I will stress "the correct way" so don't expect short cuts and rules of thumbs. Notice, I use the word "I" a lot. These will be my thoughts, my ideas. I will present facts and in instances, my interpretation of the facts. I might even editorialize. This inaugural column starts a series on Relief Valves.

The Problem

In the May 2000 issue of Chemical Engineering Progress¹ (CEP), there was an article entitled "Ease Relief System Design and Documentation". While I'm not intending to discuss the article in whole, the author stated something that appeared to get one reader's attention. The author wrote,

If you want to reduce the size of a relief device for cost savings, then design it at a higher set pressure; however, the MAWP of the weakest link should not be ignored.

This statement prompted a "Letters to the Editor" in the October 200 issue of CEP² (no, it wasn't me) where a reader wrote,

This is not true; for a certain MAWP, the capacity of the relief device is not a function of its set point, but of MAWP alone. For example, for a MAWP of 100 psig, the relief valve capacity will be same whether it is set at 80 psig or 100 psig. In both cases, the maximum relieving pressure for the ASME non-fire case (or BS 5500 fire case) is 124.7 psia and the discharge capacity will remain identical. The only difference is that if the set point is 80 psig, the allowable overpressure will be 37.5%, while, at the same for a set point of 100 psig, it will be 10%. For the ASME fire case, the values will be 51.25% and 21% respectively. These are defined very clearly in API 520, Tables 2 to 6.

I'm seeing that people do not quite understand what API 520³ and the ASME Boiler and Pressure Vessel Code, Section VIII, Division 1⁴, are really saying. (For those not familiar with API and ASME, ASME, or the American Society of Mechanical Engineers, is the organization that sets the codes in the United States that determine how pressure vessels are to be designed and protected. These codes are law and must be followed. The American Petroleum Institute, or API, sets the standards by which the codes are followed. API publishes the Recommended Practices 520 and 521, among others.)

MAWP and Design Pressure

In paragraph 1.2.3.2, para b, API 520 defines maximum allowable working pressure (MAWP) as:

... the maximum gauge pressure permissible at the top of a completed vessel in its normal operating position at the designated coincident temperature specified for that pressure.

The operative word here is "completed". {parse block="google_articles"}The vessel is completed when a fabricator, according to the code laid down by ASME, has designed it. The vessel's fabricator, not the Process Engineer, determines MAWP. (Some may try to stretch my definition of "completed" to mean that the vessel is also erected in place. Not quite because the certified vessel drawings, which are delivered way before the vessel is, contains this information).

In the same paragraph, API 520 says that the MAWP is normally greater than design pressure. The Process Engineer usually sets the design pressure at the time the vessel specification is being written. The design pressure is the value obtained after adding a margin to the most severe pressure expected during normal operation at a coincident temperature. Depending upon the company the engineer works for, this margin is typically the maximum of 25 psig or 10%. The vessel specification sheet contains the design pressure, along with the design temperature, size, normal operating conditions and material of construction among others. It is this document that will eventually end up in a fabricator's lap and from which the mechanical design is made.

Relief Valve Set Pressure

Unfortunately, project schedules may require that relief valve sizing be carried out way before the fabricator has finished the mechanical design and certified the MAWP. The Process Engineer must use some pressure on which to base the relieving rate calculations. In paragraph 1.2.3.2, para. c, API 520 states that the design pressure may be used in place of the MAWP in all cases where the MAWP has not been established. Guess what pressure the Process Engineer usually sets relief valves at? There are even times when the relief valve must be set even lower than design pressure. For example, a high design pressure may be desirable for mechanical integrity but a PSV set at the design pressure may end up with a coincidental temperature that would require the use of exotic materials of construction or that promotes decomposition and/or run-away reaction.

So, Why the Confusion?

The confusion faced by the reader who wrote the "Letters to the Editor", and probably many others, is due to a number of reasons. First and I think foremost, is the way ASME does not relate the maximum allowable pressure limits to relief valve capacity. ASME, Section VIII, Division 1, refers to MAWP throughout the entire document when talking about relief valve set pressure and allowable overpressure. I believe the reader may have been referring to and interpreting what is stated in paragraph UG-125 of ASME Section VIII, Division 1. It states in part,

All pressure vessels other than unfired steam boilers shall be protected by a pressure relief device that shall prevent the pressure from rising more than 10% or 3 psi, whichever is greater, above the maximum allowable working pressure except as permitted in (1) and (2) below.

Sub-paragraphs (1) and (2) mention cases where the pressure rise may be higher.

However, when ASME talks about certifying the capacity of a relief device, MAWP is never mentioned. ASME Section VIII, Division 1 clearly states in Paragraph UG-131, para c(1) that

Capacity certification tests shall be conducted at a pressure which does not exceed the pressure for which the pressure relief valve is set to operate by more than 10% or 3 psi, whichever is greater, except as provided in para c(2)...

Sub-paragraph para c(2) covers a fire case. {parse block="google_articles"}Again, capacity certification is based only on the set pressure of the relief valve and is unrelated to MAWP, unless of course the set pressure is MAWP.

Another area of confusion might involve the definition of capacity and how the term is used in ASME and API. Relieving rates are determined from "what can go wrong" scenarios and if allowed to go unchecked, would overpressure the vessel. Once the Process Engineer determines the controlling relieving rate from all the scenarios, the required relief valve orifice size is determined using the appropriate equation given in API. Once the required relief valve orifice size is calculated, an actual orifice size equal to or greater than the calculated orifice size is chosen from a selection available from a particular manufacturer. The maximum flow through this actual valve will be the valve's capacity.

Conclusion

The problem and solution can be summarized as follows:

CODE AS WRITTEN
Capacity based on Set Pressure + Allowable Overpressure

Code clearly requires that the relief valve's capacity be based solely on set pressure and not on the vessel's maximum allowable working pressure. Indeed, as shown above, if the relief valve's capacity was based on MAWP, then code might even force the Process Engineer into an unsafe design. A good analogy is highway speed limits. In the United Stated, many highway speed limits are set for 65 miles per hour. This does not mean a driver cannot travel slower and, under certain conditions for safety, it is almost a necessity that one does.

If it is safe to do so and the protected vessel can be allowed to pressurize to a greater extent, the relief valve set pressure can be increased, thereby reducing the relief valve's size and cost. Remember also that there is piping and possibly downstream equipment to "catch" and process the relieving fluid associated with the relief valve which may also benefit by this reduction.

One way of accomplishing a reduction in relief valve size is by increasing the vessel's design pressure. There is an economic trade off here as the vessel's cost can increase above what you may save by reducing the size of the valve. Another approach to consider is increasing the relief valve's set pressure right up to MAWP after receiving the certified vessel drawings. However, depending on project schedule, the cost savings may be offset by the high costs associated with late design changes.

Final Say

I welcome and encourage your feedback. Feel free to E-Mail me at the Internet address below. All correspondences that include a name will be published in this column. Better yet, I encourage discussion of any topic I cover utilizing The Chemical Engineers' Resource Message Board. This will enable the entire Internet community to join and learn.

References

1. Ahmad, S.
2. Letters to the Editor
3. API
4. ASME
"Ease Relief System Design and Documentation," Chem. Eng. Progress, pp 43-50 (May 2000) Chem. Eng. Progress, p 10 (October 2000) (www.api.org) Recommended Practice 520, "Sizing, Selection, and Installation of Pressure-Relieving Device in Refineries, Part 1-Sizing and Selection", 7^th Edition (January 2000) (www.asme.org) "Boiler and Pressure Vessel Code, Section VIII, Division 1" (1998)


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Author/Community Member Discussion Regarding This Article

The following discussions were added July 21, 2001.

Discussion One

From Mr. Jeffrey Niemeier:

Philip,

I am responding to your column in Cheresources.com. {parse block="google_articles"}I disagree with your interpretation of the ASME code. The capacity of a relief device that is used to determine adequacy of design is based on the allowable overpressure. If the overpressure is higher the flow will be higher. You can take credit for this. UG-125 makes it clear that the only requirement is that the pressure not exceed 110% of the MAWP (121% for a fire). The stamped capacity is there only for reference. It could not possibly be used to make a judgement on two-phase flow capacity.

Also, contrary to what you have in your article it is many times advantageous to have a set pressure much lower than the MAWP. This is especially true if runaway reaction is a possibility. A low set pressure allows the reactants to start venting much earlier, thereby removing reactant from the vessel before the temperature and reaction rate get too high. Check out the DIERS literature.

I think you should remove your column.

Philip Replies:

Mr. Niemeier,

Thank you for taking the time to read my article "Relief Valve Set Pressures" in the "Process Engineering-As I See It" section of "The Chemical Engineers' Resource Page and for your feedback.

Allow me to respond.

You write,
"I disagree with your interpretation of the ASME code. The capacity of a relief device that is used to determine adequacy of design is based on the allowable overpressure. If the overpressure is higher the flow will be higher. You can take credit for this. UG-125 makes it clear that the only requirement is that the pressure not exceed 110% of the MAWP (121% for a fire). The stamped capacity is there only for reference."

You are falling into the same trap that most people fall into and this is precisely why I wrote the article in the first place. The certified capacity which determines adequacy of design of a relief valve (which is the 'device' I assume you are referring to since the capacity of a stand alone rupture disk is different) is based on the criteria described in ASME Section VIII, Divison 1, Paragraph UG-131, not Paragraph UG-125. Paragraph UG-131 states that, except for some very specific exceptions, the overpressure is to be 10% or 3 psi, whichever is greater and is to be referenced to set pressure. If set pressure happens to be equal to MAWP, fine, but as you point out
later in your letter, this is not always the case. There are no provisions that I can find that allow a vendor to certify a relief valve for
overpressures greater than these. In addition, the certified capacity is the only flow that is guaranteed by the relief valve vendors, nothing more or less. Using greater overpressures than the 10% or 3 psi described above in calculations for required relieving rate does not alter the guaranteed capacity provided by the vendor and which is required by ASME, Paragraph UG-129.

If you require further evidence of these points, feel free to contact Farris Engineering, a well-known relief valve vendor at http://www.cwfc.com/. However, I can save you some time because I didn't write the article without doing some research first.

Now, you think capacity certification (stamped capacity) is just for reference?!? It is this certified capacity, not the calculated maximum relieving rate, which must be used when sizing inlet lines to the relief valve (using the 3% rule, ASME Section VIII, Division 1, Appendix M, Paragraph M-7) and in many cases, the outlet lines as well.

I invite you to re-read the article, specifically the paragraph titled "So, Why the Confusion? and also read ASME Section VIII, Division 1, Paragraph UG-131. The code is like a bowl of spaghetti. You have to weave through all the pertinent paragraphs to get to the end... and the full story.

For your information, the stamped capacity for relief valves in vapor/gas service is usually given in terms of SCFM of air at set pressure plus 10% overpressure and 60 degress F. Valves in steam service are stamped in terms of pounds per hour of steam at set pressure plus 10% overpressure at the saturation temperature. ASME Section VIII, Division 1, Appendix 11 gives you a means for converting to your particular vapor/gas. Valves in liquid service are given in terms of gallons per minute of water at set pressure
plus 10% overpressure and 70 degrees F.

You continued in your letter with the following (still referring to the stamped capacity):

"It could not possibly be used to make a judgement on two-phase flow capacity."

I don't know why you mention two-phase flow since I didn't in my article. However, I agree, relief valve manufacturers at the present time do not have capacity certification capability for two-phase flow.

You then added:

"Also, contrary to what you have in your article it is many times advantageous to have a set pressure much lower than the MAWP."

I invite you to re-read my paragraph titled, "Relief Valve Set Pressure" specifically the sixth line where I write, "There are even times when the relief valve must be set even lower than design pressure."

Are we talking about the same article here? Did you actually read my article?

You finish your letter with:

"I think you should remove your column."

MAWP, design pressure, relief valve set pressure, what is allowed overpressure, calculated relieving rates, certified capacity, etc. appears to be a very confusing and misunderstood concept that most people are having when dealing with safety relief systems; you included. And this is precisely why my column must not only stay but also grow. My only regret at this time is that I obviously did not get the point across to you and for that I apologize. I hope this response clarifies those points you misunderstood. I will be more than happy to answer any other concerns you may have or even further discuss the points you already brought up.

Mr. Jeffrey Niemeier Replies:

I still disagree with you. It is basic physics that you will get more flow if you have a higher driving pressure. There is no reason not to take credit for overpressures above 10% as long as you don't go above the maximum allowable pressure for the valve. See the following memo from Paul Papa, director of engineering for Farris:

As you mentioned, all of our certified capacities are based on testing that is done at 10% overpressure. Nameplates are therefore marked with capacities at 10% overpressure in the appropriate certification fluid (steam, air, or water). From a selection standpoint, valves are typically sized based on 10 % overpressure. As you correctly pointed out, Section VIII of the ASME Code indicates that ASME Code stamped pressure vessels require a relief device that must pass all of the required flow without allowing the pressure to go any higher than 10% above the maximum allowable working pressure (MAWP).

There are many cases where a vessel is being used at a pressure at much less than its MAWP. In those cases, you can size and select the valve at an overpressure greater than 10%. This will allow you to use a smaller valve as
the valves capacity will increase as the pressure increases. Per ASME Code, the pressure must still be kept to the 10% accumulation pressure, that is no more than 10% above the MAWP.

For the most part, this is not a problem for the valve as long as the higher overpressure does not exceed the maximum pressure limit of the valve at the relieving temperature.

Philip Replies:

Jeff,

I'm not saying, nor have I ever said that you can't calculate the required relieving rate for the MAWP + overpressure and then choose a valve based on this. Remember fire cases? They go to 21% overpressure (set pressure basis) don't they and are often the sizing criteria for the valve. However, this still does not alter the guaranteed (stamped capacity) flow rate of the valve which is based strictly on set pressure + 10% (for the most part). And it is this flow that your maximum calculated relieving requirement cannot exceed and which is subsequently used in the hydraulic calculations, per ASME. I don't know how many times I have to say this. And this is what my article was strictly about; the criteria used to obtain the certified (stamped) capacity. Also, I also don't see in the Farris response you sent me where they will guarantee any number other than what they are allowed to stamp even if MAWP is 100 psig and you want to allow the pressure to go to 110 psig. If a valve is set for 50 psig, it will be certified based on 55 psig, clear and simple! Neither Farris nor any other vendor will guarantee that your valve will pass a different flow rate at any other condition. Nor will they tell you to use any other value in subsequent hydraulic or downstream sizing equations.

I would love to see you prove the calculations that give you that extra margin you keep talking about. You think it's simple? Try doing the calculations at a high pressure for something very non-ideal such as ethylene. All of the equations shown in ASME and API fall apart rather severely since they are all based on ideal gas. Try convincing an OSHA inspector that at conditions which exceeded the capacity certification test at the time of equipment failure that your valve was properly sized and should have been big enough. If you can do this via flow equations derived properly for non-ideal gases from the sound use of thermodynamics, great and there is no argument from me. Heck, I'll even write about it!! After all, ASME and API specifically state that the engineer is to use good engineering judgment and if in question, work with the relief valve vendor. As a matter of fact, you can even deviate from the code if you can prove that a given relief valve will not be adversely affected from whatever you are trying to do. As an example, I am currently on a project where the inlet line loss for one particular valve is in the 6% range. Notwithstanding the fact that the 3% rule is "non-mandatory" (but is still considered good engineering practice), we are working with Farris to determine if this would adversely
affect the valve. If it doesn't, we can live with the current piping configuration.

