Catalysis Today 59 (2000) 417–422

Pilot plant performance of a SO2 to SO3 oxidationcatalyst for flue-gas conditioning

Esperanza Alvareza, Jesús Blancoa,∗, Carlos Knappa, Joaquın Olivaresb, Luis Salvadorba Instituto de Catálisis y Petroleoquımica, CSIC, Camino de Valdelatas s/n, Cantoblanco, 28049 Madrid, Spain

b Departamento de Ingenierıa Quımica y Ambiental, ESII, Universidad de Sevilla, Camino de los Descubrimientos s/n, 41092 Sevilla, Spain

Abstract

A honeycomb catalyst for the oxidation of endogenous SO2 from a coal-fired power-station flue-gas has been developed.The catalyst reached a SO2 to SO3 conversion of 60 vol.% after 200 h in operation at the pilot plant. When this catalyst isfurther treated for another 100 h at lab scale to complete its activation, a stable 80 vol.% conversion is obtained. The resultshave been used to design an industrial unit for flue-gas conditioning to improve the fly ash collection by the electrostaticprecipitator in a 220 MW coal-fired power plant. © 2000 Elsevier Science B.V. All rights reserved.

Keywords:Monolith; Sulfur dioxide oxidation; Flue-gas conditioning; SO2 abatement

1. Introduction

Emissions control techniques are some of the mainenvironmental concerns of energy supply based oncoal-fired power plants. As it is known, the sulfuroxides emitted from coal-fire combustion stations arecurrently controlled by wet lime/limestone scrubbersand spray dryers where the disposal of the wasteproducts is a major problem due to the lack of suitabledisposal sites.

Latterly, catalytic oxidation of sulfur dioxide tosulfur trioxide has emerged as one alternative to theabove mentioned processes. In these catalytic sys-tems, a useful SO3 by-product is obtained from whichvarious commercial products may be produced, suchas sulfuric acid and ammonium sulfate [1–3]. Anotherpromising application of the SO2 to SO3 catalytic ox-idation, in coal-fired power stations, is flue-gas condi-

∗ Corresponding author. Tel.:+34-1-5854-802;fax: +34-1-5854-789.E-mail address:jblanco@icp.csic.es (J. Blanco)

tioning to improve the performance of the particulateelectrostatic precipitator (ESP). Modern ESPs reachfly ash removal efficiencies of 99.9% complying withthe particulate current emission standards. Howeverwhen low-sulfur coal is used in order to fulfill thesulfur oxide emission regulations, inefficient opera-tion of these ESP unit arises, since in this situation,the SO3 concentration in the flue-gas is not enoughto adjust the fly ash resistivity. Although the amountof the required SO3 to increase the ESP’s efficiencydepends on various factors such as gas temperature orcoal ash content, the typical values for SO3 concen-tration usually vary within the range of 20–50 ppmv[4]. Most SO3 flue-gas conditioning systems cur-rently used, are based on catalytic conversion of SO2,supplied in liquid form or obtained by burning sulfur.The drawbacks of these units, besides the economicaland environmental issues, comprise the storage andhandling of those hazardous materials.

An original flue-gas conditioning method, whichovercomes these problems, has been proposed [5,6],where endogenous SO2catalytic oxidation to SO3

0920-5861/00/$ – see front matter © 2000 Elsevier Science B.V. All rights reserved.PII: S0920-5861(00)00307-2

418 E. Alvarez et al. / Catalysis Today 59 (2000) 417–422

takes place in such a way that the ESP requirementsare fulfilled. The design of an adequate installationvaries with the operational parameters of each spe-cific industrial plant to be retrofitted. The percentageof oxidized SO2 could be managed in a broad rangeof values depending on the portion of the flue-gasto be treated. In the case of total flue-gas condition-ing, Unland et al. [5] reported that a 75 MW powerplant would only need 6 vol.% of SO2 conversion.Whereas, according to the Epricon technology [6],in a 250 MW power plant where around 3% of thetotal flue-gas is treated, about 70 vol.% conversionwould be required. In both cases, the catalysts wereconformed as a parallel channel-type or honeycombstructure to avoid plugging by the fly ash, but as theywere prepared by impregnation, their deactivation dueto abrasion could be expected.

In the present study, an industrial-scale monolithiccatalyst based on vanadium pentoxide, potassium sul-fates and diatomaceous earth has been tested at pilotplant scale for the endogenous SO2 oxidation to SO3.The catalyst with honeycomb configuration has beenproduced by extrusion in order to minimize the abra-sion effect observed when impregnated catalysts areused. A 60 vol.% conversion of the SO2 to SO3 wouldallow the conditioning of a flue-gas, which contains700 ppmv of SO2 by treatment of about 3% of thetotal flue-gas from a coal-fired power plant.

