The design of a 1330 ktonnes per year low-energy catalytic reformer

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,~i

T

U

Delft

Technische Universiteit Delft

FVO Nr.

Preliminary Plant Design

Dept. of Chemical Process Technology

Subject

The design of a 1330 ktonnes per year low-energy

catalytic refonner

Authors

M.M.D. van den Berg

F .H. Janmaat

J.M. Mulder

AJJ.F. Nossent

Keywords

Telephone

015-2123654

015-2121115

015-2126655

015-2570275

Catalytic reformer, aromatics, hydrogen sulfide, hydrogen,

membrane systems.

Date of assignment :

Date of report

:

6/10/1997

24/12/1997

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FVO Nr.

Preliminary Plant Design

Dept. of Chemical Process Technology

Subject

The design of a 1330 ktonnes per year low-energy

catalytic reformer

Authors

M.M.D. van den Berg

F.H. Janmaat

J.M. Mulder

AJ J.F . Nossent

Keywords

Telephone

015-2123654

015-2121115

015-2126655

015-2570275

Catalytic reformer, aromatics, hydrogen sulfide, hydrogen,

rnernbrane systerns.

Date of assignment :

Date of report

6110/1997

24/12/1997

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Summary

The objective is to design a modem catalytic refonner on base of an existing refonner at a refmery in the Netherlands. This refonner should have a capacity of 156 ton naphta per hour. On a yearly basis (8520 operating hours) this is 133 billion ton. The naphta is converted to benzene, toluene and xylene (BTX). By-products are hydrogen and light hydrocarbons. The BTX products are mainly used as base chemicals for the chemical industry.

The existing catalytic refonner has the disadvantage of having a large recycle stream to the reactor consisting of hydrogen and light hydrocarbons, which do not usefully take part in the reaction. This which causes

unnecessarily large compression and heating costs. Also the reactor volume, which is the limiting factor in increasing plant capacity, is not used optimally. The design objective ofthis project was therefore to fmd a solution to increase the hydrogen concentration in the recycle.

The catalytic refonning process consists of four sections: Hydrogen sulfide removal

Reactor section

Hydrogen separation and recycle Product separation

1t was decided to remove the hydrogen sulfide from the feed naphta by distillation. This process has the

advantage over stripping the hydrogen sulfide offwith a treat gas th at there are no compression costs combined with the fact that no amine treatment and consequently no drying ofthe naphta is needed

The reactions take place in a series ofthree reactors with intennediate heating to make up for the temperature drop caused by ofthe overall endothermic reaction system. The initial reaction temperature is 538 °C; the pressure is 32 bara. The reactor section is based on the existing reactor section, because the exact kinetics ofthe complex reaction system are not known. Together to the three reactors a swing reactor is placed, so the process doesn't have to be stopped when the catalysts of a reactor needs regeneration. This mode of operation is called fully regenerative catalytic refonning.

The hydrogen separation is performed by a membrane system. Membranes are continuous and perfonn sufficiently weil. With the membranes a purity of 98% hydrogen and a recovery of 90% can be achieved.

Furthermore, it is a new technology ofwhich further improvement of performance can be expected. The product separation is carried out by four distillation columns, one to split the product in Cs and higher carbon numbers, one to separate the C/C4 in fuel gas and LPG, one to separate the product in CsfC7 and higher carbon numbers, and the last one to separate rest in Cs and C9+.

The total investment costs ofthe plant are 64.5 million dollar, the operating costs are 350.6 million dollar and the product income 404.1 million dollar. These figures are calculated with the fractional method [4]. A rate of return of45.7%, and discounted cash flow rate of return of55.0% are calculated. These figures are high, when th is is truly the result ofthe project it is advised to do the investment irnmediately. However one should bear in mind that these are just results of rough investrnent and cost estimations. The break-even point is after 4.1 years, when it is taken into account that it takes two years to design and built the plant.

A number of points deserve more attention and more effort would have been put into them if it wasn't for lack of time.

The reaction kinetics are not known exactly because ofthe complexity ofthe reaction system. Therefore it is very hard to simulate the reaction section properly and obtain proper estimations ofthe stream compositions between the reactors. As is known accurate models have been developed by different oil companies. The above and the fact th at improvement ofthe reactor section was not an objective led to the fact th at the authors had neither the time nor the inclination to put a lot of effort it.

The membrane system has been considered to be a black box. A vailable mode Is were either too complex or reached us too late. A company producing of membranes supplied us with the separation performance and conditions, which were used in the CHEMCAD simulation and in the economical calculations. The exact configuration ofthe system couldn't be supplied to us. The question is if it would be profitable for an oil company to put a lot of time and effort into research into a specialized subject like membrane systems.

When cooling streams below 180 °C, airfms are more favorable (cheaper) than heat exchangers using cooling water. Due to lack of time no literature search was done to find infonnation necessary for their design. Therefore, all heat exchangers for cooling or condensing we re designed to use cooling water.

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Acknowledgement:

We would like to thank the following persons for their assistance and support:

P.F. Hoeijmans, C.P. Luteijn, Prof. Sie, Jan Smit, E. Reitserna, Greg Fleming (Medal Membrane Separation Systems Du Pont Air Liquide, USA), T. Scullion (UOP Molecular Sieves, UK), dhr. Fakhri (Lab. API, TUDelft), Susanne Ehrhart (Ube Industries Deutschland), Stefan Tesselbaum ( University of Technology in Denmark), Mark Henry (Bend Research Technologies, Houston Texas), Kenny and Dolly (Nashville, Tennessee).

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Table of appendices

ApPENDIX 1. PROPERTIES OF THE COMPONENTS ... 2

ApPENDIX 2. LETHAL DOSES AND MAC V ALUES ... .4

ApPENDIX 3. SIMPLIFICATION COMPONENT LIST ... 6

ApPENDIX 4. PROCESS FLOW SHEET ... 8

ApPENDIX 5. STREAM TABLES ... 9

ApPENDIX 6. FLASH CALCULATION ... 24

ApPENDIX7. COLUMN CALCULATION ... 25

ApPENDIX 8. PUMPS ... 29

ApPENDIX 9. HEAT TRANSFER CALCULATIONS ... 30

ApPENDIX 10. EQUIPMENT LISTS ... 36

ApPENDIX 11. SPECIFICATION FORMS ... 47

ApPENDIX 12. PINCH DATA ... 85

ApPENDIX 13. HEAT INTEGRATION ... 87

ApPENDIX 14. CALCULATION EQUIPMENT COSTS BY FACTORIAL METHOD ... 89

ApPENDIX 15. INVESTMENT COSTS FACTO RIAL METHOD ... 92

ApPENDIX 16. UTILITY COSTS ... 93

ApPENDIX 17 . OPERATING COSTS FACTORIAL METHOD AND PROFIT CALCULA TION ... 94

ApPENDIX 18. INVESTMENT COSTS TA YLOR METHOD ... 95

ApPENDIX 19 . OPERATING COSTS B Y TA YLOR METHOD AND PRO FIT CALCULA nON ... 96

ApPENDIX 20. DCFRR CALCULATION ... 97

ApPENDIX 21. EVALUATION OF RESULTS ... 98

ApPENDIX 22. FAX MEMBRANE SPECIFICA nONS ... 99

ApPENDIX 23. CONTINUOUS PRESSURE SWING ADSORPTION ... 100

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Introduction, FV03212

1. Introduction

Catalytic reforming is one of the most important processes in the oil refining industry to produce gasoline fractions of higher octane number (antiknock properties) from hydrocarOOn fractions with a OOiling point in between 80°C and 180°C (naphta). Besides the reformate, catalytic reforming produces extensive amounts of hydrogen. The process is in use since the late 1940's and in 1990 the total capacity of catalytic reforming plants amounted to 8.7 million barrels per day worldwide [1]. In comparison, the 1990 world wide total capacity of catalytic cracking was 10.2 million barrels per day.

