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ABPG – Brimstone Sulfur Symposium

Vail - 2011

Presentation hard copy material

I. ABPG Mission – History

See following pages 2 -3

II. ABPG Membership

See following page 4

III. ABPG Member Correspondence: August 2010--Aug 2011

See following pages 5 - 35

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ABPG MISSION - HISTORY

Mission:

To accumulate a database on amine unit operations and become a clearing-

house for the analysis and distribution of such data to the gas processing and

refining industry on a global basis through articles in trade journals and

through symposiums.

Purpose:

To provide an open forum for exchange of operating experiences for the

purpose of benchmarking, troubleshooting and developing best practices

leading to improved unit operations and reliability.

Scope:

The focus will be primarily issues pertaining to process units that will include

amine treating, sour water treating, sulfur recovery and tail gas treating.

History:

The history of the ABPG is relatively short. It was formed in late 1992 when a group

representing the refining industry, design engineering and an independent

consultant met to discuss the possibility of developing a real-world database for

amine unit operations. The impetus for this initial meeting was a stated interest by

many refiners, both major and independent, in evaluating their amine unit

operations with respect to others in the industry.

The early meetings were devoted to developing an amine unit survey questionnaire

and the ABPG Best Practices Manual. Responses from the questionnaire were the

basis for the 1994 database and includes 75 amine units.

In subsequent meetings, ABPG members analyzed the 1994 database and developed

two articles that were published in industry trade journals. The database was also

the basis for the ABPG Amine Users Symposium held in 1995.

In 1995 ABPG members developed a new survey questionnaire that focused on

amine unit cost control. The results of this survey database were published in the

Summer 1997 issue of Petroleum Technology Quarterly.

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In 1997, the ABPG established a Data Exchange Network (DEN) to provide

members with an open forum to post questions and exchange operating experiences.

Since 1998, ABPG members have met annually to address specific topics of interest

and develop outlines for future articles to be published that focus on issues of

general interest to the industry.

In 2000, the ABPG adopted a focus project to develop a protocol, format and

procedure for standardized amine testing to present to the industry. This effort is

still in progress.

In 2002, the ABPG established a website to host the DEN. Currently, only ABPG

members and ABPG member company personnel have access, but an effort is under

way to make selected DEN data available in the public domain.

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AMINE BEST PRACTICES – MEMBERSHIP ----- August 2011

Title: Member Directory

[ADM-001]:

1. Asquith, Jim – Valero Energy Corporation 2. Bela, Frank – Member Emeritus (formerly Texaco, Shell) 3. Buziuk, Frank – Member Emeritus (formerly Chevron) 4. Crockett, Steven – BP (US) 5. Davis, Jay – Chevron 6. Eguren, Ralph – BP (US) 7. Hatcher, Nate – Member Emeritus (formerly ConocoPhillips) 8. Heeb, Dick – Marathon Petroleum Company LLC

9. Hittel, Shelley - SemCAMS 10. Keller, Al – ConocoPhillips 11. Kennedy, Bruce – Member Emeritus (formerly Petro-Canada) 12. Ritter, Nathan (“Tex”) – Flint Hills Resources 13. Schendel, Ron – Consultant 14. Smith, Conrad – DCP Midstream (formerly Duke Energy Field Services) 15. Stern, Lon – Member Emeritus (formerly Shell Global Solutions)

16. Tracy, Frank - ConocoPhillips 17. Tunnell, Duke – Business Manager 18. Way, Bill – EnCana Corporation 19. Welch, Bart – Chevron 20. Young, Mark – Suncor Energy 21. Zacher, Mike – Member Emeritus (formerly BP Refining, UOP)

Updated 23-July-11 // lhs

[ ABPG – Membership.doc ] LHS/07-23-11

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Amine Best Practices Group

Member Correspondence – August 2010 – August 2011

ARU ARU-249 Tower Vacuum Relief

ARU-250 Carbon Bed Support & Loading

ARU-251 Amine in Fuel Gas

ARU-252 Separator Sizing

ARU-253 Dimethylethanolamine (DMEA) Corrosion

ARU-254 Skid Mounted ARU Rental

ARU-255 PSV Pool Fire Case

SRU

SRU-199 CFD Modeling

SRU-200 Hot Restart

SRU-201 Rotary Packed Bed Desulfurizer

SRU-202 Co-firing Natural Gas

SRU-203 Catalyst Activity

SRU-204 Reaction Furnace Thermal Shroud

SRU-205 Steamloc Steam Traps

SRU-206 BFW Temperature

SRU-207 Fixed Area H2S monitors

SRU-208 Merox Oxidizer Vent Gas

SRU-209 Dealing with H2S

SRU-210 Electric Tail Gas Valves

SRU-211 Sulfur Pit Air Intake Piping

SRU-212 Above Grade Sulfur Collection

SRU-213 Liquid Nitrogen Cleaning

SRU-214 Low Capacity SRU

SRU-215 1st Condenser Tube Failures

SRU-216 Eductor Carbon Deposit

SRU-217 Sulfur in Gasoline

SRU-218 Turnaround Maintenance Manpower

SRU-219 Sulfur Flammability in Low O2

SRU-220 SRU-221 SRU-222

SRU-223

NH3 Destruction Sulfur Tank Emergency Vent

Pit Vent Eductor Orientation

CO Emissions

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TGT TGT-053

Haldor Topsoe TK-22X Catalyst

TGT-054 Portable H2S Process Analyzer

TGT-055 Regenerator Overhead Restriction

TGT-056 Quench Column pH Control

SWS

SWS-027 Loss of SWS Bottoms Level

SWS-028 OH Condenser Temp Control

SWS-029 NH3 Acid Gas Line corrosion

SWS-030 Water Vapor Calc

GEN

GEN-029 Darville Crude Oil RSH Control

GEN-030 New Vendors

GEN-031 Solid KOH for Propane H2S Removal

GEN-032 Jet Merox Clay Filter Usage

GEN-033 Baker-Hughes Sulfix H2S Scavenger

GEN-034 Fundamentals of Safety

LLD

LLD-001 TGU Mercaptans

LLD-002 Rich Amine Flash Drum Alarms

LLD-003 Stripped Water pH

LLD-004 Presulfided TGU Catalyst

LLD-005 Rich Amine Emulsions

LLD-006 TGU Booster Blower

LLD-007 Fixed Valve Trays

LLD-008 SO2 Emission Spikes

LLD-009 ARU LOPA

LLD-010 OSHA National Emphasis Program (NEP)

LLD-011 SCOT Catalyst Sulfiding - Problems

LLD-012 Compabloc Lean/Rich Plate Exchangers

LLD-013 Oxygen Line Fire

LLD-014 Sulfur storage pit corrosion

LLD-015 Burner Light-off Sequence

LLD-017 Hythe Plant HAZOP & LOPA

LLD-018 Legal Implications & 2011 LOPA Update

LLD-019 SRU Hot Standby Operation

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Amine Best Practices Group

Member Correspondence – August 2010 – August 2011

ARU

Carbon Bed Support & Loading ARU-250

What size screen and support media has proven effective in preventing carbon

migration out of the vessel?