2. Experimental

2.1. Catalyst preparation

The honeycomb catalyst was based on diatoma-ceous earth with potassium sulfate and ammoniummetavanadate as active-phase precursors. The compo-nents were intimately mixed before extrusion in orderto obtain a homogeneous paste following the methodreported elsewhere [7]. The catalyst was manufacturedby extrusion of a paste comprising all the compo-nents. The diatomaceous earth and potassium sulfatepowders were kneaded after adding the vanadium saltaqueous solution. The V/K atomic ratio was 3.5 with7.5 wt.% V2O5. The square-cell monoliths, after ac-tivation, had a cell density of 5 cells cm−2, pitch of0.38 cm, wall thickness of 0.08 cm, geometric surface

area of 720 m2 m−3, bulk density of 0.7 g cm−3 and acrushing strength of around 750 kg cm−2. The exter-nal monolithic dimensions were 95×95×500 cm3.

2.2. Catalytic activity test: pilot plant scaleexperiments

A schematic diagram of the pilot plant is shownin Fig. 1. The main body of the system was a cat-alytic reactor that contained a basket for the verticallyarranged monolith package. The pilot plant operatedusing combustion gases from a gas–oil burner with awell-defined composition and operating temperature.The head unit was a gas generator designed to ope-rate between 200 and 650 STP m3 h−1 of flue-gasin the range of 350–500◦C. These operational condi-tions were reached by setting the air–fuel ratio and theair-cooling water-recirculated device. A fly ash doserand a solid–gas mixer were placed upstream of the cat-alytic reactor. Two heated lined probes were connectedup and downstream of the reactor in order to withdrawaliquots for SOx analysis. A centrifugal fan locateddownstream of the bag-filter composed the tail unit.The pressure and temperature in the inlet/outlet of thereactor were monitored by a PC-program, along withthe unburned hydrocarbons, SO2 and O2 levels. Anadditional thermocouple was installed at 15 cm depthinside a selected centered monolith channel in orderto check the catalyst temperature. The required SO2inlet concentrations were obtained by sulfur carbidedoping of the gas–oil. The catalytic reactor design andcoupling accessories assured the isokinetic gas flowconditions at the entrance of the reactor.

2.3. Catalytic activity test: lab scale experiments

Catalytic activity measurements were carried out ina reactor working close to the isothermal axial pro-file, where carborundum fixed and isolated the mono-lith. A 98 wt.% sulfuric acid solution was placed at thereactor exit in order to avoid corrosion problems in-side the SO2 analyzer. Temperature and flow rate werecontrolled by two PID inlet/outlet reactor located andmassflow controllers, respectively. The gas supplierswere cylinders without significant percentages of im-purities and no fly ash dosage was added. The op-erating pressure was maintained at 120 kPa and SOx

analysis was continuously performed.

E. Alvarez et al. / Catalysis Today 59 (2000) 417–422 419

Fig. 1. Pilot plant scheme.

2.4. Analysis and characterization techniques

Textural properties of monolithic samples wereanalyzed in terms of total pore volume, mean porediameter and BET surface area. These measures co-rresponded to the fresh and used catalyst from the pilotplant performance. Total pore volume and mean dia-meter were determined using a CE Instrument Pascal140/240 porosimeter. The BET surface area was de-termined from the N2 adsorption/desorption isothermsat 77 K, which were carried out in a Carlo-ErbaSorptomatic Series 1800. In both techniques, sampleswere outgassed overnight at 140◦C. In the case ofN2 isotherm technique, samples were also outgassedto vacuum below 10−4 Torr to ensure a dry cleansurface.

The mechanical strength of the fresh and usedmonolithic channel wall was determined using aChatillon LTCM Universal Tensile Compression andSpring Tester with a cylinder test head of 1 mm dia-meter.

A Varian 3400 model GC coupled to a FPD wasused to monitor the inlet/outlet SO2 concentration atlab scale experiments.

The inlet/outlet SO2, O2 and unburned hydrocar-bons measures at the pilot plant scale were performedusing an ESSA ML9850 model ultraviolet fluores-cence analyzer.

The EPA Method 8 [8] was selected for SO3 sam-pling at the reactor outlet; the condensed sulfuric acidmist was measured by the barium–thorine titrationmethod.

