High temperature sulphur removal in a 250 MW integrated coal gasification combined cycle power plant

,~Ii

T

U

Delft

FVONr.

Fabrieksvoorontwerp

Vakgroep Chemische Procestechnologie

Onderwerp

High temperature sulphur removal in a 250

MW integrated coal gasification combined cycle

power plant

Auteurs

R.

van

de Water (Remko) M. Blokland (Marco) C.H.M. Boons (Maarten) B.W. Hoffer (Bram)

Keywords

Telefoon

015-2145623 0184-412514 076-5413062 015-2146579

IGCC, Coal Gasification, Power Plant, Regenerative Desulphurization, Monolith

Datum opdracht

Datum verslag

3 February 1997

2

May

1997

• • •

Fabrieksvoorontwerp

Vakgroep Chemische Procestechnologie

Subject

High temperature sulphur removal in a 250 MW

integrated coal gasification combined cycle

power plant

• •

Demkolec

_.

:

\

_{Demo KV-STEG}

_Buggenum

/

Summary

The objective of the design was to develop a 250 megawatt Integrated coal Gasification Combined Cyc1e (IGCC) power plant with high temperature sulphur removal. The advantage of this new type of syngas c1eaning is the potential to increase the plant efficiency by better heat management. The general plant set-up was modeled according to the existing Demkolec plant in Buggenum, the Netherlands. The high temperature sulphur removal is realized by a rotating monolith reactor with integrated regeneration of the

MnO/y-AI₂0₃ catalyst. This reactor which is designed to reduce the H₂S

content of the syngas to below 10 ppm can be continuously on-stream thanks to its self-Iubricating properties. The catalyst is regenerated with S02 leading to the direct formation of S2'

The designed plant uses 2200 ton co al dai1 and has an electricity output of

254 MW(overall efficiency 37.5 %). The Demkolec plant uses ca 1900 ton

coal dail and produces 253 MW (overall efficiency 43%). The differences

can be accounted to the insufficiently detailed design of the electricity generation section in the preliminary design (cooling water losses are too

high). However other losses in the preliminary design (7.1 %) are lower than

in the Demkolec plant (20 %). A part of this difference can be accounted for by better heat management in de desulfurization section (high-temperature) in the preliminary design.

The plant as designed here is not profitable. The Internal Rate of Return

(lRR) is -2.63 %, the Return On lnvestments (ROl) is -6.50 %. The main

cause for this loss is the large nitrogen stream used and the high equipment costs.

The NO_x emlSSlOn of the designed gas turbine is 79.24 g Gr1 (legal

maximum 95 g Gr1). The NO_x emission of the total designed plant is 66.93 g

We would like to express our sineere thanks to Dr. Van Langeveld, Ir. Luteijn, Prof. Ir. Grievink, Ir. Paaijens, Dhr Smit en Theo.

TABLE OF CONTENTS

1. INTRODUCTION ...••...•..•...••..•...•.•.•.•••.•...•...•...•...•.. 1

1.1 COAL AND ELEcrRICITY GENERATION ...••....•...•••.•...•...•... 1

1.2 INTEGRA TED GASIFICATION COMBINED CYCLE ...•....•...•.•.•.•....•...•.•...•.•...• 1

1.3 H2S CORROSION ...••...•...••.••...•...• 2

IA HIGH TEMPERA TURE SULPHUR REMOV AL ... 2

2. PROCESS INPUT INFORMATION ...•.•...•...•..•...•...•...•... 4

2.1 BA TIERY LIMITS AND EXTERNAL SPECIFICATIONS ...•...•••... 4

2.2 SELEcrION OF PROCESS ROUTE ...•..•...•.•... 5

2.2.1 Pre-treatment ..................... 5

2.2.2 Gasification ...... ; ... 5

2.2.3 Desulfurization ...... 6

2.2.4 Electricity generation ... 7

2.3 LOCATION OF THE PLANT .•...•...•...•.•... 7

204 LIST OF US EO CHEMICALS ...••...•...•.•... 7

3. PROCESS STRUCTURE AND PROCESS FLOWSHEET •...•... 8

3.1 INTROoUcrION ... 8

3.2 PRE-TREATMENT ... 8

3.2.1 The crusher ....................... 8

3.2.2 The pulverizer ... : ........................... 8

3.2.3 The dryer ...... 9

3.2.4 The bag house .......................................... 9

3.3 GASIFICATION ... 9

3.3.1 The gasifier ..................... .................. 10

3.3.2 The slag qllencher ... 11

3.3.3 The slag crusher ... 11

304 DESULFURIZATION ... 11

3.4.1 The mOllolith reactor ...... 11

3.4.2 The cold trap ... 11

3.4.3 The su/phllr combustor ... 12

3.5 ELEcrRICITY GENERATION ... 12

3.5.1 The Schllmacher filter ............................ ....... 12

3.5.2 Combilled cycle ... 12

3.5.3 The steam cycle ... 13

3.5.4 The heat exchallgers ... : ... 13

3.5.5 The pumps .................................. 14 4. EQUIPJlIIENT DESIGN ...... 16 4.1 BELTCONVEYER ... 16 4.2 THE CRUSHER (M 1 ) ... 16 4.3 BELTCONVEYER ... 17 404 THE PULVERIZER (M2) ... 17 4.5 THEDRYER(M3) ... 18 4.5.1 The cyclolle ............................. .... 19 4.6 THE BAGHOUSE (M4) ... 21 4.7 THE GASIFIER (R 1) ... 23

4.8 THE SLAG QUENCHER (VI) ... 24

4.9 THE SLAG CRUSHER (M5) ... 25

4.10 THE DESULFURIZATION REACTOR (R2) ... 25

4.11 THE COLO TRAP (M6) ... 28

4.12 THE SULPHUR COiVIBUSTOR (R3) ... 28

4.13 THE SOLIO FILTER (M7) ... 29

4.14 THE HEAT EXCHANGERS ... 30

S. 111ASS AND HEAT BALANCE ... 35

6. PROCESS CONTROL ........ 36

6.1 INTRODUCfION ... : ... 36

6.2 THEGASIFICATION SECfION ... 37

6.3 THE DESULFURIZATION SECfION ... 37

6.3.1 The monolith reactor (R2) ...... 38

6.3.2 Sulphur combustor (R3) ... 38

6.3.3 The cold-trap ...... 38

6.4 THE STEAM CYCLE ... 38

6.5 THE GAS TURBINE ... 38

6.5.1 The compressor (C2) ... 38

6.5.2 The combustor(R4) ..................... 39

6.5.3 The turbine (El) ...... 39

6.6 ALTERNATIVE OPERATION: NATURAL GAS FIRING OFTHE GAS TURBINE ... 39

7. PROCESS SAFETY, HEALTH AND ENVIRONMENT ...... 40

7.1 INTRODUCfION ... 40

7.2 SAFETy ... 40

7.3 HEALTH ... 42

7.4 ENVIRONMENT ... : ... 43

7.5 SAFETY FEATURES ... 43

7.6 HAZARD AND OPERABILITY STUDY ... 44

8. PLANT ECONOMICS ... 45

8.1 COST MODEL ... 45

8.2 THE INVESTMENT COSTS ... 47

8.2.1 The Taylor method .............. 47

8.2.2 The Zevnik-Buchanan method ................. ... 48

8.3 REVENUES AND PROFIT ... 50

8.4 PROFITABILITY ... 50

8.4.1 Return On Investment ... 50

8.4.2 Internal Rate of Return ....................... ...... 51

9. CONCLUSIONS AND RECOMMENDATIONS ... 52

REFERENCES ...... 53

-Appendices

I

Proces flowsheet

II List of components

III

Coal composition

IV

Equipment specifications

V

Equipment calculations

VI

Partic1e si ze distributions

VII

Mass and heat balance

VIII

Proces economics

IX

Power calculations

X

Streams and components

1. Introduction

An important part in the education of a chemical engineer at the Delft University of Technology is the preliminary design of a chemical plant. The subject is chosen in co-operation with the department of Industrial Catalysis of this same University.