Catalytic reforming also is an important route for the production of the aromatic base chernicals: benzene,

toluene, xylene (BTX) and ethylbenzene. These aromatics are a high value feed for the chemical industry. The derivatives from these basic chemicals make thousands of products found in every home and office. The major endproducts of these chemicals are listed below:

Benzene: source of styrenic plastics, phenolic polymers and nylon.

Toluene: a major source ofbenzene through hydrodealkylation process; also useful as a solvent in paints. Ortho- and para-xylene: sources of polyester fibers, plastic film and plastic OOttles.

Meta-xylene: usually converted to para-xylene through catalytic isomerisation processes.

When the production of these aromatics is the primary objective ofthe catalytic reforming plant, the process has to be operated under more severe conditions (high temperatures and low hydrogen partial pressure) than when the objective is a higher octane number.

The aim of the project was designing a catalytic reformer plant for the production of aromatic base chemicals.

The design of the reactor section will be based on that of an existing catalytic reformer. In the existing catalytic reformer plant recycling costs are high due to a large recycle of fuel gas and LPG from the top of the final fractionator back to the primary fractionator. Furthermore the recycle stream of hydrogen back to the reactor feed stream has a concentration of only 65 vol.% hydrogen. The rest ofthis stream consists offuel gases (Cl and C2) which are inert in the reactor. These therefore only take up reactor space and decrease possible reactor

throughput.

The plant as a whole is required to have lower energy costs and larger possible throughput. This will be done by increasing the concentration of hydrogen in the hydrogen recycle and revising the layout of the fractionating section. Furthermore, for the separation of hydrogen and light hydrocarbon of the reactor effluent, new apparatus will be considered, like membranes.

In the future, the demand for aromatics is eX"pected to grow further because of the growing use of the products the aromatics are used for. The catalytic reformer process is the biggest supplier of these aromatics. This ensures that the reformer process will hold it' s place as an important part of oil refinery and justifies further research on how the process can be improved.

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Base of design, FV03212

2. Base of design

2.1. Objective and specifications

The objective of this project was to design a catalytic reformer based on an existing catalytic reformer. The new plant should have lower energy consumption as weIl as a high purity hydrogen recycle. Besides the aromatics,

the hydrogen is an important product. At this time the hydrogen is led to a low purity hydrogen header which has to be purified elsewhere on the refinery to obtain the quality needed for the high purity hydrogen header. If the hydrogen produced in the process is purified to such extent that it can be led to the high purity hydrogen header, not only wiIl the gas recycle to the reactor be smaller but also there will be no more need of external purification.

A selection of properties for all components in the process can be found in appendix 1. The feed of the process is naphta which can one of the products of the atmospheric distillation of cTUde oi! at an oil refinery. The components that are harmful to the catalyst have to be removed fiom the naphta. The sulfur content is lowered by transfonning it to hydrogensulfide with hydrogen in the hydrofining unit, followed by the removal of the produced hydrogensulfide. The hydrofining unit is not part of the design. Instead the feed naphta is taken to have a H2S concentration of 400 weight ppm. The maximum allowed hydrogensulfide concentration in the reactor feed is 0.2 weight ppm. Any water in the feed also is to be removed.

The catalytic reformer operates for 8520 hOUTS per year and has a capacity of 156 tons naphta per hOUT. The plant is designed with a depreciation time of 15 years. Section 2.5 deals with the extaernal specifications and the process boundary values.

2.2. The four subprocesses

The design of the process was based on a literature search on available technologies. The process was divided into foUT subprocesses:

• Reactor section • H2S removal section • Hydrogen recovery section • Product separation section.

These process parts are discussed separately in sections 2.2.1. to 2.2.4.

2.2.1. Reactor section

The design and improvement of the reactor section as such was not an objective in this project. Therefore the reactor section was based on the reactor section of an existing catalytic reformer. The existing reforming process is the fully regenerative catalytic refonning process. The principles and simulation of the reactor section are outlined in section 3.1.

2.2.2. H

2

S removal section

The H2S formed in the hydrofining unit from the bounded sulfur is to be removed fiom the feed to protect the catalyst in the reactor section. This operation can be performed in different ways and two methods were

reviewed:

• Stripping with light gas • Distillation

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2.2. 1. 1. Stripping with light gas

A possible way to separate H2S from the naphta is to make use of a stripper. Firstly, the H2S rich naphta has to

be cooled to 45°C. The naphta is led through a stripping column and is stripped with gas. This gas can be, depending on the process, purified H2 gas (98 vol%) or the vapourous product ofthe flash. Tbis last stream

consists ofabout 6 vol% H2; the rest is lower alkanes (Cj-C4).

Stripping with the flash vapour product has the disadvantages of dealing with a large gas stream and of the fact that light hydrocarbons will dissolve in the naphta, wbich is undesirable. The flow of lower hydrocarbons through the reactor uselessly takes up reactor space. Therefore the lower alkanes have to be distilled from the naphta. After being stripped offthe naphta the H2S is removed from the treat gas stream by an amine treater. In

tbis column the H2S is absorbed in a solution of diethanolamine and water. The purified gaseous stream is led

through a drying column to remove any water that was earried along from the amine treater. This water is poisoning to the catalyst in the reactor beeause it strips offthe eatalysts acidic chlorine sites [3]. The drying column is a packed bed ofzeo1ites (molecular sieves). The dryer unit consists oftwo beds ofwbich one is regenerated while the other is operating. From the drying column the gaseous stream is led to the hydrogen recovery section (membranes) where the lower alkanes and the naphta, wbich was carried along in the H2S

stripper, are separated from the H2 product stream. The heavy naphta hydrocarbons may be fouling for the

membranes. If so, they should be removed first.

When stripping with purified hydrogen (98 vol%) a smaller volume stream is needed to strip the H2S off. After

being stripped off the feed naphta, the H2S is to be removed from the H2 stream. Again tbis is done in an amine

stripper after which the H2 stream is dried. The hydrogen is led to the reactor, together with some hydrocarbons

from the naphta feed it carried along from the H2S stripper. The elegant aspect oftbis process design is that the

purified H2 is used both as treat gas and as the H2 recycle to the reactor. The obvious disadvantage ofusing the

purified hydrogen stream is the fact that a considerable gas stream loses its pressure over the membrane and bas to be repressurised (20~3-720 bara).

Using purified H2 as treat gas seems awkward in the fact that an already purified product stream is

contaminated. Nevertheless tbis method has a lot of advantages, as outlined above. The treat gas stream is smaller so a smaller H2S stripper and the amine treatment unit ean be used. Also no lower alkanes have to be

distilled off the reactor feed, elirninating the need for an expensive distillation column. 2.2.1.2. Disti/lation

Another way to remove the H2S, is by distillation. The top stream of the distillation column will be small and

rich ofH2S (105 kglhr, 83mol % H2S). The amine treatrnent described above can not increase the H2S

concentration. Therefore no amine treatment and no water removal unit is needed. Investment and operating costs of these units can be avoided.

A disadvantage of the use of a distillation column in comparison to the stripper, is the reboiler duty. The condenser duty is no extra disadvantage beeause the feed naphta doesn't have to be cooled before entering the distillation. The order of magnitude of the reflux is 6660 kglhr. The naphta feed has a flow of 156 tonne/hr. This means that the condenser duty is much less than the exchanger duty before the stripper.

2.2.3. Hydrogen recovery section

Typical catalytic reformer off gas has a hydrogen concentration of approximately 65 vol. %. This stream is partly recycled to the reactor where a certain hydrogen partial pressure is required for optimum operation. In the current situation the plant is operated at maximum reactor throughput. Therefore the limiting factor for increasing production is reactor space. The lower alkanes in the gas recycle make no contribution to the production of aromatics and therefore uselessly take up reactor volume. One of the objectives of tbis study was to increase the production capacity by increasing the hydrogen concentration in the gas recycle.