What I wrote was about stamped capacity and how this is what is recognized by Code, nothing more or less. Stamped capacity being the only guaranteed flow rate is fact, not interpretation. When to use stamped capacity has been set down by both Code and interpretation. You can disagree all you want with what I wrote. That is fine and I just love getting into these types of discussions. How else can we learn and grow? But you made some rather inaccurate statements in your first letter about what was stated, or not
stated in the article and that still leads me to feel you didn't totally understand it.

I've been giving this debate some further thought and I hope this will finally put it to rest. You and I seem to be disagreeing on the concept of sizing versus what you do with the valve once it has been sized.

My article on the Web site starts off with a statement made by the author of an article published in the May 2000 issue of CEP. This person made a statement that basically said if you want to decrease the size of a given relief valve, find a way to increase it's set pressure. However, the following Letter to the Editor in the October 2000 issue inspired my article. It stated, "This is not true; for a certain MAWP, the capacity of the relief device is not a function of its set point, but of MAWP alone." Now, had you really read my article, you would have seen these quotes! As I've said before and written in my article, per ASME, capacity for a given valve (which I am defining as certified capacity) is a function of set point only, not MAWP. Therefore, to increase the certified capacity of a given
size valve, one can increase the set pressure. Alternatively, to reduce the size and cost of the purchased valve, one can increase the set pressure of a smaller valve and maintain the same certified capacity.

The whole point should be "safety", not who can buy the smallest relief valve. If you want to stretch what ASME is trying to convey just to buy a smaller relief valve, go ahead; you don't work for me. If you did work for me, I wouldn't allow it. I'll take the conservative approach, it just isn't worth the potential liability. After all Jeff, you'll never know if the valve you bought is the right one unless the system over pressure is caused by the controlling scenario and the valve does or does not work. And guess what? This rarely happens since most controlling scenarios are loaded with conservative assumptions that are never achieved in real life.

One last point if I may. I still can't understand why you would want to set the relief valve at a lower pressure (50 psig) and still allow the system to achieve the much higher pressure (110 psig). This makes no engineering sense to me. You should just specify the set pressure at MAWP and go with the allowable overpressures. Then you won't have any debates. You mention run away reaction? I'm sorry but this is bad engineering practice and if you do this at your site, I would stop the practice. Run away reactions can occur
too fast for relief valves to react. I would use a stand alone rupture disk set at the lower burst pressure (your 50 psig example) for the run away reaction scenario and use the relief valve for other scenarios set at the 100 psig MAWP. This is a much safer design.

Discussion Two

From Mr. Don Gregurich:

Dear Mr. Leckner,

I read your about Relief Valve Set Pressures with interest because the issue came up several years ago and we had quite an internal debate (at a different company). {parse block="google_articles"}I don't remember the details, or resolution, but I was on the opposite side of the argument. Essentially, the situation is that there is a vessel with a design pressure of 150 psig (and let's assume same MAWP) and we want to use a relief valve set at 30 psig (There could be a number of process reasons for doing so.) Then, if this is the only relief on the vessel, does it need to be sized to accommodate the worst credible non-fire overpressure case flow at a relieving pressure of 33 psig or 165 psig? I felt that the answer was 165 psig, if, of course, all of the design was appropriate (valve body can handle the pressure/temperature at the relieving conditions, the inlet and outlet piping are sized correctly for those relieving conditions etc.) It certainly seemed to me that this was consistent with the spirit of the code, after all, the whole point is to limit the pressure in the vessel and this would meet that requirement.

Of course we must also be consistent with the letter of the code, if we are to avoid jail time, but I didn't see anything in the code that contradicted the above interpretation. Although I see your quote regarding rating of valves, don't see where it explicitly states that a valve can not be operated above its "official rated capacity" to perform the required service. I see that the code states that "the valve must have the capacity to relieve the load*" and that "the official rated capacity is that which is stamped*" but I don't see where it says that the valve must be able to relieve the load at the official rated capacity. So if I look at this in the strictest sense (by the letter of the code) I do not see that we can't size it based on the 165 psig. I don't think this is unreasonable: rating a pump at a given pressure/flow point serves to define the pump's capacity - but the pump can operate at different flows and pressures then at the rated point. Likewise for the relief valve. In fact, we could have selected this same valve, used a spring for 150 psig, and it would be functioning the same at the relieving conditions, as if it would with the 30 psig spring.

I confess that I haven't taken the time to reread through the code, but I was convinced back then that my interpretation met the spirit and letter of the code. I'd be interested to hear what you have to say about my position.

Thanks

Don Gregurich

Phlip Replies:

Don,

First, call me Phil and thanks for taking the time to read my column. Second, I do not agree with your position even though on the surface it appears to be sound, and this is why.

First (and this really doesn't answer the question but I just couldn't help but comment) I couldn't figure out why anyone would want to have such a low set pressure and still allow the relieving pressure to go so high. It is not logical and seems to defeat the purpose of having a low set pressure. You might as well have just set the valve for 150 psig and taken the conventional 10% overpressure and eliminate the controversy. But then I thought of a run-away reaction scenario where you may want the valve to pop open early and allow the vessel to slowly relieve its contents but still allow the pressure to build-up. However, this uses the relief valve as a
depressurization valve and I am not in favor of this. There are designs with proper use of control valves to do this.

In your E-Mail, you say "Then, if this is the only relief on the vessel, does it need to be sized to accommodate the worst credible non-fire overpressure case flow at a relieving pressure of 33 psig or 165 psig? I felt that the answer was 165 psig,...".

It is very clear that the valve does not "need" to be sized for 165 psig. It only "needs" to be sized for set pressure +10% overpressure (non-fire, single device case) in order to be consistent with the calculation of the certified (stamped) capacity. However, you are correct in that on the surface (at least as far as I can tell), the code (ASME Section VIII, Division 1, 1998 Edition) does not seem to prohibit one from sizing the valve based on such a high overpressure. ASME appears to only be concerned that the MAWP is not exceeded beyond its requirements. And as you say in your E-Mail, "I don't see where it says that the valve must be able to relieve the load at the official rated capacity."

However, a relief valve vendor will only guarantee the stamped capacity and ASME is very clear how this is to be determined. Therefore, in the event of an accident, how would you have guaranteed to an OSHA inspector that the catastrophic failure of the vessel or attached piping/equipment was not a result of an improperly sized relief valve since you have no guarantee of the flow rate through the valve? I guess you could try to prove that your calculations were reasonably accurate using real properties, accurate vapor
flow equations and the correct thermodynamics. Is all this really worth the potential liability just to buy a somewhat smaller relief valve? Not where this Process Engineer stands! So Don, that's Process Engineering-As I See It!

Mr. Don Gregurich Replies:

Thanks Phil,

I appreciate your insight; I understand what you're saying about the rated capacity of the valve as the only real figure that you can hang your hat on.I honestly can't remember the details of the situation, only that it never did get to the point where we needed to make a final decision on the set point issue. It's those applications that lie outside of the norm that really make us think about what we're doing, which hopefully leads to a better understanding.

I do have another one for you, this one also seem "sensible" to me, but I don't know what the code would think:

If an ASME stamped pressure vessel is being used for an "atmospheric" operation with an open vent to atmosphere that has been properly sized, is a relief valve (or disk) required? "Proper sizing" of the vent line means application of the same analysis and calculations that would be done for sizing of a relief device. The vent line would be sized to accommodate the worst credible upset condition at a relieving pressure that is under 15 psig. The location is in a jurisdiction that requires compliance with the ASME code.

Thanks again,

Don

Philip Replies:

Don,

I am not on any committee so I can't give you an official answer. However, ASME specifically states in Section VIII, Div 1, 1998 Edition, paragraph UG-125(a)"All pressure vessels within the Scope of this Division, irrespective of size or pressure, shall be provided with pressure relief devices in accordance with the requirements of UG-125 through UG-137." The key here is what they consider part of the "Scope". The Indroduction, Paragraph U-1, goes into the definition of "Scope" and it can get rather complex. I would have to read through this very slowly and carefully to see if this exact situation is addressed. I do have some other sources I can
review for interpretations and your question may not be able to be answered without a direct interpretation from the ASME committee. I would agree with you that it doesn't make sense to need a pressure relief device for this situation. I can tell you that in the past, when facing a situation where we do have an ASME coded vessel but cannot come up with any credible scanrio, nothing at all, we put a 3/4" x 1" relief valve on the vessel and call it out for thermal expansion - end of story.

I think I have an answer for the second question you asked me concerning an ASME stamped pressure vessel being used for an "atmospheric" operation with an open vent to atmosphere. Based on ASME Section VIII, Division 1, Introduction, Paragraph U-1 para c(2)(h), vessels having an internal or external operating pressure not exceeding 15 psi (103 kPa) would not fall under the scope even though it has a stamp. As long as any pressure vessel meets all the applicable requirments of the Division, it may carry the Code U Symbol.

The Problem

Accidents not only damage equipment but also cause injury or even death to plant personnel. To reduce the number of incidents of accidents, it is the job of the Process Engineer to analyze the process design, determine the "what can go wrong" scenario and either find a way to "design" out of it or provide protection against catastrophic failure in the event an accident does occur, i.e. install a relieving device such as a relief valve and/or rupture disk.{parse block="google_articles"}

For the purposes of this discussion and those that will follow, a "what can go wrong" scenario is defined as an action that will cause a vessel containing a gas or liquid to overpressure, leading to a catastrophic failure of that vessel if it were not for the presence of a relief valve or rupture disk. To find these potentially deadly incidences, the Process Engineer goes through a type of "self HAZOP" (Hazardous and Operability Study), analyzing the process to determine what these scenarios are. For each credible scenario identified, the Process Engineer performs calculations to determine the amount of vapor or liquid that must be relieved from the vessel in order to prevent the overpressure from occurring (the relieving load).

Fair warning, this discussion and those that will follow in future installments assume the reader is at least somewhat knowledgeable in Process Design. Safety analysis should never be left up to the junior Process Engineer unless closely supervised. So if you Mr., Mrs., Ms. Reader become confused or somewhat lost in some of the terminology used, I sincerely apologize. I welcome your very direct and specific questions about anything I write and would be happy to help you understand what I am saying. This is an extremely important function for Process Engineers.

The Checklist

Since there are many potential causes of failure, it would be nice to have a checklist to make the analysis organized and somewhat standard. After all, those of us in Process Design may be working on a project for a chemical plant today and end up on a project for a pharmaceutical plant tomorrow (it happens, believe me). A pretty good checklist is given by Table 2 in Section 3 of The American Petroleum Institute (API) publication, "Guide for Pressure-Relieving and Depressuring Systems"², better known as API Recommended Practice 521 (or just API 521). For those not familiar with API, this is the organization in the United States that sets the standards by which codes (laws) are followed. API publishes the Recommended Practices 520 and 521, among others. A condensed version of the API checklist is presented in Table 1 below.

In this installment, I want to establish a framework for analyzing a given process. In future installments, I will begin to tackle the scenarios (Overpressure Cause) themselves in some detail. The ultimate goal is for the Process Engineer to identify credible "what can go wrong" scenarios, perform relieving load calculations to prevent catastrophic failure and eventually size the relieving device and system.

The Concept of Double Jeopardy

I can't begin to tell you how many people still don't understand this concept. API allows you to "ignore" failures that fall under the "Double Jeopardy" principle (See API 521, March 1997, 4^th edition, paragraph 2.2). Double Jeopardy basically means two unrelated failures occurring at the exact same time, i.e. simultaneous. This does not mean the failures occurred one minute, one second or even one millisecond apart. It means at exactly the same instant in time! Let's look at some examples.{parse block="google_articles"}

There is a loss of power to a pump causing stoppage of cooling water to a condenser on a distillation column. Because vapor from the distillation column can no longer be condensed, pressure builds up to the point of popping the relief valve, i.e. the system goes into relief. At the same time, the control room operator strokes open a steam flow control valve to the reboiler on that same distillation column causing the generation of an excessive amount of vapor. When calculating the total amount of vapor that must be relieved to prevent damage, should one take into account the excessive vapor produced by the wide-opened steam valve? Or, should we consider only the normal amount of vapor exiting the column; API 521, paragraph 2.3.2 says that the control valve should be considered to be in its normal operating position unless its function is affected by the primary cause of failure, this being loss of power to a pump.

The answer is, this is a Double Jeopardy failure, two unrelated events occurring at the same time. One has nothing to do with the other. Therefore, you only need to calculate the relief load for one scenario at a time. For the loss of power to a pump scenario, the relief load would be based on the amount of vapor generated at the "normal" rate of steam to the reboiler.

Let's look at the situation again. With the pump stopped, cooling water is lost to the condenser causing the distillation column to go into relief. However, this time the control room operator realizes that the relief valve has opened and attempts to stop steam flow to the reboiler. The operator puts the steam control valve in manual and tries to close it but it won't respond because it is stuck. To free it, he strokes it wide-open, shooting steam into the reboiler and causing the generation of an excessive amount of vapor. Now we have two failures occurring at the exact same time but are now related. The power failure stops the pump and thus the cooling water to the column condenser. This causes the column to go into relief, which then causes the operator to react, initiating the second failure.

A very obvious example of a Double Jeopardy failure would be a tube rupture in the reboiler occurring at the same time cooling water flow was lost to the condenser. Two very unrelated failures occurring at exactly the same time.

By the way, stuck opened control valves occurring simultaneously with a second failure does NOT constitute Double Jeopardy. That valve may have been stuck in its operating position for a significant amount of time before the second failure has occurred. The first failure was the mechanical failure of the valve (sticking) and that did not happen at the same time as the second failure. These are unrelated failures but they do not occur simultaneously!

Being Conservative

There are times when a failure may be obvious, cooling water stops to a condenser. Then there are times when it will take some stretching to find any, a pressure vessel operating in a nonflammable, low-pressure system with little fluid throughput. There are three approaches you can take when analyzing your process. You can be CONSERVATIVE. You can be conservative or, as I like to think of myself, you can be CONSERvative. I follow API 520 and 521 to the letter as a minimum. If my company or the client I am working for happens to have more stringent rules, these obviously supercede what is in API documents. For example, API 521, paragraph 3.15.1.1 basically allows you to "ignore" heights above 25 feet when considering how much of a vessel to include in a fire zone calculation. I worked for a company that used 50 feet for its standard.{parse block="google_articles"}

I will perform the necessary relieving load calculations for any scenario that cannot be rationally explained away, even if there is only a remote possibility that the failure will ever occur. The alternative is to perform fault tree analysis, risk assessment, etc. If it can be shown that a given scenario is indeed a 1 million to 1 long shot, then I'll label it as being not credible. Until then, it's only a relatively small amount of time that needs to be spent to perform the necessary calculations to be safe.