The objective of the assignment is to design a 250 MW Integrated coal

,Gasification Combined Cycle (IGCC) power plant with high temperature sulphur removal. The sulphur is removed continuously by chemosorption in a rotating monolith reactor.

1.1 Coal and electricity generation

About one-third of the world's energy fuels are consumed in generating electricity; and this fraction represents a wide variety of fuels. Coal is the main fuel for electric power generation; almost 40 % of this electricity is generated from co al. As newly industrialized nations expand their energy demand, world coal use is projected to at least double in the next 30 years. Together with the inevitable depletion of the world's oil supply, this provides the incentive for development of power generating processes from non-petroleum resources: coal, natural gas, biomass and nuclear sources.

1.2 Integrated Gasification Combined cycle

Current production of electricity from coal is carried out by gasification of coal. Gasification is seen as one of the technologies which offer a favorable alternative to the conventional pulverized coal combustion process. In a one step combustion of coal the energy obtained is used to generate steam which is fed to steam turbines for production of electricity. The main difference between gasification and combustion is that the syngas produced in the gasifier by partial oxidation is used further in the process for a second combustion step in a gas turbine. For this second combustion step the syngas stream has to satisfy certain specification (like a maximum inlet temperature, maximum levels of corrosive compounds like H2S and a maximum number of particles of a certain

size). Surplus heat generated in the gasifier and heat left in the offgas of the gas turbine is used to generate steam for the use in steam turbines. This set-up is called an Integrated Gasification Combined Cycle.

By producing syngas and bringing it to the operating requirements for the gas turbine the energy generated by coal gasification can be used more efficiently than in one step combustion of coal (gas turbines have a higher total efficiency than steam turbines because they can directly use the energy of the gas stream instead of having to use water as a energy transmission medium). Another advantage is that the gasification process needs only 20-30% of the oxygen required for complete combustion to carbon dioxide and water. Instead of

FVO 3198 _Introduction

oxygen steam is used to provide part of the needed oxygen in the gasifier. For the reasons mentioned above new coal fired electricity plants have a IGCC set-up instead of the traditional one step combustion set-set-up. The IGCC is presented in the following scheme:

Sleam Oxygen

Coal Pretreat- _{G asifier} ment Ash BFW Oxygen Desulphur-isation Sulphur Waste heat boiler Gas turbine Steam turbine

Figure 1.1 Scheme of integ rated gasification combined cycle

1.3 H

2

S corrosion

Elec.

Electricity

A major problem in the IGCC process is the formation of H2S in the outlet gas

stream of the gasifier. The H2S will corrode the blades of the (expensive)

turbine and therefore has to be removed from the gas stream. Usually this is done by cold (about 500 K) scrubbing. The H2S is fed to a CLAUS plant where

it is converted to sulphur. The offgas of the plant can be treated in a SCOTT plant to further boost the conversion to sulphur. The disadvantage of cold scrubbing in respect to the overall plant efficiency is that it operates at a relative low temperature. Af ter the cleaning the gas stream has to be heated again for combustion in the gas turbine. High temperature sulphur removal is therefore a logical step to increase the plant efficiency.

1.4 High temperature sulphur removal

There are several options for high temperature sulphur removal. We have chosen for the set-up that is researched at the Delft University of Technology (Department Industrial Catalysis of the Faculty Chemical Engineering): a rotating monolith reactor consisting of twelve separated tube reactors. The monolith rotates around the central axis. The gasifier offgas stream flows through one part of the monolith. In this section H2S reacts with the

impregnated MnO/'y-alumina catalyst. The H₂S concentration in the monolith off gas is low enough to not significantly damage the gas turbine. In another section of the monolith the catalyst is regenerated. There are two methods of regeneration: with steam giving a concentrated H2S stream which can be fed to

the CLAUS plant or with excess S02. The latter leads to direct formation of disulphur. The sulphur can be separated from the S02 by a cold trap which liquefies the sulphur. The excess S02 is recycled with a part of the sulphur to a sulphur combustor, where it reacts with O2 to S02 for the regeneration of the catalyst.

FVO 3198 _{Process Input Information}

2. Process Input Information

The preliminary design of the integrated coal gasification combined cycle power plant is based on the existing IGCC plant in Buggenum (The Netherlands). To design the plant according to the required specifications several decisions have been made. The total production of electricity is 250 MW. To meet this objective the coal feed will be adapted. Starting assumption is that 1900 ton day-l coal is needed (this is used by the plant in Buggenum). The process is continuous throughout the who Ie plant. The plant is on stream 90% of the year. This equals 7884 production hours a year. The rest of the time the plant is down for maintenance, replacement of equipment and other process interruptions.

2.1 Battery limits and external specifications

Figure 2.1 gives a black box representation of the entire plant with all the input

and output streams:

C 1 oa _Sulphur

02 Stack gas

IGCC

Wet ash

Steam

Slag

N2 Fly Ash

I

Electricity

Figure 2.1 Blackbox ofthe IGCC-Power plant

The design uses several streams from outside the battery limit. The conditions of these streams are taken into consideration in this chapter. Table 2.2 shows the feed and product streams and utilities used in the plant:

1 bar, 314 K mean d: 2 cm flyash 1 bar, 612 K, 2<d<9 Ilm

The plant produces its own electricity (for production purposes) and steam. The oxygen and nitrogen are obtained at the right specifications. Polluted product streams are to be delivered at the right specifications for end of pipe treatment.

2.2 Selection of process route

The process consists of four sections: pre-treatment, gasification, desulfurization and electricity generation (see figure 1.1).

2.2.1 Pre-treatment

In the pre-treatment section the coal is crushed and pulverized into small particles (90 % < 74 Jlm) to achieve a high conversion in the gasifier and dried with hot nitrogen gas. The drying is because of the variation of the coal's water content. Too much water leads to unwanted products and temperature deviations in the gasifier. The co al is assumed free of dust. The type of coal used in this design is the same as used in the Buggenum power plant and is called Drayton Coal.

The Drayton Coal is assumed to have the following composition (see appendix III [4, P 6-49]): 4.0 1.5 1.0 5.0 1.5 1.2 4.5

The minerals are assumed to consist of Si02, Fe203, CaO and S03' The minerals except S03 are inert and do not react in the gasifier. The particle size distribution of the coal can be found in appendix VI and is based on a norm al distribution function with a mean of 25 cm and a standard deviation of 8.5 cm.