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There are various methods to separate hydrogen from 10wer alkanes. In this study we investigated four of them:

• cryogenic separation • pressure swing adsorption • membrane separation • oil wash.

2.2.3.1. Cryogenic separation

In cryogenic separation the separation is based on the difference in boiling point of the different species in the mixture. The difference with distillation lies in the fact that in cryogenic technology the fluids are gases which can not be liquid at ambient temperature. The low temperatures are reached by expanding a gas or liquid over a valve to a lower temperature (Joule-Tomsom effect).

For cryogenic reformer off gas separation the typical recovery is about 90% at 95% purity. The low

temperatures require the use of metals that do not become brittie at low temperature like stainless steel, nickel steel, copper or aluminum for the cryogenic equipment. Another disadvantage of cryogenic gas separation is the necessary pretreatment of the feed to remove aromatics by adsorption and higher alkanes by absorption, which would otherwise cause problems in the low temperature equipment.

The advantage of cryogenics in reformer off gas separation is that it is continuous, that it has relatively low utility requirements [18] and the possibility ofintegrated LPG recovery.

2.2.3.2. Pressure Swing Adsorption

In pressure swing adsorption the separation is based on the difference in selectivity of the adsorbent for

different species in a mixture. The adsorbent is usually activated carbon or zeolites or a combination of the two. The hydrogen lower alkane mixture is fed at high pressure to a vessel containing a packed bed of the lean adsorbent, after which the vessel is closed. For a certain period (in the order ofminutes) adsorption ofthe alkanes takes place. After this period the purified hydrogen is let out of the vessel at high pressure leaving behind the adsorbent loaded with alkanes. After this the vessel is closed again and another valve opens. The pressure in the vessel is now lowered to allow the adsorbed alkanes to desorb and leave the vessel at low

pressure. The vessel containing lean adsorbent is repressurized using part of the purified hydrogen and the process starts again.

The PSA hydrogen product typically has very high purity (>99.999%) [20]. Because the adsorption process described above is discontinuous, there is a need for using a multiple vessel arrangement to be able to fit PSA into a continuous process. The installation costs for these installations are high because of the need for a sophisticated valve system and the multiple vessels. Another disadvantage ofPSA is the low recovery, which is 70-75% for a four vessel configuration and up to 80-85% for a ten vessel configuration [20].

The advantages of PSA are that the purified hydrogen is produced at high pressure, eliminating most compressor costs for the hydrogen recycle and the fact that the adsorbent is not spent.

In addition to the regular PSA process some time was taken to investigate the possibility of a continuous form ofthe adsorption concept. The objective was to combine the advantage ofPSA, i.e. the production ofpurified hydrogen at high pressure, with the advantage of having a continuous process.

A rough sketch and an explanation ofthe design are given in appendix 23. 2.2.3.3. Membrane Separation

In membrane gas separation the separation is based on the difference in selectivity of the membrane for different species in a mixture. The selectivity of a membrane for a species is expressed by the permeability of the membrane for the species. Membrane technology is the youngest of all gas separation processes. A lot of effort is being put into research to optimize this technology and therefore further improvement of performance and reduction of installation and maintenance cost can be expected in the future of membrane technology. Several membrane systems exist to perform the hydrogen lower alkanes separation. In all cases the hydrogen will permeate, losing its pressure for the pressure drop of the permeate is the driving force of the membrane separation. Hydrogen recoveries of 90% at 98 vol% purity are typical for membranes [19, and appendix22]. The retentate are the lower alkanes at high pressure.

The advantage of membrane separation is the straightforward and continuous process configuration combined with good separation performance and low maintenance costs.

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The disadvantages of membranes are the pressure drop of the purified hydrogen product eausing high compressor costs and the need for membrane replacement after five years.

2.2.3.4. Oil wash process

In the oil wash process the separation is based on the difference in selectivity of the so called sponge oil for the different species in a mixture. The sponge oil used in this process ean be either an external oil or the feed naphta of the reformer process itself. The lower alkanes are stripped fiom the hydrogen lower alkane mixture by the absorbent oil. The oilloaded with lower alkanes is fed to a distillation column where the lower alkanes and the oil are separated. This process is the generally used method to perform the separation of the hydrogen lower alkane mixture.

The advantages of this process are basically the same as those for the membrane separation process; namely the straightforward and continuous process configuration combined with good separation performance and low maintenance costs. Furthermore there is the possibility of integrated LPG production.

The disadvantage is the high energy need in the form of the reboiler duty in the distillation column and large compression or cooling costs for the spunge naphta.

2.2.3.5. choice ofhydrogen recovery me/hod

In most of our process alternatives we decided to make use of a membrane system to perform the hydrogen lower alkane separation. This choice was based on the following points:

- PSA recoveries for the four or ten bed configurations are too low for our purpose. There is no need for very high purity hydrogen which is the main advantage of PSA. The continuous adsorption process was not developed far enough to be considered seriously.

- The oil wash is a widely used process and therefore presents no innovative challenge.

- Already membrane technology is a good alternative for cryogenic and adsorption separation methods with its own advantages and disadvantages. Further increase of the performance of membrane systems due to

development and optimization of membrane technology ean be expected in the (near) future. Installation of a membrane system now will probably allow the plant to be easily upgraded to make use of future improvements. This makes the use of membranes very attractive.

2.2.4. Product separation section

The product has to be divided into the desired product streams with the desired purity. This ean be done with the use of multiple distillation columns. We have reviewed two different arrangements of columns for the aromatics separation:

• Aromatics divided into three products in two columns • Aromatics divided into three products in one column

In both cases a fust distillation column separates the reformate fiom the light gases (LPG and fuel gas). Also in both cases, the LPG and fuel gas are divided in a separate column. This arrangement is most convenient beeause now the first column controls the purity of all aromatic products. Furthermore the first column makes the volume flow to be handled in the downstream columns much smaller allowing the use of smaller towers.

2.2.4.1. Aromatics divided in three products in !wo columns

This arrangement is the most simple and straightforward approach. The columns are simple with relatively normal reboiler and condensor duties.

2.2.4.2 Aromatics divided in /hree produc/s in one column

The advantage of this arrangement is clear beeause it means that two columns are replaced by one. However this replacement will make the column with the three productstreams more complex. The column needs a lot of stages and has relatively a high reboiler and condensor duties.

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2.2.4.3 Final choice

Looking at the results after modelling the two cases in CHEMCAD the following choice was made. The

separation of the aromatics in two columns was energetically more suitable, although the total number of stages is higher than when one column is used.

2.3. Process alternatives

2.3.1. Introduction

In this section four ofthe processes developed on the way to a final process choice are discussed. All processes we re simulated successfully in CHEMCAD. After comparison ofthe processes, the four processes mentioned in this section were considered to be the most feasible, both technically and economically.

How the results of economical calculations for all four ofthem and further discussion led to the eventual process choice, is outlined in section 2.3.2.

All processes are based on the original catalytic reformer. The heart ofthe process, the reactor section followed by a flash, is identical to the original design for all processes. A block diagram and a short explanation for each process are given below.

2.3.2. Processes

2.3.2.1. Process 1

H2S removal by the purified H2 recycle.

~ Stripper Naphta Reactor

Section Product Separation H,+ H,S Amine Treat + Dryer H, H, (98%) r l H, + CI/C4 CI/C4

I

Membrane H2 (98%) Prod~ct H2 Separation

figure 2.1: H2S removal by the purified H2 recycle.

Section

L

Fuel Gas +

LPG C9+

In this process the feed naphta is cooled to 45°C and stripped from H2S at 20 bara using part ofthe purified

hydrogen product (98vol%) from the membrane system. While the naphta main stream flows on to the reactor,

the H2 polluted by H2S and a small amount ofnaphta form the top stream. In an amine treater the H2S is stripped from the H2 and naphta stream. The H2 stream together with some original feed naphta, is dried and fed to the

reactor.

