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Article on Modern Nitric Acid Production Process

By July 9, 2015

Modern nitric acid plants are designed mostly according to three nitric acid processes: mono medium pressure, mono high pressure and dual pressure. Medium pressure at about 4 – 6 bar will be preferred where a high ammonia efficiency is decisive for the plant economics. High pressure of more than 8 – 12 bar is an advantage with regard to the absorption performance. The dual pressure process offers both of these advantages as this process combine a medium pressure combustion with the use of an absorption at higher pressure.

Nitric acid plants are an important source of nitrous oxide (N2O). The concentration of N2O in the tail gas of nitric acid plants ranges between 800 – 2,000 ppm. Some countries have already limited the emissions of N2O to the atmosphere while many other countries will introduce restrictions in the future. Thus, emission reduction at nitric acid plants is also exercising the minds of suppliers of technology to nitric acid plants.

Mono Medium Pressure Process

In this process, the air required for burning the ammonia is supplied by an uncooled air compressor. The compressor set can be designed either as an inline train configuration or preferably as a bull gear type with an integrated tail gas turbine. The operation pressure is governed by the maximum final pressure obtainable in an uncooled compressor, i.e. 4 – 5 bar abs. in the case of radial compressor and 5 – 6 bar abs. with axial flow compressors. Plant capacities of up to 500 mtpd of nitric acid (100%) can be realized using a single ammonia combustion unit and one absorption tower. Due to the process pressure, higher capacities of up to 1,000 mtpd are feasible if a second absorption tower is used. The medium pressure process is the process of choice when maximum recovery of energy is required. The air compressor is usually driven by a tail gas expansion turbine and a steam turbine, the steam being generated within the plant. If the credit for exported steam is high then the compressor train can be driven by a high voltage synchronous or asynchronous electric motor rather than a steam turbine so that all the steam generated can be exported.

With this plant type, it is possible to produce one type of nitric acid with a max. concentration of 65% or two types of acid with different concentrations, e.g. 60% and 65%, while the NOx content in the tail gas can be reduced by absorption to less than 500 ppm. The NOx cotent has to be further reduced to the required value by selective catalytic reduction using a non-noble-metal catalyst and ammonia as the reducing agent. The catalyst gauze only need to be changed once every 6 months due to the low burner load.

The medium-pressure process is characteristic of a high overall nitrogen yield of about 95.0 % or 95.2% in conjunction with the tail gas treatment process, a low platinum consumption and a high steam export rate.

Mono High Pressure Process

In the high pressure process, a radial multistage compressor with an intercooler section is used to compress the process air to a final pressure of 8 – 12 bar abs. Preferably, the bull gear type with integrated tail gas turbine is selected but alternatively an inline machine can be used. The compressor may be driven by a steam turbine or an electric motor. Due to the higher pressure, all equipment and piping can be of a smaller size and only one absorption tower is required. The arrangement of all equipment is very compact so that the building for the machine set and the burner unit can be kept small, but this does not pose a problem for maintenance work. This type of plant is always recommended when a quick capital return is desirable.

Plant capacities between 100 mtpd (100%) and 1,000 mtpd of nitric acid can be realized. The achieved acid concentrations of up to 67% are slightly higher than with the medium pressure process. Two or more product streams with different concentrations are likewise possible.

The NOx concentration in the tail gas can be less than 200 ppm by absorption alone. If lower NOx is required, an additional catalytic tail gas treatment unit can be easily integrated. The nitrogen yield attained by the high pressure process is in the order of 94.5%.

For capacity below 100 mtpd, a low capital investment cost may be much more preferable to an optimum recovery of energy. In such a case, a simplified process, which does not include tail gas and steam turbines, can be used. The heating of the tail gas downstream of the absorption section is not necessary. Special design concepts for the process gas cooler unit, condenser and bleacher reduce the cost even further. The operating pressure required by such a process depends on the maximum permissible NOx content in the tail gas. If the specified NOx limit is below 200 ppm, a pressure of at least 7 bar abs. must be considered, depending on the cooling tower water temperature. Lower NOx values of less than 50 ppm can be achieved, if required, by integrating a catalytic tail gas treatment process.

Dual Pressure Process

The dual pressure process is developed to accommodate even more stringent environmental pollution control requirements. The process air is compressed to a final pressure of 4 – 6 bar abs. The NOx gas from the ammonia combustion unit is cooled in a heat exchange train, producing steam and preheating tail gas, and then compressed to 10 – 12 bar abs. in the NOx compressor. The final pressure is selected so as to ensure that the absorption section is optimized for the specified NOx content of the tail gas and that the compressors, driven by a steam turbine, can be operated using only the steam generated in the process gas cooler unit while ensuring that some excess steam will always be available in order to guarantee steady operating conditions at all times. Alternatively the compressor set can be driven by either a high voltage asynchronous or synchronous motor and the steam generated can be exported. The dual pressure process combines in an economical way the advantages of the low pressure in the combustion section and the high pressure in the absorption section. Plant capacities of up to 1,500 mtpd of nitric acid (100%) can be achieved in a signle-train configuration. The machine set can be designed either as an inline-shaft set or optically as a bull gear type unit with integrated air / NOx compression and tail gas expander stages. The inline machine concept is favorable in the case of high plant capacities in excess of 1,100 mtpd up to 2,200 mtpd or if an increased energy export is required. For plant capacities in excess of 1,500 mtpd the ammonia combustion and process gas cooler unit needs only to be duplicated.

The nitrogen yield in a plant of this type is more than 96% with a NOx content in the untreated tails gas of less than 150 ppm (vol.) being possible. The NOx content may be further reduced to less than 50 ppm, if required, by selective catalytic reduction using non-noble-metal catalyst and ammonia as the reducing agent. Acid concentration of more than 68% can be achieved. Two or more product streams with different concentrations are also possible. In addition, external streams of weak nitric acid (various concentrations) can be processed if required. Due to the low burner load, the catalyst gauze can remain in the burner for an operating period of 6 – 8 months or longer before being partially or completely replaced

Typical Consumption Figures of Modern Nitric Acid Plants

The following chart is the comparison of typical consumption figures for steam turbine driven and inline compressor set nitric acid plants, per ton of nitric acid (100%), with NOx content in the tail gas of less than 50 ppm.