2.2.2 Gasifïcation

In the gasifier the coal is converted with steam and pure oxygen (99.9%) into CO, H2' H2S, CH4 , COS, NH3 and slag. The slag is formed from the minerals

FVO 3198 Process Input In{ormation

Si02, Fe203 and CaO. The gasifier operates autothermic. The main products are formed by the following basic reactions:

2C(s) + 02(g)

_<=>

2CO(g) (2.1) C(s) + H2O(g)

<=>

CO(g) + H2(g) (2.2) S(s) + H2(g)

<=>

H2S(g) (2.3) C(s) + _2H2(g)

<=>

CH4(g) (2.4) .1Hr (2.1)

=

-222.0 kJ morl .1Hr (2.2)

=

+175.44 kJ morl .1Hr (2.3)

=

-20.63 kj morl .1Hr (2.4)

=

-74.90 kJ morl

The gasification reactor is an entrained flow reactor with an internal heat exchanger to generate saturated steam from the hot gas stream (Note: for simulation purposes this heat exchanger was used to superheat the saturated steam). The reactor is assumed to be a equilibrium reactor based on minimization of gibbs energy. The gasification section has the following constraints: the carbon conversion is higher than 99.9 mol% and the water in the outgoing syngas stream is below 0.2 wt%.

2.2.3 Desulfurization

The H2S and COS are removed from the gas stream by means of a reaction with a MnO/y-alumina catalyst in a rotating monolith reactor. In this reactor the H2S is converted into H20:

+

(2.5)

.1Hr (2.5) = -10.0 kj marl

The regeneration of the deactivated catalyst occurs in the same reactor by reaction with S02:

.1Hr (2.6) = +31.9 kj mOrl

The disulphur is liquefied in the cold-trap and a fraction is recycled to produce S02 in the sulphur combustor:

+ (2.7)

.1Hr (2.7) = - 296.83 kj marl

The main specification for the monolith reactor is a H2S concentration of 10 ppm in the cleaned gas stream.

2.2.4 Electricity genera ti on

The electricity is generated in a gas turbine and a steam turbine section. The desulfurized syngas is fed to the gas turbine where it is combusted with a compressed 021N2 stream to provide a stream of hot, high pressure gas. The gas expands and conducts work on the turbine blades to turn the shaft which drives both the compressor and the generator. In the combustion chamber the following reactions occur:

CO(g)

+

H2(g)

+

ÓHr (2.8) = -282.98 kJ morl ÓHr (2.9) = - 241.82 kj morl Y202(g)

<=>

Y202(g)

<=>

COzCg) H₂0(g) (2.8) (2.9)

The hot exhaust gas stream is sent to a waste heat section, to generate high and medium pressure steam by heat exchange with the boiler feed water. The two-circuit set-up is chosen to make maximum use of the generated heat. The steam section is a closed circuit.

2.3 Location of the plant

The plant is located in Buggenum (The Netherlands) at the borders of the river Maas. Raw materials (i.e. the coal) are easily supplied by shipping and large storage fields for coal are present. The produced disulphur can also be shipped and transported to the Rotterdam Harbor. Cool water is obtained from the river and the waste heat from the condenser can be disposed into this river.

Furthermore, safety measures are at a high level, well educated personnel and staff is available.

2.4 List of llsed chemicals

A list of all chemicals used in the plant with all significant properties and data is presented in appendix Il.

FVO 3198 _{Process structure and process flowsheet}

3. Process structure and process flowsheet

3.1 1ntroduction

In the previous chapter the boundaries and main decisions for the process have been made; in this chapter a closer look at each unit will be made. Motivations of the process conditions and structure, choice of the equipment and a schematic overview of the plant in form of a process flowsheet will be handled. In the next chapter the chosen equipment will be designed. Numbers of streams and equipment are si mil ar to those used in the flowsheet. This flowsheet can be found in appendix I and at the end ofthis chapter (figure 3.1).

3.2 Pre-treatment

Main requirements for the pretreatment are:

• coal particles must be pulverized to 90% < 74 Jlm (-200 mesh) [2, p 9-22] • the coal must be dried to 1 wt% water in the product stream.

3.2.1 The crusher

To achieve a high conversion of coal in the gasifier (Rl) it is necessary to feed the gasifier (Rl) with coal particles of which at least 90 % passes a -200 mesh sieve [2, p.9-22]. The raw, run of mine (ROM) coal is assumed to have a maximum size of up to 50 cm [1]. The particles are assumed to be distributed according to a norm al distribution function with a mean of 25 cm and a standard deviation of 8.5 cm (see appendix VI). A crusher is used to grind the raw coal into smaller particles. The ROM co al must be reduced to particles smaller than 5 cm, for this is the input requirement for the chosen pulverizer (M2) [3]. For the size distribution of the crushed particles see appendix VI. The distributions are calculated with CHEMCAD 3.20. The chosen crusher is a double-roll crusher of the Jeffrey Mamifacturing Company [1]. It was found [1] that this apparatus has the required throughput and product size.

3.2.2 The pulverizer

The coal obtained from the double-roll crusher (MI) has to be pulverized to 90% of the particles smaller than 74 Jlm (-200 mesh) as required for the gasification reactor (Rl). The in- and output particle size distributions as calculated by CHEMCAD 3.20 can be found in appendix VI. The coal particles in the feed stream are smaller than 5 cm [3] (see Crusher above). The chosen pulverizer is a Babcock & Wilcox Co. Type E Pulverizer. It was found [3] that this apparatus has the required throughput and product size. The pulverizer can

accept a 25 per cent moisture content, which is weIl above the 9.5% moisture content of Drayton Coal.

3.2.3 The dryer

Drayton Coal has a moisture content of approximately 9.5 wt% [4]. To control the exact amount of water in the gasifier (Rl) it is necessary to reduce the water content in the coal to a minimum (1.0 wt% of the total flow). It was found that if the water content in the input coal is allowed to deviate 50 wt%, the weight percentage water in the dried coal must not be higher than 1 wt% because of the problems this will present in the gasification (undesired temperature changes and undesired products). If the coal is dried to less than 1.0 wt% water the required energy for drying will be much higher. Nitrogen gas is used to dry the coal. Nitrogen is chosen because of the availability on-site (oxygen is also delivered by an air separation plant) and the inherent safer operation compared to air drying (with small partic1es and oxygen th ere is the danger of dust explosions). The gas is heated from 300 K to 450 K in a heat exchanger (H5). The nitrogen which is considered to be at 20 bar and 300 K (same conditions as the oxygen feed) is heated in a methane fueled heater to 450 K. The needed nitrogen to achieve a 1 wt% water content in the exit coal was calculated with

CHEMCAD 3.20 : 6176 ton N2 dai 1

• To heat this gas to 450 K 19.8159 ton methane dai1 has to be combusted. This equals 23590 m3 methane dai1 (density 0.84 kg m-3 [9]). The cost of methane is taken as $ 0.15 m-3. Yearly

cost of methane are thus estimated as: 0.9 [on-line fraction]

*

365 [days year-1]

*

23590 [m3 dai1]

*

0.15 [$ m-3]= $ 1,162,397 year-1. Based on this cost it is assumed to be cheaper to heat exchange the needed nitrogen with the exit gas stream (STR.49). The chosen dryer is single-stage pneumatic-conveyor dryer

(Raymond Division, Combustion Engineering Inc.) [2, p.20-51]. This apparatus is suitable for the large required throughput and can dry the coal to the required moisture content, without losing too much particles (because of the cyclone which is included in the dryer). Transport of the small coal partic1es to the gasifier is done by conveying though pipes with nitrogen gas.