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Base of design, FV03212 2.3.2.2. Process IJ H2S removal by distillation. HS C7-2

rlC'

~

Naphta Reactor Product _Separation

Distillation Section Section

I

C9+ H₂+ CI/C4 H₂(98%)

--.

Fuel Gas + LPG Membrane H2 (98%) Product ~ H2 Separation _CI/C4 ~

Figure 2.2: H2S removal by distillation.

In this process the H2S is distilled directly from the naphta feed stream. The top product is a very small H2S rich

stream (3kmollhr, 63 vol% H2S). The naphta flows on to the reactor. The hydrogen product is purified by a

membrane system similar to that in process I.

As suggested the use of lift gas was tried in an attempt to reduce the columns reboiler duty. When operating at the conditions ofprocess I1, part ofthe lift gas dissolved in the naphta and in total major no advantageous effect was seen. By using lift gas one is combining this process II with process 1.

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2.3.2.3. Process 111

Oil wash with H2S removal by alkane treatgas.

To Amine treater Fuel gas LPG Products Reactor H" CI/C4 Naphta C 11 C4 Stripgas F d N hta

ee H2S Stripper 1----~"""---tI H, Separator

H, (95%) Product

L...-_ _ --' Naphta Recycle

figure 2.3: Oil wash with H2S removal by alkane treatgas.

C7-r

AIomatics Separator C9+ Fuel Gas CI/CZ Base of design, FV03212

In tbis process no membranes are used to purify the H2 rich stream; the design was made because the oil wash

is the generally used method. Instead the vapourous flash product is led to the oil wash. In the oil wash process the lower alkanes are separated fiom the H2 by absorpton in the so called sponge oi!. This sponge oil is the

naphta feed together with a naphta recycle (1:3 ratio). After absorption at high pressure (69 bara) the lower alkanes are separated fiom the naphta by a flash followed by a distillation tower (6 bara). The naphta is divided between the recycle and the reactor feed. The lower alkanes from the flash and the distillation are used as the treat gas to remove H2S fiom the feed naphta. After the amine treater and a drying column, the lower alkanes,

fuel gas and LPG, are separated in a distillation column.

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2.3.2.4. Process IV

Original catalytic reformer with membrane H2 purification.

C7

-Feed

Stripper N hta Reactor Product Separation

Section Section H, + ClIC4 H, + H,S H,(98o/.) +CJ/C4 ClIC4 Fuel Gas + LPG

Amine Trea! H,+CJ/C4 _Mem_b_ran_e _{H2 (}₉₈_%)_Prod_uct

+ Dryer _H,_S_e_p_arn_ti_o_n

figure 2.4: Original catalytic reformer with membrane H2 purification.

This process is mentioned in tbis section because it is very similar to the original catalytic reformer. The only difIerence is that H2 is purified by a membrane system before being recycled to the reactor. As in the original process the vapourous flash product is used as treat gas in the H2S stripper (45°C, 20bara). After the H2S stripper the lower alkanes are removed from the feed naphta stream by a debutaniser. Tbe vapourous flash stream, polluted with H2S is led to an amine treater, a drying column and is separated by the membrane system

into an CI-4 stream and a purified H2 stream.

2.3.3. Discussion

All four processes are built around the same reactor section and flash and are followed by a similar separation section consisting of four distillation columns. Therefore an economical comparison of the four processes could be and was based only on the difIerences between the processes. The investment costs for the processes are the equipment costs for all equipment minus tbe reactor section, tbe flash and the distillation section calculated using the metbod 00 pp218-223 [4]. The utilities needed for operating with tbis equipmeot were calculated using the methods outlined in the chapter 4. Table 2.1 shows the results of these calculations. In the calculations the possibilities of heat transfer between tbe process strearns was not taken into account.

Unfortunately the calculations tbernselves have been lost and were oot redone due to lack of time and therefore one can choose to ignore the results shown in table 2.1. Nevertheless the results are shown because the eventual decision was partly based 00 them. When examining tbe processes, an experienced eye can see that tbe results of the calculations are not far fetched_

Table 2.1: Results of equipment cost and utility_ need calculations for the fo ur process altematives.

Process I Process II Process III Process IV Investment

[Mi]

2.42 2.48 5.54 3.08 Electricity [MW] 8.56 8.45 12.41 8.31 Cooling water [kg/sj 430 327 446 485 HP steam [kg/sj 59.8 66.3 182 148

Process

m

will be the most expensive process. Tbe disadvantages of tbis process are the large naphta and lower alkane recycles, the inefficient use of compressor energy (69

-7

6

-7

69 bara) and the fact that bigh pressure equipment like towers, pumps and compressors will be required. In the case of compressors spare units will be needed because of the poor reliability of reciprocating compressors. Process I and IV are fairly sirnilar. Tbe discussion about the advantages of usiog H2 as tbe treat gas is io section 2.2.l.l. It was decided to use purified

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hydrogen as the treat gas because in process IV a distillation column is needed to remove the lower alkanes from the naphta reactor feed and the fact that a larger gas recycle to the H2S stripper is used.

This left the choice between process land process

n.

The economical and utility calculations show no large difference between the two. The fact that the removal of H2S fiom the naphta feed by distillation was not

mentioned in any literature found on the subject made it a more interesting process to study and design.

Therefore it was decided to continue the project focusing on process

n

.

2.4. Reaction and stoechiometry

The composition ofthe naphta feed stream can be found in the stream tables (appendix 5). Naphta consists of the following components:

Paraffins, which are very stabie and have the general formula CNH2Nt2, with straight or branched chains but

. without any closed ring structure.

Naphthenes are also called alicyclic hydrocarbons or cycloparaffins and are expressed by the general formula CNH2N. They are saturated hydrocarbons with a closed ring structure that can have one or more paraffinic side

chains. The naphthenes in naphta have rings of five or six carbon atorns.

Aromatics are non-saturated hydrocarbons containing one or two aromatic nuclei (benzene, naphthalene),

which may be linked up with paraffinic side chains. The general formula is CNC2N-6 .

The list of components was converted into a more simple list so that the calculations would he more

comprehensive without losing sight of the objective of the plant. This was done looking at the properties of the different components and taking similar components as one component. For more details see appendix 3.

2.4.1. Reactions

The most important reactions in the reactor can be divided into four classes, which are shown below. 1. Dehydrogenation

This type of reaction produces the most aromatics and hydrogen and is metal catalyzed.

Example:

()CHJ

_-

_{+3 H2}

-Methylcyclohexane Toluene

The reaction is highly endothermic, extremely rapid and controlled by the equilibrium. The conversion to aromatics is favored by low pressure, high temperature and a low

Hi

naphta ratio. Therefore, most ofthe dehydrogenation reactions take place in the first reactor, where the pressure is the highest.

2. Isomerisation

This type of reaction changes the geometric arrangement of a molecule without changing its composition.

Example:

Dimethylcyclopentane Metylcyclohexane

The reaction is acid catalyzed, slightly endothermic, very rapid and controlled by the equilibrium.

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This reaction is favored in the second reactor 3. Dehydrocyc/ization

Conversion of a paraffin to a 5 or 6 member saturated ring. Example:

C

I

c-c-c-c-c-c

Isoheptane

-Dimethylcyc10pentane

This type of reaction is endothermic, catalyzed by both acid and metal and favored by high temperature and low pressure. In the third reactor, where the pressure is 10west, the dehydrocyc1ization will take place most.