Plant Type Medium Pressure Process High Pressure Process Dual Pressure Process
Operating Pressure (abs.) 5.8 bar 10.0 bar 4.6 / 12.0 bar
Ammonia 284 kg 286 kg 282 kg
Electricity 9.0 KWH 13.0 KWH 8.5 KWH
Platinum, Primary Losses 0.15 g 0.26 g 0.13 g
Platinum losses after recovery 0.04 g 0.08 g 0.03 g
Cooling water (Δt = 10 k), including water for steam turbine condenser 100 t 130 t 105 t
Process Water 0.3 t 0.3 t 0.3 t
LP heating steam, 8 bar, saturated 0.05 t 0.20 t 0.05 t
HP excess steam, 40 bar, 450 ºC 0.76 t 0.55 t 0.65 t

Leading Technology Licensors

  • Weatherly, a company of Chematur Engineering Group of Sweden;
  • ESPINDESA, a company of Técnicas Reunidas of Spain;
  • Borealis of Austria;
  • Uhde (now ThyssenKrupp Industrial Solutions) of Germany;
  • MECS Technology of USA;
  • KBR of USA;
  • Technip of France.

Cost and Cost Effectiveness for Emission Control

Nowadays a big challenge to all nitric acid plants is emission control. The following table identifies the costs and cost effectiveness calculations included in EPA’s ACT document, which was developed for three model plants that cover most of the range of U.S. nitric acid plants.

As shown, cost effectiveness ($/ton of NO, removed) varies from 576-5297/ton for extended absorption, to 5507-5715/ton for NSCR, to $231-$305/ton for SCR. In all cases, cost effectiveness improves with plant size. NSCR is considerably less cost effective thin extended absorption largely because of NSCR’s higher utility costs and lack of any nitric acid recovery credit, which exists in the case of extended absorption.

Phoenix Equipment sells second hand nitric acid plants that are immediately available for purchase and relocation. Buying a used plant can save you significant capital and drastically shorten the time required to build a plant. Here are some nitric acid plants we currently have for sale:

If you are interested in learning about any of the above plants, please contact Edward Zhang (Tel: 732-520-2187; email:

Posted in: Uncategorized

Syngas and Hydrogen Production Plants

By June 16, 2015

Synthesis Gas (Syngas) / Hydrogen Production from Natural Gas

Syngas, or synthesis gas, is a fuel gas mixture consisting primarily of hydrogen and carbon monoxide. Syngas is usually a product of gasification of natural gas, coal, petroleum coke, biomass, municipal wastes, liquid petroleum gas, naphtha and refinery off-gas. Syngas is mainly used as fuel to generate electricity and as feedstock to produce various chemicals such as ammonia, methanol, methane, dimethyl ether, etc. The hydrogen content in syngas, after purification, has broad applications in chemical, petrochemical and oil refinery industries. Typical examples are hydrogenation and hydrotreating processes.
Modern production of syngas integrates two process sections: steam methane reforming (SMR) of hydrocarbon feedstock and separation of hydrogen from carbon monoxide. Below is the process diagram illustrating the production using natural gas, the most common feedstock to make syngas.

In the reforming reaction, natural gas is mixed with steam, heated to over 1,500 degrees Fahrenheit, and reacted with nickel catalyst to produce hydrogen (H2) and carbon monoxide (CO).

CH4 + H2O < ——– > 3H2 + CO

To produce additional hydrogen, CO from the reforming reaction interacts with steam in the water gas shift reactor.

CO + H2O < ——– > CO2 + H2

A reformer is the core of this process. The gas-and-steam mixture travels down into reformer tubes that hang in vertical rows surrounded by gas burners that heat the mixture. The reformer tubes are full of nickel catalyst, which triggers a reaction, causing the methane in natural gas to react with water vapor to form hydrogen, carbon monoxide, and carbon dioxide.

Additional hydrogen is created in the water gas shift reactor. The water gas shift reactor is filled with an iron-chrome based catalyst that causes steam (H2O) to break into oxygen and hydrogen. The hydrogen moves through the plant, while the oxygen joins carbon monoxide from the furnace (reforming reaction) and becomes carbon dioxide (CO2).

If high purity hydrogen is the main product, raw syngas will be treated by pressure swing adsorption (PSA) section to filter out remaining traces of carbon monoxide, carbon dioxide, steam, and methane from the hydrogen.

In the PSA process, raw syngas is pumped into a cylinder at pressure. The cylinder contains beads of adsorbent material. The impurities in the syngas, such as carbon dioxide, are adsorbed onto the internal surfaces of the adsorbent beads, leaving hydrogen in the vessel, most of which is removed as purified hydrogen product. Pressure in the cylinder is reduced, releasing the impurities from the adsorbent material. A small amount of product hydrogen is used to flush the waste gas through an exhaust port, preparing the vessel for another production cycle.

The purified hydrogen from PSA can be delivered to a variety of customers in the electronics, biomass and steel industries. More importantly, the hydrogen product is a necessary supply to refineries, which require large volume for hydrocracking, dearomatisation and desulphurisation processes. In addition, the purified hydrogen is the imperative feedstock for chemical plants, which have processes such as methanol and ammonia synthesis.

At present, the most widely used and cheapest method for hydrogen production is the steam reforming of methane (natural gas). This method includes about half of the world hydrogen production, and hydrogen price is about US$ 7 / GJ.

Below listed are several well-known companies which provide process technology and engineering of syngas / hydrogen production on the basis of steam reforming. The number in parenthesis shows the units installed worldwide approximately.