3.2.4 The baghouse

The nitrogen / water stream (STR.5) from the dryer (M3) contains too much partic1es and needs further c1eaning. Because of the small average partic1e size <80 ~m c1eaning with cyclones is difficult and thus expensive. A relatively easy solution is cleaning by means of a bag filter. Collected dust can be fed to the outgoing stream of the dryer. The chosen baghouse is a shaker type fabric filter from Wheelabrator-Frye Inc. [2, p 20-99]. A shaker cleaned filter was chosen for its simple operation compared to other options like pulse flow and reversed flow c1eaned filters.

3.3 Gasification

The main requirements for the gasification section are: • more than 99.9 mol% conversion of carbon

FVO 3198 _{Process structure and process flowsheet}

• the gasifier exit temperature must be around 1123 K (for optimal H₂S removal in the monolith)

3.3.1 The gasifier

The Demkolec plant in Buggenum makes use of the Shell coal gasification technology. Based on this a Shell-koppers entrained flow was chosen. The choice for this entrained flow reactor also allows for relatively simple modeling of the gasification section. An extra advantage of this reactor compared to the moving and fluidized bed reactors is the possibility of handling all types of coal. Because of the high temperature high pressure reaction (no activation energy barriers) the reactor can be modeled as a minimal Gibbs energy reactor. The reaction is assumed to take place at 1800 K and 28 bar [2, p 9-21]. These high temperatures and pressure are chosen because at these conditions only simple molecules are formed [4, p62] (H2S, COS are the only problem

components in the exit stream). High pressure also leads to faster reaction rates because of the higher density. Furthermore high pressures are favorable in the desulfurization section because absorption occurs best at high pressures [4, p64]. In the next section of the reactor the products are cooled to the exit specification of 1123 K and 28 bar. In this section components are assumed to be at their exit concentration. The temperature and reaction profile in the modeled reactor is given in figure 3.2.

1800 K reactions 2.1 to 2.4 Gasifier burners

o reactions

Cooling section

1123 K no reaction product stream from gasifier

dj ~ 330 K no reactions <= à:i reactants to gasitïer 0. >= v E-Main stream

•

Figure 3.2 : Temperatllre profile in gasifier

The ingoing steam rate (reactant) must be high enough to keep the temperature level of the reactor on the desired level, supply oxygen for the reaction (maintaining a carbon conversion of >99.9 mol %) and low enough to suppress the formation of undesired reaction products like methane, carbon dioxide and water.

3.3.2 The slag quencher

The hot (1800 K) molten slag is assumed to exit the bottom of the gasifier in large chunks (30 cm). The slag flows into a vessel with flowing cooling water. Because of the glassy state of the slag the material doesn't exhibit leaching. The input cooling water doesn't have to be clean and the output water can be fed back to the environment as long as its below the aIlowed 314 K.

3.3.3 The slag crusher

The slag can be used for construction purposes (i.e. as landfiIl). The assumed product specification for molten slag is an average si ze of 2 cm. To meet this requirement the exit slag from the quencher has to be crushed.

A single-roIl crusher is adequate for the desired crushing operation. The chosen crusher is a Jeffrey Manufacturing Company Single-roIl Crusher.

3.4 Desulfurization

The mainrequirements for the desulfurization section are: • less than 10 ppm H2S in the exit stream

• the produced S2 must be 99.99 % pure

3.4.1 The monolith reactor

To achieve the high temperature removal of H2S (and COS) a rotating monolith

reactor is used. The use of this technology has several advantages over the traditional colq scrubbing sulphur removal. The temperature in the monolith reactor is around 1123 K which means no intensive cooling or heating of the gas stream is necessary. The rotating set-up of the reactor allows for continuous operation (the monolith is self-Iubricating) and simultaneous production of S2.

The bottleneck for the H2S chemisorption is assumed to be the diffusion of the

molecules in the center of the gas stream to the reaction surface. Because of the experimental nature of the rotating monolith and to keep the process flexible the reactor is over-designed (four times the minimal required length is taken). For regeneration of the reactor there are two options:

• with steam leading to production of H2S

• with S02 leading to production of S2'

The latter variant was chosen because of the simple process (H2S must be fed to

a CLAUS plant to convert it to sulphur). The formed sulphur is 99.99 % pure S2

as a consequence of thermodynamic circumstances in the regenerator. Because the conversion in the regeneration section is low, an excess amount (90%) of

S02 is used. The pressure drop is estimated at 0.5 bar.

3.4.2 The cold trap

The S2 leaves the regeneration section of the monolith reactor together with the excess S02. The different boiling points (S2: 717 K, S02: 263 K) aIlow for separation of the two components by cooling the stream to 600 K in a adapted

FVO 3198 _{Process structure and process flowsheet}

liquid-gas vessel. The cooling water can be used to provide a large part of the input steam (reactant) needed for the gasifier.

The separation vessel has a disengagement height equal to the vessel diameter. The liquid level in the vessel depends on the hold-up time required for smooth operation. A typical value for the hold-up time is 10 minutes.

3.4.3 The sulphur combustor

To provide the needed 502 for the regeneration section of the monolith part of

the liquid 502 out of the cold trap is fed to a combustor where it is completely

burned. The combustion of the liquefied disulphur into sulphur dioxide is done in a combustion chamber in which the sulphur is being sprayed in a nozzle and burned with pure oxygen. The production of 502 from 52 is chosen instead of

obtaining 502 from outside the battery limit because it is estimated to be less

expensive and because the combustion reaction also provides the necessary heat to keep the desulfurization section on temperature. The excess 50₂ is used to cool the gas stream in the combustor until both have a temperature of 1123 K. This is done by adding 502 in small amounts. This requires a certain hold-up

time in the combustor. The chosen hold-up time is 10 minutes.

3.5 EZectricity generation

The main requirements for the electricity generation section are: • net production of 250 MW of electricity

• delivering the main exit streams, stack gas and cooling water, at the desired specifications.

3.5.1 The 5chumacher filter

The gas stream into the gas turbine may not contain particles larger than 2 )lm.

The gas stream af ter the desulfurization step contains coal particles with a diameter between 2 and 9 )lm; therefore the gas has to be cleaned. This cleaning is done by a Schumacher Candle Filter, DIA-Schlllnlith T10-20.

3.5.2 Combined cycle

The electricity is generated in a combined cycle, which comprises a gasturbine and a steam cycle. The power plant is equipped with a Siemens V94.3 gas turbine because of its high thermal efficiency (36.5 %). The gas turbine exists of a compressor, a combustion chamber and the actual turbine, where the compressor and turbine are working on the same shaft, so the pov.:er needed for the compression is generated by the turbine. The 02/N2 stream from the air

separation unit entering the compressor is further compressed to 27,5 bar. Between the compressor and the combustion chamber approximately 20% (mass) of the stream is extracted and fed to the air separation plant. This prevents an unusual high difference between the compressor and the turbine mass flow, which could lead to an increased pressure ratio. The restrictions for the gas turbine are the turbine inlet temperature, which is limited to approximately 1600 K, and the maximum NO formation of 95 g/GJ. Also the

particles entering the turbine may not exceed the si ze of 2 ~m, but this is being taken care of by the Schumacher filter.

NOx can be produced from two sources :

1. Fuel-bound nitrogen : This type of nitrogen is completely transformed into NOx, but the gas which has been subjected to the purification, has been

removed practically of all of the nitrogen components like NH₃ and HCN, so the fuel bound nitrogen will be virtually zero.