4. Hydrocracking

This type of reaction is undesirable because of the cracking of higher octane molecules and the loss of H2 . Example:

C-C-C-C-C-C-C +H 2

~

C-C-C + C-C-C-C

Heptane Propane Butane

The reaction is exothermic and irreversible. The reaction is slow but the rate increases with increasing temperature, residence time and carbon number. These reactions wiIl mostly take place the third reactor. 2.4.2. Stoechiometry

The determination of the exact stoechiometry of a complex reaction system of a component mixture as complex as the one studied was not done. Instead the data on the reactor feed and reactor effluent fiom the original process were used to determine a stoechiometric model based on one component which was assumed to be totally spent in the reaction. The conversions of all other components were based on the conversion of this component. The stoechiometric coefficients are shown in table 2.2. These coefficients were used in the stoechiometric reactor model in CHEMCAD.

Table 2.2: Stoechiometric coeeficients for all components to be used in the CHEMCAD stoechiometric reactor

Hydrogen 1381,0 MthCyc10hexane -104,6 Methane 238,5 Eth-Cyc1opentane -187,6

Ethane 243,3 Toluene 326,4

Hydrogen Sulfide 0,0 N-Octane -114,9 Propane 247,6 I -PropCycPentane -30,6 I-Butane 84,9 Ethylcyc10hexane -86,4 N-Butane 130,6 P-Xylene 220,9 I-Pentane 112,8 2-Methyloctane -88,5 N-Pentane 73,3 N-Nonane -68,2 Cyc10pentane 0,0 Isopropylbenzene 107,6 2-Methylpentane 41,6 Isopropyl-CycC6 -37,7 N-Hexane 14,4 Butylcyc1pentane -28,1 MthCyc10pentane 3,3 N-Decane -15,9 Benzene 19,2 N-ButylCycHexane -2,8 Cyc10hexane -25,0 N-Undecane -9,2 2-Methylhexane -47,4 1,2,4,5-TetMthBz -13,9 N-Heptane -132,9 Naphthalene -2,1

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Base of design, FV032 12

2.5 External specifications and boundary values for products.

2.5.1. Specifications

All streams have specifications that should be met.

Tbe maximum acceptable H2S concentration in the reactor feed is 0.2 weight ppm and no water is allowed.

In table 2.3 the required properties and maximum acceptable impurities of tbe product streams are tabulated. T _abI 23 M _{e .} _{aX1mum accepta e lmpunhes an propertles}. bI . _d 0 fth e pro uc streamsd t .

Stream Specification Temperature Pressure [0C] [bara] reactor feed 0.2 wppmH2S 538 32 reactor feed

o

wppmH20 538 32 reformate 2.5 g/kg C4_ C7- 25 glkg CsA 40 7 CsA 10 glkg toluene, 8.0 glkg C9+ 40 7 C9+ 40 g/kg CsA 40 7

Hydrogen 95 vol% hydrogen 40 43

Fuel gas 50 glkgLPG 40 6

LPG 25 g/kg Cs+, 25 glkg fuel gas 40 20

Tbe specifications for the utilities are shown in table 2.4. T bI a e 2 . 4 : ut lhes spec cahons il"· ifi

Utility Pressure [bara] temperature in [0C] temperature out [0C]

coolingwater 3 20 40

hp steam 40 410 250

I

(Tsat)

lp steam 3 190 133.5

I

(Tsat)

2.5.2 Location and location related restrictions

Tbe location of the catalytic reformer would be tbe location of the original plant, which is situated in a major industrial area in the Netherlands. No further effort was therefore put into finding a suitable site. Because of tbe fact that the process conditions, streams and operations are comparabie to tbose of the original process it was assumed that no special or extra restrictions applied.

2.6 physical properties list of all components

In tbis section a selection of useful physical properties is given. Tbe components are tabulated in the order of descending boiling points in table 2.5.

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T a e bI 25 C . . omponent propertles . . m t e h or d er 0 fd escen di ng boT 1 mgpomts.

no name structure normal molar melting liquid temp. 1 liquid temp. 2

boiling weight point density density

point 1 2 2

[0C] [kglkmol]

_lrc]l

[kglm3] rOCll [kglm3] [oCj

1 Hydrogen H2 -252.76 2.02 -259.20 77.63 -259.20 31.25 -239.97 2 Methane CH4 -161.49 16.04 -182.48 452.91 -182.48 159.98 -82.57 3 Ethane C2H6 -88.60 30.07 -182.80 650.11 -182.80 200.87 32.27 4 Hydrogensulfide H2S -60.35 34.08 -85.47 983.46 -85.47 348.25 100.38 5 Propane C3H8 -42.05 44.10 -187.69 810.93 -250.00 755.10 -187.71 6 2-Methylpropaan C4H1O -11.72 58.12 -159.61 730.90 -159.61 222.81 134.99 7 Butane C4H1O -0.50 58.12 -138.29 732.47 -138.29 228.27 152.03 8 2-Methylbutane C5H12 27.84 72.15 -159.90 777.49 -159.90 237.25 187.28 9 Pentane C5H12 36.07 72.15 -129.73 756.20 -129.73 231.44 196.50 10 Çycl~ntane C5HI0 49.25 70.13 -93.84 859.70 -93.87 272.36 238.61 11 2-Methylpentane C6Hl4 60.26 86.18 -153.60 794.81 -153.60 234.66 224.35 12 Hexane C6H14 68.73 86.18 -95.31 754.75 -95.31 232.42 234.28 13 Methylcyclo~ntane C6H12 71.81 84.16 -142.42 883.02 -142.42 263.91 259.64 14 Benzene C6H6 80.09 78.11 5.53 892.22 5.53 300.52 288.90 15 Cyclohexane C6H12 80.72 84.16 6.54 789.41 6.54 273.76 280.39 16 2-Methylhexane C7H16 90.05 100.20 -118.25 782.15 -118.25 237.93 257.22 17 Heptane C7H16 98.43 100.20 -90.58 771.46 -90.59 232.37 267.11 18 2,2,4- C8H18 99.24 114.23 -107.37 790.06 -107.37 245.63 270.81 Trimethylpentane 19 Water H20 100.00 18.02 0.00 1001.33 0.01 985.47 60.00 20 Methylcyclohexane C7H14 100.93 98.19 -126.57 885.46 -126.57 266.88 299.04 21 Ethylcyclopentane C7H14 103.47 98.19 -138.44 885.46 -138.44 262.20 296.37 22 Methylbenzene C7H8 110.63 92.14 -94.97 966.28 -94.97 298.54 318.60 23 Octane C8H18 125.68 114.23 -56.77 759.14 -56.77 231.64 295.68 24 iso-Propylpentane C8H16 126.45 112.22 -112.65 776.50 20.00 25 Ethylcyclohexane C8H16 131.80 112.22 -111.31 881.36 -111.31 249.42 336.00 26 Ethylbenzene C8HlO 136.20 106.17 -94.95 959.82 -94.95 284.23 344.00 27 1,4-Dimetylbenzene C8HI0 138.36 106.17 13.26 866.47 13.26 277.86 343.05 28 1,3-Dimetylbenzene C8HI0 139.12 106.17 -47.85 918.13 -47.85 280.42 343.85 29 2-Methyloctane C9H20 143.28 128.26 -80.37 785.48 -80.37 237.07 313.60 30 1,2-Dimetylbenzene C8HI0 144.43 106.17 -25.17 915.47 -25.17 284.11 357.15 31 Nonane C9H20 150.82 128.26 -53.49 771.04 -53.49 236.15 322.50 32 iso-Propylbenzene C9H12 152.41 120.19 -96.01 954.18 -96.01 275.83 357.85 33 iso-Propylhexane C9H18 154.76 126.24 -89.39 882.36 -89.39 272.06 353.85 34 Butylcyclopentaan C9H18 156.60 126.24 -107.97 876.12 -107.97 261.40 347.85 35 l-Methyl,3- C9H12 161.33 120.19 -95.54 955.49 -95.54 281.49 364.00 ethylbenzene 36 l-Methyl,4- C9H12 162.01 120.19 -62.32 931.07 -62.32 281.48 367.05 ethylbenzene 37 1,3,5- C9H12 164.74 120.19 -44.73 915.37 -44.73 277.72 364.15 Trimethylbenzene 38 l-Methyl,2- C9H12 165.18 120.19 -80.80 960.89 -80.80 281.55 376.85 ethylbenzene 39 ter-Butylbenzene C1OH14 169.15 134.22 -57.88 929.71 -57.88 272.80 386.85 40 1,2,4- C9H12 169.38 120.19 -43.82 924.17 -43.82 279.67 375.95 Trimethylbenzene 41 iso-Butylbenzene CI0H14 172.79 134.22 -51.45 909.80 -51.45 280.79 376.85 42 sec-Butylbenzene ClOH14 173.33 134.22 -75.43 930.38 -75.43 270.07 391.39 43 Decane CI0H22 174.15 142.29 -29.64 765.83 -29.64 235.04 345.30