In the case of analyzing for Double Jeopardy, this is where I get my most grief; my conservatism tells me to error on the safe side, the client's money tells me to make the scenario go away. If I feel that even a Double Jeopardy failure can lead to a loss of life or major equipment damage, I might go ahead and do the relieving load calculations anyways (this does not usually go over too well with the powers-to-be because this usually results in larger relief systems).

API 521, paragraph 3.4 states that one can take credit for operator response after 10-30 minutes. I stick with the higher end at all times.

When analyzing a system for failures of control valves, I will always assume all my valves will fail as they are intended (fail close will indeed fail close, fail open will indeed fail open) except for the one control valve that will cause an overpressure hazard! This valve I assume to fail in the opposite direction (fails closed if it is intended to fail open).

When analyzing a system for "what can go wrong" scenarios, plant instrumentation sometimes may be used to justify the elimination of some scenarios. For example, if I have a hard wired (opposed to a Distributed Control System-DCS) pressure interlock that will shut steam off to the reboiler when the column pressure rises to some predetermined value, and there are redundant pressure switches, I might consider cooling water failure to the condenser as not being a credible scenario. On the other hand, once a credible scenario has been established, you are never to take into account the use of instrumentation as a means of reducing the relieving load.

Final Say for this Installment

Always analyze the system as a whole. Don't get tunnel vision on one area of the process. Of course you will consider individual "what can go wrong" scenarios but remember the example I used when discussing Double Jeopardy. I considered two actual failures that combined into one. I found these two by considering the system as a whole.{parse block="google_articles"}

"What can go wrong" scenario analysis is a very important but complex process. I do not intend to cover every nuance associated with it (simply because it will be impossible). I also don't expect everyone to agree with my analysis for every API RP521 Item number (Table 1 above). That's the fun and scary part of doing this type of work. Some of it can be highly subjective as to what constitutes a credible scenario. I strongly suggest you get a copy of API 520³ and 521² and the ASME Boiler and Pressure Vessel Code, Section VIII, Division 1⁴ and try to read through them (the operative word here is "try").

I welcome and encourage your feedback. Feel free to E-Mail me at the Internet address below. All correspondences that include a name will be published in this column. Better yet, I encourage discussion of any topic I cover utilizing The Chemical Engineers' Resource Message Board. This will enable the entire Internet community to join and learn. And, don't forget to check back for future installments on this series.

References

1. "Boiler and Pressure Vessel Accidents Soar," Chem. Eng. Progress, p 13 (August 2000)
2. API (www.api.org) Recommended Practice 521, "Guide for Pressure-Relieving and Depressuring Systems", 4^th Edition (March 1997)
3. API (www.api.org) Recommended Practice 520, "Sizing, Selection, and Installation of Pressure-Relieving Device in Refineries, Part 1-Sizing and Selection", 7^th Edition (January 2000)
4. ASME (www.asme.org) "Boiler and Pressure Vessel Code, Section VIII, Division 1" (1998)

Before I begin, let me point out that most of what is included in this series of articles can be found in API RP520¹ and API RP521², and ASME Section VIII, Division 1³.  Much of what is found in these documents can also be found in vendor literature.

Why and When to Use Rupture Disks

Why Do We Use a Stand-Alone Rupture Disk?

A rupture disk is just another pressure relieving device. It is used for the same purpose as a relief valve, to protect a vessel or system from overpressure that can cause catastrophic failure and even a death.

When Do We Use a Stand-Alone Rupture Disk?

Some of the more common reasons are listed below. You may think of others.

Figure 1: Basic Components of a Rupture Disk

1.  Capital and Maintenance Savings: Rupture disks cost less than relief valves. They generally require little to no maintenance.

2.  Contents will be lost, but who cares? A rupture disk is a nonreclosing device, which means once it opens, it doesn't close. Whatever is in the system will get out and continue to do so until stopped by some form of intervention. If loss of contents is not an issue, then a rupture disk may be the relief device of choice. 

3.  Benign service: It is preferable that the relieving contents be non-toxic, non-hazardous, etc. However, this is not a requirement when deciding to use, or not use, a stand-alone rupture disk.

4.  Rupture disks are extremely fast acting: Rupture disks should be considered first when there is a potential for runaway reactions. In this application, relief valves will not react fast enough to prevent a catastrophic failure. A relief valve may still be installed on the vessel to protect against other relieving scenarios. Some engineers prefer to use rupture disks for heat exchanger tube rupture scenarios rather than relief valves. They are concerned that relief valves won't respond fast enough to pressure spikes that may be experienced if gas/vapor is the driving force or liquid flashing occurs.

Figure 2: Rupture Disk Mounted on a Vessel

5.  The system contents can plug the relief valve during relief: There are some liquids that may actually freeze when undergoing rapid depressurization. This may cause blockage within a relief valve that would render it useless. Also, if the vessel contains solids, there is a danger of the relief valve plugging during relief.

6.  High viscosity liquids. If the system is filled with highly viscous liquids such as polymers, the rupture disk should seriously be considered as the preferable relieving device. Flow through a relief valve will be very difficult to calculate accurately. Also, very viscous fluid may not relieve fast enough through a relief valve.

Cost Comparison Between Comparable Stand-Alone Rupture Disk and Relief Valve

Rupture disk manufacturers burst at least two disks per lot before shipping them to a customer. {parse block="google_articles"}As a consequence even if you want just one rupture disk you will be buying three. Therefore, the first usable rupture disk is comparatively expensive. Also for new installations, each installed rupture disk must be purchased along with a holder. However, the same holder may be used for replacement purchases as long as you buy the exact same rupture disk from the same manufacturer.

Below is a capital cost comparison between Continental Disc Corp. (www.contdisc.com) 3" Ultrx Hastelloy C rupture disks with holders and Farris Engineering (www.cwfc.com) 2600 series relief valves, based on a budget estimate in year 2001 dollars.

 

This capital cost comparison will vary considerably with size and material of construction but you get the point. However please note that everything has a value and the loss of contents should be considered in the overall cost difference between a rupture disk and a relief valve.

When Do We Use a Rupture Disk-Relief Valve Combination?

Rupture disks are often used in combination with and installed just upstream and/or just downstream of a relief valve.  You may want to choose the combination option if:

Figure 3: Rupture Disk- Relief Valve Combination

 

1.  You need to ensure a positive seal of the system (the system contains a toxic substance and you are concerned that the relief valve may leak). Application: rupture disk installed upstream of the relief valve.

2.  The system contains solids that may plug the relief valve over time. Remember, the relief valve is continuously exposed to the system. Application: rupture disk installed upstream of the relief valve.

3.  TO SAVE MONEY!  If the system is a corrosive environment, the rupture disk is specified with the more exotic and corrosion resistant material. It acts as the barrier between the corrosive system and the relief valve. Application: rupture disk installed either upstream and/or downstream of the relief valve.

Below is a capital cost comparison between combination Hastelloy C rupture disks with stainless steel relief valves and three stand-alone Hastelloy C relief valves. Again, this is based on a budget estimate in year 2001 dollars using Continental Disc Corp. rupture disks and holders and Farris Engineering relief valves.

Combination of Hastelloy C Disk and SS Relief Valve

Single Installation Total = $10,200

Total for three installations = $27,140

Three stand-alone Hastelloy C relief valves = $40,200

Summary

A stand-alone rupture disk is used when:

1. You are looking for capital and maintenance savings

2. You can afford to loose the system contents

3. The system contents are relatively benign

4. You need a pressure relief device that is fast acting

5. A relief valve is not suitable due to the nature of the system contents

A rupture disk / relief valve combination is used when:

1. You need to ensure a positive seal of the system
   
2. The system contains solids that may plug the relief valve over time
   
3. TO SAVE MONEY!  If the system is a corrosive environment, the rupture disk is specified with the more exotic and corrosion resistant material

References

1. API (www.api.org) Recommended Practice 520, "Sizing, Selection, and Installation of Pressure-Relieving Device in Refineries, Part 1-Sizing and Selection", 7^th Edition (January 2000)
2. API (www.api.org) Recommended Practice 521, "Guide for Pressure-Relieving and Depressuring Systems", 4^th Edition (March 1997)
3. ASME (www.asme.org) "Boiler and Pressure Vessel Code, Section VIII, Division 1" (1998)

Before I begin, let me point out that most of what is included in this series of articles can be found in API RP520¹ and API RP521², and ASME Section VIII, Division 1³. Much of what is found in these documents can also be found in vendor literature

Sizing

Sizing the rupture disk is a two-part procedure. First, determine how much flow the rupture disk needs to pass. Then determine how big it needs to be.{parse block="google_articles"}

How much flow does it need to pass?

Answering this question is the same as determining the required relieving rate for the system. There is no difference between determining the relieving rate for a rupture disk and a relief valve. They both require a set pressure (burst pressure for rupture disk), an allowable overpressure, an evaluation and calculation of the required relieving rate for each credible scenario and then choosing the flow rate associated with the worst-case scenario. Determining the controlling relieving rate is a paper in of itself and I will not attempt to get into details here.

How Big?

There are two recognized methods that can be used to answer this question, the Resistance to Flow Method or the Coefficient of Discharge Method.

Resistance to Flow Method

The Resistance to Flow Method analyzes the flow capacity of the relief piping. The analysis takes into account frictional losses of the relief piping and all piping components. Calculations are performed using accepted engineering practices for determining fluid flow through piping systems such as the Bernoulli equation for liquids, the Isothermal or adiabatic flow equations for vapor/gas and DIERS methodology for two-phase flow.

Piping component losses may include nozzle entrances and exits, elbows, tees, reducers, valves and the rupture disk (note that the rupture disk and its holder are considered a unit). Let me emphasize that in this method, the rupture disk is considered to be just another piping component, nothing more, and nothing less. Therefore the rupture disk's contribution to the over all frictional loss in the piping system needs to be determined. This is accomplished by using "Kr", which is analogous to the K value of other piping components. Kr is determined experimentally in flow laboratories by the manufacturer for their line of products and is certified per ASME Section VIII, Division 1³. It is a measure of the flow resistance through the rupture disk and accounts for the holder and the bursting characteristics of the disk.

Below is a list of some models of Continental Disc Corporation rupture disks with their certified Kr values⁴.

 

For comparison, the following is a list of some models of Fike rupture disks with their certified Kr values⁵.

If at the time of sizing the manufacturer and model of the rupture disk are unknown, there are guidelines to help you choose Kr. API RP521² recommends using a K of 1.5. However, ASME Section VIII, Division 1³ states that a Kr of 2.4 shall be used. Which one? Remember that ASME is Code (meaning LAW for the most part) and API is a recommended practice. In addition, as can be seen in the tables above, even ASME may not be as conservative as you may think. Therefore, it is in the engineer's best interest to determine ahead of time the manufacturer and model of the rupture disk that eventually will be purchased. This can be done without knowing the exact size, as Kr is more manufacturer and model specific than size specific (see above tables). If a number of manufacturers are on the allowable purchase list, then at the very least choose the most likely models you would buy from each manufacturer and use the largest Kr from that list. This will be a significantly better guess than just using guidelines.

Once the piping system is laid out and all the fitting types are known, including the rupture disk, the engineer can proceed with the calculations in the following manner (presented here as a suggestion, there are many ways to do it).

1. Known are the two terminal pressures, these being the relieving pressure (upstream) and the downstream pressure (a knock-out pot, atmosphere, etc.).
   
2. Also known are the fluid properties and required relieving rate (the flow the rupture disk needs to pass).
   

3. Choose a pipe size. This will be the size to use for all components, including the rupture disk.

4. For vapor/gas or two-phase flow, use one of the accepted calculation methods to determine the maximum flow through the system. The maximum flow through the system is commonly known as critical flow or choked flow. For liquids, use the Bernoulli equation to calculate the flow that will balance the system pressure losses.

5. Per ASME Section VIII, Division 1, multiply this flow by 0.9 to take into account inaccuracies in the system parameters. Compare the adjusted calculated flow to the required relieving rate. If it is greater, then the calculation is basically done. However, the next smaller line size should also be checked to make sure the system is optimized; you want the smallest sized system possible. If the adjusted calculated flow is less than the required relieving rate, the pipe is too small, choose a larger size and repeat the calculations.

Why not just choose a large Kr? Isn't that more conservative?

Many times, relief is not to atmosphere but to some downstream collection and treatment system, e.g. knockout drums and flares or thermal oxidizers. These are more often than not specified at a time period in the design that predates the actual purchase of the rupture disk. The flow used to size this equipment will be based on the capacity of your relief system as determined above.

If the rupture disk contributes a significant portion of the frictional losses to the system, a fictitiously large Kr might result in an oversized piping system. Sounds all right on the surface but once the actual rupture disk is chosen, the calculation must be repeated with the "real" Kr and this may be a much lower value than originally used. More fluid will flow through the system than previously determined because there will actually be less resistance to flow. The result is that the downstream processing equipment may have been undersized.

The opposite is also true. An initial guess of a fictitiously small Kr might ultimately result in oversized downstream equipment and the excessive expenditure of a significant amount of money.

Atmospheric discharge must also be similarly analyzed because the flow capacity determined after rupture disk selection may have a major impact on the emissions reported for permitting if they were based on the initial value of Kr.

Coefficient of Discharge Model

The second calculational method is the Coefficient of Discharge Method. The rupture disk is treated as a relief valve with the flow area calculated utilizing relief valve formulas and a fixed coefficient of discharge, ‘Kd', of 0.62. This method does NOT directly take into account piping so there are restrictions in its use. These restrictions are known as the "8 & 5 Rule" which states that in order to use this method to {parse block="google_articles"}size the rupture disk ALL of the following four conditions MUST be met³:

The rupture disk must be installed within 8 pipe diameters of the vessel or other overpressure source.

1. The rupture disk discharge pipe must not exceed 5 pipe diameters.
2. The rupture disk must discharge directly to atmosphere.
3. The inlet and outlet piping is at least the same nominal pipe size as the rupture disk.

A sketch of the "8 & 5" rule starting with a 2" nominal sized pipe is shown at the below.

The flow area calculated with this method is called the Minimum Net Flow Area or MNFA. The MNFA is the rupture disk's minimum cross

Figure 1: Diagram Showing the "8 & 5" Rule

sectional area required to meet the needed flow.

This is not the area (and thus the size) you specify. Just like a pipe with a nominal size and an actual inside diameter, the rupture disk has a nominal size and an actual Net Flow Area or NFA. The rupture disk purchased must have a NFA equal to or greater than the MNFA. The manufacturer publishes the NFA for every rupture disk model and size they sell. The NFA also accounts for bursting characteristics of the disk and the holder.

Below is a list of some Continental Disc Corporation rupture disks with their NFA⁴.

 

Once the actual NFA of the rupture disk is determined, the calculations must be repeated, basically for the same reasons discussed above for the Resistance to Flow Method.

