 5374,013 83,11095 0 8,66E-08
BENZENE KG/HR 38,96331 888,1696 4,202689 3,62E-06 3,50E-09
Gasoline 5 mass% Vanillin into CDU
N-BUTANE I-PENTAN N-PENTAN TOLUENE
PSEUDO COMPONENT 1 PSEUDO COMPONENT 2 PSEUDO COMPONENT 3 PSEUDO COMPONENT 4
Figure A2. A pie chart showing the mass fraction of different components in the outlet gasoline, with lignin added directly to the CDU.
39
WATER KG/HR 107,4533 1665,619 0,036413 4326,963 0,784335
METHY-01 KG/HR 0,001067 0,001649 8,62E-07 7,37E+00 1,16E-12
H2 KG/HR 165,1526 0,419656 7,73E-05 0 0
N-HEP-01 KG/HR 0 0 0 7992,254 0
OLEIC-01 KG/HR 0 0 0 8196,989 0
THIOP-02 KG/HR 0 0 0 0,00E+00 5144,029
CARBO-01 KG/HR 0 0 0 725,7389 0
N-HEP-02 KG/HR 0 0 0 1548,991 0
1-EIC-01 KG/HR 9,91E-08 1,367237 27,57985 2,36E+04 0
1-HEP-01 KG/HR 0 0 0 2396,228 0
PS1 KG/HR 399,4301 43177,03 821,3552 0,000775 4,76E-07
PS2 KG/HR 11,27841 20880,13 3434,682 1,698874 2,16E-05
PS3 KG/HR 0,535366 15108,12 12784,35 724,7593 0,001764
PS4 KG/HR 0,00583 3,21E+03 12187,73 15851,71 0,266936
PS5 KG/HR 7,27E-07 1,16E+01 204,1315 29176,85 52,67275
PS6 KG/HR 5,71E-12 0,002955 0,24401 1,97E+04 1603,829
PS7 KG/HR 0 6,96E-11 4,67E-08 6643,025 5952,973
PS8 KG/HR 0 0 0 0 86723,96
Mass Flow KG/HR 4496,122 98878,34 29556,82 120835 99478,52
Volume Flow M3/DAY 110127 55316,16 1054,46 910101 3176,498
Component Std. Vol. Fraction (M3/DAY)
METHANE 0,151205 3,19E-04 2,70E-07
ETHANE
0,004733 5,93E-05 5,28E-08
PROPANE 0,103981 0,004694 5,42E-06 3,32E-15
I-BUTANE
0,048834 0,005625 9,67E-06 3,19E-14
N-BUTANE 0,148315 0,024478 5,04E-05 1,96E-13 5,30E-16
I-PENTAN 0,062518 0,025956 1,05E-04 1,99E-12 4,57E-15
N-PENTAN 0,070989 0,0405 0,0002 3,47E-12 7,57E-15
CYCLO-01
0,007112 0,005932 4,35E-05 1,67E-12 2,63E-15
TOLUENE
0,00539 0,047623 0,002659 9,45E-13
BENZENE
0,003442 0,007758 0,000133 2,86E-11 3,76E-14
WATER
0,008394 0,012865 1,02E-06 0,030275 7,46E-06
METHY-01 9,55E-08 1,46E-08 2,76E-11 5,90E-05
H2
0,342118 8,60E-05 5,72E-08
OLEIC-01
0,06407
THIOP-02
0,046483
CARBO-01 0,006168
N-HEP-02
0,013871
1-EIC-01
9,70E-12 1,32E-05 0,000964 0,206724
1-HEP-01
0,021266
DIPHE-01
PS1
0,041794 0,446679 0,030683 7,26E-09 6,06E-12
PS2
0,001121 0,205268 0,121926 1,51E-05 2,61E-10
PS3
5,12E-05 0,142867 0,436536 0,006208 2,05E-08
PS4
5,36E-07 0,029176 0,400104 0,130548 2,99E-06
PS5
6,55E-11 0,000103 0,006572 0,235642 0,000578
40
PS6
4,96E-16 2,54E-08 7,58E-06 0,153313 0,016987
PS7
1,42E-12 0,050526 0,061532
PS8
0,87441
Component Mass Fraction