ARU-052 is a good reference for carbon selection, but I do not recall a discussion on

loading guidelines.

2-4-11

RESPONSES

1) 2-4-11

Most common in my experience is a lateral Johnson Screen distributor buried in the

carbon: http://www.johnsonscreens.com/content/laterals-1 I also suggest providing

room for bed expansion in case you need to backwash.

2) 2-4-11

Following probably repeats many of my previous DEN comments:

We usually go with the bed design provided by the carbon vendor. The trick with

loading is to initially fill the vessel with water to minimize generation of fines. Once the

vessel is filled, back-flush the bed (from the bottom up) with enough water to fluidize the

bed and float out the fines. This requires LOTS of water and a very high flow rate. But,

if you do not fully execute this step, you will have carbon fines fouling your downstream

filters and equipment, and possibly promote foaming as well.

It is then recommended to let the carbon sit in the water filled vessel for 24 hours before

draining the vessel and placing it in service. This "activates" the carbon and also

attempts to get as much of the air out of the pores as possible. Avoid initial high

velocities which may generate fines.

3) 2-4-11

Some of our plants load the beds with a slurry, from the supplier. Otherwise we do the

same – backflush with condensate, wait 24 hours, etc.

4) 2-4-11

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I also try to follow vendor recommendations. Attached (ARU-250A) are three references

from my files.

5) 2-5-11

Ditto previous comments.

6) 2-7-11

We mostly have the lateral Johnson Screen distributor buried in the carbon.

7) 2-7-11

Our loading procedure is similar to that previously stated:

Fill the vessel half full with RO treated water.

Load carbon from bags via the top manway.

Soak for 12 hours.

Backwash from the bottom of the bed with more RO water for 20-30 minutes to a

drain or water truck.

We use granular 8 x 30 mesh Petrodarco carbon.

8) 2-7-11

We typically get our vessels from the vendors (e.g. Calgon) and don't specify internal

details. The ones I've looked in all have the lateral Johnson Screen inlets and outlets. I

don't know what the screen size is, but be have a 5μ post-filter to catch fines. Loading is

done as others have described (for the ones that don't slurry load from the carbon supplier

truck): Load into water, let it soak 12-24 hours, backflush to sewer to carry away fines

and oxygen, then line it up to the process and gradually establish flow so you don't send

a slug of water to the absorber. Granular 8 x 30 mesh carbon is typically used.

9) 2-9-11

We utilize a Johnson screen support (wire size 93, rod 158, slot size = 0.010 in, support

bar 0.19 in thick) in the carbon filter followed by an after filter with a cartridge rating of

5μ absolute.

10) 2-9-11

I am sure that you will find low cost vessel venders who will offer outdated distribution

methods such as lateral pipe distributors, wire mesh screens and other tried-and-failed

methods. But, they are either cheaper than the Johnson screen, or the vender has not

kept up on technology. Based on all the comments and past experience, don't accept

these lower cost or outdated alternatives. They can or may work in the short term, but

will most likely become problematic, usually at the worst possible time.

11) 2-9-11

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I agree Johnson screens are the best.

12) 2-14-11

I'll also throw in my vote for Johnson Screen. With the exception of forgetting an anchor

bolt (which was quickly discovered following startup), I have not heard of any incidents

with Johnson Screens. Can't say this for the alternatives though.

If you ever become weary of changing carbon, there is always the HCX in-situ

regenerable oil removal system from MPR services. Pushing a button to flip a rinse valve

lineup is much less of a pain to deal with.

Separator Sizing ARU-252

Our design approach for sizing a knockout drum is to use the Souders-Brown equation

where max velocity through the demister pad is:

Vmax = K * sq rt [(liquid droplet density – vapor density) / (vapor density)]

where K is a constant provided by the demister manufacturer, typically ~ 0.35 at 100

psig.

The GPSA Engineering Databook uses the same method, but says to multiply K by 0.6-

0.8 for glycol and amine solutions. I looked through the references listed by GPSA and

found a 1980 paper by Laurence Reid (himself) where it says to multiply K by 0.82.

Does anyone understand why amine droplets are so much harder to knock out than

hydrocarbon or water, and what is the appropriate fraction of K to use? The 0.6-0.8 is

a pretty wide range.

3-10-11

RESPONSES

1) 3-10-11

I don't think it has much to do with droplet properties. My guess is this is a way to deal

with the fact that foam travels through demisters; essentially it is applying a foam factor

to the sizing. Adding diameter arrests foam height growth by spreading out the foam

structure over a wider area.

2) 3-11-11

Droplet diameter, physical properties and foaming tendency are all part of the equation.

If you just need a knock-out and are willing to allow droplets of a certain size pass

through this unit, then that is one matter. If you really need a clean stream such that

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aerosol liquid droplets (like smoke) will not even pass through this unit then you need a

coalescer. Depending on the pressure and temperature, a knock-out may end up being

quite a bit larger and more expensive than a properly sized coalescer.

3) 3-11-11

The approach you describe is what I have always used with a de-rate factor of 0.75,

regardless of whether for amine, glycol, etc. I've come to the conclusion that anything

further (save CFD modeling) is just polishing the art.

4) 4-11-11

The key to selecting the right K factor is to realize that everything has some measure of

importance. I would agree that the correction factor for glycol or amine should be 0.75.

Don't forget to add the correction for operating pressure shown in Fig. 7-9.

It is also important to understand that as the gas rates drop off, there is a minimum

turndown condition. If the gas velocity through the mesh pas is less than about 25% of

the full design velocity, the liquid will start to bypass the mesh pad and carry over. We

have been surprised in some cases by reasonably sized mesh pads which were really

over-sized for the operating conditions.

SRU

Hot Restart SRU-200

For hot restart following a Claus trip, what is the minimum Reaction Furnace

temperature for going directly from pilot ignition to acid gas (i.e.; without natural gas

heat up)?

We generally require at least 1800°F. Given the wide range of H2S flammability limits,

however, couldn't one safely introduce acid gas at, say, 1000-1200°F?