2. High flame temperature : Part of the combusted N₂ will be transformed into so-called thermal NOx'

The amount of therm al NOx is primarily a function of the flame temperature,

while the adiabatic (stochiometric) flame is determined by the fuel gas composition. As a result of its composition (60% CO and 30% H2) coal gas

tends to bum at higher temperatures than natural gas. This would result in twice to three times higher NOx production. To prevent this, the syngas is diluted with

nitrogen and water (vapor), which lowers the flame temperature and thus NOx

formation. By this measure the turbine inlet temperature is also on an acceptable level. The gas turbine is designed to handle a flow of 624 kgs-I.

3.5.3 The steam cycle

The steam cycle consists of four heat exchangers, two flashers and three turbines. The water of the cycle is heated in the h~at exchangers (H2, H3, H4) to its boiling point (675 K at 300 bar). Then the stream is flashed, producing a liquid and a vapor stream. The steam from flasher 1 is high pressure steam (601K at 125 bar) and before it enters the steam turbine, it is superheated (707 Kat 125 bar). The steam from flasher 2 is medium pressure steam (523 Kat 40 bar) and is mixed with the effluent of steam turbine 1 before it enters steam turbine 2. The stream leaving steam turbine 2 (443 K at 8 bar) is directed to steam turbine 3 and expands to 3 bar. The steam is then completely condensed in a condenser (H6), and pumped back to the other heat exchangers. The critical parameter in the steam cycle is the liquid fraction of the stream leaving the turbines. If this fraction is too high (> 15%), the turbine is likely to break down. Only in steam turbine 2 and 3 there is liquid present, but their fractions (resp. 4,8% and 8,5%) are lower than the critical value.

3.5.4 The heat exchangers

To cool or heat the process streams to the desired temperature, heat exchangers are used. For a good performance of the heat exchangers the temperature difference between both streams must at least be 15°C.

The chosen heat exchangers are of the shell and tube modeis, because they will be exposed to extreme conditions, large streams with high temperature and pressures, which excludes the other modeis. Furthermore they have the following characteristics:

FVO 3198 Process structure and process flowsheet • They have a good mechanicallayout, a good shape for pressure operations • Easily cleaned

• Uses well-established fabrication techniques • Can be constructed from a wide range of materials • Well-established design features

In the design different types of shell and tube types are used. Heat exchanger Hl is a membrane heat exchanger and is integrated with the gasifier. For heat exchanger H2 a floating head model with packing is used because they operate counter-currently and can easily be cleaned. The disadvantage of these types is that they are expensive. Exchanger H4 uses a fixed tube sheet because it has to operate counter-currently. The other exchangers (H3, H5 and H6) are simple U-tube models because of the larger temperature difference. The advantage of the last two models is that they are relatively cheap. The disadvantage, that the shell for H4 and the tube side for the U-tubes are difficult to clean, is not a major problem because the process streams are relatively clean.

3.5.5 The pumps

For the transporting of the fluids in the process, centrifugal pumps are chosen. These are available in a wide range of sizes, capacities and performances up to 480 bar. The main advantages of the centrifugal pumps are their simplicity, small size, low cost, uniform (non-pulsating) flow and low maintenance expense. Because the head required is higher than can be generated by a single impeller, a multistage centrifugal pump is used. All impellers are in series and the total head is the sum of the heads of the individu al impellers. The pump is of the diffuser type which have vertically split barrel-type exterior casings with inner casings containing diffusers. The efficiencies of the pumps can be derived from figure 10.62 [6, p428], and are all in the range of 75- 85%. Ta be on the safe side an efficiency of 80% is chosen sa the pumps have a safety margin.

f

_t

_{~---i._---___,} coolwater ~ M3 ST

~,-M3 coolwoler I

-~---

make-up water ....

TI"

nitrogcn/oxygen ~

M 3 DRYER P 3 PUMP - moln streom _{Process flow scheme of a 250 MW}

C 1 COMPRESSOR

M 4 BAGHOUSE FILTER R 1 GASIFIER ~ steam/water circuit _{powerplant with continuous}

W~ ~~ÄtU~~g'JANGER M 6 COLD TRAP R 2 MONOUTHIC REACTOR _---- - ₌ regeneration cycle _{high temperoture desulphurization}

H 3 HEA T EXCHANGER _{M 6} CRUSHER

R 3 SULPHUR COMBUSTOR D nltrogen

H 4 HEA T EXCHANGER

M 7 SCHUMACHER - - - _{M. Blokland} B.W. Hoffer FVO nr. 3198

FVO 3198 Equipment design

4. Equipment design

4.1 Belt conveyer

Transportation of the coal to the crusher is done by a belt conveyer. It is assumed that the distance between the belt loading point and the belt discarding point (the crusher) is 20 m. The needed power for this transportation belt is estimated using table 7-7 [1, p.7 -10]. For transportation of 98 ton h-I along 30 meters 760 Watt is needed. Assuming a proportional relation between power, loading and distance about 470 Watt will be needed for transportation of 92 ton h-I _{(2200 ton dai ') along 20 meters.}

4.2 The crusher (MI)

The chosen crusher is a double-roll crusher of the Jeffrey Manufacturing

Company [1] (figure 4.1 [1]). The specifications of the apparatus were determined with table 7-2 [1, p.7-15]. The maximum si ze of the feed is taken as 50 cm. Product size must be about 3.8 cm (the lowest value in the tabie; to meet the required input of the pulverizer (M2) ). The coal throughput to meet the designed electricity production is 2200 ton dai '. This equals 92 ton h-I. The

specifications for this crusher are:

Roll size diameter x width :76 x 122

Speed of rolls : 115

[cm x cm] [RPM]

The required energy to crush 85.5 wt% of the coal was calcu1ated in

CHEi\1CAD 3.20. It is assumed that the rest of the material requires a

proportional amount of energy.

Required crushing energy per kg coal: 1.41 kJ kg,l.

Required crushing energy per day for 2200 ton coal per day : 3103 MJ dai'

=

0.03592 MW.

4.3 Belt conveyer

To transport the coal from the crusher (MI) to the pulverizer a belt conveyer is used. A distance of 10 meters between the installations is assumed. The needed power is 270 Watt. Calculation similar as in section 4.1.

4.4 The pulverizer (M2)

The chosen pulverizer is a Babcock & Wilcox Co. Type E Pulverizer (figure 4.2 [2]). The B& W pulverizer is a ring-roller mill and consists of a single row of balls operating between a rotating bottom ring and a stationary top ring. Externally adjusted spring apply pressure to the top ring to give the required loading for proper pulverizing. The mills, with their spiraling balI movement, are high suitable for grinding coal. The required throughput of 2200 ton day"1 (92 ton h-I) is possible for this machine.