METHANE 0,124619 1,25E-04 9,84E-08 0 4,72E-22
ETHANE
0,004623 2,76E-05 2,28E-08 0,00E+00 2,11E-21
PROPANE 0,144771 0,003116 3,33E-06 1,99E-15 6,45E-18
I-BUTANE
0,075439 0,004143 6,60E-06 2,12E-14 5,12E-17
N-BUTANE 0,237874 0,018719 3,57E-05 1,35E-13 3,27E-16
I-PENTAN 0,107158 0,021212 7,99E-05 1,47E-12 3,02E-15
N-PENTAN 0,122898 0,033431 0,000153 2,58E-12 5,05E-15
CYCLO-01
0,014648 0,005825 3,95E-05 1,48E-12 2,09E-15
TOLUENE
0,012901 0,05435 0,002812 0 8,71E-13
BENZENE
0,008359 0,008982 0,000142 3,00E-11 3,51E-14
WATER
0,023052 0,016845 1,23E-06 0,035809 7,88E-06
H2
0,035431 4,24E-06 2,61E-09 0 0
N-HEP-01
0 0 0 0,066142 0
OLEIC-01
0 0 0 0,067836 0
THIOP-02
0 0 0 0 0,05171
N-HEP-02
0 0 0 0,012819 0
1-EIC-01
2,13E-11 1,38E-05 0,000933 0,195068 0
1-HEP-01
0 0 0 0,019831 0
PS1
0,085691 0,436668 0,027789 6,41E-09 4,79E-12
PS2
0,00242 0,21117 0,116206 1,41E-05 2,17E-10
PS3
0,000115 0,152795 0,432535 0,005998 1,77E-08
PS4
1,25E-06 0,032456 0,412349 0,131185 2,68E-06
PS5
1,56E-10 0,000117 0,006906 0,24146 0,000529
PS6
1,22E-15 2,99E-08 8,26E-06 0,162795 0,016122
PS7
0 7,04E-16 1,58E-12 0,054976 0,059842
PS8
0 0 0 0 0,871786
41
Appendix A3
Gasoline 5 mass% Vanillin into DHT1
N-BUTANE I-PENTAN N-PENTAN TOLUENE
PSEUDO COMPONENT 1 PSEUDO COMPONENT 2 PSEUDO COMPONENT 3 PSEUDO COMPONENT 4
Figure A3 A pie chart showing the mass fraction of different components in the outlet gasoline, with lignin added directly to DHT1.
42
Table A3 Component mass flow, Standard vol. fraction and mass fraction data for different streams taken from Aspen. (Addition of vanillin directly to DHT1).
Component Mass Flow LPG Gasoline Kerosene Diesel Heavy Fireing Oil
METHANE KG/HR 1861,965 17,06011 0,003881 0 6,92E-17
ETHANE KG/HR 22,97774 1,300989 0,000283 0,00E+00 3,09E-16
PROPANE KG/HR 815,1312 167,8243 0,041253 2,60E-10 9,45E-13
I-BUTANE KG/HR 500,066 261,35 0,081243 2,71E-09 7,50E-12
N-BUTANE KG/HR 1689,01 1271,277 0,438354 1,72E-08 4,80E-11
I-PENTAN KG/HR 901,5168 1696,805 0,972099 1,84E-07 4,43E-10
N-PENTAN KG/HR 1080,045 2801,027 1,852832 3,22E-07 7,40E-10
CYCLO-01 KG/HR 134,5736 510,4058 0,476826 1,79E-07 3,06E-10
TOLUENE KG/HR 321,5257 12452,08 76,3579 0,000602 1,28E-07
BENZENE KG/HR 82,80707 846,8376 1,690875 3,41E-06 5,15E-09
WATER KG/HR 295,8293 5774,021 0,052304 13,54668 0,800235
H2S KG/HR 0 0 0 0 0,00E+00
METHY-01 KG/HR 0,001731 0,001216 3,89E-07 7,54E+00 1,72E-12
H2 KG/HR 566,599 0,608924 0,000112 0 0
OLEIC-01 KG/HR 0 0 0 8399,068 0