9-29-10

RESPONSES

1) 9-29-10

I basically use 1000-1200°F, or at least slightly glowing refractory. At lower

temperatures when potentially considered advisable to first heat up on natural gas in a 2-

zone furnace, my preference is to initially establish a minimum natural gas fire with

excess air and sufficient amine acid gas to zone 2 (only) to consume residual O2.

2) 9-29-10

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We go to H2S at 680°C (~ 1250°F).

3) 9-29-10

For a hot restart, we have recommended a maximum of 30-minutes since shutdown.

After that the normal start-up is required, with the purges, etc.

4) 9-29-10

Historically, we generally permitted going straight from pilot to acid gas at 1400°F. With

the improved reliability of Stackmatch pilot ignitors, however, at least some of our

refineries now seem to see less incentive to bypass the transition through natural gas

firing.

5) 9-30-10

The only place I know of does not allow hot relight below 1900°F and allows

a maximum of three attempts before re-purging.

6) 10-4-10

Although NFPA 86 is the Standard for Ovens and Furnaces, it says 1400°F is the

minimum temperature for restarting without requiring a purge or flame scanner in

service.

7) 10-5-10

We have a minimum of 1800°F.

8) 10-6-10

Be sure that the auto-ignition and other combustion temperatures you use are appropriate

for the situation at hand. When you reference a "wide range of flammability", does this

imply that you are going to do your hot restart in an excess oxygen atmosphere? I think

not. The temperature of concern here is that required to sustain combustion at the H2S/air

ratio to be used for the restart. The presence of the ignitor flame isn't likely to provide

sufficient supplemental heat to sustain the sub-stoichiometric H2S combustion you desire,

if the system has cooled too far. I would use a conservative approach and not use a lower

hot restart temperature with the ignitor in service than I might consider acceptable with a

questionable or non-functional ignitor.

9) 10-20-10

A number of the newer burner management systems make this question moot. They

automatically kick in timed purges etc.

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10) 1-14-11

First, let me emphasize that, as I understand it, there is generally no intention of

bypassing the purge cycle and pilot ignition, just the transition through natural gas firing.

Most of the responses appear directed at explosion avoidance, which I don’t think is the

issue. Lighting even a sub-stoich H2S flame in a relatively cool furnace is arguably safer

than lighting natural gas by virtue of its wider explosive range, lower ignition

temperature and lower heat release. The real issue is the risk of acid condensation at

abnormally cool refractory and metal surfaces.

Natural Gas Co-firing SRU-202

An idled 1968 vintage sulphur plant with HGBP reheat was recently recommissioned.

The acid gas load is only 40% of design, and natural gas co-firing is required to achieve

the necessary reheat.

Has anyone co-fired long term?

10-28-10

RESPONSES

1) 10-28-10

Natural gas is co-fired normally for one of two reasons – either to increase mass flow at

high turndown, or raise the Reaction Furnace temperature for hydrocarbon destruction

(BTX in particular). BP in Los Angeles practices co-firing for turndown, and Tony

Houlemard gave a nice paper on the subject at the 2006 Brimstone conference. (ARU-

202A attached). The Wilson Creek plant also co-fires lean acid gas, but the burner was

specifically designed for the purpose. (ARU-202B, 2003 LRGCC, attached) I'm pretty

sure that Aramco practiced co-firing to raise Furnace temperatures to destroy

hydrocarbons in the DGA acid gas.

2) 10-28-10

My experience has been to discourage long-term co-firing. Because we highly insist on

always being in reducing mode in the burner/front end of the thermal reactor, there is a

battle for the limited air. If the natural gas is at least partially oxidized to CO, then all

can be OK, but there is a significant risk of making COS, CS2 and soot. Operating the

burner with local excess air results in it leaving the well designed mixing area and relying

on the residual O2 to be consumed in the thermal reactor, where there is very poor mixing

at best.

3) 10-28-10

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ASRL has demonstrated that the methane does not have to "find" O2 to be completely

destroyed; it can also react with SO2. There are a number of successful co-firing

operations. (as per Tony's Brimstone paper). The key is to use natural gas, NOT plant-

produced fuel gas with heavier hydrocarbons. Also ensure sufficient flow rates for good

mixing and relatively high furnace temperatures.

4) BP, 10-28-10

We recommend against long-term co-firing due to potential sooting. However, most

refineries have to do it at times, to keep enough load on the SRU to operate in a steady

manner. I recommend that you consider using either steam or nitrogen to get your load /

velocities up to try and balance the heat load. If that doesn't work or there is not enough

energy release, then use natural gas.

5) 10-29-10

I concur with previous comments, but have also been forced into co-firing at one facility

to achieve sufficient ammonia destruction temperatures. For a couple of years, frequent

sooting of the first bed was rather common, and rakes were actually installed to

periodically break up the carbon crust while online. However, the problem was largely

mitigated by burner improvements and modernization of the 1970s vintage

instrumentation.

6) 10-30-10

I concur that there are a large number units successfully co-firing, but I also agree that

instrumentation and controls are key. I know of one site that had some success with H2,

but even CCR reformer gas that has been compressed and knocked-out has potential for

heavies.

7) 11-15-10

Comments from one of the authors of the Wilson Creek Gas Plant paper:

Suggest you contact HEC about Harmattan gas plant which was recently modified for co-

firing. Wilson Creek is likely the Cadillac as there is split flow and an HEC burner with

fuel gas on the inside followed by air and acid gas mixing gun. There was an issue where

the hot gas by-pass was too cold so we had to insert ceramic ferrules into the tube at both

ends to reduce heat transfer. This created a higher outlet temperature but caused more

pressure drop so we had to add another valve to pinch the main gas out to force more into

the HGBP. The plant has been co-firing for over 9 years and no catalyst changes to date.

8) 1-14-11

If soot is made, it will be readily evidenced by off-color sulfur in the first condenser. The

other downside of co-firing is reduced recovery efficiency, but the upside is reduced

Converter dew points.

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9) 2-18-11

I agree with others’ comments and concerns. We have had to co-fire with natural gas due

to less sulfur in the plant from time to time and have not had any issues. No real

experience with long term as some have mentioned.

Reaction Furnace Thermal Shroud SRU-204

As follow-up to SRU-189 (Reaction Furnace loss of containment due to high

temperature), the plant has removed the thermal shroud to facilitate periodic thermal

imaging scans in order to detect hotspots before another failure occurs. As a result of the

incident, Operations is more vigorously monitoring the view ports and reporting anything

that looks like refractory failure. The thermal scan is used to confirm if there is a hot spot

before making a decision to bring SRU down for further inspection.