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Base of design, FV03212 isopropylbenzene 45 1,2,3- C9H12 176.12 120.19 -25.36 928.84 -30.00 290.71 391.35 Trimethylbenzene 46 1-Methyl-4- C1OH14 177.13 134.22 -67.90 922.80 -67.90 272.52 378.85 isopropylbenzene 47 1-Methyl-2- C1OH14 178.18 134.22 -71.51 945.64 -71.51 273.92 383.85 isopropylbenzene 48 Butylcyclohexane CIOH20 180.98 140.27 -74.73 864.77 -74.73 262.68 393.85 49 1,3-Diethylbenzene C1OH14 181.14 134.22 -83.89 942.46 -83.89 275.05 389.85 50 Butylbenzene C1OH14 183.31 134.22 -87.85 943.09 -87.85 270.23 377.35 51 1,2-Diethylbenzene C1OH14 183.46 134.22 -31.22 918.57 -31.22 274.66 394.85 52 1,4-Diethylbenzene C1OH14 183.79 134.22 -42.83 910.41 -42.83 270.07 384.75 53 Undecane CllH24 195.93 156.31 -25.58 772.29 -25.58 231.42 365.61 54 1,2,4,5- CI0H14 196.84 134.22 79.23 839.18 79.23 280.82 402.85 Tetramethylbenzen e 55 Naphtalene CI0H8 217.99 128.17 80.28 993.93 60.00 314.90 475.25 14

(22)

Process structure and flowsheet, FV03212

3. Process structure and flowsheet

3.1 Reactor design

3.1.1. Reaction thermodynamics

The feed naphta contains paraffins, naphtenes with five or six carbon atom cycles and aromatics. Since the main goal of the process is to increase the aromatic concentration, the desired reactions are dehydrocyclyzation of paraffins, dehydrogenation of cyclohexanes, and dehydroisomerisation of alkyl cyclopentanes to aromatics. Even the best catalyst are not completely selective for these reactions. Some undesired reactions that decrease the liquid yield (gas formation by cracking) and deactivate the catalyst (formation of carbonaceous deposits on the catalyst) will occur to some extent.

Paraffins, mostly linear, are the main components ofvirgin naphtas. On the reforming catalyst the main reaction of the paraffins are the following: dehydrogenation to olefins, isomerization, dehydrocyclyzation to aromatics and naphtenes, and hydrocracking or hydrogenolysis to lighter paraffins. Naphthenic hydrocarbons are alkylcyclopentanes and alkylcyclohexanes , and the possible reactions are dehydroisomerization of alkylcyclopentanes and dehydrogenation of alkylcyclohexanes to produce aromatics, isomerization, and ring opening to produce paraffins.

In table 3.1 the reaction enthalpies [pISI, 15] at 500°C ofthe most important reactions that occur during catalytic reforming of naphta are tabulated.

Table 3.1: Thermodynamic data at 500°C of some naphta reforming reaction s

Reaction Mfr (kJ/mol) cyclohexane <=> benzene

+

3 H2 221.1 methylcyclohexane <=> toluene

+

3 H2 215.6 methylcyclopentane <=> cyclohexane -15.9 methylcyclopentane <=> benzene

+

3 H2 205.2 n-hexane <=> l-hexene

+

3 H2 129.8 n-hexane <=> 2-methylpentane -5.9 n-hexane <=> 3-methylpentane -4.6 n-hexane <=> benzene

+

4 H2 266.3 n-heptane

+

H2 <=> butane

+

propane -51.5 n-heptane

+

H2 <=> n-hexane

+

methane -62.0

Present-day reforming operations require very clean feedstock for obtaining continuously high-quality

reformate combined a long catalyst life. Naphta hydrotreatment is practiced in order to lower the concentration of the sulfur which is poisoning for the platinum reforming catalysts. The catalysts used for hydrotreatment of naphta consists of fieshly prepared oxide mixtures, typically CoO-MoÛ:3 or NiO-MoÛ:3, on activated alumina. They catalyze the reaction fiom bounded sulfur to dihydrogensulfide through hydrogen addition.

Although hydrogen is harmfu1 for the thermodynamics and kinetics of the desired reactions, the process is carried out in the presence ofhydrogen to decrease the catalyst deactivation produced by coverage ofthe active surf ace with coke.

3.1.2. Reactor layout

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swing reactor. This swing reactor can be coupled to each of the main reactors in order to regenerate one of the main reactor's catalyst without stopping the process. The catalysts form a packed bed in this reactor.

Because the overall reaction system is endothermic (tabie 3.1), the reactors are operated at high temperature (538°C). These high reaction temperatures will cause the equilibrium to shift towards the products. The temperature of the reactant flow will decrease due to the endothermic reactions. Therefore reactor section consists ofthree reactors with intermediate heating. In table 3.2 the properties ofthe flows in and out ofthe reactors as weIl as reactor volumes and catalyst loading are given.

T bi 32 a e : propertles 0 f t e ows 10 an h fl d out 0 f tereactors an h d reactor vo umes an catalyst oa d I ding reactor volume [m3] catalyst mass [tonne] Tin Tout Pin Pout

[0C] _rC] _[bara] _[bara] Reactor 1 10 6.5 538 448 32 31 Reactor 2 23.7 15 538 488 31 29.5 Reactor 3 69.2 45 538 528 29.5 27.5 Swing Reactor 69.2 45

Height

With the ratio = 2

Diameter

In the first reactor most of the dehydrogenation, a very endothermic reaction, takes place. Therefore, the temperature drop is largest in this reactor. In the second reactor, after reheating ofthe reactant stream by fired heater to reaction temperature, isomerisation of cyclopenta-alkanes to cyclohexa- alkanes takes place most. Beside this slightly exothermic reaction, other endothermic reactions take place. Therefore a smaller

temperature drop is measured. The effluent of the second reactor is also led through a fired heater. To obtain a sufficiently large amount of BTX' s, a third reactor is placed. In this reactor, cracking and other undesirable reactions occur more than in other reactors. The first reactor has the smallest volume. This is because of the highly endothermic reaction which quickly causes a large temperature drop so that a small residence time is wanted.

3.1.3. Catalyst

3.1.3.1. Function ofthe cata/yst

The desired reactions require two types of function of the catalyst: a hydrogenation-dehydrogenation function (metal fiom Group VIII ofthe periodic tabie) and a cyclization or isomerisation function. The last function is provided by oxides with acid properties.

The catalyst in the catalytic reformer process consists of:

metal sites: platinum and optional promoter metals (Re, Ir., San). acid sites: alumina with chloride

The catalyst is bifunctional: both acid and metal sites are catalytic and both sites are needed for good

conversion. Nevertheless acid and metal sites catalyze different reactions. Both proximity and balance between the acid and metal function are needed for optimum performance.

During catalytic reforming, hydrogen LPG (C3-C4) and fuel gas (C1-C2) are produced in addition to the

reformate. Light gaseous products (C1-C4) have a relatively low economic value, consequently, the selectivity of the catalyst towards the heavy products is very important.