* Foster Wheeler (100)

* Lurgi (30)

* Linde (120)

* KBR (100)

* Uhde (60)

* Praxair (140)

* Haldor Topsoe (40)

* Technip (240)

* CB&I (175)

Numerous hydrogen PSA units have been installed worldwide. Leading technology licensors include:

* Air Products & Chemicals




* Lurgi

A modern syngas / hydrogen plant typically has capacity range from 5 MMSCFD to 200 MMSCFD. For a 90 MMSCFD plant, the capital cost to build such a plant is estimated at approximately US$ 55 million nowadays. Average hydrogen production cost is US$ 1.60 per 1,000 standard cubic feet, or US$ 53 million per year to operate a plant of this capacity. This estimate is based on the following utility consumption.

Utilities consumption per 1,000 ft3 of contained hydrogen

  • Natural gas feed, million Btu LHV 317
  • Natural gas fuel, million Btu LHV 126
  • HP export steam, lbs 90
  • Boiler feedwater, lbs 120
  • Power, kWh 52
  • Cooling water, gal 8

Phoenix Equipment sells second hand hydrogen (hyco) and synthesis gas (syngas) plants, as well as hydrogen purification plants that are immediately available for purchase and relocation. Buying a used plant can save you significant capital and drastically shorten the time required to build a plant. Here are some hydrogen and syngas plants we currently have for sale –

If you are interested in learning about any of the above plants, please contact Edward Zhang (Tel: 732-520-2187; email:

Posted in: Chemical Plants, Uncategorized

Article on Large Scale Production of Methanol

By January 6, 2015

The capacity of methanol plants is increasing to reduce investments, taking advantage of the economy of scale. The capacity of a world scale plant has increased from 2500 MTPD a decade ago to about 5000 MTPD today. Even larger plants up to 10,000 MTPD or above are considered to further improve economics and to provide the feedstock for the Methanol-to-Olefin (MTO) process. A methanol plant with natural gas feed can be divided into three main sections. In the first part of the plant natural gas is converted into synthesis gas. The synthesis gas reacts to produce methanol in the second section, and methanol is purified to the desired purity in the tail-end of the plant.

The capital cost of large scale methanol plants is substantial. The synthesis gas production including compression and oxygen production when required may account for 60% or more of the investment. In many plants today either tubular steam reforming or two-step reforming (tubular steam reforming followed by autothermal or oxygen blown secondary reforming) is used for the production of synthesis gas. However, stand-alone Autothermal Reforming (ATR) at low steam to carbon (S/C) ratio is the preferred technology for large scale plants by maximising the single line capacity and minimising the investment. ATR combines substoichiometric combustion and catalytic steam reforming in one compact refractory lined reactor to produce synthesis gas for production of more than 10,000 MTPD of methanol. The ATR operates at low S/C ratio, thus reducing the flow through the plant and minimising the investment. The ATR produces a synthesis gas well suited for production of both fuel grade and high purity methanol. The design of the methanol synthesis section is essential to ensure low investment. The optimal design and the choice of operating parameters depend on the desired product specification. In many plants Boiling Water Reactors (BWR) are used. The use or incorporation of adiabatic reactors may be advantageous.

Several reforming technologies are available for producing synthesis gas:

  • One-step reforming with fired tubular reforming
  • Two-step reforming
  • Autothermal reforming (ATR)

In one-step reforming, the synthesis gas is produced by tubular steam reforming alone (without the use of oxygen). This concept was traditionally dominating. Today it is mainly considered for up to 2,500 MTPD plants and for cases where CO2 is contained in the natural gas or available at low cost from other sources.

The synthesis gas produced by one-step reforming will typically contain a surplus of hydrogen of about 40%. This hydrogen is carried unreacted through the synthesis section only to be purged and used as reformer fuel. The addition of CO2 permits optimization of the synthesis gas composition for methanol production. CO2 constitutes a less expensive feedstock, and CO2 emission to the environment is reduced. The application of CO2 reforming results in a very energy efficient plant. The energy consumption is 5–10% less than that of a conventional plant. A 3,030 MTPD methanol plant based on CO2 reforming was started up in Iran in 2004.

The two-step reforming process features a combination of fired tubular reforming (primary reforming) followed by oxygen-fired adiabatic reforming (secondary reforming). By combining the two reforming technologies, it is possible to adjust the synthesis gas to obtain the most suitable composition. The balance required to obtain a desired value of M depends on the natural gas composition. The heavy gas requires more steam reforming and less oxygen compared to the requirements for lean gas. The same is true for gas containing CO2. The secondary reformer requires that the primary reformer is operated with a significant leakage of unconverted methane (methane slip). Typically 35 to 45% of the reforming reaction occurs in the tubular reformer, the rest in the oxygen-fired reformer. As a consequence the tubular reformer is operated at low S/C ratio, low temperature and high pressure. These conditions lead to a reduction in the transferred duty by about 60% and in the reformer tube weight by 75 to 80% compared to one-step reforming. The two-step reforming lay-out was first used in a 2400 MTPD methanol plant in Norway. This plant was started up in 1997. A 5000 MTPD plant based on similar technology was started up in Saudi Arabia in 2008.

