Design of a plant in Brunei producing 500,000 tonnes/annum synthetic oil products from natural gas, using Fischer-Tropsch technology

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CPD NR

3299

Conceptual Process Design

Process Systems Engineering DelftChemTech - Faculty of Applied Sciences

Delft University of Technology

Subject

Design of a plant in Brunei producing 500,000

tonnes/annum synthetic oil products from natural gas,

using Fischer-Tropsch technology

Authors

(Study

nr.)

Telephone

A.A.J. Breijer 9124024 015-2126799

J.W.W. van Ganswijk

9250039 015-2850790

A.J. Greidanus 1020404 06-47466786

M.L. Oudshoorn 9635304 06-42117759

M.W. Pannebakker

9641145 015-2138102

Keywords

Gas-to-Liquid, Fischer-Tropsch,

Slurry reactor, Cobalt catalyst,

Transportation fuels

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This report is the result of a three months study to design a gas to liquid plant in Brunei, South-East Asia. The plant is based on Fischer-Tropsch technology, a technology that makes the production of high quality, clean fuels possible on a large scale. The production of the plant amounts to 500,000 tonnes/annum of clean diesel, kerosene and naphtha. The major feedstock of the plant is natural gas, the reserves of which in Brunei are currently estimated to be 315 billion m3. This means the plant could theoretically operate on these reserves for about 12,000 years. Except for natural gas the only other feedstock of the plant is pure oxygen. A company having experience producing industrial gases will deliver the oxygen. This plant will be outside the battery limits.

The economic basis of the plant is to use relatively cheap natural gas to produce good transportable liquid fuels. Natural gas from the wells in Brunei is relatively cheap: due to the remote location of the gas, transportation of it is relatively expensive. With the still increasing crude oil prices, transportation fuels from natural gas become economically more attractive. Another positive prospect of the synthetic fuels from natural gas are the sulphur free, clean fuels resulting from it. The products already satisfy fuel regulations for the far future, while maximum sulphur levels are reduced frequently. Because of the strict requirements for fuel cleanliness and quality it is plausible that clean fuel prices will increase within the coming decennium. Clean fuels can be blended with fuels not meeting requirements to obtain sellable fuel and in this way produce added value.

The design of the plant is based on economic considerations. The scale of the plant implies the use of technology already proven to work on an industrial scale. Promising laboratory technologies have been considered to be used in this design. However, when performing an economic evaluation, they turned out not to be feasible due to the risks of scaling up. All process units in the final design have been tested and proved their performance in industry. The creativity of this process lies in the combination of the different process units and the adjustment of the design to the location, market and the requirements of present time. The on-stream factor of the designed plant is 0.91 operation hours/available hours per year. This number is based on realistic figures of the shutdown time for comparable

petrochemical plants.

A first calculation has been made for the economics of the plant for a plant lifetime of 17 years. The total investment costs amount to 63 million US$, the annual production cost amount to 114 million US$. In this first calculation the net cash flow is negative and amounts to -7.8 million US$ annually. The sensitivities of the economic criteria have been reviewed and it seems that an error of –10% in the operating cost or an error of +10% in the total income can still result in a positive cash flow. When the break even has to be reached in the plant lifetime of 17 years, the product prices have to increase by 17%. As the product prices are now based on normal fuels the prices can probably be set higher for the produced clean fuels. If cheaper natural gas is available the plant also becomes

profitable. It can be concluded that although the designed plant has a negative cash flow, many opportunities are present to make the plant feasible. Especially if better utilisation of the by-products and wastes can be achieved the economics look promising.

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1. Introduction

In this chapter the background to the design will be given. The objectives will be summarized and the context it was made in will be briefly discussed. Also the situation on the field of the specific technology and the markets for the products are mentioned.

1.1 The assignment

This design was made as an assignment for the course CE3811 (Conceptual Process Design) at Delft University of Technology. The designers are all 5th-7th year students Chemical Engineering at that university. The assignment was issued by the section PSE of Delft ChemTech and is for educational purposes. There are no concrete plans to build a plant like the one that is designed.

The objective of this conceptual process design project is to design a plant producing 500,000 tons/year products out of natural gas using Fischer-Tropsch technology1. The main products are diesel (C15-C20), kerosene (C10-C14). Naphtha (C5-C9) and LPG

(C2-C4) will be accepted as by-products. The Fischer-Tropsch reaction should be

designed in such a way that the “heavy product route” is used: a product distribution where, next to the products, a lot of waxes are produced. By hydrocracking the heavy by-products the amount of diesel and kerosene can be further increased.

The production of diesel, kerosene and naphtha is included in the 500,000 tons/year production. LPG may be produced but doesn’t add up to the production. The location of the plant is a petrochemical complex in Brunei and natural gas from that region should be used as feedstock, which is delivered by pipeline.

Furthermore only literature information from 1998 and before should be used. Any information from after 1998 should be discarded. This will allow a fair comparison between this and alternative designs.

The main products diesel and kerosene are mainly used as transportation fuels. Naphtha can be used as feedstock for a cracker to produce e.g. ethylene/propylene or can be further processed to gasoline. LPG could be used for house warming, as transportation fuel or as a feedstock for chemicals production, too. Environmental issues involved are reviewed in chapter 9. All of the process streams are (complex) mixtures of a large range of hydrocarbons. Therefore a selection was made of the main components to be

considered in the design. A list of them is given in appendix 1.

The world market for transportation fuels is very large; the products from this plant will have no effect on the market situation. Main markets for the products are South East Asia and the United States. The market in Brunei is too small for all products to be sold there, but the plant is located there because of the presence of large natural gas reserves.

Only two other commercial plants for the production of synthetic fuels from synthesis gas are operated in the world. One by Shell in Bintulu, Malaysia, the other one by Mossgas in Mossel Bay, South Africa. In South Africa, Sasol also operates a Fischer – Tropsch plant that uses coal-derived synthesis gas. Because of the promise for the production of clean, high-quality fuels a lot of research in the field has been done in recent years and

numerous patents have been filed for all stages of the process. This might have serious implications if the plant would be commercialised.

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1.2 History

In 1923, Franz Fischer and Hans Tropsch discovered the Fischer-Tropsch process, in which synthesis gas (CO and H2) are converted catalytically into hydrocarbons and

alcohols2. During World War II the Germans used the technology to produce high quality fuel from coal-derived synthesis gas. Recently, there is a high interest in this technology for it produces one of the most desirable diesel fuels: with a high cetane number and negligible sulphur content. In the future the Fischer-Tropsch process will probably gain increasing importance as a clean liquid fuel supplier.