I pointed out to Operations that the shroud has an important thermal insulation role in the

protection of the furnace shell and that removal may solve one problem but create other

issues such as acid dew point corrosion and refractory pinch spalling. The decision to

remove the shroud was justified based on two older (20+ years) SRUs at the same facility

that do not have shrouds. One has an expanded metal screen for personnel protection,

and the other has a roof 10 ft above.

Has anyone willingly or unwillingly operated their furnaces without a shroud?

Does anyone use removable shrouds that permit on-line inspection?

Does anyone employ thermal imaging to monitor the skin temperature and how

effective is it?

11-9-10

RESPONSES

1) 11-9-10

I suspect that, depending upon your weather environment, you may be able to get along

without the shield. As you mentioned, the purpose is to protect the shell from the cooling

effects of wind, rain and snow which can result in acid condensation. I do know of one

plant that installed a roof which enabled them to perform thermal scans, but do not know

of anyone that periodically removes a conventional shroud to perform scans.

2) 11-9-10

We have not used a removable system to be able to make thermal scans. We had one

plant in a relatively rainy climate (US Gulf Coast) claim that a leaking rain shield

accelerated failure on the top of a reaction furnace shell, but the analytical evidence did

not support this theory.

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Two years ago, I would have said a rain shield is an absolute requirement. It's my new

opinion that CFD modeling is needed to confirm or deny this assertion now.

3) 11-9-10

A rather simple way to assess how much risk you are taking without the shroud is to

check the shell temperature during your coldest days and in a heavy rainstorm. I had

experience at a plant with occasional rain or snow and wintertime temperatures of -40°.

We did remove some of the shroud on one of the units to monitor hot spots. Our

maintenance folks modified the section of shroud they cut away and were able to more or

less set it back in place for some weather protection.

4) 11-9-10

The SRU furnaces that I am aware of that do not have weather shrouds are in warm, dry

climates – specifically, not too much rain to worry about.

5) 11-10-10

We have two Parsons SRUs in NJ operating for 25+ years without a rain shield. I was

concerned upon first seeing this since I've always believed a shield was rather important,

but review of maintenance records indicated reliability had been very good. Only repairs

to the shell were at the furnace manway due to internal refractory problems unrelated to

lack of a shield. They do periodically perform IR scans on the unit.

6) 11-10-10

We had one thermal reactor without a weather shroud where we had to replace over 50%

of the vessel during the last turnaround. We installed a shroud during that turnaround.

We do not do routine monitoring of the shell temperature using thermal imaging. We

have removed some of the shroud to check the shell temperature if we thought there was

a problem, but replaced the shroud when this was completed. We have one refinery that

has skin TIs on the shell.

7) 11-11-10

Most of our furnaces are older Parsons designs and do not have thermal shrouds. These

are California and Gulf Coast refineries and there haven't been any problems associated

with the shell on these vessels. We do thermal imaging routinely, usually quarterly or

after startup. One refinery found a hot spot around a new nozzle this way which turned

out to be an installation issue but allowed them to run with a steam ring on the nozzle to

the turnaround. Our refinery located where it is temperate but rains all the time does not

have a shroud. Only our northern US units have shrouds.

8) 11-11-10

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In my experience, there are vast differences in the actual performance of shrouds. In our

refineries, there are usually ridge vent openings and bottom louvers, ostensibly to

maintain skin temperatures within 400-600°F. Of course, Operations does not really

make many adjustments to the louvers. That being said, I do not recall many serious

metal corrosion incidents on these types of shroud installations.

Following a burn-thru at one of our Gulf coast refinery SRU Thermal Reactors, a fiber

optic wrap of the vessel was purchased and installed to provide continuous, complete

monitoring of the shell temperatures...all for a cost of less than $1 million (but I do

believe the cost was > $200M).

One location has hinged inspection panels on their shroud to provide access for thermal

imaging, etc. I had never seen this design before.

9) 11-12-10

Our Alberta plants are located in a relatively harsh weather zone and all our furnaces

have shrouds. We have not had any issues with hot spots to failure, and generally

refractory is good in the firing chamber when we do our inspections. We do have PM’s

to check for hot spots but they don’t pull the shroud to conduct the thermal tests. I think

the operators are using a hand-held temperature gun and simply look for spots that are

warmer than expected. They do the same on the power boiler fireboxes.

10) 1-3-11

When a small pile of refractory (about a 5-gallon bucket's worth) was observed to have

accumulated directly in front of the checker wall of one of our Gulf Coast thermal

reactors, the weather shield was removed on Thursday for thermal scanning of the shell

on the following Monday. I have had the glorious opportunity to work for a man named

"Murphy" for many years, and believe me when I say if it can happen it will. Following

a severe rain storm over the weekend, on Monday morning there were about four

wheelbarrows worth of refractory bricks laying on the bottom of the furnace, presumably

due to undue thermal cycling. As it turns out we did not have any hot spots, but I would

not recommend removing a weather shield prior to a heavy rain storm.

11) 1-3-11

I have seen everything from no shroud to shrouds with access to permanent shrouds that

require complex work to remove for a thermal scan. They all have had refractory issues

but, unfortunately, I don't have enough info to draw correlations.

Dealing with H2S SRU-209

I would like to better understand how we all are addressing personnel exposure to H2S in

the refinery and gas patch. In the past we have used the HAZOP, revalidation and MOC

processes along with the stream compositions (including in particular % H2S) to

determine whether a change would be necessary or required.

-17-

Is this still the fundamental method for the protection of employees from H2S

exposure, and are there any other methods being used?

3-3-11

RESPONSES

1) 3-3-11

In the gas field, where H2S concentrations are generally minor, we have a hard time even

getting samples taken and analyses run. This is a problem because even a few ppm H2S

will concentrate in the still overhead and become a health, safety and environmental

hazard. It is usually not that hard to obtain concentration information in facilities with

SRUs, where H2S is a well known problem. Then it just pretty much goes back to

Company, OSHA, PSM and other agency requirements as to monitoring, control and

mitigation of the exposure. This is basically one of the things PSM is all about. We

always include nodes in the PHA dedicated to facility siting and human factors to try to

cover all possible ways that the presence of H2S can become a problem. We typically do

an MOC for changes in gas composition at the inlet of plants when new production

comes on line to a facility.

2) 3-3-11

I think we are using tools and processes similar to what you describe. In our sour

facilities, we make the general assumption that H2S could be pretty much anywhere, even

in normally sweet or purged systems. Any intrusive work generally requires precautions

that assume toxic, or at least environmentally undesirable, emissions until proven

otherwise. I'm not sure if this answers your question. If not, perhaps you could provide

an example of what you are thinking.

