2 ECN-C--00-087

AbstractThe nitric acid industry in the Netherlands is a major industrial source of nitrous oxide (N2O), animportant contributor to the greenhouse effect. The project aims at determining the feasibility of theuse of a catalyst developed by ECN for the direct decomposition of N2O and the Selective CatalyticReduction (SCR) of N2O with natural gas in the off gas of a nitric acid plant. For plants with a tail gastemperature of at least 400°C in front of the expander, direct decomposition is an economic option(from NLG 1 per ton CO2 equivalents removed). For plants with a temperature below 400°C in frontof the expander, SCR using natural gas can be used. More economic, however, is end-of-pipe SCR ofN2O using 500 ppmv LPG (upwards of NLG 3 per ton CO2 equivalents removed).

Keywordsnitrous oxide, N2O, greenhouse effect, nitric acid production, abatement, catalysis, directdecomposition, SCR, natural gas, LPG, elevated pressures.

SamenvattingDe salpeterzuurindustrie is een substantiële Nederlandse industriële bron van lachgas (N2O), eenbelangrijk broeikasgas. Het doel van het project is het onderzoeken van de haalbaarheid van hettoepassen van een door ECN ontwikkelde katalysator om N2O te verwijderen uit de afgassen van eensalpeterzuurfabriek door middel van directe N2O ontleding en selectieve katalytische reductie (SCR)met aardgas. In salpeterzuurfabrieken met een temperatuur van ten minste 400°C kan N2O verwijderdworden door middel van ontleding met de ECN katalysator (vanaf NLG 1 per ton verwijderde CO2equivalenten). Voor fabrieken met een lagere temperatuur kan SCR met aardgas toegepast worden bijhoge druk. Het is echter kosteneffectiever om SCR van N2O met LPG na de expander toe te passen(vanaf NLG 3 per ton verwijderde CO2 equivalenten).

Trefwoordenluchtemissies, lachgas, distikstofoxide, N2O, broeikasgas, salpeterzuurproductie, klimaatbeleid,katalyse, directe decompositie, SCR, aardgas, LPG, hoge druk.

VerantwoordingAan dit project is in het kader van het Besluit Milieusubsidies, Regeling Milieugerichte Technologieeen subsidie verleend uit het programma Reductie Luchtemissies Bedrijven 1999 dat gefinancierdwordt door het ministerie van Volkshuisvesting, Ruimtelijke Ordening en Milieubeheer. Novembeheert dit programma.

BegeleidingscommissieW. Ch. Glasz (DSM)M. Pottier (Hydro)K. Schöffel (Hydro)J. Williams-Jacobse (VROM)W.C. van der Lans (Novem)

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CONTENTS

SUMMARY 5

SAMENVATTING 7

1. INTRODUCTION 91.1 N2O abatement in a nitric acid plant 91.2 Nitric acid plants in the Netherlands 10

2. PROJECT OVERVIEW 112.1 Aim of the project 112.2 Catalyst tests 112.3 Calculation of the cost efficiency 11

3. RESULTS OF THE CATALYTIC TESTS 123.1 Direct decomposition of N2O 123.2 Selective Catalytic Reduction with methane 123.3 Selective Catalytic Reduction at atmospheric pressure 13

4. TECHNICAL AND ECONOMIC EVALUATION 134.1 Direct decomposition of N2O 134.2 SCR using natural gas in front of the expander 144.3 End-of-pipe removal of N2O 14

5. GENERAL CONCLUSIONS 16

REFERENCES 17

ATTACHMENT A: CATALYST TESTS 19Direct catalytic decomposition of N2O. 19Selective Catalytic Reduction of N2O with methane 20Effect of space velocity and pressure 20Temperature rise 21SCR of N2O downstream the expander: LPG-SCR. 22

ATTACHMENT B: BASIS OF DESIGN 23General tail gas composition 23Inlet conditions and outlet requirements direct decomposition 24Inlet conditions and outlet requirements SCR using natural gas 24Production plant information 25CO2 reduction efficiency (TEWI guidelines) 25

ATTACHMENT C: INVESTMENT AND EXPLOITATION COSTS 27Total investment costs 27Site specific aspects - retrofit 27Exploitation cost 28Annual Cost and Cost Efficiency 28Fluctuation of exploitation cost 29

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SUMMARY

ECN has developed a catalyst for the selective catalytic reduction (SCR) of N2O from the tail gases ofnitric acid plants. In a previous project (358510/0310, see ref. 6) the technical and economic feasibilityfor SCR using propane as a reductant has been investigated. The costs of propane (or LPG) werefound to be responsible for 50 to 70% of the total annual costs. The catalytic system can be used athigh pressure in front of the expander and as an end-of-pipe configuration at atmospheric pressure.The cost efficiency was around NLG 5 per ton CO2 equivalents removed in both cases, with a N2Oconversion efficiency of 90 to 95%.In the current project the direct composition of N2O (without any reductant) and the replacement ofpropane by methane for SCR with the ECN catalyst were studied. For both options the ECN catalystwas tested under realistic conditions, a conceptual process set-up was made, and the cost efficiencywas calculated.

Direct decomposition of N2O to nitrogen and oxygen with the ECN catalyst is possible at a tail gastemperature of 400°C and a space velocity of 13,000 h–1 (N2O destruction is higher than 70%). Athigher temperatures N2O conversion increases and a higher space velocity can be used. N2Odestruction efficiency improves at higher pressure. Higher concentrations of water have a smallnegative effect on N2O destruction. The cost efficiency for 70% N2O destruction at a tail gastemperature of 450°C in front of the expander is around NLG 1 per ton CO2 equivalents removed. At atemperature of 400°C the cost efficiency is NLG 2.2 per ton CO2 equivalents removed, because alarger catalyst volume has to be used.

SCR of N2O with methane occurs at lower temperature than direct decomposition, but methane is notas effective a reductant as propane. Moreover, at temperatures where N2O conversion is high, not allmethane has reacted. The emission of unreacted methane is unwanted, because methane is agreenhouse gas itself. Methane conversion increases strongly with pressure, so for N2O removal infront of the expander (pressure is 10 bara) methane-SCR can be used. The cost efficiency ranges fromNLG 4 to 6 per ton CO2 equivalents removed, based on a 75% N2O conversion efficiency anddepending on the complexity of process integration.

For N2O removal from the tail gases of plants with a temperature lower than 400°C in front of theexpander, end-of-pipe selective catalytic reduction using 500 ppmv of propane (C3H8:N2O ratio 1:3) ismore cost efficient than SCR in front of the expander using methane. Because this end-of-pipe N2Oremoval technique is located behind the expander, the effect on the nitric acid production process isvery limited. The heat generated by the SCR reaction is used to pre-heat the tail gases from about100°C to 380°C. At this reactor inlet temperature a N2O destruction efficiency of 90% is reached at aspace velocity of 20,000 h–1. The cost efficiency of this system is about NLG 3 per ton CO2equivalents removed.

Consumption of reducing agent (natural gas or LPG) is – apart from the capital costs – the maincontributor to the cost efficiency of SCR (contribution of 15 up to 40%). The prices for natural gas andLPG are subject to fluctuations of the oil price. The impact of a ca. 20% higher or lower price for thereducing agent has been determined, resulting in accuracy ranges of ± NLG 0.3 for the cost efficiencyof SCR.

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SAMENVATTING

ECN heeft een katalysator ontwikkeld voor de selectieve katalytische reductie (SCR) van N2O, datvrijkomt bij de productie van salpeterzuur. In een voorgaand project (358510/0310, zie referentie 6)zijn de technische en economische haalbaarheid van SCR met propaan als reductant onderzocht. Dekosten van propaan (of LPG) blijken verantwoordelijk te zijn voor 50 tot 70% van de totale jaarlijksekosten. De katalytische reactor kan zowel bij hoge druk, upstream de expander, als bij atmosferischedruk, downstream de expander, ingezet worden. Voor beide gevallen bedraagt de kosteneffectiviteitrond NLG 5 per ton verwijderde CO2-equivalenten bij een N2O verwijderingsrendement van 90-95%.In het onderhavige project zijn de directe decompositie van N2O (dus zonder gebruik te maken vaneen reductant) en de vervanging van propaan door methaan onderzocht. De ECN katalysator is voorbeide opties getest onder realistische omstandigheden. De resultaten zijn gebruikt voor het opstellenvan een concept-procesontwerp en het berekenen van de kosteneffectiviteit.