1.3 The design

Some key factors in the design will now be mentioned and references will be provided to the chapter where elaboration on that subject is given.

1.3.1 Process selection

In this project much attention was paid to the selection of the process alternatives. The plant can be divided in three parts: the syngas production, the F-T reaction section and the downstream processing (see appendix 2). For both the first two stages of the process a number of commercially established processes are available, as well as a number of processes that are promising but still in the research phase. For the downstream

processing basically every process applied in the oil refining industry can be considered, since the products from the first part of the process are somewhat similar to crude oil. A very large number of combinations are thus possible and choices for one a part influences the choices for the rest of the plant. For an overview of all alternatives considered and the selection process see chapter 2. The selected alternatives for the main parts of the plant are:

- Syngas production: Autothermal reformer

- F-T reactor: Slurry reactor with cobalt catalyst

- Separation section: Atmospheric crude distillation

- Downstream processing: Hydrocracker

For the alternatives selected for the smaller processes in the plant see chapter 2. See chapter 5 for specifications of all operating conditions.

1.3.2 Detailed designing

Due to the limited amount of time and the relatively complex process not every unit could be designed in detail, so the focus was on the main units, as described in chapter 3. Because of the limited literature available on quantitative kinetic data of especially the hydrocracker and the autothermal reformer the design of these units was very difficult. No experiments were conducted to check the results of the calculations on these units, but where possible literature was checked to see if the results were as expected. The mass and heat flows gave reasonable results. The sizing of these units was roughly estimated. For other units shortcut or detailed design methods were applied. See chapter 8 for more details.

In the process modelling with Aspen Plus a problem was to choose a suitable

thermodynamic model to use for the complex mixtures in the process. The mass and heat flows, though, seemed reliable if inspected with common sense. (see chapters 4 and 8).

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One unit within the battery limit was not designed at all, the water processing. To make an estimation for the economics, a value for a standard water processing facility was retrieved from literature and estimations for the streams were made, but no other designing was done. Also an air-separation plant has to be built, but together with the principal it was decided that the design and operation of this plant would be outsourced to a specialised company.

All specifications set in the assignment are met with the final design. The annual production is slightly higher than the target. The naphtha, kerosene and diesel meet specifications. The fourth hydrocarbon product stream is fuel gas, which is a mixture of waste streams from different units that have a significant caloric value. No LPG is produced but part of the fuel gas might be purified and sold as LPG. The economics for this look promising but this was not designed (see chapter 2). Besides these products a large amount of water is produced that is discharged to the sea. A stream of CO2 is vented

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2. Process options & selection

In this chapter a brief description of the different alternatives for the general process that were generated during literature research and creativity sessions will be given. Also the criteria that were used to evaluate these options will be discussed. The alternative that was chosen will be discussed in more detail in chapter 5.

2.1 The complete process

According to the method of Douglas3, the first choice in a design process is to choose between a batch and a continuous process. In this case the choice for a continuous process was made quickly. All options proposed in literature were operated in the continuous mode and none of the reasons that Douglas gives to consider using a batch process apply to our assignment.

In the history of the Fischer – Tropsch technology many process concepts and many layouts per concept have been proposed. The method that was chosen by the team to find the “best” concept was to create many different possibilities and then select one of them according to objective criteria, at first the economic margin. To generate possibilities, creative sessions were held (see chapter 12) and an extensive literature study was performed. Since the assignment clearly states that the process route where heavy hydrocarbons are produced in the Fischer – Tropsch reaction (that are then processed to the desired products) had to be chosen, all options that did not comply with this were ignored.

Very soon it was recognized that in the Fischer – Tropsch technology three main parts could be identified. These three parts are the production of synthesis gas, the Fischer – Tropsch reaction section and the separation and further processing of the products from the Fischer – Tropsch reactor. Different processes for al these steps were found and all possible combinations were reviewed.

2.1.1 Alternatives for synthesis gas production

When synthesis gas is produced from natural gas, two (pseudo) reactions are responsible for the conversion of methane into CO and hydrogen. These reactions are:

CH4 + H2O → CO + 3H2 (∆Hº = 206 KJ/mol) (2.1)

CH4 + ½ O2 → CO + 2H2 (∆Hº = -36 KJ/mol) (2.2)

Reaction (2.1) is usually referred to as steam reforming; reaction (2.2) is called partial oxidation. As can be seen clearly from the stoichiometry of these reactions, steam reforming gives synthesis gas with a H2/CO ratio of 3 and partial oxidation gives

synthesis gas with a ratio of 2. Based on these reactions and combinations of them, a number of industrial processes have been developed. For the industrial production of synthesis gas from natural gas there are basically six options.

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2.1.1.1 Steam reforming

In steam reforming steam and natural gas react in the presence of a Ni-based catalyst at residence times of several seconds. Syngas is formed at 1200 K and pressures of 15-30 bar. Excess steam is added to avoid carbon deposition on the catalyst, and the feed H2O/CH4 molar ratios are typically 2-5. The conversion at the outlet of the reactor is

typically 90-92%.

2.1.1.2 Partial oxidation

Oxygen and natural gas are preheated, mixed and ignited in a burner. Only the partial oxidation reaction takes place, although this is a rough approximation (see kinetics, chapter 4). The reactor temperature must be high enough to reach complete CH4

conversion, and at these temperatures combustion products like CO2 and H2O are also

formed. The reactor outlet temperature is 1300-1400 K. At the reactor outlet the gas composition is near thermodynamic equilibrium. The stoichiometric O2/CH4 consumption

is 0.5, but actual operation requires a O2/CH4 ratio of 0.7, due to complete oxidation of

part of the methane.

2.1.1.3 Autothermal reforming

Autothermal reforming is a combination of non-catalytic partial oxidation, complete combustion of methane and steam reforming in one reactor. As the feed enters the reactor, a part of the methane is partially oxidized and a part is completely combusted in the combustion zone that is operating at temperatures of ca. 2200 K. The heat produced in these exothermal reactions is used to provide the heat needed in the endothermic steam reforming reaction. This section is in the second part of the reactor. In the Ni-based catalyst bed, operating at temperatures of 1200-1400 K, the steam reforming and

equilibrating reactions take place. The O2/CH4 ratio in the feed is 0.55-0.6. Excess steam

is added to the feedstock to prevent coke formation in the reactor. By controlling the H2O/CH4 and O2/CH4 ratios the H2/CO ratio can be precisely adjusted.