3) 3-4-11

I also don't quite understand the question. Do you mean direct employee protection tools

and methods, or identifying where and when such protections should be in place and/or

implemented? In either case, the answers will cover a broad spectrum of approaches.

4) 3-4-11

I am specifically looking to see if the way that we each identify the hazards of H2S and

the solutions to minimize employee exposure are similar. I suspect that most of us use

our HAZOPs and other facility reviews to determine if we have the potential for

employee exposure to H2S and then the solution. So, is anyone doing the analysis and

solution in other ways?

-18-

5) 3-5-11

I've seen several non-HAZOP approaches used. Some of them are:

Fundamental design criteria – Locations with high H2S, such as a sulfur recovery

unit, will have fixed point H2S sensors, fixed location escape SCBA units, restricted

access and other protections in place. No HAZOP required – this will be

implemented throughout the process unit.

Plant-wide approach – One of the places I worked advertised itself as the first total

hydroprocessing refinery. One of the basic requirements was that any person entering

the plant boundaries was required to have a personal H2S detector on their person at

all times. A separate facility required every person entering the plant area to carry a

MSA short term escape breathing apparatus. Of course, there were also other

protections established by design criteria, HAZOPS, MOC reviews and other means.

Leak release and dispersion studies

6) 3-8-11

Typically our approach is similar to what has been stated and somewhat subjective. As

part of our process safety management implementation, we are supposed to do toxicity

studies (dispersion modeling) to further quantify the risk and better justify the appropriate

design safeguards and for emergency response planning. However, I am not aware of the

modeling actually being done yet for any of our facilities. There is also discussion of

implementing this in real time to help with the emergency response during an actual

release event. I understand this is already being done for stack emissions at some plants.

7) 3-8-11

There is another process that I overlooked in my previous response. Designed by the

regulatory agency for San Francisco area refineries, it is called Inherently Safer System

(ISS) review, which was performed in conjunction with HAZOPs. Basically, wherever

there is a significant hazard, it is necessary to identify if there is an inherently safer

alternative. There were no binding requirements or obligations to make changes, but it

was necessary to identify where things could be inherently safer. A simplified example is

that it would be safer to switch from sour to sweet crude, but this would be uneconomic.

Another example would be to substitute chloramine or onsite chlorine generation for bulk

liquid chlorine for cooling tower water treatment. In fact, integration of ISS with

HAZOP did lead to implementation of improvements missed by best practice sharing and

other processes.

8) 3-9-11

Our approach to protecting employees in sour gas environments is similar to the other

responses. We have moved to the use of personal H2S and other gas detection monitors.

Our work permit system includes the use of personal monitors by employees and

contractors for all related work.

-19-

The scope of HAZOP has been expanded to include LOPA or SIL analysis. This added

level has increased the quality and severity of our safety-related discussions and provided

more stringent control of our MOC projects.

Finally, we have mandatory education for all employees and contractors about the

hazards and protection from H2S.

9) 3-9-11

Having worked startups in North/South America, Middle East and Asia, the universal

weak link I see is the need for better operator training in process fundamentals and job

skills. Deficient operating procedures are often a factor, for various reasons. Often,

procedures are non-existent or disregarded. More often, verbose procedures are so

cluttered with obvious fluff as to discourage use. Effective procedures must be concise

and well organized. Often, the author appears more concerned with fulfilling a legal

requirement than providing a useful tool for the operator.

Another key is to involve the operators in safety rules. Typically, the safety department’s

philosophy is “You can’t be too safe.” They are wrong. Burdensome unwarranted safety

requirements are like crying wolf. When an operator becomes convinced that an

inconvenient safety rule is unfounded, often the tendency is to disregard any such

inconvenience. Make sure the operators agree with the rules, and encourage them to

discuss any they consider unnecessary. Where dispute exists, evaluate the need for

special PPE or procedures first-hand, then enforce the rules. Whenever possible, of

course, minimize the need for special PPE with engineered solutions.

10) 4-11-11

Attached SRU-209A describes a program that we have been piloting in one of our

refineries called the Man Down system. Personnel are equipped with a personal monitor,

capable of detecting up to six gases including H2S and SO2, which transmits a wireless

alarm signal to the control room. The system can determine the victim’s location and if

he is immobile. The user can also trigger a panic signal. To date the system has been

piloted in five units and will be expanded to the entire refinery within the year.

11) 7-1-2011

I suggest Marathon's Man Down system

would make a good Vail presentation.

-20-

Sulfur Flammability in Low O2 SRU-219

A client asked if liquid sulfur can support combustion in a N2 atmosphere containing up

to 1.5% O2. I expect not, but does anyone know for sure?

7-9-11

RESPONSES

1) 7-10-11

I think the question may be too simplistic to address the underlying issues. I don't have

definitive proof, but do not believe that an atmosphere of 1.5% O2 can support

combustion. But, I have seen sulfur fires in storage tank vapor spaces where the bulk O2

level sampled at a tank vent was very low, or not detectable.

What is the basis for the question and what is the source of the O2? In the situation I

reference, corrosion led to some small holes that allowed O2 to enter the vapor space.

Iron sulfide, presumably, on the tank wall started small localized smoldering, consuming

most or all of the O2. In some cases, this was found before significant damage was done.

In other cases, they developed into rather worrisome sized fires, as the fire near the tank

wall increased the size of the opening until an open flame fire occurred.

2) 7-10-11

In response, PSA N2 used to blanket the vessel can reportedly contain up to 1.5% O2.

3) 7-10-11

I checked the temperature rise assuming 1.5% O2 in N2 at 280°F over a sulfur pit (liquid

sulfur with H2S also at 280°F) and just enough O2 to meet stoichiometry. Turns out

under these ideal conditions the temperature goes up to about 590°F – well above the

auto-ignition temperature of sulfur (and H2S). This would suggest that, once ignited,

even 1.5% O2 would produce self sustaining combustion. If you assume that heat then

raises the temperature of the liquid sulfur producing more sulfur vapor and you then

provide more of the 1.5% O2 the temperature goes even higher. Of course, these calcs

assume just the right amount of 1.5% O2 gas and do not account for dilution with either

excess gas or excess sulfur vapor. Nonetheless, it seems to me liquid sulfur COULD (not

necessarily WOULD) support combustion in a N2 atmosphere with 1.5% O2.

On the other hand, the 1.5% N2 would likely be a pad rather than a purge and, even if

combustion occurred, the temperature of the vapor space would only rise to a max of

about 590°F before all the O2 was consumed. This would most likely not result in

damage.