Autothermal reforming (ATR) features a stand-alone, oxygen-fired reformer. The autothermal reformer design features a burner, a combustion zone, and a catalyst bed in a refractory lined pressure vessel. The burner provides mixing of the feed and the oxidant. In the combustion zone, the feed and oxygen react by sub-stoichiometric combustion in a turbulent diffusion flame. The catalyst bed brings the steam reforming and shift conversion reactions to equilibrium in the synthesis gas and destroys soot precursors, so that the operation of the ATR is soot-free. The catalyst loading is optimized with respect to activity and particle shape and size to ensure low pressure drop and compact reactor design. The synthesis gas produced by autothermal reforming is rich in carbon monoxide, resulting in high reactivity of the gas. The synthesis gas has a module of 1.7 to 1.8 and is thus deficient in hydrogen. The module must be adjusted to a value of about 2 before the synthesis gas is suitable for methanol production. The adjustment can be done either by removing carbon dioxide from the synthesis gas or by recovering hydrogen from the synthesis loop purge gas and recycling the recovered hydrogen to the synthesis gas. When the adjustment is done by CO2 removal, a synthesis gas with very high CO/CO2 ratio is produced. This gas resembles the synthesis gas in methanol plants based on coal gasification. Several synthesis units based on gas produced from coal are in operation, this proves the feasibility of methanol synthesis from very aggressive synthesis gas. Adjustment by hydrogen recovery can be done either by a membrane or a PSA unit. Both concepts are well proven in the industry. The synthesis gas produced by this type of module adjustment is less aggressive and may be preferred for production of high purity methanol. The first ATR in operation in commercial scale at H2O/C ratio of 0.6 was an industrial demonstration in South Africa in 1999. The first commercial plant was started in Europe in 2002. Very large units were started in South Africa in 2004 and in Qatar in 2006. A methanol plant with a single line capacity of 10.000 MTPD is in the engineering phase for start-up in Nigeria in 2012.

State-of-the-art ATR and two-step reforming for syngas production are compared for large scale (10,000 MTPD) fuel grade methanol plants. The same natural gas composition and plant inlet pressure has been used for all calculations. The methanol synthesis takes place at 100 bar in a standard loop with parallel BWR installed with adiabatic top layers. Typical operating parameters for large scale plants in commercial operation have been selected. The module of the synthesis gas to the loop has in both cases been adjusted to 2.05. In the two-step reforming case this is done be adjusting the exit temperature of the primary reformer. For the ATR case hydrogen recovered from the loop purge gas is recycled to the synthesis gas. The two-step case has been designed with a loop carbon efficiency of 95% by setting the loop purge.

Mainly due to the low S/C ratio, the synthesis gas in the ATR-based case has a high CO/CO2 ratio and is lean in hydrogen. This makes the gas more reactive for MeOH synthesis and requires less recycling in the loop than with two-step reforming. The high loop carbon efficiency (95%) in the two-step reforming case requires a relatively high recycle ratio. However, decreasing the loop efficiency would further reduce the single line capacity, and is not seen as an attractive choice. The factors listed above have significant impact on the equipment size in the loop. For the recycle compressor the power (or size) increases by about 70% for the two-step case compared to the ATR case. Likewise, the size of other equipment in the loop (reactors, feed/effluent heat exchanger, piping, etc.) increases considerably with the two-step synthesis gas generation system.

ATR at low S/C-ratio is the preferred technology for large scale methanol production at low cost as illustrated above. Several plants for production of synthesis gas for various applications operate today at S/C=0.6. Further reduction to S/C=0.4 would increase the single line capacity.

The introduction of heat exchange reforming (HTER) either upstream (HTER-s) or in parallel with the ATR (HTER-p) has the potential to further reduce the costs. In both cases, the methanol module increases and the specific oxygen consumption is reduced. In the series scheme, all the feed gas passes through the HTER-s and the ATR. This results in a high feed conversion even at a low S/C ratio. The required heat for the endothermic reaction is provided by the combined stream of the ATR effluent and the stream leaving the catalyst bed of the HTER-p. This scheme has a higher single line capacity than stand alone ATR and the series scheme. However, a higher S/C ratio in the HTER-p is required than in the ATR to obtain the same feed conversion.

Further development of the methanol synthesis technology has the potential for reducing the overall plant cost. A large part of the cost arises from the reactor itself and from the recycle loop including the reactor feed/effluent heat exchanger typically used. It has been demonstrated that a higher conversion can be obtained by adjusting the operating conditions to condense the methanol in the reactor itself. Methanol is removed from the gas phase allowing a higher conversion of the reactants according to reactions 1-3. The once-through conversion can be increased to above 90%, thus avoiding the need for reactant recycle. Process economic evaluations indicate that the condensing methanol technology may reduce the overall investment by 10% or more for a 10,000 MTPD fuel grade methanol plant


To learn more about Phoenix Equipment and our Methanol Plants for sale, visit our Plants Page!

1500 TPD Methanol Plant

190 TPD Methanol Plant

1300 TPD Methanol Synthesis Loop

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A Look at Aurora Filters

By September 24, 2014

Aurora Filters has developed a line of portable Filter Systems that combine a single plate vacuum filtration device with a jacketed glove box. This design allows for the filtration, optional drying, discharging and packaging of product in an enclosed, inert environment, and ensures product integrity and safe working conditions. AURORA® FILTERS offers a range of Models to suit specific needs and ranges of product being produced. The Model Numbers correspond to the size of the filter, the material that it is manufactured from and the style of flanges used for the process nozzles. The P-Series Models are designed for +/- 14.7psi (+/- 1bar) pressure throughout the filter. Typical applications are the kilo lab and pilot plant production of pharmaceutical bulk actives, intermediates and fine chemicals that are sensitive to the atmosphere, or are produced in volatile solvents. (View PDF) (more…)

Posted in: Aurora Filters, Chemical Plants, Filters, Fine Chemicals Filters, News, Nutsche Filters, Pharmaceutical Equipment, Portable Filter Systems, Process Equipment, Used Process Equipment

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Unused (3) Alfa Laval Compabloc Heat Exchangers for Sale at Phoenix Equipment

By September 5, 2014

Two 2,063 Sq Ft and One 2,751 Sq Ft Unused Alfa Laval Compabloc 316L Stainless Steel Heat Exchangers are for sale at Phoenix Equipment. These unused Model CPL75-V-300 and CPL75-V-400 Compabloc Plate Heat Exchangers are top of the line and utilize Alfa Laval’s breakthrough technology that allow you to operate with a wide range of aggressive media and at high temperatures and pressures.