De ECN katalysator katalyseert de directe ontleding van N2O vanaf een afgastemperatuur van 400°Cbij een space velocity van 13.000 h–1. De N2O-verwijderingsgraad is dan groter dan 70%. Bij eenhogere afgastemperatuur vindt er meer N2O-omzetting plaats en kan een hogere space velocity (=minder katalysator) gebruikt worden. De ontleding van N2O verloopt beter bij een hogere procesdruk.Een hogere concentratie van water in het gas heeft een klein negatief effect op de N2O-verwijdering.De kosten effectiviteit bij een N2O verwijderingsrendement van 70% bedraagt ongeveer NLG 1 perton verwijderde CO2-equivalenten bij een afgastemperatuur van 450°C. Als de temperatuur van hetgas 400°C is, wordt de kosteneffectiviteit NLG 2.2 per ton verwijderde CO2-equivalenten, omdat eengroter katalysatorvolume nodig is.

SCR van N2O met methaan vindt plaats bij een lagere temperatuur dan directe decompositie, maarmethaan is minder effectief als reductant dan propaan. Niet al het methaan heeft geregeerd bij detemperatuur waarbij de N2O omzetting hoog is. Emissie van niet-gereageerd methaan is ongewenstomdat methaan zelf een broeikasgas is. Bij een verhoogde procesdruk is de methaanomzetting hoger,zodat methaan-SCR upstream de expander (procesdruk ca. 10 bara) wel toegepast zou kunnen worden.De kosteneffectiviteit voor een N2O verwijderingsrendement van 75% ligt tussen NLG 4 en 6 per tonverwijderde CO2-equivalenten, afhankelijk van de moeite die het kost de installatie in een bestaandefabriek in te bouwen (retrofit).

In fabrieken waar de temperatuur upstream de expander lager is dan 400°C, is een nageschakelde SCRinstallatie met een lage hoeveelheid propaan (500 ppm, C3H8:N2O-verhouding 1:3) kosteneffectieverdan SCR met aardgas upstream de expander. Omdat deze techniek downstream de expander kanworden toegepast is de invloed op het salpeterzuurproces minimaal. De reactiewarmte van de SCR-reactie wordt gebruikt om de afgassen op te warmen van ca. 100°C tot 380°C. Bij deze temperatuuraan de ingang van de katalytische reactor is de N2O-omzetting 90% bij een space velocity van 20.000h–1. De kosteneffectiviteit van dit systeem ligt rond de NLG 3 per ton verwijderde CO2-equivalenten.

Het verbruik van het reductant (aardgas of LPG) levert – op de kaptitaalslasten na – de grootstebijdrage aan de jaarlijkse kosten van N2O verwijdering door middel van SCR (15 tot 40%). De prijzenvoor aardgas en LPG zijn afhankelijk van schommelingen in de olieprijs. Een ca. 20% lagere of hogere prijs van de reductanten leidt tot variaties in de kosteneffectiviteit van ongeveer NLG 0.3 perton CO2-equivalent verwijderd.

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1. INTRODUCTION

Nitrous oxide (N2O, laughing gas) has been identified as one of the main compounds contributing tothe greenhouse effect next to CO2 and methane. The Global Warming Potential of N2O is about 310,i.e. on a 100-years time scale nitrous oxide is 310 times more effective per kg at trapping heat in theatmosphere than CO2. Nitrous oxide is also active in the destruction of the ozone layer.The nitric acid industry is one of the major industrial sources of N2O. In the Netherlands thecontribution of the nitric acid production to the total N2O emissions is about 40%. Nitrous oxide is abyproduct of the oxidation of ammonia to NOx; its concentration in the off-gases ranges from 500 to3500 ppmv [1].Nitrous oxide is one of the greenhouse gases of which the emissions have to be reduced according tothe Kyoto protocol. If all 8.5 Mton CO2 equivalents of nitrous oxide are removed from the off gases ofthe Dutch nitric acid plants, more than 15% of the targeted 50 Mton emission reduction for 2010 isachieved. However, to date, there are no techniques commercially available that can be readilyimplemented in an existing nitric acid plant [1].

1.1 N2O abatement in a nitric acid plantA nitric acid plant has two separate reaction sections. First, an ammonia combustion reactor, in whichammonia and air are converted to NOx over a Pt/Pd gauze at around 900°C. Subsequently, the gas iscooled in a series of heat exchangers and led into the NOx absorption section, where nitric acid isformed. The exit gases are heated using the heat produced in the ammonia combustion. Finally, thegases are led through an expander for energy recovery (Figure 1).

Several abatement techniques are being developed.BASF is currently testing a catalyst for thedecomposition of N2O directly behind the Pt/Rhammonia oxidation gauze catalyst in a nitric acidplant. The high temperature (850 to 950°C) mayaffect the catalyst stability. Furthermore, the catalystmay decompose NOx as well, which would reduce theefficiency of the nitric acid production process [2,3].

A safer option, that does not influence the NOxproduction efficiency, may be the direct catalyticdecomposition behind the absorption section of anitric acid plant, where the conditions are less severe.In the literature, many materials have been screenedfor the direct decomposition of nitrous oxide [4]. Temperatures for complete conversion of N2O of aslow as 250°C have been reported [5]. However, these tests have not been performed under conditionsthat are relevant for the nitric acid industry. Especially the presence of water in the off gases has animportant effect on the ignition temperature and on the N2O destruction efficiency of the catalysts.To date, no catalyst has been found that shows sufficient N2O destruction efficiency below 400°C inpresence of water. However, in many of the nitric acid plants in Europe the temperature behind theabsorption section is lower than 400°C.

Destruction of N2O can be carried out at lower temperature by adding a hydrocarbon as a reductant.ECN developed an iron-containing zeolite catalyst for the selective catalytic reduction of N2O usingpropane as a reductant [7]. Nitrous oxide reduction efficiencies of higher than 90% are achieved at areactor inlet temperature below 300°C. The SCR system can be installed in front of the expander orbehind the expander, directly in front of the stack (T = 100 - 200°C). In the latter case the gases can bepreheated using the heat of the reaction of propane with N2O.

Absorption

Pt/RhGauzecatalyst

Air Ammonia

����������������������������������������������������������������

Stack

Expander

HNO3

Direct Decomposition

•DirectDecomposition• SCR

Figure 1: Possible locations for N2Oremoval in a nitric acid plant.

1

Using these results, a conceptual design was made for the SCR process, as well as a calculation of theinvestment and operational costs. The cost efficiency of propane-SCR ranges between NLG 5 and 6.5per ton CO2 equivalents removed. Between 50 and 70% of the annual costs are related to the use of thereducing agent. Propane or LPG is generally not available on site of a nitric acid plant, whichnecessitates transport and storage facilities [6].

The current project aims at a further improvement of the cost-effectiveness of N2O removal in thenitric acid industry.

Direct catalytic decomposition of N2O seems an attractive option, because the temperature in front ofthe expander is 400°C or higher in all nitric acid plants in the Netherlands. Preliminary results showedthat the catalyst for propane-assisted SCR is also active in the decomposition of N2O [6].

Another option is the replacement of LPG by an alternative reductant. Methane is well known as areductant for the (N)SCR of NOx [8,9]. Moreover, it is readily available on site of nitric acid plants inthe form of natural gas.

1.2 Nitric acid plants in the NetherlandsThe emission of N2O from the nitric acid industry in the Netherlands takes place in six plants. DSMhas 4 plants (2 in Geleen and 2 in IJmuiden) and Hydro Agri owns two plants in Sluiskil. Table 1gives further information on the Dutch nitric acid plants [1].Kemira Agri also operates two nitric acid plants that are equipped with an NSCR installation and havevery low N2O emissions. Furthermore, they will be (or are already) closed.

Table 1: Data on Dutch nitric acid plantsPlant A B C D E FN2O-emission (ppm) 1250 1500 1150 1250 920 1500Annual N2O emission (kton/yr)2 4.4 4.5 - 6 1.5 - 2 5.4 1.7 2CO2 equivalents (Mton/yr)2 1.3 1.4 - 1.8 0.4 - 0.6 1.6 0.5 0.6Capacity (kton HNO3/yr) 584 500 210 730 255 245Plant type1 dual dual mono dual mono dualp in front of expander (bara) 10 11 5 11 4 10T in front of expander (°C) ? 500 450 ? 400 ?T behind expander (°C) ? 220 220 ? 220 ?H2O concentration (%) ? 0.5 1 ? 1 ?

