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

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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)

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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)

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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)

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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).

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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)

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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)

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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)

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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

References Arend, M., Nonnen, T., Hoelderich, W., Fischer, J., & Groos, J. (2011). Catalytic deoxygenation of oleic

acid in continuous gas flow for the production of diesel-like hydrocarbons. Applied Catalysis

A: General 399, 198-204.

Blunk, S., & Jenkins, B. M. (2000). Combustion Properties of Lignin residue from Lignocellulose

fermentation. California: University of California.

Calvo-Flores, F., Dobado, J., Isac-García, J., & Martín-Martínez, F. (2015). Lignin and Lignans as

Renewable Raw Materials: Chemistry, Technology and Applications. Wiley.

Demirbas, A., Balubaid, M. A., Basahel, A. M., Ahmad, W., & Sheikh, M. H. (2015). Octane Rating of

Gasoline and Octane Booster Additives. Petroleum Science and Technology 33:11, 1190-1197.

(2013). European Chemistry for Growth. Brussels, Belgium: Cefic.

Gary, J., & Handwerk, G. (2001). Petroleum Refining: Technology and Economics. New York: Marcel

Dekker, Inc.

GIL-ROBLES, J. M., & C, E. (1998). EUROPAPARLAMENTETS OCH RÅDETS DIREKTIV 98/70/EG.

Europeiska gemenskapernas officiella tidning.

Grupta, R. B. (2009). Hydrogen: Fuel, Production, Transport and Storage. USA: CRC Press.

Hatch, L. F. (1969). Hydrocarbon Process.

Heinemann, H., & Speight, J. (2011). Fundamentals of Petroleum and Petrochemical Engineering.

Boca Raton: CRC Press.

Hulteberg, C., Tunå, P., & Abdelaziz, O. (2016). Information given to us by our tutors. Lund.

Hydrogen Fuel Cost vs Gasoline. (n.d.). Retrieved May 06, 2016, from Hydrogen Energy Systems:

http://heshydrogen.com/hydrogen-fuel-cost-vs-gasoline/

Joffres, B., Laurenti, D., Charon, N., Daudin, A., Quignard, A., & Geantet, C. (2013). Thermochemical

Conversion of Lignin for Fuels and Chemicals: A Review. Oil Gas Sci. Technology, 753-763.

Johnson, D., Chornet, E., Zmierczak, W., & Shabtai, J. (2002). Conversion of Lignin into a Hydrocarbon

Product for Blending with Gasoline. ACS Division of Fuel Chemistry, 1ed, Vol. 47, 380-381.

Konnov, A. (2016). Fundamental Combustion Course Compendium . Lund , Sweden.

Moulijn, J. A., Makkee, M., & van Diepen, A. E. (2013). Chemical Process Technology. Chichester, UK:

John Wiley & Sons Ltd.

Natural Gas Price Statistics. (n.d.). Retrieved May 08, 2016, from Eurostat: Statisitcs Explained:

http://ec.europa.eu/eurostat/statistics-

explained/index.php/Natural_gas_price_statistics#Natural_gas_prices_for_industrial_consu

mers

OK-Q8. (2009). OKQ8 Bensin - Miljöklass 1.

Pagliaro, M., & Konstandopoulos, A. G. (2012). Solar Hydrogen - Fuel of the Future. Cambrigde, UK.

Preem. (2016). Inside information given by Preem. Gothenburg, Sweden: Preem.

34

Stenström, S. (2015). Energy and Environment Course Compendium. Lund, Sweden.

Stournas, S. L. (1992). Correlations Between Diesel Fuel and Engine Emissions. Athens, Greece.

Svenskt Vatten. (n.d.). Retrieved May 06, 2016, from http://www.svensktvatten.se/om-oss/in-

english/

T.A. Albahri, M. A. (2002). Octane Number and Aniline Point Of Petroleum Fuels. Fuel Chemistry

Division Preprints, 47(2), 711.

Thringa, R., Katikaneni, S., & Bakhshib, N. (2000). The production of gasoline range hydrocarbons

from Alcell® lignin using HZSM-5 catalyst. Fuel Processing Technology Vol. 62, Issue 1, 17-30.

Ulrich, G. (1984). A Guide to Chemical Engineering Process Design and Economics. New Hampshire,

USA: Wiley.

Umney, N. (1992). Corrosion of Metals in Association with wood. Conservation Journal.

Vishtal, A., & Kraslawski, A. (2011). Challenges in Industrial Applications of Technical Lignins.

BioResources: Vol. 6 Issue 3, , 35-47.

YU, Y.-x., XU, Y., WANG, T.-j. L.-l., ZHANG, Q., ZHANG, X.-h., & ZHANG, X. (2013). In-situ

hydrogenation of lignin depolymerization model. JOURNAL OF FUEL CHEMISTRY AND

TECHNOLOGY, 443-448.

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.