2.1.1.4 Combined reforming

In industrial practice steam reformers are frequently combined with an autothermal reformer. The autothermal reformer takes away some duty of the steam reformer, reducing the required size and operation severity of the steam reformer. Primary

reformers can be operated around 970 K and 40 bar. Methane conversion in the primary reformer is ca. 75% because of the lower operating temperature. The secondary

autothermal reformer operates at the conditions mentioned above. The overall methane conversion can be complete. Disadvantage of this setup for Fischer – Tropsch synthesis is that more methane reacts to CO2 instead of CO and that two units are needed instead of

one.

2.1.1.5 Membrane reforming

A promising development in the syngas production is the use of membrane technology. In processes using steam reforming the reaction temperature has to be very high in order to achieve reasonable CH4 equilibrium conversions. Membrane technology offers the

means to achieve conversions higher than the equilibrium value by selectively removing hydrogen from the reaction mixture, thereby pushing the equilibrium to the right.

Membrane steam reformers therefore can operate at much lower temperatures, while the membrane ensures an acceptable CH4 conversion. Platinum, palladium and silica are

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examples of membrane materials. Palladium membranes are 100% selective for hydrogen, and membrane reactors using palladium membranes have been tested at laboratory scale7.

2.1.1.6 Nitrogen-rich synthesis gas

For all options where pure oxygen is used there is also the possibility to use (enriched) air. This means a large stream of inert nitrogen is added to the process. The processes mentioned above do not change significantly if this is done, but of course it strongly influences the heating and cooling requirements, compression costs and the equipment sizing. There may also be major implications on the Fischer – Tropsch reaction section, because the nitrogen cannot be removed easily from the synthesis gas. An advantage is that no air separation is necessary to produce oxygen.

From the 6 option mentioned above 4 are very similar. The partial oxidation, steam reforming, autothermal reforming and combined reforming can all produce synthesis gas more or less of the specifications that are required. Also in literature, for instance Pena4

concludes that the economic margin differences between these various syngas routes are small. The choice of a syngas production route depends on local conditions, such as electricity prices, availability of oxygen, steam, exact syngas specifications, etc. At this stage, it is thus impossible to decide between one of the syngas options at this level of detail, and for the choice of the overall process the four similar options for syngas production can be regarded as one. So there are 3 options left: the grouped one, the nitrogen-rich synthesis gas and the membrane steam reforming.

2.1.2 Alternatives for the Fischer – Tropsch reaction section

In the past, a large number of different reactor designs have been proposed for the F-T reaction. In our literature study we focused on the ones that were designed or

commercially used in the last 2 decades. For a Fischer – Tropsch process that favours the production of heavy hydrocarbons a large value of the ASF-parameter α (see chapter 4) is required. High values of this parameter are favoured by specific catalyst design and reactor conditions, mainly temperature.

In literature, five different processes were found that favour the production of heavy hydrocarbons. A short summary of these options is given:

2.1.2.1 Slurry reactor

Different companies developed this concept, but the principle is similar for all of them. The slurry-reactor is a vessel where the synthesis gas enters at the bottom. The liquid is a mix of the produced waxes that are in the liquid phase at the operating temperature of about 200-250 °C. In this liquid phase the solid catalyst is present as small solid particles. The synthesis gas is bubbled through the liquid and reacts to form products. Cooling in the reactor is done by means of steam coils. Unreacted syngas and gaseous products leave the reactor at the top, a part of the liquid mixture is taken out continuously at the side. From this liquid stream the solid catalyst is separated and recycled to the reactor, the product stream (‘syncrude’) is further processed. Main advantages of the slurry reactor are the good heat transfer and the practically ideal mixing of the liquid phase. This means the large amounts of heat produced in the F-T reaction can be removed easily and no hotspots can develop in the reactor. This also means that the temperature in the reactor

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can be controlled more precise. The pressure drop over the reactor is quite low, what reduces compression costs for the synthesis gas recycle. Disadvantages are the need for catalyst separation and the possibility of diffusion limitation because the syngas has to diffuse from the gaseous bubbles to the catalyst particles through the liquid phase. Within the general concept of a slurry reactor a number of companies developed their own version, of which two will be discussed in more detail below.

Sasol8,9

Sasol was the first company to operate a slurry reactor on a commercial scale. After starting research in the early 1980’s and performing tests on a pilot-reactor with a

diameter of 1 m, in 1993 a commercial slurry reactor of 5 m diameter and with a capacity of 2,500 bbl was commissioned. The feedstock for this reactor is coal-based syngas prepared with Sasol’s Lurgi-gasifiers. The reactor uses an iron-based catalyst but research is done on cobalt catalysts, mainly for applications where the feedstock will be natural gas instead of coal. Separation of the catalyst particles from the reactor wax proved to be difficult, but Sasol developed and patented a very successful process to achieve this.

Exxon10,11

Exxon started research on F-T processes in the early 1980’s and invested over US$ 300 mln. They have a pilot-plant in Baton Rouge in the US, but do not operate commercial reactors. The Exxon process is developed especially for Gas-to-Liquid applications, so with natural gas as a feedstock. The process is based on a cobalt catalyst and is operated at a lower temperature than the Sasol process to favour the production of heavy

hydrocarbons.

A number of other companies developed slurry reactor concepts. Most of them use cobalt catalysts and are designed for Gas-to-Liquid applications. In general, cobalt is considered the favoured catalyst for these applications in a slurry reactor.

2.1.2.2 Fixed-bed reactor

Fixed-bed reactors for Fischer-Tropsch synthesis have been used for a long time. The first commercial reactors commissioned by Sasol in the 1950’s used this technology. It is a low-temperature process and operates at temperatures between 200-240 °C. Because of the large heat production in the F-T reaction the fixed bed must have a small diameter to enable heat removal from the catalyst, so a multitubular reactor is used. Cooling is achieved by evaporating cooling water on the outer shell of the tubes. This cooling mechanism is less efficient than the one used in the slurry reactor, so temperature gradients are present in the tubes. This means the average temperature in the reactor has to be lowered which has a negative effect on the productivity.

In the fixed-bed reactor, syngas is fed from the top and led through the tubes filled with catalyst. At the bottom all products, gaseous and liquid, are collected. Catalyst is retained in the reactor and has to be replaced after a certain time, which means the reactor has to be off-line. To prevent large pressure drops the catalyst particles have to be larger than in the slurry reactor.