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3. Results and discussion

3.1. Pilot plant scale experiment

The green catalyst must undergo an activation step,which requires SO2 and SO3 enrichment. In this acti-vation process, which has been previously studied [7],the SOx sorption gives rise to eutetic VK pyrosulfatecomplexes that were liquid under reaction conditions.The activation step could last from 60 to 100 h depend-ing on the SO2 and O2 concentration in the flue-gas.This activation step has been shown to be a crucialfactor to achieve satisfactory catalytic performance.

Moreover, interesting studies about the nature of themolten liquid active-phase have been carried out inorder to identify which species were responsible forthe catalytic activity [9].

In the pilot plant experiment, the activation of thecatalyst was carried out using the flue-gas accordingto the conditions previously selected [7]. The oper-ating conditions at the pilot plant are summarized inTable 1.

The concentration of unburned hydrocarbons wasconstant at 750 ppm (by vol.) and ca. 300 ppm forCO2, CO and NOx . Particulate matter was added ina concentration of about 15 g N m−3. The monolithicpieces were distributed in two beds of 50 cm heighteach, with a total catalyst volume of 180 l. The pres-sure drop through the catalytic bed was in the rangeof 22–26 mm of H2O m−1.

In Fig. 2, the outlet concentration of SO2, SO3 andSO2+SO3 against time in operation is plotted. As

Table 1Operational conditions (NTP) for the pilot plant scale testa

Average temperature (◦C) 470Pressure (kPa) 120GHSV (h−1) 3200LV (m s−1) 1.4AV (m h−1) 4.4

Gas composition (by vol.)SO2 700 ppmO2 3%H2O 7.5%N2 Balance

a GHSV: gas hourly space velocity; LV: linear velocity; AV:area velocity.

Fig. 2. Outlet concentration of SO2, SO3 and SO2+SO3 as afunction of time in operation of the monolithic catalyst for thepilot plant experiment.

can be observed, different physico-chemical processestook place during the course of the experiment.

During the first 10 h, a steeply decreasing valueof the outlet SO2 was observed, but SO3 was notobserved. The SO2 was transformed into SO3 andretained in the molten active-phase, starting the for-mation of VK-pyrosulfates, the active-phase. The SO2sorption/oxidation phenomena and negligible SO3desorption during this period merely corresponded tothe activation step of the green catalyst.

During the next 40 h in continuous operation, therate of the activation process slowed down and theoutlet SO2 concentration started to increase untila maximum value of about 440 ppm was reached.Meanwhile, the SO3 desorption was slightly takingplace. Therefore, as SO2 increased, only a smallamount of SO3 started to be released from the SO3saturated molten active-phase.

After 50 h in operation, the SO2 catalytic oxida-tion started to be the main process in progress asthe active-phase is more developed and less SO3 wasneeded to be transformed into VK-pyrosulfates. Dur-ing this period the SO2+SO3 outlet concentration wasgetting closer to the inlet SO2 value. Therefore, the ox-idation reaction rate increased as further active-phasewas formed from the simultaneous SOx absorptionprocess.

After 200 h in operation, most of the active-phase had been formed although the outlet SO2+SO3value was lower than 700 ppm, indicating that the

E. Alvarez et al. / Catalysis Today 59 (2000) 417–422 421

activation process had not been accomplished. How-ever, the SO2 to SO3 conversion achieved the requiredrange of 60 vol.% and the pilot plant experimentwas stopped. The analysis of the textural propertiesof the fresh and used catalyst showed significantchanges. TheSBET decreased from 26 in the freshsample to 2 m2 g−1 in the used one. A reduction ofthe total pore volume was also observed from 0.27to 0.18 cm3 g−1. Drastic change of the mean pore di-ameter was recorded: the 65 nm diameter in the freshsample shifted to 200 nm after reaction. The mechan-ical strength of the used catalyst steeply increased to750 from the value of 114 kg cm−2 of the brittle freshsample.

Data from this physical analysis did not informabout the actual situation when SO2 oxidation reac-tion was proceeding. As mentioned before, this typeof catalyst — denominated supported liquid phase(SLP) — presents the characteristics of having amolten active-phase under reaction conditions. Whenthe reactor was cooling down, the VK molten com-pounds precipitated and therefore, the porous structurechanged. Nevertheless, a general evaluation could bedrawn from the textural characterization results. Asa consequence of the SOx uptaking and simultane-ous melting processes during the activation step, thesmaller pores were filled up. Chemical changes ofthe VK-bearing species also produced variations onmechanical properties.