N2O + CH4 + 1.5 O2 N2 + CO2 + 2 H2Ocatalyst

N2O N2 + 0.5 O2catalyst

1 In a dual pressure nitric acid plants the ammonia oxidation step and the NO absorption step arecarried out at different pressures (about 3 and 10 bara respectively). In mono pressure plants onepressure is used in the whole plant (about 5 bara).2 Estimated.

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2. PROJECT OVERVIEW

2.1 Aim of the projectThe project aims at establishing the feasibility of direct catalytic decomposition and methane-assistedSCR for N2O emission reduction in a nitric acid plant. To this end the catalyst is tested for bothreactions in a simulated off gas of a nitric acid plant. A minimum N2O destruction efficiency of 70% isrequired by Novem guidelines. The results of the tests have been used to develop a conceptual processdesign and to calculate the cost-effectiveness of these N2O removal strategies.

2.2 Catalyst testsThe tests of the catalyst were performed by ECN using a fully automated flow set-up. The set-upallows tests under simulated nitric acid plant tail gas conditions and can be operated under pressure.Table 2 summarises the test conditions.

Table 2: Reaction conditionsConditionsPressure range 1 to 4 baraTemperature range 150 - 500°CReactor diameter 25 mmCatalyst particles 0.71 - 1.4 mmSpace velocity 10,000 to 45,000 h–1 (STP)

The composition of the tail gas is different for every nitric acid plant. In agreement with Hydro Agriand DSM a representative gas composition was established, which was used in the tests. The influenceof variations in the concentration of several compounds was also investigated.

Table 3: Gas compositionCompound Representative

concentrationVariations tested

N2O 1500 ppm 1000 - 2000 ppmNO 100 ppm -NO2 100 ppm 0 - 240 ppmH2O 0.5 vol% 0 - 1 vol%O2 2.5 vol% 1 - 4 vol%

The results of the catalyst test are described in chapter 3 and presented in more detail in Attachment A.

2.3 Calculation of the cost efficiencyStork Engineering Consultancy B.V. (SEC)1 performed the technical and economic feasibility studyfor a DeN2O system, capable of reducing N2O in tail gas originating from existing nitric acid plants.The following configurations have been evaluated:

1. process-integrated direct decomposition (high pressure case);2. process-integrated N2O reduction using selective catalytic reduction with natural gas (high pressure

case);

1 SEC has been incorporated into Jacobs Comprimo Nederland, a division of Stork Engineers & Contractors B.V.

12 ECN-C--00-087

3. end-of-pipe N2O reduction using both selective catalytic reduction and direct decomposition(atmospheric pressure case).

In co-operation with DSM and Hydro Agri, SEC has composed a typical Basis of Design for the highpressure case (10 bara) and for the atmospheric pressure case, based on a tail gas flow of 200,000 Nm3

/ hour and an N2O concentration of 1,500 ppmv (see Attachment B). A conceptual process and reactordesign has been developed based on kinetic data of the catalyst performance.The conceptual process set-up and calculations of the cost efficiency are summarised in Chapter 4.Details on the cost calculations are described in Attachment C.

3. RESULTS OF THE CATALYTIC TESTS

3.1 Direct decomposition of N2OUnder the representative conditions (Table 3), N2Oconversion is 80% at 400°C and 4 bara when aspace velocity of 10,000 h–1 (STP) is used (seeFigure 2).At higher temperatures, a higher space velocity (i.e.a smaller amount of catalyst) can be used: forplants with a temperature of 450°C in front of theexpander, a space velocity as high as 45,000 h–1 canbe applied to reach 70% N2O destruction. At 500°CN2O conversion is higher than 95% with this spacevelocity (see Figure 2).A higher concentration of water in the tail gas hasonly a small effect on N2O destruction. For a plantwith a water concentration of 1% and a temperatureof 400°C (e.g. plant E in Table 1), 70% N2Odestruction efficiency is reached using a spacevelocity of 13,000 h–1.Direct decomposition also works at atmospheric pressure, but conversion of N2O improves when thepressure is increased from 1 to 4 bara. It is expected that performance at 10 to 11 bara (the pressure infront of the expander in dual pressure plants) will be somewhat higher than at 4 bara.Higher or lower concentrations of oxygen, NOx and N2O (see Table 3) have no influence on the N2Odestruction efficiency.Temperature increase due to the heat of the N2O decomposition reaction is about 4°C.

3.2 Selective Catalytic Reduction with methaneMethane-assisted N2O destruction takes place at about 50°C lower temperature than directdecomposition under the same conditions (Figure 3). However, methane is not as effective as areductant as propane. This is due to the lower reactivity of methane, which also causes the lowconversion of methane even at relatively high temperatures (see Figure 3). The emission of unreactedmethane is unwanted, because methane is a greenhousegas itself (GWP = 21).An N2O conversion of 70% is reached at a pressure of 4 bara and a tail gas temperature of 350°C – arepresentative temperature in front of the expander for European nitric acid plants – with a spacevelocity of 10,000 h–1 and 1500 ppmv of added methane.Emission of unreacted methane is about 300 ppmv at those conditions.Conversion of N2O and especially methane are much lower at atmospheric pressure. High methaneemissions make methane-SCR not a viable option for N2O abatement behind the expander.

0

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250 300 350 400 450 500

T (°C)

N2O

des

truct

ion

(%)

SV = 10,000

SV = 45,000

Figure 2: Direct decomposition of N2O.Conditions: p = 4 bara, 1500 ppmv N2O,200 ppmv NOx, 0.5% H2O, 2.5% O2.

ECN-C--00-087 13

Emission of unreacted methane is strongly decreased at higher pressure. When the system is applied ina dual-pressure plant at 10 bara, methane emission is expected to be lower (about 150 - 200 ppm).

3.3 Selective Catalytic Reduction at atmospheric pressureBecause methane-SCR is no viable option for N2O destruction behind the expander, an alternative wasdeveloped. SCR with 500 ppmv of propane provides an N2O destruction efficiency that is higher thanwith 4500 ppmv methane (Figure 4). Propane conversion is almost complete at a temperature wheremethane conversion does not yet occur; propane slip and CO emission are below 10 ppmv above350°C.Using 500 ppmv propane instead of 1500 ppmv propane (as used in project 358510/0310, see ref. 6)has only a small effect on N2O destruction efficiency. The adiabatic temperature rise due to thereaction of 500 ppm propane, N2O and oxygen is about 34°C.

4. TECHNICAL AND ECONOMIC EVALUATION

4.1 Direct decomposition of N2OThe cost calculation is based on a tail gas temperature of 450°C in front of the expander, a spacevelocity of 45,000 h–1, and 70% N2O conversion. The process set-up is simple: two tie-ins should bemade in the existing tail gas line between the nitric acid process and the expander. A lateral flowcatalytic reactor (catalyst volume about 5 m3), which is in essence a fixed bed plug flow reactor wherethe tail gas is forced to pass through a flat layer of catalyst contained between gauze sheets, isconnected to the tie-ins.− The pressure drop is estimated at 40 mbars, which causes a lower efficiency of the expander.

The costs due to a lower energy recovery are included in the cost calculations.− The costs per ton CO2 equivalents removed are about NLG 1.2. The total investment costs are

about NLG 5 million (extra costs due to retrofit of 50% of the investment costs are included).− For plants with a higher temperature in front of the expander, N2O conversion is higher and cost

efficiency is improved to below NLG 1 per ton CO2 equivalents removed.− For plants with a temperature of 400°C a lower space velocity of 13,000 h–1 (i.e. a larger DeN2O

reactor) is necessary to reach 70% N2O conversion. This has a considerable impact on the costefficiency: NLG 2.2 to 2.5 per ton CO2-equivalents removed.

− Variations in the catalyst costs have a relatively small impact on cost efficiency.

0

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250 300 350 400 450 500

T (°C)

N2O

and

CH

4 des

truct

ion

(%)

CH4-SCR:N2O destruction

N2O Decomposition

CH4 destruction

Figure 3: SCR of N2O with methane.Conditions: p = 4 bara, S.V. = 20,000 h–1,1500 ppmv N2O, 200 ppmv NOx, 0.5%H2O, 2.5% O2, 4500 ppmv CH4.