THIOP-02 KG/HR 0 0 0 0 5,57E+03
CARBO-01 KG/HR 0 0 0 607,2385 0
N-HEP-02 KG/HR 0 0 0 1416,169 0
1-EIC-01 KG/HR 9,97E-07 2,561267 26,38582 23184,37 0,00E+00
DIBEN-01 KG/HR 0 0 0 0,00E+00 0,00E+00
1-MET-01 KG/HR 0 0 0 0,00E+00 0,00E+00
VANIL-01 KG/HR 0 0 0 0,00E+00 0,00E+00
1-HEP-01 KG/HR 0 0 0 1885,886 0
PS1 KG/HR 937,1792 43137,53 323,1034 0,000597 7,02E-07
PS2 KG/HR 33,60936 22883,29 1410,012 0,876444 3,17E-05
PS3 KG/HR 2,486181 20927,72 6865,043 822,5201 0,002588
PS4 KG/HR 0,022265 2929,709 4,83E+03 2,35E+04 0,387035
PS5 KG/HR 1,55E-06 4,752443 43,19919 29325,26 7,20E+01
PS6 KG/HR 6,12E-12 0,000502 0,025776 19453,69 1821,756
PS7 KG/HR 0 2,86E-12 3,94E-09 6198,455 6397,558
PS8 KG/HR 0 0 0 0 86723,96
Mass Flow KG/HR 8678,746 115686 13575,67 114807 100589
Volume Flow M3/DAY 304678 3702,633 485,3914 4509,621 3213,033
Component Std. Vol. Fraction (M3/DAY)
METHANE 0,193405 0,000384 7,84E-07
ETHANE
0,002014 2,47E-05 4,83E-08
PROPANE 0,05012 0,002238 4,93E-06 3,82E-15
I-BUTANE
0,027712 0,003142 8,76E-06 3,59E-14
N-BUTANE 0,090154 0,014718 4,55E-05 2,19E-13 7,72E-16
I-PENTAN 0,045027 0,018382 9,44E-05 2,20E-12 6,67E-15
N-PENTAN 0,053408 0,030043 0,000178 3,81E-12 1,10E-14
CYCLO-01
0,005594 0,004602 3,86E-05 1,78E-12 3,84E-15
43
TOLUENE
0,0115 0,096601 0,005312 5,15E-09 1,38E-12
BENZENE
0,002919 0,006476 0,000116 2,87E-11 5,48E-14
WATER
0,009222 0,039042 3,17E-06 1,01E-04 7,53E-06
METHY-01 6,18E-08 9,42E-09 2,70E-11 6,43E-05
H2
0,468366 1,09E-04 1,80E-07
OLEIC-01
0,069851
THIOP-02
0,049808
CARBO-01 0,005491
N-HEP-02
0,013493
1-EIC-01
3,89E-11 2,17E-05 0,002006 0,216346
PS1
0,039131 0,390679 0,026242 5,95E-09 8,84E-12
PS2
0,001334 0,196937 0,108824 8,30E-06 3,80E-10
PS3
9,49E-05 0,173246 0,509654 0,007497 2,98E-08
PS4
8,17E-07 0,023317 0,344447 0,205859 4,28E-06
PS5
5,58E-11 3,71E-05 0,003024 0,251999 0,000782
PS6
2,13E-16 3,78E-09 1,74E-06 0,16132 0,019085
PS7
2,59E-13 0,050162 0,065408
PS8
0,864905
Component Mass Fraction
METHANE 0,201395 1,47E-04 2,86E-07 0 6,88E-22
ETHANE
0,002485 1,12E-05 2,08E-08 0,00E+00 3,08E-21
PROPANE 0,088167 0,001451 3,04E-06 2,27E-15 9,40E-18
I-BUTANE
0,054088 0,002259 5,98E-06 2,36E-14 7,46E-17
N-BUTANE 0,182688 0,010989 3,23E-05 1,50E-13 4,77E-16
I-PENTAN 0,09751 0,014667 7,16E-05 1,61E-12 4,40E-15
N-PENTAN 0,11682 0,024212 0,000136 2,81E-12 7,36E-15
CYCLO-01
0,014556 0,004412 3,51E-05 1,56E-12 3,04E-15
TOLUENE
0,034777 0,107637 0,005625 5,25E-09 1,27E-12
BENZENE
0,008957 0,00732 0,000125 2,97E-11 5,12E-14