3.2. Lab scale experiment

As the activation process had not been completedduring the pilot plant run, because the SOx inlet/outletbalance was not closed, the catalyst activation wascontinued for a further 150 h at lab scale. The experi-ment was carried out using a 28 cm length 4 channelmonolith taken from that used in the pilot plant. Inthis run, neither fly ash nor water vapor was added.The remaining operational conditions were as similarto those used in pilot scale as the scaling-down pro-cedure allowed. The operating parameters are shownin Table 2.

The evaluation of the outlet SO2, SO3 andSO2+SO3 concentrations with time in operation isshown in Fig. 3. After 70 h, the SOx reached steadystate values of 175 ppm for SO2 and 540 ppm for SO3,indicating that the activation process was concluded.

Table 2Operational conditions (NTP) for the lab scale test

Average temperature (◦C) 470Pressure (kPa) 120GHSV (h−1) 3200LV (m s−1) 0.4AV (m h−1) 4.4

Gas composition (by vol.)SO2 700 ppmO2 3%N2 Balance

At this point, 80 vol.% SO2 to SO3 conversion wasobtained, which remained constant until the reactionsystem was shut down 80 h later.

The textural and mechanical properties were alsoanalyzed. TheSBET and total pore volume did notchange from 2 m2 g−1 and 0.18 cm3 g−1, respectively,for the used catalyst at the pilot plant scale experi-ment. Nevertheless, the mean pore diameter increasedfrom 200 to 600 nm. As the activation process had notfinished after the 200 h in operation at the pilot plantexperiment, the small increase in SOx uptake duringthe lab scale run was thought to cause the alterationin the porous structure of the catalyst.

The experimental results for both pilot plant andlab scale runs, which summarize the catalytic activityperformance, are plotted in Fig. 4. The SO2 to SO3conversion versus time in operation is represented and

Fig. 3. Outlet concentration of SO2, SO3 and SO2+SO3 as afunction of time in operation of the monolithic catalyst for the labscale experiment.

422 E. Alvarez et al. / Catalysis Today 59 (2000) 417–422

Fig. 4. Catalytic activity in terms of SO3 percentage as a functionof time in operation for both pilot plant and laboratory experiments.

the dotted line indicated the break point between thedifferent operating-scale systems. In spite of variationsin some of the operational parameters, such as linearvelocity and the absence of fly ash and water vaporat lab scale conditions, a continuity of the catalystperformance was observed. In Fig. 4, at least threedifferent steps can be distinguished:1. In the first 50 h, the SO2 to SO3 conversion

smoothly increased along with the formation ofthe active-phase through capturing the SOx by thestarting VK compounds. Moreover, these newlyformed compounds must also be transformed intoa molten material.

2. Between 50 and 300 h, the outlet SO3 concentra-tion steeply increased at different rates. In this step,the amount of the molten VK-pyrosulfates wasaugmented on the support and consequently, theSO2 to SO3 catalytic reaction increased. The phasetransformation, which accompanied the chemicalchange, might also alter the porous structure be-cause the liquid layer in motion was occupyingthe inter and intraparticulate space. Due to thismovement, the starting VK compounds would beexposed to further activation processes. This factmight explain why the released SO3 did not havea linear trend.

3. The steady state condition for the catalytic activitywas achieved after 300 h in operation, where theactive-phase was completely formed and neitherSO2 nor SO3 was retained in the catalyst.

4. Conclusions

An industrial-scale honeycomb catalyst based onvanadium–potassium–diatomaceous earth was manu-factured. A pilot plant was designed and built to simu-late the real conditions of a coal-fired power plant.

The catalytic activity in the pilot plant scale ex-periment showed satisfactory results. The 60 vol.%requirement of SO2 to SO3 conversion was achievedafter 200 h in operation.

A further experiment carried out at lab scale con-ditions showed a residual catalytic activity and whenthe activation process finished, a stable 80 vol.% SO2to SO3 conversion was achieved.

The results have been used to design an industrialunit for flue-gas conditioning to improve the perfor-mance of the electrostatic precipitators in the 220 MWcoal-fired power plant at Puertollano (Spain) run byCompañıa Sevillana de Electricidad-Endesa Company.

Acknowledgements

This work has been sponsored by the EuropeanCommunity for Steel and Coal and public andprivate Spanish institutions related to the electricenergy generation field and coal mining (Endesa,Ocicarbon and Compañıa Sevillana de Electrici-dad). The authors are thankful to CICYT, ECSCand CSE-ENDESA-OCICARBON for their financialsupport.

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