-21-

4) 7-10-11

In my opinion, this low O2 capability of supporting sulfur combustion is yet another

reason for opting to keep an oxidizing atmosphere in sulfur storage, to continuously

oxidize any FeS as it tries to form.

5) 7-11-11

In a massive state (the pool of liquid sulfur) the contact area would be hard to generate

and I think it would be almost impossible. My guess is that a catalytic oxidation in the

vapor area via iron oxides is quite probable as Lon warns about.

6) 7-11-11

We have done some research on the Limiting Oxygen Concentration (LOC) for H2S and

typical hydrocarbons to ensure a safe environment after purges or vacuum removal before

opening equipment for maintenance. I can share that the LOC for H2S is about 7% O2,

below which combustion would not occur. I have no such data for sulfur, but would

expect that 1-2% O2 is too low for combustion.

7) 7-11-11

I believe that at higher temperature (280°F in the vapor space) the LOC would be quite a

bit lower than at the normal test conditions (25°C). At the higher temperature you have

both H2S and sulfur vapor.

8) 7-12-11

I don't expect that liquid sulfur can sustain combustion in only 1.5% O2. My main

concern is that the vapor space will also contain H2S, which has a wide explosive range

(4-44%) and would probably be a more likely source of combustion than liquid sulfur.

9) 7-18-11

I concur with the concerns over potential formation of pyrophoric FeS acting as the

catalyst to ignite the sulphur.

10) 7-20-11

The following paper states that pyrophoric sulfides in a 2% O2 atmosphere oxidize along

a different reaction pathway than in air – without the incandescence required for an

ignition source – and proposes blanketing the empty holds of crude transport ships with

2-4% O2 to safely oxidize FeS without generating excessive heat.

11) 7-20-11

Along those lines, I had suggested that nominal 1% O2 in the blanket gas might actually

be advantageous in preventing FeS accumulation.

12) 8-12-11

My earlier comment that because combustion would lead to a temperature above auto

-22-

ignition sustained combustion was possible is incorrect. The reaction would have to

result in a temperature sufficient not only to reach auto ignition but heat the incoming

gases sufficiently to ignite them as well. The minimum flame temperature is much

higher than auto ignition temperature. My conclusion is that N2 with 1 1/2% O2 purge or

pad will not support combustion.

TGT

Regenerator Overhead Restriction TGU-055

Has anyone experienced gradual restriction of the SCOT 6” overhead vapor line?

Any ideas on the potential fouling mechanism?

Restriction has increased Stripper operating pressure to 22 psig at times. The restriction

appears to be downstream of the PCV which is at the high point, piping is not pocketed.

Top of vapor line is running hotter than the bottom. They tried to steam-out the line

while on-line. Little change in DP was noticed. Any ideas on potential fouling

mechanism?

11-22-10

RESPONSES

1) 11-22-10

I have not heard of plugging there except at one refinery that added a chemical to uplug

the stripper packing and ended up plugging everything up, including the packing.

2) 11-22-10

One plant complained about their check valve on the line going back to the front end of

the SRU plugging, but they pressured up the TGU regenerator and blew through the

restriction. No deposit was recovered for analysis. Difficult to speculate on.

3) 11-22-10

Is NH3 buildup possible? Measure reflux H2S and pH. If NH3 is the problem, purge

reflux and increase condenser outlet to > 160°F. I assume you have bypassed the PCV to

rule out the valve.

4) 11-22-10

The only times that I have observed this have been where sulfur pit vents or O2 bearing

lines are tied in somewhere. The pit vent is probably self explanatory, but O2 is a bit

more subtle. In the case of the latter, the oxygen can cause direct oxidation of the

remaining H2S in the overhead to sulfur and/or pipe steel to iron sulfate.

-23-

5) 11-22-10

I have seen SCOT overhead line plugging from two phenomenon, both of which

responded extremely quickly to steaming/condensate wash:

H2S hydrate build-up at low ambient temperatures

NH3 salt build-up in the control valve, as Frank noted

6) 11-22-10

Does the "restriction" occur during periods of high throughput to Claus unit? You state

gradual restriction, but then state 22 psig “at times.” At what pressure does the stripper

normally operate? What is the Claus front end pressure during normal operation and

periods of high throughput? I'm thinking that during periods of high Claus throughput

with high front end pressure, and higher than normal H2S being absorbed for higher than

normal recycle flow, you just run out of umph.

7) 12-8-10

We have only seen this during the winter – attributed to hydrates which were quickly

cleared by a water wash.

8) 1-14-11

Bear in mind that NH3 is a product of amine degradation, which is accelerated as reboiler

temperatures increase with increased tower backpressure.

9) 2-8-11

I seem to recall that we once plugged the demister in one of our reflux drums with

ammonia salts and had to steam the heck out of it in order to get the pressure back to

normal.

Quench Column pH Control TGU-056

My spotty memory from past Brimstone discussions is that during an SO2 excursion from

the hydrogenation reactor someone continuously injected caustic into the quench column

because of low pH. This went on for a few days/weeks (?) and eventually they corroded

out the column overhead line leading to a fatality.

Did this really happen or was I dreaming?

Can someone share a safety bulletin for this or a similar incident?

What are your thoughts on how the pH should be controlled?

-24-

My opinion is that I don’t want to make it too easy to just add caustic and move on to the

next issue. I’m OK with adding caustic to bring the pH back to above 6 or 7 but THE

NEXT STEP should be to determine the cause of the SO2 excursion – bad air demand

analyzer, sulfur carryover from the last condenser, low H2, etc.

Is this too conservative?

Is closed loop pH control safe?

6-8-11

RESPONSES

1) 6-8-11

Our pH adjustment is typically via NH3 which, in my opinion, makes it much easier to

avoid an overdose. Also, no issues with chlorides that will be present if using electrolytic

cell caustic (rather than low-chloride “Rayon grade”). My recommendation is to NOT go

to closed loop control. Reliability of pH measurement is too dicey.

2) 6-8-11

We have some plants that do have pH on control, but we have also found that pH meters

in this service require a lot of maintenance to maintain reliability. Those without

automatic control use either NH3 or caustic, and a couple operate their SRUs well enough

that the caustic line is blocked in because it leaked and they rarely ever needed it.

I recommend remote pH control of some sort, and if you have good pH analyzers then

you may consider automating, but the normal operator rounds should include pH testing

twice/shift. As you pointed out, the issue is almost always upstream and the air-demand

controls have to work along with the tail gas analyzer, otherwise you will be fighting it.