0

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T (°C)

N2O

des

truc

tion

(%)

500 ppm C3H8

4500 ppm CH4

1500 ppmC3H8

Figure 4: SCR of N2O. Conditions:atmospheric pressure, S.V. = 20,000 h–1,1500 ppm N2O, 200 ppm NOx, 0.5%H2O, 2.5% O2.

14 ECN-C--00-087

4.2 SCR using natural gas in front of the expanderAs a base for the cost calculations, a DeN2O reactor with a space velocity of 10,000 h-1 used at apressure of 10 bar(a), an inlet temperature of 350°C and a natural gas dosing of 1500 ppmv is used.The extrapolated slip of natural gas is maximum 10% (=150 ppm). At those conditions, a 75% N2Oreduction is obtained. Temperature increase due to catalytic conversion of natural gas is about 25-30°C.

Figure 5 shows a conceptual process set-up. Natural gas is added to the nitric acid plant tail gas beforeentering the DeN2O reactor, a lateral flow reactor with a catalyst volume of 21.5 m3. The reaction ofN2O with natural gas leads to a temperature increase over the DeN2O reactor and a reactor outlettemperature of 380ºC. Downstream the DeN2O reactor the cleaned tail gas is cooled in tail a gas coolerto 350ºC. This tail gas cooler is a steam boiler using boiler feed water of 105ºC producing saturatedsteam with a pressure of 21 bar(a) and a temperature of 215°C. The cooling step is necessary forproper functioning of the existing expander.− The pressure drop over the SCR DeN2O system using natural gas is approximately 160 mbar.− The cost efficiency ranges between NLG 4 and NLG 6 per ton CO2-equivalents removed,

depending on the costs for retrofit in an existing nitric acid plant.− Total investment costs are about NLG 15 million (retrofit factor 50%) and 25 million (retrofit

factor 150%). The investment costs are higher than for direct decomposition due to a largerreactor and catalyst volume (S.V. = 10,000 h–1 compared to 45,000 h–1 for direct decompositionat 450°C). Also the costs of extra equipment, such as the tail gas cooler, add to the investmentcosts.

− The annual costs of this SCR of N2O using natural gas are NLG 4 to 6 million. The contributionof the costs of natural gas to the annual costs is 20 to 25%.

− Variations in the catalyst and the natural gas costs have a relatively small impact on the costefficiency. A 100 mbar higher pressure drop adds NLG 0.4 to the costs per ton CO2 equivalentsremoved.

4.3 End-of-pipe removal of N2OAt atmospheric pressure, the use of natural gas as a reductant was no feasible option. Selectivecatalytic reduction using 500 ppmv of LPG was investigated instead. Over 70% reduction of N2O canbe reached with a space velocity of 20,000 h–1 and an inlet temperature of 340°C.Tail gas originating from the expander outlet (1.1 bar(a), 100ºC) is heated to 340ºC in a recuperatorusing the heat of the tail gas leaving the DeN2O reactor. During start-up the tail gas is heated up withan in-line burner. A fixed flow of LPG is added from an LPG storage tank and the gas enters thelateral flow reactor (catalyst volume: 11 m3). The surplus heat of this reaction (about 30°C) is used tomake up for the heat losses over the recuperator and to sustain an adequate temperature level in the

Figure 5. Conceptual process set-up for N2O removal using natuaral gas in front ofthe expander.

������

SCR Reactor

Cooler

tail gasabsorber

natural gas

tail gasexpander

steamcondensate

In-line mixer

ECN-C--00-087 15

DeN2O reactor. The cleaned tail gas is cooled to 130ºC in the recuperator and vented to theatmosphere via the existing stack (see Fig. 6).− The pressure drop over the SCR DeN2O system using 500 ppmv LPG is approximately 120 mbar.− This process variant for the DeN2O system can be considered as an end-of-pipe technique. It

gives minimum interference with the production process rendering the tail gas.− The cost efficiency is about NLG 4.5 per ton CO2 equivalents for 70% N2O conversion.− The cost efficiency can be improved to approximately NLG 3 per ton CO2 equivalents, by

increasing the N2O conversion to 90%, which takes little extra investment costs.− Total investment costs are NLG 11 million, with the costs of the recuperator as a major

contributor (about 25%).− The annual costs are about NLG 4 million, 30% of which is due to the costs of LPG (including

transportation costs).− Variations in the catalyst and the LPG costs have a relatively small impact on cost efficiency. A

pressure drop of 220 mbar instead of 120 mbar, adds NLG 0.5 to the costs per ton CO2equivalents removed. Using a cheaper or a more advanced recuperator can improve the costefficiency.

− An alternative for the use of LPG is operating the DeN2O reactor at a temperature high enoughfor direct N2O decomposition. Burning a sufficient amount of natural gas in the in-line burnermakes up for the heat loss over the recuperator.

− Cost efficiency of this option is about NLG 3.5 to 4 per ton CO2 equivalents removed. The fuelcosts are lower and the prize of the recuperator is higher than for SCR with 500 ppmv LPG.

����Reactor

tail gasexpander

natural gas

stack

In-line burner (start-up)

Air-air-preheater

LPG

Figure 6. Conceptual process set-up for end-of-pipe N2Oremoval.

16 ECN-C--00-087

5. GENERAL CONCLUSIONS

− For plants with a tail gas temperature of at least 400°C in front of the expander, directdecomposition of N2O is the most economic choice. The ECN catalyst works from about 400°C.At higher temperatures catalytic activity increases strongly, so that a reactor with a smallercatalyst volume can be used. Cost efficiency ranges between NLG 1 (high temperature, 95% N2Oreduction) and NLG 2.2 (tail gas temperature of about 400°C, 70% N2O reduction) per ton CO2equivalents removed (see Figure 7).

− For plants with a temperature of 350°C in front of the expander, selective catalytic reduction ofN2O with natural gas can be used. This option is only possible in high pressure plants, for at lowpressure emission of unreacted methane (a greenhouse gas itself) is high. Cost efficiency for 75%N2O reduction ranges between NLG 4 and 6 per ton CO2 equivalents removed, depending on thecomplexity of retrofitting the installation.

− For N2O removal from the tail gases of plants with a temperature lower than 400°C, end-of-pipeselective catalytic reduction using 500 ppmv of propane is more cost efficient than SCR in frontof the expander using methane. The costs per ton CO2 equivalents removed range from NLG 3(90% N2O reduction) to NLG 4.5 (70% N2O reduction). Direct decomposition at atmosphericpressure (using natural gas to heat the tail gases) is expected to have a comparable costefficiency.

− The cost efficiency for SCR using 500 ppmv propane (from NLG 3 per ton CO2 equivalentsremoved) is better compared to the SCR option using 1500 to 1900 ppmv propane from theprevious project (around NLG 5 per ton CO2 equivalents removed). (see Fig. 7).

− Before implementation of the N2O removal system is possible, pilot plant tests are necessary.Pilot plant tests include a long run catalyst test testing under real conditions in, preferably in aside stream of a nitric acid plant.

Figure 7: Comparison of cost efficiency of the various options for N2O removal in anitric acid plant (blue = from this study, green = from project 358510/0310, see ref. 6).Ranges in cost efficiency are mainly determined by:− N2O destruction efficiency. Generally a higher N2O conversion is more cost

efficient, although annual costs may be higher.− The cost of the reducing agent (LPG or methane) has a strong effect on the cost

efficiency of SCR.− The cost efficiency of DeN2O techniques situated upstream the expander are

influenced by the complexity of retrofit in an existing plant.

Cost Efficiency (NLG per ton CO2 equivalents removed)0 1 2 3 4 5 6

T > 400°C

T < 400°C

Direct decompositionT = 400°C

In front of the expander

End-of-pipe

SCR, LPG

SCR, natural gas

SCR, 500 ppm LPG

SCR, 1500 ppm LPG

Direct decomposition

Direct decompositionT > 450°C

ECN-C--00-087 17

REFERENCES

1. M.T. Taal, Dutch Notes on BAT for the production of nitric acid, December 1999, Infomil,Den Haag, the Netherlands.

2. G. Kuhn, V. Schumacher, E. Wagner (BASF), Catalytic reduction in nitric acid plants ofN2O from adipic acid, The International Fertilizer Society, proceeding No. 435a, 1999.

3. M. Schwefer, R. Maurer, and M. Groves (Krupp Uhde), Reduction of Nitrous oxideemissions from nitric acid plants, paper presented at "Nitrogen 2000" conference, 12-14march, 2000, Vienna, Austria.