WATER
0,031998 0,049911 3,85E-06 0,000118 7,96E-06
S
0 0 0 0 0
H2S
0 0 0 0 0
METHY-01 1,87E-07 1,05E-08 2,86E-11 6,56E-05 1,71E-17
H2
0,061285 5,26E-06 8,23E-09 0 0
OLEIC-01
0 0 0 0,073158 0
THIOP-02
0 0 0 0 0,055399
CARBO-01 0 0 0 0,005289 0
N-HEP-02
0 0 0 0,012335 0
1-EIC-01
1,08E-10 2,21E-05 0,001944 0,201941 0
DIBEN-01
0 0 0 0,00E+00 0
1-MET-01
0 0 0 0,00E+00 0
VANIL-01
0 0 0 0,00E+00 0
1-HEP-01
0 0 0 0,016427 0
DIPHE-01
0 0 0 0 0
PS1
0,101368 0,372884 0,0238 5,20E-09 6,97E-12
PS2
0,003635 0,197805 0,103863 7,63E-06 3,15E-10
44
PS3
0,000269 0,180901 0,505687 0,007164 2,57E-08
PS4
2,41E-06 0,025325 0,355484 0,204628 3,85E-06
PS5
1,68E-10 4,11E-05 0,003182 0,25543 0,000716
PS6
6,62E-16 4,34E-09 1,90E-06 0,169446 0,018111
PS7
0 2,48E-17 2,90E-13 0,05399 0,063601
PS8
0 0 0 0 0,862161
45
Figure X4 A pie chart showing the mass fraction of different components in the outlet gasoline, with lignin being pretreated before being added directly to the CDU.
Appendix A4
Gasoline 5 mass% Vanillin Pretreated
N-BUTANE I-PENTAN N-PENTAN TOLUENE
PSEUDO COMPONENT 1 PSEUDO COMPONENT 2 PSEUDO COMPONENT 3 PSEUDO COMPONENT 4
Figure A4 A pie chart showing the mass fraction of different components in the outlet gasoline, with lignin being pretreated before being added directly to the CDU.
46
Table A4 Component mass flow, Standard vol. fraction and mass fraction data for different streams taken from Aspen. (Pretreatment of vanillin before addition to the CDU).
Component Mass Flow LPG Gasoline Kerosene Diesel Heavy Fireing Oil
METHANE KG/HR 1848,014 23,85418 0,01308 0,00E+00 4,04E-17
ETHANE KG/HR 22,52746 1,750589 0,000962 0,00E+00 1,80E-16
PROPANE KG/HR 768,3627 214,4932 0,140772 0,00E+00 5,50E-13
I-BUTANE KG/HR 445,7484 315,4706 0,278144 2,34E-10 4,35E-12
N-BUTANE KG/HR 1465,45 1493,771 1,50329 1,49E-09 2,78E-11
I-PENTAN KG/HR 729,8248 1866,12 3,348744 1,61E-08 2,57E-10
N-PENTAN KG/HR 856,9593 3019,577 6,388497 2,82E-08 4,29E-10
CYCLO-01 KG/HR 103,1865 540,6174 1,652384 1,58E-08 1,77E-10
TOLUENE KG/HR 219,9972 12375,63 265,2656 1,04E-03 7,35E-08
BENZENE KG/HR 60,24698 865,2016 5,886987 3,12E-07 2,97E-09
WATER KG/HR 185,9989 158,0911 0,00995 7,745697 0,633342
S KG/HR 0 0 0 0 0,00E+00
H2S KG/HR 0 0 0,00E+00 0 1,02E-24
METHY-01 KG/HR 0,003167 0,00306 2,82E-06 0 9,97E-13
H2 KG/HR 227,3939 0,376922 0,000152 0 5,18E-34
OLEIC-01 KG/HR 0 0 0 7010,312 0
THIOP-02 KG/HR 0 0,00E+00 0,00E+00 0 4,65E+03