Although pH control is a common problem, we have not had the problem of mechanical

failure in the quench column overhead.

3) 6-8-11

Either NaOH, Na2CO3 solution or NH3 is acceptable for pH control. Often it depends on

disposal implications and whether a caustic delivery system is existing. I favor automatic

control, but an HIC is acceptable. My former plant only had a local block valve.

Automatic control must be modulating, not on/off, and alarms should alert the operator

when caustic is being added.

In any event, there is no excuse for prolonged SO2 breakthrough. Our TGUs typically

have no supplemental H2, and my advice is to adjust SRU tail gas H2S/SO2 to maintain 2-

3% H2 in the TGU tail gas. The need for actual off-ratio operation is less likely with a 3-

stage SRU than a 2-stage. Extremely important is a reliable stand-alone thermal

conductivity H2 analyzer – not a GC or thermal conductivity analyzer in series with a GC

sample train.

-25-

The only serious corrosion incidents I am aware of were the result of operating acidic for

several hours, and I am unaware of any personal injury as a result. Even with a basic pH,

I suspect prolonged caustic injection could result in thiosulfate corrosion, potentially

compounded by salt deposition with under-deposit corrosion from thiosulfate and

elemental sulfur.

4) 6-10-11

I don't recall the Brimstone conversation, but we have had our share of ongoing corrosion

experiences. I do heartily agree that ending the SO2 excursions is the real solution to the

problem. However, playing with some solution chemistry brought out something

interesting.

The following summarizes computations made from one of Nate Hatcher's sour

electrolyte VLE computation sheets.

pH Concentration – mole/kg x 10

-6

H2SO3- NH3 NH4+

9 2.2 303.0 227

8 18.3 25.1 187

7 69.5 1.79 133

6 95.8 0.14 104

5 99.5 0.01 90

The chart indicates why pH can be a misleading indicator of corrosivity – mainly because

it only tells how much free H+ is in the solution, not how much total H

+ is available in the

solution to do damage. The table shows that even up to pH = 8 (a solution with 10-4

mole/kg of dissolved SO2), a considerable amount of the original SO2 dissolved will have

an available H+ for corrosion via:

2 HSO3- + Fe → H2 + Fe

++ + 2 SO3

=

In this pH range, the free H+ diminishes rapidly, but there is still H

+ available for

corrosion associated with the bisulfite ion. We find this analog with amine systems and

the HS- species as in the corrosion paper Nate and I did at Brimstone in 2008.

In order for the caustic or NH3 control to totally effect corrosion relief, the HSO3- has to

be totally neutralized to sulfite or SO3=. Complete neutralization does two things: (1) It

takes the H+ away from bisulfite, rendering it non corrosive, and (2) it allows more H2S

to dissolve in the solution to facilitate conversion of the remaining sulfite/bisulfite to

thiosulfate, a highly soluble salt of both Na+ and NH4

+:

2 H2S + 2 SO3= + 2 HSO3

- → 3 S2O3

= + 3 H2O

-26-

Conversion to thiosulfate insures prevention of bisulfite conversion to sulfur via aqueous

Claus reaction.

Attached below is a data plot of pH-analyzer-driven caustic injection to both stages of a

WP Desuperheater/Contact Condenser in response to an SO2 breakthrough. The

complete neutralization knocked out the corrosivity as well as prevented any plugging as

can be seen in the corrosion meter lines on the plot.

The original design provided a day pot of caustic that has to be put in to replace the spent

caustic from a breakthrough, all manually operated. We had to unplug and replace

equipment frequently in those days. While it did take a while to get the bugs out of the

pH analyzer (mainly keeping flow through the sampling system), this set up works really

well now.

5) 6-10-11

All of our TGUs have caustic or NH3 injection that is manually controlled by the operator

based on pH analyzer indication with further confirmation from a lab sample. Haven’t

had good experience with reliable caustic injection systems and have gone to NH3 for our

newest unit.

6) 6-12-11

My thanks to Al and Nate for the nice work/chemistry lesson on SO2 neutralization – just

one more example why it is so nice to have you two beautiful minds with us in ABPG!

-27-

SWS

Loss of SWS Bottoms Level SWS-027

1) 10-26-10

In general, redundant level indication is necessary to take credit for alarms, since false

level indication is a potential cause of abnormal level (in which case there will be no

alarm from that device).

2) 10-26-10

Sounds like a good place for a level transmitter tied into the DCS. We don't have any

SWSs though.

3) 10-26-10

The majority of our Sour Water Strippers send stripped water directly to the suction of

wash water or desalter booster pumps. I can’t think of any that go to the sewer in any

significant volume. Your case may be unique in this aspect, thus installing a second level

device is probably justified. Alarm the low level and deviation between the two levels.

4) 10-26-10

I agree with the previous recommendation, which mirrors our typical IPF schemes.

5) 10-26-10

We also typically go to the suction of a booster pump with a LCV on the discharge. All

of the stripped sour water is consumed in either the desalters or wash water systems and

then to the sewer via those LV's.

6) 10-27-10

Our newer strippers have an LC with local gauge and a SIS LALL/LAHH trip with a

deviation alarm. The dual LTs utilize different technologies – one magnetic float and one

DP cell.

7) 10-27-10

Sour water stripping is usually not critical as it is typically fed from tankage, thus tieing

feed to hold level is a common safeguard. Regarding stripped water reuse, this has been

a standard since my first NPRA Q&A (don't ask the year), but unfortunately it is not yet a

true standard industry-wide.

-28-

8) 10-28-10

Our standard is redundant LI's on different taps (plus a sight glass on one of these taps).

For PHA purposes, this gives the operators an alarm if the LC indication goes bad and

drives the level valve wide open. The pumpout time from the low level is supposed to

give the operators time to respond. To my knowledge, we have no systems with remote

operated block valves or where the bottoms level is tied into a trip.

9) 1-1-11

As an aside, attached (SWS-027A) is a worksheet I prepared to calculate the suppression

of DP-cell level transmitters – prompted by an incident involving a high TGU quench

tower level which failed to trigger an alarm due to miscalibration. (Alarm trip point was

80%, and transmitter pegged at 75%.)

10) 2-18-10

We have redundant level indication with low-low level bottoms pump trip on one SWS.

Another has full trap-out for the reboiler which would maintain clean bottoms. Loss of

level on the others would result in loss of pump suction and loss of flow indication.