The dimensions of the pulverizer are estimated according to the following assumptions:

• the throughput is 92 ton h-I equals 1528 kg min-I;

• the residence time for one kg coal (to be grinded to 90% < -200 mesh) is assumed to be 3 minutes;

• the density of the coal is 1345.68 kg m-3 (bituminous coal [2, p.3-96]); • the pulverizing chamber has a cylindrical shape;

• the diameter (D) / height (h) ratio of this chamber equals 7 (deducted from figure 4.2 [2]);

• 1/3 of the chamber is filled with coal;

• the size ratios of the chosen pulverizer are equal to the size ratios in figure 4.2 [2];

• the pulverizer dimension in the figure are 6 x 6 x 6 cm. Calculations:

Coal in the chamber: Volume of the coal: Volume of the chamber: Dimensions of the chamber: Size ratio: 1320 [kg min-I]

*

3 [min]

=

3960 [kg] / 1345 [kg m-3] = 2.95 [m3]

*

3 [-] = 10.22 = 0.25rcD2h; D = 7h; h = 0.64 mand D = 4.5 m; 4.5 m equals 3.5 cm in the picture;

1 cm == 1.286 m;

4583 ka' _{0 '} 3.41 m3;

10.22 m3;

Dimensions: 6 x 6 x 6 [cm] equals 7.7 x 7.7 x 7.7 m. The required energy to pulverize 85.5 wt% of the coal was calculated in

CHEMCAD 3.20. It is assumed that the rest of the material requires a proportional amount of energy.

FVO 3198 _{Equipment design}

Required pulverizing energy per kg co al : 49.65 kJ kg-I.

Required pulverizing energy per day for 2200 ton coal per day : 109222 MJ day-I = l.264 MW. .

I

,Rofaling classif~,

,

Figure 4.2: B& W Pulverizer type E. [2, p 8-43

J

4.5 Tlze Dryer (M3)

The chosen dryer is a single-stage pneumatic-conveyor dryer (Raymond Division, COl71hustiorz Engineering lnc:.) [2, p.20-SI]. A picture of this dryer is shown in figure 4.3 [2]. This figure illustrates a single-stage dryer employing a long tube carrying the gas at high velocity, a fan to propel the gas, acyclone collector, a paddie mixer, recycle, and a CE-Raymond cage mill for fine grinding and dispersion of the mixed feed in the nitrogen stream.

The significant velocity effect in the pneumatic conveyor is the difference in velocities between gas and solids, which is the reason why a major part of the total drying actually occurs in the feed section. By assuming a duct diameter of 0.3 mand a gas velocity of 23 rnJs, if the solids velocity is taken as percent of this speed, the velocity difference between the two is 4.6 mis. The evaporation capacity, based on the caJculation in CHEMCAD 3.20 and a water reduction of 90 wt%, is 7838 k~ h- ' . A dryer capable to handle this capacity has the following dimensions (see tïgure 4.3):

UI Flash-Drying Systems A 5.49 B 11.58 C 6.10 D 23.16 E 5.94 F 8.23 G 6.86 H 3.05 clone 5.49

1

J

I

i

I

~!---~Hr~~--~~~----JU~~L---. I i

I

I !

al

~i

I

!

'

I

t

(Mixer

support I~Vfl

I

L---lI~32~~~rLi

Oil or gos burner

I

₁

I <li

I

.

I I

Figure 4.3: Single-stage pneumatic conveyor dryer [2, p 20-53)

4.5.1 The cyclone

The diameter of the cyc10ne (5A9 m) is the Dr (figure 4A [5]). The dimensions are based on the design of the high efficiency cyc10ne [5, pA02]:

FVO 3198 Equipment design

2.745 2.745 8.235 21.960 Dust out let diameter 2.059 No. of turns 5.000

The cyclone efficiency is calculated in CHEMCAD 3.20, and is 0.7101, based on particle size distributions (appendix VI) and the given cyclone dimensions. The pressure drop over the cyclone is estimated according to the following equation:

P'r { ,[

,(2r

J]

,}

M

=

203 ll~ 1

+

2<P- r~' -1

+

2u

z

(4.1)

where:Pr

=

gas density [kg m-3]

UI

=

inlet duct velocity [m S-I]

u,

₌

exit duct velocity [m S·I]

rt

=

radius of circle to which the center line of the in1et is

tangential [m]

re

=

radius of the exit pipe [m]

<l>

=

cyclone pressure drop factor [-] ([S]).

The pressure calculated drop is: 1.36 kPa

CHEMCAD 3.20 calculated with the Koch and Litch method: 1.30 kPa 0'5C\: 0'5DcX°' \ Collectinq I"Ocoer dlometer De

~ ~

~~

"0'375 De

4.6 The Baghouse (M4)

This fabric filter is a collector in which dust is removed from the gas stream by passing the dust-Iaden gas through a fabric [2, p.20-97]. The chosen baghouse is a shaken-cleaned filter. The open lower ends of the bags are fastened over openings in the tube sheet that separates the lower dirty-gas inlet chamber from the upper clean-gas chamber. The bag supports from which the bags are suspended are connected with a shaking mechanism. The dirty gas flows upward into the filter bags, and dust collects on the inside surfaces of the bags. When the pressure drop rises to a chosen upper limit as the result of dust accumulation, the gas flow is stopped and the shaker is operated, giving a whipping motion to the bags. The dislodged dust falls into the dust hopper located below the tube sheet and is fed to the output of the dryer. The filter is constructed with multiple compartments, so that the individual compartments can be sequentially taken off line for cleaning while the other compartments continue in operation (figure 4.5):

10 fan

(3) (4)

Figure 4.5: Baghouse shaking mechanism [2, p 20-100J

The chosen baghouse is a shaker type fabric filter from Wheelabrator-Frye Inc.

(figure 4.6). The pressure drop across the fabric and the collected dust layer may be expressed by:

where:Óp Kç

=

=

pressure drop [kP a]

cloth-resistance coefficient [lIm]

FVO 3198

=

=

=

=

gas viscosity [Pa.s]

superficial velocity through filter [mis]

dust-Iayer resistance coefficient [mlkg] dust loading on filter [kg/m2]

Equipment desi gn

The resistance factors are ca1culated in CHEMCAD 3.20. The bag filter specifications are:

T bI 43 B h a e

. .

.

agJ ouse specl Ica IOns 'ti f

Cloth resistance factor 0.87 [l/m]

Dust resistance factor 0.10 [mIkg]

Number of cells 2.00 [-] Cells cleaned 1.00 [-] Number of bags/cell 78.00 [-] B a~ diameter 0.15 [m] Filter arealba~ 1.49 [m-] 1 Efficiency 0.98 [-]

Pressure drop 0.53 [bar]

Floor space 3.62 [m2]

Gas velocity 0.04 [mis]

The filter is shaken every 20 minutes.

4.7 The Gasifier (Rl)

The gasifier is the heart of the combined cycle power plant. In this reactor occurs the partial oxidation of coal. The chosen gasifier is a Shell-Koppers Entrained Flow Reactor (see figure 4.7). The gasifier consists of an outer

pressure vessel and an inner gasification chamber with a water-cooled membrane wall [6]. The inner gasifier wall temperature is controlled by circulating water through the membrane wall to generate saturated steam. The membrane wall encloses the gasification zone, from which two outlets are provided. One opening at the bottom of the gasifier is used for the removal of slag; the other outlet allows hot raw gas and fly slag to exit from the top of the gasifier. The entrained-flow gasifier consists of a plug-flow system in which the fine coal particles concurrently react with steam and oxygen. The residence time is a few seconds.