CARBO-01 KG/HR 2,45E-05 1,26E-06 5,68E-10 538,0809 4,22E-31
N-HEP-02 KG/HR 0 0 0 1147,905 0
1-EIC-01 KG/HR 1,65E-07 1,122335 27,82476 2,40E+04 0
1-HEP-01 KG/HR 0 0 0 1777,179 0
DIPHE-01 KG/HR 0 0 0 0,00E+00 0,00E+00
PS1 KG/HR 642,1499 42639,18 1116,491 5,51E-05 4,05E-07
PS2 KG/HR 17,85633 20042,92 4266,954 0,056183 1,82E-05
PS3 KG/HR 0,839132 13887,66 14610,01 119,2684 0,001486
PS4 KG/HR 0,013601 4221,238 19686,33 7341,097 0,224341
PS5 KG/HR 3,54E-06 30,55681 666,5071 28703,45 44,75662
PS6 KG/HR 7,84E-11 2,16E-02 2,23E+00 19839,93 1433,32
PS7 KG/HR 1,01E-19 2,36E-09 1,63E-06 7430,879 5165,134
PS8 KG/HR 0 0 0 0 86723,96
Mass Flow KG/HR 7367,178 101698 40660,83 97933,5 98017,28
Volume Flow M3/DAY 190764 29441,72 1441,242 433465 3126,463
Component Std. Vol. Fraction (M3/DAY)
METHANE 0,291491 0,000602 8,87E-07
ETHANE
0,002999 3,73E-05 5,51E-08
PROPANE 0,071743 0,003203 5,65E-06
I-BUTANE
0,037511 0,004245 1,01E-05 3,63E-15
N-BUTANE 0,118782 0,019362 5,24E-05 2,23E-14 4,60E-16
I-PENTAN 0,055353 0,022633 0,000109 2,25E-13 3,97E-15
N-PENTAN 0,06435 0,036259 0,000206 3,91E-13 6,57E-15
CYCLO-01
0,006513 0,005457 4,49E-05 1,84E-13 2,27E-15
TOLUENE
0,011948 0,107484 0,006195 1,05E-08 8,14E-13
47
BENZENE
0,003225 0,007407 0,000136 3,08E-12 3,24E-14
WATER
0,008805 1,20E-03 2,03E-07 6,77E-05 6,11E-06
H2
0,285439 7,57E-05 8,19E-08
1-EIC-01
9,79E-12 1,06E-05 0,00071 0,263116
PS1
0,040715 0,432327 0,030442 6,45E-10 5,23E-12
PS2
0,001076 0,193112 0,110555 6,25E-07 2,24E-10
PS3
4,86E-05 0,128709 0,36412 0,001276 1,76E-08
PS4
7,58E-07 0,037612 0,471701 0,07552 2,55E-06
PS5
1,94E-10 0,000267 0,015661 0,289572 0,000498
PS6
4,14E-15 1,82E-07 5,05E-05 0,193149 0,015405
PS7
5,20E-24 1,94E-14 3,62E-11 0,070599 0,054174
PS8
0,887283
Component Mass Fraction
METHANE 0,243334 0,000235 3,22E-07 0 4,12E-22
ETHANE
0,002966 1,72E-05 2,37E-08 0,00E+00 1,84E-21
PROPANE 0,101173 0,002109 3,46E-06 0,00E+00 5,61E-18
I-BUTANE
0,058693 0,003102 6,84E-06 2,39E-15 4,44E-17
N-BUTANE 0,19296 0,014688 3,70E-05 1,52E-14 2,84E-16
I-PENTAN 0,096098 0,01835 8,24E-05 1,64E-13 2,62E-15
N-PENTAN 0,112838 0,029692 0,000157 2,88E-13 4,38E-15
CYCLO-01
0,013587 0,005316 4,06E-05 1,61E-13 1,80E-15
TOLUENE
0,028968 0,12169 0,006524 1,07E-08 7,50E-13
BENZENE
0,007933 0,008508 0,000145 3,19E-12 3,03E-14
WATER
0,024491 1,55E-03 2,45E-07 7,91E-05 6,46E-06
H2
0,029942 3,71E-06 3,73E-09 0 5,28E-39
OLEIC-01
0 0 0 0,071582 0
THIOP-02
0 0 0 0 0,047433
CARBO-01 3,22E-09 1,24E-11 1,40E-14 0,005494 4,31E-36