Why I Don't Like the Coefficient of Discharge Model

• It's too restrictive! During the basic design phase of a project, actual piping configuration is unknown. You may think you are within the "8 & 5" rule at first but may not be when the final details are worked out. Remember, the "5" means 5 pipe diameters. For a 3" line, that is only a nominal 15". For a 6' vertical vessel with a rupture disk discharge being piped to a drain hub on the floor, the 15" maximum length is exceeded without even thinking.
  

• Using the Resistance to Flow Method is valid for all cases. All sizing calculations can be standardized.{parse block="google_articles"}

• The Kr used in the Resistance to Flow Method is obtained by actual flow data for a given model of rupture disk and holder. Its use will provide a much more accurate calculation. The 0.62 coefficient of discharge used in the Coefficient of Discharge Method is very general and independent of rupture disk manufacturer model and type, holder, disk bursting characteristics and flow restrictions of the total relief system.

• Two-phase flow can be a major concern when using this method. The coefficient of discharge was established mainly for true vapors. Its application to liquids is questionable and its application to two-phase flow is totally fictitious. Granted, for the Resistance to Flow Method the Kr is not particularly applicable to two-phase systems either but one can easily compensate for this in the system calculations. Also, the rupture disk is only a part of an entire piping system and its overall contribution to the system frictional losses may not be greatly significant. Therefore, errors in Kr may not be very significant. In the Coefficient of Discharge Method, it is the only device considered. If the coefficient of discharge is grossly in error, the MNFA calculated will also be grossly in error.

• The same argument can be made for highly viscous liquid systems such as polymers.

In Summary

• Obtain the required relieving rate using good sound "what can go wrong" scenario analysis.

• Use the Resistance to Flow Method to calculate the size of the rupture disk (use the Coefficient of Discharge Method if you really must and you fall within the "8 & 5" rule).
  
• For the Resistance to Flow Method, try to choose the manufacturer and model of rupture disk you intend to purchase ahead of time or at least have a list of acceptable manufacturers and a good idea of the model you intend to use from each.
  
• For the Resistance to Flow Method use the ASME Kr value of 2.4 if you have no idea who the manufacturer(s) will be at the time of sizing.

References

1. API (www.api.org) Recommended Practice 520, "Sizing, Selection, and Installation of Pressure-Relieving Device in Refineries, Part 1-Sizing and Selection", 7^th Edition (January 2000)
2. API (www.api.org) Recommended Practice 521, "Guide for Pressure-Relieving and Depressuring Systems", 4^th Edition (March 1997)
3. ASME (www.asme.org) "Boiler and Pressure Vessel Code, Section VIII, Division 1" (1998)
4. Continental Disc Corporation (www.contdisc.com), Certiflow^TM Catalogue 1-1112
5. Fike (www.fike.com), Technical Bulletin TB8104, December 1999
6. Another good rupture disk manufacturer to investigate would be Oseco (www.oseco.com).
7. A good reference source for calculating flow through the system for liquids and gas/vapors is CRANE Technical Paper 410, "Flow of Fluids Through Valves, Fittings, and Pipe"
8. A great source and one that I feel should be the bible on two-phase flow is: Leung, J.C. "Easily Size Relief Device and Piping for Two-Phase Flow", Chemical Engineering Progress, December, 1996

Before I begin, let me point out that most of what is included in this series of articles can be found in API RP520¹ and API RP521², and ASME Section VIII, Division 1³. Much of what is found in these documents can also be found in vendor literature.

Problem{parse block="google_articles"}

1. What is the maximum allowable specified burst pressure?
2. What should the expected stamped (rated) burst pressure of the rupture disk be?
3. At what pressure(s) can we expect the delivered rupture disk to actually burst at?
4. What is the maximum allowable operating pressure in the vessel?

All these questions must be considered in order to properly set the burst pressure of a rupture disk.

What is the Maximum Allowable Specified Burst Pressure?

Burst Pressure

What do we mean by burst pressure? This is the pressure at which the rupture disk will open or burst. It is analogous to the set pressure of a relief valve and is specified by the process engineer.

Design Pressure

To find the maximum allowable specified burst pressure, the process engineer first needs to define a vessel design pressure. The design pressure is an arbitrary value above the vessel maximum operating pressure. One guideline used by many process design engineers is to increase the maximum operating pressure by 25 psig or 10% whichever is greater. For example, if the maximum operating pressure is 70 psig, then 25 psig should be added to arrive at the design pressure since 10% is only 7 psig. The design pressure would then be set at a nice round 100 psig. Other criteria to determine design pressure may be used but I recommend that the margins never be less than what I described above (the reason will become apparent later).

Maximum Allowable Working Pressure (MAWP)

The next step is to determine the Maximum Allowable Working Pressure (MAWP) of the vessel. A vessel specification stating design pressure, the coincident design temperature and other parameters is sent to the manufacturer. The manufacturer performs a series of calculations utilizing these parameters, amongst others, to determine material thickness for use in vessel fabrication. A standard material thickness (greater than or equal to what was calculated) is chosen. With the actual material thickness known, the true MAWP is calculated. The vessel design documents are then stamped (certified) at this pressure in accordance with code. However, for one reason or another, the MAWP calculation is not always done and the vendor will just stamp the vessel at the specified design pressure.

The Maximum Allowable Specified Burst Pressure

So, what is the maximum allowable specified burst pressure? Theoretically it is the MAWP. However, rupture disks are typically specified during basic engineering, which is performed way before the vessel is mechanically designed. This, combined with the fact that the true MAWP may never really be known (as mentioned above), the maximum allowable specified burst pressure will more typically be the vessel's design pressure.

Note if the rupture disk is to be used in conjunction with another relief device to fulfill the total required relieving capacity, the maximum allowable specified burst pressure could be 5% or even 10% greater than the design pressure (or MAWP). See ASME Section VIII, Division 1 paragraphs UG-125 and UG-134.

Also note that the specified burst pressure can be lower than the maximum allowable. Indeed, this is often the case if the rupture disk is used to protect reactor vessels against over pressure due to run-away reactions.

Stamped Burst Pressure

What should the expected stamped (rated) burst pressure of the rupture disk be?

What do we mean by "stamped or rated" burst pressure? Per code, the rupture disk vendor must provide a tag containing, amongst other things, the rated or what is typically called the stamped burst pressure. This is a guaranteed value so the user knows (within an allowable tolerance; more on this later) the exact bursting pressure of the rupture disk. Also this stamped burst pressure must never exceed the design pressure (or MAWP); except for the special case mentioned above.

So, the rupture disk vendor stamps the disk with the burst pressure specified by the process engineer? Not necessarily!

Manufacturing Range (MR)

Figure 1: Graphical Representation of the Manufacturing Range

A rupture disk is made out of a sheet of material, e.g. stainless steel, high alloys, ceramics, etc. Like all things in this world, this sheet of material is not perfect. To quantify the inaccuracies in sheet material thickness, the vendor uses what is called the Manufacturing Range (MR).

Figure 2A: Specified and Stamped Burst Pressure

 

The MR is expressed as ±% of the specified burst pressure. It determines the highest pressure above the specified burst pressure or the lowest pressure below the specified burst pressure that the disk can be stamped at. This is shown graphically in Figure 1.

Figure 1 shows the two extremes, a MR of ± 0% and a MR of ± some value%. Note that other combinations may be used such as + 0% and - some value% or - 0% and + some value%.

Let's look at an example. If the specified burst pressure is 100 psig with a MR of ± 0%, the stamped or rated burst pressure will be 100 psig (see Figure 2A). However, if the MR is +5% and - 10%, the disk can be delivered with a stamped burst pressure of 105 psig, 90 psig or anywhere in between (see Figure 2B). That's right, if the MR is anything but ± 0%, the user won't know the stamped burst pressure until the rupture disk is ready for shipment!

Figure 2B: Differences in Specified and Stamped Burst Pressures

Do you see anything wrong with this rupture disk as specified?

Remember, the stamped or rated burst pressure must never exceed the vessel's design pressure or MAWP (assumes a single device, no special cases). Since the specified burst pressure is the design pressure, this particular rupture disk is not acceptable because the delivered rupture disk may have a stamped burst pressure of 105 psig or 5 psig greater than design!

How can we avoid this problem? There are a number of ways.

The process engineer specifies the Manufacturing Range, not the manufacturer. You can ask for any range within the capability of fabrication including ± 0%. Considering the potential problems, why specify anything other than ± 0%? Cost. A MR of +5% and -10% can save as much as 40% off the cost of a similar rupture disk with a MR of ± 0%. Even if you demand +0% (which you should), you can still realize some cost savings if a stamped burst pressure lower than specified is acceptable (not always a good idea as will be discussed later). Note that code only affects the upper stamped limit, not the lower.

Another way to avoid the potential violation of code and still get a cheaper rupture disk is to specify a burst pressure that will be lower than the vessel design pressure. Thus, when the MR is added the stamped burst pressure will not exceed the design pressure. The maximum allowable specified burst pressure could be determined in the following manner:

P_{spec_max} = (DP) - (+MR/100) x (P_{spec_max})

Where DP = Design pressure

So:

P_{spec_max} = (DP) / [1+(+MR)/100]

Since DP = 100 psig and the upper value of MR = +5%,

P_{spec_max} = 100 /[1+(+5/100)] = 100/(1+0.05) = 100/1.05 = 95.2 psig

This rupture disk would be specified with a burst pressure no higher than 95.2 psig while the stamped burst pressure may be as high as 100 psig.

Note that the standard Manufacturing Range for most manufacturers is ± 0% and this is reflected in the base price you will be quoted.

Where Will the Disk Burst?

At what pressure(s) can we expect the delivered rupture disk to burst at?

Trick question? The answer should be the stamped burst pressure. But again the world isn't perfect.

Burst Tolerance

 

Figure 3A: Burst Tolerance

The Manufacturing Range is applied to the specified burst pressure but there is yet another unknown due to imperfections in the material used to fabricate the rupture disk. This is accounted for in the burst tolerance. Burst tolerance is applied to the stamped burst pressure and is set by code. For stamped burst pressures of 40 psig and lower, the burst tolerance is ± 2 psi. For stamped burst pressures above 40 psig, the burst tolerance is ± 5%.

Let's look at the examples again but apply the burst tolerance. For this discussion, I'm changing the specified burst pressure for the case of a rupture disk with a Manufacturing Range of +5% and -10% to 95.2 psig (see Figure 3B) so the stamped burst pressure can't exceed code.

The important thing to notice is that in both Figures 3A and 3B, the upper limit of the stamped burst pressure is equal to the design pressure but the maximum bursting pressure is 105 psig, or 5 psig over design pressure. Unlike the stamped burst pressure, which by code cannot exceed the design pressure (or MAWP), the maximum expected burst pressure can if it is caused by the burst tolerance.

Figure 3B: Specified, Stamped, and Maximum Burst

Maximum Allowable Operating Pressure

What is the maximum allowable operating pressure in the vessel?

Up to now, the discussions focused on the upper limit of the stamped burst pressure because this is governed by code. But the lower limit is extremely important to consider as well because of the possible affect it has on the maximum allowable operating pressure in the vessel.

Operating Ratio (OR)

The operating ratio is defined as the ratio of the maximum operating pressure to the lowest stamped burst pressure. The OR is used to protect against premature bursting of the rupture disk. If the operating pressure is too close to the lowest stamped burst pressure, or the system pressure cycles (pressure rises and falls during operation) too close to the stamped burst pressure, the material will fatigue and can {parse block="google_articles"}eventually loose its structural integrity. This is a classic reason for premature bursting of a rupture disk.

The manufacturer publishes the Operating Ratio for every rupture disk model they sell. For example, the Continental Disc Corporation's ULTRX rupture disk has an operating ratio of 90%⁴. This means the system pressure can operate to within 90% of the lowest stamped burst pressure without the fear of premature bursting. However, it's always best to operate as far away from the lowest stamped burst pressure as you can to avoid material fatigue.

From Figure 3B above, the lower limit or minimum stamped burst pressure is 85.7 psig:

P_{stamped_min} = (P_spec) - ABS [(-MR/100)] x (P_spec)

Where ‘ABS' stands for Absolute Value.

So:

P_{stamped_min} = (P_spec) x {1- ABS [(-MR/100)]}

Since P_spec = 95.2 psig and the lower value of MR = -10%,

P_{stamped_min} = 95.2 x {1 - ABS [(-10/100)]} = 95.2 x {1-ABS [(-0.1)]} = 95.2 x (1-0.1) = 95.2 x 0.9 = 85.7 psig

Therefore based on an OR of 90%, the maximum allowable operating pressure should not be greater than:

P_op = P_{stamped_min} x OR = 85.7 x 0.9 = 77 psig.

Since our discussions have been based on a maximum operating pressure of 70 psig, this rupture disk is acceptable. But note that this 10% cushion exists only because of the design pressure margin used (25 psig). Had the margin been less, say only 10%, the rupture disk we would want to use would be unacceptable.

How to avoid this problem?

• Set the design pressure appropriately
• Choose a rupture disk with a MR of ± 0%
• Choose a rupture disk with a OR of 90% (they don't really go much higher)

There is one more point to consider. Although I have never seen any mention of checking the maximum allowable operating pressure against the minimum expected burst pressure (arrived at by taking into account the burst tolerance), I think it only makes good engineering sense to do so. After all, if the disk can burst at this lower pressure, one certainly does not want to operate too close to it!

Getting back to our question, what is the maximum allowable operating pressure in the vessel? In this case, it is 77 psig.

Summary

• What is the maximum allowable specified burst pressure?

- Design Pressure or MAWP if the rupture disk is the only relief device

OR

- For special cases, 105% (or even 110%) of design pressure or MAWP if the rupture disk is a secondary device

 

• What should be the expected stamped (rated) burst pressure of the rupture disk?

- As specified by the process engineer for a Manufacturing Range of ± 0%

OR

- As specified by the process engineer but could be adjusted per the Manufacturing Range if other than ± 0%

• At what pressure(s) can we expect the delivered rupture disk to actually burst at?

- ± 5% of stamped burst pressure for stamped pressures greater than 40 psig

OR

- ± 2 psi for stamped pressures 40 psig and lower

• What is the maximum allowable operating pressure in the vessel?

- Specified by the process engineer based on operating need but must be checked against the Operating Ratio of the rupture disk
- I strongly suggest you also check against the minimum expected burst pressure as well.

• Manufacturing Range is applied to the specified burst pressure
• Burst Tolerance is applied to the stamped burst pressure
• Set the design pressure appropriately
• Choose a rupture disk with a MR of ± 0%
• Choose a rupture disk with a OR of 90%

References

1. API (www.api.org) Recommended Practice 520, "Sizing, Selection, and Installation of Pressure-Relieving Device in Refineries, Part 1-Sizing and Selection", 7^th Edition (January 2000)
2. API (www.api.org) Recommended Practice 521, "Guide for Pressure-Relieving and Depressuring Systems", 4^th Edition (March 1997)
3. ASME (www.asme.org) "Boiler and Pressure Vessel Code, Section VIII, Division 1" (1998)
4. Continental Disc Corporation, ULTRX ^® Catalogue 3-2210-3

 

Before I begin, let me point out that most of what is included in this series of articles can be found in API RP520¹ and API RP521², and ASME Section VIII, Division 1³.  Much of what is found in these documents can also be found in vendor literature.