Stock #9200: Unused 2,751 Sq Ft Alfa Laval Compabloc 316L Stainless Steel Plate Heat Exchanger, Model CPL75-V-400 (more…)

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Unused PMMA (Acrylic) Extrusion Line – 750-850 Kg/Hr for Sale

By August 6, 2014

Phoenix Equipment has for sale an unused 750-850 Kg/Hr OMIPA OM-150 (Acrylic) Extruder for processing raw PMMA and other acrylics. This extrusion line has never been used and all equipment is still in original packing. OM150The useful width is 2050 mm, thickness 2-12 mm, production/hour is 750-850 Kg/Hr, Power Supply: 3 x 480 V, 60 Hz. The control is on the left hand side in extrusion direction. The total installed power is 1025 KW, installed motive power is 555 KW, and installed heating power is 470 KW. The Extrusion Line’s contemporaneity factor is 0.6, the average power used is 615 KW, Max. Water Consumption is 26.6 m3/hour, Max. Air Consumption is 8.8 Nm3/hour, Machine Color: RAL 7032.

Click Here to View Other Process Plants Available at Phoenix Equipment

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Liquid phase oxidation processes for the removal of hydrogen sulfide (H2S) from gas streams

By July 29, 2014

Phoenix Equipment has several liquid phase oxidation units for removing H2S from Gas Streams. Below is a link to our refinery units where you can find several oxidation processes and other refinery units as well as a well-written piece by Gary Nagl called, “The State of Liquid Redox”.

Phoenix Equipment Corporation Refinery Units for Sale

The State of Liquid Redox – By Gary J. Nagl


As early as the beginning of the century, work has been ongoing to develop a liquid phase, regenerative process for converting hydrogen sulfide (H2S) into pure, elemental sulfur. This work has lead to the introduction of over 25 different processes, most of which with very little commercial success. However, in the late 1940’s, the North Western Gas Board and the Clayton Aniline Company developed the Stretford Process, which utilized an aqueous solution of vanadium and anthraquinone (ADA). Although the Stretford process had some serious process, operational and environmental problems, the process filled a much-needed niche and became fairly popular throughout the 50’s, 60’s and 70’s.

In the late 1960’s the CIP process, which employed an aqueous solution of chelated iron, was introduced in the United Kingdom. The process failed miserably; however, its failure did lead to the successful introduction of the LO-CAT® Process in the late 1970’s, which solved many of the problems encountered with the CIP Process and the Stretford Process. In the late 1980’s, another chelated iron process, the Sulferox Process, was introduced; however, the developers of the LO-CAT and Sulferox processes have recently combined efforts to improve liquid redox processing even further.

Throughout the last half of the century, liquid phase oxidation has played an important role in the recovery of sulfur from various sources of hydrogen sulfide. Of all the processes available for converting H2S to sulfur, current liquid phase oxidation systems are the most versatile. They are able to treat any type of gas stream containing H2S, at a wide variety of operating conditions and all at removal efficiencies exceeding 99%. This paper will discuss the major liquid oxidation processes, describing in detail their advantages and disadvantages. In addition, the current R&D efforts in the field of liquid phase oxidation will be discussed with a glance of what future developments may occur.

Keywords: Desulfurization, sulfur recovery, liquid redox, chelated iron, LO-CAT


Liquid phase oxidation processes for the removal of hydrogen sulfide (H2S) from gas streams were initially developed to correct certain problems associated with dry oxidation processes such as iron sponge. The problems being mainly large plot requirements, replacement of the oxidation media on a frequent basis, and safety problems. This development work led to the utilization of various oxygen carriers dissolved or suspended in a liquid phase, which could be regenerated continuously at ambient temperatures.

Most early development work was done for systems processing coal gas or town gas with the objective of removing both hydrogen sulfide and ammonia by the formation of ammonium sulfate and elemental sulfur. Some of these early processes included the Feld Process, the Gluud Process, and the Koppers C.A.S. Process all of which employed polythionate solutions. Because of the complex chemistry, none of these processes achieved any commercial success.

After the failure of the polythionate processes, development shifted towards utilizing suspensions of iron oxide in aqueous solutions, which in essence was an attempt at a continuous iron sponge process. The hydrogen sulfide reacts with an alkaline compound to form hydrosulfide, which reacts with iron oxide to form iron sulfide, which in turn reacts with oxygen to form iron oxide and sulfur. Development work in this area lead to the introduction of the Burkheiser, Ferrox and Manchester processes.

During the 1920’s the Thylox and Giammarco-Vetrocoke processes met with some commercial success. However, both of these processes employed thioarsenate solutions, which resulted in toxicity problems caused by the arsenic.

Another group of processes which showed technical promise but were limited by toxicity problems were those employing iron cyanide solutions. The Fischer Process and the Autopurification process are examples. These processes employed ferric and ferrocyanide complexes as oxidizing agents.

A summary of the various liquid phase oxidation processes, which have been developed throughout the years, is contained in Table I.

The Stretford Process

The first liquid phase, oxidation process, which gained widespread commercial acceptance, was the Stretford process. The process was developed by the North Western Gas Board and the Clayton Aniline Company in England to remove H2S from town gas. The original process utilized an aqueous solution of carbonate/bicarbonate and anthraquinone disulfonic acid (ADA). However, the process suffered from several inherent problems. The solution had a very low capacity for dissolved sulfides resulting in large liquid circulation rates and hence, high power consumption. In addition, the sulfur formation reaction was very slow requiring large liquid inventories and resulting in high byproduct formation (thiosulfates).

These problems were corrected to a certain extent by the addition of alkali vanadates to the solution, which, in essence, replaced dissolved oxygen as the oxidant in the conversion of hydrosulfide ions (HS-) to elemental sulfur. In the reaction the vanadium ions undergo a valance change from +5 to +4. To reoxidize the Va+4 ions back to the +5 State, ADA is added as an oxygen carrier, and the ADA is subsequently regenerated with air. The process chemistry can be summarized as follows;

CO3-2 + H2S HS-+HCO3-
4VO3-+2HS-+H2O V4O9-2+2S+4OH-
V4O9-2+2OH-+H2O+2ADA 4VO3-+2ADA(reduced)
2ADA(reduced)+O2 2ADA+H2O

A typical flow diagram of a Stretford unit is illustrated in Fig. 1.