4. F. Kapteijn, J. Rodriquez-Mirasol, and J.A. Moulijn, Heterogeneous CatalyticDecomposition of Nitrous Oxide, Appl. Catal. B 9 (1996) 25.

5. K. Yuzaki, T. Yarimizu, S. Ito and K. Kunimori, Catalytic decomposition of N2O oversupported rhodium catalysts: high activities of Rh/USY and Rh/Al2O3 and the effect of Rhprecursors, Catal. Lett 47 (1997) 173.

6. R.W. van den Brink, M.J.F.M. Verhaak, J.W.N. van Lijssel, and M.M.C. Gent, Catalyticdestruction of N2O in the nitric acid industry – Executive Summary, report ECN-C--00-003,January 2000, Petten, the Netherlands.

7. R.W. van den Brink, S. Booneveld, J.R. Pels, D.F. Bakker, and M.J.F.M. Verhaak, Catalyticremoval of N2O in model flue gases of a nitric acid plant using a promoted Fe zeolite, Appl.Catal. B, submitted.

8. M. Ogura, M. Hayashi, S. Kage, M. Matsukata, and E. Kikuchi, Determination of activepalladium species in ZSM-5 zeolite for selective reduction of nitric oxide with methane,Appl. Catal. B 23 (1999) 247.

9. Y. Li and J.N. Armor (Air Products), Catalytic reduction of NOx using methane in thepresence of oxygen, U.S. patent 5,149,512 (1992).

18 ECN-C--00-087

ECN-C--00-087 19

ATTACHMENT A: CATALYST TESTS

Direct catalytic decomposition of N2O.In Figure A-1 the dependence of the conversionof N2O on the space velocity is shown, i.e. thetotal flow (calculated at room temperature andatmospheric pressure) divided by the catalystbed volume. To reach a N2O destruction level of70% at 400°C (the lowest available temperaturein front of the expander in a Dutch nitric acidplant), a space velocity of 13,000 h–1 isrequired. At 450°C – which is considered as anaverage for the Dutch nitric acid industry – aspace velocity as high as 45,000 h–1 can beused.It must be noted that conversion is plottedagainst the average temperature between inletand outlet of the reactor. The calculatedadiabatic temperature rise due to decompositionof N2O is only 4°C at 450°C.

Direct catalytic decomposition of N2O is only an option in front of the expander in a nitric acid plant,where the pressure ranges from 4 to 11 bara. Our test facility operates at pressures up to 4 bara. Inorder to make extrapolation to 10 bara possible, the pressure dependence was investigated. Figure A-2shows that N2O conversion improves when the pressure is raised from 1 to 4 bara. The differencebetween 3 and 4 bara is such, that the N2O conversion is expected to increase further when pressure israised to 10 bara. However, for the economic evaluation the conversion values at 4 bara were used.

The presence of water is of important influence on the ability of catalysts to remove N2O from a gasstream. In absence of water the N2O conversion curve is shifted to about 50°C lower temperaturecompared to when 0.5% of water is present (Figure A-3). In a gas stream containing 1% of water, theconversion of N2O is somewhat lower. The difference in conversion between 1% and 0.5% of water ismuch smaller than between 0% and 0.5%. Figure A-3 shows that at a space velocity of 13,000 h–1 70%

Figure A-3. Effect of water concentration on thedirect catalytic decomposition of N2O Conditions:p = 4 bara, SV = 13,000 h–1, 1500 ppmv N2O, 100ppmv NO, 100 ppmv NO2, 2.5% O2, balance N2.

0

20

40

60

80

100

250 300 350 400 450 500

T (°C)

N 2O

con

vers

ion

(%)

0% water0.5% water1% water

0

20

40

60

80

100

250 300 350 400 450 500

T (°C)N 2

O c

onve

rsio

n (%

)

SV = 10,000SV = 13,000SV = 20,000SV = 30,000SV = 45,000

Figure A-1. Direct catalytic decomposition of N2O.Conditions: p = 4 bara, 1500 ppmv N2O, 100 ppmvNO, 100 ppmv NO2, 0.5% H2O, 2.5% O2, balance N2.

Figure A-2. Effect of pressure on the directcatalytic decomposition of N2O Conditions: SV =13,000 h–1, 1500 ppmv N2O, 100 ppmv NO, 100ppmv NO2, 0.5% H2O, 2.5% O2, balance N2.

0

20

40

60

80

100

250 300 350 400 450 500

T (°C)

N 2O

con

vers

ion

(%)

1 bara3 bara4 bara

20 ECN-C--00-087

N2O conversion is not yet reached at 400°C when 1% H2O is present (conditions of nitric acid plant Ein Table 1). When the space velocity is decreased to 10,000 h–1, the conversion of nitrous oxide isabove 70% at 400°C in the presence of 1% of water.The concentration of oxygen was varied between 1 and 4% with completely identical N2O conversioncurves as a result. The concentration of NOx (NO2 concentration between 0 and 240 ppm) had noinfluence on catalyst activity either. Finally, the concentration of nitrous oxide itself (varied between1000 and 2000 ppm) also did not have a significant effect on N2O conversion.

Selective Catalytic Reduction of N2O with methane

Effect of space velocity and pressureMethane is a more stable molecule than propane and more difficult to activate by the catalyst. This isreflected in an about 50°C highertemperature for N2O destruction undersimilar conditions (Figure A-4a). Comparedto direct decomposition the addition of4500 ppmv of methane to the feed reducesthe N2O conversion temperature by approx.50°C.Figure A-4b shows that there is a largedifference in the conversion of thereductant. The conversion of propane isalmost complete at 350°C, while methaneconversion is only 18% at the sametemperature. Obviously, emission ofmethane has to be limited, because methaneis a greenhouse gas (global warmingpotential is 21 CO2 equivalents).

With a space velocity of 20,000 h–1,conversion of N2O in presence of methaneis still below the required 70% at 350°C, thetemperature chosen as representative fornitric acid plants. In order to lower thetemperature for nitrous oxide destruction

Figure A-5. Nitrous oxide conversion (full markers)and methane conversion for CH4-SCR at differentpressures. Conditions: SV = 10,000 h–1, 1500 ppmvN2O, 100 ppmv NO, 100 ppmv NO2, 0.5% H2O, 2.5%O2, 4500 ppmv CH4, balance N2.

0

20

40

60

80

100

250 300 350 400 450 500

T (°C)

N 2O

and

CH 4

con

vers

ion

(%)

N2O, 1 bara CH4, 1 baraN2O, 2 bara CH4, 2 baraN2O, 4 bara CH4, 4 bara

0

20

40

60

80

100

250 300 350 400 450 500

T (°C)

Hydr

ocar

bon

conv

ersi

on (%

)CH4C3H8

0

20

40

60

80

100

250 300 350 400 450 500

T (°C)

N 2O

con

vers

ion

(%)

CH4-SCRC3H8-SCRDecomposition

Figure A-4. Nitrous oxide conversion (a) and hydrocarbon conversion (b) for CH4-SCR, C3H8-SCRand direct decomposition. Conditions: p = 4 bara, SV = 20,000 h–1, 1500 ppmv N2O, 100 ppmv NO,100 ppmv NO2, 0.5% H2O, 2.5% O2, 4500 ppmv CH4 or 1900 ppmv C3H8, balance N2.

a b

ECN-C--00-087 21

and to increase methane conversion, tests at a space velocity of 10,000 h–1 were performed (Figure A-5). N2O conversion is higher than 80% at 350 °C, but CH4 conversion is still only about 40% at 350°C.

Pressure is an important parameter to assess the feasibility of N2O removal with natural gas upstreamand downstream the expander. Downstream the expander, the pressure is atmospheric and from FigureA-5 it is clear that the conversion of methane remains very low in this case. From 1 to 4 bara, methaneconversion increases almost linearly with pressure. This is of importance for the extrapolation to 10bara, which is the pressure upstream the expander in dual pressure nitric acid plants (see Table1).Higher pressures also have a positive effect on N2O destruction efficiency, albeit that the effect issmaller than with methane conversion (Figure A-5).The N2O destruction efficiency depends on the concentration of the added methane. With 4500 ppmv

CH4 the N2O conversion is higher than 80% at anaverage reactor temperature of 350°C. An addedmethane concentration of 1500 ppmv yields a lowerN2O destruction efficiency. Methane conversion ishigher when lower amounts of methane are used(Figure A-6).In presence of methane also NOx reduction takesplace. Between 350 and 450°C NOx conversionranges from 10 to 50% dependent on reactionconditions. Low space velocity, high pressure, andhigh methane concentration favour NOx destruction.Formation of carbon monoxide – regarded as adrawback in the SCR of N2O [2] – was very low inmethane-SCR over the ECN catalyst. At 350°C lessthan 5 ppmv of CO was found in any of theexperiments.