Calculation of the volume of the gasification chamber: Total throughput:

The throughput is based on volume flow of the syngas stream (from

CHEMCAD 3.20). This is the maximum amount of volume which has to be in

the reactor. The residence time is assumed to be 5 seconds, so the volume of the gasification chamber is 248 m3. According to figure 4.7 has the reactor a

dimension ratio of din/h in = 2. The shape of the reactor is cylindrical, so the resulting dimensions are:

din

=

8.60 m hin = 4.30 m

As said, the generated heat is used to generate steam from water. The exchange of heat occurs inside the gasifier by a membrane wall. The mem bra ne has to be resistant to a pressure of 28 bar and a temperature of 1800 K. The syngas flow stream has to be cooled down from 1800 K to 1123 K. This is done with the heat exchanger by converting water into saturated steam. The amount of water is 48.23 kg/s and 216055 m3/h. The water flows around the cylinder with the dimensions dout-din and hout-hin' Knowing the volumetrie flow and the shape the gasifier has the following dimensions:

dout = 9.25 m hout = 4.60 m

The slag flows through a compartment which is about 5.0 m long (compared to the gasification chamber) and the crud gas output tube is about 4.0 m high. The total height of the reactor is thus: 5.0 + 4.6 + 4.0

=

13.6 m. The two inlets of the gasifier contain a mixing point of the steamloxygen stream and the coal feed.

FVO 3198 _{Equipment design}

pressure steamloxygen stream. This mixing section is about two times 5.0 m. The total width of the reactor is thus: 5.0

+

9.25

+

5.0 = 19.25 m.

Crvd qos

t

Figllre 4.7: Entrained-flow reactor [2. p 9-21

J

The particles in the gasifier are burned according to the shrinking core model [7]. This theory states the following relationship between conversion and radius of aparticle: where:ç r R = = = degree of conversion

particle radius after gasification particle radius before gasification

(4.3)

Therefore the partiele size distribution has changed (see appendix VI).

4.8 The Slag Ollenclzer (Vi)

To quench the slag flow of 241 ton day-I (equals 10 ton h-I) an amount of 270 ton da{1 water (T

=

287 K) is needed. The water from the river flows in a bath. The water flows into the top of the tank and leaves the tank (at 314 K) at the bottom of the top with a sieve (> 1 cm). The slag particles are removed by means of a belt conveyor and led into the crusher. The residence time of water and slag is assumed to be 5 minutes, so there has to be a tank volume of (270*103*5/(24*60))/998

+

(241*103*5/(24*60))/4000= 1.16 m3. Assuming the

dimensions as I x w x h, with w = 1

=

0.9h the tank has the following dimensions (the cubic 1 x w x O.lh is empty):

Table 4.4: Tank dimensions

4.9 The Slag Crusher

(MS)

The slag that has been cooled down by the water quench has asolid shape and has to be grinded by the crusher into partic1es with a smaller size. The partic1es can be sold as a product. It is assumed that the partic1es from the quench have a mean size of 30 cm and have to be crushed to partic1es with a mean size of 2 cm. The total slag flow is 241 ton dail and 10 ton h-I. To handle this capacity a

single -ro11 crusher can be used. The chosen crusher is a Jeffrey Manufacturing Company Single-roll Crusher.

The specifications for this crusher are:

RoU size diameter x width :45 x 45 [cm x cm] With this device 99.9 wt% of the slag partic1es have a diameter of 2 cm.

Required crushing energy per day for crushing the slag in a plant capacity of 2200 ton coal per day: 22.47 MJ dai1 == 0.2601 kW.

4_10 The Deslllfllrization Reactor (R2)

Because of the corrosive nature H2S and COS have to be removed from the co al

gas before it enters the gas turbine. For the desulfurization a rotating monolith reactor was chosen. The reactor consist of 16 parallel tubular reactors containing monolith catalyst [17,18, 19]. These 16 tubular reactors are divided into three sections:

• Sulphur uptake section • Purging section

• Regeneration section

In the sulphur uptake section the gas from the gasifier is desulfurized, in the purging section the coal gas is removed and in the regeneration section the sorbent is regenerated. The sorbent used during desulfurization is a manganese oxide on y-alumina monolith with a channel diameter of 2 cm. The gas used in the purge section is nitrogen. Regeneration is do ne using sulphur dioxide yielding sulphur, which can be sold. The main concern for this reactor is effective sealing at the high working temperature. This is achieved by the use of alumina ceramics. These ceramics are to certain extent 'self-lubricating'. The whole reactor section is constructed out of this material. At both sides of the reactor the tubes are mounted in disks, which are pressed to two other disks by

FVO 3198 Equipment desi gn

adjustable springs on each side. The middle disks distribute the gas flows in the four sec ti ons (figure 4.8: disk 3). The outer disks contain the four gas connections of the reactor (figure 4.9: disk 1). This reactor is designed for counter current operation.

2

2 ₃

o

Figure 4.8: Constructing (~f rhe monolith reactor [19

J

Uptake section

clean gas

Regeneration section

Figure 4.9: Schematic viel\.' ofrhe mOl1olith reactor [17]

This type of reactor is only in its experimental phase and because of th at a lot of the data needed to calculate the size of the reactor is not available. That is why

only an estimation based on bench scale expèriments can be made. Bakker et al.

[17, 19] already made such an estimation for a similar plant. They used the following assumptions:

• Shell process for a 250 MW demonstration coal based power plant

• 0.5 mol % H2S in coal gas.

• The acceptor is 32 wt 90 manganese on y-alumina monolith.

• Maximum capacity for sulphur is 17 wt %.

• The rotation speed is 10 rotations per hour.

• Outlet concentration H2S is <40 ppm.

Bakker estimated that 2 m3 acceptor would be sufficient for this situation. In

this preliminary design the outlet concentration is taken 10 ppm just to be on the safe side when it comes to corrosion of the turbines. Assuming only 1/4 of

the acceptor was effectively used for H2S absorption, because the concentration

of hydrogen sulfide is not always stabie, 8 m3 acceptor will be needed in the

reactor. The estimation of the reactor dimension is based on keeping the average velocity in the reactor the same as in the laboratory experiments. Due to the high temperature and pressure the diffusion coefficient will decrease. The relation used to determine the difference between atmospheric diffusion and diffusion at high temperature and pressure is:

D ex;

FT

p (4.4)

~

T Po JWl123 1 -2 1 D

=

-·_·D = - · - · D =6.99·10 ·D = - · D (4.5) T,P T o P 0 293 28 0 0 14.3 0

The diffusion will be 14.3 times slower than at T = 293 K and P

=

1 bar. This

means that the residence time will have to increase 14.3 times to maintain the

clean up efficiency. Bakker et al used a monolith with a channel diameter of 2

mmo The monolith used here has a channel diameter of 2 cm. It will take the

molecules 10 times longer to reach the monolith surface, which means th at the residence time will have to increase 10 times. The total sealing factor for the residence time is 143. This means that the reactor length will increase with the

same amount. The length of the reactor is: 143

*

2 cm

=

2.86 m. The volume of

the acceptor is 8 m3. This is the volume of the tubes. The diameter of the tubes

can be calculated using the following equation:

D =

~

=

I

4·8 =047 (4.6)

Illhe ~~ ~ 18.TC .2.86 .

To calculate the reactor diameter the following empirical formula is used:

(16 )

Drellctor = -;-+ 1 . D tuhe = 2.87

The results of the calculations above are given in tab Ie 4.5.

Table 4.5: Dimensions of the monolithic reactor

Length

Diameter reactor Diameter single tube Rotation speed

st volume

2.86m

2.87 m

0.47 m

10 rotations per hour

8 m3

FV03198 _{Equipment design}

4.11 The cold trap (M 6)

An important parameter in sizing a vertical separation vessel, which the cold trap is, is the maximum design velocity (uv). This velocity can be estimated using the following equation.