Although the addition of vanadium to the Stretford Process increases the reaction rate of hydrosulfide ions to sulfur sufficiently to make the process commercial, it still produces a significant amount of byproduct thiosulfate. The reaction is still slow enough that air streams cannot be treated due to the high rate of thiosulfate formation.

Hundreds of Stretford units have been installed throughout the world; however, their popularity vanished in the mid 1980’s after the introduction of the chelated iron processes which addressed many of the deficiencies of the Stretford process

Chelated Iron-Based, Liquid Redox Processes

Iron is an excellent oxidizing agent for the conversion of H2S to elemental sulfur; however, due to the very low solubility of iron in aqueous solutions, the iron had to be present in the dry state (iron sponge) or in suspensions (the Ferrox process) or compounded with toxic materials such as cyanides. In the 1960’s development work was begun in England to increase the solubility of elemental iron in aqueous solutions. This work led to the introduction of the CIP process, CIP being an acronym for “Chelated Iron Process”.

In this process, iron, in its’ ferric state (+3), is held in solution by a chelating agent, namely ethylenediaminetetraacetic acid (EDTA). The intent of the process was to oxidize sulfide (S=) and hydrosulfide (HS-) ions to elemental sulfur by the reduction of the ferric (Fe+3) iron to ferrous (Fe+2) iron, and the subsequent reoxidation of the ferrous ions to ferric ions by contact with air. The chemistry of all chelated iron processes is summarized as follows with (l) and (v) representing the liquid and vapor states, respectively;

H2S(v) + H2O(1) H2S(1) (1)
H2S(1) H++HS- (2)
HS-+2Fe+3+H+ S0+2Fe+2 (3)
1/2O2(air)+ H2O(1) 1/2O2(1) (4)
2Fe+2+1/2O2(1)+H2O 2Fe+3+2OH- (5)

Overall Reaction

S0+H2O (6)

Equations 1 and 2 represents the absorption of H2S into the aqueous, chelated iron solution and its subsequent ionization, while equation 3 represents the oxidation of sulfide ions to elemental sulfur and the accompanying reduction of the ferric iron to the ferrous state. Equations 4 and 5 represents the absorption of oxygen into the aqueous solution followed by oxidation of the ferrous iron back to the ferric state.

Equations 3 and 5 are very rapid. This is in contrast to the oxidation reactions in the Stretford process when using vanadium. Consequently, iron-based systems generally produce relatively small amounts of byproduct thiosulfate ions, and in properly designed units, air streams can actually be processed. However, as in the Stretford process, equations 1 and 4 are relatively slow and are the rate controlling steps in all chelated iron processes.

It is interesting to note that the chelating agents do not appear in the process chemistry, and in the overall chemical reaction, the iron cancels out. So the obvious question is why is chelated iron required at all, if it doesn’t part take in the overall reaction. The iron serves two purposes in the process chemistry. First, it serves as an electron donor and acceptor or in other words, a reagent. Secondly, it serves as a catalyst in accelerating the overall reaction. Because of this dual purpose, the iron is often called a “catalytic reagent”. The chelating agent(s) do not part take at all in the process chemistry. All the chelating agents do is to increase the solubility of iron in water, thus reducing the circulation rates required to furnish the two moles of iron required in equation 3.

Although it appears that chelated iron would solve many problems associated with previous liquid oxidation processes, the CIP Process failed miserably. In very short order after starting up, all of the iron precipitated out of solution as iron sulfide, FeS. The problem was that the chelation strength of many chelating agents varies considerably with solution pH, and unfortunately, the pH’s experienced in the CIP process were outside the range of EDTA.

In the early 1970’s, a small company in the Chicago area (ARI Technologies) started development work on a process, which employed multiple chelating agents. The idea being that by employing chelating agents with overlapping pH ranges where the chelation strengths were high, the iron would stay in solution at all times. This development work led to the introduction of the LO-CAT® process in the late 1970’s. The process worked well for units processing small quantities of H2S; however, in the first unit processing tons per day of sulfur, the process, in essence, failed. It was found that the chelating agents were disappearing very rapidly requiring extremely high chemical makeup rates.

Research found that the chelating agents were being oxidized to useless byproducts by a free radical mechanism. After a few years of experimentation, the chelate oxidation rate was reduced to an acceptable level by the introduction of free radical scavengers and by switching to chelating agents, which were much more resistant to oxidation.

Iron-based, liquid oxidation has developed into a very versatile processing scheme for treating gas streams containing moderate amounts of H2S. Advantages of these systems include the ability to treat both aerobic and non-aerobic gas streams, removal efficiencies in excess of 99.9%, essentially 100% turndown on H2S concentration and quantity, and the production of innocuous products and byproducts.

The three most common processing schemes encountered in iron-based, liquid oxidation systems are illustrated in Fig. 2 through 4. Fig. 2 shows a “Conventional” unit, which is employed for processing gas streams, which are either combustible or cannot be contaminated with air such as carbon dioxide, which is being treated for beverage purposes. In this scheme, equations 1 through 3 are performed in the Absorber while equations 4 and 5 are performed in the oxidizer. Fig. 3 illustrates an “Autocirculation” unit, which is used for processing acid gas (CO2 and H2S) streams or for other non-combustible streams, which can be contaminated with air. In this scheme, equations 1 through 3 are performed in the “Centerwell” which is nothing more than a piece of pipe open on each end. The purpose of the centerwell is to separate the sulfide ions. from the air to minimize byproduct formation. The volume within the centerwell is essentially the same as the absorber in a conventional unit. The other unique feature of the Autocirculation scheme is that no pumps are required to circulate solution between the centerwell (absorber) and the oxidizer. In these units there is a larger volume of air than acid gas; consequently, the aerated density on the outside of the centerwell is less than on the inside resulting in a natural circulation from the oxidizer into the centerwell.