Temperature riseAll the figures presented above show conversions plotted against average reactor temperature. Thereaction of N2O with methane and oxygen is exothermic, which causes an increase of the outlettemperature of the catalyst bed. The calculated adiabatic temperature rise for the reaction of 4500ppmv CH4 with N2O and oxygen at 4 bara is 120°C for full conversion of methane. For the experimentat a space velocity of 10,000 h–1, 4500 ppmv methane and 4 bara, CH4 conversion is approx. 40%, sothe temperature rise is about 50°C.In a representative nitric acid plant the tail gaseshave a temperature of 350°C entering thecatalytic converter. Using the temperature rise of50°C, an inlet temperature of 350°C would yieldan average temperature of 375°C. At the lattertemperature N2O conversion is well above 90%.With 1500 ppmv of added methane, theconversion of N2O is around 75% at an inlettemperature of 350°C.

When the adiabatic temperature rise is taken intoaccount, the conversion curve for N2O inpresence of 4500 ppmv CH4 comes close to thatof direct decomposition. Yet, the action of CH4is not only to raise the temperature. When 4500ppmv of methane is fed to the catalyst in absence

Figure A-7. Methane conversion in presence andabsence of 1500 ppmv N2O, 100 ppmv NO and 100ppmv NO2. Conditions: p = 4 bara, SV = 10,000 h–1,1500 ppmv CH4, 0.5% H2O, 2.5% O2, balance N2.

0

20

40

60

80

100

250 300 350 400 450 500

T (°C)

CH4 c

onve

rsio

n (%

)

with N2Owithout N2O

Figure A-6. Nitrous oxide and methaneconversion for CH4-SCR with differentconcentrations of methane. Conditions: p = 4 bara,SV = 10,000 h–1, 1500 ppmv N2O, 100 ppmv NO,100 ppmv NO2, 0.5% H2O, 2.5% O2, balance N2.

0

20

40

60

80

100

250 300 350 400 450 500

T (°C)

N 2O

and

CH 4

con

vers

ion

(%)

N2O (4500 ppm CH4)N2O (1500 ppm CH4)CH4 (4500 ppm CH4)CH4 (1500 ppm CH4)

22 ECN-C--00-087

of N2O (in other words: catalytic combustion of methane), methane conversion takes place at a highertemperature. The combustion of methane starts at 400°C, while the reaction with N2O starts already at300°C (Figure A-7). As in the case of propane, methane acts as a reagent rather than just to raise thetemperature.

SCR of N2O downstream the expander: LPG-SCR.The use of natural gas as a reductant for N2O abatement at atmospheric pressure seems no realisticoption, because of the high emissions of greenhouse gas methane. Also the heat production (approx.25°C) due to the reaction of methane is too low. In the atmospheric pressure option, the heat producedis recirculated in order to pre-heat the gases entering the reactor to the required reactor-inlettemperature. A typical heat exchanger needs a temperature difference of about 30°C to operate.Adding 500 ppmv of propane to the gas stream can create a temperature difference of 30°C over thereactor. In the previous calculations on N2O abatement with LPG (mixture of propane and butane)1500 ppmv of propane was used. Lowering the LPG input to 500 ppmv will improve the costeffectiveness of LPG-SCR.Figure A-8 shows that N2O conversion takes place at lower temperatures with 500 ppmv propanecompared to 4500 ppmv methane. Furthermore, propane has almost completely reacted at 350°C,while methane conversion does not yet take place at this temperature.

Figure A-8. Nitrous oxide conversion (left) and hydrocarbon conversion (right) vs. averagereactor temperature at atmospheric pressure .Conditions: p = 1 bara, SV = 20,000 h–1, 1500 ppmvN2O, 100 ppmv NO, 100 ppmv NO2, 0.5% H2O, 2.5% O2, balance N2.

0

20

40

60

80

100

250 300 350 400 450 500

T (°C)

N 2O

con

vers

ion

(%)

500 ppm C3H84500 ppm CH4decomposition 0

20

40

60

80

100

200 250 300 350 400 450 500

T (°C)

hydr

ocar

bon

conv

ersi

on (%

)

500 ppm C3H8

4500 ppm CH4

ECN-C--00-087 23

ATTACHMENT B: BASIS OF DESIGN

The inlet conditions, physical constraints and the outlet requirements that have been used as basis forthe conceptual design of the DeN2O system, are summarised, thus forming the Basis of Design (BOD).Two process variants are considered:− high temperature direct decomposition of N2O, upstream the expander (high pressure) without

addition of reducing agent;− selective catalytic reduction (SCR) of N2O by addition of natural gas as a reducing agent, either

upstream (high pressure case) or downstream (atmospheric pressure case) the expander.

General tail gas compositionThe tail gas composition varies with each nitric acid plant being considered. In Table B-1 thefluctuation of the tail gas composition in the Netherlands is presented

Table B-1: Fluctuation of tail gas compositionComponent ConcentrationO2 1 - 4 vol.%H2O 0.2 - 1 vol.%NO 50 - 200 ppmNO2 50 -200 ppmN2O 1000 - 2000 ppmNH3 0 - 25 ppmN2 95 - 99 vol.% Based on these figures and in consultation with DSM and Hydro Agri, one typical tail gas compositionwas determined, thus forming a representative average composition for the Dutch nitric acid producingindustry. This composition is presented in Table B-2:

Table B-2: Tail gas compositionComponent Concentration[-] [vol.%] [ppm] [g/Nm3]O2 2.5 35H2O 0.5 5,000 4NO 1001

NO2 1001

N2O 0.2 1,500 2.9NH3 0N2 96.8 1193sulphuric components 0 0dust 0 0total 100.0 12351 Concentrations for NOx have been derived, assuming aDeNOx facility has already been integrated in the nitric acid plant.

For the conceptual design it has been assumed that for normal operation, fluctuations in tail gascomposition have no impact on the technical and economic feasibility results.

Inlet conditions and outlet requirements direct decompositionThe given inlet conditions and required outlet conditions determine the required performance of theDeN2O system, based on high temperature direct decomposition, and are presented in Table B-3.

Table B-3: Inlet conditions and outlet requirements direct decompositionProcess parameter ConditionsInlet flow [Nm3/h] 200,000Inlet pressure [bar(a)] 10Maximum allowable pressure drop1 [mbar] 200Inlet temperature [°C] 450Minimum required N2O reduction [%] 70

1 An increased pressure drop influences the power recovery and theproduction capacity of the production plant, thus decreasing the efficiency ofthe production plant. Based on figures provided by Hydro Agri, 200 mbarpressure drop will lead to a production and a power recovery loss (about 1MW), equivalent to 1 million NLG / year (reference: Minutes of Meeting 18-11-99 at Novem, Utrecht, doc. no. 7.2729-PC 99/16 dated 23 November1999).

24 ECN-C--00-087

Inlet conditions and outlet requirements SCR using natural gasThe given inlet conditions and required outlet conditions determine the required performance of theSCR DeN2O system and are presented in Table B-4 for both the atmospheric pressure case and thehigh pressure case.

Table B-4: Inlet conditions and outlet requirements SCR using natural gasProcess parameter Pressure case:

atmospheric high pressureInlet flow [Nm3/h] 200,000 200,000Inlet pressure [bar(a)] 1.11 10Maximum allowable pressure drop2 [mbar] 200 200Inlet temperature [°C] 100 350Maximum allowable outlet temperature3 [°C] - 360Minimum required N2O reduction [%] 70 701 During normal operation the outlet pressure of the expander is 1 bar(a). To enable tail gas flow through theSCR DeN2O system, an overpressure of about 0.1 bar(g) is required. This leads to a reduced efficiency of theexpander (see note 2);2 An increased pressure drop influences the power recovery and the production capacity of the production plant,thus decreasing the efficiency of the production plant. Based on figures provided by Hydro Agri, 200 mbarpressure drop will lead to a production and a power recovery loss (about 1 MW), equivalent to 1 million NLG /year (reference: Minutes of Meeting 18-11-99 at Novem, Utrecht, doc. no. 7.2729-PC 99/16 dated 23 November1999);3 A temperature increase of 7°C leads to 1% power recovery and production losses of the downstream expander.