The main advantages of the fixed-bed reactor are the easy scale-up, the lack of need for product/catalyst separation. Diffusion limitation is not likely to occur because the syngas only has to diffuse from the gaseous phase to the solid particles and not through a dense liquid phase as in the slurry reactor. Disadvantages are the large mechanical structure of

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the reactor and resulting capital costs, the large catalyst usage and the large pressure drop over the reactor. Since conversions-per-pass are generally lower in a fixed-bed reactor than in a slurry reactor, the compression costs for the recycle are significantly higher. The only companies who have developed new processes that use a fixed-bed reactor are Shell and BP. The latter has plans to build a pilot-plant in Alaska, the former has been

operating a commercial plant in Malaysia since 1993. Research institutes around the world are doing research to develop other concepts using a fixed-bed reactor. A few of these technologies and the specific processes used in commercial operation are discussed below.

Shell2,8,9

The Shell Middle Distillates Synthesis (SMDS) plant in Malaysia uses natural gas as a feedstock. From the syngas a range of mainly heavy hydrocarbons is produced in a multitubular fixed-bed reactor with a cobalt catalyst. The products are hydrotreated to produce high-quality chemicals and transportation fuels.

Sasol2,9

Sasol started operating Arge fixed-bed reactors in South Africa in 1955. They used an iron-based catalyst and syngas produced from coal. After operating the Arge reactors for over 30 years, Sasol decided to choose the slurry reactor concept for their

low-temperature F-T processes. (The option of a multitubular fixed-bed reactor with an iron catalyst is not considered in this process selection.)

2.1.2.3 Supercritical

A process is being developed for a F-T process where a supercritical solvent is introduced12,13,14,15. Hydrocarbons such as hexane are used as a solvent and at reactor temperature and pressure of 200-250 °C and 30-40 bar they are in the supercritical phase. The solvent and the syngas are fed to the multitubular reactor the same way as in the non-solvent setup. The products that are formed dissolve in the supercritical phase or leave the reactor in the vapour phase. Advantages of the supercritical process are that the

selectivity shifts towards heavier hydrocarbon products and the better heat removal from the catalyst by the solvent. Disadvantages are the need for a larger reactor, the higher compression costs because of the high pressure and the larger stream, and the fact that the concept has only been tested on a laboratory scale.

2.1.2.4 Nitrogen-rich synthesis gas

As described in the paragraph about synthesis gas production, one option is to produce it with air, resulting in syngas with large amounts of (inert) nitrogen. This syngas can be used for F-T reaction in a fixed-bed reactor16. Of course, introducing these amounts of nitrogen means that a much larger reactor is needed. Because of the amount of nitrogen that is also recycled, compression costs would be significant when a recycle is used. For this reason no recycle is applied, meaning a large amount of unreacted syngas is lost. The advantages are that no air-separation plant is needed and that the nitrogen removes a part of the heat that is produced in the reaction. This concept is not applied in commercial processes, only in research projects on a small scale. The key question is whether the savings on air separation are larger than the costs for the lost syngas, larger reactor and increased required compression capacity.

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2.1.2.5 Fluidised-bed reactor

The fluidised-bed reactor (or fixed fluidised-bed reactor, because in older versions of this design the catalyst is circulating) is generally used for high-temperature F-T processes, in the range of 300-350 °C. The reason is that if a fluidised bed is used in conditions that favour the formation of heavy waxes, these waxes stick to the surface of the catalyst particles and cause clotting. Therefore the temperature at which a normal fluidised-bed Fischer - Tropsch process can be operated has a lower limit.

Converting the heavy waxes in-situ can solve the problem of clotting between catalyst particles and thus a lower operating temperature (260-280 °C) can be applied2,8,17_{. In this}

setup an amount of hydrocracking-catalyst is added (e.g. ZSM-5 zeolite) to directly convert the heavy waxes formed to products that do not cause catalyst clotting. If the right amount of hydrocracking catalyst is added and the process conditions are right, these waxes are converted directly to the desired middle distillates. Advantages of this setup are that no wax/catalyst separation is needed and the product treatment does not have to include a large hydrotreatment step. A major disadvantage for the production of heavy hydrocarbon products is that, because of the relatively high temperature, always a large fraction of (unwanted) light products is produced. Although Sasol in South Africa is operating the fluidised-bed concept on a commercial scale, the concept where it is

combined with hydrocracking is only tested at a very small scale. After scale-up the advantages might decrease or new problems might surface.

2.1.3 Alternatives for the downstream processing

After the F-T reaction there are two main product streams. These are a complex mixture of liquid products (‘syncrude’) and a mixture of unreacted syngas and light products. Of course the unreacted syngas is recycled, but the rest has to be processed to products or recycled otherwise. For the downstream processing of these products there is a large number of options2,18,19,20,21,22_{. The unreacted syngas can be recycled to the reactor feed}

after the gaseous hydrocarbon products have been separated.

In general, a significant fraction of the F-T products can be purified and sold directly as diesel, kerosene or naphtha. In the F-T reaction also a large amount of water is produced, that has to be separated from the other products. After separation it can be purified to use as a recycle stream or discharged to the environment. The remaining heavy wax can be processed in different cracking processes and then be refined to increase the production of the main products.

2.1.3.1 Gaseous product processing

The gaseous product stream could be separated into methane, ethane/ethylene, LPG and light naphtha fractions. The light hydrocarbons (C1-2) could then be reprocessed to

syngas, the other products could be sold or processed to heavier hydrocarbons and thus increase the amount of desired products. For this separation a cryogenic unit would be necessary.

Another option is to recycle the light gases such as methane, ethane and propane with the synthesis gas. Then a significant purge would be necessary to prevent build-up, but compression costs would be less and no cryogenic separation would be needed.

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2.1.3.2 Liquid product processing

The liquid product stream from the reactor combined with the condensed part of the gaseous products can be treated in a few different ways. If hetero-atoms such as sulphur or nitrogen are bound in the hydrocarbons, a hydrotreater can be used to remove them. Since no hetero-atoms are present, a hydrotreating unit is not further taken into account. The liquid product stream could first be sent to a hydrocracker and then to a fractionation column, or the other way around. The main differences are the sizes of the hydrocracker and the fractionator and in the first option there is a risk of unwanted hydrocracking of the diesel and kerosene fractions.

After the separation, the naphtha could be sent through a reformer to improve the quality and convert it into gasoline. The C3-C8 alkanes and alkenes (if present) in the LPG and

naphtha streams can be processed in an oligomerisation unit to increase the yield of heavier hydrocarbons.