3.1.3.2. Cata/yst deactivation

In the oil refining and petrochemical industry, deactivation ofthe catalyst by carbonaceous deposits is an important technological problem. Deactivation in the reforming process is caused by coke formation and deposition on the catalyst. In fact, for reforming reactions of naphta, the thermodynamics are such that it wou1d be desirabie to work at high temperature and low pressure. However, such operating conditions favor coke

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formation, and therefore many reforrning units operate under high pressure in order to increase the lifetime of the catalyst.

3.1.3.3. Catalyst regeneration

The catalyst, which has the form of smal! cylindrical stretched particles, is loaded in the reactor in a packed bed. During operation the catalyst deactivates and has to be regenerated. The reactor of which the catalyst has to be generated is taken off the process by replacing it with the swing reactor. This operation is called the reactor swing. There are 120 swings a year; the first reactor is swinged 30 times a year the second and the third are swinged 40 and 50 times per year respectively. To regenerate the catalyst fisrtly the coke is removed by burning it off with diluted air. Afterwards chloride is added to restore the chloride concentration on the catalyst. A little bit of sulfur is also added to decrease the great initial hydrogenic activity of the metal function.

The catalyst has a lifetime of 5 years after which it is replaced by a new batch.

3.2 Flash

The flash is the fust separation step ofthe product. The cooled reactor effluent is flashed to separate the hydrogen. The top product consists of most ofthe hydrogen and light hydrocarbons (Cl -C4). These are led to

the membrane system to be separated. The bottom product is led to the distillation section. Flash operating conditions and design results are given in table 3.3. The temperature of 45°C was determined by trial and error. It proved that at this temperature a good hydrogen recovery was combined with little flow of higher (C5+) alkanes.

T bI 33 FI h a e

..

as operatmg conditions and design results.

Pressure (bara) 20 Temperature (0C) 45 Volume (m3) 30.8

3.3 Membranes

The membranes are used to separate the hydrogen fiom the light hydrocarbons. A literature search did not provide enough information to completely design a membrane module. Therefore several companies known to produce industrial size membrane systems have been contacted in the search for more detailed information on which to base the design of a membrane unit. Understandably not much information was released, but most companies were willing to use their expertise and modeling ability to provide operation data, performance and costs of a membrane unit based on the specifications for this process. The data used are those provided by a company called Medal, which is a daughter company of Air Liquide in the United States (appendix22). The operating conditions and design results ofthe membranes are given in table 3.4.

T a e bI 34 . : mem rane operatmg con hons a nd design resultsb .

H2 purity % 98

H2 recovery % 90 Pressure in [bara] 20 Permeate (H2) pressure out [bara] 3 Retentate pressure out [bara] 20

It must be noted that membrane systems can operate at different permeate purities. However a higher permeate purity results in a lower recovery. Also the permeate output pressure is not the minimum of 1 bara, which would maximize the driving force for the separation. This is done because the compressor work is calculated as the integral ofthe pressure times the change in volume as shown in equation 4.29. The change in volume flow when compressing fiom 1 to 3 bara is relatively large requiring relatively high compression costs.

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3.4 H

2

S Separation

3.4.1 H

2S distillation column

This column is designed to separate the hydrogensulfide fiom the naphta feed. This feed is coming in at a temperature of 200°C and a pressure of 20 bara. It was chosen to operate at the pressure of the feed naphta to have no extra compression costs. The hydrogen sulfide should he removed fiom the naphta so that the amount ofhydrogensulfide in the reactor feed doesn't exceed the 0.2 ppm. At feed pressure and at given specifications the tower was designed in CHEMCAD. The operating conditions and design results can he found in table 3.5. T bI a e 35

..

T ower Tl operatmg con hons an d design results.

Top Temperature

roq

153.7 Bottom Temperature [0C] _269.1 Numher of Trays 23 Feed Tray 1 Condenser Duty [kW] -27437 Reboiler Duty [kW] 16845

3.5 Product separation section

3.5.1 Debutanizer

The maximum amount oflower alkanes (C4-) in the reformate stream is 2.5 grfkg C4. The tower operates at 20 bara. It is undesirable to operate at lower pressure hecause condensation wiIl then he at such a temperature that cooling water can not he used as the coolant. Cryogenic cooling wiIl he necessary in that case. At given

pressure and specifications the tower was designed in CHEMCAD. The operating conditions and design results can he found in table 3.6.

T bI 36 T a e

..

ower T2 operatmg con ditlOns a nd design results. '

Top Temperature

[oq

39.7 Bottom Temperature [0C] 245 Numher of Trays 28 Feed Tray 10 Condenser Duty [kW] -6651 Reboiler Duty [kW] 23826 18

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3.5.2 Fuel gas / LPG splitter

This is the only column that needs cryogenic cooling. This column is necessary in every process design. The specifications for the separation is: in the fuel gas, a maximum of 50 gr/kg C3 is allowed, in the LPG a maximum of 25 gr/kg CS and 25 gr/kg C2-. It operates at 20 bara. Decreasing pressure wiIl cause the

condensation temperature to he even lower. An expander is used after the column to depressurize the topstream to meet with the product specifications. At given pressure and specifications the tower was designed in

CHEMCAD. The operating conditions and design results can he found in table 3.7.

conditions and design results. r---~----~~----~ -44.9 75.2 19 9 -3260 1570

3.5.3 CT /

CS+ splitter

The pressure of the feed of tbis tower is lowered to 7 bara, the specified product pressure. Tbis is done hecause at low pressure the reboiler and condenser duties are lowest. The design specifications involving the separation of the components are: the maximum amount of C8A in the CT stream is 25 gr/kg. At given pressure and specifications the tower was designed in CHEMCAD. The operating conditions and design results ean he found in table 3.8.

T bI a e 38

..

T ower T4 operatmg con hons an d design results Top Temperature [0C] 173.9 Bottom Temperature [0C] 232.8 Numher of Trays 67 Feed Tray 25 Condenser Duty [kW] -12600 Reboiler Dlltr [kW] 13901

3.5.4 CS / C9+ splitter

The last column in the product separation section also operates at 7 bara. The design specifications involving the separation ofthe components are: a maximum of 10 gr/kg toluene and a maximum of 8.8 gr/kg C9A. The C9+ stream may contain a maximum of 40 gr/kg C8A. At given pressure and specifications the tower was designed in CHEMCAD. The operating conditions and design results can he found in table 3.9.

T bI a e 39

..

T ower T5 operatmg con hons an d design results Top Temperature [0C] 227.7 Bottom Temperature [0C] 244.8 Numher of Trays 62 Feed Tray 30 Condenser Duty [kW] -22060 Fired heater [kW] 24647

(27)

3.6 Pumps

The pumps that are used are centrifugal pumps. These pumps are widely used in the process industry and can be used for the operating conditions of this process. All pumps are single-stage.

3 .

7 Compressors

As with the pumps the compressors that are used are centrifugal compressors. The operating conditions of this process do not require the use of reciprocating or other high cost compression equipment. The assurned maximum compression duty of one compressor is 500 kW. When higher compression duty is required a series of multiple compressors is used. The number of compressors in series is equal to the total compressor duty divided by the maximum duty of one compressor.

3.8 Thermodynamics

The therrnodynarnic model that has been used in the flowsheeting software CHEMCAD is the Soav e-Redlich-Kwong model. This model will apply because all components are hydrocarbons which are processed at moderate temperatures and pressures. Only CHEMCAD has been used to calculate phase behavior and enthalpies.

3.9 Process flow scheme

In appendix 4 the process flowsheet can be found. Reactor feed preparation

The feed is coming into the H2S removal column. A H2S rich stream is the top product, a small amount of

hydrocarbons flows along. The bottom stream is the purified naphta, which is mixed with the high purity (98 vol%) hydrogen recycle. The top stream is the only waste stream produced by the process. Because it was not known in what kind of process this stream would be treated at what conditions, it was chosen to leave the stream as it is.