ECN-C--00-087 25

Production plant informationFor process and cost calculation, a reference plant is selected resulting in the following starting points:

Site information:The DeN2O system will be located outside.

Location:atmospheric pressure case: clear area next to existing tail gas ducting (utilities present at battery limit).high pressure case: no typical location or existing facility has been selected, but for cost estimatingtypical retrofit factors will be used for low and high complexity for process integration.

Production information:Annual operation hours : 8,000 hours / year

Utilities:The following utilities are considered to be present at site: instrument air, electricity, nitrogen, boilerfeed water and steam.For the ‘natural gas’ case it is assumed that also natural gas is present at site. No separate unloadingand storage facilities for natural gas are therefore included in the design either the investment costestimate. For the LPG case transport and storage were included.For the natural gas, calculations have been used on basis of Dutch “Slochteren” natural gas. Theamounts of methane used in the catalytic tests, which were carried out using pure methane, have beenconverted to amounts of natural gas using the composition of “Slochteren” gas. On site however, alsoBritish and Industrial gas are being used. As for cost calculations a natural gas price is used correctedfor the heating value (and thus methane concentration), no major impact on the cost efficiency isexpected.

Start-up / shutdown:After a general shutdown of the nitric acid plant, start-up includes a gradual increase of the operatingtemperature, including the DeN2O system.

CO2 reduction efficiency (TEWI guidelines)The contribution of one kilogram N2O to the greenhouse effect has a factor 310 more impact than onekilogram of CO2. Using direct decomposition only, no extra CO2 is emitted except an extra emissiondue to increased power consumption and less efficient power regeneration by the existing expander.The CO2 reduction efficiency of a SCR DeN2O system using natural gas and a SCR DeN2O systemusing 500 ppmv LPG is slightly less efficient due to extra CO2 production and/or slip of the reducingagent. The overall positive effect of reducing 1 kilogram N2O on the CO2 reduction potential for thethree cases as defined above, is presented in Table B-5.

26 ECN-C--00-087

Table B-5: CO2 reduction efficiencyComponent CO2 equivalent

[kg CO2 /kg comp]

Inlet per

100 kg N2O

[kg component]

Outlet per

100 kg N2O

[kg component]

CO2 reduction

[kg CO2]

1) Direct Decomposition (high pressure case)N2O 310 100 30 21,700Overall CO2 reduction per 100 kg N2O at inlet 21,700Net CO2 reduction per kg N2O reduced [kg CO2 / kg N2O] 310

2) SCR using natural gas (high pressure case)N2O 310 100 25 23,250CO2 1 0 75 - 75CH4 21 36 4 - 84Overall CO2 reduction per 100 kg N2O at inlet 23,091Net CO2 reduction per kg N2O reduced [kg CO2 / kg N2O] 308

3) SCR DeN2O system using 500 ppmv LPG (atmospheric pressure case)N2O 310 100 30 21,700CO2 1 0 99 - 99propane - 33 0 0Overall CO2 reduction per 100 kg N2O at inlet 21,601Net CO2 reduction per kg N2O reduced [kg CO2 / kg N2O] 309

For cost calculation the reduction of CO2-equivalents is corrected for the extra CO2 emission due toincreased power consumption, including the recovery losses of the expander. A correction factor of0.61 kg CO2 per kWh was used (reference: “TEWI Richtlijnen”).

ECN-C--00-087 27

ATTACHMENT C: INVESTMENT AND EXPLOITATION COSTS

For direct decomposition, SCR using natural gas and an SCR DeN2O system using 500 ppmv LPG,the investment costs (with an accuracy of 30%) and the exploitation costs have been determined. Thebasic assumptions for this cost estimation are presented in this chapter.

Total investment costsTo determine the total investment, the required equipment has been estimated using supplierinformation and SE&C databases. Scale factors for non-equipment related hardware and services, likepiping, electrical, instrumentation, erection etc., obtained from existing facilities with comparablefunctions like DeNOx systems, and the assumptions summarised below, have been used to estimate thetotal investment cost:− utilities as mentioned in the BOD (Appendix B) are present at the battery limit;− clear construction area is available with utilities present at the battery limit;− the price for the catalyst is estimated to vary between NLG 50 and NLG 100 per kilogram

catalyst. For this study, an average price of NLG 75 per kilogram catalyst is used;− construction interest with an interest of 7% and a duration of 1 year.

The following items are excluded for the investment cost estimate:− area acquisition costs;− building and environmental permit acquisition costs;− site specific aspects and retrofits;− sewer system for discharge of rain water;− extra N2O, NOx, propane and CO analysers for system performance monitoring;− connection of instrumentation to the central DCS system;− production losses which might occur due to down-time during construction period.

For unforeseen items, a contingency factor of 15% has been used.

Site specific aspects - retrofitAs the high pressure cases are meant to be incorporated in an existing nitric acid production facility,the site specific aspects are essential to the cost estimate. To cover those specific aspects, SEC hasapplied retrofit factors.In accordance with the VROM methodology for application of retrofit factors2, the following range offactors is normally used:− low complexity for the process integration of the new facility : retrofit factor 1.5 - 2− high complexity for the process integration of the new facility : retrofit factor 2 - 3.5

In order to verify the applicability of these VROM retrofit factors, SEC has investigated the best avail-able retrofit factor for this DeN2O technology within the nitric acid industry, based on SEC’s experi-ence for DeNOx facilities within the Dutch chemical industry. This investigation has resulted into thefollowing:− average retrofit factor for DeNOx facilities within the Dutch nitric acid industry : 2− average retrofit factor for DeNOx facilities within the Dutch chemical industry : 2.1− range for retrofit factor for DeNOx facilities within the Dutch chemical industry : 1.7 - 2.9.

2 VROM publication: ‘Kosteneffectiviteit van milieumaatregelen in de industrie - beschrijving van de methodiek’,

deel A, Publicatiereeks Lucht & Energie, nr. 119, 1995

28 ECN-C--00-087

Taking into account the results of this investigation, a minimum retrofit factor of 1.5 (rise factor 50%)and maximum retrofit factor of 3 (rise factor 200%) have been used to determine the minimum andmaximum Total Investment Costs.

Exploitation costThe exploitation cost and the cost effectiveness of the DeN2O reactor system has been set upaccording to the VROM method3 and has been based on the following assumptions:

Capital cost:− the depreciation is calculated using the Capital Recovery Factor method (“annuïteitenmethode”);− the interest rate is 7%;− the depreciation period for the electro-mechanical part of the DeN2O system (including posts like

Instrumentation & Control, Engineering services etc., but excluding Catalyst and Civil) is 10years;

− the depreciation period for the civil part of the DeN2O system is 30 years;− the life time of the ECN catalyst is not yet known, but estimated to maintain an optimum

performance for at least 5 years. Therefore, the depreciation period for the catalyst part of theDeN2O system (including Installation Cost, other services excluded) is 5 years.

Operational cost - fixed:− the yearly cost for maintenance for electro-mechanical equipment is estimated using SE&C

experience, at a percentage of 3% of the electro-mechanical TIC;− the yearly cost for maintenance for civil constructions is estimated using SE&C experience, at a

percentage of 1% of the civil installation TIC;− as the DeN2O reactor system is automatically operating, the cost for personnel is estimated to

amount maximum half an operator man year;− insurance and general cost is set to 2% of the TIC.

Utility consumption:− the used electrical power, mainly originating from the operating system, recuperators and air

blowers during start-up, is incorporated at a price of NLG 0.15 per kWh;− the use of LPG is incorporated at a price of NLG 0.75 per kg, including delivery at site by truck,

excise and taxes4;− natural gas consumption is incorporated at a price of NLG 0.25 per Nm3;− other utilities, like water and waste disposal, are not applicable.