2.1.4 Conclusions on the process alternatives

As said before, five alternatives for the F-T reactor have been reviewed in the selection process. The slurry reactor with iron or cobalt catalyst has been treated as one option, so the five options were:

1. Slurry reactor

2. “Normal” multitubular fixed-bed reactor

3. Multitubular fixed-bed reactor with nitrogen-rich syngas 4. Multitubular fixed-bed reactor with supercritical solvent 5. Fluidised-bed reactor with added hydrocracking catalyst

Combined with the three options for the synthesis gas production, this gives 15

theoretical process options so far. This is immediately reduced to nine though, because if the option to produce syngas with air is chosen, there is only one option for the reactor design and vice-versa, the multitubular fixed-bed reactor. To speed up the selection process this number was further reduced to six, because the syngas production with membranes was only reviewed in combination with a multitubular fixed-bed reactor, assuming that if that option doesn’t seem feasible for that type of reactor, the same holds for the other reactor types. If that combination would come out best, the combination between syngas production with membranes and other reactors could be reviewed later. So the remaining options were:

1. “Classical” syngas production – Slurry reactor (Slurry)

2. “Classical” syngas production – Multitubular fixed-bed reactor (MTFB) 3. “Classical” syngas production – Multitubular fixed-bed reactor in supercritical

solvent (SCMTFB)

4. “Classical” syngas production – Fluidised-bed reactor with added hydrocracking catalyst (CFFB)

5. Syngas production with air – Multitubular fixed-bed reactor with nitrogen-rich

syngas (N2MTFB)

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Combined with the options for the downstream processing this gives many possibilities for the entire process. But for all 6 options mentioned above the feed to the downstream processing is similar, only in the alternative where hydrocracking catalyst is added to the F-T reactor no separate hydrocracker is needed. Furthermore for every product mixture from the F-T reaction section the downstream processing can be adapted in such a way that the end products are about the same. For these reasons the choice of an exact layout for the downstream processing is not yet made and one of the 6 options for the synthesis gas production and the F-T reactor has to be selected first.

2.2 Selection method

As a method of selecting a process the method of Douglas3 was followed and the

economic margin was taken as the main criterion. If this was not sufficient, other criteria could be reviewed if necessary, such as utility costs and investment costs. The economic margin gives an indication for the value that the plant produces. This margin is defined as:

Margin = Total value OUT (products, wastes, fuel) – Total value IN (feed, fuel) Other costs, such as utilities, investment, operation and maintenance costs are not

considered here, these will be calculated at later design levels. The streams that enter and leave the process are given in appendix 2, the simplified block scheme. To make a first selection the margin is used as a selection criterion of process alternatives. It is a tool for rejecting alternatives in an early stage of the design. In this way a preliminary figure for the economic performance is calculated with an accuracy of about ± 40%3. So only if the difference between two options is more than 40% an option can be discarded.

This method has some consequences for the 6 options that are reviewed. The main advantage of the membrane reforming are the lower utility costs. But because all in- and outgoing streams are approximately the same as for the steam reforming, the economic margin is not separately calculated for the membrane reforming. Since the different alternatives for the downstream processes also differ mainly in utility and installation costs, the most favourable product distribution from the F-T reaction mixture is chosen in the calculation of the economic margin.

The calculations for the economic margins for the 5 options are given in appendix 3, the results are presented in table 2.1.

Table 2.1: Economic margin for the different process alternatives

Process option Margin [mln US$] Difference with slurry [%]

MTFB 51.8 -31.1

CFFB 66.1 -12.1

Slurry 75.2 0.0

N2MTFB 0.285 -99.6

SCMTFB 46.4 -38.3

From these results it is clear that the supercritical option and the option with the nitrogen-rich synthesis gas can be discarded. The difference between the slurry reactor, the

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fluidized-bed reactor and the multitubular reactor is not large enough to make a choice based just on the economic margin.

2.2.1 Other selection criteria

With the economic margin no other options can be discarded, so for the 3 remaining options other aspects have to be taken into account. Although it actually is not in this design level, some estimations about the investment costs, the plant downtime, compression costs and catalyst separation can be made.

Option 2, the fluidised bed with added catalyst, is a very novel technology that has not been applied on a large scale before. This means scale-up problems are likely to occur, that extra capacity has to be installed to cope with production losses in the start-up phase and that research would have to be done which takes time and resources. These risks are thought to be high. The amount of desired products produced is also smaller than in the other options, so the advantages are not very large. Another disadvantage is that this option does not produce heavy hydrocarbons as was stated as the desired option in the assignment. For these reasons this option will not be taken further into account. For the two remaining options the criteria mentioned earlier are evaluated. The results are given in table 2.2.

Table 2.2: Selection of a process option with the extra criteria taken into account. Option Margin [mln US$] Capital investments18 Operating time Compression costs Catalyst separation Slurry 75.2 + + + - MTFB 51.8 - - - +

Since there are clearly more advantages to the slurry reactor and this option also has a higher estimated economic margin, this option is chosen.

2.2.2 Catalyst selection

As seen in the short description of the slurry-reactor earlier in this chapter, both cobalt and iron based catalysts can be used. For our design a choice will have to be made. Both catalysts have advantages and disadvantages.

The main difference between the iron and cobalt catalyst is that iron catalyzes the water-gas shift reaction. This means that if synwater-gas with a low H2/CO ratio is used, the iron

catalyst has an advantage because no separate reactor has to be built to raise this ratio. The cobalt catalyst is more expensive than the iron catalyst, but has a much longer possible lifetime. An advantage of cobalt is the much higher activity in the F-T reaction, which allows lower operating temperatures and thus heavier hydrocarbon products. Since these heavy products are the most desired ones and the syngas produced from natural gas has a H2/CO ratio of about 2, the cobalt catalyst is chosen for this design.

2.2.3 Process selection for the synthesis gas production

Although the slurry reactor has been chosen as the preferred option for the Fischer-Tropsch reaction, there are still several possibilities for the synthesis gas production left that were grouped earlier. First we will review the membrane reforming.

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2.2.3.1 Hydrogen membrane reforming (HMR)

To make a quick and easy comparison between the HMR process and the established synthesis gas production routes, one of the “classical” production routes was taken. The combined reforming process was chosen. Some advantages and disadvantages for the HMR were found, a summary is given in table 2.3.