(4.8)

Now the vessel area (A) can be calculated using equation 4.9:

(4.9)

From the vessel area the diameter can be calculated: D = 1.16 m. The dis engagement space (LD ) is taken equal to the vessel diameter.

The liquid hold-up time (tH) is 10 minutes, so the liquid volume (V_{L )} in the vessel is:

(4.10) This can be used to calculate the liquid depth (LL):

L

=

VL

=

₁₅₅

=

_1.48 L A 1.05

(4.11)

The totallength of the vessel is: L = LD + LL + 0.30 = 1.16 + 1.48 + 0.30 = 2.94

m. The 0.30 is the estimated height needed for the inlet and outlet stream.

4.12 The sulphur combustor (R3)

The main difficulty in designing the sulphur combustor is the fact that the sulphur enters the reactor as a liquid. One of the objects of this preliminary design was designing a more efficient process for hot gas clean up. To keep the costs low a simple solution must be found. To avoid the sulphur dioxide from being converted to sulphur trioxide two options are available:

• spray the sulphur into a vessel to form a mist of sulphur drop lets in oxygen and inject small amounts of sulphur dioxide until the outlet temperature is reached.

• bum the sulphur oxygen mixture in the reaction section of the combustor. Then transport it to the next section wh ere it is mixed with the sulphur dioxide. For this case the latter was chosen, because of the lower risk of formation of sulphur trioxide. To achieve a good heat transfer good mixing and a hold-up time is needed. The chosen hold-up time was 5 minutes, because the gas streams need

a certain amount of time to reach the outlet temperature of 1123 K and this cannot be too long to limit the size of the combustor. For the reaction section a

gas velocity of 20 mis is chosen. The reaction between sulphur and oxygen is

very exothermic and is assumed to take place instantaneously. The reaction section is a tubular reactor in which perfect mixing due to the high gas velocities needed for spraying the sulphur, is assumed.

The residence time in the relaxation section of the combustor is 5 minutes. The

gas flow is 0.25 m3/s, so this section has a volume of 3.7E-2 m3. The diameter of

this section can be ca1culated using the following formula:

D =

~

4 .

~

v =

~

4 . 0.25 = 0.13

v·n 20·n (4.12)

Now the length can be ca1culated: L = 3.0 m.

The same ca1culation is performed for the reaction section. Table 4.6 gives the results of this ca1culation.

Reaction Relaxation

3.75

*

10.2 362.68

4.13 The Solid Filter (M7)

3.00 7.73

0.13 7.73

The cleaning is done by a Schumacher Candle Filter, DIA-Schumlith T10-20.

Pulse pipes

Gas out

Gas in

Elements

Dust out

Figure 4.10: Schumacher filter [20}

In the Schumacher filter (figure 4.10) [20], the pressure vessel is divided in two by a tube sheet from which clusters of elements are suspended. The

FV03198 Equipment design Schumacher has a device which holds elements in place using individu al springs which are, in turn, secured by a single holding plate. A single venturi/pulse pipe combination is used to clean a group of filter elements. The purpose of the venturi is twofold: flIstly, to enhance cleaning pressure and, secondly, to entrain clean gas and thus increase the volume and temperature of the pulse gas. The number of candles is 800. The filter medium is granular and made of silicon carbide (SiC).

4.14 The heat exchangers

The general equation for heat transfer across a surface is :

where

Q

=

U

=

A =

b.Tlm =

Q

=

U ·A·/)'T. Im

heat transfer per unit time [W]

the overall heat transfer coefficient [W m-2 Oei] heat transfer area [m2]

logarithmic temperature difference [OC]

(4.13)

The prime objective in the design of a heat exchanger is to determine the heat transfer area required for the specified duty, which can be done if the other three variables are known. Q and b.Tlm are calculated with the values of the heat transfer and the temperature of the streams given by CHEMCAD 3.20. The overall heat transfer coefficient depends on the heat transfer resistance of the individual coefficients and is derived from [10, table a.7-1, p487] and [6, p567-568] , depending on the system used.

Shell and bundie diameters

If the heat-transfer area is known or estimated, the number of tubes can be calculated. With an estimate of the diameter, thickness and length of the tubes, the bundie diameter Db (in mm) can than be obtained from the equation:

where :Db =

do =

Nt =

Kl,nl =

bundle diameter [mm] outside tube diameter [mm] number of tubes [-]

empirical constants.

(4.14)

Kl and nl are empirical constants which depend on the tube arrangements and the number of tube passes and can be found in [6, table 12.4, p577]. The shell diameter can be obtained by determining the shell-bundle clearance and add this to the bundle diameter.

Tube si de pressure drop

Due to friction in the tube, there will be a pressure drop, which can be calculated with the following formula :

where : 6Pt = Np = Jf = L = I1P = N [.8j .

(.!::...)(J::.-)-m

+ 2.5.] P . U,2 (4.15) , P f d. " 2 I t" w

tube-side pressure drop (Pa) number of tube passes (-)

tube side friction factor (is obtained from figure 2.24 [8]) leng th of one tube (m)

inside diameter (m) viscosity ratio::::: 1

density of the fluid (kg m-3)

tube side velocity (m S-I)

Shell side pressure drop

Due ta the baffles in the shell there will also be a pressure drop, which can be ca1culated with the following formula :

Where :L1P_s

=

Ds

=

de

=

lb =

Us =

M.

=

8j

.(DI'

).(~).

P

.u;

.(~)-O.I.t

.1 f d l 2 "

e B t"w

shell-side pressure drop (Pa) shell inside diameter (m)

shell side equivalent diameter (m) baffle spacing (m)

shell side velocity (ms-I)

(4.16)

For the calculatian of the shell-side equivalent diameter the following formula is used for a triangular pitch arrangement:

d

=

l.lO.(p2_0.917'd2 )

e d ' 0 "

(4.17)

where: Pt = tube pitch (mm)

A full design of heat exchanger H2 and ca1culations of the other exchangers are

FVO 3198 Equipment design H2 35.9 200 212.01 - 846.67 H3 212.40 200 340.23 3121.41 H4 110.30 200 47.05 11722.31 H5 11.31 50 99.38 2276.01 H6 341.80 1200 105.4 2702.39

4.15 The Pumps

The theoretical power required for pumping an incompressible fluid is given by: where Pt = ~P =

Qm

=

p

=

Qv

=

M·Qm p= =M·Q· I P v Theoretical power [W]

Pressure differential across the pump [Pa] Mass flow rate [kg S-I]

Density of the fluid [kg m-3]

Volumetrie flow rate [m3 S-I]

(4.18)

The actual power required is obtained by combining the efficiency with forrnula 4.18 : where: Pa =

11

=

M·Q P = v (J Tl Actual power [W] Overall efficiency [-] (4.19)

The calculations are given in appendix V, the values of the actual power are given in table 4.8

Table 4.8 : Calculated theoretical and actual

PI 4.47 5.58 5.59

P2 1.44 1.80 1.81

P3 5.48 6.85 6.86

4.16 Compressors and Turbines

In the gasification process two compressors and four turbines are used. The compressors are used to compress the process-streams to the right pressure. The first one is used to supply the O₂ for the gasifier and the second one supplies an