Temperature and Backpressure Considerations

In Part 3, I discussed how to set the burst pressure of the rupture disk. However, the discussion is not complete without considering the affects of temperature and backpressure on the bursting pressure.

Temperature

The rupture disk manufacturer uses both the specified burst pressure and the specified temperature when designing and stamping the disk. (In this instance, I use the term design to mean arriving at the correct burst pressure, not mechanical integrity). However, it is more than likely that the temperature of the rupture disk will not be at the specified temperature when it is called into service. Why is this so?{parse block="google_articles"}

The temperature most commonly specified is that of the relieving fluid coincident with the burst pressure, i.e. relieving conditions. Sounds logical, but remember that the disk is continuously exposed to the process stream for hours, days, weeks or even months before it may ever be needed. Or, the disk may be exposed to ambient conditions. Therefore, expect the disk temperature to be approximately equal to its environment during normal operation of the system. When a process upset occurs, system pressure rises until it reaches relief (burst). The temperature of the relieving fluid also rises per thermodynamics. However, the time interval between normal system operation and relief is usually so small that the rupture disk's temperature hardly has time to come to equilibrium with the higher process fluid temperature. Therefore the disk can actually be colder than it's specified temperature. The affects?

In general, burst pressure varies inversely with temperature. For some rupture disks, the burst pressure can be as much as 15 psi greater than stamped if the actual temperature is 100^oF lower than specified, e.g. a disk specified with a burst pressure of 350 psig at a temperature of 400^oF will actually burst at 365 psig if its temperature is only 300^oF⁴. This doesn't sound like a big difference but if 350 psig were the design pressure (or MAWP) of the vessel, then a burst pressure of 365 psig would be in violation of code (LAW). The opposite is also true. A disk at a temperature hotter than specified when called into service will burst at a pressure lower than stamped. Although this is considered to be the more conservative approach because code can't be violated and there is no risk of catastrophic failure of the vessel, specifying too low of a temperature can lead to the not so desirable action of premature bursting.

The bottom line is that the specified burst temperature must be carefully considered. Specify the lowest temperature at the time the disk is expected to burst. Consider that this might be the normal process operating temperature or even ambient rather than the calculated relieving temperature.

Note that different materials and different types of rupture disks have different sensitivities to temperature. This is an excellent topic of discussion for your rupture disk manufacturer!

Backpressure

A rupture disk is actually a differential pressure device where the specified burst pressure is equal to the difference between the desired upstream pressure (vessel) at the time of rupture disk burst and the downstream pressure (backpressure):{parse block="google_articles"}

P_burst = P_vessel - P_backpressure

Or alternately the desired upstream pressure (vessel) at the time of rupture disk burst is equal to the sum of the specified burst pressure and the downstream pressure (backpressure):

P_vessel = P_burst + P_backpressure

Either way, it is apparent that the vessel pressure at the time the rupture disk bursts (commonly called the relief pressure) is directly dependent on backpressure.

When discussing relief systems, three types of backpressure are considered, these being constant, built-up and superimposed.

Figure 1A: Single Vessel, Single Rupture Disk Protection, Expected Constant Back pressure = 0 psig

Figure 1A shows a system comprised of a single vessel protected by a single rupture disk with a specified burst pressure of 100 psig. The relief pipe discharges a few inches below the liquid surface in a knockout drum, which is held at a constant 0-psig pressure. Therefore, the rupture disk sees a constant (fixed) backpressure of 0 psig. If the vessel were to go into relief, this disk will burst at 100 psig and the vessel relief pressure will be 100 psig (100 + 0 = 100).

Figure 1B: Single Vessel, Single Rupture Disk Protection, Actual Constant Back pressure >  Expected

Figure 1B is the same system however for some reason the pressure in the knockout drum is to be maintained at 5 psig instead of 0 psig. The constant (fixed) backpressure against the rupture disk is now 5 psig. If the vessel were to go into relief, the rupture disk would still burst at 100 psig but the vessel relief pressure would now be 105 psig (100 + 5 = 105) rather than the 100 psig expected. This situation could result in a violation of code³.

Figure 1C: Single Vessel, Single Rupture Disk Protection, Actual Constant Back pressure <  Expected

Figure 1C is again the system however for some reason the pressure in the knockout drum is to be maintained at -5 psig instead of 0 psig. The constant (fixed) backpressure against the rupture disk is now -5 psig. If the vessel were to go into relief, the rupture disk would still burst at 100 psig but the vessel relief pressure would now be only 95 psig (100 + (- 5) = 95) rather than the 100 psig expected. There is no particular safety concern here because the vessel can't over pressure. However, the Operating Ratio is affected, which can result in a very premature bursting of the rupture disk.

For the vessel relief pressure to be specified correctly, the rupture disk vendor must be told the constant back pressure so that the rupture disk can be designed accordingly. And, if you truly want the vessel relief pressure to be at a specific value then the "constant" backpressure given to the vendor must be maintained at all times.

The key point is that during design, be aware of the constant backpressure and ensure that the vessel relief pressure will not violate code or affect normal operation.

 

Figure 2A: Two Vessel System - Common Discharge Built-up and Superimposed Back Pressures

Now let's look at the system shown in Figure 2A. A second vessel with a single rupture disk also specified to burst at 100 psig is added in close proximity to the first vessel. The relief piping from the two vessels is tied into a common header before discharging into a knockout drum in the same manner as before, the tie-in occurs near the vessels. At the exact moment Vessel No. 2 goes into relief and its rupture disk bursts, Vessel No. 2's relief pressure is 100 psig due to the constant 0-psig backpressure as described above. After the disk bursts, flow is established causing pressure to build up in the piping system (built-up backpressure). The amount of built-up backpressure is dependent on the system pressure drop and possibly even the phenomenon of choked flow.  For the purpose of this discussion, assume total built-up backpressure is 10 psig after rupture disk No. 2 bursts and the pressure in Vessel No. 2 is about 110 psig. Because of the proximity of the two discharge pipes and vessels, the pressure near vessel No. 1 will also be at about 110 psig. This pressure, which is exerted or imposed onto rupture disk No. 1, is called the superimposed backpressure with respect to rupture disk No. 1. If vessel No. 1 were to go into relief shortly afterwards, then for rupture disk No. 1 to burst, the pressure in vessel No. 1 would have to build to about 210 psig (100 + 110)!   This is clearly unacceptable!!

One solution to this potentially catastrophic condition is to separate the two relief lines so that one cannot directly affect the other (see Figure 2B below). Of course the answer may very well be that this is not an application for rupture disks but for relief valves! The key point is, avoid combining multiple rupture disk piping into a common relief header.

Figure 2B: Two Vessel System - Common Discharge Built-up and Superimposed Back Pressures

Note that built-up backpressure is variable and depends on the relieving rate, which is a function of the relieving scenario. Also, built-up backpressure has no affect on the vessel's relief pressure for systems such as those shown in Figure 1 above. Built-up backpressure is the result of fluid flow only and there is no fluid flow before the rupture disk bursts.

Therefore, along with the Manufacturing Range (MR), Operating Ratio (OR) and Burst Tolerance (BT) that were discussed in Part 3, the process design engineer must also strongly consider the backpressure (especially superimposed backpressure) when specifying the rupture disk.

 

In Summary

• Generally, burst pressure varies inversely with temperature so the specified burst temperature must be carefully considered.
  - Specify the lowest temperature at the time the disk is expected to burst.
  - Different materials and different types of rupture disks have different sensitivities to temperature effects.{parse block="google_articles"}
  
• The rupture disk is a differential pressure device.
  - The specified burst pressure is a value equal to the vessel relief pressure minus the backpressure.
                                                                   Or
  - The vessel relief pressure equals the specified burst pressure plus the backpressure.
  
• There are three types of backpressure to consider, these being constant, built-up and superimposed.
  - Constant backpressure is the pressure in the system that does not vary. It is generally a predictable component of the superimposed backpressure.
  - Built-up backpressure is the pressure created in the system as a result of fluid flow. It is a varying component of the superimposed backpressure.
  - Superimposed backpressure is the total pressure exerted (imposed) on the rupture disk by other sources. It is a variable that directly increases or decreases a vessel's relief pressure. It can also interfere with the expected operating ratio of the disk.
  
• Do not pipe multiple vessel relief systems into a common header; keep the piping separate. However, the individual piping may go to a common disposal system.
  
• Along with the Manufacturing Range (MR), Operating Ratio (OR) and Burst Tolerance (BT), the process design engineer must also consider backpressure when specifying the rupture disk.

References

1. API (www.api.org) Recommended Practice 520, "Sizing, Selection, and Installation of Pressure-Relieving Device in Refineries, Part 1-Sizing and Selection", 7^th Edition (January 2000)
2. API (www.api.org) Recommended Practice 521, "Guide for Pressure-Relieving and Depressuring Systems", 4^th Edition (March 1997)
3. ASME (www.asme.org) "Boiler and Pressure Vessel Code, Section VIII, Division 1" (1998)
4. Nazario, F. N., "Rupture Discs, A Primer", Chemical Engineering Magazine, June 20, 1988.

For the relief valve/rupture disk combination (Figure 1), rupture disk sizing is totally dependent on relief valve sizing, regardless whether the rupture disk is installed upstream or downstream of the relief valve. Consequently, the discussion at this point must turn to a brief overview of relief valves.

Relief Valve Sizing Overview

Basically, the relief valve is treated as an ideal nozzle, i.e. isentropic (constant entropy) flow. A correction factor, the coefficient of discharge, is incorporated into the sizing equations to take into account the fact that this is not an ideal nozzle. The sizing equations themselves can be found in one or more of the references presented at the end.

To size a relief valve, the process engineer first determines the required relieving flow and fluid properties based on an analysis of "what can go wrong" scenarios. The flow and properties are then inserted into the appropriate sizing equation to arrive at a calculated relief valve area. If this were a stand-alone relief valve, the process engineer would use this calculated relief valve area to choose an actual relief valve from

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 Figure 1: Relief Valve/Rupture Disk Combination

a vendor catalog. But since this is a discussion of the relief valve/rupture disk combination, adjustments should be made to the calculated relief valve area before the actual relief valve is chosen.

The Rupture Disk Affect

The presence of a rupture disk acts to de-rate the relief valve capacity. This de-rating factor, called the Combination Capacity Factor (CCF), may or may not be implicitly included in the sizing formulas. Nevertheless, it is the responsibility of the process engineer to apply the factor correctly.

The Combination Capacity Factor (CCF) 

The Combination Capacity Factor (CCF) is a calculated value that is derived from data obtained during certified capacity testing of the stand-alone relief valve and the relief valve/rupture disk combination. The manufacturer first determines the capacity of the stand-alone relief valve. The rupture disk is then added, close-coupled, to the inlet of the relief valve and the capacity of the relief valve/rupture disk combination is determined. Finally, the CCF is calculated as the ratio of the relief valve/rupture disk combination capacity to the stand-alone relief valve capacity:

CCF = Flow _{Combination Capacity} / Flow _{Stand-Alone Relief Valve Capacity}

 

Below is a list of certified Combination Capacity Factors for the Continental Disc Corporation model ULTRX ^® rupture disks with the Crosby JOS/JBS Relief Valve ⁴.

For comparison, the following is a list of certified Combination Capacity Factors for the Fike model MRK rupture disk with the Crosby JOS/JBS Relief Valve⁵.

Note that the CCF is a certified value and is only good for the design of the relief valve and the rupture disk that are used in the test. Since it is in the best interest of the rupture disk manufacturer to certify as many of their rupture disk designs with as many different types of relief valve designs as possible, it is typical for the rupture disk manufacturer to perform this testing and reporting of the CCF. The certified CCF will always be less than or equal to 1.0.

If the manufacturer and/or model of the rupture disk and relief valve are unknown at the time of sizing, or there is no published value for a relief valve/rupture disk combination, ASME³ requires that the CCF is not to exceed 0.9.

Applying the CCF

API Recommended Practices 520¹ shows the CCF as being applied to the denominator of the relief valve sizing equation. For example, a typical sizing equation for gas relief might look something like this:

Eq. (1)

Where:
W = required relieving rate, mass flow
T  = relieving temperature, absolute
Z  = compressibility factor
M = molecular weight
C = gas constant = a function of (C_p / C_v)
C_p = specific heat at constant pressure (consistent units)
C_v = specific heat at constant volume (consistent units)
K_d = coefficient of discharge, dimensionless
K_b = backpressure correction factor, dimensionless
P1 = relief pressure (absolute)

Note that this is the same as dividing the calculated, stand-alone relief valve area by the CCF to arrive at a required relief valve area for the combination unit:

Eq. (2)

And:

The process engineer will use this required relief valve area as the basis for choosing a relief valve from the vendor catalog.

One important thing to note is that the preceding methodology is not a requirement of code (ASME). ASME only requires that the stand-alone relief valve's certified flow capacity be de-rated by the CCF:

There is no mention of using the CCF to arrive at a relief valve area. Indeed, prior to the most recent edition of API RP520¹, the sizing equations themselves did not explicitly include a correction factor for the relief valve/rupture disk combination.

Note also that de-rating the certified flow capacity is only required if the rupture disk is installed upstream of the relief valve, it is not required if installed downstream of the relief valve.

Certified (Rated) Capacity

As stated above, each stand-alone relief valve will have associated with it a certified flow capacity, which is a function of both the relief valve area and the set pressure. This flow is determined by certified capacity testing procedures and is to be considered the guaranteed flow rate that can be achieved through the particular valve. With very few exceptions, this flow is used in determining both the relief valve inlet and outlet (tail pipe) line sizes. The certified flow capacity is officially stamped on the relief valve documentation. For relief valve/rupture disk combinations, the de-rated certified flow will also be stamped on the documentation.

Although the relief valve is chosen based on area, the process engineer must still ensure that the certified flow capacity is greater than or equal to the required relieving flow:

Certified Flow Capacity ³ Required Relieving Flow

If it is not, the chosen relief valve is too small. For the relief valve/rupture disk combination, the required relieving flow would be compared to the de-rated or combination certified flow capacity:

Combination Certified Flow Capacity ³ Required Relieving Flow

Relief Valve Sizing

Inlet Line

The relief valve inlet line is defined as the piping between the inlet to the system (e.g. the inlet to a vessel nozzle) and the relief valve inlet flange. Sizing this inlet line is a trial-and-error procedure. First, the process engineer chooses a line size using guidelines set by code; code requires that the flow area of the line and all associated piping components be at least equal to the relief valve inlet flow area. Then, using {parse block="google_articles"}accepted fluid flow equations (e.g. Darcy for single phase liquid or gas/vapor and DIERS for two phase) and the certified flow capacity of the stand-alone relief valve the non-recoverable frictional losses in the line are determined. The sum of all non-recoverable losses should be less than 3% of the relief valve set pressure, this criteria is commonly referred to as the 3% Rule. In general, if the 3% Rule is exceeded then the chosen line size is too small.