Table 2.3: Pros and cons for the H2 membrane reforming process

Advantages Disadvantages

1. Less heat needed for reaction 1. High risk technology

2. No pure oxygen needed 2. Cost of palladium membrane

3. One reactor vessel 3. Lack of technical knowledge

4. In situ hydrogen separation 4. Scale-up

Advantages of HMR

The HMR process uses a palladium membrane that selectively transports H2. In the steam

reforming reaction, an equilibrium reaction, a constant removal of H2 will result in a

conversion higher then the equilibrium conversion. As a result, the reformer reactor can be operated at much lower temperatures. Balachandran6 describes the temperature dependence of the methane conversion using a palladium membrane. Already at a temperature of 920 K, a methane conversion of 100% can be achieved. This results in a lower energy need for the production of synthesis gas. Due to the high-energy

requirement of conventional reforming processes, a large difference in economic potential is expected if the heat effects are included in the economic margin.

The combined reforming process uses an autothermal reformer after the steam reformer to oxidize unconverted methane. (Part of the methane is oxidized partially for the

production of synthesis gas; part is oxidized completely for the generation of heat needed in the steam reforming process) The use of the H2 membrane thus eliminates the need of

an autothermal reformer, which results in two advantages: the reaction can be carried out in one reactor vessel, and no expensive oxygen separation plant is needed to produce the pure oxygen used for oxidizing methane. It is expected that the oxygen plant is more expensive then the autothermal reactor.

In the downstream processing section of the plant, H2 might be needed for hydrotreating

and -cracking. The HMR process produces pure H2 if the sweep gas is cleaned. The HMR

process integrates this hydrogen production into the reforming process.

Disadvantages of HMR

The implementation of new technology only tested at pilot plant scale brings along a bigger risk then the use of conventional technologies. This bigger risk has to be accounted for in the design, bigger safety margins have to be included. These safety margins come with a price.

Palladium is a rare noble metal, which is very costly. The relatively low diffusion rate of H2 through the membrane (compared to the industrial space velocities) means a high active surface area is needed. This makes the implementation of a palladium membrane at industrial scale a costly affair.

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The lack of technical knowledge about the membrane reforming process means some research probably has to be conducted, to verify some of the experimental data from literature.

Scale up from pilot plant scale to industrial scale usually is a difficult process. The easiest way to achieve safe scale up is the multi tubular reactor. In this configuration, only one tube has to be tested at laboratory scale. This tube is then scaled up in parallel.

Conclusions on HMR

From the advantages, the savings in heating costs and oxygen were estimated. If more design detail is needed to choose one of the process options, the costs of the extra reactor can be calculated later. In this stage however, it is omitted. From the input/output

structure of the HMR process and the combined reforming process the extra cost of pure oxygen can be calculated. For estimation of the extra heat requirement of the CR

(Combined Reforming)of the HMR process, a simplified flow sheet appendix 2 has been used. For the heat requirement of the combined reforming process, stream tables of the Lurgi low-pressure methanol process are used23. From the disadvantages, only the cost of the palladium membrane can be estimated, all other disadvantages are qualitative. The results are given in table 2.4. The revenues are the same for both options, so the difference in costs also represents the difference in economic margin.

Table 2.4: Estimation of the costs of CR and HMR

Costs

CRM

Material cost 54.75 M$/yr

Extra energy cost 9.97 M$/yr

Total 64.72 M$/yr

HMR

Material cost 43.20 M$/yr

Palladium cost 19 M$/yr

Total 62.20 M$/yr

As shown above, estimating the costs for the two options did not result in a significant difference in economic potential. The economic potential of the HMR process is an optimistic value, since the HMR reactor needs extra CH4 to deliver the energy needed for

the steam reforming reaction, which is neglected here.

The calculated results combined with the qualitative disadvantages, high risk technology, lack of technological knowledge and potential scale up problems lead to the conclusion that the CRM is the best alternative, and thus hydrogen membrane reforming is discarded.

2.2.3.4 Selection of a “classical” synthesis gas production

Now the alternative of synthesis gas production using hydrogen membranes is discarded, a choice has to be made between the four other available options. All of these options have been described shortly earlier in this chapter. This choice cannot be based on the economic margin, because it was recognized earlier that the economic margin for these options is approximately equal. The best option depends on local conditions and requirements of the rest of the process. The first criterion used is the syngas ratio. The

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Fischer – Tropsch reactor needs a syngas ratio of approximately 2.1, to prevent hydrogen depletion. This means for two of the process options, partial oxidation and steam

reforming, that if they were chosen, extra process equipment had to be installed to adjust the syngas ratio. Since there are no large advantages with these options, this is a reason to discard them. This leaves only the combined reforming and the autothermal reforming as available alternatives. The combined reforming is actually a steam reformer with an autothermal reformer behind it. The advantage of this process would be that the steam reformer is smaller than a single steam reformer unit and is operated at a lower

temperature. The main disadvantage of this setup is that the steam reformer produces a higher synthesis gas ratio and more CO2. This CO2 is not converted to syngas in the

autothermal reactor. The syngas ratio could be adjusted subsequently to a lower value in the autothermal reactor, but not to the desired value. Because of these disadvantages the autothermal reformer is chosen as process for he production of synthesis gas. The synthesis gas ratio can be adjusted easily to the desired value and no excessive CO2

formation is occurring.

2.2.4 Synthesis gas purification

Although an autothermal reformer produces significantly less CO2 than the combined

reforming, still some is produced. Because this CO2 does not react in the F-T reaction

(see chapter 4), it will only act as a diluent and cause equipment to be larger. For this reason it has to be removed from the synthesis gas. There is also still a very large amount of water present in the syngas stream leaving the autothermal reactor because of the excess steam in the feed and the water produced in the complete combustion of methane. This water can easily be removed by condensing and this is also desired for the same reasons as the CO2 removal. The water can easily be removed by condensation. For CO2

removal the most common process is absorption with an amine24. This process is almost 100% selective to remove CO2 and leaves only about 50 ppm in the product stream. The

question is now which component to remove first. This is easily answered, since the temperature at which the CO2 removal works, is approximately 40°C, and at that

temperature and the given pressure the water is completely condensed. So when the syngas is cooled to 40°C the water can be removed in a simple V/L separator and after that, the syngas is led through the absorber to remove the CO2. This leaves syngas with

mainly hydrogen and CO, some unreacted methane and traces of water, CO2 and nitrogen

(see process stream summary, appendix 5). The amines with the CO2 absorbed in it are

regenerated with steam. The CO2 could then be processed to a product and sold, or vented

to the air. Because there is no market for CO2 in Brunei it will be vented.