N-HEP-02
0 0 0 0,011721 0
1-EIC-01
2,17E-11 1,10E-05 0,000684 0,245244 0
1-HEP-01
0 0 0 0,018147 0
PS1
0,084554 0,419274 0,027459 5,63E-10 4,13E-12
PS2
0,002351 0,197083 0,10494 5,74E-07 1,86E-10
PS3
0,00011 0,136558 0,359314 0,001218 1,52E-08
PS4
1,79E-06 0,041508 0,48416 0,07496 2,29E-06
PS5
4,67E-10 0,0003 0,016392 0,293091 0,000457
PS6
1,03E-14 2,12E-07 5,48E-05 0,202586 0,014623
PS7
1,33E-23 2,32E-14 4,02E-11 0,075877 0,052696
PS8
0 0 0 0 0,884782
48
CDU
FRNICE
FLASH1
STR1
STR2
STR3
MIXER
CONSPLIT
MIXERVAP
DHT1
DHT1SSEP
MULT
SCALER
SPLITSID
DHT2
DHT1SEP
DHT2SEP
FLASH11
FLASH13
FLASH12
MIXBENS
GHT
MIXDIESE
STARTMIX
MIXHV
SEP
GHTSEP
CDUSTREA
FEED
VAPOR
SIDE1
SIDE21
SIDE3
BOTTOM
CDUSTEAM
VAPPROD
FREEWATE
RECCOND
STESTR1
STESTR2
STESTR3
LLGO
LHGO
HGO
LLGOVAP
LHGOVAP
HGOVAP
SIDE22
LHGOIN
FREEWTCD STRWTR1
STR2WAT
STR3WAT
CONPRE
MIXVAP
VAPOUT
DHT1H2
NULLDHT1
DHT1PROD H2SUT
DHT1DESU
LPGIN
CRUDEFLO
SIDE212
SIDE211
DHT2LIQ
DHT2VAP
DHT2H2
DHT1COMP
DHT1IN
DHT1SEP1
DHT2COMP
DHT2IN
DHT2SEP2
FLASH11L
LPG
ISO
REF
KEROSINE
GASOLINE
TALLOIL
GHTV
GHTL
GHTH2IN
DIESEL2
LIGNIN
SCALERFE
HVFOIL
OUT
DIESEL
GHTOUT
LLGO2
GHTCOMP
Figure B1 A process flow diagram taken from Aspen, showing a detailed view of the simulation. (No vanillin added)
Appendix B. Detailed Illustration of Aspen Simulation for Different Scenarios
B1. No Vanillin Added
49
CDUQC=0
QR=0
QF=0FRNICE
Q=76
FLASH1
Q=-48STR1
QC=0
QR=0
QF=0
STR2
QC=0
QR=0
QF=0
STR3
QC=0
QR=0
QF=0
MIXER
CONSPLIT
MIXERVAP
DHT 1
Q=24
DHT 1SSEP
Q=-0
MULT
SCALER
SPLIT SID
DHT 2
Q=-4
DHT 1SEP
Q=0
DHT 2SEP
Q=0
FLASH11
Q=-14
FLASH13
Q=3
FLASH12
Q=-8
MIXBENS
GHT
Q=2
MIXDIESE
STARTMIX
MIXHV
MULT
SCALER2 FURNICE2
Q=3 STOIC
Q=-1
DIESSEP
Q=-0
GHT SEP
Q=0
CDUSTREA
FEED
VAPOR
SIDE1
SIDE21
SIDE3
BOT T OM
CDUSTEAM
VAPPROD
FREEW AT E
RECCOND
STESTR1
STESTR2
STESTR3
LLGO
LHGO
HGO
LLGOVAP
LHGOVAP
HGOVAP
SIDE22
LHGOIN
FREEW TCD STRWT R1
STR2W AT
STR3W AT
CONPRE
MIXVAP
VAPOUT
DHT 1H2
NULLDHT 1
DHT 1PROD H2SUT
DHT 1DESU
LPGIN
CRUDEFLO
SIDE212
SIDE211
DHT 2LIQ
DHT 2VAP
DHT 2H2
DHT 1COMP
DHT 1IN
DHT 1SEP1
DHT 2COMP
DHT 2IN
DHT 2SEP2
FLASH11L
LPG
ISO
REF
KEROSINE
GASOLINE
T ALLOIL
GHT V
GHT L
GHT H2IN
DIESEL
SCALERFE
HVFOIL
FALSELIG
CRUDEMIX
LIGNIN LIGNIN2 LIGNIN3
DIESELUT
H2DIES
OUT
DIESELOU
GHT COMP GHT OUT
LLGO2
Q Duty (Gcal/hr)
Figure B2 A process flow diagram taken from Aspen, showing a detailed view of the simulation for the scenario where vanillin is added directly to the CDU (See red area).