Outlet Line (Tail Pipe) Sizing Overview 

The sizing of the tail pipe is done in a similar manner to that outlined above for the inlet line. The process engineer first chooses a pipe size. Then, using accepted fluid flow equations (e.g. Darcy for liquids, Isothermal or Adiabatic for gas/vapor and DIERS for two-phase) and the same certified flow capacity as used for the inlet line, a built-up (variable) backpressure is calculated. The built-up backpressure is converted to a percentage of the relief valve set pressure and is then compared to some maximum value that is set by the particular relief valve manufacturer. For example, tail pipes on conventional style relief valves would be sized such that the built-up (variable) backpressure does not exceed 10% of the relief valve set pressure. For balanced bellows style relief valves, tail pipes would be sized such that the built-up (variable) backpressure does not exceed 30% to 55% of the relief valve set pressure, depending on manufacturer. If the calculated percentage is less than or equal to these maximums, the line size is acceptable. If the calculated percentage is greater, the line size may or may not be acceptable. This is because the only requirement of code is that the built-up backpressure does not affect the relief valve's ability to relieve the required amount of flow necessary to protect the system. Built-up backpressures greater than the stated maximums require a de-rating of the relief valve based on curves developed by the manufactures. As long as the de-rated valve can still relieve the required relieving flow, the line size chosen is acceptable. If not, then the line is too small.

Now that we've sized the relief valve in the relief valve/rupture disk combination, what about sizing the rupture disk? Actually, we already did!

The Rupure Disk

Sizing

You will recall from Part 2 of this series that sizing the rupture disk is a two-part procedure. First, determine how much flow the rupture disk needs to pass. Then determine how big it needs to be.

Both criteria have been met with the relief valve sizing. How much flow? The rupture disk must be able to pass the certified flow capacity of the relief valve. How big? The rupture disk must be big enough so that its contribution to the frictional losses does not pose a significant impact on the ability of the relief valve to protect the system. For a rupture disk installed in the inlet line, the rupture disk's net flow area must be at least equal to the relief valve inlet flow area; it may be larger. Also, its contribution to the non-recoverable frictional losses should be minimal so as to ensure that the piping system meets the 3% Rule. Indeed, you may even find that the rupture disk must be one-size larger than the inlet to the relief valve in order to satisfy the 3% Rule. For example (Figure 2), a 2" x 3" relief valve (2" being the inlet flange size and 3" being the outlet flange size) may require a 3" rupture disk!

Figure 2: 2" x 3" Relief Valve

Figure 3: Rupture Disk is Larger than Outlet Flange

For a rupture disk installed in the tail pipe, the rupture disk size should be large enough so that it contributes minimally to the built-up backpressure. And again, the rupture disk may very well have to be a size larger than the relief valve outlet flange to accomplish this (Figure 3).

For both the inlet line and tail pipe calculations, the rupture disk's certified K_r is used in the friction loss calculations.

What the Code Says About...

Bursting

The stamped (certified) burst pressure of the rupture disk must be between 90% and 100% of the relief valve set pressure. Also, the bursting of the rupture disk and the opening of the relief valve must be essentially coincident with each other.

Backpressure

When specifying a rupture disk that will be used upstream of a relief valve, it is expected that the superimposed backpressure will be constant and essentially zero (after all, there should be nothing between the rupture disk and the relief valve but some trapped air). However, over time the rupture disk may leak for a variety of reasons. This leakage will cause a build-up of pressure between the rupture disk and the relief valve. As we saw in Part 4, unexpected backpressure on the rupture disk will change the relieving pressure of the vessel or system. To guard against this, code requires the use of a "tell-tale". The "tell-tale" must consist of, as a minimum, a pressure gage and a vent line inserted between the rupture disk and the relief valve. Typically, a valve is put into the vent line for a more controlled design (Figure 4). In installations where the rupture disk holder is close-coupled with the relief valve, this system is inserted into a chamber within the holder. Note that a better tell-tale design would include a pressure transmitter with an alarm as well as the pressure gage.

 

Figure 4: Valve in Vent Line

For rupture disks installed after the relief valve, the disk's bursting pressure must not be affected by any backpressure affects nor can there be allowed a pressure build-up between the relief valve and rupture disk that may affect the operation of either device. A "tell-tale" should be used to protect against pressure build-up between the devices due to leaks through the relief valve. The only way to protect against backpressure affects is to make sure the superimposed backpressure is well defined and constant (see Part 4 of this series).

Obstructions

A bursting rupture disk must not cause obstruction of the relief valve or the relief piping. Therefore, the non-fragmenting rupture disk is used in this service. This disk will break cleanly, with no material being broken off.

Final Thoughts

• Above I discuss the fact that the rupture disk needs to be able to pass the certified flow capacity of the relief valve! But which flow capacity, the stand-alone relief valve or the relief valve/rupture disk combination? Unfortunately, the way ASME³ reads, there is plenty room for interpretation. For example, paragraph UG-127 (a) (3) (b') (5) basically says the rupture disk must be able to pass the certified capacity of the relief valve/rupture combination. However, for the rupture disk installed in the tail pipe, paragraph UG-127 (a) (3) (c') (4) says the rupture disk must be able to pass, "...the rated capacity of {parse block="google_articles"}the attached pressure relief valve without exceeding the allowable overpressure." Now, for individual cases where the rupture disk is installed only upstream of the relief valve or only downstream of the relief valve, I can buy into this as not being contradictory, i.e. use rated capacity of the relief valve/rupture combination for the inlet line or use the rated capacity of the stand-alone relief valve for the tail pipe. But what about the case where the rupture disk is installed both upstream and downstream of the relief valve?
  
  The flow used to evaluate the inlet line is to be the same flow used to evaluate the tail pipe. And, the 3% Rule clearly wants you to use the certified capacity of the stand-alone relief valve with the rupture disk being treated as just another piping component.
  
  So which do I suggest we Process Design Engineers use? The certified flow capacity of the stand-alone relief valve in all instances; it will be a little more conservative.

• The code requirements discussed above help to emphasize the importance of the material presented in Parts 3 and 4 of this series, i.e. the maximum allowable specified burst pressure, the Manufacturing Range, the Burst Tolerance, the Operating Ratio, and superimposed, built-up and variable backpressures; especially as they relate to the relief valve/rupture disk combination

Summary

• Rupture disks may be installed upstream and/or downstream of a relief valve.

• The rupture disk acts to de-rate the relief valve capacity. This de-rating factor is called the Combination Capacity Factor. Standards call for the use of this factor in determining relief valve area and in de-rating the stand-alone relief valve's certified capacity. Code only requires the use of this factor in de-rating the stand-alone relief valve's certified capacity.

• The size of the rupture disk in this application is totally dependent on relief valve sizing.

• The rupture disk must be able to pass the certified flow of the relief valve.

• The size of a rupture disk installed at the inlet of the relief valve should have minimal affect on the 3% Rule and must have a flow area of at least equal to the inlet flow area of the relief valve.

• The size of a rupture disk installed at the outlet of the relief valve should provide minimal contribution to the built-up backpressure.

• Code governs how a rupture disk is applied to a relief valve installation and the general type of rupture disk to use (non-fragmenting).

• Code addresses rupture disk bursting requirements.

• Code addresses backpressure affects and what must be done to avoid it.

• When specifying a rupture disk, especially in combination service with a relief valve, the maximum allowable specified burst pressure, the Manufacturing Range, the Burst Tolerance and the Operating Ratio all must be considered very carefully.

References

1. API (www.api.org) Recommended Practice 520, "Sizing, Selection, and Installation of Pressure-Relieving Device in Refineries, Part 1-Sizing and Selection", 7^th Edition (January 2000)
2. API (www.api.org) Recommended Practice 521, "Guide for Pressure-Relieving and Depressuring Systems", 4^th Edition (March 1997)
3. ASME (www.asme.org) "Boiler and Pressure Vessel Code, Section VIII, Division 1" (1998)
4. Continental Disc Corporation (www.contdisc.com), ASME Combination Capacity Factors, Catalogue 1-1111
5. Fike (www.fike.com), Technical Bulletin TB8103, July 1999

 

I will also touch upon the type of rupture disks you can purchase. Before I begin, let me point out that most of what is included in this series of articles can be found in API RP520¹ and API RP521², and ASME Section VIII, Division 1³.  Much of what is found in these documents can also be found in vendor literature.{parse block="google_articles"}

We've answered the two questions required to size a rupture disk, how much flow and how big. Now it's time to specify the rupture disk so that it can be purchased for our process. Although API RP520¹ provides a specification sheet that can be adapted by any company as a standard, there are fifty-three separate items asked for in this specification sheet. Much of what is on this specification sheet is not required by the manufacturer to be able to provide you with the correct disk. Let's look at the basic minimum information you, the Process Design Engineer must provide.

Must Haves

Project Identifier/Company Information/Device identifier/Number of Devices

The vendor will want to know who you are. It is also necessary to "name" the relief device for proper documentation. A unique instrument Tag number should suffice for each device ordered.

Code/Standard Requirements

Various codes and standards dictate how the rupture disk is to be marked and stamped.

Maximum Operating Conditions

Temperature

The maximum operating temperature is used to determine materials compatibility.

Pressure

The maximum operating pressure will be used with the stamped burst pressure to determine the Operating Ratio. The Operating Ratio will help determine the type of disk to purchase.

Rupture Disk Burst Conditions

Temperature

This must be coincident with the bursting pressure and will also be stamped on the disk. You will recall from Part 4 that this parameter is extremely important in making sure the disk will burst at the pressure you need it to burst, not less or greater. Also remember that it is not necessarily the same as the maximum operating temperature of the system.

Pressure

This is the pressure that meets system protection requirements, taking into account the Manufacturing Range. The vendor will stamp this value on the disk. It is also used with the Maximum Operating Pressure to determine the Operating Ratio.

Process Media (Liquid/Gas/2-Phase)

Some rupture disk models are designed according to the media in which they are used. Process media is also used to determine materials compatibility.

Backpressure/Vacuum

The manufacturer uses the backpressure to help determine disk type and how it is to be supported in the system. Vacuum service will either require the use of a special support for disk installation or even dictate the type of disk to use. Note that exposure to vacuum conditions must be considered both upstream and downstream of the disk.

Service Conditions (Status/Cyclic/Pulsating)

This typically refers to the upstream conditions. Cyclic service is considered to be large changes in pressure over a relatively long period of time. Pulsating service is considered to be small changes in pressure but occurring frequently or even rapidly. Both of these can have a major affect on the Operating Ratio. The manufacturer uses the service conditions to help determine disk type and how the disk is to be supported in the system.

Rupture Disk and Holder Material Requirements

Many installations require the rupture disk to be mounted inside a holder. The holder is then bolted onto a vessel nozzle or between pipe flanges. Make very certain the materials of construction of both the disk and its holder is totally compatible with the system media and operating conditions.

Disk Size

This is the nominal size you determined when answering the question, how big?

Flange Connection Details

These tell the manufacturer how big the holder needs to be (connection size), the pressure rating of the system it will be installed in (class) and the type of connection, e.g. raised or flat faced flanges, sanitary connections, etc.

The pressure rating or class can be a most confusing concept. This refers to the flanges in the piping system. More common flange ratings are 150 and 300 pounds (pressure pounds, not weight) but they can go very much higher. A major difference in these classes is the thickness of material, number of boltholes and the bolthole pattern you would get in the flange.

Required Accessories for Rupture Disk

Options can be added to the basic design. For instance to enhance corrosion protection, coatings or linings can be applied.  Some types of rupture disks can withstand upstream vacuum conditions without doing anything special to them others may need special supports.

Required Accessories for Holder

Options can be added to the rupture disk holder as well. For instance to enhance corrosion protection, coatings or linings can be applied. Tell-tales may be specified under this header or can be specified under the heading of "Special Considerations".

Other Special Considerations

You can specify just about anything under this heading including the need for a tell-tale. You may want to give more specific detail of a particular design item. You can ask for burst detection and alarms, etc., etc. and etc. The best reference source would be your manufacturer and/or their catalog.

Again, the above should be considered just the minimum amount of information the manufacturer needs to provide the proper rupture disk. Of course your particular manufacturer, or even your company standards, may require much more.

Should you stop here, perhaps not? Below is some information that I consider to be "should haves".

Should Haves

MAWP (or Design Pressure) of the Vessel or System

A vendor does not necessarily require this information (they were already told what to stamp the disk for). However a good vendor will actually be your second set of eyes and make sure that this, along with the other information given, is consistent with Code requirements.

Manufacturing Range

One would think that this should fall under the "must haves" but not really. When the burst pressure was specified in the "must haves", the manufacturing range had to be taken into account. All the vendor needs to know is what to stamp the rupture disk at and will therefore design the disk with the appropriate manufacturing range to accommodate. However, it never hurts to spell it out so there are no misunderstandings.

In Combination with a PSV

With this information, the rupture disk vendor will be able to recommend the proper type of rupture disk to use for this service. They will also be able to recommend proper installation techniques. And again, the vendor is your second set of eyes and may be able to tell whether your specification data is consistent.

Calculate and Report the Operating Ratio

I could never quite figure out why the vendor cannot just do the simple math but I've seen this as requested information on a number of vendor's specification sheets.

What about all the rest of the information usually included in many specification sheets, e.g. required relieving flow, molecular weight, specific heat ratio, specific gravity, compressibility factor, viscosity, etc.? These are definitely important, but really only to the Process Design Engineer. You need this information to answer the two questions, how much flow and how big? The vendor doesn't need these but we all seem to include them on our specification sheets nevertheless!

The best suggestion I can make is to talk to the vendor first, find out exactly what they need and provide it. But of course, never violate your own company standards.

Types of Rupture Disks

The manufacturer can recommend the type of rupture disk that will best suit your application based on the information supplied. However, it doesn't hurt to have some knowledge {parse block="google_articles"}of the type of rupture disks that can be purchased. There are a multitude of different types and the following only represents the most common types you will most likely come across.

Forward Acting Solid Metal

This rupture disk is domed shape and installed such that the media is on the concave side of the disk (Figure 1). It can be used in systems where the Operating Ratio is at about 70% or less. It has a random bursting pattern which means it can be fragmenting (loose material) and thus cannot be used in combination with relief valves. This type of rupture disk can be used in vacuum or larger backpressure services but will require special supports to prevent reverse flexing. Its number one advantage is that it is cheap.

Figure 1: Forward Acting Solid Metal Rupture Disk

 

Forward Acting Scored Metal

This rupture disk is similar to its solid metal cousin (Figure 1) except that the disk is scored (Figure 2). Unlike the ill-defined bursting pattern of the solid metal design, this rupture disk has scored lines that will force the disk to burst along a fixed pattern. This design is a little more expensive but increases the useful Operating Ratio to about 85 to 90%. It also eliminates fragmenting, which means it can be used in combination with a relief valve. Also, there are many designs that allow this type of disk to be installed in vacuum environments without requiring special supports; it will still need special supports in high backpressure service to prevent reverse flexing.