To the bottom stream approximately 50 mole% hydrogen is added. The flow is heated to assure that there is no liquid left and then led to a compressor where it is compressed to reactor pressure (32 bara). After compression,

the stream is heated to 508°C by the hot reactor effluent stream. To heat the feed it to the reaction temperature (538°C), it is heated in a fired heater.

Reactor section

The naphta is led through in total three reactors with intermediate reheating to 538°C (reaction temperature). The reactor effluent (528°C) is used to pre-heat the reactor feed due to which it cools down to 432.5

oe.

Hereafter the stream is used as the hot fluid in the reboiler of tower 1 due to which it cools down to 313

oe.

The stream is now cooled by cooling water in a heat exchanger to 192.3 °C which is its dewpoint temperature and then it is partly condensed to 45°C. The product stream now has a pressure 20 bara due to pressure losses in the four heat exchanging steps.

Hydrogen purification

After the flash, the top stream containing H2 (65 vol%) and lower alkanes are led to the membrane system

where it is separated into a H2 permeate stream with a purity of 98 vol% and a retentate stream. The pressure of

the perrneate stream is low due to a pressure drop in the membranes. The H2 stream is compressed to 20 bara and split into a H2 product stream and the hydrogen recycle. The split ratio is approximately 60/40 %. The H2

product is compressed to 40 bara and cooled to 40°C. Hydrocarbon product purification

The retentate stream fiom the membranes is led together with the bottom stream of the flash to the debutanizer,

which is operated at 20 bara. The LPG and fuel gas, which are the top product ofthe debutanizer, are separated

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from each other in a fuel gas / LPG separator. The fuel gas, the top product (-44

oe

,

20 bara), is expanded to 7

bara and -75

oe

and is used to partially condense the top stream ofits own tower. It is then heated to 40

oe

in two steps. The first step is a heat exchanging step to 20

oe

with used cooling water (40°C) . The second step is a heat exchanging step to 40

oe

using low pressure steam. The LPG, the bottom, is cooled to 40

oe

by cooling

water.

From the bottom stream of the debutanizer, the

eT

is separated in a next column after the stream is expanded to 7 bara. The top stream (174°C) is first condensed to 150

oe

by cooling water and then cooled to 400

e

by cooling water. The bottom stream is led to the final distillation column which is operated at 7 bara, where

es

and

e9+

are separated. The top stream (228°C) is condensed to 228

oe

by cooling water and then cooled to 40

oe

by cooling water. The bottom stream (245°C) is cooled to 40

oe

by cooling water.

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Equipment calculations, FV03212

4. Equipment calculations

4.1 Reactor

A method is proposed for calculating the equilibrium constants of the many complex reactions that occur during catalytic reforming of C6-C9 mixtures [14]. The method uses individual equilibrium constants for each carbon number group. These individual equilibrium constants give a simple and fairly precise way to calculate the parrafins and aromatics fractions.

The method assumes that C6 paraffins and naphthenes do not change their carbon number when they are converted to aromatics during catalytic reforming. The C7-~ paraffins and naphthenes are converted to aromatics and cracked to lower paraffins. The ratio in which the paraffins and naphthenes are cracked are derived from experimental data. CIO is assumed to be completely cracked. The cracking ratios are given in table

4.1:

T bI a e 41 .. C rac e k dh eavy components. Carbon number Percentage cracked

C7 37

Cg 37

C9 50

CIO 100

These cracked paraffins and naphthenes are converted to lower components, the ratios in which these lower components are formed are given in table 4.2:

Table 4.2: Formed lighter com ponents. Carbon Percentage formed

Methane 9 Ethane 17 Propane 21 Butane 14 2-Methylpropane 10 Pentane 10 2-Methylbutane 18

The remaining components are assumed to take part in the reforming reactions. The key reaction of catalytic reforming converts paraffins to aromatics and hydrogen as follows:

i = 6, 7,8 or 9

For which the equilibrium reaction is:

Where: Ki,p Xi,A Xi.?

PH

2

K

I,p

=[~}4

H2 xi,p

Equilibrium constant for the paraffins in the ith carbon group. Mole fraction aromatics of the ith carbon group of the total mixture. Mole fraction paraffines of the ith carbon group of the total rnixiure. Partial pressure of hydrogen.

(4.1) (4.2) [bara4] [-]

r -]

[bara] 22

(30)

Equipment calculations, FV03212

Temperature effects on the equilibrium constant can be calculated as an integrated form ofthe Van't Hoff equation:

(

1000)

lnKj,p

=

a

T

+b

(4.3)

The equilibrium constants for dehydrocyclization of paraffins to aromatics for temperatures in the range of 650-8200

K are obtained by using the derived coefficients oftable 4.3 to calculate equation 2. T a e bI 4 3 C ffi

..

oe Clen s or e lyC ocyc lization of paraffins' t ti d h dr .

Carbon number a b C6 -32.60 52.17

C7 -31.24 53.36

Cg -29.70 53.76

C9 -28.94 53.78

An easier way for equation 4.4 is by letting B symbolize the mole fraction of each carbon number group. Then the material balance is:

,,9

B"

=

1.0 L...,

=6

When the naphthenes concentration is neglected: or

Substituting equation 4 in equation 2 gives:

[ X'A }

K -

I, 4

j,p-

B _

H₂ j Xj,A Rearranging gives:

[

K

]

- I,p

B

Xj,A - 4

+K

.

j

PH

2 I,p X1, .p

=B

J -X'I, A

From the aromatic fraction that has been derived, the paraffins fraction can be found:

(4.4)

(4.5)

(4.6)

(4.7)

(4.8)

From these fractions the total mol quantities of aromatics and paraffins can be calculated, for each carbon number group. Out of these results the hydrogen balance is corrected by increasing the quantity of hydrogen. When this method was used for the reactor feed, a realistic result appeared as it was compared with

experimental data of the reactor effluent. However in the simulated process a stoeichiometric reactor was used, for which no kinetic models are necessary. Data were available for experiments in which the same feed and process conditions were used as in the simulated processes. From these experimental data ofthe reactor feed and effluent the stoeichiometry was calculated and used.

The design of a 1330 ktonnes per year low-energy catalytic reformer

,~i

T

U

Delft

FVO Nr.

Preliminary Plant Design

Dept. of Chemical Process Technology

Subject

The design of a 1330 ktonnes per year low-energy

catalytic refonner

Authors

M.M.D. van den Berg

F .H. Janmaat

J.M. Mulder

AJJ.F. Nossent

Keywords

Telephone

015-2123654

015-2121115

015-2126655

015-2570275

Catalytic reformer, aromatics, hydrogen sulfide, hydrogen,

membrane systems.

Date of assignment :

Date of report

:

6/10/1997

24/12/1997

FVO Nr.

Preliminary Plant Design

Dept. of Chemical Process Technology

Subject

The design of a 1330 ktonnes per year low-energy

catalytic reformer

Authors

M.M.D. van den Berg

F.H. Janmaat

J.M. Mulder

AJ J.F . Nossent

Keywords

Telephone

015-2123654

015-2121115

015-2126655

015-2570275

Catalytic reformer, aromatics, hydrogen sulfide, hydrogen,

rnernbrane systerns.

Date of assignment :

Date of report

6110/1997

24/12/1997

Summary

Acknowledgement:

Table of contents

Table of appendices

1. Introduction

2. Base of design

2.1. Objective and specifications

2.2. The four subprocesses

2.2.1. Reactor section

2.2.2. H

S removal section

2.2.3. Hydrogen recovery section

2.2.4. Product separation section

2.3. Process alternatives

2.3.1. Introduction

2.3.2. Processes

I

L

rlC'

~

I

--.

C7-r

2.3.3. Discussion

[Mi]

m

-7

-7

_-

_lrc]l