B2. Vanillin Added Directly to the Crude Distillation Unit
50
CDUQC=0
QR=0
QF=0FRNICE
Q=76
FLASH1
Q=-45STR1
QC=0
QR=0
QF=0
STR2
QC=0
QR=0
QF=0
STR3
QC=0
QR=0
QF=0
MIXER
CONSPLIT
MIXERVAP
DHT1
Q=18
DHT1SSEP
Q=-0
MULT
SCALER
SPLITSID
DHT2
Q=-4
DHT1SEP
Q=0
DHT2SEP
Q=0
FLASH11
Q=-13
FLASH13
Q=1
FLASH12
Q=-10
MIXBENS
GHT
Q=8
MIXDIESE
STARTMIX
MIXHV
MULT
LIGNINSCSEP
Q=-1
GHTSEP
Q=0
CDUSTREA
FEED
VAPOR
SIDE1
SIDE21
SIDE3
BOTTOM
CDUSTEAM
VAPPROD
FREEWATE
RECCOND
STESTR1
STESTR2
STESTR3
LLGO
LHGO
HGO
LLGOVAP
LHGOVAP
HGOVAP
SIDE22
LHGOIN
FREEWTCD STRWTR1
STR2WAT
STR3WAT
CONPRE
MIXVAP
VAPOUT
DHT1H2
NULLDHT1
DHT1PROD H2SUT
DHT1DESU
LPGIN
CRUDEFLO
SIDE212
SIDE211
DHT2LIQ
DHT2VAP
DHT2H2
DHT1COMP
DHT1IN
DHT1SEP1
DHT2COMP
DHT2IN
DHT2SEP2
FLASH11L
LPG
ISO
REF
KEROSINE
GASOLINE
TALLOIL
GHTV
GHTL
GHTH2IN
DIESEL2
LIGNIN
SCALERFE
HVFOIL
LIGNSCIN
FALSELIG
TRUCDUST
OUT
DIESEL
GHTT IO
LLGO2
GHTCOMP
Q Dut y (Gca l/hr )
Figure B3 A process flow diagram taken from Aspen, showing a detailed view of the simulation for the scenario where vanillin is added directly to DHT1 (See red area).