 

Figure 2: Forward Acting Scored Rupture Disk

 

 Forward Acting Composite

This rupture disk can be flat or domed and is comprised of a top section preceded by a bottom seal (Figure 3). The burst pressure is a function of these two sections. It is not uncommon for the bottom section to be of a totally different material of construction from that of the top section, even non-metallic. The domed disk design will burst due to pressure applied to the concave side whereas the flat disk design may bedesigned to burst in either direction! 

Figure 3: Forward Acting Composite Rupture Disk

 

 

 

Slits and tabs in the top section control burst pressure and the bursting pattern. The flat construction can be used for the protection of low-pressure systems. Operating ratios are typically around 80% for the dome construction and 50% for the flat construction. This disk may require special supports to be used in vacuum or high backpressure conditions. Some designs are non-fragmenting, which means they can be used in relief valve combination.

Reverse Acting

This rupture disk is domed shape and installed such that the media is on the convex side of the disk (Figure 4). It is designed such that pressure pushes against the disk causing it to flex back into a forwarding acting disk and then burst. This rupture disk can be used in systems where the Operating Ratio is at about 90% or less. It can be, and very often is, manufactured to be non-fragmenting and thus is a good choice for use in combination with relief valves. This type of rupture disk can be used in vacuum or larger backpressure services without special supports.

Figure 4: Reverse Acting Scored Rupture Disk

 

Final Thoughts

Liquids

Liquids are treated the same way as gases/vapors in all aspects of determining those two questions, how much and how big. However, do not forget to take the hydraulic pressure into account. Pressure in the system will not be equal throughout. If the rupture disk is installed on a nozzle or in a pipe at the top of a liquid filled vessel, the pressure at the rupture disk will be less than all points below it. If the rupture disk is installed on a nozzle at the bottom of a liquid filled vessel, the pressure at the rupture disk will be greater than all points above it.{parse block="google_articles"}

What are the implications of this? If the rupture disk is located at the top of the vessel, the vessel pressure will be greater than the bursting pressure so specify the burst pressure to be less than the vessel's MAWP or design pressure. If the rupture disk is at the bottom of the vessel, the vessel pressure will be less than the bursting pressure. However, the rupture disk cannot be specified at a pressure higher than MAWP or design. Therefore, realize that the disk will burst even though the pressure at the top of the vessel will be less than design or MAWP.

Also note that normal variations in level will cause normal variations in the pressure, i.e. the rupture disk will experience pressure cycling or pulsing. Unlike gases/vapors where normal system pressure cycling or pulsing is usually minimal, it may be significant in liquid filled systems.

One More Option to Consider

Ask your manufacturer if they provide a "Fail Safe" design. This design will provide pressure relief at or below the certified burst pressure even if the disk is damaged or installed improperly. It will function in this capacity equally well in gas/vapor or liquid service. The major drawback is that it is only available in forward acting non-composite rupture disks.

Other Non-Close Relief Devices

Figure 5A: Rupture Pin Relief Device End of Pipe with Atmospheric Discharge

Figure 5B: Rupture Pin Relief Device Discharge to Header

There are other options to consider for non-closing relief devices other than rupture disks. Although details are beyond the scope of this article, there is one particular device I wish to bring to your attention and which is gaining in popularity, the Rupture Pin^{6, 7}. Although ASME will not allow what is called a Breaking Pin device to be used as a primary relief device, as of May 1990, it will allow the use of the Rupture Pin device. The two are similar but for the Breaking Pin device to work, the pin must completely break but for the Rupture Pin device to work, the pin only needs to bend or buckle. Another name for this device is the Buckling Pin.  Figures 5A and 5B show two types of rupture pin devices. Device "A" might be used directly on a vessel and will relieve to atmosphere. Device "B" might relieve into a piping header.

The rupture pin device usually consists of a piston or plunger on a seat, kept in position by a slender, usually cylindrical pin. At set point, axial forces caused by system pressure acting on the piston or plunger area causes the pin to buckle. The unrestrained pin length, the pin diameter and the modulus of elasticity of the pin material determine the buckling point of the pin.

There is virtually no device size limitation. They have been manufactured as small as 1/8" and as large as 48". There are virtually no pressure or vacuum limits either. They can be designed for a set pressure as low as 2" of water to as high as 35,000 psi and vacuums to as low as 1 psi. Unlike rupture disks, which are solely differential devices, the rupture pin can be designed to sense system pressure only, or differential pressure.

You are now ready to sit through one of those manufacturer's presentations and hopefully understand what he is talking about!

Summary

• API RP520 provides a specification sheet that can be adapted by any company as a standard
• Not all of the information asked for in the API specification sheet is actually required by the manufacturer in order to design the correct rupture disk. This information can be broken down into "must haves", "should haves" and "what is needed to size the disk".
• The manufacturer will always be provided with the "must haves".
• The manufacturer should also be given the "should haves" as this is a way to utilize them as a second pair of eyes and for a consistency check of the sizing.
• There are many different types of rupture disks on the market. Before selecting the correct rupture disk for your particular application, always discuss this with the manufacturer.
• Liquid service has its own set of potential problems for rupture disk design. It is highly recommended that you discuss liquid service with your manufacturer.
• There are other "non-closing" relief devices that can be considered for use. Some can only be used as secondary relief devices. However the one that can be used as a primary relief device and is gaining in popularity is the Rupture Pin.

References

1. API  (www.api.org) Recommended Practice 520, "Sizing, Selection, and Installation of Pressure-Relieving Device in Refineries, Part 1-Sizing and Selection", 7^th Edition (January 2000)
2. API (www.api.org) Recommended Practice 521, "Guide for Pressure-Relieving and Depressuring Systems", 4^th Edition (March 1997)
3. ASME (www.asme.org) "Boiler and Pressure Vessel Code, Section VIII, Division 1" (1998)
4. Continental Disc Corporation (www.contdisc.com), ASME Combination Capacity Factors, Catalogue 1-1111
5. Fike (www.fike.com), Technical Bulletin TB8103, July 1999
6. www.burstpressuresystems.com
7. www.rupturepin.com

Blanketing may prevent liquid from vaporizing into the atmosphere. It can maintain the atmosphere above a flammable or combustible liquid to reduce ignition potential. It can make up the volume caused by cooling of the tank contents, preventing vacuum and the ingress of atmospheric air.{parse block="google_articles"}

Blanketing can simply prevent oxidation or contamination of the product by reducing its exposure to atmospheric air. It can also reduce the moisture content. Gas such as nitrogen is supplied in a very pure and dry state.

The list of products blanketed is extensive ---- everything from adhesives, catalyst, chemicals, fats and oils, foods, fuels, inks, pharmaceuticals, photographic chemicals, soaps, and water.

Pressure Points

In order to effectively perform any of these functions the blanketing system must be capable of pressurizing the vapor space and accurately maintaining that pressure.

Further, conserving the amount of gas used requires that the blanketing pressure be very low. Additionally, it must be less than the tank's pressure capability. This pressure also must be below the normal tank venting pressures top prevent unnecessary actuation of these devices and the subsequent discharge of blanketing gas, as well as product loss.

These functions have been performed in various ways over the years. One method involves continuous purging, whereby blanketing gas introduced into the tank as a continuous flow exits through a vent or other opening. This method is wasteful of the blanketing gas and not always effective in maintaining an inert atmosphere.

Another approach employs a simple, direct-operated pressure-reducing valve (PRV) to blanket the tank. However, these devices are the best suited to a continuous flow rate. When used to blanket a tank they must throttle over a wide flow range. Ranging from a shut-off to a full flow condition varies the controlled pressure significantly due to droop and lockup conditions.

Pressure variations can be as much as 30 percent below or 20 percent above setpoint. The results are the poor control, wasting of blanketing gas, and possibly not maintaining the necessary atmosphere within the tank.

Another Approach

One of the more effective techniques relies on a blanketing system with the necessary controls and valving to sense and maintain the set pressure within the tank to as close as ± 0.25 inch w.c. [water column] (± 0.009 psi). Set pressures of 0.5 inch w.c. are possible and common.

Such a system will directly sense the tank pressure, control the inlet blanketing gas pressure if required and, through a main control valve, throttle blanketing gas into the tank.

The system may also provide a purge in the sensing and main supply piping, plus reverse flow protection to keep product out of the blanketing system. The installation may also require a pressure switch to monitor tank pressure and perform alarm or control functions in response to over- or under-pressure conditions.

Operation, Activation

When blanketing gas enters the typical system (figure 1), the first-stage regulator (A) reduces inlet pressure from high to low. (This is not always required.) From here, the gas flows to the main control valve (E), which is closed when the tank is at or above its setpoint.

Supply is also piped to a filter regulator ( supplying the pilot loading pressure - typically 10 psi - for actuation of the main flow control valve. This reduced pressure then floes through an orifice © and to the lower case of the main flow control valve as well as to the inlet of a sensing regulator (D).

Figure 1: Typical Tank Blanketing System Schematic

The sensing regulator includes a diaphragm, which directly senses tank pressure. Its design, weighted and inverted, allows sensitivity to as close as 0.1 inch w.c. as well as the ability to sense and control vacuums at any pressure up to the inlet pressure. The outlet of the sensing regulator connects to the tank vapor space.

In operation, if tank pressure is below setpoint, the spring will move the sensing regulator diaphragm upward, opening the seat and allowing flow from the filter regulator through the orifice and into the tank. The size of the orifice limits the volume flowrate of gas to the tank. This minimal flow may be all that's necessary to satisfy a small demand.

However, as demand increases due to pump-out or thermal cooling, the pressure drops in the pilot loading line that connects © to main valve (E). This then reduces the pressure applied to the underside of the main valve diaphragm, which allows the spring to move the diaphragm, throttling the main valve open and supplying blanketing gas to the tank.

Setpoint Response

When the tank reaches setpoint, tank pressure operating against the sensing regulator diaphragm causes the seat to close. This blocks pilot loading flow and pressure rises to the 10 psi setting. This pressure then operates against the main valve diaphragm, causing the main flow control valve seat to close. Flow to the tank ceases.{parse block="google_articles"}

At pressures below the setpoint the sensing regulator will throttle, varying pressure in the sensing system and then throttling the main valve. The amount of throttling is proportional to the offset from setpoint. Therefore, as tank pressure rises and approaches the setpoint, the main valve throttles and is closing as the tank reaches setpoint. This maintains accurate setpoints and reduces overshooting.

In addition to a blanketing system, the tank will contain the emergency vent as well as normal pressure and vacuum vents --- all of which must be properly sized for outgassing from pump-in, increases in temperature, drop in barometric pressure, or fire. The vacuum relief provides redundant tank protection for unforeseen conditions, loss of blanketing gas supply or system failure.

Figure 2: Sample Operational Stages of Blanketing Process

The selection of a blanketing pressure is very important. It must be less than the tank's pressure capability, and also should allow some dead-band below the tank vent operating pressure to avoid interaction.

Figure 2 illustrates a tank blanketed to 0.5 inch w.c. having an allowable working pressure of 6 ounces per square inch. A pressure of 0.25 inch w.c. added for shut-off results in a maximum blanketing pressure of 0.75 inch w.c. The tank has a normal pressure breathing vent that starts to open at 1.5 ounces per square inch (2.6 inches w.c.) and is full open at 1.8 ounces per square inch (3.1 inches w.c.).

Further, an emergency vent opens at 3 ounces per square inch (5.2 inches w.c.) and is fully open at 3.5 ounces per square inch (6.1 inches w.c.). There is a dead-band of 2.1 inches w.c. between the two vents. The dead-band between the blanketing system and the normal vent is 2.6 -0.75 = 1.85 inches w.c. The presence of an adequate dead-band will ensure stable operation and avoid unnecessary loss of blanketing gas or product going out an open tank vent.

Proper sizing is important for stable operation. Under-sizing can result in tank pressures dropping well below the setpoint as the main flow control valve struggles to keep up with demand. Over-sizing is less critical. However, taken to an extreme, it can result in unstable operation due to exceeding setpoint.

Summary

A gas blanketing system will effectively contain gases within the vapor space of a tank. Such a system can contain the gas, or protect the tank contents form contact with atmospheric air. It also can maintain the vapor space contents at an oxygen concentration level low enough to prevent ignition of flammable vapors.

The system are simple to use and require little, if any, maintenance. They are capable of blanketing at pressures as low as 0.5 inch w.c., minimizing consumption of blanketing gas.

Nitrogen is the most common blanketing gas, but others are used if compatible with the process. From a virtually endless list of products that qualify for blanketing, those that are volatile as well as those sensitive to contact with moisture or air are all candidates.

Crane Eq. 2-24 / Eq. (1)

Crane Eq. 2-23 / Eq. (2)

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For metering applications the ASME equation is often used to determine the Expansion Factor Y:

Eq. (3)

for flange taps, where P and T are in absolute units.

For high accuracy metering (e.g. AGA 3) use of these calculations is often limited to .

Eq. (4)

where P and T are in absolute units.

It is then often assumed that the orifice flow goes critical (i.e. sonic velocity) at P₂ = 0.528 x P₁ (for air) and the choked flow equation is then applied:

Eq. (5)

Eq. (6)

(Crane Pg 2-15), where P_cr = 0.528 x P₁

For P₂ less than P_cr this method assumes that the flow does not increase further.

Download additional information on the derivation of this method in our Article Supplements Repository

Cunningham Method

Experiments carried out by RG Cunningham and published by ASME in July 1951 clearly demonstrated that the assumption of a fixed limit to critical flow through thin square edged orifice plates is not correct. {parse block="google_articles"}The flow continued to increase as P₂ was reduced below the expected critical condition. Limiting flow was not evident even with P₂ as low as 0.1 x P₁.

Cunningham’s work included tests with air and steam with the results and conclusions presented as tables, charts and formulas. Limited information is provided for the tests with steam.

The results demonstrated that with suitable corrections to the Expansion Factor Y, the formula for non-critical flow should be used in all cases for thin square edge orifice plates. Critical flow can, however, be expected for thick orifice plates with t ? 6 x the orifice diameter.

Cunningham’s paper also includes an equation for the Flow Coefficient, though this appears to provide only a rough approximation and other methods may be preferable (e.g. AGA 3):

Eq. (7)

The ASME formula for Y was shown to be appropriate only down to P₂ = 0.63 x P_1; (not the normally expected 0.528 from thermodynamic analysis) at which point there is a distinct discontinuity in the flow to lower discharge pressures. Continued use of the ASME formula for Y produces errors of up to 12% if used for lower discharge pressures. Alternative methods reviewed involved errors of up to 40%.

Analysis of the Cunningham data suggests the following formula may be used to determine Y for flange taps at discharge pressures below 0.63 x P₁:

Eq. (8)

where Y_0.63 is Y from the ASME formula at P₂ = 0.63 x P₁.