NH3 Acid Gas Line Corrosion SWS-029

We experience chronic corrosion of the 10” sour water acid gas lines between SRU feed

KO and the junction of amine and NH3 acid gas to the SRU Reaction Furnace. Areas of

high corrosion at steam tracers forced replacement of the lines last year. Similar

corrosion leaks surfaced again last week on the new section of line.

The steam tracing is on the side of the line, insulation is 2” calsil, no indication of CUI.

The KO Drum was also replaced due to corrosion under some of its steam tracing runs.

No corrosion evident on the lines from the SWS Overhead Condenser to the KO Drum,

which have the same tracing and insulation. Inspectors confirmed corrosion is on the

process side.

Has anyone seen this, and if so what was the cause?

5-24-11

RESPONSES

1) 5-24-11

Assuming the line is carbon steel and the tracing is stainless, was there any galvanic

activity between the two at the elevated temps? What is the nature of the

corrosion? How much CO2 is there? How hot is the steam?

2) 5-24-11

-29-

My guess is localized condensation occurs at cooler pipe wall areas and flashes at the

tracers. Controls Southeast can design ControTracing to maintain entire pipe wall above

the dew point.

3) 5-24-11

I have recently had reinforced to me that you generally need water present to have

corrosion. A good question in this case is where is the water coming from. Is the

corrosion external between the piping and the steam tracer? If so, you are somehow

getting moisture in this location. We have had issues with internal corrosion on lines that

have a steam tracer running along the bottom or the side (4:00 / 8:00 position) especially

if there is any period of stagnant flow.

4) 5-24-11

Basis some prior experience with dew point corrosion in some of our SRU gas lines, I

would agree with some previous responses that one likely explanation is inadequate heat

to prevent condensation on the pipe, and have also experienced mitigation of high

corrosion rates with ControTrace.

One big problem with steam tracers is poorly applied heat transfer material between the

tracers and pipe. And obviously, poor insulation installation and weather-proofing.

5) 5-25-11

Did you have any carryover events (that anyone will admit to)? Going above 180°F with

liquid water with this amount of NH3 holding a large quantity of H2S in the water will

really accelerate corrosion rates.

6) 5-25-11

Followup: The NH3 acid gas is nominally equimolar H2S, NH3, H2O with ~ 0.2% CO2.

See attached photos, followed by a P&ID markup showing areas of relative corrosion

activity (SWS-029A).

7) 5-25-11

From your pictures, I agree with the ControTrace recommendation.

8) 5-26-11

I also support ControTracing. This is similar to "wet-dry, erosion/corrosion" that is

common in a number of other processes. The protective scale is probably continually

forming, when the pipe wall is dry, and being washed away when condensation from the

colder sections of pipe pass over the affected areas. This is one of the more aggressive

forms of corrosion and can lead to leaks in a matter of weeks or months, in some cases.

-30-

9) 5-27-11

We recently had what may be a similar corrosion mechanism occurring on steam traced

lean amine lines, internally right where the tracer was touching the line. The materials

specialist indicated the mechanism likely due to nucleate boiling from the high localized

temperature resulting in repeated formation and stripping of the FeS layer. In the SWAG

application we utilize ControTrace or have also had good success with EHT to ensure a

more uniform heat distribution around the pipe wall.

Water Vapor Calc SWS-030

Does anyone have a handy formula for approximating partial pressure of water in

NH3 acid gas as a function of temperature?

I have only recently learned that steam tables overstate water partial pressure because

dissolved H2S/NH3 suppress the water vapor pressure.

7-19-11

RESPONSES

1) 7-19-11

Hysys, ProMax or other simulation package should give a fair representation. I don't

have any quick fixes.

2) 7-19-11

I don't have a formula, but do use the following graph:

-31-

3) 7-19-11

I developed the following approximation by regressing a load of simulations to determine

how much NH3 and H2S would be held in the reflux vs. temperature:

PH2O = XH2O * PSAT H2O

PSAT H2O is the vapor pressure of water, which can be found by the Antoine equation.

XH2O is the mole fraction of water in the liquid, which depends upon the solubility of

NH3 and H2S in the liquid. I found this to be a fairly simple relationship.

4) 7-23-11

A colleague graciously prepared the following plot basis equimolar H2S/NH3:

-32-

0

5

10

15

20

25

30

35

40

160 165 170 175 180 185 190 195 200

Temperature - °F

Mo

l %

Wate

r

13psig

15psig

17psig

19psig

LLD

Excerpt -- Burner Light-off Sequence LLD-015

Basis discussions while conducting my SRU training, there can be a lack of

understanding by some operations/technical personnel of the critical importance of

correct sequencing of lighting a burner. The key issue can be simply addressed by

considering the auto-ignition temperatures and explosive ranges of the following

combustibles:

NH3: 1200°F, 16-25 vol-% in air

H2: 1075°F, 4-75 vol-% in air (and very high flame speed!)

CH4: 1000°F, 5-15 vol-% in air

H2S: 500°F, 4.3-46 vol-% in air

Liquid sulfur: 450-480°F

Sequencing

From the table, it becomes clear that when lighting a natural gas fired burner, it is

imperative to start the air flow and then the gas flow. For example, if air is started

-33-

first, the natural gas has only to increase its concentration to ~ 5 vol-% in order to reach a

combustible mixture. If, on the other hand, natural gas is started first, there can be a significant “gas bomb”

created before the concentration is reduced by air down to 15 vol-% or less.

One paradigm trap that may have led people to start natural gas flow first is the

admonition that the SRU must not have free oxygen in contact with the catalyst, hot

sulfur, pyrophoric iron, etc. While permitting free O2 to enter the SRU is bad, avoidance

of an explosion is more important. Actually, if natural gas is started first, a significant

amount of air must flow to lower the gas concentration from 100 % all the way down to ~

15 fol-%. Likewise, as soon as the much larger air valve starts to open, the smaller gas

valve can quickly catch up and raise its concentration to the 5 vol-%+ required for

ignition.

Ignition

From the table, it is apparent that attempting to light a natural gas/air mixture without a

pilot/ignitor is significantly more dangerous than with H2S. Attempting to so light an

H2S/air mixture is less dangerous due to both the relatively low auto-ignition temperature

and broader explosive range. Still, I would not support attempting to “light off the wall,”

particularly with the highly reliable pilot/ignitors now on the market.

One additional point

When setting up the temperature control loop, my experience is that the burner

control/responsiveness is significantly improved by having the TC set the air flow and

have the natural gas ratio controlled to the air flow. After all, the air flow has to move

about 9.5 times more flow than the natural gas flow.

8-27-10

[ABPG DEN Details – Brimstone Sulfur Symposium at Vail 2011.doc ]

LHS/08-25-11