2.2.5 Catalyst/reactor wax separation

The discussion given below is mainly based on one report25.

The waxy slurry leaving the F-T reactor contains about 30 wt% catalyst particles that have to be recycled to the reactor. The particles are assumed to have approximately a normal distribution (see appendix 6).

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In general, there are a couple of mechanisms for the separation of solid particles from the wax:

1. Vacuum distillation of the wax 2. Thermal cracking

3. Gravitational sedimentation

4. Chemical methods (conversion of catalyst) 5. Other chemical/physical methods (e.g. clogging) 6. Centrifuging / hydrocyclones

7. Use of a solvent (reduce liquid viscosity) 8. High-gradient magnetic separation (HGMS) 9. Filtration (including membranes)

A couple of technical conditions have to be fulfilled:

• The amount of particles that remains in the wax feed to the hydrocracking unit must not exceed 2 to 5 ppm particulates (as indicated by manufacturers)

• The cleaned wax phase that is recycled to the FT reactor must not be too heavy in composition (buildup of higher molecular weight waxes in the reactor).

• The flow returning to the reactor should not exceed 60 wt% in solids due to difficulties in pumping and handling

Additionally, there are some other preferences:

• The separation process should differ as little as possible in temperature and pressure conditions from the reactor conditions.

• Almost all the catalyst particles should be recycled to the reactor due to the high price.

Another general consideration is whether to place (part of) the separation in the reactor versus placing it outside the reactor.

Separation of the particles in the reactor has the advantage that the particles are not transported out of the reactor, are then separated and finally recycled back again. On the other hand, a separation unit in the reactor will increase the complexity of the reactor design and make maintenance and troubleshooting more difficult. Finally, iron catalyst particles are much more sensitive to attrition than cobalt on alumina particles, which will undergo almost no mechanical wear. So the disadvantage of transporting the particles only holds for iron which makes in-reactor separation less attractive.

Now, options 1 to 9 will be discussed subsequently.

Options 1 and 2 are not feasible. Vacuum distillation is not possible because the heaviest wax has too high a boiling point. Thermal cracking will convert a too large fraction of the wax to coke and gases.

Gravitational sedimentation (3) is no option due to the very long time required (see app. 7). For example, a particle of 5 microns will require three quarters of an hour to fall one centimetre.

The chemical conversion of the catalyst (4) is too costly and too complex. The clogging mechanism (5; addition of a chemical to establish attracting forces between the particles) creates larger particles, which allows quicker sedimentation. However, this will require

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an extra process step in order to return the particles to their original state before returning them to the reactor, and to recover the chemical component.

Hydrocyclones (6) are known to be suitable for removing at least the bulk of the particles. These are low-capital-cost items with a relatively low energy requirement.

Centrifuges (6) are able to remove also smaller particle sizes. However, it is shown (app. 7) that even for a centrifuge with 3000 G, a particle of 0.5 micron will need 2 minutes to move 1 centimetre.

Another technique is based on the reduction of the wax viscosity by dissolving it in a supercritical solvent. The viscosity of the liquid will thus be reduced considerably, enabling easier separation of the solids. Butane or hexane (both readily available at our plant) could be used as supercritical solvent. However, this enhanced sedimentation will neither be sufficient to remove the smallest particles (lowest possible: 100 ppm in cleaned wax). So also here a second unit would be required. The major disadvantage is the

accumulation of high molecular weight wax in the reactor. This occurs when the heaviest fraction of the wax, containing the catalyst, is recycled.

High-gradient magnetic separation has been proved to be a feasible technique. This technique separates components on the basis of differences in magnetic susceptibility. The 2 to 5 ppm requirement could be met, which makes this technique attractive. Filtration can be divided into three techniques:

• Deep-bed filtration

o Sub-micron solids can be removed.

o Spent filters (containing fines) have to be replaced. o Small particles are permanently removed.

• Filtration with screens

o Practical bulk removal of particles >1 micron.

o Filter aids will increase removal of small catalysts particles.

o Removed catalysts particles (in the filter aid) will need an extra recovering step.

• Membranes processes

o Microfiltration (>0.1 micron) and ultrafiltration (<0.1 micron). o Advantages: simplicity, low costs and high product purity. o Disadvantages: lack of theoretical understanding.

According to the discussion above, feasible options for our process are: 1. Centrifuging / hydrocyclones

2. Use of a solvent (reduce liquid viscosity) 3. High-gradient magnetic separation (HGMS) 4. Filtration (including membranes)

Some of them, however, are only suitable for bulk separation and have to be combined with a residual separation. In table 2.5 an overview is given.

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Table 2.5: Overview of possibilities for the catalyst/reactor wax separation

Suitable for

Process Capital cost Bulk Residual Both

Centrifuge Medium X

Hydrocyclone Low X

Solvent High X

HGMS High X X

Deep-bed Low X

Screen filters Low X

Membranes Low X From the economical point of view the techniques of the high-gradient magnetic separation and the supercritical solvent were not an option. Furthermore deep-bed and screen filters have some disadvantages (as mentioned), which are not present with the membrane filters. So the membrane filter can be an option for residual separation of the catalysts from the wax. For the bulk separation we can choose between the centrifuge and the hydrocyclone. The centrifuge can not remove particles whose size are less than 2 to 0.5 micron and the hydrocyclone could possibly remove particles only down to 5 to 8 micron, but the centrifuge is more expensive than the hydrocyclone. On the other hand, in the centrifuge-membrane configuration, the size of the membrane is smaller due to fewer particles that have to be removed in the residual step in contrary to the hydrocyclone-membrane configuration.

The following assumption is made: the price for scaling up the membrane is within the price difference between centrifuge and hydrocyclone. Therefore the configuration for this design will be a hydrocyclone as bulk separation followed by a membrane as a residual separation.

2.2.6 Selection of processes for downstream processing

By now, choices have been made for both the synthesis gas production and the Fischer – Tropsch reactor. Therefore, the process alternatives for the downstream processes can be reviewed. As described in chapters 4 and 5, the product distribution and operating conditions of the reactor were fixed. The options for downstream processes can now be reviewed step-by-step.