B3. Vanillin Added Directly to the Dehydrotreater
51
CDUQC=0
QR=0
QF=0FRNICE
Q=76
FLASH1
Q=-54STR1
QC=0
QR=0
QF=0
STR2
QC=0
QR=0
QF=0
STR3
QC=0
QR=0
QF=0
MIXER
CONSPLIT
MIXERVAP
DHT1
Q=31
DHT1SSEP
Q=-0
MULT
SCALER
SPLITSID
DHT2
Q=-3
DHT1SEP
Q=-0
DHT2SEP
Q=0
FLASH11
Q=-16
FLASH13
Q=3
FLASH12
Q=-8
MIXBENS
GHT
Q=5
MIXDIESE
STARTMIX
MIXHV
MULT
SCALER2
FURNICE2
Q=3
PRETRE
Q=-5
SPLITPRE
DECOMP
W=-2928
DSD
Q=4
SEP
Q=-0
GHTSEP
Q=0
CDUSTREA
FEED
VAPOR
SIDE1
SIDE21
SIDE3
BOTTOM
CDUSTEAM
VAPPROD
FREEWATE
RECCOND
STESTR1
STESTR2
STESTR3
LLGO
LHGO
HGO
LLGOVAP
LHGOVAP
HGOVAP
SIDE22
LHGOIN
FREEWTCD STRWTR1
STR2WAT
STR3WAT
CONPRE
MIXVAP
VAPOUT
DHT1H2
NULLDHT1
DHT1PROD H2SUT
DHT1DESU
LPGIN
CRUDEFLO
SIDE212
SIDE211
DHT2LIQ
DHT2VAP
DHT2H2
DHT1COMP
DHT1IN
DHT1SEP1
DHT2COMP
DHT2IN
DHT2SEP2
FLASH11L
LPG
ISO
REF
KEROSINE
GASOLINE
TALLOIL
GHTV
GHTL
GHTH2IN
DIESELPR
SCALERFE
HVFOIL
FALSELIG
CRUDEMIX
LIGNIN
LIGNIN2
LIGNIN3
PRE1
PRE2
RECIRK
ENTER
H2PRETRE
PRETRPR
HYRODGEN
DIESEL2
UTNER
H2
DIESEL
GHTSEPOU
LLGO2
GHTCOMP
Q Duty (Gcal/hr)
W Power(kW)
Figure B4 A process flow diagram taken from Aspen, showing a detailed view of the simulation for the scenario where the vanillin is pretreated before being added to the CDU (See red area).
B3. Vanillin Pretreated Before Being Added to the Crude Distillation Unit
52
Appendix C. Aspen Parameters and Data.
C1. Table of Aspen Data
Table C1 Data and results for the Aspen simulations.
Pretreated Lignin-CDU DHT1 No Lignin
Flow (m3/day) kg/s m3/day kg/s m3/day kg/s m3/day kg/s
Inlet CDUSTREAM 9114.7 9.4 9114.7 90.4 9114.7 90.4 9465.9 93.8
Vanillin (Stg 20) 577,2 5.1 360.3 4.8 0 0 0 0
Side streams
Side 1 1800 18.3 2000 20.9 1950 19.2 2000 19.7
Side 21 300 3.4 600 6.5 750 8 800 8.5
Side 22 580 9.5 1100 12.3 1000 11.1 1000 11.2
Side 3 20 0.2 50 0.5 50 0.5 50 0.5
Duty (MW) Duty (MW) Duty (MW) Duty (MW)
Energy Heating 92 93 76 92
Cooling -62 -55 -45 -54
53
C2. Temperature Profile without the Addition of Vanillin
Block CDU: Temperature Profile for Main Column
Stage
Temp
eratu
re C
1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29190
195
200
205
210
215
220
225
230
235
240
245
250
255
260
265
270
275
280
285
290
295
300
305
310
315
320
325
330
335
340
345
350
355
360
Temperature C
Figure C1 Temperature profile in the CDU for the simulation without the addition of lignin.
54
C3. Temperature Profile with the Addition of Vanillin Directly to the Crude Distillation Unit
Block CDU: Temperature Profile for Main Column
Stage
Tem
pera
ture
C
1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29210
215
220
225
230
235
240
245
250
255
260
265
270
275
280
285
290
295
300
305
310
315
320
325
330
335
340
345
350
355
360
Temperature C
Figure C2 Temperature profile in the CDU for the simulation with the addition of lignin directly to the CDU unit.
55
C4. Temperature Profile with the Addition of Vanillin Directly to the Dehydrotreater
Block CDU: Temperature Profile for Main Column
Stage
Tem
pera
ture
C
1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29180
200
220
240
260
280
300
320
340
360
Temperature C
Figure C3 Temperature profile in the CDU for the simulation with the addition of lignin directly to DHT1.
56
C5. Temperature Profile with Preatment of Vanillin Before Addition to the Crude Distillation Unit
Block CDU: Temperature Profile for Main Column
Stage
Tem
pera
ture
C
1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29205
210
215
220
225
230
235
240
245
250
255
260
265
270
275
280
285
290
295
300
305
310
315
320
325
330
335
340
345
350
355
360
Temperature C
Figure C4 Temperature profile in the CDU for the simulation where lignin was pretreatment before being added to the CDU unit.