The use of a formula similar to the form of the ASME equation is based on an expectation that there is a reasonable probability that the flow to lower pressures will be similarly sensitive to the same geometric and process parameters. The use of ? (beta) to the 4th power provides a reasonable fit to the experimental data. Since the relationship between Y and the pressure ratio is linear the (0.63-P₂/P₁) component is clearly appropriate. The inclusion of k as a direct divisor in the equation is less obvious and difficult to confirm from the limited data available.

The chart first chart below is extracted directly from the Cunningham report and clearly shows the discontinuity at a pressure ratio of 0.63 and the potential for error if the ASME formula is used beyond this point. The chart second chart below is generated using the proposed method.

Figure 1: Results with Cunningham Method	Figure 2: Results with Proposed Method

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References

1. Technical Paper 410M, Crane, 1983
2. Orifice Meters with Supercritical Compressible Flow, RG Cunningham – ASME, July 1951
3. Perry’s Chemical Engineering Handbook – 7^th Edn – RH Perry, DW Green
4. Flow Measurement Engineering Handbook, 3^rd Edn – RW Miller
   (Note: formula 9.58 is correct only for air with ß = 0.15 and could be replaced with the method proposed above.)

Most commonly, pyrophoric iron fires occur during shutdowns when equipment and piping are opened for inspection or maintenance. Instances of fires in crude columns during turnarounds, explosions in sulfur, crude or asphalt storage tanks, overpressures in vessels, etc., due to pyrophoric iron ignition are not uncommon.{parse block="google_articles"}

Often the cause of such accidents is a lack of understanding of the phenomenon of pyrophoric iron fires. This article aims to explain the basics of pyrophoric iron fires and to provide ideas for developing safe practices for handing over equipment for inspection and maintenance.

What is Pyrophoric Iron Oxidation?

The word "pyrophoric" is derived from the Greek for "fire-bearing". According to Webster's dictionary, "pyrophoric material" means "any material igniting spontaneously or burning spontaneously in air when rubbed, scratched, or struck, e.g. finely divided metals".

Iron sulfide is one such pyrophoric material that oxidizes exothermically when exposed to air. It is frequently found in solid iron sulfide

scales in refinery units. It makes no difference whether these pyrophoric sulfides exist as pyrite, troilite, marcasite, or pyrrhotite. It is formed by the conversion of iron oxide (rust) into iron sulfide in an oxygen-free atmosphere where hydrogen sulfide gas is present (or where the concentration of hydrogen sulfide (H₂S) exceeds that of oxygen). The individual crystals of pyrophoric iron sulfides are extremely finely divided, the result of which is that they have an enormous surface area-to-volume ratio.

When the iron sulfide crystal is subsequently exposed to air, it is oxidized back to iron oxide and either free sulfur or sulfur dioxide gas is formed. This reaction between iron sulfide and oxygen is accompanied by the generation of a considerable amount of heat. In fact, so much heat is released that individual particles of iron sulfide become incandescent. This rapid exothermic oxidation with incandescence is known as pyrophoric oxidation and it can ignite nearby flammable hydrocarbon-air mixtures.

Basic chemical reactions: Iron sulfide is one of the most common substances found in refinery distillation columns, pressure vessels, etc. It is formed by the reaction of rust or corrosion deposits with hydrogen sulfide as shown below:

Eq. (1)

There is a greater likelihood of this reaction occurring when the process involves a feedstock with high sulfur content. This pyrophoric iron sulfide (PIS) lays dormant in the equipment until the equipment is shutdown and opened for service, exposing the PIS to air, allowing the exothermic process of rapid oxidation of the sulfides to oxides to occur, as shown in the equations below:

Eq. (2) Eq. (3)

The heat usually dissipates quickly unless there is an additional source of combustible material to sustain combustion. The white smoke of SO₂ gas, commonly associated with pyrophoric fires, is often mistaken for steam.

Pyrophoric Iron Oxidation in Distillation Columns

In petroleum refineries, the equipment most prone to pyrophoric combustion induced fires is the distillation columns in crude and vacuum distillation units. Deposits of iron sulfide are formed from corrosion products that most readily accumulate at the trays, pump around zones, and structured packing. If these pyrophoric iron sulfide (PIS) deposits are not removed properly before the columns are opened up, there is a greater likelihood of PIS spontaneous ignition. The trapped combustible hydrocarbons, coke, etc. that do not get adequately removed during steaming/washing often get ignited, leading to fires and explosions inside the equipment. These fires not only result in equipment damage but can also prove fatal for the personnel who are performing inspection and maintenance work inside the columns.{parse block="google_articles"}

The accidents due to pyrophoric iron oxidations are entirely avoidable if safe procedures for column handover are followed. The targets of these procedures should be twofold:

1. First, to remove all the combustibles
2. Second, to remove or neutralize pyrophoric iron sulfide deposits

The basic distillation column oil-cleanup procedure is discussed in steps below.

Distillation Column Oil Cleanup Procedure

1. Steaming: The steaming is done after all liquid hydrocarbons have been drained from the column and associated piping. The objective of steaming is to make the column and associated piping free of residual hydrocarbons. The column vent and pump strainers in the side draw piping are de-blinded and steaming is started from utility connections at the bottom of the column. Generally, steaming is continued for about 20 to 24 hrs, ensuring the column top temperature remains more than 100 ^°C throughout the operation.

2. Hot Water Washing: When clear steam is observed exiting the column vents, water washing of the column should be started. With steam still in commission, water is sent to the column, usually via reflux lines, and it is drained from the column bottom, associated pump strainers, etc. The water flow rate should be adjusted so that steam still comes from the vent (i.e. water should not result in condensing of all steam before it reaches the column top). Water flow should be stopped for 2-3 hrs and then resumed. This cycle of steaming and washing should be repeated several times for a total of about 15 to 20 hours. Injection of a turpene-based detergent into the steam can also be considered. The condensate-soap solution can be collected and circulated through the various side cuts.

3. Blinding: When clear water is observed at side draw pump strainers, etc., associated piping should be isolated by installing blinds wherever isolation is possible.

4. Cold Water Washing: The hot water wash should be followed by a cold water wash (i.e. steam should be fully closed). The cold water washing is done for about 20-24 hrs.

5. Chemical Injection for Removal and Neutralization of PIS Deposits: During the cold-water wash or after washing is over, chemical injection for removal of pyrophoric sulfides should be considered. The various options for chemical treatment are discussed below:

• Acid cleaning - This procedure involves pumping in an acid with some corrosion inhibitor. The acid dissolves sulfide scale and releases hydrogen sulfide gas. It is effective and inexpensive, however, disposal of hydrogen sulfide gas can be a problem, as can corrosion (when the system contains more than one alloy). Dilute hydrochloric acid solutions may also be used. The resulting iron chloride turns bright yellow, acting as an indicator for removal of the iron sulfide.

• Acid plus hydrogen sulfide suppressant - Additional chemicals can be added to the acid solution to convert or scrub the hydrogen sulfide gas.

• Chelating solutions - Specially formulated, high pH, chelating solutions are quite effective in dissolving the sulfide deposits without emitting hydrogen sulfide, but this is an expensive application.
• Oxidizing chemicals - Oxidizing chemicals convert sulfide to oxide. Potassium permanganate (KMnO₄) has been used commonly in the past to oxidize pyrophoric sulfide. Generally the potassium permanganate is added to the tower during the cold water washing as a 1% solution. At various intervals, samples are taken and checked for color. The colors of the fresh KMnO₄ and the spent MnO₂ are purple and brown respectively. If the color of the solution becomes brown, additional KMnO₄ is needed. The reaction is judged complete when the solution color remains purple. It takes approximately 12 hours to complete the job.

  The cost of potassium permanganate treatment is more expensive than acid cleaning and traditional oxidizing agents such as sodium hypochlorite or hydrogen peroxide. Nevertheless, it is less corrosive to equipment than acid cleaning and used properly can be safer than other oxidizing agents.

  The following conditions should be avoided when using potassium permanganate:

  · Do not add KMnO₄ to acids or use in a low pH environment

  · Combustible materials should not be allowed to contaminate KMnO₄ stocks

  · Residual MnO₂ may remain in vessels after treatment and cause combustion or flammability issues in equipment with large surface areas such as packed towers

  · KMnO₄ cannot be used in conjunction with most detergents

  · KMnO₄ may have a “bad reputation” in some processing plants, but this is often times the result of misuse by contractors or plant personnel.

Alternative oxidation technologies are being developed with a focus on

• increasing safety in application
• saving water
• eliminating odor problems
• minimizing wastewater problems
• reducing wastes

One such alternative is Zyme-Flow®. Zyme-Flow® offers unique chemistry which is patented and offered by license from United Laboratories International, LLC as Zyme-Flow® and related products. The Zyme-Flow® chemical applications are administered by a highly trained staff of technicians provided by United and sold only by license from United Laboratories International, LLC.

The Zyme-Flow® generic Vapour Phase® method is apparently unique in that the de-oiling and oxidizer composition that is being dispersed actually may be vaporized in the steam (instead of being just atomized).  This allows the Zyme-Flow® composition to travel (in easily measurable concentrations) extensive distances and throughout equipment with high efficiency for contacting and condensing on internal surfaces.  The composition may expend quickly, but the application technicians can measure its progress.  This prevents over-dosing.

A very generic Vapour Phase® procedure may include:

1. Stop feed and de-inventory the unit per normal procedures.
2. Perform initial isolation per the pre-established plan with the Zyme-Flow® specialists.
3. Establish a thorough steam path throughout the equipment.
4. Add Zyme-Flow® to the incoming steam (commonly only 1-3 points are needed even for very large units).
5. Continue the injection and steaming for 8-12 hours.
6. Perform isolation per pre-established plan with the Zyme-Flow® specialists.
7. Possibly perform a final rinse with cold water to cool the columns quickly.
8. The unit is then ready for opening, ventilation, and hot work.

One major advantage for oxidizing pyrophoric iron sulfide is that the distribution dynamics of the Vapour Phase® applications are often more efficient than atomized distribution methods.  This is why Zyme-Flow® is often used for decontamination of flare lines and overhead systems where few injection points can be utilized.   The same dynamics allows for full treatment of most tight packing structures in refinery columns.

In some situations, Zyme-Flow® applications may be combined simultaneously or in sequence with compatible solvent and oxidizer products to target specific challenges such as monomer / polymer coatings and other challenges.    These applications require specialty design consideration from the Zyme-Flow® specialists.  This is especially important in structured packing situations where polymerization tends to coat and protect the pyrophoric deposits from contact with oxidizers until cooling promotes cracking of the polymer.

Solvent/Surfactant Steam Dispersion Methods

There are alternatives to the steam, wash, blind, and wash again technologies.  These include steam dispersion technologies which are sometimes combined with oxidizer washing technologies.   These alternatives may include steam dispersion of organic solvent products and can be very good to excellent de-oiling and degassing compositions (which expose pyrophoric iron sulfide to subsequent oxidizer treatments).{parse block="google_articles"}

For critical path process units in a turnaround, a very generic procedure may include:

1. Stop feed and de-inventory the process equipment per normal procedures.  Sometimes this involves steaming and sometimes water displacement.
2. Establish a steam flow throughout the equipment.  The design of the flow path is critical.
3. Add chemicals to the incoming steam to promote de-oiling and degassing of the equipment (this may involve numerous injection points).
4. When the outflow vapors are within safety and environmental specifications, blow down the equipment to the atmosphere for a short time (or continue to flare or condenser as needed).
5. Perform isolation as needed prior to final washing for oxidation (per pre-designed isolation plan).
6. Perform a thorough water wash with oxidizer until the oxidation requirement of the fluid path is complete.  Sometimes this must involve total fluid fill of the equipment to obtain positive contact of pyrophoric surfaces.
7. The unit is then ready for opening, ventilation, and hot work (unless other chemical treatments are required).

This generic procedure may allow for 24-48 hours of savings over the extended steaming, isolation, and washing approaches, and may be safer to perform.  There are several products currently offering this type of service.  Many of them are strong encapsulators and require secondary treatment to break the emulsions.

The oxidizers Zyme-Flow and Zyme-Ox are proprietary products from United Laboratories International, LLC for most refinery and petrochemical decontamination applications.  For more information on Zyme-Flow^SM Process Technology, the readers can visit www.zymeflow.com or contact.

Case Studies: Pyrophoric Iron Fires

The history of refining is replete with cases of fires and explosions due to pyrophoric iron ignitions. A few of these cases are discussed below (details like the location and date of the incidents are not included), to give the reader an idea of the nature of pyrophoric iron fires and the lessons learned.

 

 

 

 

General Precautions to Avoid Pyrophoric Iron Fires

1. The scraps and debris collected from cleaning of filters in naphtha / crude service must be kept wet and disposed of underground.{parse block="google_articles"}
2. Tanks, reactors, columns, and exchangers in high-sulfur feed service must be kept properly blanketed with N₂ during idle periods.
3. All equipment and structured packing must be properly water washed and kept wet when exposed to the atmosphere.
4. In processes where catalyst handling is required (such as in Hydrotreating and fluid catalytic cracking) caution must be taken during catalyst recharge or disposal. When unloading any spent coked catalyst, the possibility exists for iron sulfide fires. If the spent catalyst is warm and contacts oxygen, iron sulfide will ignite spontaneously and the ensuing reaction may generate enough heat to ignite carbon deposited on the catalyst. Therefore catalyst must be stripped of all hydrocarbons, cooled to about 50 ^o C and wetted with water to prevent it from igniting vapors. Once cooled, the used catalyst may be emptied into drums for later shipment to a regenerator or a disposal site. As the catalyst may be highly pyrophoric (containing iron sulfide, etc.), it should be dumped into drums containing an internal liner for shipment. The drum and liner should first be filled with inert gas, which is then displaced by the catalyst. The liner should be tied off and a small chunk of dry ice placed inside the drum before sealing. These precautions should protect against catalyst auto ignition.

References

1. "Pyrophoric Materials Handbook, Flammable Metals and Materials", By Charles R. Schmitt, P.E., C.H.C.M., Edited By Jeff Schmitt
2. "Pyrophoric Fires and Column Shutdown", Refineries Quarterly Safety Bulletin, April-June 1997.
3. “Oxidizer Use in Refinery Chemical Cleaning: Selection Considerations and Case Histories”, presented at NACE, Chris Spurrell (Chevron) and Ron Kenyon (Delta Tech Services)
4. "Methods for Removal of Iron sulphide", Mr. Fu Huang, Chinese American Association of Corrosion & Materials Engineers
5. "Formation and Oxidation of Sulfides on Pure Iron and Iron Oxides", Masatoshi Watanabe1, Minoru Sakuma, Takeshi Inaba, and Yasutaka Iguchi, Department of Metallurgy, Graduate School of Engineering, Tohoku University, Sendai 980-8579, Japan
6. “Basic Technology of Zyme-Flow Process”, Bevan Collins, International Technical Director, United Laboratories, LLC , www.zymeflow.com
7. TECHNICAL BULLETIN: Safe Handling of CRITERION Hydrotreating Catalysts
8. www.worlfuels.com, NPRA Q&A Minutes, 1999 Session I