2.2.6.1 Gaseous reactor product

The first stream that was reviewed is the gaseous stream that leaves the reactor. At operating conditions this stream contains unreacted syngas, light hydrocarbon gases, a large amount of water, but also evaporated heavier hydrocarbon products. The desired situation is to recycle the synthesis gas, discharge the water to the treatment, and process the hydrocarbons. The easiest way to separate the water and the heavy hydrocarbons is to cool the gaseous product stream at reactor pressure and then separate the vapour and the liquid stream. This gives a vapour stream with mainly C1-C5 hydrocarbons, syngas and

some nitrogen. The liquid is a mixture of water and the desired products naphtha, kerosene and diesel. The vapour stream can be further separated or recycled untreated to the reactor feed. In case of a direct recycle the purge stream has to be substantial to prevent build up of light gases. In case of further separation this has to be done by cryogenic separation. The largest component that causes build-up in the recycle is

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methane. Because the normal boiling point of methane is -161.5 °C26_{the gas recycle has}

to be cryogenically separated below this temperature, which is very expensive. The only advantages are the lower compression costs for the recycle and the somewhat lower raw material usage. The net effect was not calculated but estimated not to be economical. Because the main components in the reactor gas will still be carbon monoxide and hydrogen, the advantages for the compression costs are not that large. Furthermore the raw material is not very expensive and the savings would be very small. For these reasons it was chosen not to separate the gaseous stream further and recycle it at a temperature of approximately 40 °C. To determine the size of the purge stream an economic evaluation was done. If the purge is smaller less syngas is lost, but the compression costs increase. If the purge increases the compression costs decrease but more syngas is lost and more raw materials have to be used. The result of this evaluation can be found in appendix 7, an optimum was found at approximately 40% of the gas purged. At this moment it was recognised that because of the large purge stream it was probably possible to retrieve all the hydrogen needed for the downstream processes from the purge, meaning no excess hydrogen has to be produced in the syngas production.

2.2.6.2 Liquid reactor product

The condensed liquid can be mixed with the liquid wax product from the reactor and the large amount of water can easily be removed in a settler. This gives a liquid C5+

hydrocarbon product. This liquid has to be kept at a sufficiently high temperature to prevent the formation of solid wax and thus at sufficiently high pressure to prevent a vapour phase in the settler. As mentioned earlier this stream can be separated first and the heavy wax can be hydrocracked afterwards, or the total stream can be hydrocracked and then separated. If the latter option is chosen the flow through the hydrocracker is larger and the risk of cracking of diesel and kerosene is present. The first option is also mentioned in literature as the best option to produce middle distillates27_{. This is why it}

was chosen to separate the liquid stream first and send the wax fraction to the

hydrocracker. The liquid hydrocracker product is fully recycled to the separation column. No purge stream is necessary because all heavy components will be hydrocracked and all light components will be removed in the separation. Because of the high pressure at which the hydrocracker operates (~100 bar) some gaseous product is formed when the liquid is flashed to low pressure. This vapour consists mainly of C3-C7 hydrocarbons.

Because hydrocracking requires an excess of hydrogen28,29, the hydrogen is separated from the hydrocracker product stream and recycled. Here a small purge is necessary because of the build up of light hydrocarbons such as propane and butane.

2.2.7.2 Separation section

In the separation section the liquid hydrocarbon product is separated in heavy wax (C20+)

and the products diesel, kerosene, naphtha and LPG. The most common option for these types of mixtures is an atmospheric crude distillation. In that way it is possible to separate the liquid hydrocarbon stream to our desired products at desired purity in one single unit2

.

An atmospheric crude distillation consists of a large column with a reboiler and condenser and a number of side strippers to take out the products. The column can be operated in such a way that the bottom product is C20+ wax and the side products are

diesel, kerosene and “heavy” naphtha. The top-product is a mixture of “light” naphtha and light hydrocarbon gasses, LPG, water from the stripping steam and some ethane and

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methane that was left in the wax. The diesel and kerosene products are on specification in a single run, the naphtha might be sent trough a naphtha reformer to improve the quality and make it possible to sell it as gasoline. It was chosen not to do this because it is likely that there is a market for the naphtha product and gasoline was not a desired product. The top-stream should be further purified to increase the yield of naphtha and meet the product specifications of LPG. The vapour stream that is formed in the flash of the hydrocracker recycle stream is of about the same composition as the top-stream of the atmospheric crude distillation, so they can be mixed and purified together. This is done in a second distillation column, where the stream is split in a bottom product containing the C5+ hydrocarbons (naphtha) and the water and a top product with all the light gases. After

the water is removed this bottom product can be mixed with the naphtha stream from the atmospheric crude distillation and be sold as finished product. The top-product from this column contains C1-C5 hydrocarbons and some hydrogen. This stream could be further

purified to LPG, butane and fuel gas, but because these are not our desired products this was not designed. The stream is of significant size though, and the economic margin for processing this stream to LPG instead of using it as fuel was calculated (see appendix 8) and found to be 1.6 mln US$. It is thus recommended to look to possibilities to purify this stream and sell LPG as an end product, because the economics look promising.

2.2.7.3 Other downstream processing options

The only process that was mentioned earlier as a possibility for the downstream processing that is not reviewed yet is the oligomerisation unit. This unit can be used to convert light alkanes and alkenes to heavier hydrocarbons and thus increase the yield of desired products (diesel and kerosene). Since from the kinetics of the F-T reactor follows that only alkenes up to pentene are formed, only hydrocarbons up to around C10 could be

formed in the oligomerisation unit. The amount of alkenes formed is quite small so the increase in the production would be marginal. Furthermore the presence of alkenes in our products is allowed, so there is no need to remove or convert them. For these reasons it was chosen not to incorporate an oligomerisation unit in the process.

2.2.8 Remaining process units

In this chapter, so far the selection of processes for the production and processing of the main products were discussed. There are some units left though where a process had to be selected for the processing of wastes (water) of reactants (hydrogen). These will be briefly discussed in the rest of this chapter.

2.2.8.1 Hydrogen production for the hydrocracker

As one of the feed streams for the hydrocracker hydrogen is needed. This hydrogen has to be produced on-site because hydrogen is hard to transport and store. Possibilities are to build a separate hydrogen production facility or extract hydrogen from one of the process streams. In the latter case a surplus of hydrogen would have to be produced in the syngas production. When reviewing the process streams it is clear that two streams are present where hydrogen could be removed easily. These are the purified synthesis gas from the syngas production and the purge of the syngas recycle. Both streams have generally the same components so the same process(es) could be used to remove the hydrogen. Because the purge is in principle a waste stream this is the preferred stream to take the hydrogen from. This stream could also be manipulated (pressure, temperature) without major implications for the rest of the process. From the process stream summary