Annual Cost and Cost EfficiencyIn Table C-1 an overview is presented of the use of utilities and revenues for all cases as basis forcalculating the annual costs.Table C-1: Overview use of utilities and revenuesUtilities Use / Revenues Units Direct

DecompositionSCR usingnatural gas

SCR using500 ppmv LPG

Retrofit factor - 1½ 1½-3 -Personnel man / year - 0.5 0.5Natural gas Nm3 / year - 3,200,000 9,600 3LPG ton / year - - 1,544Electricity kWh / h - 40 40Pressure drop1 mbar 40 160 120Steam production2 ton / h - 4.7 -

3 VROM publication: ‘Methodiek Milieukosten’, Publicatiereeks Milieubeheer, nr. 1994, Mei 1995;4 CBS, VEMW magazines, Petroleum - Economist September 1999

ECN-C--00-087 29

Ad 1: The annual costs for operating the DeN2O system are determined by the exploitation costsincluding losses due to a decreased efficiency of the existing expander of the nitric acid plant.Based on figures provided by Hydro Agri, 200 mbar pressure drop will lead to a recovery lossof 1 MW and production losses, equivalent with 1 million NLG / year (ref. Minutes ofMeeting 18-11-99 at Novem, Utrecht, doc. no. 7.2729-PC 99/16 dd. 23 November 1999).Recovery losses for the DeN2O systems are linear extrapolated from this base figure.

Ad 2: The steam generated is saturated steam at low pressure (low quality steam). Production costsfor this type of steam are in the range of NLG 20-28 per ton steam. However, as the nitric acidplants already have a surplus steam production, the steam has to be delivered to externalconsumers. Therefore, for this study a net revenue of NLG 10 per ton steam produced is used.In practise, the actual revenues might vary from NLG 0 up to NLG 20 per ton steam produced.

Ad 3: For start-up only, natural gas consumption for the SCR DeN2O system using 500 ppmv LPGhas been based on 4 start-ups per year, each start-up lasting for 4 hours.

Table C-2: Exploitation costs and cost efficiency (base cases)Post Unit Direct

DecompositionSCR using natural gas

RF = 1½ RF = 3

SCR using500 ppm

TIC 106 NLG 5.4 14.7 24.4 11.2Exploitation CostCapital Cost 103 NLG / year 752 2,125 3,458 1,597Fixed Operation 103 NLG / year 278 831 1,348 650Energy2 103 NLG / year - 993 993 1,329Recovery Losses 103 NLG / year 200 800 800 600Steam Revenues 103 NLG / year - 376 376 -Annual Cost 103 NLG / year 1,230 4,313 6,163 4,176Efficiency N2O % 70 75 75 70Reduction N2O ton / year 3,248 3,480 3,480 3,248Cost per ton N2O NLG / ton N2O 379 1,239 1,771 1,286CO2 equivalents 103 ton / year 1 1,006 1,068 1,068 1,001Cost per ton CO2 NLG / ton CO2 1.2 4.0 5.8 4.21 Including a correction factor of 0.61 kg CO2 per kWh according to “TEWI Richtlijnen”.2 Including the costs of LPG and/or natural gas.

Fluctuation of exploitation costTo get an impression of the possible fluctuations of the exploitation costs, for all base cases thefollowing characteristics have been varied to check the impact on the cost and cost efficiency:− impact of catalyst price of NLG 50 and NLG 100 per kilogram;− impact of fluctuation of the natural gas price. Costs for natural gas consumption up to 170,000

Nm3 per year are around NLG 0.50 per Nm3. Costs for natural gas consumptions around3,000,000 Nm3 per year vary between NLG 0.20 and NLG 0.25 per Nm3;

− impact of fluctuation of the LPG price. Over the period 1992 up to 1999, a fluctuation of ± 20%in the average LPG price has been observed. Therefore, the impact of a minimum price of NLG0.60 per kg LPG and a maximum price of NLG 0.90 per kg LPG has been determined.Fluctuations of the LPG price in 2000 were found to lie in the same range;

− impact of an increased pressure drop.Per base case also an indication is presented of the cost consequences to improve the N2O reductioncapacity or to optimise the DeN2O system design.

Fluctuation direct decompositionBesides the general fluctuations, direct decomposition can be optimised to the following processconditions:− at an operating temperature of 500°C, the N2O conversion is increased up to 95% at a space

velocity of 45,000 h-1. No drastic increase in investment costs is expected;

30 ECN-C--00-087

− at an operating temperature of 400°C, a space velocity of 13,000 h-1 is required to maintain a N2Oconversion of 70%. This will have a drastic increase of both investment and catalyst costs.

Table C-3: Influence on Cost and Efficiency - direct decomposition caseOriginal Deviation Direct Impact on Cost Efficiency

[NLG / ton CO2]Base Case 1.2Catalyst Price NLG 75 / kg Catalyst NLG 50 / kg Capital Cost 1.1

Catalyst NLG 100 / kg Capital Cost 1.2Gas Price N.A.LPG Price N.A.Pressure drop 40 mbar Pressure drop 140 mbar Expander Losses 1.7Operating temperature 450°C Operating temperature 500°C N2O reduction 0.9

Operating temperature 400°C Capital Cost 2.2

Fluctuation SCR using natural gasThe effect of the general fluctuations on the cost efficiency are presented in Table C-4. Only theimpact on the low retrofit case (factor 1½) is presented:

Table C-4: Influence on Cost and Efficiency - SCR using natural gas (low retrofit)Original Deviation Direct Impact on Cost Efficiency

[NLG / ton CO2]Base Case 4.0Catalyst Price NLG 75 / kg Catalyst NLG 50 / kg Capital Cost 3.8

Catalyst NLG 100 / kg Capital Cost 4.1Natural Gas NLG 0.25 / Nm3 Natural Gas NLG 0.20 / Nm3 Energy Cost 3.8LPG N.A.Pressure drop 160 mbar Pressure drop 260 mbar Expander Losses 4.4

Fluctuation SCR DeN2O system using 500 ppmv LPGBesides the general fluctuations, the cost efficiency of the SCR DeN2O system using 500 ppmv LPGcan be influenced by:− increasing the operating temperature of the DeN2O reactor to an average of 380°C, thus

increasing the N2O conversion to 90%. The duty of the recuperator has to be increased slightly(about 10%);

− increasing the performance of the recuperator to an outlet temperature of the tail gas of 120°C instead of 130°C. The LPG use for making up the heat losses of the recuperator is then reducedfrom 500 to about 330 ppm.

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Table C-5: Influence on Cost and Efficiency - SCR DeN2O system using 500 ppmv LPG caseOriginal Deviation Direct impact on Cost Efficiency

[NLG / ton CO2]Base Case 4.2Catalyst Price NLG 75 / kg Catalyst NLG 50 / kg Capital Cost 4.1

Catalyst NLG 100 / kg Capital Cost 4.2LPG NLG 0.75 / Nm3 LPG NLG 0.60 / Nm3 Energy Cost 3.9

LPG NLG 0.90 / Nm3 Energy Cost 4.4Pressure drop 120 mbar Pressure drop 220 mbar Expander Losses 4.7Operating temperature 355°C Operating temperature 380°C N2O reduction 3.3Outlet temperature 130°C Outlet temperature 120°C Energy Cost 3.8

Direct decomposition at atmospheric pressureAn alternative for a SCR DeN2O system using 500 ppmv LPG could be direct decomposition atincreased temperatures. In stead of using the heat originating from the reaction of LPG with N2O, thetail gas can be heated directly by continuous combustion of natural gas in the in-line burner (seeFigure 6 on page 13). To minimise consumption of natural gas, also an optimised heat recovery systemis required.Compared with the SCR DeN2O system using 500 ppmv LPG, at the same space velocity (20,000 h–1)an operating temperature of ± 440°C is required to obtain 70% conversion of N2O. Assuming aminimum temperature approach of 30°C over the recuperator, about 250 Nm3/h natural gas is required.This will lead to a decrease of utility costs (about 33% of the costs for dosing 500 ppmv LPG).However, as the required operating temperature of the DeN2O reactor is almost 100°C higher than forthe SCR DeN2O system using 500 ppmv LPG, the duty of the recuperator increases with about 25% (±21 MW). This will result in increased investment costs.Based on these assumptions and the available test results of direct decomposition at atmosphericpressure, an extrapolation from the SCR DeN2O system using 500 ppmv LPG results in increaseddepreciation costs and decreased utility costs. Based on a N2O conversion ratio of 70%, cost efficien-cies in the range of 3.5-4.0 NLG/ton CO2 can be expected. Further improvement of the N2O conver-sion ratio up to 90% requires higher operating temperatures and increased reactor configurations andtherefore different requirements to the DeN2O process set-up. For a proper cost estimate, a separateconceptual design for direct decomposition at atmospheric pressure is required.