The last type of processing scheme (Fig. 4) is the aerobic unit (air contaminated with H2S) in which equations 1 through 5 all occur within the same vessel, at the same time and without separation of the absorber and the oxidizer. These are generally less expensive units than the other two schemes; however, because there is always oxygen in the presence of sulfide ions, consequently, these units produce the most byproducts.

The Present

Although references are continuously made that iron-based, liquid redox systems are plagued with plugging and foaming problems and that the process cannot be operated at high pressure due to pump and foaming problems, these problems for the most part have been solved for some time.

Foaming occurrences are either a start-up phenomenon or the result of large amounts of liquid hydrocarbons entering the unit. During the initial start-up of a unit, the surface tension properties of the fresh solution are such that the foaming may occur during the first few days of operation. However, by following proper start-up procedures, this foaming is easily avoided. In addition, this foaming tendency is only experienced when the entire unit is filled with fresh solution, which only happens during the initial start-up of the unit. Foaming does not occur during subsequent start-ups.

Continuous incursions of small amounts of liquid hydrocarbons are frequently experienced with no adverse effect on the operation of a unit; however, the introduction of large amounts of liquid hydrocarbons can present foaming and plugging problems. This would also be true of Claus units, selective oxidation processes and hydrocarbon-based, redox systems. However, for aqueous-based redox systems, “Designer” surfactants1 have been developed, which in essence totally alleviates the problems caused by the introduction of large amounts of liquid hydrocarbons.

During the early days of liquid redox, sulfur plugging was a severe operating problem. Packing plugged, static mixers plugged, pipes plugged, heat exchangers plugged and distributors plugged. For the most part, all of these plugging problems have been eliminated. Vessels with random packing are no longer used, on-line cleaning procedures have been developed for static mixers1, which require very little operator attention, proper pipe design has eliminated pipe plugging, proprietary heat exchanger designs and proper operating procedures have minimized heat exchanger plugging, newly designed absorber spargers are being installed, which have greatly extended the life of sour gas spargers and improved quality control of oxidizer sparger materials and proper operation of the process has minimized oxidizer sparger plugging.

Operation of aqueous-based liquid redox systems at high pressure has been a problem due to difficulties with keeping the liquid circulation pumps running. Circulation pumps were always specified as ANSI, open-impeller centrifugal pumps. The logic being that closed-impeller pumps would plug with sulfur particles or possibly erode. Consequently, for high head applications in which open-impeller pumps would not apply, plunger type pumps were chosen. The plunger pumps had no difficulty supplying the required head; however, seal rings had extremely short lives. To solve this problem, a multi-staged, closed-impeller, centrifugal pump was installed in one high pressure application with excellent results. The pump was in continuos operation for approximately1_ years without any signs of plugging or erosion. For all future high-pressure applications, closed-impeller single or multi-stage centrifugal pumps will be specified. Obviously, the original concern about plugging had no basis.

The Future

Although iron-based, liquid redox processes have gained acceptance as evidenced by over 150 units being licensed worldwide, there are still areas in the process, which need to be improved upon. Current areas of R&D efforts are reduction in operating costs, reduction in equipment size and improvement in molten sulfur color.

Operating costs for aqueous, iron-based redox systems are composed of replacing chemicals which are either oxidized in the unit or which are physically lost from the unit and of electrical power required for circulating solution and injecting air. For any iron-based system there is an economic tradeoff between iron concentration and the solution circulation rate. Since 2 moles of iron are required for every mole of hydrogen sulfide (equation 3), the amount of circulating solution required is dependent on the iron concentration in the solution and the amount of H2S in the sour gas stream — the higher the iron concentration, the lower the circulation rate and hence, the lower the power consumption. Conversely, the higher the iron concentration, the higher the catalyst makeup rate required to replace iron from physical losses such as solution lost with sulfur withdrawal. There is an optimum iron concentration based on the incremental cost of power and the amount of solution, which is normally lost from the system. This relationship is shown in Figure 5.

Besides replacing iron, which is physically lost from the system with the sulfur and blowdown streams, chelating agents are chemically oxidized into useless, non-toxic byproducts within the system and must be replaced. As stated previously, different chelating agents have different resistances to chemical oxidation. In addition, chemicals may be added to the system or made in the unit, which act as free radical scavengers, thus retarding chelate oxidation. Research continues on the development of oxidation resistant chelates and on economical, free radical scavengers. For example, many compounds from the polyamine family have proven to be excellent free radical scavengers reducing chelate degradation to essentially zero. Unfortunately, the oxidation rates of the polyamines are extremely high and consequently, uneconomical. Also new chelating agents have been developed which have very high resistances to oxidation; however, they currently are uneconomical to manufacture. The search for stable chelating agents and/or economical free radical scavengers continues.

A large portion of the electrical consumption in an iron-based, liquid oxidation system is associated with blowing air through the solution to satisfy the oxygen demand of equation 5. Due to the low solubility of oxygen in water, a large excess of air is generally employed depending on oxidizer design. Although the air is free, the back pressure on the blowers is usually between 0.5 and 1.0 bar (g), thus the blower usually represents a large portion of the unit’s electrical consumption.

A considerable amount of research is currently underway to develop new mass transfer devices which will improve the oxygen utilization in liquid oxidation systems with the aim of reducing the quantity of air required (operating cost) and reducing the size of the oxidizing vessels (capital cost). Currently, there are two types of oxidizers employed -low head and high head oxidizers. In the low head oxidizers air is sparged through approximately 3 meters of solution at superficial air velocities of less than 3.5 M/min by means of distributors equipped with EPDM sleeves which are perforated with very small holes. Solution flow is perpendicular to the airflow. These oxidizers are relatively poor mass transfer devices; however, they do provide much need solution inventory for proper operation of the system.

In high head oxidizers air is sparged through approximately 7 meters of solution at superficial velocities of greater than 10 M/min by means of course bubble pipe distributors. Solution flow is co-current to the airflow. These oxidizers provide mass transfer coefficients which are approximately 4 times better than the low head oxidizers; however, this is at the expense of higher discharge heads on the air blowers.

Neither the low or high head oxidizers are very good mass transfer devices; however, they generally do not plug with sulfur and they do supply solution inventory required for proper operation of the system. A main goal of current research is to develop a mass transfer device, which will reduce the amount of air required to approximately stoichiometric quantities, will reduce the oxidizer volume and will not plug with sulfur. To this end, special synthetic membranes are being tested. Initial results indicate that the first two objectives of the research — stoichiometric air and high mass transfer coefficients — can be obtained. Long term testing is currently being carried out to determine the plugging tendencies of the membranes. If this last phase of testing is successful, liquid oxidation systems will become much smaller and less expensive to operate.

Sulfur produced from liquid redox systems has the same chemical assay as Claus sulfur, and it does have several commercial uses2 in its unmelted form. In fact Lubrizol in France has been recycling its produced sulfur with no adverse effect for quite some time. However, there is always a desire to improve the appearance (color) of redox sulfur, which is degraded due to the presence of iron polysulfides. Development work is being conducted on a variety of filter/wash systems with the goal of removing the chelated iron from the sulfur cake prior to final use. However, due to the extremely low price of sulfur and due to the relatively low quantities of sulfur produced in liquid redox plants, it is difficult justifying much work in this area.


Liquid phase oxidation systems have undergone considerable evolution during the 20th century, and this will continue into the 21st century. Foreseeable developments for the near future will be smaller equipment sizes and lower operating costs which will be achieved by the development of better oxygen mass transfer devices reducing the amount of air required and the size of oxidizers and by the addition of free radical scavengers into the systems.


1. Reicher, Myron, Niemiec, Bill & Katona, Tamas, “The Use of Liquid Redox Desulfurization Technology for the treatment of Sour Associated Gases at Elevated Pressure”, presented at the GPA European Chapter Continental Meeting, 1999

2. Nagl, Gary, ” Emerging Markets for Liquid Redox Sulfur”, presented at Sulphur97, Vienna, Austria

Table I. Liquid Phase Oxidation Processes



Polythionate Solutions Feld Process
Koppers C.A.S. Process
Iron Oxide Suspensions Burkheiser Process
Ferrox Process
Gluud Process
Manchester Process
Thioarsenate Solutions Thylox Process
Giammarco-Vetocoke Process
Iron Cyanide Solutions Fischer Process
Autopurification Process
ADA Solutions Stretford Process (4)
Takahax Process (4)
Iron Complex Solutions Cataban Process
Konos Process
CIP Process
Sulferox Process (4)
LO-CAT Process (4)
LO-CAT II Process (4)
(4) Indicates commercial success

Here are links to different Refinery Units available right now at Phoenix Equipment:


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Complete Refinery Units for Sale at Phoenix Equipment

SulFerox Unit for Sale at Phoenix Equipment

By May 21, 2014

View our SulFerox (SFX) Unit from within a Refinery, 4 TPD, used for oxidizing Hydrogen Sulfide within crude oil to create elemental sulfur. View this SulFerox Refinery Unit for sale and a description below about the SulFerox Process.

SulFerox®  Process

SulFerox® is a reduction-oxidation process developed by Shell. In this process H2S is directly oxidized to elemental Sulfur while the ferric ions are reduced. This is subsequently regenerated by oxidation with air. The recovered sulfur cake normally has moisture content of approximately 30%. A higher purity sulfur product can be obtained by additional washing and melting of the cake. The SulFerox® process can be used for the treatment of H2S, in gases such as natural gas, amine unit regenerator of gas, syngas treatment, geothermal gas streams and carbon dioxide purification, refinery fuel gas, hydrogen recycle, coke oven gas desulfurization.

SulFerox®  Process has many advantages.

  • Direct conversion of H2S to Sulfur in one step. Thus, avoiding the combination of absorption and Sulfur recovery units.
  • Less than 1 ppm of H2S in the treated gas can be achieved without the need of further treatment for environmental impact.
  • Large turndown in gas flow and/or inlet H2S concentration.
  • Very selective H2S removal as hardly any CO2 will be removed.
  • The lowest investment cost in Redox type technology.

Phoenix Equipment is selling a used SulFerox Unit from a refinery. Buying a used plant can save significant investment and make production commissioning much quicker. Check our link for more details.

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125,000 Liter Stainless Steel Fermenter for Sale – Stock #8728

By April 17, 2014

Recently added to inventory, Stock #8728, is a 316L Stainless Steel Fermenter for Sale, Manufactured by Delta, with Vertical Agitator and Coil Jacket (Quantity: 5 Available). Check out our other Fermenters and Reactors on our Website.






Full Description:

Used 125,300 liter Delta Fermenter with intermix vertical agitator, shell rated 3 bar @ 143 C, half coil jacket rated 6 bar @ 160 C, made with 316 stainless steel, surface area approx. 120 cubic meter, weight 43,000 kg, Fermenter Measures 3880mm inside dia x 9100 M T/T, Built 2002. (Quantity: 5 available)


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Phoenix Equipment specializes in buying and selling used chemical process plants and machinery, serving the chemical, petrochemical, pharmaceutical, refinery, gas processing and power generation industries. We carry over 40 categories of process equipment that is immediately available, as well as over 50 complete plants that can be relocated quicker and for less capital investment that building a new process plant. In addition to buying and selling, we have the in-house capability to safely perform comprehensive site closure and dismantling projects. Because we understand the value of chemical processing machinery, we can pay you for your shut down plant and safely perform the removal of all assets, structures and associated parts at no cost.